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 Coal gas residuals 
 
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COAL GAS RESIDUALS 
 
minium uiiiiiiiuiiiiiiiiiiiiiiiiiiiiiiiiiiiiiiiniiiiiiiiimiiiiiiiiiiiiiiiiiiiiiiiiiiiimiiiiiiiiiiiiii iij 
 
 KflillBodiQxh 
 
 raw-Ma Doon uu im 
 
 PUBLISHERS OF BOOKS FOIU 
 
 Coal Age ▼ Electric Railway Journal 
 Electrical World v Engineering News-Record 
 Railway Age Cazette " American Machinist 
 Electrical Merchandising v The Contractor 
 Engineering 8 Mining Journal <*■ Po we r 
 Metallurgical 6 Chemical Engineering 
 
COAL GAS 
 RESIDUALS 
 
 BY 
 
 FREDERICK H. WAGNER, M.E. 
 
 MEM. AM. SOC. M.E., AM. GAS INSTITUTE, FRANKLIN INSTITUTE, 
 
 AM. SOC. NAVAL ENGINEERS. 
 
 AUTHOR OP " BLAST-FURNACE GASES," 
 
 " COAL AND COKE," ETC. 
 
 Second Edition 
 Revised and Enlarged 
 
 McGRAW-HILL BOOK COMPANY, Inc. 
 
 239 WEST 39TH STREET, NEW YORK 
 
 LONDON: HILL PUBLISHING CO., Ltd. 
 
 6 4 8 BOUVERIE ST., E. C. 
 1918 
 
4io\ 
 
 COPYRIGHT, 1914, 1918, BY THE 
 
 McGraw-Hill Book Company, Inc. 
 
DEDICATED TO THE MEMORY OF 
 
 WALTHER FELD 
 
 of whom nothing of greater praise can be said than 
 by repeating the following lines which appeared in 
 "Die Chemische Industrie,'' published in Germany, 
 namely: 
 
 " The barium industry as well as industries connected 
 with tar and ammonia must thank him for the new 
 paths which he has broken, and in the history of the 
 development of German chemical industries the life 
 work of Walther Feld must for all time remain an 
 imperishable page of fame." 
 
PREFACE TO SECOND EDITION 
 
 Conditions arising during this period when there is a general 
 patriotic endeavor to supply the War Department with the many 
 products needed to carry the war to a successful issue, have 
 necessitated the addition of certain special features to this book : 
 hence the new edition. 
 
 The important additions are the treatment of the process of 
 distillation and some general information on tar products to the 
 chapter on tar; additional information on the product derived 
 from spent oxide in the chapter on cyanogen; data on the pro- 
 duction of nitric acid to the chapter on ammonia; additional data 
 to the chapter on naphthalene; some new material on the produc- 
 tion of benzol and toluol; and a chapter on the manufacture of 
 sulphuric acid when spent oxide is used as a base. 
 
 The test for cyanogen has been revised so that the chemist 
 engaged in this work finds it simple and practical. 
 
 I wish to thank the friends who have called attention to typo- 
 graphical errors in the previous edition. These have been cor- 
 rected. 
 
 I trust that this revised edition may find as much favor among 
 operators as did the former one. 
 
 FREDERICK H. WAGNER 
 
 Baltimore, Md., November, 1917 
 
PREFACE TO FIRST EDITION 
 
 The lack of general information in concise form covering the 
 various methods of securing the residuals pertaining to the car- 
 bonization of coal, and the necessity of calling attention to the 
 waste of our natural resources in so many instances, was the 
 incentive leading to the writing of the following pages. All of the 
 residuals except coke are treated, the latter product being reserved 
 for a separate volume, and in presenting the Feld theory and 
 practice to the coal gas producer, as well as to the coke oven 
 operator, I am taking the liberty of quoting the two appended 
 letters, first by Camillo J. Guttmann, of London, which appeared 
 in the Gas World on July 18, 1914, and the second by W. R. 
 Ormandy from the Gas World of August 8, 1914. 
 
 The first letter says: 
 
 "The advantages of oil over coal as a source of power for many pur- 
 poses, but particularly for naval purposes, have lately beeen so vigorously 
 canvassed and so exhaustively discussed by engineering experts, that 
 there is now no purpose in dilating upon them. In view of the fact that, 
 whether for good or ill, the die has now been cast in favor of oil, and 
 that the nation (English nation) has already committed itself to far-reach- 
 ing schemes of foreign exploitation, it becomes much more pertinent at 
 the moment to consider what resources we have for securing our supplies 
 at home; for besides the more obvious advantages, there are others 
 hardly less important in finding our sources of supply in this country. 
 
 In articles which have recently appeared in the Press, public attention 
 was directed to this aspect of the question, and to the possibility of pro- 
 ducing oils for the Navy from bituminous coal, shale, peat, or even sewage 
 matter. These articles mainly drew attention to the del Monte and 
 similar processes of obtaining motor spirit and fuel oil from coal. While 
 these processes will certainly promote the solution of the oil fuel problem, 
 they are somewhat dependent for their ultimate fruition on the success- 
 ful creation of a market for their residual product — a semi-fuel for house- 
 hold and like purposes. 
 
 These interesting articles, however, omitted to mention one process 
 which, to my mind, whether alone or in combination with other methods, 
 
 viii 
 
PREFACE ix 
 
 opens up a wide vista of possibilities in this connection. I refer, as most 
 experts will have guessed, to the process perfected by the late Walther 
 Feld. This process without the intervention of any other, produces 
 not only fuel oils, but, at the same time, and in one operation, a pitch 
 containing practically no free carbon (in other words, precisely the pitch 
 that is now in ever growing demand for road-making and other purposes), 
 a benzol which is valuable as a substitute for petrol, and, finally, the usual 
 yield of ammonium sulphate. This last is produced direct by a beauti- 
 fully simple process which entirely obviates the necessity of purchasing 
 or manufacturing sulphuric acid. 
 
 The general introduction of this process in coke-oven plants and in 
 gas-works throughout the country would, apart from the economy in 
 working and prime costs, greatly assist in providing a natural, national 
 and adequate reserve of oil fuel, and is, therefore, worthy of earnest con- 
 sideration jointly with other processes direct towards this end. 
 
 It should be noted that the success of the Feld process is not in any 
 way dependent on the sale of a residual product, constituting, as it does, 
 an improvement on existing processes in by-product recovery. 
 
 Our agricultural interests would derive all the advantage of a more 
 rational production of ammonium sulphate. The fertility of land may, 
 in many cases, be trebled by the judicious use of this, one of the most 
 generally valuable of our artificial manures. Ammonium sulphate is 
 now recognized to be so necessary that its use has almost become a cri- 
 terion of the progress of agriculture, and of the degree of a country's 
 enlightenment. 
 
 The writer recently had the privilege of inspecting a trial Feld plant 
 installed at the works of one of the foremost Westphalian coal and iron 
 firms, and intended to deal with the gases from a battery of coke-ovens 
 with a daily capacity of 800, and ultimately 1600 tons of coal. This 
 plant is, I understand, the first of several which are on order, and in the 
 writer's opinion, represents the first step by which Germany will ulti- 
 mately become largely independent of foreign oil for her Navy, and of 
 petrol for her motor-propelled vehicles. 
 
 The above considerations are not intended to put forward the Feld 
 process as a panacea for the fuel problem, but to show that the available 
 field is a broad one, and that the Feld process must necessarily play an 
 important part in any future developments. 
 
 It will be seen that the present needs of our Admiralty furnish the 
 Government with a rare opportunity for combining the solution of a 
 departmental economy with a new industrial departure influencing our 
 national economy, and which will not only increase our power for offensive 
 defence, but — and this is interesting — enhance our power for passive 
 resistance just where it is weakest, by reinforcing our home supplies of 
 food. 
 
X PREFACE 
 
 It is much to be hoped that our rulers will show themselves sufficiently 
 enlightened to embrace this opportunity and, what is quite as essential, 
 embrace it promptly and vigorously." 
 
 The second communication, that from Mr. W. R. Ormandy, 
 
 says: 
 
 "From the correspondence which has been taking place in The Gas 
 World it is quite obvious that the problem which Mr. Feld set out to 
 solve is not realized by all readers. No doubt from time to time descrip- 
 tions of Feld plant and processes have appeared in these columns; but 
 in view of the increased attention which is now being devoted towards 
 the better utilization of the national fuel resources it might not be 
 out of place to describe briefly the general principles which Mr. Feld 
 enunciated. 
 
 Regarded from a scientific standpoint, even the most modern methods 
 of utilizing fuel are carried out in a primitive and, in most cases, bar- 
 barous manner. In coke-oven and gas-works practice coal is distilled and 
 the heterogeneous products of such distillaton are condensed together 
 into an abominable mixture upon which subsequently much time, skill, 
 fuel and capital has to be spent to bring about even a partial separation 
 of the tarry mass into its separate constituents. Maybe Mr Feld was 
 not the first to realize the waste that was involved in such a process 
 and to point out how much more rational and economical it would 
 be to separate the various constituents of the distillate by means 
 of an adequate system of fractional cooling and washing. On the other 
 hand, he was undoubtedly the first to make this problem a life-work, 
 and to set out to devise apparatus and to work out reactions to overcome 
 the difficulties with which every step in this great problem was beset. 
 A comparatively superficial examination of the problems involved showed 
 the impossibility of dealing with the complicated products of the dis- 
 tillation of coal in an adequate manner. Making use of liquid reagents 
 and the utilization of insoluble reagents in suspension brought with it the 
 necessity of devising some new form of washing apparatus wherein such 
 solid suspensions could be used continuously and' efficiently. This led 
 to the evolution of the Feld washer, which alone rendered possible the 
 utilization of many of the reactions which had subsequently to be made 
 use of. In spite of the many difficulties under which he labored, Mr. Feld 
 succeeded in perfecting and proving the suitability of the intensive type 
 of centrifugal washer, and lived long enough to find the principles in- 
 volved therein adapted by a large number of makers who flooded the 
 markets with modifications as close to the original as the law permitted. 
 
 The next stage of development necessitated an intimate knowledge of 
 the vapor tension and vapor pressure of solutions of gases and of oils. 
 Much of this was non-existent, and it necessitated a vast amount of work 
 
PREFACE 3d 
 
 before a basis was available upon which to build the superstructure of 
 the new process. Instead of cooling the gases from whatever form of 
 distilling apparatus was being employed, these were kept at such a tem- 
 perature that no deposition of even the most easily condensable of them 
 took place before the products were allowed to enter the first washer. 
 So far as regards dealing with the oils, the principle involved was similar 
 throughout, namely, to reduce the temperature by stages in a number of 
 washers, so that the gas was cooled in any one stage to such a temperature 
 as allowed of condensation of whatever product was desired in that stage. 
 Thus, in the plant visited by the writer, in the neighborhood of Teplitz, 
 in Austria, where the products of distillation of a large number of by- 
 product coke ovens distilling brown coal were being treated, the first 
 washer was kept at such a temperature as permitted only of the con- 
 densation therein of pitches and tars and oils boiling above the tem- 
 perature of 240 degrees C. The fluid tar running away from the bottom 
 of the first cooling washer was allowed to cool and re-circulated, by means 
 of a pump, through this washer at such a rate as enabled the necessary 
 reduction in temperature to take place. In each of the subsequent coolers 
 a similar chain of operations took place. It is obvious that it is possible 
 by increasing or diminishing the number of coolers and by varying the 
 rate of cooling, to obtain oils from any one washer boiling between any 
 desired range of temperatures. In this wise, at Teplitz tar, heavy oil, 
 and light oil were separated free from water and ammonia. By this time 
 the temperature of the gases has been brought so low that the water 
 vapor contained therein is on the verge of condensing. By using the 
 counter-current system and allowing the gases to meet a stream of water 
 in the Feld washer it is possible to reduce the temperature of the outgoing 
 gas to about 60 degrees C, and yet to have water leaving the washer at 
 such a temperature that no appreciable amount of ammonia remains 
 therein. By similar ingenious means the temperature of the gases was 
 lowered still further without loss of ammonia until a product was left 
 containing the whole of the ammonia and so little water that it was pos- 
 sible to recover the ammonia as sulphate free from tar and coloring 
 matter by means of some of the many reactions which Feld worked out 
 for this purpose. 
 
 It would lead us too far at this point to go into the details regarding 
 the recovery of cyanogen, but it is worth pointing out that the Feld sys- 
 tem of fractional cooling permits of the removal of cyanogen at a tem- 
 perature and under conditions which prevent the formation of the sulphur 
 compounds of cyanogen, which lead to such heavy losses of this valuable 
 material, in operations carried out under the old regime. The cooling of 
 the gases to a degree which permits of the removal of the ammonia salts 
 still permits of the benzene compounds remaining in the form of vapor, 
 and subsequent washing in a final Feld washer with heavy oil results in 
 
xu PREFACE 
 
 the recovery of practically the whole of the benzene group boiling below 
 140 degrees C. in a form which permits of their recovery in great purity 
 with the minimum of labor and expense. 
 
 The subject of the recovery of the ammonia as sulphate at the expense 
 of the sulphur contained in the gas with the ammonia has been in the 
 public eye to such an extent during the last few years that the work of 
 Feld in this direction has no doubt caused the general public to overlook 
 the service which he has rendered by not only pointing out better theo- 
 retical lines upon which to attack the great problem of dealing with coal 
 tar distillates, but in proving that such theoretical methods were well 
 within the practical province. 
 
 There is no doubt that the immediate futiu-e will witness great ad- 
 vances in the more intelligent utilization of our fuel resources and such 
 improvement is bound to be accompanied by the production of increased 
 amounts of both volatile and heavy oils. It is almost impossible to con- 
 ceive that such advances can take place without benefiting in a large 
 degree from the pioneer work carried out by Mr. Feld, and thus far Mr. 
 Guttmann is more than justified in associating Mr. Feld's name with fuel 
 oil from national sources." 
 
 The author would greatly appreciate any suggestions from the 
 readers of this book with the view of extending its usefulness in 
 ensuing editions, and in thus keeping its contents abreast of any 
 developments made in this field of conservation and industry. 
 
 FREDERICK H. WAGNER 
 
 Baltimore, Md., August, 1914. 
 
CONTENTS 
 
 PAGE 
 
 Preface vii 
 
 Introduction 1 
 
 CHAPTER I 
 Tar 7 
 
 CHAPTER II 
 Naphthalene 52 
 
 CHAPTER III 
 Cyanogen 66 
 
 CHAPTER IV 
 Ammonia 89 
 
 CHAPTER V 
 Benzol 164 
 
 CHAPTER VI 
 Sulphuric Acid 191 
 
 CHAPTER VII 
 Tests 200 
 
 Appendix 223 
 
 Index 237 
 
COAL GAS RESIDUALS 
 
 INTRODUCTORY 
 
 A hundred years of development have passed since Murdoch's 
 discoveries permitted the lighting of the city of London with 
 coal gas, and these years have placed gas lighting in the fore- 
 front of industrial progress, while the history of these years 
 plainly shows how the proper utilization of by-products can revo- 
 lutionize the magnitude of an industry. The early years of coal 
 carbonization could depend upon no, or but very little, assist- 
 ance from by-product sales in the matter of cheapening the cost 
 of production, and not until the aid of modern chemistry was 
 called in, did the production of residuals make any appreci- 
 able change in this condition. 
 
 The recovery of residuals is a very important conservation 
 of resources, and it forms one of the principal means of 
 revenue to the coal-gas producer, the sale of these residuals 
 reducing the cost of gas production in a degree corresponding 
 to the efficiency of the recovery methods adopted and the 
 market value of the product. 
 
 The principal residuals recovered today are tar, naph- 
 thalene, cyanogen, ammonia, and in the case of coke-oven 
 gas, also benzol by a direct recovery method, and they will 
 be treated in this consecutive order in the following pages. 
 The recovery of benzol is confined almost entirely to the coke- 
 oven plant, where a direct method of recovery is adopted, 
 and to the tar distiller, benzol being one of the most valuable 
 of tar constituents. While the recovery of naphthalene can- 
 not exactly be termed one of profit in a pecuniary sense at 
 present, the removal of a certain portion from the gas is of 
 distinct advantage, and the methods adopted for its removal 
 will, therefore, be given. 
 
 The gas produced during the carbonization of coal is a 
 mixture of fixed gases, vapors of various kinds, and, at times, 
 
2 COAL GAS RESIDUALS 
 
 also globules of liquids, which are held in suspension, and 
 are thus carried forward by the gas; these gases and vapors 
 also carry forward some solid carbon in the shape of dust. 
 
 The principal fixed gases are hydrogen, H 2 ; methane, 
 CH 4 , also known as "marsh" gas; ethane, C 2 H 6 ; propane, 
 C 3 H 8 ; butane, C 4 Hi ; ethylene, C 2 H 4 ; small amounts of 
 butylene, C 4 H 8 ; propylene, C 3 H 6 ; acetylene, C 2 H 2 ; carbon 
 dioxide, C0 2 ; carbon monoxide, CO; hydrogen sulphide, 
 H 2 S; nitrogen, N 2 ; oxygen, 2 ; and ammonia, NH 3 ; while 
 the principal vapors in the mixture are benzol, C6H 6 ; toluol, 
 C 6 H 5 CH 3 ; xylol, C 6 H 4 (CH 3 ) 2 ; carbon disulphide, CS 2 ; and 
 aqueous vapors. These latter vapors are those of substances 
 which become liquid at ordinary temperatures, but the vapors 
 of naphthalene, CioH 8 , phenols, etc., are those of substances 
 which become solid at ordinary temperatures, and must, 
 therefore, be subjected to a special treatment. 
 
 As will be seen later, some of these constituents are of 
 inestimable value to the coal-gas producer, and consequently 
 the treatment of the gas after it is produced in the carboniz- 
 ing plant, is of great importance; this treatment should not 
 only consist of a method of cooling the gas, and thus condens- 
 ing and precipitating the vapors as a fluid, but the method of 
 treatment should be such as to retain in the gas those valuable 
 illuminating constituents which may be lost to a greater or 
 lesser extent in the usual condensing plant, and thus the Feld 
 system, which embraces successive cooling with a fractionation 
 of the products, appears most attractive from the standpoint 
 of efficiency and simplicity. 
 
 The usual condensing system in coal-gas practice embraces 
 the use of a primary condenser, exhauster, tar extractor, 
 secondary condenser, and ammonia washer, the tar, together 
 with quite an amount of illuminants, being thus removed in 
 a great measure by cooling, while in the Feld system the gas 
 is not cooled below the point where any volatile hydro-carbons 
 are precipitated or absorbed by the effluent, the tar being 
 fractionated in three washers into pitch, heavy oils, and 
 middle or light oils, and the gas is treated for cyanogen, 
 combined hydrogen sulphide and ammonia, naphthalene, and 
 finally benzol, the entire process being carried out in Feld 
 vertical centrifugal washers, as shown in diagram A. 
 
INTRODUCTORY 
 
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4 COAL GAS RESIDUALS 
 
 As will be seen from Diagram A, the entire condensing 
 plant consists of eleven washers, the first, or pitch washer 
 operating at a temperature of from 160° to 200° C, the gas 
 being kept at this temperature in its progress from the 
 hydraulic main by insulating both the main and the pipe 
 connections, or by the application of external heat, heavy 
 oils from washer (2) being pumped into washer (1), where 
 they are used as the active pitch-extracting medium. 
 
 Washer (2) removes the heavy oils, and it is operated at 
 a temperature of from 160° to 80° C, all operating tem- 
 peratures being determined from the actual dew point of 
 the gas for the constituent to be removed, the extracting 
 medium being the middle oils from washer (3). 
 
 Washer (3) removes the middle oils, due to contact be- 
 tween the gas and the oils separated by cooling in washer 
 (4), this cooling being effected by bringing the gas into 
 intimate contact with water, the effluent from washer (4) 
 being run into a separating tank where the oil is separated 
 from the water by specific gravity. 
 
 The light oils are removed in washer (5), where the gases 
 are washed by means of heavy oils which, are previously 
 cooled in a special cooler, these light oils being run into a 
 reservoir from whence they are pumped into the first benzol 
 washer. 
 
 The washers for ammonia and hydrogen sulphide operate 
 under a temperature of from 40° to 36° C, the hydrogen 
 sulphide being combined with the ammonia in these washers 
 with the consequent formation of ammonium sulphate, while 
 the cyanogen washer operates at about 36° to 34° C. 
 
 The naphthalene is removed in washer (9) at a tempera- 
 ture of about 18° C, while washers (10) and (11) serve to 
 extract the benzol and its homologues, the partially satu- 
 rated oil coming from washer (10) being pumped into washer 
 (ID- 
 
 For the complete theory of condensation the reader is 
 referred to "The Cleaning of Blast-Furnace Gases," 1 only 
 such portions of this theory as are applicable to the removal 
 of tars and oils being given in the following chapters. 
 
 The usual method of condensation as practised in America 
 1 Published by the McGraw-Hill Book Co., Inc., New York. 
 
INTRODUCTORY 5 
 
 has as its natural result the cooling of the gas and, in so 
 doing, the removal of those constituents which cannot re- 
 main permanently in the gas without fouling the purifying 
 apparatus and the distributing system. This condensation 
 starts as soon as the gas leaves the mouthpiece and enters 
 the ascension or stand pipe delivering the gas to the hy- 
 draulic main, and it has been estimated that from 30 to 50 
 per cent of all condensible constituents are deposited in 
 this main. 
 
 Heavy tars are thus deposited in the hydraulic main, and 
 these should be removed as quickly as possible because their 
 presence acts as an absorbent upon the illuminants, and thus 
 reduces the candle-power of the gas. All subsequent con- 
 densation should proceed slowly in order to retain as many 
 of the hydro-carbons in the gas as is possible, because sudden 
 chilling not only removes these valuable constituents, but it 
 also causes the deposition of some of them in the shape of 
 naphthalene, with resultant stoppages. 
 
 The tar produced from coal gas is one of its most impor- 
 tant residuals, and while it is used to some extent in its 
 original form, it finds its most important application in the 
 arts, where it forms the basis of aniline color production, 
 900 different colors and shades being listed as coming from 
 tar, these being produced by distilling the tar in tar stills 
 and, after a certain temperature has been reached, water-like 
 oils result, these oils being the base of all of the beautiful 
 coal-tar colors. Carbolic acid, naphthalene, anthracene, and 
 benzol are also produced in like manner, and each of these 
 constituents in turn produce a long series of other products. 
 By treating anthracene we produce alizarin, one of the most 
 valuable coloring matters, and which forms the base of 
 indigo. Many blue dyes in early days were produced from 
 the juice of the indigo plant, native to India, which juice, 
 when exposed to the atmosphere, oxidized and precipitated 
 indigo. 
 
 Indian indigo is still used to a great extent in the dyeing 
 industry, but the blues produced from coal-tar are rapidly 
 displacing it. The alizarin produced from'coal-tar is identical 
 in chemical composition and coloring properties with madder, 
 and this has become a very valuable commercial product. 
 
6 COAL GAS RESIDUALS 
 
 Substances termed mordants by dyers, owing to their fixing 
 qualities, are mixed with the true coloring matter, and are thus 
 employed as vehicles for the latter. By adding the mordant 
 alumina to alizarin, red dyes result, and by adding the mor- 
 dant iron, darker tones are produced, and by adding chrome 
 the resultant colors are violet. 
 
 Chemists are able to split up quinine and produce quinoline, 
 the latter being also a coal-tar derivative, but up to the present 
 time, in spite of the fact that quinoline is known as the skeleton 
 of quinine, we have as yet been unable to produce the latter from 
 quinoline. 
 
CHAPTER I 
 TAR 
 
 Tar is a thick, dark-brown, viscid, oily liquid, produced 
 during the dry distillation of coal; its chemical nature is 
 very complex, and it contains a large number of compounds. 
 The crude gas leaving the retorts is a mixture of permanent 
 gases, but this mixture carries with it quite a number of 
 hydro-carbon and other vapors, and as the illuminating, as 
 well as the calorific, quality of coal gas is, in a great measure, 
 dependent upon its hydro-carbon constituents, it is of prime 
 importance to so treat the gas as to retain as many of these 
 hydro-carbons as is possible. A reduction in temperature, 
 however, soon reduces the hydro-carbons of greater density 
 to liquid form, and this liquid is usually termed "tar." In 
 spite of the fact that the gas temperature at the hydraulic 
 main outlet is perhaps never lower than 145° F., large quan- 
 tities of tar are deposited at this point, and it is, therefore, al- 
 most impossible to retain the hydro-carbons of this class in the 
 gas with the usual method of condensation. 
 
 Both the nature of the coal used and the temperature of 
 carbonization exercise a great influence upon the quantity, 
 as well as upon the quality of the tar produced; thus both 
 quantity and quality regulate the amount of revenue re- 
 ceived, and in this manner govern to a great extent the 
 final cost of gas production. 
 
 It appears from experiments conducted by Dr. Bunte, and 
 afterwards confirmed by later experiments made by the 
 Paris Gas Company, that the quantity of tar produced varies 
 with the percentage of oxygen contained in the coal, the net 
 results of these experiments leading to the statement that the 
 greater the amount of oxygen present in the coal, the greater 
 will be the amount of tar produced, this statement seeming 
 also to hold good for ammonia production; on the other 
 hand, it is stated that the greater the percentage of "oxygen 
 present, the less will be the production of coke. 
 
8 COAL GAS RESIDUALS 
 
 Tar produced by high temperature carbonization contains 
 but traces of paraffines and small quantities of defines, but 
 the acetylenes, benzenes, hydro-carbons, and naphthalene are 
 present in large quantities, while the nitrogen is usually 
 found in pyridine bases, and carbolic acid and cresols are the 
 principal phenols; this tar also contains quite a quantity of 
 what is known as "free carbon." Tar produced in vertical 
 retorts contains paraffinoids in large measure, and its "free 
 carbon" content, as well as that in coke-oven tar, is low, in 
 spite of the fact that high carbonization temperatures prevail 
 here; the cause of this difference in the quality of the tar is 
 probably due to the travel of the gas, and which travel is in- 
 fluenced principally by the shape of the carbonizing chamber. 
 "Mass" carbonization, as practiced in verticals and coke-ovens, 
 as a rule produces better tar than "layer" carbonization as 
 practiced in the older types of carbonizing apparatus, particu- 
 larly where light charges are in use. 
 
 The term "free carbon" applied to this usual constituent 
 of coal-gas tar is not exactly correct, as this constituent is 
 not really a pure carbon, but a composition containing several 
 other substances as well as volatile matter; experiments in a 
 number of cases showed that after being precipitated, washed, 
 and dried it usually appeared as a brown-black powder yield- 
 ing about 8.5 per cent soluble matter of an oily, rather heavy 
 nature, while the remaining ash contained quite an amount 
 of oxide of iron. Against this we find, according to Allen, Vol. 
 II, part II, page 55, that Behrens found "free" carbon to consist 
 of 91 to 92 per cent of carbon, 3.1 per cent of hydrogen, and 
 0.4 to 0.9 per cent of ash. 
 
 As opposed to high temperature carbonization, the tars 
 produced by low carbonization temperatures usually possess 
 less naphthalene and hydro-carbons of the benzene series, 
 but a larger proportion of paraffines and olefines, while the 
 nitrogen usually appears as anilines. 
 
 The result of experiments conducted by L. T. Wright on 
 the effect of temperature of carbonization as related to tar 
 yield is shown in Table I, the temperatures of carbonization 
 varying between 1112° F. and 1472° F. 
 
 The quality of tar produced also seems somewhat dependent 
 upon the type of retort used, and very extensive tests were 
 
TAR 
 
 9 
 
 made by R. 0. Wyne-Roberts on this premise, the results of 
 these tests being given in Table II, carbonization having been 
 
 TABLE I. 
 
 -TAR 
 
 YIELDS 
 
 
 
 Experiment number 
 
 1 
 
 2 
 
 3 
 
 4 
 
 5 
 
 Specific gravity of tar 
 
 6600 
 1.086 
 
 7200 
 1.102 
 
 8900 
 1.140 
 
 10,162 
 1.154 
 
 11,700 
 1 206 
 
 
 
 Composition of tar (by per cent of 
 weight) 
 
 % 
 
 % 
 
 % 
 
 % 
 
 % 
 
 
 
 Ammoniacal liquor 
 
 1.20 
 9.17 
 10.50 
 26.45 
 20.32 
 28.89 
 
 1.03 
 
 9.65 
 
 7.46 
 
 25.83 
 
 15.57 
 
 36.80 
 
 1.04 
 
 3.73 
 
 4.47 
 
 27.29 
 
 18.13 
 
 41.80 
 
 1.05 
 
 3.45 
 
 2.59 
 
 27.33 
 
 13.77 
 
 47.67 
 
 0.383 
 
 Crude naphtha 
 
 0.995 
 
 
 0.567 
 
 
 19.440 
 
 
 12.28 
 
 Pitch 
 
 64.08 
 
 
 
 carried out in horizontal retorts, vertical retorts, inclined 
 retorts, and chamber ovens. 
 
 TABLE II. — COMPOSITION OF TARS 
 
 Composition 
 
 Horizontal retorts 
 
 Vertical retorts 
 
 In- 
 clines 
 
 Cham- 
 ber 
 
 
 1 
 
 2 
 
 3 
 
 4 
 
 5 
 
 6 
 
 7 
 
 1.119 
 
 S 
 
 9 
 
 10 
 
 Spocific gravity at 
 60° F 
 
 1.20 
 
 1.25 
 
 1.22 
 
 1.25 
 
 1.10 
 
 1.12 
 
 1.13 
 
 1.095 
 
 1.18 
 
 
 
 Composition by 
 distillation .... 
 
 % 
 
 % 
 
 % 
 
 % 
 
 % 
 
 % 
 
 % 
 
 % 
 
 % 
 
 % 
 
 Light oils, below 
 338° F 
 
 3.10 
 
 7.68 
 
 j 10.15 
 1 11.64 
 62.00 
 
 20.00 
 
 1.10 
 13.10 
 13.20 
 72.60 
 
 1.40 
 10.50 
 16.14 
 71.80 
 
 6.00 
 
 1.10 
 
 13.10 
 
 13.20 
 
 72.10 
 
 9.00 
 
 28.70 
 
 5.85 
 
 12.32 
 
 (11.95 
 1 15.96 
 49.75 
 
 2 to 4 
 
 1.70 
 
 21.90 
 
 21.60 
 
 54.80 
 3.60 
 0.40 
 
 3.70 
 
 20.20 
 
 23.10 
 
 52.60 
 4.80 
 4.20 
 
 2.30 
 
 16.50 
 
 20.80 
 
 60.00 
 3.60 
 0.60 
 
 4.40 
 
 28.50 
 
 19.20 
 
 47.50 
 
 2.60 
 2.20 
 
 0.40 
 
 Middle oils, 338 to 
 518° F 
 
 10.20 
 
 Heavy oils, 518 to 
 662° F 
 
 30.10 
 
 Pitch 
 
 53.90 
 
 Naphthalene 
 
 4.70 
 11.10 
 10.10 
 
10 
 
 COAL GAS RESIDUALS 
 
 AMMONIA LIQUOR riR5TLiqHT0lL3 5ECOND01LS 
 DISTIL 
 
 » WTH 
 
 CARBOLIC Ot 
 
 CARBOLIC CEEOSOTE ANTHRACENE PITCH 
 
 aLTOKS" OlLToZTO° OIJ 
 
 COOLFlLTEB 
 
 a. Pros 
 
 wash v 
 
 IZ%HiSOjC%1»; 
 
 WASH WITH WATCB 
 
 I 
 
 WASH With 6%NA0H(%1I0) 
 Pistil Pubihed Oil 
 
 WTCEMtOIATi: 
 
 oil 
 
 SOLVENT ILLUM. 
 
 NAPHTHA OIL 
 
 Oil Distil 
 
 1 
 
 ANTHRACENE 
 
 Cfl*cE 
 Luobic Anthracene Pitch 
 
 OIL 
 
 WASH wrrH«A0H 
 
 i n 1 1 
 
 Benzols Solvent Illuminat Bottoms 
 
 NAPHTHA OILTO ftECASEJ 
 
 ToO.B7S ! l<S 0.9«3 
 
 r 
 
 a 
 
 90* BCNIOLTO KXjfc 50%8ENZOLTo 14CT SOLVENT NAPHTHA 
 
 TO no* 
 
 Neutral Oils 
 
 DOTIL 
 
 pkewtes 
 
 nccoMPOSE ... 
 ^DILUTE H 2 s6 
 
 Deco»^ise With 
 
 SOLVENT NAPHTHALENE 
 
 /(APHTHA 
 
 (Ursolic Acid 
 Diagram I. — Scheme showing the Refining of Fractions. 
 
 The distillation, or refining of coal-gas tar, is usually carried 
 out to the fractions shown in Table III, while Diagram I shows 
 the products secured by refining these fractions. 
 
 TABLE III. — TAR DISTILLATION 
 
 No. 1 
 
 No. 2 
 
 No. 3 
 
 Fractions 
 
 Temperature 
 C. degrees 
 
 Fractions 
 
 Temperature 
 C. degrees 
 
 Fractions 
 
 Temperature 
 C. degrees 
 
 Crude naphtha, 
 or light oils . . 
 
 Heavy oils, dead 
 oils, or creo- 
 sote oils 
 
 Anthracene oils. 
 
 Pitch 
 
 to 170° 
 
 170 to 270° 
 about 270° 
 
 First runnings, 
 or first light 
 oils 
 
 Second light 
 oils 
 
 Carbolic oils . 
 
 Creosote oils. . 
 
 Anthracene 
 oils 
 
 Pitch 
 
 to 110° 
 
 110 to 210° 
 210 to 240° , 
 240 to 270° 
 
 about 270° 
 
 Light naphtha 
 Light oils .... 
 
 Carbolic oils . 
 Creosote oils . 
 Anthracene 
 
 oils 
 
 Pitch 
 
 to 110° 
 110 to 170° 
 
 170 to 225° 
 225 to 270° 
 
 
 
 The Removal of Tar. — ■ In the usual systems as applied at 
 present, the removal, or the precipitation, of tar is accom- 
 plished by condensation or cooling; quite an amount of 
 condensation occurs between the retort and the hydraulic 
 
TAR 
 
 11 
 
 main, a further amount of tar being deposited in the hydraulic 
 main proper, after which the gas is conducted to the con- 
 densers, where gradual cooling further reduces the tar content. 
 After leaving the condensers the gas still contains quite a 
 number of tar globules in suspension, 
 the removal of these last traces of tar 
 being only possible by means of friction, 
 a Pelouze and Audouin condenser being 
 the usual apparatus employed for this 
 purpose. 
 
 This condenser is subject to quite a 
 variety of forms, but the construction of 
 the various types all depend upon the 
 same principle (see Fig. 1). In this 
 machine a friction bell (A) is suspended 
 in a cylindrical casing (B), the bell 
 dipping into a reservoir (C) containing 
 ammoniacal liquor. This bell consists of 
 several sections arranged in pairs con- 
 centrically, the inner section of each 
 pair being perforated with small round 
 holes facing a blank space on the next 
 outer section, the perforations in this 
 outer section being somewhat larger than 
 those of the inner. The gas enters at 
 (D) and passes up and under the friction U^^-r I i ' 
 drum, or bell, finding emission from the !/- — ■+' 
 latter through the holes in the various 
 shells or sections; the tarry vapors are 
 
 I 
 
 -K*- 
 
 " wire-drawn" and condensed in passing Fig. l. — Pelouze and Au- 
 through the perforations, striking against douin Conden « ,r - 
 the blank space opposite the holes and being thus deposited 
 on the shell of the drum, flowing to the bottom and 
 thence to the overflow, the gas passing on and out at (E). 
 The friction drum is suspended and so balanced as to act 
 as its own regulator, the drum rising or falling in the hy- 
 draulic seal as the make of gas increases or diminishes, the 
 amount of seal for maximum make being regulated by the 
 counterweight (F), thus causing a larger or smaller number 
 of openings to be exposed to the passage of gas. 
 
12 
 
 COAL GAS RESIDUALS 
 
 Another type of tar extractor is built in the form of 
 an injector (Fig. 2), and is used in connection with the 
 Otto by-product coke-oven system; here the gas is sprayed 
 with a stream of tar as it enters the apparatus at (A), the 
 tar being collected in a drum and a certain portion being 
 continuously pumped back to the spray by means of a 
 rotary pump. The gas leaves the drum through a series 
 of perforations in the outlet pipe. 
 
 The Simon-Carves Company uses what is known as a 
 cyclone tar extractor, the shell or body of which is cylin- 
 
 Fig. 2. — Injector Type 
 of Tar Extractor. 
 
 Fig. 3.— Cyclone Type of Tax 
 Extractor. 
 
 drical at the top and ending in ■ a deep cone at the bottom. 
 The gas enters at (A) and is whirled around in the body 
 of the machine (B), leaving through the central pipe (C); 
 the condensed tar leaves at the bottom of the apparatus 
 and enters a seal at (D) (see Fig. 3). 
 
 The centrifugal tar extractor is built up of a fan casing 
 containing a series of blades mounted upon a rapidly revolv- 
 
TAR 
 
 13 
 
 ing shaft; the gas enters the casing at (.A) (Fig. 4), and is 
 whirled about through the medium of the fan blades, the 
 heavier tar particles being 
 hurled out against the cas- 
 ing by centrifugal force. 
 The rotary tar extractor 
 (Fig. 5) is built up of a j 
 cylindrical, drum-like shell, 
 a horizontal shaft passing 
 through its center; a per- 
 forated drum is mounted 
 on the shaft, and a second 
 perforated drum, inclosing 
 
 Fig. 4. — Centrifugal Type of Tar Extractor. 
 
 and concentrically surrounding the former, is mounted on a 
 hollow spindle, the two drums revolving in opposite directions 
 through the medium of a set of planetary gears. The per- 
 
 Fig. 5. — Rotary Type of Tar Extractor. 
 
14 COAL GAS RESIDUALS 
 
 forated drums are built in sections, the flanges connecting the 
 sections forming a portion of a helix, so arranged that as 
 the drums revolve in the liquid tar kept in the bottom of 
 the casing, they are caused to pick up tar and spread it 
 over the revolving drums, thus causing the gas to pass 
 through a thin film of tar in finding egress from the 
 interior of the drums. The bottom of the casing is provided 
 with a steam pocket for the purpose of heating the deposited 
 
 Fig. 6. — Centrifugal Type of Static Tar Extractor. 
 
 tar if this should be found necessary. The gas enters the 
 machine at (A) and leaves at (B). 
 
 Another centrifugal tar extractor, without moving parts, 
 is shown in Fig. 6, and it consists of two compartments 
 marked (1) and (2) which communicate with each other 
 through the central orifice (3). The gas enters the device 
 at (4) and is then thrown tangentially against the periphery 
 of the first compartment (1), centrifugal motion being im- 
 parted to the gas by the circular shape of the casing, this 
 centrifugal or swirling motion increasing in rapidity as the 
 gas approaches the central orifice (3), thence passing into the 
 second chamber (2), where centrifugal motion is again 
 imparted to the gas, but in an expanding ratio instead of a 
 contracting one, as is the case in compartment (1). 
 
TAR 15 
 
 M. J. Mazeron (see "Gas World," June 27, 1914) states 
 that the centrifugal force in both the contracting and expand- 
 ing swirl varies inversely as the cube of the radius of the gas 
 path, thus attaining high values. The entering gas is pro- 
 jected against the tar already deposited on the periphery of 
 the inlet compartment (1), while in compartment (2) this 
 action is reversed, thus causing the tangential velocity to 
 decrease as the gas travels towards the periphery of compart- 
 ment (2), the pressure being thus restored. The gas leaves 
 the extractor at (5), while the deposited tar flows out at (6) 
 into the seal pots (7) and (8). Mazeron also states that 
 from the point of view of tar separation, compartment (2) is 
 of greater importance than compartment (1), but that the 
 latter chamber has the effect of making the tar vesicles 
 coalesce. 
 
 The speed of the gas just before it passes through the 
 central orifice (3) is said to be about 500 feet per second, and 
 the pressure absorbed by the device is about 2.4 inches of water. 
 
 The tar, after its removal from the gas, usually contains 
 quite an amount of ammoniacal liquor, from which it is 
 separated in a separating well, the ammonia passing off on 
 one side to storage, and the tar on the other. In order to 
 separate the tar from any liquor or water which may still 
 remain in combination, the tar may be worked up in a cen- 
 trifugal dehydrating plant, shown in Fig. 7. The tar is 
 cleared of water and heavy particles in this machine, and is 
 thus purified. The tar is heated to from 85° to 95° F. in the 
 heater (.A); after which it passes through a filter (B), and 
 thence by means of a special valve into the drum of the 
 centrifugal, the latter making about 2500 revolutions per 
 minute. The heavier tar is thrown out by centrifugal force 
 against the shell of the machine and rises up the sides to the 
 top of the disc (C), where pipe (D) takes up the tar and 
 pipe (E) the liquor, this liquor being returned to the ammonia 
 well. All of the heavier particles in the tar, such as coke 
 dust and other substances which are insoluble in benzol, are 
 collected in the lower portion of the drum and can readily 
 be removed when the drum is being cleaned, it being necessary 
 to clean the drum every six to twelve hours, depending upon 
 the condition of the tar. 
 
16 
 
 COAL GAS RESIDUALS 
 
 At one plant where these centrifugals are in use, the separa- 
 tion of liquor from tar has produced an increased revenue of 
 $5.55 per day with the daily treatment of 2800 gallons of 
 tar, producing 2352 gallons of liquor free tar, selling at 17.5 
 per cent more than the former price, and separating 448 
 gallons of ammonia liquor which is sold independently of the 
 tar, this additional net revenue being secured after deducting 
 
 Fig. 7. — Tar Dehydrating Plant. 
 
 labor, operating cost, and interest and amortization on the 
 capital expended. 
 
 All of these methods for the removal of tar, with the ex- 
 ception of the Otto system, are dependent upon cooling or 
 condensing the gas, and the final product, including that of 
 Otto, must later be distilled in order to secure its various 
 constituents. 
 
 Opposed to this system is the one devised by Feld, in which 
 the heat contained in the gas is utilized by a system of 
 fractional coolings to separate, or wash out of the gas, tar in 
 several of its principal constituents. This treatment may 
 
TAR 17 
 
 also be varied in such manner as to prevent the formation of 
 tar as much as possible by maintaining the heat of the gas 
 coming from the retorts at a temperature above the dew 
 point of the gas for the constituents of high boiling points. 
 This may be done by covering the connections from the 
 retorts to the first washer with insulating material, or by 
 applying heat to the exterior of the pipes. The gas thus 
 protected is led to the washers, where it is subjected to 
 fractional cooling and successive washings at successively 
 lower temperatures, so that the various tar constituents are 
 separated from each other by employing the temperature of 
 the gas itself, and without the necessity of employing ex- 
 traneous heat after previous cooling. 
 
 The Feld System of Fractional Separation of Tar Products. 
 — As no tar separating plants of the Feld system have as yet 
 been erected in the United States, statements showing results 
 obtained in this branch of by-product recovery will neces- 
 sarily have to be confined to European practice. 
 
 In Europe we find that the distillation of 1000 kilos (2200 
 pounds) of coal produces on an average 
 
 30 kilos (66 pounds) of tar, 
 120 kilos ('204 pounds) of water, and 
 300 cbm. (10,590 cu. ft.) of gas, 
 
 measured dry at 0° C. and 760 mm. pressure, and that this 
 average will give 
 
 100 grammes of tar, and 
 
 400 grammes of water per cubic meter of gas. 
 
 The dew point of this gas for water lies between 70° and 
 75° C. (158° and 167° F.), or, to be more explicit, at 72° C. 
 (162° F.), which means that the gas can be cooled to 
 several degrees above 72° C. without condensing the water 
 vapors. 
 
 As stated before, tar is a mixture containing more or less 
 so-called free carbon, with hydro-carbons and other organic 
 substances of various boiling points, generally designated by 
 Feld as follows: 
 
 Light Oils, with a boiling point up to 170° C. = 338° F. 
 
 Creosote Oils, with a boiling point from 170° to 230° C. = 
 338° to 446° F (also termed medium oils). 
 
18 COAL GAS RESIDUALS 
 
 Heavy oils, with a boiling point from 230° to 270° C. = 446° 
 to 518° F. 
 
 Anthracene Oils, with a boiling point from 270° to 320° C. = 
 518° to 608° F., while the residue after distillation, consist- 
 ing of from 20 to 30 per cent of free carbon and hydro- 
 carbons, with a boiling point above 320° C. (608° F.), is known 
 as pitch. 
 
 The products due to tar distillation are of a secondary 
 nature, that is, they are produced by means of a second more 
 or less disintegrating process of distillation, and their condi- 
 tion and respective volumes are not at all identical with the 
 primary tar ingredients contained in the gas. This statement 
 seems to be borne out by the fact that soft pitch produced 
 from the gas by Feld's process only contains from 1 to 3 
 per cent of "free carbon," while the usual pitch produced by 
 distillation contains at times from 20 to 30 per cent of "free 
 carbon." 
 
 Due to the fact that these products are produced by such 
 secondary distillation, it becomes almost impossible to deter- 
 mine by computation how much of these primary constituents 
 were contained in the gas based upon the amount of oils and 
 pitch produced during the distillation of the tar. 
 
 Just as in the case of water vapors, the vapors of tar 
 constituents, which are capable of vaporization at variable 
 temperatures, produce variable tensions, and this vapor ten- 
 sion increases the closer the temperature approaches the 
 boiling point of the liquid in question. According to this 
 vapor tension, the gas, coming in contact with one of these 
 fluids, will absorb a certain portion of the fluid in vapor 
 form at a certain temperature, or the gas will saturate itself 
 with the vapor of the fluid in question. The amount neces- 
 sary for saturation increases with the temperature of the gas, 
 and therefore has a fixed relation to the latter. 
 
 If, on the other hand, a gas having a temperature of B°, 
 saturated with the vapors of any fluid, be heated to a higher 
 temperature of C°, without remaining in contact with the 
 vapor producing fluid, the gas will at the temperature C, 
 or at any other temperature above B, no longer saturate 
 itself with the vapors of the fluid, and the gas will then be 
 superheated with reference to the vapors of the fluid in question. 
 
TAE 19 
 
 If we bring this gas, or a gas of a hydro-carbon character, 
 in a superheated condition into intimate contact with hydro- 
 carbons, it will, just as in the above case of saturation, 
 evaporate hydro-carbons and the gas will saturate itself with 
 the product of this vaporization. The vaporization will be 
 complete in proportion to the heat content due to super- 
 heating, and there will also be a decrease in volume as well as 
 in temperature. 
 
 This reduction in temperature is not, however, always 
 accompanied by cooling, or by the extraction of heat, as the 
 heat content existing before superheating has boon converted 
 into latent heat in the vapors of the hydro-carbon. Only 
 when the fluid to be vaporized is much cooler than the 
 caloric compensating point (see "Cleaning of Blast Furnace 
 Gas," McGraw-Hill Book Co., Inc., New York), and when 
 the mass of the fluid is very large, the heating of the mass 
 which cannot be brought to the point of vaporization will 
 absorb heat and produce a corresponding cooling of the gas. 
 By regulating the mass and the temperature of the fluid to 
 be evaporated it becomes possible to permit or avoid a 
 reduction in the heat contents of the gas. 
 
 The following will explain what Feld terms the "caloric 
 compensating point." The heating of water saturated gas 
 at and from 40° C. to 250° C, without the additional absorp- 
 tion of water vapor, or the superheating of a gas saturated at 
 40° C. to a temperature of 250° C, causes the heat content 
 to increase from 51.7 calories by 75.6 calories, giving a total 
 of 127.3 calories, or just as much additional heat as if the gas 
 had been heated from 40° to 58° C. with the simultaneous 
 absorption of water vapor. This higher temperature, or 58° 
 C. in the above example, is the "caloric compensating point" 
 for a gas superheated to 250° C. and whose dew point for 
 water is at 40° C. 
 
 The changes in condition are therefore, in the case of 
 saturation with the vapors of tar constituents, considerably 
 less than would be the case if the gas had been saturated 
 with water vapors; the heat of vaporization of the hydro- 
 carbons under consideration is about one-seventh that of 
 water vapor. For example, if the heat of vaporization of 
 benzol is 92.9 calories, of toluol 83.6 calories, of xylol 78.3 
 
20 COAL GAS RESIDUALS 
 
 calories, the heat of vaporization of such tar constituents as 
 have a higher boiling point will be still less. 
 
 If the superheated gas is simultaneously mixed with several 
 hydro-carbons, or with such as will permit of mixing or of 
 being absorbed, the one with the lower boiling point will also 
 act as a gas and saturate itself with the vapors of the one 
 having the higher boiling point; but the conditions governing 
 a case such as this are not so readily explained as if we were 
 dealing with water vapors, or with the vapors of non-mixing 
 fluids. In the case of mixing fluids the tension of vaporiza- 
 tion changes for the various constituents, and the final 
 results, which of course are of the most importance, due to 
 the mixing qualities of the hydro-carbons with the constitu- 
 ents of low boiling points, will cause the vaporization of a 
 greater mass of the fluids of higher boiling points than would 
 be the case if the latter fluids were present in an unmixed 
 condition. 
 
 In accordance with the heat contents corresponding to the 
 degree of superheat of the gas, and to the mass of constituents 
 of low boiling points active in the gas, the gas will wholly 
 or only partially saturate itself with the vapors of the fluids 
 in question; it is therefore possible that the gas may saturate 
 itself with the vapors of the high boiling constituents (as 
 their tension of saturation is lower), but will remain super- 
 heated with reference to the low boiling constituents. 
 
 If, for example, the superheated gas should be cooled by 
 means of outside coolers, a portion of the high boiling con- 
 stituents with which the gas had saturated itself, will immedi- 
 ately condense, and the amount of condensation will be 
 proportional to the decrease in temperature; but the lower 
 boiling constituent can only separate itself after the gas has 
 been cooled to the temperature corresponding to the dew 
 point of the fluid in question. 
 
 In this case we also find the conditions which govern the 
 mixable fluids altered, as the dew point of the lower boiling 
 constituents rises if at the same time a high boiling constitu- 
 ent, which will mix with the lower one, is separated from the 
 gas. 
 
 Coal gas and similar gases are in this unsaturated condition 
 with reference to the oily tar constituents contained in the 
 
TAR 21 
 
 gas when it leaves the retort, but it will be saturated with 
 the major portion of the bitumen constituents. The dew 
 points of this gas for these various constituents lie at varying 
 lower temperatures, and are dependent upon the boiling 
 points and the masses of the various constituents of the gas. 
 
 Feld determined that the dew points of the usual coal gas 
 for the greater portion of the bitumen tar constituents lie 
 above 200° C. (392° F.), and for the greater portion of the 
 oil constituents below 200° C. As long as the gases are not 
 cooled below 200° C, only unimportant quantities of the 
 oily tar constituents will be separated; of course, this temper- 
 ature of extraction varies in accordance with the quality of 
 the coal used and of the gas produced. 
 
 In order, then, to extract the tar constituents directly and 
 fractionally from the gas, and which has heretofore only 
 been accomplished by secondary distillation of the tar, Feld 
 proceeds as follows: 
 
 The gases coming from the benches or the coke-ovens, as 
 the case may be, are protected against condensation by 
 wholly or partially insulating the hydraulic main and pipe 
 connections, so that the temperature of the gas does not drop 
 below the point at which tar products of low boiling points 
 are separated. For the removal of the other tar products, 
 Feld employs his centrifugal washer; in these washers progres- 
 sive cooling, or progressive cooling combined with progressive 
 washing, is attained through a progressive drop in tempera- 
 ture, and in this manner various tar products are obtained 
 and separated from each other through the medium of the 
 heat of the gas itself. 
 
 This treatment of the gas can be employed to secure any 
 number of tar constituents; in case it is desired to secure 
 pitch and tar oils of higher boiling points, two washers are 
 employed, and they are operated in such manner that the 
 temperature in the second washer remains above the dew 
 point of the gas for water; or, for example, if it is not 
 desired to separate pitch from the other constituents, those 
 constituents of the tar which have higher boiling points can 
 be separated from the gas in one washer by maintaining a 
 temperature near the dew point of the gas for water. 
 
 Cooling down to the desired temperature can be wholly or 
 
22 COAL GAS RESIDUALS 
 
 in part accomplished in surface coolers, but it is of prime 
 importance and advantage to cool the gas to a given separa- 
 tion point by washing the gas with a liquor produced from 
 the products secured during a later period of separation, or 
 with liquids which contain tar constituents with which the 
 gas cannot saturate itself at the given temperature. 
 
 The drop in temperature of the gas is thus produced by the 
 vaporization of the tar constituents in question. Generally 
 speaking, the gases are cooled at one point of progressive 
 separation by washing them with the fluids of a later period 
 of progressive separation; it is therefore immaterial whether 
 the washing liquids are taken from the gas itself, or whether 
 they are secured from outside sources. In the following 
 description the separation of pitch is included, and it is 
 assumed that the pitch, or bitumen-charactered constituents, 
 shall not contain any products with a boiling point below 320° 
 to 350° C. (608° to 662° F.). 
 
 The gases are cooled down to the dew point of such con- 
 stituents as have a boiling point of from 320° to 350° C. 
 either in the first washer, or just before entering the washer. 
 As is well known, the dew point for the greater number of 
 constituents lies above 140° to 160° C. (248° to 320° F.); 
 within these temperatures such constituents as have boiling 
 points of 320° C. and above are extracted, or such con- 
 stituents with which the gas will not saturate itself at from 
 140° to 160° C. 
 
 The constituents extracted at the given temperature leave 
 the washer as a hot fluid which congeals upon being cooled to 
 about 70° C. (158° F.), depending upon its content of hydro- 
 carbons. 
 
 The gases coming from the first washer are thus practically 
 freed of carbon and pitchy tar constituents, and they contain 
 principally tar oils with boiling points below 350° C. (662° 
 F.); in addition to this, the gases contain such portions of 
 the constituents having boiling points at and above 350° C. 
 as are equivalent to the vapor tensions of the constituents 
 corresponding to a temperature of 160° C. (320° F.), while 
 the vapor tension of the bitumen tar constituents is very low. 
 
 Those tar constituents which have a dew point from 
 160° C. down to that for water, in this case above 75° C, 
 
TAR 23 
 
 are removed from the gas in the second washer, assuming 
 that the dew point of the gas for water is at 72° C; the 
 gas is consequently cooled down from 160° to 75° C. The 
 lowest boiling point of any product separated during this 
 drop in temperature is, for example, at 250° C. (482° F.), 
 and in addition to constituents with boiling points of from 
 250° to 350° C. the product will contain such an amount of 
 constituents with boiling points above 350° C. as will be equiva- 
 lent to the reduction in vapor tension of the gas for these 
 constituents. 
 
 In order to separate as far as possible those constituents 
 having boiling points above 350° C. from those of lower 
 boiling points, as much of the separated product coming 
 from the second washer is pumped into the first washer as 
 is required to reduce the temperature in the first washer 
 down to 140° to 160° C. (284° to 320° F.). Upon the en- 
 trance of this liquid into the first washer, the constituents 
 with boiling points below 350° C. (662° F.) vaporize, because 
 the gases at 160° C. and above are not saturated with them. 
 Those constituents with boiling points above 350° C. with 
 which the gases are saturated at 160° C. cannot vaporize. 
 
 This method of operation is best employed in cooling the 
 gas and at the same time causing the greatest possible 
 separation of the various tar constituents from each other; 
 thus the greatest possible number of constituents with boil- 
 ing points at and above 350° C. are separated in the first 
 washer, while the separated product from the second washer 
 will experience a reduction in contents of such products as 
 have boiling points at and above 350° C. 
 
 In some cases Feld advises the insulation of the washer 
 proper, in order to prevent chilling the outer surface of the 
 gas, and also in order to be in a position to pump as great 
 a quantity of the product coming from the second washer 
 into the first one as is possible. 
 
 Instead of using the tar products which have been taken 
 from the gas for cooling and washing, it is possible to use 
 tar taken from other sources, which tar is thereby separated 
 into its constituents and later recovered in its separated 
 state during the process of cooling and washing. It is well 
 to note here that this tar, which is thus distilled through 
 
24 COAL GAS RESIDUALS 
 
 the agency of the heat of the gas itself, does not take up 
 any free carbon because the distillation is accomplished at 
 very low temperatures. 
 
 Instead of fractionately separating the tar constituents 
 of high boiling points as described above, the constituents 
 can be jointly washed out at a temperature slightly above 
 the dew-point temperature of the gas, in this case from 72° 
 to 80° C. (162° to 176° F.), but in both cases the gas leaving 
 the last, or heavy oil washer as explained above, or upon 
 leaving the hot tar washer described below, will be free of 
 tar and at the same time retain a temperature above the 
 dew point for water. As far back as 1903 Feld freed the 
 gas of tar by hot tar washing before a temperature equiv- 
 alent to the dew point for water was reached. 
 
 A Feld washer especially arranged for hot tar washing is 
 shown in Fig. 8, and in order that condensation may be 
 avoided in the washer by reducing the temperature of the 
 gas to the dew point for water, the washing tar should be 
 heated to from 40° to 60° F. above the temperature of the 
 gas at the inlet to the washer, this hot tar being the active 
 tar-extracting medium. 
 
 Referring to Fig. 8, which shows a portion of a Feld 
 washer with superimposed Pelouze condenser, 
 
 A = Pelouze condenser, 
 B = Gas port in washing chamber, 
 C = Upper washing chamber, 
 D = Lower washing chamber, 
 E = Gas inlet, 
 
 k = Basin, or tank containing wash tar, 
 
 r = Radiator, or steam coil, 
 
 I = Rotary tar pump, 
 m = Preheater for wash tar, 
 
 s = Insulated connection from tank (fc) to pump (l). 
 
 In order to keep the wash tar at the desired temperature, 
 the steam pipe from (m) is located beneath the pipe (s). 
 
 The hot wash tar is pumped from tank (k) through the 
 insulated connection (s) and preheater (m) to the Pelouze 
 condenser (A) through the overflow pot (n); any surplus 
 tar overflows at (o) and (p) and returns to tank (fc). 
 
 The tar overflows from the Pelouze through (b) and (d), 
 (e) and (g) into the upper chamber (C) of the washer, where 
 
Missing Page 
 
TAR 25 
 
 it is picked up by the washing cones within the chamber 
 and spread out over the entire gas space, the depth of the 
 spray being from 8" to 14", depending upon the size of the 
 washer. The gas is thus brought into intimate contact with 
 the hot wash tar and is subjected to a thorough washing, 
 the wash tar, as well as the tar expelled from the gas, over- 
 flowing through the gas port to the next lower washing cham- 
 ber; here the wash tar is again placed in circulation by the 
 cones in this chamber, and the total combined tar finally 
 leaves the washer through the pot (i) attached to the gas 
 inlet connection, and by means of overflow (h) enters tank 
 (/»'), all surplus tar from this tank flowing off to the tar 
 storage. 
 
 In order that the maximum results may be secured, the 
 wash tnr should be constantly circulated through the washer, 
 but it is not always necessary to circulate it through the 
 Pelouze condenser, this omission being dependent upon the 
 condition of the tar in the gas. The amount of hot tar to 
 be circulated through the washer may vary from 0.5 to 2.5 
 gallons per 1000 cubic feet of gas per hour; larger volumes 
 of gas require less circulating tar than do smaller volumes, 
 but the amount of washing tar circulated must be so regu- 
 lated as to prevent the temperature of the gas reducing the 
 temperature of the wash tar. 
 
 The circulating wash tar should be pumped through the 
 preheater (m) before it enters the washer, this preheater 
 being supplied with a steam coil, and in some cases it may 
 be necessary to insulate the tank containing the wash tar, 
 depending upon its location, and even the washer proper 
 with some insulating material in order to increase the effi- 
 ciency, as many of these washers are set up in the open in 
 Europe, being thus exposed to the weather. 
 
 The tank containing the wash tar is also provided with a 
 steam coil, or radiator, made in one piece in order to prevent 
 leaks and thus mix water with the tar. Care must be 
 exercised in operation to the extent of preventing a deposit 
 of thick tar in the Pelouze condenser, and if the washer 
 should be stopped for any reason it must be emptied of tar 
 at once in order to avoid cooling the tar and thus clogging 
 up the passages. 
 
26 
 
 COAL GAS RESIDUALS 
 
 TABLE IV.— FELD HOT TAR WASHING 
 
 Washer without Pelouze Condenser 
 
 Gas tem- 
 perature 
 
 Tar content in grains 
 per cu. ft. 
 
 Grains 
 removed 
 
 Efficiency 
 
 % 
 
 Gas passed per 
 24 hours cubic 
 
 inlet to 
 washer 
 
 Inlet 
 
 Outlet 
 
 feet 
 
 131° F. 
 127° F. 
 129° F. 
 129° F. 
 126° F. 
 
 11.512 
 10.386 
 10.818 
 13.518 
 9.907 
 
 0.2930 
 0.2160 
 0.2930 
 0.4170 
 0.5250 
 
 11.219 
 10.170 
 10.525 
 13.101 
 9.382 
 
 97.46 
 97.92 
 97.29 
 96.92 
 94.70 
 
 2,463,000 
 2,453,832 
 2,341,800 
 2,397,000 
 2,400,000 
 
 Washer with Pelouze Condenser 
 
 140° F. 
 194° F. 
 
 10.050 
 13.100 
 
 0.0436 
 0.0872 
 
 10.0064 
 13.0128 
 
 99.56 
 99.34 
 
 2,600,000 
 2,600,000 
 
 Table IV gives the results secured in washing with hot 
 tar, the washer being supplied with a Pelouze condenser in 
 the one instance, but not in the other. 
 
 This method of operation has made it possible to extract 
 ammonia, hydrogen-sulphide, and cyanogen by chemical 
 washing at high temperatures, and which would not be 
 possible if the gas contained tar; this high temperature 
 washing not only prevents the production of weak solutions 
 of these valuable gas constituents, but it also prevents losses 
 to a great extent. The tar-free gas, at a temperature with 
 which it leaves the washer, is subjected to further washing 
 in a series of centrifugal washers, and as explained in the 
 following pages, where the ammonia, cyanogen, and, if the 
 plant is so arranged, also hydrogen sulphide, is removed. 
 
 After leaving the last ammonia washer the gas still retains 
 a temperature close to that of its dew point for water, and 
 is therefore still saturated with water vapors; besides thisj 
 the gas also carries the vapors of those tar constituents 
 which have boiling points of 250° C. and above, these having 
 been separated from the tar in the washers preceding those 
 for ammonia, etc. 
 
TAR 27 
 
 This gas also contains all tar constituents having boiling 
 points below 250° C, that is, middle oils, naphthalene, light oils, 
 xylol, toluol, and benzol. In order to expel these middle 
 oils, that is fractional distillation at boiling points down to 
 200° C, the gas is so far cooled as to retain a temperature 
 still somewhat higher than the dew point for light oils, which 
 latter have their boiling points at and below 200° C Accord- 
 ing to Feld, after having been subjected to hot tar washing, 
 coal gas can be cooled down to 18° C. ((55° F.) without 
 expelling any appreciable quantity of light oils. 
 
 This cooling down from 72° to 18° C. can now, as the gas 
 no longer contains any valuable constituents which are soluble 
 in water, be performed in an additional centrifugal washer by 
 bringing the gas into direct contact with water. With the 
 middle oils a volume of water, corresponding to the drop in 
 temperature, is also expelled, or in the case described above, 
 400 — 15 = 385 grammes of water per cubic meter of gas. 
 
 This water readily separates from the oil expelled with it, 
 and this oil contains nearly all of the naphthalene emulsified 
 in water; after the water has been separated from the mix- 
 ture the naphthalene can be directly removed. If the expul- 
 sion of naphthalene by cooling down to 18° ('. should be so 
 great as to cause stoppages, the cooling should be performed 
 in two fractions; first, the gas should be cooled to slightly 
 below the dew point for naphthalene, 25° to 35° C (77° to 
 95° F.), and then the naphthalene washed out of the gas by 
 employing the middle oils expelled in the preceding washer 
 as a washing medium. 
 
 After complete cooling the gas now still contains all of the 
 light oils, especially the low boiling hydro-carbons like 
 benzol and its homologues, and with which the gas has not 
 saturated itself, that is, their partial tension in the gas is 
 lower than the vapor tension of those constituents at the 
 lowest temperature to which the gas has been cooled. 
 
 A gas thus treated, and one which is to be used for illu- 
 minating purposes, possesses a higher illuminating power than' 
 if it had been condensed and purified in the usual manner, 
 or by the common method of condensation with the com- 
 mon expulsion of the various constituents. 
 
 In the present practice, due to common separation and 
 
28 COAL GAS RESIDUALS 
 
 expulsion, the tar constituents of high boiling points act as 
 solvents for those of lower boiling points, but, as is well 
 known, the vapor tension of a solution is always lower than 
 the vapor tension of the pure constituent, and therefore, in 
 accordance with the old practice of purification, low boiling 
 tar constituents, such as benzol, are to a greater or less 
 extent expelled from the gas, thus reducing its illuminating 
 power in direct proportion to the amount thus removed. 
 
 These low boiling constituents, which are so valuable to 
 the illuminating quality of the gas, will, due to the present 
 ■method of condensation, pass off or be expelled in part with 
 the tar, from which they must be regained by separate dis- 
 tillation, while in the Feld process of condensation they 
 remain in the gas. 
 
 If it should be desired to also expel these constituents 
 from the gas, the latter should be washed after final cooling 
 with tar oils of high boiling points, and which were expelled 
 from the gas at a previous period of the process. Such high 
 boiling tar oils have, at the low temperature in question, a 
 very low tension, and as this tar oil was taken from the gas 
 itself the gas is already saturated with these oils in an amount 
 equivalent to the low vapor tension. In consequence, due 
 to the use of these high boiling tar oils for washing purposes, 
 or for removing benzol and its homologues from the gas, 
 the gas will not absorb any of the oil, but on the contrary, 
 as the gas suffers a reduction in volume when the benzol is 
 expelled, a very small volume of this wash oil is also expelled, 
 together with other high boiling tar constituents from the 
 gas, and the latter is thereby made cleaner. 
 
 The light oils can now be distilled from the products 
 having high boiling points, and which were used for washing 
 out the light oils, after which they are again returned to the 
 washers as washing media; as soon as this wash oil has 
 saturated itself with impurities, and thus becomes valueless 
 for benzol washing, the impure oils of high boiling points 
 are used for washing in a preceding stage, for instance in 
 the first or second washer for cooling and washing purposes, 
 where the low boiling constituents are vaporized and then 
 recovered in their original and pure condition further on in 
 the process, while the constituents of higher boiling points 
 
TAR 29 
 
 remain in the product coming from the washer in question. 
 
 Feld's process for the separation of tar from gas has been 
 in operation at various European plants since 1906, orig- 
 inally as an experiment but finally as a permanent method 
 of operation. Below are given several operating results 
 secured in practice. 
 
 At a coke-oven plant in Upper Silesia the gases produced 
 by the carbonization of 395 tons of coal per 24 hours were 
 treated in two Feld washers operated in parallel, the tem- 
 perature of the gas averaging 150° C. (302° F.) with the 
 following results: 
 
 Temperature of gas, 150° to 160° C. = 302° to 320° F. 
 
 Temperature of pitch expelled from washer 150° C. = 302° F. 
 
 Product obtained, 5500 pounds of thin soft pitch. 
 
 Upon heating this pitch to 340° C. (644° F.) the following 
 products were obtained: 
 
 Test No. 1 Test No. 2 
 
 Oil and water at 315° C. (599° F.) 2.6 % 7.0 % 
 
 Residue after 340° C. (644° F.) 97.4 % 93.0 % 
 
 The Feld process was installed at a Bohemian "brown" 
 coal works in 1907; here the gases pass through a number of 
 Feld washers and surface coolers. The results given during 
 a test run showed that due to the large percentage of water 
 in the "brown" coal, the dew point of the gas was found to 
 be between 90° and 92° C. (194° and 198° F.). The gases 
 are brought from the ovens with a temperature of 110° C. 
 (230° F.), and enter the first Feld washer at 89° C. (192° F.); 
 owing to the large amount of condensed water, three Feld 
 washers were used in series in order to wash out the sepa- 
 rated bitumen and heavy oils, which, at 89° C, were present 
 in the water in the shape of fog. The expelled products at 
 various temperatures are given in Table V. 
 
 The constituents gained upon distilling these products are 
 given in Table. VI; columns 2, 3, and 4 give the quantities 
 secured in a percentage of the whole distilled off at the tem- 
 peratures given in column 1. 
 
30 COAL GAS RESIDUALS 
 
 TABLE V. — BROWN COAL PRODUCTS 
 
 
 Gas tem- 
 perature in 
 degrees C. 
 
 Expelled product 
 
 Efficiency 
 
 % 
 
 Apparatus 
 
 Bitumen 
 lbs. 
 
 Oil I, 
 
 lbs. 
 
 Oil II, 
 
 lbs. 
 
 1st washer 
 
 2d washer 
 
 3d washer 
 
 1st cooler 
 
 89 
 
 89 
 
 89 
 
 76.8 
 
 35 
 
 4849.0 
 
 132.0 
 
 3.5 
 
 4984.5 
 
 1032 
 1932 
 
 1320 
 1320 
 
 97.4 
 2.5 
 0.1 
 
 Total in 24 hours . . 
 
 
 TABLE VI. — PRODUCTS OF DISTILLATION 
 
 Aggregate 
 
 Bitumen 
 
 Oil I 
 
 Oil. II 
 
 Oil III 
 
 condition 
 
 Hard, melts 
 at 30° C. 
 
 Light, running fluid 
 
 Very light 
 
 Spec, gravity 
 
 0.9647 at 36° C. 
 
 0.92 at 16° C. 
 
 0.896 at 16° C. 
 
 
 Temperature of 
 saturation 
 
 to 89° C. 
 
 From 
 98° to 76.8° C 
 
 From 
 76.8 to 35° C. 
 
 From 
 78° to 20° 
 
 Products of Distillation 
 
 To 120° C. 
 
 " 140° C. 
 
 " 160° C. 
 
 " 180° C. 
 
 " 200° C. 
 
 " 220° C. 
 
 " 240° C. 
 
 " 255° C. 
 
 " 270° C. 
 
 " 285° 
 
 " 300° 
 
 " 320° 
 
 " 340° 
 
 " 360° 
 
 " 365° 
 Water 
 Residue of soft 
 pitch . . . 
 
 'C. 
 
 • c. 
 •c. 
 
 'C. 
 'C. 
 'C. 
 
 % 
 
 1.0 
 
 1.5 
 
 1.8 
 
 3.0 
 
 10.5 
 
 17.5 
 
 26.0 
 
 35.0 
 
 43.5 
 
 57.0 
 
 66.0 
 
 76.5 
 
 80.0 
 
 1.0 
 
 19.0 
 
 % 
 
 0.5 
 1.5 
 5.0 
 24.5 
 56.0 
 78.5 
 86.5 
 92.5 
 96.0 
 
 % 
 0.2 
 2.5 
 6.0 
 25.0 
 58.5 
 83.5 
 94.5 
 98.0 
 
 % 
 9.0 
 15.0 
 40.0 
 57.5 
 87.0 
 95.0 
 
 0.5 
 
 3.5 
 
 0.2 
 
 1.8 
 
TAR 
 
 31 
 
 The last column of Table VI shows a light oil which was 
 expelled or separated when the gases were cooled down from 
 7S° to 20° C. 
 
 It should be noted here that before the introduction of 
 the Feld process of fractional distillation jn washers, the 
 production of pitch by the usual system of condensation and 
 distillation amounted to as much as 35 per cent, and after 
 the introduction of this process the pitch production in the 
 same works was as low as 7.5 to 12 per cent, the production 
 of valuable oils increasing in the same ratio. The tar pro- 
 duced by the old process sold for $5 per ton, while, with the oils 
 produced directly from the gas by Fold's system, the equiv- 
 alent in oils and pitch sold for $16 per ton. 
 
 At the coke works of the Ilsider Huette, fractional dis- 
 tillation, by Feld's process, has been in operation for some 
 time with the results given in Table VII. 
 
 TABLE VII. — ILSIDER RESULTS 
 
 
 Product No. 1 
 
 Product No. II 
 
 Temperature of 
 separation 
 
 to 70° C. 
 
 70° to 30° C. 
 
 
 Products of 
 
 % 
 
 Distillation 
 
 
 
 To 135° C 
 
 
 3.75 
 
 " 200° C 
 
 1.1 
 
 10.75 
 
 " 230° C 
 
 1.9 
 
 2.4 
 
 
 " 240° C 
 
 37.75 
 
 " 250° C 
 
 
 58.55 
 
 " 270° C 
 
 12.4 
 
 68.95 
 
 " 300° C 
 
 17.6 
 
 80.05 
 
 " 340° C 
 
 24.3 
 75.7 
 
 
 Residue after distillation 
 
 19.95 
 
 Product No. I, Table VII, is absolutely free of naphtha- 
 lene and only contains from 1 to 3 per cent of "free carbon"; 
 according to reports received from the Huette this product 
 is an excellent material for lining the steel converters, owing 
 to its high percentage of bitumen. While the use of the 
 tar ordinarily produced and containing from 20 to 30 per 
 cent of "free carbon" only permitted of from 25 to 30 
 charges in the converter before relining became necessary, a 
 lining of this new product permitted of 02 charges, all other 
 
32 
 
 COAL GAS RESIDUALS 
 
 conditions remaining the same, and the life of the converter 
 lining was thus increased by more than 100 per cent. 
 
 Product No. II is a middle oil from which the greater 
 portion of the naphthalene can be pressed; after cooling the 
 gases down to 18° C. the oil of product II is used to wash 
 out the light oils and benzol. 
 
 The tar-washing plant erected by Feld at Moncieu, Bel- 
 gium, is being operated with the results shown in Table 
 VIII. 
 
 TABLE VIII. — MONCIEU RESULTS 
 
 Temperature of separation 
 
 Bituminous tar 
 
 Creosote oil 
 
 Product I 
 
 Product II 
 
 Product III 
 
 Specific gravity 
 
 I at 25° C 
 
 II at 20° C 
 
 III at 17° C 
 
 1.175 
 
 1.135 
 
 0.96 
 
 Viscosity at 50° C. (H 2 at 50° = 1) . 
 
 Carbon contents (Anyline, Pyridine, 
 
 Methane) 
 
 8.0 
 0.23% 
 
 6.9 
 0.32% 
 
 
 Products of Distillation 
 
 
 % 
 
 J 
 
 % 
 
 To 140° C 
 
 
 
 
 " 190° C 
 
 
 
 
 10.0 
 
 " 200° C 
 
 
 
 
 29.5 
 
 " 210° C 
 
 
 
 
 47.5 
 
 " 220° C 
 
 
 
 
 66.0 
 
 " 230° C 
 
 
 
 
 81.5 
 
 " 240° C 
 
 
 
 
 88.0 
 
 " 250° C 
 
 
 
 3.5 
 
 93.0 
 
 " 260° C 
 
 
 
 
 96.0 
 
 " 270° C 
 
 2.0 
 
 9.5 
 
 
 " 290° C 
 
 9.5 
 
 11.5 
 
 
 " 300° C 
 
 11.5 
 
 13.0 
 
 
 " 310° C 
 
 13.0 
 
 16.5 
 
 
 " 320° C 
 
 15.0 
 
 
 
 " 330° C 
 
 
 18.5 
 
 
 " 350° C 
 
 
 24.5 
 
 
 " 360° C 
 
 36.5 
 
 45.5 
 
 
 Residue 
 
 60.0 
 
 45.3 
 
 40.0 
 
 Melting point of residue 
 
 53.5° C. 
 
 60.5' 
 
 C. 
 
 Acid oils 
 
 1.0% 
 
 2.9 
 
 7c 
 
 Bases 
 
 1.8 
 
 2.2 
 
 
 
 
 
TAR 
 
 33 
 
 Another system of tar extraction, and one which has been 
 used with success on producer gas, consists in cooling the 
 gas to a degree sufficient to condense the tar vapors, after 
 which the gas, together with the tar, is passed through a 
 porous diaphragm of spun glass, placed between two metal 
 screens. It is stated that in passing this diaphragm an 
 important change occurs in the physical state of the tar. 
 On the inlet side, the tar exists as a large number of minute 
 particles ordinarily known as tar fog, and in passing the 
 diaphragm these particles coalesce, so that on the discharge 
 side the tar particles are of relatively large dimensions which 
 cannot be carried along with the gas current and they im- 
 mediately separate out by gravity. 
 
 This extractor is shown in Fig. 9; after the gas has been 
 cooled, it is delivered under pressure by means of the pump 
 (B) into the main (C), then passing through the diaphragm 
 (E) and discharged 
 from there into the 
 main (F), while the 
 tar accumulates in the 
 separator (G), from 
 whence it is with- 
 drawn to storage. 
 
 In the electrical pre- 
 cipitation of tar the 
 duty of the electric 
 field is to agglomerate 
 the fine tar particles 
 so that they can read- 
 ily be removed later 
 in some form of mechanical separator, and it is stated 
 that this agglomeration can proceed without regard to 
 temperature, thus retaining the light oils in the gas and 
 increasing the illuminating power. This latter system closely 
 approaches the results secured with that of Feld in that 
 tar can be removed at a temperature above that of the 
 dew point of the gas for water, but the Feld is the only 
 system which can secure these results with a simultaneous 
 separation of the tar into several of its fractions. 
 
 Experiments covering the electrical separation of tar from 
 coal gas have been under way for some time, and the principal 
 
 Fig. 9. — Static Type of Tar Extractor. 
 
34 COAL GAS RESIDUALS 
 
 difficulty to be surmounted seemed to lie in the selection and 
 design of the electrodes. 
 
 In a report made to the Michigan Gas Association by Prof. 
 A. H. White, R. B. Rowley and C. K. Wirth, on the experi- 
 ments conducted at the plant of the Ann Arbor Gas Com- 
 pany, Ann Arbor, Mich., it is stated that the tar separator 
 consisted of a six-inch gas pipe, felt-jacketed, about six feet 
 long, set vertically; the gas was conducted into the separator 
 by means of a 2" diameter pipe attached to the bottom, and 
 it left the separator on the side near the top. 
 
 The electrode, consisting of a cylindrical series of fine wires 
 mounted on the circumference of cast-iron discs, was ce- 
 mented into a 3" nipple, this nipple being screwed into the 
 top flange of the extractor. 
 
 Gas at the rate of 11.2 cubic feet per minute was admitted 
 into the extractor, thus causing the suspended particles in the 
 gas to be in contact with the electrical field for a period of 4.7 
 seconds, the current used amounting to 1.2 milleamperes at 
 10,000 volts, which was sufficient to remove all of the tar 
 carried by the gas into the extractor. 
 
 This experiment was followed by another of greater capacity 
 in which the extractor consisted of an inverted "U" made of 
 8" cast-iron pipe and tees, the electrodes being attached to 
 the tops of the pipes. These electrodes were four inches in 
 diameter and were strung with 16 No. 27 steel wires, each 
 electrode having an effective length of 5 feet 8 inches. 
 
 Twenty-eight thousand cubic feet of gas were passed into 
 the extractor per hour, and if the entire cross-section were 
 free of obstruction, the speed of the gas would be 22.6 feet per 
 second, giving the suspended particles a contact period of 
 0.5 second in the electrical field. If the space outside of the 
 electrode only is considered, the speed of the gas would be 30.2 
 feet per second, and the contact period 0.38 second. 
 
 The difference in potential between the electrode wires 
 and the pipe forming the extractor, the latter being grounded, 
 was maintained at 20,000 volts, with a current consumption 
 of 2.6 to 3 milleamperes, while the- watt-meter on the input 
 side of the transformer varied in its readings from 200 to 230 
 watts, it being stated that the results, as regards tar removal, 
 were very satisfactory. 
 
TAR 35 
 
 TAR PRODUCTS 
 
 The many valuable tar products, so much in demand at 
 present, are secured by distilling the raw tar into its various 
 constituents in stills of either the vertical or horizontal type, the 
 former finding great favor in Europe, while the latter is used to 
 a large extent in this country. 
 
 Before running the raw tar into the still, it should be sub- 
 jected to some character of dehydration process, and while 
 many processes have been proposed to accomplish this, the 
 following two methods are those most extensively used. In the 
 one case the raw tar is subjected to a prolonged rest in large 
 reservoirs, followed by a centrifugal separating process, as ex- 
 plained on page 15, and in the other it is pumped into stills 
 of large dimension until the tar begins to flow from a cock 
 located immediately below the top dome, after which the flow of 
 tar is stopped and the still is subjected to heat. 
 
 The heat applied to the still causes the tar to become very 
 thin before it reaches the boiling point, the greater portion of 
 the water being collected on the upper surface of the thin tar 
 and, as the heating continues, the level of the liquid is caused 
 to rise above the overflow cock, thus allowing the water to be 
 drawn off from time to time by opening the water cock; any 
 vapors which may be produced during this heating process are 
 passed through a condenser and recovered. 
 
 The vertical type of still is to be preferred over the horizontal, 
 as its rate of evaporation is more rapid; this difference in time of 
 evaporation causes about one-half the amount of fuel to be con- 
 sumed per unit of tar distilled in the vertical still as compared 
 with the horizontal, but it has many operating disadvantages 
 not inherent in the horizontal type. 
 
 The vertical still and its apurtenances is shown in Fig. 9a, 
 while the horizontal type is shown in Fig. 9b. The vertical 
 still is usually constructed of a diameter about equal to its 
 height, the shell being supplied with a concave bottom and a 
 domed top, both the bottom and the top being constructed in 
 sections for the purpose of adding strength to the structure and 
 to facilitate repairs. 
 
 The concave bottom is not sprung directly from the lower end 
 of the shell, but it is curved into a channel groove, the outer 
 
36 
 
 COAL GAS RESIDUALS 
 
 upper edge of this groove being attached to the shell. The 
 concave bottom considerably increases the heating surface of the 
 still, and it also permits the metal to expand and contract with 
 ease. The direct flame does not come in contact with the still 
 bottom, this being protected by a brick arch, but the hot gases 
 reach the bottom and lower half of the still through flues built 
 into the brickwork for this purpose. 
 
 The horizontal still is set in brickwork as shown in Fig. 9b, 
 and it is heated on somewhat less than half of its cylindrical 
 surface, the metal being protected by an arch immediately 
 
 Pitch ciofar 
 
 Fig. 9a. — Vertical Still. 
 
 above the fire and so designed that the portion of the shell sub- 
 jected to heat may be quickly replaced if damaged. 
 
 All exposed portions of stills should be lagged; this is not 
 only on account of the saving of fuel, but it also prevents undue 
 corrosion, because the latter is prone to occur with great rapidity 
 at points where excessive condensation is possible. 
 
 The condensing tanks for the condensing worms are made of 
 either cast or wrought iron, and may be of any shape suitable 
 for the purpose. These tanks are provided with a drain cock 
 for the purpose of running off the water when required for re- 
 pair purposes, while an outlet is provided near the top to permit 
 the cooling water to escape during the first period of distillation. 
 
 The condensing worms consist of a series of coils of either 
 cast iron or wrought iron pipe, the latter being preferable, owing 
 to the more rapid transmission of either heat or cold. 
 
TAR 
 
 37 
 
 From the condenser the distillate is run to the receiving and sep- 
 arating vessels, the receiver having several outlets, provided with 
 cocks, each outlet communicating with a separate main in order 
 to conduct the different distillates to their proper storage vessels. 
 
 Rich still should be provided with a pitch cooler, usually a 
 cylindrical tank, from which the cooled pitch is run into the 
 pitch bay. 
 
 Fig. 9b. — Horizontal Tar Still. 
 
 The fuel used under the still may be either coal, coke, or 
 producer gas, the latter permitting of better control of the dis- 
 tillation process and therefore being preferred to the solid fuel. 
 
 When the still is properly charged, and all connections are 
 found clear, the firing is started, especial care being exercised 
 during the first period of heating to prevent priming or boiling 
 over. When the tar is thoroughly heated and drops of dis- 
 tillate appear at the ends of the condensing worms, the heating 
 is somewhat lessened, and great caution must now be exercised, 
 as it is during this period when priming or boiling over usually 
 occurs. If the still should prime the fires must be dampened 
 and cold water poured on the dome of the still, this procedure 
 usually stopping the trouble. 
 
38 COAL GAS RESIDUALS 
 
 The first portion of the distillate is ammonia water and crude 
 naphtha, which passes from the collecting vessel into a separator, 
 where the oil is separated from the water by difference in gravity, 
 the water going to one side and the oil to the naphtha receiver. 
 The temperature can now be raised a little as the volume of 
 water becomes less; this first fraction, known as "fight oil," 
 comes off at between 70° to 170° C. (158° to 338° F.), and is 
 run into the light oil receiver. During the production of this 
 fraction the water in the condenser tank should be kept cold. 
 
 This fraction is usually about 3 per cent of the whole, and the 
 intermediate products produced from it are Benzene, Toluene, 
 and Xylene, while the crude commercial products are Benzol 
 and Solvent Naphtha used as solvents, paint thinners, motor fuel, 
 and gas enrichment. The intermediate chemical products are 
 Nitrobenzene, Aniline Salts, Aniline Oil, and Carbolic Acid, 
 while the refined chemical products produced from these are the 
 Nitrotoluenes, Diphenylamine, other explosive ingredients, Ani- 
 line Dyes, Hydroquinone, and other photographic developers, 
 as well as drugs and medicines. 
 
 The distillate coming from the worm will now be practically 
 free of water, the fire is gradually increased, and the distillate is 
 thus kept running at an even rate. The flow of cooling water 
 through the condensing tank i^ cut off, and the temperature of 
 the water in the tank is gradually allowed to increase by ab- 
 sorbing heat from the distillate as the latter passes through the 
 worm. The condenser tank is provided with an open steam 
 connection to permit of increasing the temperature of the water 
 when heat absorption from the condenser worm is not sufficiently 
 rapid to produce the required results. This necessity for heat- 
 ing the water is due to the increasing quantity of naphthalene 
 salts in the distillate, and the worm must therefore be kept 
 warm to prevent the deposit of the salts and the consequent 
 choking of the worm coils. 
 
 This second fraction is produced between 170° and 230° C. 
 (338° and 446° F.) and at the end of the period it is again cut, 
 this second fraction being known as "middle, or dead oil," and 
 the run of the succeeding distillate is diverted to the next 
 receiver. 
 
 The intermediate products from the middle oil are Phenol, 
 Cresols, Naphthalene, and Heavy Hydrocarbons. These inter- 
 
TAR 39 
 
 mediates in turn produce Creosote Oil, Lamp Black, and some 
 disinfectants. The intermediate chemical products are Carbolic 
 Acid, Picric Acid, Phthalic Acid, Naphthols, Naphthylamines, 
 and Salicylic Acid, while the refined chemical products consist of 
 Picric Acid, Picrates and other Nitro-compound explosive ingredi- 
 ents, Naphthol dyes and colors, Indigo, and refined Carbolic Acid. 
 
 The yellow color of the distillate now increases and the odor 
 of hydrogen sulphide predominates; the intensity of the fire is 
 again increased and the temperature of the water surrounding 
 the condensing worm is also allowed to increase. This cut is 
 taken off between 230° and 360° C. (446° to 680° F.) and is 
 known as Heavy, or Anthracene Oils. Another cut is sometimes 
 made at this point, the first between 230° and 270° C. (446° and 
 518° F.) and the second above 270° C. (518° F.). At the end of 
 this fraction some trouble may be experienced by the oil, while 
 cooling, becoming solid, due to the presence of naphthalene 
 salts, and the water in the condensing tank must be brought 
 to the boiling point. 
 
 As the distillation proceeds the fire is slackened and steam is 
 introduced into the still, this procedure lengthening the life of 
 the still and also somewhat reducing the time required to distill 
 the tar. The time for admitting the steam as well as the time 
 for stopping operations cannot be pre-determined, as it depends 
 on the character of the tar distilled and upon the character of 
 the pitch required. 
 
 The intermediate products due to this fraction are Cresols, 
 Naphthalene, Anthracene, Heavy Hydrocarbons, and Quinoline 
 Bases, which in turn give Creosote Oil, Lamp Black, Road Oil, 
 Timber Impregnators, Roofing and Paving Tars as crude com- 
 mercial products, the intermediate chemical products being 
 Anthraquinone and Alizarin. 
 
 Above 360° C. (680° F.) the residue is pitch, amounting to 
 about 65 per cent of the total tar. Owing to the heat still 
 stored in the still after operations are completed, the pitch 
 residue is retained in the still for several hours, after which it is 
 run into the coolers and from there into the pitch bay, where it 
 is allowed to solidify. 
 
 The above fractions and the temperature of "cutting" are 
 taken from an average sample of tar, but the three kinds of 
 coal tar produced vary greatly in their characteristics. Horizon- 
 
40 COAL GAS RESIDUALS 
 
 tal retort tar is the heaviest, and contains a low percentage of 
 oils and a high percentage of pitch, while its "free" carbon 
 content is also high. Coke-oven tar is lighter than that from 
 horizontal retorts and contains a higher percentage of oils and 
 a lower one of pitch. Vertical retort tar is the lightest of the 
 three, and contains more oils and less pitch and "free" carbon 
 than do the others. 
 
 From these differences it will readily be seen that the chemical 
 constituents also vary greatly, and that the quantity of any 
 particular constituent evolved from the tar will depend upon 
 its original constitution. 
 
 Continuous distillation has also been tried both in Europe 
 and in this country, and while success has varied in differ- 
 ent localities, encouraging results have lately been announced, 
 the principal systems being those of Hirzel, Kubierschky and 
 Borrmann, Sadewasser, and Raschig, all of these having been 
 devised in Europe. 
 
 In the Hirzel system two or more small direct-heated stills 
 are employed, a thin layer of tar flowing over the hot bottoms 
 of the stills. The first still drives off the ammonia water and 
 light oil, the middle oil being produced in the second still, and 
 the heavy oil in the third, providing the tar is to be fractionated 
 in this manner, while soft pitch leaves the last still in a con- 
 tinuous stream. Distillation in this case is effected by the use 
 of superheated steam. 
 
 The process devised by Kubierschky and Borrmann causes 
 the tar to flow in continuous finely divided jets down a column 
 against a counter-current of superheated steam, in the presence 
 of which the high-boiling oils evaporate at a comparatively low 
 temperature. The resultant composition of the vapors of course 
 depends upon the ratio of the vapor tensions of water and the 
 oils at any temperature. 
 
 The tar in Sadewasser's process is highly heated in a direct- 
 heated tubular superheater, superheated steam being used as 
 the heating medium; the heated tar from the superheater is 
 forced through a spray nozzle into a vacuum vessel, entering 
 the latter in a finely atomized state, thus causing the tar parti- 
 cles to volatilize immediately. Liquid pitch is first drawn off, 
 followed upon cooling by the crude heavy and middle oils, and 
 finally by the light oil and ammonia water. 
 
TAR 41 
 
 In the system devised by Raschig three stills are employed, 
 the first two being heated by steam and the third by hot water, 
 but neither the steam nor the hot water comes in contact with 
 the tar to be distilled. In the first still the ammonia water and 
 light oils are driven off under atmospheric pressure, while the 
 middle oil is driven off in the second still under a vacuum. The 
 third still, also under vacuum, volatilizes the constituents form- 
 ing the heavy oil, and the liquid pitch runs off from this still 
 in a continuous stream. 
 
 The demand for coal-tar products, to be used in the pro- 
 duction of explosives, dyes, perfumes, and drugs, has been brought 
 home with ever increasing force to all gas works and coke- 
 oven operators, and in order that these operators may consider 
 the immensity of the proposition, and that every faculty must 
 be strained to the utmost at the present moment to give the 
 Nation what it requires, some of these finished products are 
 briefly described below. 
 
 The total consumption of artificial dyes in the world attained 
 a value of $92,000,000 in 1913, the United States consuming 
 about $12,000,000 worth of the total. Of this latter amount 
 about $3,000,000 worth was produced domestically, chiefly, 
 however, from imported raw materials; $7,850,000 worth was 
 imported from Germany; $910,000 worth from Switzerland, and 
 $370,000 worth from Great Britain and France. 
 
 The outbreak of European hostilities put a stop to these im- 
 ports, and we were consequently thrown upon our own resources 
 to supply this much needed commodity, in many cases being 
 compelled to go back to mineral, vegetable, and wood dyes, 
 the production of which had been greatly reduced after the ad- 
 vent of the aniline product, and this need has therefore given 
 a tremendous stimulus to the coal-tar chemical industry. 
 
 As stated above, this industry does not only include the manu- 
 facture of dye-stuffs, but also quite a number of very valuable 
 drugs and medicines, besides the high-powered explosives. 
 
 The industry as a whole is based upon the use of about 10 
 crude compounds which are removed from the original tar in 
 combination with some 150 other substances, being separated 
 from the latter by fractional distillation. These 10 crude com- 
 pounds produce about 300 very complex substances from which 
 about 900 dyes, besides the drugs, explosives, and perfumes are 
 
42 COAL GAS RESIDUALS 
 
 made by means of refined and complicated mechanical and 
 chemical processes. 
 
 We have an abundance of raw material in the form of tar 
 available, but a rapid development of this industry cannot be 
 expected unless an effective law protecting the home manu- 
 facturer against ruinous competition by foreign makers is given 
 by Congress. 
 
 The advent of aniline dyes has been of the greatest value to 
 our cotton manufacturer, for while silk and wool yield readily 
 to vegetable dyes, the inert character of the cotton fiber to 
 assimilating the colors of vegetable origin, and the difficulty 
 presented in mordanting cotton with metallic salts, caused our 
 cotton fabrics to be limited to but a few colors. The aniline 
 color dyes directly, without a mordanting substance on the cotton 
 fabric, and has thus led to the present enormous development 
 of colored cottons. 
 
 The 10 crude substances mentioned above are benzene, 
 toluene, xylene, phenol, cresol, naphthalene, anthracene, methyl 
 anthracene, phenanthrene, and carbazol, while the principal 
 "intermediates" are aniline oil, aniline salts, pure aniline, tolui- 
 dine, nitrobenzol, naphthol, phthalic acid, salicylic acid, resorcin, 
 anthraquinone, and many others up to the number of about 300. 
 
 Benzene and its homologues, toluene and xylene, besides 
 being fractionated from tar, may be washed directly from the 
 gas, and the same also holds good for naphthalene. The frac- 
 tionation process as explained above produces four principal 
 substances, namely Light Oil, Middle Oil, Heavy Oil, and Pitch, 
 the distillation of 100 tons of bituminous coal giving on an 
 average of 34 gallons of light, 205 gallons of middle, 100 gallons 
 of heavy oils, and 8400 pounds of pitch. 
 
 The 34 gallons of light oil give 12J gallons of benzol, 186 
 pounds of naphtha and residues; the benzol in turn producing 5 
 gallons of benzene, 1J gallons of toluene, and 6 gallons of 
 residues. 
 
 The 205 gallons of middle oil give 1512 pounds of fuel oil, 522 
 pounds of naphthalene salt, which in turn give 261 pounds of 
 naphthalene, and 261 pounds of residues, and 168 pounds of 
 crude phenols, the latter producing 37 pounds of carbolic acid, 
 56 pounds of cresylic acid, and 74 pounds of residues and 
 losses. 
 
TAR 43 
 
 The 100 gallons of heavy oils produce 928 pounds of green oil, 
 and 186 pounds of anthracene cake, the latter giving 19 pounds 
 of anthracene and 168 pounds of residues and losses. 
 
 The transformation of the ten crude coal-tar products into the 
 "intermediates," is based on eleven processes, or 
 
 1. Nitration: This process requires that an aromatic com- 
 pound shall be treated with a mixture of nitric and sulphuric 
 acids, this treatment usually resulting in one, and sometimes 
 two or three, atoms of hydrogen being replaced by the nitro 
 group N0 2 . The factors of time, temperatures, and propor- 
 tion of mixture largely effect the extent and nature of this 
 reaction, and results, besides the main product, in producing 
 a residual in the form of dilute sulphuric acid which can be 
 concentrated and used over again. 
 
 2. Chlorination: Chlorination is dependent upon the action of 
 dry chlorine gas, but this reaction is not so readily controlled as 
 is nitration, the end result often producing quite a variety of 
 substitution products which are very difficult to separate. The 
 residual in this process consists in the recovery of approximately 
 one-half the chlorine used in the shape of hydrochloric acid. 
 
 3. Sulphonation: This process is dependent upon the action of 
 fuming sulphuric acid, a large excess being usually employed; 
 while this reaction is usually more readily controlled, it often 
 results in the production of a variety of sulpho-acids which re- 
 quire separation, and uses for the discarded products is often 
 a problem. The excess sulphuric acid is lost. 
 
 4. Reduction: Reduction consists of substituting hydrogen 
 for oxygen in the nitro-compounds, thus forming the correspond- 
 ing amido-bodies. For instance, benzol is changed into nitro- 
 benzol by nitration, this latter product being then reduced to 
 amidobenzol, or aniline, by the reduction process. The usual 
 reducing agents consist of iron borings and filings, in conjunc- 
 tion with sulphuric or acetic acid. 
 
 5. Oxidation: The usual reagents employed in this process 
 are chlorate or bichromate of potassium, lead peroxide, manga- 
 nese dioxide or a permanganate, in conjunction with sulphuric 
 or hydrochloric acid. Useful residues in the shape of potassium 
 chloride, chrome alum, and several others usually result. 
 
 6. Fusion: The operation of caustic fusion is usually per- 
 formed with sulpho-acids, resulting in replacing the sulpho- 
 
44 COAL GAS RESIDUALS 
 
 group by hydroxyl. As an example, when benzol monosul- 
 phonic acid is fused with caustic soda, it results in the production 
 of phenol. This operation is quite difficult, and requires great 
 care, the results secured often being quite variable. 
 
 7. Alkylation: Alkylation consists of introducing the radicals 
 methyl or ethyl, present in both grain and wood alcohols, into 
 the hydroxyl- or amido-groups. This operation is conducted at 
 elevated temperatures and pressure in autoclaves, hydrochloric 
 acid in conjunction with the alcohols, or methyl or ethyl chloride, 
 being used as reagents. 
 
 8. Liming: Lime or chalk is used to effect separations in 
 mixtures through the varying solubilities of the calcium salts, 
 this holding especially good for the sulpho-acids. It is also used 
 to decompose chlorides and thus to effect the separation of 
 resultant acids and aldehydes. 
 
 9. Condensation: This process embraces a large number of 
 operations in which two molecules of different substances, or 
 of the same substance, unite to form a new compound, with the 
 consequent elimination of ammonia, water, or hydrochloric 
 acid, sulphuric acid in excess being the usual condensing agent. 
 Chlorides of zinc, antimony, sulphur, aluminum, and phosphorus 
 are also used as condensing agents. 
 
 10. Carboxylation: This process consists of introducing the 
 acid carboxyl group, by the combined action of caustic soda and 
 carbon dioxide on phenol and its homologues. As an example, 
 the carboxylation of phenol produces salicylic acid. 
 
 11. Diazotizing and Coupling: The diazotizing process is 
 based on the fact that an aromatic amine will react with nitrous 
 acid and produce a diazo-compound, while the coupling opera- 
 tion requires that the diazo-compound in the presence usually of 
 sodium acetate will readily unite with quite a number of aro- 
 matic substances. This reaction splits up a product and yields 
 the original amine and the amido-derivative of the second 
 substance employed. As an example, salicylic acid when coupled 
 with a diazo-compound, and then subjected to reduction, will 
 be changed into amido-salicylic acid. 
 
 In the above eleven processes I have attempted to explain 
 the chemical transformation required to produce the inter- 
 mediates from the crudes, but besides these principal chemical 
 changes there are many minor ones, as well as quite a number of 
 
TAR 45 
 
 mechanical operations, such as precipitating, filtering, baking, 
 boiling, and blowing off. 
 
 These processes result in producing 17 separate chemical 
 classes, from which 921 dyes are produced, and it may be well 
 now to refer to some concrete example of dye production. 
 
 Of all the coloring matters employed by man it is an assured 
 fact that indigo is probably the oldest and most valuable, and it 
 is recorded that the blue ribbons, removed from the wrappings 
 of mummies over 5000 years old, were colored with indigo. 
 Indigo, as formerly employed, was secured from indigo plants, 
 and many India merchants made fortunes in days gone by from 
 its importation, but the synthetic production of this product 
 came on the market in 1897, and it has practically displaced the 
 plant product entirely. Prof. Adolph von Bayer, of Munich, 
 was the leading spirit and master mind in its final production, 
 and he and his students labored on this subject for almost half 
 a century. Before Bayer began his monumental task it was 
 known that the oxidation of indigo would convert it into isatin, 
 and that the distillation of indigo with potassium hydroxide 
 would produce aniline; also that fusion with the same alkali 
 would produce anthranilic acid, and that nitric acid would 
 produce nitrosalicylic and picric acids. 
 
 All of these products resulted from the disintegration of the 
 indigo molecule, and this led to positive knowledge that indigo 
 blue contained a benzene ring, and that both carbon and nitro- 
 gen in neighboring positions to each other were attached to this 
 ring. Isatin then being a product of the oxidation of indigo, 
 Bayer reasoned that by a reverse process of reduction he could 
 cause the indigo molecule to form again. His original attempts 
 proved fruitless, but during the procedure he discovered other 
 substances, such as indol, oxindol, and dioxindol. He firmly 
 believed that indol was a close relative to indigo, and thought 
 that by starting with the substance which had given him indol 
 he could form indigo blue by some modification in his former 
 process. 
 
 This modification consisted in an attempt to secure the indigo 
 molecule through isatin instead of indol as the intermediate step, 
 and although he failed to secure indigo blue by simple reduction 
 from isatin, he was successful with the use of acetyl chloride, 
 phosphorus trichloride, and phosphorus. A great many inter- 
 
46 COAL GAS RESIDUALS 
 
 mediate steps were followed, and finally after 15 years of research 
 he had accomplished one of the greatest triumphs in synthetic 
 chemistry. 
 
 Since Bayer's original discovery, four different methods have 
 been devised for producing indigo, three of these starting with 
 benzene, and one with naphthalene. The usual procedure with 
 benzene consists in first nitrating the benzene to form nitroben- 
 zene, which is reduced to aniline; the aniline is then treated 
 with chloracetic acid to form phenyl glycine, this product being 
 then fused with sodium oxide to form indoxyl, which is finally 
 blown out with an air blast to form indigo. 
 
 If naphthalene is used to start with, it is first necessary to 
 oxidize the naphthalene to phthalic anhydride, this product 
 being treated with ammonia to form phthalamid; by the action 
 of sodium hypochlorite on the phthalamid, anthranilic acid is 
 formed, which in turn is treated with chloracetic acid to form 
 phenyl glycine, and carboxylic acid; this is now fused with 
 caustic soda to form indoxyl, and the air blast again gives indigo. 
 
 The manufacture of magenta is due to benzene and toluene. 
 The toluene is nitrated to form o- and p-nitrotoluene, this 
 being reduced to o- and p-toluidine. The benzene is nitrated 
 to form nitrobenzene, and is then reduced to aniline; this 
 aniline is now treated with o- and p-toluidine, and o- and p- 
 nitrotoluene, or the toluene products, to form the magenta base; 
 the latter after being treated with hydrochloric acid forms 
 magenta dye. 
 
 Hydroquinone, so much used in photography as a developer, 
 starts with benzene, which is nitrated to form nitrobenzene, the 
 latter being reduced to aniline, oxidized to quinone, and then 
 reduced to hydroquinone. 
 
 In the manufacture of the high-powered explosives we find 
 the nitro-substitution products of the coal-tar derivatives used as 
 bursting charges for explosive projectiles, torpedoes, mines, 
 detonators, and primers. They are also used in low freezing or 
 non-freezing mining explosives, in plastic explosives, and in other 
 types for the purpose of imparting some desired characteristic. 
 
 The two high-powered explosives depending upon coal-tar 
 products most commonly used are picric acid and trinitrotoluol. 
 
 Picric acid, besides being an explosive, is also a yellow dye, 
 and is principally used on wools and silk. 
 
TAR 47 
 
 In the manufacture of picric acid, phenol or pure carbolic 
 acid is used as the base; an equivalent weight of concentrated 
 sulphuric acid is mixed with the phenol, and this mixture is 
 heated to slightly above 212° F., after which it is allowed to 
 cool slowly; after cooling it is dissolved in twice its weight of 
 water, and then slowly added to three times its weight of nitric 
 acid, the nitration being completed by heating after all fuming 
 has ceased. This final mixture is allowed to cool very slowly, 
 thus permitting the picric acid to crystallize out, after which it 
 is purified by washing with warm water, and then re-crystallized. 
 
 Picric acid, as an explosive proper, has its disadvantages, as, 
 due to its acid nature, it readily combines with the metals 
 with which it may be in contact, thus forming picrates which 
 are more sensative to percussion and friction, often causing a 
 shell containing the picric compound to explode by a jolt while 
 being transported. The melinite used by France, the lyddite 
 used by England, and Shimose powder used by Japan are all 
 picric-acid products. 
 
 Trinitrotoluene, another nitro-substitution product produced 
 from coal-tar derivatives, although much more difficult to manu- 
 facture, is by far a safer explosive to store and handle, and it 
 forms the base of the explosives used by Germany, although the 
 other countries are also using this compound to some extent at 
 present. 
 
 This explosive is produced by first converting the toluene into 
 a mononitro-compound, after which it is treated with nitric 
 acid at a slight increase in temperature, thus converting the 
 mononitro-compound into dinitrotoluene, after which the latter 
 is nitrated with concentrated nitric acid in a sulphuric acid 
 solution, the final product thus formed being trinitrotoluene. 
 This compound is also produced from orthonitrotoluene by 
 means of a sulphonating process similar to that used in the 
 production of picric acid, after which it is nitrated and crystal- 
 lized, the resultant crystals being then purified in alcohol; this 
 product is re-crystallized to the final compound, the alcohol 
 being removed in a centrifugal hydro-extractor. 
 
 The manufacture of synthetic perfumes, while still in its in- 
 fancy, has' nevertheless found a great deal accomplished in this 
 line, but much remains to be done. One of the most widely 
 used perfumes is vanillin, which is synthetically prepared and 
 
48 COAL GAS RESIDUALS 
 
 which is identical in composition with the vanillin extracted 
 from vanilla pods. Other perfumes coming under this head are 
 heliotropine, essence of bitter almonds and aubepine, these 
 perfumes also being identical in chemical composition with the 
 same perfumes obtained from plants. 
 
 A different category exists, however, covering perfumes which 
 chemically have nothing in common with the natural product, 
 artificial essence of violet being a prominent example of this 
 class, and the discovery of this compound opened a wonderful 
 path for the chemist in that he has found it possible to manu- 
 facture a perfume which has an odor identical with nature's 
 product, but which in its chemical composition differs entirely. 
 Another perfume of this class is artificial musk, this latter product 
 having almost entirely supplanted the natural article. Artificial 
 geraniums also come under this classification. 
 
 Light tar oils, with their content of benzene, are of prime 
 importance in the perfumer's art; the benzene is converted into 
 nitrobenzene, which in turn yields essence of mirbane with its 
 odor of bitter almonds. Toluene is convertible into benzoic 
 aldehyde, or essence of bitter almonds, identical with nature's 
 article found in peach kernels, cherries, and cherry laurel leaves. 
 This product can be converted into cinnamic aldehyde, the 
 latter being the chief constituent of Chinese or Ceylon essence 
 of cinnamon. It has a strong cinnamon odor, but by reduction 
 it yields cinnamic alcohol, or artificial hyacinth. 
 
 Other constituents of light oils, such as the cumenes and 
 xylenes, form the basis of the perfume manufacturers' art. 
 Para-cumene is identical with that in the essence of thyme, and 
 metaxylene is converted into artificial musk. 
 
 With the use of the middle oils, containing phenol, we are 
 able to produce artificial geranium, or phenyl oxide, by a direct 
 catalytic process, and the cresols, also by catalytic action, give 
 cresol oxides, also of geranium odor. By mixing phenols and 
 cresols various geranium odors are produced, but the process is 
 rather difficult. 
 
 B-naphthol, in conjunction with methyl and ethyl alcohol, by 
 a catalytic process yields yara-yara and neroline respectively, and 
 this forms the base for the so-called Eau-de-cologne. 
 
 Salicylic aldehyde, produced from phenol by the action of 
 potash and chloroform, also forms an important perfume, and 
 
TAR 49 
 
 methyl salicylate is the essence of wintergreen, while salicylic 
 aldehyde, when acted upon by acetic anhydride, gives cou- 
 marine, a remarkably soft perfume. 
 
 The mention of the above will demonstrate the importance 
 of coal-tar derivatives for manufacturing perfumes, but a great 
 deal remains to be done in order to place this industry on the 
 same plane with the manufacture of coal-tar dyes. 
 
 In the manufacture of drugs, starting with benzene, we find 
 that the latter yields phenol, diphenol, chloronitro-benzene, and 
 aniline. The phenol in turn yields salicylic acid, which gives us 
 salol, the salicylates, methyl salicylate, and aspirin; parani- 
 trophenol, which gives phenacetine; orthonitrophenol, which 
 gives guaiacol and thiocol; oxyphenyl-arsenic acid, which gives 
 606 or slavarsan; chlorophenol, which also gives guaiacol. The 
 diphenols produce resorcin and pyrocatechine, the latter giving 
 guaiacol. The aniline produces phenylhydrazine, which in turn 
 gives antipyrin and pyramidon; from aniline we also secure 
 piperazine, acetanilide, exalgin, and atoxyle, the latter also giving 
 606 (arseno-benzol), and hectine. Then diethyaniline, which 
 gives novocaine, and dimethylaniline, which gives stovaine and 
 alypine. 
 
 Toluene is converted into benzoic acid and benzaldehyde. 
 The benzoic acid produces eucaine, stovaine, phthaleine, ami- 
 dobenzoic acid, which in turn gives novocaine and orthoform; 
 benzonaphthol, and cryogenine. The benzaldehyde gives cin- 
 namic acid. 
 
 Naphthalene gives us benzonaphthol, naphthol, and betol. 
 
 The manufacture of synthetic phenol, or pure carbolic acid, 
 is being practised in connection with several gas or coke works, 
 and this enterprise became of immense value when imports were 
 stopped early in 1915. In 1913 the United States imported 
 4077 tons of this valuable article, and its synthetic production 
 therefore became an absolute necessity, as the stock on hand soon 
 became exhausted. 
 
 The manufacture of synthetic phenol embraces five separate 
 and distinct processes, the benzene required being secured from 
 coal-tar distillation or from a direct washing of the gas. 
 
 These processes are: 
 
 1st. — The sulphonation of Benzene to Monosulphonic acid, 
 or Benzene Sulphonic Acid. 
 
50 COAL GAS RESIDUALS 
 
 2d. — The conversion of the Benzene Sulphonic Acid into a 
 Calcium Salt Solution. 
 
 3d. — The conversion of the Calcium Salt Solution into a 
 Sodium Salt Solution. 
 
 4th. — Fusing the dried Sodium Salt with Caustic Soda, 
 producing Sodium Phenolate. 
 
 5th. — Decomposing the Sodium Phenolate with a mineral 
 acid to liberate the Phenol. 
 
 In the operation of a plant for the production of pure phenol, 
 the benzol must first be purified by distillation and rectification, 
 the pure benzene being then admitted into the sulphonating 
 kettle (D, Fig. 9c), where it is treated with concentrated sul- 
 phuric acid from tank (E); fuming sulphuric acid cannot be 
 used, as it is liable to change the character of some of the 
 benzene, and also change some of the resultant monosulphonic 
 acid into a mixture of meta- and paradisulphonic acids. 
 
 This resultant monosulphonic acid is next processed in the 
 lime treater (F) with hydrated lime to form calcium sulphonate, 
 the excess acid being rendered inert by excess lime; the hydrated 
 lime is prepared in the vessel (H) and forced into the measur- 
 ing eggs (G), passing from there into the lime treater (F). 
 
 The calcium sulphonate is now forced through the filter 
 press (K), where the insoluble sulphate is retained and the 
 filtrate run into the soda tank (L); sodium carbonate is dis- 
 solved in the dissolver (M) and admitted into the vessel (L), 
 where it combines with the filtrate from the press (K) to form 
 sodium sulphonate. 
 
 This liquor is now forced through the filter press (N), where 
 the insoluble calcium carbonate is retained and the sodium 
 sulphonate run into the concentrator (0); after concentration 
 the liquor is admitted into (P), where it is crystallized, the 
 crystals being ground fine in the mill (Q), after which they are 
 ready for the fusion process. 
 
 The dried sodium salt is fused with caustic soda in the auto- 
 clave (R) in order to produce sodium phenolate, this melt being 
 dissolved in (S), the resulting mixture being then treated in- the 
 acidifying tank (T), where sulphuric acid is added in order to 
 decompose the phenolate, thus producing phenol with sodium 
 sulphate as a waste product. 
 
 This latter mixture is now admitted into the separator (U), 
 
SYNTHETIC F 
 (CARBOLIC 
 
 A Crude Benzol 5roroqe 
 
 B Ben-zol SMI 
 
 B' Rectifying Column 
 
 B 7 Deph/egmafor 
 
 B 3 Condenser 
 
 C Purified £>enzol Storage 
 
 D Sulphonahnq Kettle 
 
 D' Condenser 
 
 E Sulphuric Acid Storage 
 
 Lime. Treater 
 
 Measuring Eqqs 
 
 Lime Tank 
 
 Pumps 
 
 Filter Press for Calcium Sulpha* 
 
 Spda Tank 
 M Sodium Carbonate Dissolve* ' 
 
 Ca(C i H s SoJz + Mzj Co 3 - Z Na C t H t S0 3 + Co CO t 
 K N_ 
 
 •<>"< 
 
 mm 
 
 lHa} 
 ZA4»1 
 
 CM, 
 
 0)'2HC t H s 5O } i-Ca(oHl - Co(C L H s SO^ + lty 
 C2)-H t SO^i-Ca(OH) z 
 
 Ca 5o 4 + Z H 2 0. 
 
 \ Sulphuric ffpd. J 
 
 Fig. 9c— SYNTHEB; 
 
:nol plant 
 cid) 
 
 N Filhr Press for Calcium Carbonate 
 
 O Soda Soluhon Concent-rater 
 
 P Crystollicer 
 
 O Burr Mill 
 
 R Auhoclove 
 
 S Pissolver 
 
 T Acidifying Tank 
 
 U Separator 
 
 V Washer 
 
 W Filter Press for Sodium Sulphate 
 
 X Carbolic Sfi/I 
 
 X' Shi/ Condenser 
 
 Y Carbolic Distillate Sroraoe 
 
 Z Carbolic Distillate Concentrator 
 
 Z' Carbolic Crystoll/cer 
 
 Z* Carbolic Crystal Drying Table 
 
 H,Ot tt,S0 4 " 2C t rl s OHi-Ha J 50+ 
 't H 2 90 4 • Na t SO4+2 HjO 
 
 "Na C t H s SO, t ZNoOH = NaC & r/ s O t A/a z SOj-f r/ 3 I y i j 
 
 PHENOL PLANT 
 
 (Facing Page 50) 
 
TAR 51 
 
 the phenol, as well as the nitrate from the sodium sulphate 
 press (11'), being next run into the washer (F), the insoluble 
 sodium sulphate being retained in the press while the washed 
 solution enters a second separator ([/) and passes from thence 
 into the carbolic still (X). 
 
 The vapors from the still (X) are condensed in the twin 
 condensers (X 1 ), and run into the storage tank (}"); the weak 
 distillates from tank (F) are forced into the double effect con- 
 centrator (Z), the concentrated solution being run into the 
 crystallizer (Z 1 ) and, after being drained off, the purified crys- 
 tals are placed on the drying table (Z 2 ), after which they are 
 ready for packing. 
 
 From this it is seen that the raw material required to pro- 
 duce synthetic phenol is Benzol, Sulphuric Acid, Caustic Soda, 
 Soda Ash, and Lime, and the amounts required to produce one 
 ton of phenol are approximately: 
 
 Benzol 340 gallons 
 
 98 % Sulphuric acid 4 tons 
 
 Caustic soda 1.5 " 
 
 Soda ash (sodium carbonate) 1 " 
 
 Slaked lime 2 
 
 During the process a sufficient amount of calcium carbonate 
 is produced as a by-product to reduce the amount of lime given 
 by about one-half. The other by-products recovered are about 
 3 tons of calcium sulphate (gypsum) cake, and about 3 tons of 
 mixed sodium sulphite and sodium sulphate. 
 
CHAPTER II 
 NAPHTHALENE 
 
 Naphthalene (CioH 8 ) is a hydrocarbon and has a melting 
 point of 174° F., a boiling point of 424° F., and it sublimes at 
 lower temperatures. The deposition of naphthalene in a solid 
 state in the mains or apparatus causes many operating diffi- 
 culties, as it decreases the cross-sectional area of the gas 
 conduits and thus produces considerable back pressure on 
 the works. • 
 
 The presence of naphthalene in the gas produced during the 
 carbonization of coal, is undoubtedly due to the condensation 
 or polymerization of paraffin hydrocarbons, particularly those of 
 the Olefine or Ethylene series of the general formula C„H 2re . 
 
 The synthesis of naphthalene, which is an aromatic hydro- 
 carbon, from a paraffin hydrocarbon has been accomplished in 
 laboratory investigation, and the similar synthesis of benzene 
 and toluene by the Rittman process seems to give a further 
 proof of such changes in constitution. 
 
 The fact that naphthalene is produced in varying amounts 
 in both carbureted water gas and in oil gas, dependent upon the 
 operating conditions, tends to substantiate its paraffin origin. 
 
 Naphthalene formation is influenced by the land of coal used, 
 the type of carbonizing chamber employed, and the degree of 
 superheat to which the gas is subjected on exiting from the 
 charge. In the older type of benches, viz., horizontal and 
 inclined retorts, the gas is in contact with a considerable heated 
 surface, and while the production of naphthalene may be re- 
 duced by working with heavy charges, its total elimination is 
 impossible. 
 
 With the vertical retort systems, particularly those of the 
 continuous type, this superheating effect is to a very great ex- 
 tent eliminated, since the gas passes through the core of un- 
 carbonized coal, and little or no naphthalene is produced. The 
 tar produced in these retorts is distinctive, as it contains paraffin 
 
NAPHTHALENE 
 
 53 
 
 bodies significant particularly with the absence of naphthalene 
 in the gas. 
 
 Attempts have been made to produce records showing the 
 varying constitution of the gas in connection with high or low 
 naphthalene content, but these are of little value, as at best 
 they are only comparative but not absolute. In the same plant, 
 any change in the operation of the retort house would naturally 
 affect the quality of the gas, and it is therefore doubtful if 
 adequate comparison on the basis of gas quality should be 
 made between the gas from several retort systems in determin- 
 ing the amount of naphthalene that should be present. 
 
 Coal gas has a definite saturation for naphthalene, or in other 
 words at a certain temperature it can carry a certain quantity 
 of this impurity, and consequently if naphthalene is produced in 
 the retort house, it will always be found in the gas. 
 
 The greater portion of the naphthalene produced, especially 
 with high distillation or carbonization temperatures, goes 
 over into the tar, the gas containing only a portion of the 
 naphthalene vapor. 
 
 The amount of naphthalene required to saturate a gas 
 varies according to different investigators, and Table IX 
 
 TABLE IX. — NAPHTHALENE SATURATION 
 
 Temperature 
 
 Grains per 100 cubic feet of 
 gas 
 
 Centigrade 
 
 Fahrenheit 
 
 Schlumberger 
 
 Allen 
 
 0° 
 
 32° 
 
 2.0 
 
 6.0 
 
 5° 
 
 41° 
 
 3.2 
 
 9.8 
 
 10° 
 
 50° 
 
 6.7 
 
 14.1 
 
 15° 
 
 59° 
 
 10.9 
 
 19.0 
 
 20° 
 
 68° 
 
 16.5 
 
 24.6 
 
 25° 
 
 77° 
 
 24.7 
 
 30.9 
 
 gives this saturation according to both Schlumberger and 
 Allen. 
 
 The results of these two independent investigations vary 
 considerably, but it is readily seen that a gas saturated with 
 naphthalene must give up the latter with falling temperature. 
 
54 COAL GAS RESIDUALS 
 
 For higher temperatures than those given in the table, 
 the amount of naphthalene in saturated gas averages per 
 100 cubic feet at 
 
 90° F., 71.00 grains 100° F., 116.97 grains 
 110° F., 182.10 " 120° F., 276.23 
 
 130° F., 404.32 " 140° F., 629.62 " 
 
 The committee appointed by the American Gas Institute, 
 1913, to examine into the naphthalene problem made this the 
 subject of exhaustive study, and the committee agrees with the 
 statement that the production of naphthalene is a function of 
 the temperature the gas assumes before it leaves the retort, 
 rather than that of carbonization, and that gas works with 
 ample condensing facilities are able to confine the naphathlene 
 stoppages to the plants. 
 
 Naphthalene troubles begin in summer, and they are ag- 
 gravated by light and irregular charging; a district once af- 
 fected will give trouble for some seasons afterward through the 
 gas mains storing up the naphthalene and the gas depositing it 
 elsewhere. Efforts were formerly made to reduce the trouble by 
 charging large percentages of cannel coal, and later by carboniz- 
 ing petroleum in special retorts, although when the heats in- 
 creased, the latter method caused the formation of more 
 naphthalene. It was also found that the introduction of water 
 gas in many cases made no differences in the deposits. 
 
 A list of 29 questions was sent to 16 different plants asking 
 for particulars as to details of equipment, method of operation, 
 and naphthalene troubles. From the replies received it appears 
 that naphthalene is produced at all the works in question; the 
 heats ran from 1850° to 2200° F., except at one coke-oven plant, 
 where the rich gas during the first part of the carbonizing period 
 leaves the ovens at 1200° F., and it is free from naphthalene, 
 while the lean gas produced during the latter period of carboniz- 
 ing leaves the ovens at 1800° F. and contains naphthalene in 
 considerable quantities. 
 
 Two plants, a vertical retort system and a coke-oven re- 
 spectively, produced a crude gas with a remarkably low naphtha- 
 lene content, and hence depended entirely upon the tar, which 
 was also low in naphthalene, to absorb it. 
 
 Conditions governing the hydraulic mains and the tempera- 
 
NAPHTHALENE 55 
 
 turns in them, as well as the character of seals, dip pipes, etc., 
 were also examined into, but without result. The condensing 
 apparatus in the works answering the questions is of two types — 
 the multitubular water-cooled condenser, and the cataract 
 washer cooler. Either of these was capable of doing the re- 
 quired work, but the latter has the advantage over the other in 
 being self-purging. 
 
 Three of the works reporting employ naphthalene extractors, 
 two of the grid scrubber type and one a rotary washer. The 
 solvent in all cases was water-gas tar, the amount used being 
 0.14 and 0.25 gallon per 1000 cubic feet of gas in two of the 
 works respectively, and 0.05 for inclines and 0.072 gallon for hor- 
 izontals in the third plant, the wash oils containing varying 
 percentages of naphthalene up to 11.29. It was found that 
 every water-gas tar is not suitable for the purpose, especially 
 the tar produced at high temperatures and which is apt to be 
 saturated with naphthalene, and therefore incapable of further 
 absorption. Other tests indicated that absorption with anthra- 
 cene was better at low temperatures than at high, and the same 
 holds good for any solvent or absorbing medium. 
 
 Several of the works made use of steam and gasoline for 
 clearing stoppages, but this is not good practice; steam will 
 melt naphthalene, and gasoline will dissolve it, but the resulting 
 liquid must be immediately removed, otherwise it will solidify 
 and become so hard that nothing but a steel bar can touch it. 
 A copious stream of boiling water is preferable to steam, and a 
 heavy oil of low vapor tension to gasoline. 
 
 A study of the conditions covering the works in question led 
 to the statement that in order to prevent naphthalene troubles, 
 a system of carbonization should be selected in which the gas is 
 exposed to as little secondary heating as possible; the gas, if 
 possible, should be cooled in the works to a temperature as low 
 or lower than it will meet outside, and then finish the gas 
 by passing it through a naphthalene washer. By exercising 
 extreme care, but little naphthalene will escape beyond the works, 
 and a few grains can be taken care of in a mixed coal and water 
 gas, or where benzol is used as an enricher. Where straight 
 coal gas is made, a condenser at the outlet of the works might 
 be considered with a view to removing moisture, in the absence 
 of which naphthalene does not precipitate. 
 
56 COAL GAS RESIDUALS 
 
 Another method for obviating this trouble is to introduce 
 into the gas a sufficient quantity of the vapor of oils having a 
 somewhat lower boiling point than that of naphthalene, and this 
 method results also in a gain in candle-power and calorific value, 
 besides requiring little attention and is fairly certain in its 
 action. The drawback inherent in this method, however, is 
 found in the pernicious action of the solvent deposited by con- 
 densation on the paint of gas holders and also occasional trouble 
 with pipe joints, meters, and fittings due to condensation of 
 the oil. Dr. W. B. Davidson states that the improvement of 
 the quality of the gas treated by the addition of oil vapors is 
 not always conceded, and that proper allowance should be made 
 not only for candle and calorific value due to the oil vapors, 
 but also for the increase in volume. 
 
 The solvents most commonly used for this purpose are solvent 
 and heavy naphtha, crude heavy naphtha, condensed liquid 
 secured during the compression of oil gas, light paraffin oil, and 
 kerosene. Of these solvents, those derived from coal-tar or 
 carbureted water-gas tar are probably the most efficient; if 
 the condensate due to oil gas compression is used, an additional 
 advantage is secured from the enriching quality of this product. 
 If paraffin oils are used, they must be added to the gas in the 
 shape of a mist, or in such a finely divided state that there will 
 be no question of this mist being mechanically carried throughout 
 the entire system, a difficult performance. 
 
 Present prospects point to the necessity of removing benzol 
 and its homologues from the gas, as they are very necessary for 
 the manufacture of explosives, and as these constituents become 
 less, the more trouble can be expected from naphthalene; while 
 the removal of benzol and toluol in themselves would not have 
 a material effect on the absorption of naphthalene, the removal 
 of these two constituents also effects the removal of higher 
 boiling vapors, thus reducing the amount of solvents in the gas 
 capable of taking up the naphthalene. In such cases where 
 benzol is not extracted, from 10 to 15 gallons of solvent per 
 million cubic feet will generally be sufficient to relieve the trouble, 
 but in the opposite case the amount of solvent must be increased 
 up to about 20 gallons. 
 
 The removal of the naphthalene at the works, or the addition 
 of a solvent, will not entirely remove all troubles, providing the 
 
NAPHTHALENE • 57 
 
 distribution system contains naphthalene deposits, as although 
 the pas may leave the works practically free of naphthalene, or 
 if it contains sufficient solvent to prevent the naphthalene from 
 crystallizing out, the gas will take up naphthalene from that in 
 the mains. If this gas, carrying an increased burden of naphtha- 
 lene, reaches a colder main, it will again deposit the excess of 
 naphthalene beyond that which the solvents can hold, and this 
 condition will continue until the mains have been cleaned out. 
 These local deposits can be removed by spraying a solvent into 
 the mains in the district affected. 
 
 Spraying a solvent into the gas must be done with care, and 
 a spray must be selected which will produce a mist, and not 
 large drops, as in the latter case it will not be carried with the 
 gas for a distance sufficiently far to be effective. Care must 
 also be exercised not to spray the solvent into a main near a 
 sharp bend, as the impingement of the gas on the surface of the 
 pipe at such points will act to strip the mist out of the gas. 
 
 If the solvent is to be vaporized, and not sprayed into the gas, 
 the solvent is usually treated in an apparatus which embodies a 
 large vaporizing plate, heated by steam, the solvent flowing 
 over this heated pla.tc in a thin stream. About 10 per cent of 
 the gas to be enriched is by-passed and shunted through the 
 vaporizer, the vapors of the solvent joining the gas in its passage, 
 the mixture being then returned to the gas main in such manner 
 that the flow of the mixture is in the direction of the gas flow, 
 thus thoroughly mixing with the main volume of the gas and 
 passing on with it. 
 
 If possible, the solvent should be given to the gas at a point 
 between the station meter and the holder, this point having 
 advantages over any others. The gas is practically constant 
 in volume at this point, and the amount of solvent given to the 
 gas need, therefore, not be varied during the day, while if it 
 were added after the holder outlet, where the flow is irregular, 
 due to varying consumption, the addition of the solvent would 
 have to undergo constant regulation, owing to the differences in 
 the volume consumed during various hours of the day. Be- 
 sides this, if the solvent is added ahead of the holder, the dep- 
 osition of naphthalene in the drips at the holder inlets and 
 outlets will be prevented. 
 
 If the finished gas is a mixture of coal gas and carbureted 
 
58 
 
 COAL GAS RESIDUALS 
 
 water gas, a less amount of solvent will be required than with 
 straight coal gas, because the carbureted water gas usually 
 contains a higher percentage of solvents than does coal gas of 
 the same quality. 
 
 Washing the gas under ordinary conditions, 100 grammes of 
 anthracene oil will absorb from 10 to 25 grammes of naphtha- 
 lene, according to temperature, or 100 grains of anthracene oil 
 will absorb from 10 to 25 grains of naphthalene, but before 
 using this oil from 3 to 4 per cent of benzol should be added, 
 this addition leading to greater extraction efficiency, this effi- 
 
 A Oil Still. 
 B Condenser. 
 
 C COLLECTOR. 
 
 D Settling Tank. 
 E Wash Oil Storage, 
 
 TtXlT 
 
 Fig. 10. — Oil Regeneration Plant. 
 
 ciency being further increased by thoroughly and slowly cooling 
 the gas, as the absorption is most complete at a temperature 
 of from 60° to 70° F. 
 
 Creosote oil, under ordinary conditions, has about the same ab- 
 sorption efficiency as anthracene oil, but in both cases the absorp- 
 tion will depend upon the amount of phenols present in the oil. 
 
 Water-gas tar will absorb from 18 to 20 per cent of its 
 own weight in naphthalene, this absorption efficiency also 
 being dependent upon the amount of phenols present in the gas. 
 
 In Germany it is usually estimated that anthracene oil 
 will absorb about 40 per cent of its own weight in naphtha- 
 lene, and as their gas usually contains about one gramme of 
 
NAPHTHALENE 
 
 59 
 
 naphthalene per cubic meter, or 43.6 grains per 100 cubic 
 feet, they use from 0.6 to 0.9 kilogram of anthracene oil 
 per 100 cubic meters of gas, or 3.75 to 5.6 pounds of oil 
 per 1000 cubic feet of gas. 
 
 If water-gas tar, or if water-gas oil is used as the absorb- 
 ing medium, it can be run back to storage bearing its naph- 
 thalene burden, as the latter is not detrimental to the further 
 use of the tar or oil, but if anthracene or creosote oil is used, 
 this oil must be regenerated in order to fit it for further use; 
 a diagram of a plant for regeneration is shown in Fig. 10. 
 
 Fig. 11. — Horizontal Type of "Standard" Washer. 
 
 The saturated naphthalene absorption oil is run into the 
 still (A), where it is distilled, the distillate being condensed 
 in the condenser (B), the oil then entering the collector (C), 
 from whence it flows into the settling tanks (D), where after 
 complete cooling of the oil the naphthalene crystallizes out. 
 The mother liquid is now free of naphthalene and it is run 
 into the storage tank (E), from whence it is pumped back 
 into the washers or it is stored for other disposition. 
 
 The removal of the naphthalene from the gas is usually 
 accomplished in some type of mechanical washer, either of 
 the Standard horizontal type, as shown in Fig. 11, or in 
 the Feld vertical centrifugal type, shown in Fig. 12. 
 
 The Standard rotary washer is constructed of one double 
 
60 
 
 COAL GAS RESIDUALS 
 
 chamber, three single chambers, and two or three double 
 chambers, according to whether anthracene or creosote oil 
 is to be used as a washing medium, the larger number of 
 ■washing chambers permitting greater unit absorption. The 
 wash oil enters the last chamber of the washer, where it 
 meets the cleaner gas, and after a certain degree of ab- 
 
 
 Fig. 12. — Feld Vertical Centrifugal Washer. 
 
 sorption has been obtained the oil is pumped into the next 
 forward chamber, and so on until it is saturated, thence 
 leaving from the gas-inlet chamber and flowing to the point 
 of disposition. The chambers are interconnected with 
 piping and cocks to a pump which receives motion direct 
 from the revolving shaft of the washer proper. 
 
NAPHTHALENE 
 
 61 
 
 In the Feld centrifugal washer (Fig. 12), the wash oil 
 enters the top section through a syphon, being then pumped 
 up on the inside of the spraying cones and hurled out through 
 the perforations in a very fine mist or spray, thus filling the 
 
 Fig. 13. — Vertical Type of "Standard" Washer. 
 
 entire gas space with a mist of wash oil; the oil after partial 
 saturation flows into the next lower section through the gas 
 ports, being again brought in contact with the gas in the 
 same manner here, finally leaving the bottom of the washer 
 in a saturated condition. 
 
62 
 
 COAL GAS RESIDUALS 
 
 The action of the Standard centrifugal washer (Fig. 13) 
 is similar to that of the Feld, but instead of revolving cones, 
 the vertical shaft is provided with a perforated basket, the 
 wash oil being picked up and sent to the basket by means 
 of the bent tubes attached to its bottom. 
 
 The vertical centrifugal washers are more efficient than 
 the slow-moving rotaries, because they do not depend upon 
 skin contact for absorption, the oil in the verticals completely 
 surrounding the naphthalene vapors and causing their ab- 
 sorption. The natural flow of the liquid being downward, 
 no pumps are required for transferring the liquid from 
 chamber to chamber, and as the inflow of wash oil can be 
 regulated, the oil may be kept in the washer for any required 
 period of saturation, the outflow of oil being entirely de- 
 pendent upon the quantity of oil admitted into the top 
 chamber of the washer. 
 
 The removal of naphthalene seems to be primarily a func- 
 tion of the gas temperature, as is seen from the following 
 data secured at one of the large coal-gas works. As orig- 
 inally designed, the gas arrived at the naphthalene washers 
 without any other cooling than that produced by the atmos- 
 phere around the overhead mains, this condition causing 
 the temperature of the gas at the naphthalene washers 
 to vary greatly with the season of the year and the 
 amount of gas produced, the extraction results being given 
 in Table X. 
 
 TABLE X. — NAPHTHALENE EXTRACTION 
 
 
 Grains naphthalene per 
 
 
 Gas tempera- 
 
 100 cu. ft. of gas 
 
 Per cent 
 
 ture F.° 
 
 
 removed 
 
 
 
 
 Washer inlet 
 
 Washer outlet 
 
 
 134 
 
 229 
 
 156 
 
 32 
 
 113 
 
 121 
 
 70 
 
 42 
 
 91 
 
 61 
 
 17 
 
 72 
 
 87 
 
 39 
 
 14 
 
 64 
 
 This condition led to a great deal of trouble in naphthalene 
 stoppages, principally in the condensers which immediately 
 
NAPHTHALENE 63 
 
 followed the naphthalene and cyanogen washers, especially 
 during the summer months, necessitating the installation of 
 primary condensers between the exhausters and the tar 
 extractors, the gas temperature being thereby maintained 
 at between 85° and 90° F.; no further trouble due to stop- 
 pages was experieneed since treating the gas at these 
 temperatures, water-gas tar being used as the washing 
 medium in the naphthalene washers with results as given 
 below : 
 
 Water gas tar used per ton of coal 0.25 to 0.38 gallon 
 
 Saturated tar produced per ton of coal 0.25 to 0.37 gallon 
 
 Per cent of naphthalene by weight in saturated tar. . . . 12.71 to 18.60% 
 
 Specific gravity of saturated tar at 60° F 1.0S6 to 1.0S2 
 
 Naphthalene per 100 cu. ft. of gas at wash"- inlet . .. 40.1 to 59.4 grains 
 Naphthalene per 100 cu. ft. of gas at washer outlet. . . . 23.3 to 21.1 grains 
 Naphthalene in water gas tar before using Only traces 
 
 An attempt has lately been made to use naphthalene as 
 fuel in internal combustion engines, and some success has 
 been met with in this direction; the present drawback 
 seems to be the fact that it is necessary to start the engine 
 with gas or some liquid fuel, and so to operate it until the 
 heat generated by the engine is sufficient to melt the naph- 
 thalene, after which the latter is fed to the engine, vaporized, 
 and exploded. Complete success in this direction would 
 soon open up another source of revenue to the producer of 
 coal gas. 
 
 It may be interesting to state that a locomotive operating 
 on naphthalene has recently been constructed in France; 
 this locomotive is provided with four cylinders, each two 
 cast en bloc, 5.5 inches bore by 8 inches stroke, 70 horse- 
 power being developed at 950 revolutions per minute. This 
 machine is provided with two carburetors, one used on start- 
 ing with spirit, and the other for the naphthalene. The car- 
 buretor used in connection with the naphthalene is cast 
 en bloc with the reservoir in which the naphthalene is melted, 
 double walls for the circulation of cooling water being pro- 
 vided, the temperature of this water being maintained at 
 212° F., thus maintaining a constant melting temperature. 
 
 Naphthalene engines have been built in Germany for some 
 
64 COAL GAS RESIDUALS 
 
 years, and these engines usually consist of a single, or at the 
 most, a twin cylinder four-cycle horizontal machine, following 
 in its principal lines the model of the Otto engine. Above or 
 alongside the cylinder with its cooling jacket, there is a double- 
 walled naphthalene fusing tank, from which a heated tube 
 leads to the carburetor through a float valve and a sprayer 
 inserted in the pipe supplying air to the cylinder. 
 
 The naphthalene fusing tank is heated by waste heat from 
 the engine, while in the earlier engines the naphthalene was 
 rapidly melted by the exhaust gases; this latter method was 
 attended with the disadvantage that overheating and too free 
 an evolution of vapor were liable to occur. The Deutz gas 
 engine works adopted the plan of melting the naphthalene by 
 the heat of the cooling water, and thus avoided overheating 
 the naphthalene. 
 
 In the older Deutz engines the cooling jacket of the engine 
 cylinder was connected with a large cooling water tank in which 
 the water after a time came to a boil, the naphthalene re- 
 ceptacle being suspended in this tank, and the naphthalene 
 being thus liquefied by the heat of the cooling water and the 
 escaping steam. The hot water surrounded all the adjacent 
 parts in order to prevent the melted naphthalene from solid- 
 ifying; but even so, it was found that the naphthalene 
 solidified in the sprayer unless the air supply had been previously 
 warmed, and the heat of the exhaust gases was therefore used 
 to raise the temperature of the air in the suction pipe to 250° 
 to 300° F. 
 
 The later Deutz engines are provided with a small water- 
 tube boiler, heated by the exhaust gases, and which boiler 
 communicates with the cooling jacket of the engine, and so 
 raises the temperature of the water to the boiling point more 
 quickly after the engine is started than is the case when the 
 cylinder heat alone is relied upon for the heating of the water. 
 The naphthalene is thus melted more quickly than in the old 
 type of engine. K. Bruhn reports the efficiency of naphthalene 
 as compared with other fuels as follows. 
 
NAPHTHALENE 
 
 65 
 
 
 Gasoline 
 
 Benzo 
 
 
 Naphthalene 
 
 Consumption 
 
 
 
 
 
 
 
 
 
 
 
 at 
 
 Consump- 
 
 Cost 
 
 Consump- 
 
 Cost 
 
 Consump- < Cost 
 
 
 tion oz. 
 
 i 
 
 tion oz. 
 
 i 
 
 tion oz. 
 
 t 
 
 Half load. . . . 
 
 16.6 
 
 4.32 
 
 14.1 
 
 2.88 
 
 13.4 
 
 1.10 
 
 Three Quarter 
 
 
 
 
 
 
 
 load 
 
 13.1 
 
 3.36 
 
 10.6 
 
 2.16 
 
 11.3 
 
 0.94 
 
 Full load 
 
 11.3 
 
 2.88 
 
 9.5 
 
 1.92 
 
 10.2 
 
 0.S4 
 
CHAPTER III 
 CYANOGEN 
 
 Cyanogen (C 2 N 2 ) is a gas composed of carbon and ni- 
 trogen, and it is probably produced during the period of 
 carbonization in the form of hydrocyanic acid by the decom- 
 position of some of the ammonia, due to contact with the 
 hot coke in the retort, as per the equation 
 
 C + NH 3 = HCN + H 2 , 
 
 and the amount of cyanogen produced by the carbonization 
 of any particular coal bears a certain relation to the amount 
 of nitrogen contained in the coal. 
 
 Cyanogen in its gaseous form is colorless and very poison- 
 ous. When in combination with potassium it forms potas- 
 sium cyanide, which has found its largest application in the 
 cyanide process of gold extraction, this process depending 
 upon the solubility of gold in a dilute solution of potassium 
 cyanide in the presence of air or of some other oxidizing 
 agent, and it is best adapted for use with free milling ores 
 after the bulk of the gold has been removed by amalgama- 
 tion, thus recovering such portions as were not taken up by 
 the amalgam, the gold being later usually separated from 
 the cyanide solution by an electrolytic process, or some 
 other metallic precipitation process. 
 
 Potassium ferrocyanide is the base of a great many cyano- 
 gen products, one of the principal ones of which is prussian 
 blue; this coloring matter is obtained by mixing the potas- 
 sium ferrocyanide with a solution of ferrous sulphate, a 
 precipitate of ferro-potassic-ferrocyanide, having a grayish- 
 white color, resulting, the mother liquor, which contains 
 potassium sulphate, being removed from the precipitate by 
 evaporation. After the precipitate has been allowed to settle, 
 it is washed with a large amount of water, and then oxidized, 
 this oxidation giving it the beautiful blue color required in 
 the dyeing industry. 
 
 The amount of nitrogen going to the formation of cyanogen 
 is problematical, and the tests as to the ultimate disposition 
 
CYANOGEN 
 
 67 
 
 of the nitrogen, according to the different authorities, is given 
 in Table XI, the coal used in carbonization containing from 
 1 to 2 per cent of nitrogen. 
 
 In these six tests Knublach shows the largest percentage, 
 1.80 per cent, as going to cyanogen formation, and Desmarets 
 the lowest, or 1.30 per cent. 
 
 TABLE XI. — DISTRIBUTION OF NITROGEN 
 
 Nitrogen 
 
 Foster 
 
 Leybolt 
 
 Knublach 
 
 Desmarets 
 
 
 % 
 49.90 
 14.50 
 
 1.56 
 
 | 34.04 
 
 % 
 
 33.75 
 
 12.00 
 
 1.75 
 
 1.50 
 
 51.00 
 
 % 
 
 30.00 
 
 11.90 
 
 1.80 
 
 1.30 
 
 55.00 
 
 % 
 
 35.60 
 
 14.10 
 
 1.80 
 
 1.40 
 
 47.10 
 
 % 
 
 63.90 
 
 11.60 
 
 1.80 
 
 1.30 
 
 21.10 
 
 % 
 33.00 
 
 In ammonia .... 
 In cyanogen .... 
 In tar 
 
 13.50 
 1.30 
 2.00 
 
 In gas 
 
 50.20 
 
 
 
 The removal of cyanogen from coal gas is of decided ad- 
 vantage for two reasons, viz.: it is an impurity which is very 
 objectionable, and it is a valuable by-product under certain 
 conditions. If cyanogen is permitted to pass on into the 
 purifiers it will be found to enter into combination with the 
 iron in the purifying material, thus rendering a certain portion 
 of this iron inactive for hydrogen sulphide extraction; on 
 the other hand, the sale of spent oxide containing a high 
 percentage of cyanogen is more remunerative than if it con- 
 tained only sulphur, as this oxide is often purchased on the 
 basis of its "prussian blue" contents only, "prussian blue" 
 being the name applied to this cyanogen combination. 
 
 In spite of the removal of a certain percentage of the 
 cyanogen in the purifiers (there are also times when it is 
 thus entirely removed), quite an appreciable portion passes 
 on with the gas, and this cyanogen has a considerable corro- 
 sive action on all iron or steel with which it comes in contact; 
 this is especially noticeable in the case of gas holders, where a 
 blue color is often seen on the outside of the holder sheets, 
 indicating that corrosive action is taking place. 
 
 Cyanogen passing on into the purifiers always produces an 
 ammonia loss amounting to from § to f grain per 100 cubic 
 feet if the oxide in the purifiers is to be kept in its most 
 active form. This small amount of ammonia should be 
 
68 COAL GAS RESIDUALS 
 
 retained in the gas in order that it may neutralize its* acid 
 qualities and thus maintain an alkaline state in the oxide, 
 because this alkaline condition is conducive to the formation 
 of ammonium sulphocyanide, due to a combination between 
 ammonia and sulphur, this compound being removed from 
 the boxes through the sealed drains, while if the gas should 
 retain its acid character the cyanide-compounds would com- 
 bine with the iron and form the double iron cyanide com- 
 pound known as "prussian blue"; this latter compound is 
 insoluble and it rapidly coats the oxide with a layer which is 
 then impervious to the action of sulphur. 
 
 Many attempts have been made to produce a combination 
 between the nitrogen carried in the cyanogen and hydrogen, 
 with the object of forming additional ammonia; the most 
 successful of these attempts seems to be the process developed 
 by Mr. Charles Carpenter, of England, and the liming process, 
 both of which claim an efficiency of 98 per cent. It must be 
 stated, however, that up to the present the author does not 
 know if the financial returns would warrant the application 
 of either of these methods, and their further development is 
 therefore awaited with great interest. 
 
 The revenue to be derived from the extraction of cyanogen 
 from coal gas is dependent upon the amount of coal carbon- 
 ized and upon the market price of the product, the latter 
 fluctuating constantly; it is certainly not a remunerative prop- 
 osition for works carbonizing less than 250 tons of coal per 
 day, as the same labor expended on a plant of this size can 
 readily handle the proposition in one of double the capacity. 
 
 The estimated revenue, prior to 1915, in a plant carbonizing 
 250 tons of coal at one works, and 750 tons at another, the 
 sludge from both being worked up in the one plant, after de- 
 ducting operating expenses and 6 per cent interest and 6 per 
 cent for depreciation, amounts to $27,935 per year, and this 
 without the additional revenue to be derived by the production 
 of ammonium sulphate from the cyanogen press liquor, this 
 liquor being sold on the basis of concentrated ammonia. 
 Besides this revenue, the extraction of cyanogen in this case 
 shows an added efficiency of 22 per cent in the action of the 
 purifiers, or the life of the oxide is increased by 22 per cent. 
 
 If the cyanogen plant is combined with the direct sulphate 
 
CYANOGEN 69 
 
 recovery plant, the revenue to be derived from a combined 
 plant carbonizing as low as 250 tons of coal is unquestionable, 
 as will be shown later. 
 
 The cyanogen process operated by the British Cyanides 
 Company consists in washing the gas with ammoniacal liquor 
 in combination with sulphur. In this case the ammonium 
 sulphide present in the liquor dissolves the sulphur to form a 
 polysulphide which reacts with the hydrocyanic acid in the 
 gas and forms ammonium sulphocyanide, or 
 
 (NH0»S + 2HCN = 2NH.,CNS + H 2 0. 
 
 This ammonium sulphocyanide is converted into either 
 potassium cyanide or potassium ferro-cyanide. 
 
 The British Cyanides Company, 1 in developing its process, 
 adopted a policy in contradistinction to that prevalent on the 
 European Continent, in that they were not averse to the forma- 
 tion of sulphocyanide of ammonia, but stimulated the pro- 
 duction of polysulphides of ammonia by washing the gas with 
 gas liquor in which granulated sulphur was kept agitated, before 
 it reached the ammonia scrubbers, and while it still contained 
 hydrogen sulphide and ammonia in large quantity. 
 
 Polysulphide of ammonia is an active absorbent of cyanogen, 
 and readily combines to form sulphocyanide of ammonia. The 
 system of operation proved very simple, and the liquor thus 
 obtained was sent to a chemical plant for conversion. This 
 system was installed in a British gas plant in 1905, a "Stand- 
 ard" washer-scrubber, fitted with Holmes brushes being used. 
 The sulphur was periodically admitted into the scrubber through 
 gas-tight cups, but trouble was soon experienced with tar 
 particles, from the gas, coating the sulphur, thus interfering 
 with the necessary chemical reactions by preventing the gas 
 from coming in contact with sulphur, and finally clogging the 
 washer. This washer was replaced by two cast-iron machines 
 designed for the purpose and consisting of one chamber above 
 a lower one, the upper one being provided with a vertical 
 rotating shaft carrying stirring paddles by means of which the 
 sulphur was kept in a thorough state of agitation, the liquor 
 overflowing from the upper into the lower chamber. The gas 
 entered the lower chamber and, through the medium of dip 
 1 A. E. Broadberry. "Jnl. of Gas Lighting." Oct. 1, 1912. 
 
70 COAL GAS RESIDUALS 
 
 pipes, bubbled through the liquor contained in it, passing thence 
 up a connecting pipe into the upper chamber where it again 
 bubbled through the liquor contained in same. The upper 
 chamber formed the polysulphide of ammonia, while the lower 
 one extracted tar particles, in order to keep them from reaching 
 the sulphur, and also dealt with any polysulphides which had 
 not absorbed cyanogen in the upper chamber. This machine 
 was successful except for the back pressure thrown, and occa- 
 sional tar troubles. 
 
 Mr. P. E. Williams, engineer of the Poplar Station of the 
 Commercial Gas Company, London, where Holmes washers 
 were used for this process, found it desirable to try some other 
 form of apparatus for the production of the polysulphides in 
 order to relieve the washer of this portion of the work, and in 
 1909 he decided to make use of a special design of purifier box 
 containing spent oxide, with about a 50 per cent sulphur con- 
 tent, kept in a moist condition by means of liquor sprays. The 
 gas entered ^his box at the bottom, and was caused to bubble 
 through the polysulphide liquor, kept at a given level in the 
 bottom of the box, before it reached the oxide. 
 
 The object in using spent oxide was due to the desire to avoid 
 purchasing sulphur, and it was anticipated that by combina- 
 tion with ammonium sulphide the sulphur contained in the 
 oxide would gradually be absorbed. This apparatus proved 
 very effective, and it was soon learned that it was not necessary 
 to transfer the polysulphides, produced in the purifier, to the 
 Holmes washer, as was originally intended, as the contact with 
 the gas in the purifier was so complete as to cause the nascent 
 polysulphides to combine with the hydrocyanic acid and thus 
 immediately to produce the desired sulphocyanide of ammonia. 
 The spent oxide was sprayed each day for about 20 minutes 
 with water or liquor in order to maintain its moist condition. 
 
 It was at first presumed that the spent oxide, owing to its 
 gradual loss of sulphur, would have to be renewed from time 
 to time, but upon making a test on the material it was found 
 that the sulphur content had slightly increased, due possibly 
 to a decomposition of the hydrogen sulphide in the gas, thus 
 not only adding its molecule of sulphur to the ammonium sul- 
 phide, but also leaving additional sulphur in the oxide. 
 
 Mr. P. E. Williams further reports that a reversal of the gas 
 
CYANOGEN 71 
 
 flow for a very short period is beneficial, this reversed or down- 
 current being for the purpose of allowing the supernatant 
 liquor to drain away from the oxide. At the Poplar Station a 
 purifier 20 feet square was used to treat 3? million cubic feet 
 of gas, one layer of oxide 24 inches deep being sufficient to re- 
 move all the cyanogen from this quantity of gas. 
 
 In this process from 10 to 15 per cent of hydrogen sulphide 
 and from 15 to 20 per cent of ammonia contained in the gas 
 are removed in the cyanide vessel, these quantities of course 
 varying with the amount of cyanogen in the gas. The ammo- 
 nium sulphocyanide recovered is between 3 and 5 pounds per 
 ton of coal, dependent upon the method and temperature of 
 carbonization, while the resultant liquor contains about 2.75 
 pounds of ammonium sulphocyanide per gallon, or about 26 per 
 cent. 
 
 Before leaving the subject of sulphocyanides, it is of im- 
 portance that the conversion of these cyanides into ammonia 
 be taken up, and, as stated before, the Carpenter process as 
 well as the liming process seem to give promise of the best results 
 in this direction. The patented process of Mr. Charles Carpen- 
 ter has not been explained in detail, but it is claimed it has been 
 worked with "extremely satisfactory yields, and appears to be 
 peculiarly adapted to all gas works and coke ovens which work 
 up their own sulphate of ammonia." 
 
 This process differs from the earlier attempts of Burgevin 
 and Burkheiser, and it appears to consist essentially of treating 
 the sulphocyanide with an excess of sulphuric acid at some 
 definite temperature, when ammonia, carbon oxysulphide, and 
 acid sulphate of ammonia result. 
 
 According to the equation 
 
 KCNS + H 2 S0 4 = HCNS + KHS0 4 
 
 the action of sulphuric acid upon potassium sulphocyanide will 
 produce sulphocyanic acid, and if the latter is decomposed by 
 water in the presence of an excess of sulphuric acid, ammonia 
 and carbon oxysulphide will result, or 
 
 HCNS + H 2 = NH 3 + COS. 
 
 These two reactions will be very similar in the case of am- 
 monium sulphocyanide, ammonium bisulphate resulting in the 
 first place, or 
 
72 COAL GAS RESIDUALS 
 
 , NH 4 SCN + H 2 S0 4 = HCNS + NH4HSO4, 
 
 and the second reaction, or 
 
 HCNS + H 2 = NH 3 + COS, 
 
 will again give ammonia. 
 
 This process will depend greatly upon a proper working 
 temperature, as if the temperature should be too high the 
 cyanogen would possibly be volatilized and in consequence 
 lost as sulphocyanic acid. It is stated, 1 however, that, working 
 under proper conditions, a conversion of 98 per cent of the total 
 nitrogen into ammonia should be possible. In addition to evolv- 
 ing carbon oxysulphide, it is very probable that some hydrogen 
 sulphide will also result from the reaction, this being possibly 
 due to the fact that carbon oxysulphide is readily absorbed by 
 ammonia, which on evaporation will yield hydrogen sulphide. 
 It is also possible that a portion of the carbon oxysulphide, in 
 the presence of water and steam, might split up and yield 
 carbon dioxide and hydrogen sulphide, or 
 
 COS + H 2 = C0 2 + H 2 S. 
 
 Great attention has been aroused by the application of this 
 process, and definite commercial results are therefore awaited 
 with interest. 
 
 The second recent method of conversion, or the Lime method, 
 is based on the lines laid down by Burgevin, and is applicable 
 to such cases where calcium sulphocyanides are made. The 
 calcium sulphocyanide is first evaporated to dryness, then 
 mixed with an excess of slaked lime, and finally subjected to 
 a heat of not less than 1000° F. The cyanides, in the presence 
 of steam, are broken up under this temperature, and yield 
 ammonia and carbon oxysulphide. The ammonia may be 
 drawn off to an absorption vessel containing sulphuric acid, 
 while the calcium oxysulphide will be in great measure removed 
 by the excess of lime. It is stated that the manner in which 
 the lime and cyanide are admixed is of great importance for the 
 proper operation of the process. 
 
 The principal reactions are about as follows: 
 
 Ca(CNS) 2 + Ca(OH) 2 + 2H 2 = 
 2NH 3 + COS + CaC0 3 + CaS. 
 1 " Jnl. of Gas Lighting." May 5, 1914. 
 
CYANOGEN 73 
 
 The carbon oxysulphide thus formed is removed through a 
 layer of lime 2 or 3 inches thick, this lime being placed over 
 the mixture in the heating chamber. The reaction covering 
 the removal of the carbon oxysulphide will then be 
 COS + 2Ca(OH) 2 = CaS + CaC0 3 + 2H 2 0. 
 
 It is very essential that the carbon oxysulphide be removed 
 in this manner, as otherwise it might combine with the ammonia 
 and form urea, according to 
 
 COS + 2NH 3 = H 2 S + CO(NH 2 ) 2 . 
 
 It is claimed that this process will also give a conversion of 
 98 per cent, but the application of either method will of course 
 depend upon local operating conditions. 
 
 The two systems of cyanogen recovery which have met 
 with the most pronounced success on the European Continent 
 and here, are that of Bueb and the one devised by Feld, the 
 former making an insoluble ferrocyanide cake, and the latter a 
 soluble cyanide sludge. 
 
 THE BUEB PROCESS 
 
 In the Bueb process iron sulphate, or copperas, and am- 
 monia are used as the two reagents necessary to combine 
 with the cyanogen, and a certain portion of the ammonia 
 thus removed from the gas remains in the cyanide cake, the 
 final resulting compound being ammonium ferro-ferro-cyanide. 
 
 The plant required for the extraction of cyanogen and 
 the production of the cyanide cake is shown in Fig. 14, the 
 various apparatus being enumerated from 1 to 22, and the 
 method of operation is as follows: 
 
 Iron sulphate is stored in a bin, and the amount necessary 
 for producing the requisite copperas solution is admitted into 
 tank (2), where it is mixed with fresh warm water, or wash 
 water from the filter press (10), the solution being agitated 
 by means of the mechanically driven stirring arrangement in 
 tank (2). The amount of iron sulphate (FeS0 4 ) required in 
 the solution can be determined exactly after the cyanogen 
 content of the gas is known, but during very warm weather 
 a further amount of FeS0 4 is required to take care of the 
 condensation from the gas, and thus to prevent a dilution of 
 the liquor by the addition of water. 
 
74 COAL GAS RESIDUALS 
 
 The copperas solution (FeS0 4 .7H 2 0) is pumped by pump 
 (3) into the overhead supply tank (4), from whence a regu- 
 lated quantity, depending on the strength of the solution, is 
 permitted to flow into the top of the washer (5) ; the resultant 
 liquor is permitted to remain in the washer until the copperas 
 solution is saturated, after which it is run into the storage 
 tank (6). 
 
 From the storage tank the sludge is pumped through the 
 still (12), where a portion of the free ammonia is boiled off 
 and condensed in the condenser (13), the sludge then passing 
 to the neutralizing tank (7); this latter tank is lead-lined 
 and hooded, and is provided with a steam coil as well as an 
 agitator. A small quantity of sulphuric acid (oil of vitriol) 
 is added to the sludge in this tank in order to neutralize it 
 while the sludge is being heated to about 200° F., the acid 
 being diluted with about three times its volume of water. 
 The hydrogen sulphide liberated during the neutralizing 
 period is caught under the hood of the tank and withdrawn 
 through a pipe by means of a fan and conducted to a con- 
 venient boiler stack. 
 
 After neutralization the sludge is an insoluble double ferro- 
 cyanide of ammonium and iron, plus a solution of ammonium 
 sulphate, a little ferrous sulphate, and a trace of free acid, 
 and it is now run into the neutralized liquor storage tank (9), 
 and is pumped from here to the filter press (10). 
 
 The sludge is pressed into a cake in the press, and after 
 being washed is ready for the market, while the ammonium 
 sulphate solution from the press is run into the precipitating 
 tank (14), where it meets the concentrated ammonia, or am- 
 monium sulphide from the condenser (13), this ammonia pre- 
 cipitating any iron which might be carried in the solution, 
 thus preventing the discoloration of the resultant salts. 
 
 The sulphate solution is then pumped into the vacuum 
 evaporator (16), where it is boiled to the point of crystalliza- 
 tion, after which the crystals are dried in the centrifugal (19) 
 and rotary dryer (20). 
 Chemistry of the Process: 
 
 The following reaction takes place in the upper section of 
 the washer, where the FeS0 4 is precipitated and FeS, or iron 
 sulphide, is formed, or 
 
"♦W "W»T» 
 
 (Facing Page 74) 
 
Missing Page 
 
CYANOGEN 75 
 
 Reaction I: 
 
 2FeS0 4 .7H 2 + 2H 2 S + 4NH, = 2FeS + 2(NH 4 ) 2 S0 4 . 
 
 The next reaction takes place in the lower portions of the 
 washer, where the CN reacts or combines with the FeS, or 
 Reaction II: 
 
 2FeS + 2NH 3 + 6HCN = (NH 4 ) 2 Fe 2 (CX) 6 + 2H 2 S. 
 
 Due to the contact between the (NH 4 ) 2 Fe:(CN) 6 and the 
 NH 3 and H 2 S in the gas, a partial decomposition of the 
 former takes place, resulting in the formation of some soluble 
 (NH 4 ) 4 Fe(CN) 6 , or ammonium ferro-cyanide, or 
 Reaction III: 
 
 (NH 4 ) 2 Fe 2 (CN) 6 + H 2 S + 2NH 3 = (NH4)«Fe(CN), + FeS. 
 The next reaction takes place in the still, or 
 Reaction IV: 
 
 (NH 4 )4Fe(CN) 6 + FeS + A = (NH 4 ) 2 Fe 2 (CN) 6 + (NH 4 ) 2 S, 
 the resultant ammonium sulphide being found in the con- 
 denser. 
 
 The fifth reaction occurs in the neutralizing tank, where 
 the sulphuric acid is added, or 
 Reaction V: 
 
 (NH 4 ) 4 Fe(CN) 6 + FeS + H 2 S0 4 = 
 (NH 4 ) 2 Fe 2 (CN) e + (NH 4 ) 2 S0 4 + H 2 S, 
 
 and the last reaction takes place in the sulphate liquor pre- 
 cipitation tank, where the iron is precipitated from the liquor 
 through the agency of the (NH 4 ) 2 S produced in the still and 
 condenser, or 
 Reaction VI: 
 
 FeS0 4 + (NH 4 ) 2 S = FeS + (NH 4 ) 2 S0 4 . 
 
 The salable product in this case is an insoluble ammonium 
 ferro-ferro-cyanide cake and ammonium sulphate. 
 
 Assuming the carbonization of 300 tons of coal per day, 
 and a cyanogen content of 120 grains per 100 cubic feet of 
 gas, allowing 10,000 cubic feet of gas to the ton of coal, the 
 estimated net revenue due to the extraction of cyanogen and 
 the production of a cyanide cake, as well as ammonium sul- 
 phate from the press liquor, based on yellow prussiate of 
 potash having a value of sixteen cents per pound, will be 
 given by the following: 
 
76 COAL GAS RESIDUALS 
 
 Cyanogen removed per h day 514 pounds 
 
 FeSO* required per day 1845 
 
 Ammonia removed per day 345 
 
 Sulphate made per day 770 
 
 NHs in press cake per day 119 
 
 Sulphuric acid required per day 270 
 
 Gross Income 
 
 Cyanogen* 514 pounds at 13.25* = $68.10 
 
 Sulphate 770 " at 3 .00* = 23.10 
 
 Press Cake NHa 119 " at 7.00* = 8.33 
 
 Total gross income per day $99.53 
 
 Operating Expense 
 
 FeS04 per day '. . . .0.923 ton at $ 9.50 = $ 8.77 
 
 Sulphuric acid 0.135 ton at 11.50 = 1.55 
 
 Steam for neutralizer \ - 3 00 
 
 Steam for evaporator J 
 
 Power, 870K.W.H at 1* = 8.70 
 
 Labor = 12.00 
 
 Miscellaneous supplies — 4.00 
 
 Total operating expenses per day $38.02 
 
 Net daily income $61.51 
 
 In a series of eight tests made on a plant such as described 
 here, the average result was as follows: 
 
 Cyanogen extracted from the gas 95.59% 
 
 Cyanogen recovered and fixed 98.93% 
 
 Ammonium sulphate recovered 95.76% 
 
 THE FELD PROCESS 
 
 A diagram of the Feld cyanogen-extracting process is shown 
 in Fig. 15, and the operation proceeds as follows: 
 
 Lime is slaked in the vessel marked (3), the resultant CaO 
 being mixed with water in (4), or in place of fresh water the 
 unfiltered wash liquor from the wash-liquor tank (5) may be 
 used. A concentrated solution of copperas, FeS04.7H 2 0, is 
 made in tank (2) by dissolving the iron sulphate with steam 
 and water. The milk of lime in conjunction with the cop- 
 peras solution is run into tank (5), this latter tank also being 
 provided with an agitator, and the mixed solution thus 
 formed is pumped into the top of the washer. The valve on 
 the liquor inlet pipe is provided with a regulator so as to 
 govern the exact quantity of washing medium admitted into 
 the washer, thus insuring a saturated liquor at the overflow. 
 
 The resultant sludge is run from the overflow into tank (7); 
 when this tank is almost full some fresh milk of lime is added, 
 after which the liquor is thoroughly boiled, tank (7) being 
 supplied with a steam coil for this purpose; care should be 
 taken to keep the agitator in motion during this boiling 
 
 1 The price given for cyanogen was prevalent prior to 1915; it has sold 
 for as much as 85 cents per pound since then. 
 
CYANOGEN 
 
78 COAL GAS RESIDUALS 
 
 period. The liquor is now pumped into the filter press (8), 
 the clear liquor coming from the press being divided in its 
 flow, 75 per cent going to the storage tank (10), and 25 per 
 cent to the wash-liquor tank (5). The press is washed out 
 with hot water coming from the tank (9), this wash water 
 flowing from the press to tank (10). 
 
 The product is calcium ferro-cyanide, and it can be shipped 
 in tank cars or in barrels; it may also be concentrated by 
 boiling if desired, thus reducing its shipping volume. If the 
 liquor should be permitted to stand in storage for any pro- 
 tracted time, some milk of lime should be added from time 
 to time in order to keep it alkaline. 
 Chemistry of the Process: 
 
 In order to explain the practical chemistry of this process, 
 a plant producing 3,500,000 cubic feet of gas per day will be 
 taken as an example under the following premises: 
 
 1 cubic meter of gas = 35 cubic feet. 
 1000 cubic feet of gas = 28.3 cubic meters. 
 1 gramme = 15.43 grains. 
 1 grain = 0.0648 gramme. 
 
 It is also assumed that the gas contains 120 grains of 
 cyanogen per 100 cubic feet, or 77.75 grammes per 28.3 cubic 
 meters, therefore one cubic meter will contain 2.75 grammes 
 of cyanogen. 
 
 The reactions are as follows: 
 Reaction I: 
 
 FeS0 4 .7H 2 + Ca(OH) 2 - Fe(OH) 2 + CaS0 4 .2H 2 + 5H 2 0. 
 
 Reaction II: 
 
 Fe(OH) 2 + 2Ca(OH) 2 + 6HCN = Ca 2 Fe (CN), + 6H 2 0, 
 
 and by adding reactions I and II together we have 
 
 Reaction III: 
 
 FeS0 4 .7H 2 + 3CaO + 6HCN = 
 Ca 2 Fe(CN) 6 + CaS0 4 .2H 2 + 8H 2 
 
 for which the molecular weights are 
 
 278 + 168 + 162 = 292 + 172 + 144, 
 
CYANOGEN 79 
 
 and in order to maintain the constituents of the washing 
 medium in excess of the theoretical quantities required by 
 the HCN, we would make these molecular weights 
 
 300 +200 +150 = 250 + 220. 
 
 Under the above premises 55 cubic meters of gas would 
 contain 150 grammes of HCN, and for 1000 cubic meters of 
 
 .. . 1000 „ . 
 
 gas we would require -r^- = 18 times the above constants, or 
 
 18 X 300 = 5400 grammes of FeS0 4 .7H 2 
 
 18x200=3600 " " Ca() 
 
 18 x 150 = 2700 " " HCN 
 
 18 x 250 = 4500 " " Ca.Fe(CN), 
 
 18x220=4000 " " CaSO,.2H 2 3 
 
 besides some CaCOs and FeS. 
 
 Therefore for 100,000 cubic meters, or 3,500,000 cubic feet 
 
 of gas in 24 hours, we would have: 
 
 560 kg. = 1232 pounds of FeSO„.7H 2 0, equivalent to 
 
 2000 liters = 500 gallons of iron sulphate liquor of 28 per cent 
 
 strength. 
 360 kg. =792 pounds of lime, CaO, equivalent to 
 1100 liters = 275 gallons of milk of lime of 30 per cent CaO. 
 270 kg. = 594 pounds of HCN. 
 
 450 kg. = 990 pounds of Ca 2 Fe(CN) 6 , equivalent to 
 3000 liters = 750 gallons of calcium ferro-cyanide liquor of 
 
 15 per cent Ca 2 Fe(CN) 6 ; 
 
 but as about one cubic meter of the liquor would be used 
 daily to thin out the iron sulphate solution, it is necessary to 
 filter 4000 liters, or about 1000 gallons of calcium ferro- 
 cyanide liquor of 15 per cent strength; we also have 400 kg. 
 = 880 pounds of CaS0 4 .2H 2 0, equivalent to 800 kg. = 1760 
 pounds of filter press cake containing 50 per cent H 2 0. 
 
 The relative mixtures given above are only approximate; 
 the greater the amount of filtered cyanide liquor which is 
 mixed with the dissolved iron sulphate, the more concentrated 
 will be the resultant calcium ferro-cyanide liquor. In Ham- 
 burg, Germany, a 15 per cent solution is used, but it is 
 possible to work with from 18 to 20 per cent if the concentra- 
 tion of the FeSO<.7H 2 liquor is increased to 35 per cent, and 
 then mixed with 1.5 cubic meters of filter liquor. It is also 
 
80 COAL GAS RESIDUALS 
 
 permissible to mix the liquor with non-filtered calcium ferro- 
 cyanide liquor; it is, however, better practice to only partly 
 mix with unfiltered liquor, or on the other hand to entirely 
 use filtered liquor in order to prevent the wash liquor becom- 
 ing too thick. 
 
 The product coming from either the Bueb or Feld plant, is 
 usually worked up into potassium ferro-cyanide, and in this 
 case the working up of the sludge produced by the Bueb system 
 requires additional apparatus for the removal and concentration 
 of the ammonia carried in the sludge, while the working up of 
 Feld sludge obviates the necessity of this ammonia apparatus, 
 as the Feld system embraces the removal of cyanogen from the 
 gas after the ammonia has been extracted. 
 
 Another system of cyanogen extraction, and one which is 
 used in Great Britain, is that of Davis-Neill, in which the 
 cyanogen is removed from the gas by washing with a solution 
 of soda and ferrous carbonate, washers of the "Standard" hori- 
 zontal or vertical type being used for this purpose. The 
 resultant cyanogen liquor coming from the washers is run into 
 a still, where the ammonia carried by the water is distilled off, 
 the ammonia vapors thus produced being condensed and then 
 run to the ammonia cisterns. 
 
 After being freed of ammonia, the cyanide liquor is pumped 
 into a filter press, where the insoluble matter is separated and 
 retained, the clear solution, as well as the wash water, from the 
 press being then conducted into an evaporator, where the 
 liquors are concentrated, after which they are run into crystal- 
 lizing vats. These crude crystals, either sodium ferro-cyanide, 
 or prussiate of soda, are now washed with mother liquor from 
 the evaporator in order to remove excess soda and any insol- 
 uble matter, after which they are dissolved by means of hot 
 water, no more water being used than is absolutely necessary 
 to produce the required solution. 
 
 Insoluble matter carried in the crystals is then allowed to 
 settle, after which the clear liquor is run into another set of 
 vats, where the final crystals are formed on strings suspended 
 in the liquid. 
 
 The wash liquor, consisting of soda and ferrous sulphate in 
 solution, is prepared in mixing vats, and the precipitated ferrous 
 carbonate is separated from the sodium sulphate solution in 
 filter-presses, after which it is mixed with soda and water. 
 
CYANOGEN 81 
 
 THE MANUFACTURE OF POTASSIUM FERRO-CYANIDE 
 
 A diagram showing a plant arranged for the manufacture 
 of potassium ferro-cyanide or yellow prussiate of potash, 
 using Bueb sludge, is shown in Fig. 16. 
 Bueb Sludge Plant: 
 
 The (NH 4 ) 2 Fe 2 (CN)6 coming from the cyanogen washers is 
 stored in tank (1), from whence it is transferred by pump (2) 
 into the neutralizing vessel (3); here the sludge is neutralized 
 with a small amount of sulphuric acid (oil of vitriol), as is 
 done in the Bueb cyanogen extraction system, the acid being 
 fed in small quantities from the overhead supply tank (4), 
 this latter tank receiving its supply from the sulphuric acid 
 storage tank (5), the delivery being effected by means of 
 compressed air. After the sludge has been neutralized it is 
 delivered by pump (6) to the storage tank (7). 
 
 By means of pump (8) the sludge is now transferred to the 
 digester (9), where it is boiled with steam, the issuing vapors 
 being condensed in the condenser (10) and collected in tank 
 (11) as concentrated ammonia. 
 
 The boiled sludge is then run into the monteju (12), and 
 is forced from thence by means of compressed air into the 
 filter press (13); the ammonium sulphate liquor from the 
 filter press flows into the settling boxes (14), and from thence 
 into the sulphate liquor storage tank (15), and from here the 
 clarified sulphate solution is transferred by pump (16) into 
 the vacuum evaporator (17), where it is evaporated to 
 crystallization. 
 
 The screw conveyor (18) carries the crystalline mash to the 
 centrifugal (19), the mother liquor from the centrifugal 
 returning to tank (15), while the salts are dried in the dryer 
 (20). The "blue" cake left in the filter press (13) is carried 
 by the conveyor (21) to the cake storage (22), and from 
 thence over the cake chutes (23) to the disintegrators (24). 
 Here the "blue" cake is broken up and mixed with a weak 
 ferro-cyanide wash liquor and then run into the decomposer 
 (25), where it is boiled with lime from the agitator (26), 
 potash also being added at this point, until all of the ammonia 
 is driven off. 
 
 The vapors from the decomposer (25) are condensed in 
 
82 COAL GAS RESIDUALS 
 
 (27) and collected in the ammonia receiver (28) as concen- 
 trated ammonia. After being boiled, the ammonia free mash 
 is allowed to settle in the decomposer, after which the clear 
 liquor is drawn off into tank (29), and from there into the 
 collector (30), while the mash is run into the monteju (31) and 
 forced from there by compressed air into the filter press (32). 
 
 From this filter press all of the ferro-cyanide, in the form 
 of potassium ferro-cyanide liquor, flows to the collector (30), 
 while the press cake removed from the filter is principally 
 iron. This weak potassium ferro-cyanide liquor is trans- 
 ferred by means of pump (33) from the collector to the vacuum 
 evaporator (34), where it is evaporated to the point of 
 crystallization, and then run into the salt cooler (35); here, 
 while cooling, the greater portion of the potassium ferro- 
 cyanide solidifies to a crystal mass which, after drawing off 
 the mother liquor into tank (36), is transferred by the screw 
 conveyor (37) to the centrifugal (38), the mother liquor from 
 the centrifugals also being collected in tank (36). All of the 
 mother liquor is finally transferred by means of pump (39) 
 to the vacuum evaporator (34). 
 
 The crude salt is now dissolved in the dissolver (41), and 
 the resultant solution is run into the concentrator (42), 
 where the solution is again evaporated to the point of crystal- 
 lization and the pure concentrated potassium ferro-cyanide 
 liquor, ready to crystallize, is run while still hot into the 
 crystallizer (43). After being completely cooled here, the 
 potassium ferro-cyanide crystals deposit on the sides of 
 the crystallizer and upon strings suspended in it; after 
 drawing off the mother liquor and returning it to the 
 concentrator (42) by means of pump (44), the crystals are 
 removed to the dryer (45), and are then packed ready for 
 shipment. 
 
 Yellow prussiate of potash, or potassium ferro-cyanide, 
 K 4 Fe(CN)6.3H 2 0, has a specific gravity of 1.8533, and its 
 molecular weight is 422.36, made up by 
 
 Potassium, K, =39.10 x 4 = 156.40 
 Iron, Fe, = 55.85 x 1 = 55.85 
 
 Cyanogen, CN, = 26.01 X 6 = 156.06 
 Water, H 2 0, = 18.016 X 3 = 54.05 
 
 422.36 
 
Cyanogen 
 Wa$her. 
 
 1 Cyanogen Sludqe Storage 
 
 2 Sludge Pump. 
 
 3 Neutralizer. 
 
 4 Sulphuric Acid Supply. 
 
 5 Sulphuric Acid Storage. 
 
 6 Neutralized Sludqe Pump. 
 
 7 Neutralized Sludge Storage. 
 
 8 Pump to Digester. ' 
 
 9 Digester. 
 10 Condenser, 
 
 11 Concentrated AmmoniaTank. 
 
 12 Monteju. 
 
 13 Filter Press. 
 
 14 Settling Boxes. 
 
 15 Sulphate Liquor Storage. 
 
 16 Sulphate Liquor Pump. 
 IT Vacuum Evaporator. 
 
 18 Sulphate Conveyor. 
 
 19 Sulphate Centrifu&au 
 
 20 Sulphate Dryer. 
 
 21 CakeC( 
 
 22 CAKE 51 
 
 23 Cake 51 
 
 24 DE8INTD 
 
 25 Decomf 
 
 26 Lime Ac 
 
 27 Conoen 
 
 28 Ammok 
 
 29 Clear I 
 
 80 COLLEC 
 
 Fig. 16.— POTASSIUM FERRO CYANIDE 
 
/EYER. 
 tAQE. 
 
 rES. 
 
 *T0R8 
 
 :r. 
 \tor. 
 
 Receiver. 
 uorTank. 
 *<i Tank. 
 
 AXT OPERATING WITH 
 
 31 Monte: ju. 
 
 32 Filter Press 
 
 33 Cyanide Liquor Pump. 
 
 34 Vacuum Evaporator. 
 
 35 Salt Cooler. 
 
 36 Mother Liquor Tank. 
 
 37 Crystal Conveyr. 
 
 38 Crystal Centrifugal, 
 
 39 Mother Liquor Pump. 
 
 40 Salt Receiver. 
 
 BUEB" SLUDGE 
 
 41 Salt Dissolved. 
 
 42 Salt Concentrator. 
 
 43 Crystaluzer. 
 
 44 Mother Liquor Puivir 
 
 45 Crystal Dryer. 
 
 46 Tail Pipe Seals. 
 4T Air Compressor. 
 
 48 Air Receiver. 
 
 49 Vacuum Pump. 
 
 50 Condensers. 
 
 (Facing Page 82) 
 
CYANOGEN 83 
 
 In the Bueb process of cyanogen extraction, 1.95 pounds 
 of "blue," or (NH„) 2 Fe2(CN) 6 , are produced for every pound 
 of cyanogen extracted from the gas, and to produce 
 K 4 Fe(CN)«.3H 2 we must add 2K 2 C0 3 , or two molecules of 
 potassium carbonate to one of the "blue," the commercial 
 compound known as "black salt" and containing K 2 C0 3 being 
 used for this purpose. 
 
 The molecular weight of K 2 C0 3 is 138, and that of 
 
 (NH 4 ) 2 Fe 2 (CN) 6 is 304, therefore each pound of "blue" will 
 
 require 
 
 138 X 2 
 3Q4 = 0.908 pound of K 2 C0 3 . 
 
 As the commercial "black salt" only contains 75 per cent 
 of K 2 C0 3 , we will require ~r- = 1.211 pounds of black salt 
 
 for each pound of "blue" produced. 
 
 The amount of lime, CaO, required is 1.65 pounds per 
 pound of fixed ammonia contained in the sludge, or using a 
 commercial article containing 95 per cent CaO, we will re- 
 quire 1.73 pounds of lime per pound of fixed ammonia. 
 
 The amount of iron sulphate or copperas required will be 
 the same as mentioned in the Bueb cyanogen extraction 
 process. 
 
 Feld Sludge Plant: 
 
 As stated before, the production of yellow prussiate of 
 potash from insoluble sludge is a roundabout procedure and 
 requires much more labor and apparatus, and leads to possibly 
 more ammonia losses than if a soluble sludge were used. 
 Making calcium ferro-cyanide, which is soluble, in the washer 
 does away with extra filter presses, sulphuric acid, and 
 removal of ammonia, and it also requires less steam and 
 improves the final product; besides this, less copperas will 
 be required in the extraction process. 
 
 A plant to produce potassium ferro-cyanide, or, if desired, 
 sodium ferro-cyanide from soluble sludge is shown in Fig. 
 17. Here, as explained under the Feld cyanogen process, the 
 alkali constituent of the gas, ammonia, is replaced by an 
 external alkaline medium, and the first half of the operation 
 proceeds as explained in the Feld extraction process. 
 
84 COAL GAS RESIDUALS 
 
 The ferrous sulphate solution is mixed in a suitable vat or 
 tank with the proper equivalent of the alkaline agent, princi- 
 pally milk of lime, Ca(OH 2 ), in such proportion that the 
 alkali is in excess prior to its conveyance to the washer, 
 which in this instance is located subsequent to the apparatus 
 for ammonia removal. 
 
 In contact with the counter-current of gas, the ferrous 
 hydrate, Fe(OH) 2 , is converted by the means of hydrogen 
 sulphide, H 2 S, into ferrous sulphide, FeS, which absorbs 
 the cyanogen compounds, the presence of the calcium com- 
 pounds causing the formation of calcium ferro-cyanide, 
 Ca 2 Fe(CN) 6 ; this reaction, already given under the extrac- 
 tion process, may also be expressed as 
 
 FeS + 2CaC0 3 + 6HCN = 
 Ca 2 Fe(CN) 6 + H 2 S + 2C0 2 + 2H 2 0. 
 
 The resultant sludge, a mixture of insoluble calcium sul- 
 phate, carbonate, etc., with traces of ferrous sulphide in a 
 solution of calcium ferro-cyanide, is conveyed to a suitable 
 vat or storage tank. 
 
 This liquid is filter-pressed and the clear calcium ferro- 
 cyanide liquor is conveyed to a steam-heated digester, or still, 
 in which, by the addition of alkaline potassium or sodium 
 derivatives, principally carbonates, it decomposes, forming 
 sodium or potassium ferro-cyanide and insoluble calcium 
 carbonate, CaC0 3 . 
 
 After separating the latter compound by means of filter 
 pressing, the clarified ferro-cyanide solution is concentrated 
 and the product purified by re-crystallization. The "slimes" 
 or mud from the first filter press may be returned and mixed 
 with the washing liquid, so proportioning the admixture as 
 to prevent vitiation of its efficiency due to the presence of 
 too large a proportion of an inactive agent. 
 
 Cyanogen from Spent Oxide. In plants not provided with a 
 direct cyanogen-extraction system, the cyanogen-compounds will, 
 to a great extent, be found in the spent oxide coming from the 
 purifiers; the cyanogen content of this oxide makes the latter 
 valuable, and although the oxide also contains sulphur varying 
 from 30 to 50 per cent, the spent material in this country is 
 rarely sold for its sulphur content, the cyanogen content being, 
 as a rule, the basis on which the value is fixed. 
 
receiver: 
 
 DlSSOLVER. 
 
 Concentrator. 
 
 ■ALIZER. 
 
 ier Liquor Pump. 
 tai_ dryer. 
 s \pc Seal, 
 ier Liquor Tank, 
 ier Liquor pumr 
 um Pump. 
 
 JENSER. 
 SLD" SLUDGE 
 
 (Facing Page 84) 
 
Missing Page 
 
CYANOGEN 85 
 
 The chemical works operating on spent oxide may prefer to 
 remove the sulphur before attacking the cyanogen problem, 
 but if this is not done, the sulphur must be carried through the 
 entire process of cyanogen extraction; the latter method is, as 
 a rule, the one usually pursued, because the removal of the sul- 
 phur is not only a difficult operation, and one requiring constant 
 supervision, but it is also subject to a high fire risk and the action 
 of the carbon disulphide used as the solvent is very injurious 
 to the human system. 
 
 If the sulphur is to be removed, the spent oxide is ground to 
 a fine powder and then placed within a cylindrical vessel having 
 a perforated false bottom over which a filter cloth is spread to 
 receive the oxide after the vessel has been filled nearly to the 
 top, another filter cloth is placed over the top of the oxide, and 
 the vessel closed tight. Carbon disulphide vapors from a steam- 
 heated still are now charged into the closed vessel containing 
 the spent oxide, the disulphide vapors condensing and passing 
 down through the mass, relieving the latter of the sulphur 
 burden during the passage, the liquor from the extracting 
 vessel then being returned to the still. The apparatus is also 
 provided with a condenser for taking care of any carbon disul- 
 phide vapors not condensed in the extractor, this condensed 
 liquor being run to storage for charging the still when required. 
 
 When all sulphur has been removed from the oxide, the vapor 
 connection to the extractor is closed off, and the carbon disul- 
 phide is driven off of the sulphur in the still, the resultant 
 vapors being sent to the condenser and the carbon disulphide 
 liquor from there to storage, after which it is again ready for 
 the extraction process. The sulphur is now free of the solvent, 
 and can be drawn off when melted, steam being blown through 
 the residue in order completely to remove all traces of carbon 
 disulphide 
 
 The sulphur thus obtained may be converted into sulphuric 
 acid by roasting. 
 
 Quite a number of processes are in use for the extraction and 
 conversion of the cyanogen compounds, but they are all based 
 on three principals, viz.: 
 
 I. Lixiviation of the oxide for the removal of the soluble 
 salts and the recovery of ammonia and sulphocyanides. 
 
 II. The conversion of the insoluble ferro-cyanides into sodium 
 or calcium ferro-cyanides. 
 
86 COAL GAS RESIDUALS 
 
 III. Conversion of sodium or calcium ferro-cyanide into 
 potassium ferro-cyanide, if the latter is to be the end product. 
 
 Lixiviation is conducted in a series of filtering vats contain- 
 ing a false bottom of wooden beams, over which a layer of 
 twigs is spread to a depth of about 4 inches, the top of the twigs 
 being covered with a filter-cloth. Each vat communicates by 
 means of a trough with at least three receivers. The vats are 
 now filled with the spent, oxide, and water is run into vat No. 1, 
 filling same to a point about one inch above the oxide. 
 
 After being allowed to stand for about 24 hours, the re- 
 sultant liquor is drawn off into the first receiver, and pumped 
 from there into vat No. 2, vat No. 1 being again filled with 
 fresh water. After another 24 hours this process is repeated, 
 the liquor .from No. 2 vat going to the second receiver and 
 that from No. 1 to the first. From the second receiver the 
 liquor now goes to vat No. 3, and after the proper period of 
 settlement this' series of changes is repeated until the original 
 water from No. 1 vat has passed through at least eight vats in 
 series, after which the liquor contains its burden of ammonia. 
 This ammonia liquor is pumped to a storage tank, and from 
 there to a still where it is distilled with lime, and the ammonia 
 is recovered. The liquor after distillation contains calcium 
 sulphocyanide in solution. 
 
 The second process consists in converting the insoluble fer- 
 ro-cyanides into soluble compounds, and this may be done in a 
 digester, or in a series of filtering vats similar to those described 
 above. 
 
 The digester is provided with a vertical shaft operating a 
 stirring paddle, and after the solid material from the lixiviation 
 vats has been dried and pulverized, it is mixed with soda and 
 lime and then charged into the digester. The digester is heated 
 with direct steam, and the mass is cooked under constant 
 agitation, the sludge from the bottom of the digester being run 
 into filter-presses where the solution is separated. 
 
 If the filter vats are used, the process is similar to that 
 described for lixiviation, but the solutions will be of a higher 
 degree of concentration than those secured from the digester. 
 
 It will now be necessary to precipitate the ferro-cyanides, 
 and while the solution from the digester contains both ferro- 
 cyanide of sodium and of calcium, the solution from the vats 
 contains all of the ferro-cyanide in the shape of calcium ferro- 
 
CYANOGEN 87 
 
 cyanide, because lime only was used in this mixture, while 
 both lime and soda were used in the mixture placed in the 
 digester. Both solutions contain some sulphocyanide of cal- 
 cium and ammonia. The ferro-cyanides may be precipitated by 
 means of iron salt, ammonium salts, or potassium chloride. 
 
 If iron salts are used (FeCl 2 or FeS0 4 ), the ferro-cyanides 
 will be precipitated in the form of prussian blue, ferrous or 
 ferric chloride being used if the solution contains calcium ferro- 
 cyanide, while ferrous sulphate is used if the solution contains 
 sodium ferro-cyanide. 
 
 Owing to the presence of free ammonia, the sulphides, which 
 are always found in the solutions coming from the digester, 
 are alkaline, and they are acidified in a special vat, after which 
 they are allowed to stand in order that the sulphur may separate 
 out. The clear liquor can now be decanted, and it will then be 
 ready for precipitation with the iron salt. This latter operation 
 is performed in wooden vats, the solution being constantly 
 stirred while the iron is being added; when precipitation is 
 complete the supernatant liquor is drawn off, the prussian blue 
 is filter-pressed and then converted into potassium ferro-cyanide 
 by treating with potash or potassium carbonate, if this is the 
 end product desired. 
 
 If the precipitation is to be performed by means of ammoni- 
 acal salts, the operation should be conducted in such manner 
 that the salt is produced in sufficient quantity at the time of 
 precipitation, by using lixiviated solutions which carry ammonia 
 in such quantity as to make its neutralization with hydro- 
 chloric acid necessary. In order to accomplish this, some un- 
 lixiviated material, plus lime, should be added to the lixiviated 
 solution, this procedure giving the amount of ammonia required 
 for the reaction. 
 
 This operation is conducted in a vat provided with means 
 for agitating the solutions, hydrochloric acid being added until 
 the acid reaction is complete. The mixture is then heated to 
 about 175° F., after which the double salt, according to the 
 reaction, 
 
 Ca 2 Fe(CN) 6 + 2NH 4 C1 = Ca(NH 4 )2Fe(CN) 6 + CaCl, 
 
 will separate out in the form of a white powder. When pre- 
 cipitation is completed, agitation is stopped, and the precipitate 
 is allowed to settle, the clear liquor being drawn off and the 
 
88 COAL GAS RESIDUALS 
 
 precipitate filter-pressed. Some of the double salt will remain 
 in the mother liquor, and this can be precipitated as prussian 
 blue by means of iron salts. 
 
 The double salt, Ca(NH 4 ) 2 Fe(CN) 6 , may now be treated with 
 lime in an agitating vessel in order to secure pure calcium 
 ferro-cyanide according to the reaction 
 
 Ca(NH 4 ) 2 Fe(CN) 6 + CaO - Ca 2 Fe(CN) 6 + 2NH 3 + H 2 0, 
 
 and the calcium ferro-cyanide may then be converted into potas- 
 sium ferro-cyanide by using potassium chloride to precipitate 
 the double salt CaK 2 Fe(CN) 6 , and then converting this salt 
 into potassium ferro-cyanide by boiling with potassium car- 
 bonate according to the reaction, 
 
 CaK 2 Fe(CN) 6 + C0 3 K 2 = K 4 Fe(CN) 6 + CaC0 3 . 
 If it is desired to make the conversion with potassium car- 
 bonate in the presence of lime, the reaction will be 
 
 Ca(NH 4 ) 2 Fe(CN) 6 + 2C0 3 K 2 + CaO = K 4 Fe(CN) 6 + 2CaC0 3 + 
 
 2NH 3 + H 2 0. 
 The potassium ferro-cyanide solution thus obtained is now 
 concentrated to 30° Be., and then allowed to crystallize out. 
 
 If potassium chloride is to be used as the precipitating 
 medium, according to equation 
 
 Ca 2 Fe(CN) 6 + 2KC1 = CaK 2 Fe (CN) 6 + CaCl 2 , 
 
 the double salt thus formed will be found to be slightly soluble. 
 The calcium ferro-cyanide solutions are concentrated to 25° Be\, 
 and then admitted into a small digester provided with an agita- 
 tor, the precipitation in this digester being performed at a 
 temperature of 175° F. The potassium chloride, in crystalline 
 form, is used in excess, and the double salt CaK 2 Fe(CN) 6 will 
 be precipitated in the form of a pale-yellow powder, being later 
 separated by filtration or decantation, washed with water, and 
 finally treated with potassium carbonate in order to obtain 
 potassium ferro-cyanide, the reaction covering this conversion 
 being 
 
 CaK 2 Fe(CN) 6 + C0 3 K 2 = C0 3 Ca + K 4 Fe(CN) 6 . 
 
 The calcium sulphocyanides in the liquor from the first lixivia- 
 tion will be found in the residual solutions coming from the 
 ammonia still, and they may be precipitated by the use of a 
 potassium or sodium salt, and recovered by concentration and 
 crystallization. 
 
CHAPTER IV 
 AMMONIA 
 
 Ammonia is a product derived from the destructive distilla- 
 tion of coal, and it results from the union of nitrogen with 
 hydrogen, or N 2 + 3H 2 = 2NH 3 ; besides being one of the 
 principal sources of revenue, the removal of ammonia from 
 gas is necessary on account of its destructive influence on 
 the brass and copper work of meters and gas-fittings; when 
 burned with gas it gives off noxious fumes of oxides of 
 nitrogen. 
 
 The usual method employed for the removal of ammonia 
 depends upon its solubility in water, and at ordinary tempera- 
 tures water can absorb about 708 times its volume of am- 
 monia gas, the absorption increasing with a decrease in water 
 temperature, or 
 
 One volume of water at 32° F. will absorb 1050 volumes of NH 3 
 One volume of water at 50° F. will absorb 813 volumes of NH 3 
 One volume of water at 59° F. will absorb 727 volumes of NH 3 
 One volume of water at 68° F. will absorb 654 volumes of NH 3 
 One volume of water at 77° F. will absorb 586 volumes of NH 3 
 One volume of water at 183° F. will absorb 180 volumes of NH 3 
 
 providing the pressure of the ammonia gas is equal to that 
 of the atmosphere, and if the gas carries no tar burden. 
 
 If it is assumed that the amount of ammonia in crude gas 
 is 2 per cent by volume, the ammonia will be capable of 
 exerting only 2 per cent of the total pressure, of the gas. 
 It must also be remembered that the strength of a solution 
 of ammonia in water is dependent upon the pressure of 
 the ammonia gas with which it is in contact; therefore the 
 strength of the resultant liquor will be dependent upon the 
 pressure exerted by the ammonia in the crude gas coming in 
 contact with the liquor. 
 
 Liquors of various strengths and temperatures have definite 
 pressures, or tensions, and these pressures are capable of giv- 
 ing up ammonia to the gas if the ammonia pressure of the 
 
90 
 
 COAL GAS RESIDUALS 
 
 gas is less than that of the liquor, but if the ammonia pressure 
 of the gas is greater than that of the liquor, the liquor will be 
 able to extract more ammonia and gain in strength until a 
 point is reached at which the pressures in gas and liquor are 
 equalized. 
 
 The usual gas-works method of operation as practised in 
 America removes the ammonia by condensation and washing, 
 the first portion of the ammonia being removed in the hy- 
 draulic main due to the condensation of water vapors, which 
 thus absorb ammonia. Cooling the gas in the condensers 
 also causes the deposition of quite an amount of water which 
 in turn absorbs more ammonia, the final ammonia being 
 removed in the scrubbers or mechanical washers, using as 
 little wash water as is consistent with the result desired, viz., 
 strong liquor. Table XII gives the strength, weight, etc., 
 of ammonia liquors. 
 
 TABLE XII. — AMMONIA 
 
 Strength 
 
 Weight 
 
 per gal. 
 
 Per cent 
 
 of 
 
 NH 3 
 
 Twaddle 
 
 at 
 
 60° F. 
 
 Specific 
 gravity- 
 water 
 basis 
 
 Weight per 
 gal. in lbs. 
 
 in oz. 
 
 oz. 
 
 lbs. 
 
 at 60° F. 
 
 
 
 0.000 
 
 0.0000 
 
 0.0000 
 
 0.00 
 
 1.0000 
 
 8.3328 
 
 1 
 
 0.347 
 
 0.0217 
 
 0^2591 
 
 0.50 
 
 1.0025 
 
 8.3560 
 
 2 
 
 0.689 
 
 0.0434 
 
 0.5182 
 
 1.00 
 
 1.0050 
 
 8.3770 
 
 4 
 
 1.388 
 
 0.0868 
 
 1.0364 
 
 2.00 
 
 1.0100 
 
 8.4180 
 
 6 
 
 2.082 
 
 0.1301 
 
 1.5546 
 
 3.00 
 
 1.0150 
 
 8.4600 
 
 8 
 
 2.786 
 
 0.1735 
 
 2.0728 
 
 4.00 
 
 1.0200 
 
 8.5020 
 
 10 
 
 3.469 
 
 0.2169 
 
 2.5910 
 
 5.00 
 
 1.0250 
 
 8.5440 
 
 24 
 
 8.328 
 
 0.5205 
 
 6.2184 
 
 12.00 
 
 1.0600 
 
 8.8350 
 
 28 
 
 9.726 
 
 0.6079 
 
 7.2548 
 
 14.00 
 
 1.0700 
 
 8.9180 
 
 32 
 
 11.104 
 
 0.6840 
 
 8.2912 
 
 16.00 
 
 1.0800 
 
 9.0020 
 
 36 
 
 12.492 
 
 0.7807 
 
 9.3276 
 
 18.00 
 
 1.0900 
 
 9.0850 
 
 40 
 
 13.878 
 
 0.8676 
 
 10.3640 
 
 20.00 
 
 1.1000 
 
 9.1680 
 
 44 
 
 15.265 
 
 0.9544 
 
 11.4004 
 
 22.00 
 
 1.1100 
 
 9.2530 
 
 48 
 
 16.654 
 
 1.0411 
 
 12.4368 
 
 24.00 
 
 1.1200 
 
 9.3360 
 
 52 
 
 18.041 
 
 1.1276 
 
 13.4732 
 
 26.00 
 
 1.1300 
 
 9.4200 
 
 56 
 
 19.429 
 
 1.2143 
 
 14.5096 
 
 28.00 
 
 1.1400 
 
 9.5030 
 
 60 
 
 20.816 
 
 1.3014 
 
 15.5460 
 
 30.00 
 
 1.1500 
 
 9.5860 
 
 70 
 
 24.286 
 
 1.5179 
 
 18.1370 
 
 35.00 
 
 1.1750 
 
 9.7950 
 
 80 
 
 27.755 
 
 1.7352 
 
 20.7280 
 
 40.00 
 
 1.2000 
 
 10.0040 
 
AMMONIA 91 
 
 This table is to be used in connection with the "saturation" 
 test with H2SO4, used as a check on scrubber operation as 
 performed in many gas works, and is not indicative of the total 
 NH 3 content of the liquor, as only free NH 3 will combine with 
 the acid. The "ounce strength" in col. 1 means the ounces 
 of H2SO4 that will neutralize one gallon of the liquor used, and 
 the other data is a measure of the NH 3 equivalent of the acid 
 used for neutralization. 
 
 In order to determine the total NH 3 content of the liquor, 
 it is necessary to distill in the presence of an alkali, such as 
 caustic soda or hydrated lime, and catch the vapors in a 
 measured quantity of normal H2SO4 solution, each c.c. of which 
 is equivalent to 0.01703 gramme of NH 3 . The excess acid is 
 titrated with normal NAOH solution, and the calculation for 
 per cent of NH 3 will be: 
 
 (c.c. NAH 2 S04 - c.c. NANAOH) 0.0173 
 
 r— — ; = per cent NH 3 , or 
 
 weight of sample 
 
 per cent NH 3 x 3.88 = per cent (NH4>:S0 4 . 
 
 Assuming that the nitrogen in the coal amounts to 2 per 
 cent, and that of this 15 per cent forms ammonia, the follow- 
 ing example will illustrate how to calculate the amount of 
 ammonia present in the raw gas, under the above conditions, 
 also the amount of ten-ounce liquor to be expected provided 
 all of the ammonia produced is recovered, also the amount of 
 sulphate which this quantity of ammonia should produce. 
 
 As 2 per cent of the coal is assumed to be nitrogen, the 
 amount present in one ton of coal will be 2000 lbs. x 0.02 = 
 40 pounds, and if 15 per cent of this nitrogen is converted 
 into ammonia, we would have 40 X 0.15 = 6 pounds of nitro- 
 gen as ammonia per ton of coal, but as ammonia is a com- 
 pound consisting of 14 parts of nitrogen and 3 parts of 
 hydrogen, or if every 14 parts of nitrogen are equal to 17 
 parts of ammonia, we would have 
 
 —n — = 7.28 pounds of ammonia, 
 14 
 
 equivalent to 
 
 — — — = 509.6 grains per 100 cubic feet of gas. 
 
92 COAL GAS RESIDUALS 
 
 This amount of ammonia expressed in percentage by 
 volume can be determined by multiplying the weight in 
 grains of a cubic foot of hydrogen by half the molecular 
 weight of ammonia, and dividing this product into the number 
 of grains of ammonia per 100 cubic feet of gas, or we will 
 have in this case 
 
 509. 6 , A1 
 
 37.15 x 8.5 -I-** Per**"*. 
 
 The ounce strength of liquor is based on the sulphate 
 radical, or 2NH 3 + H 2 S0 4 = (NH 4 ) 2 S0 4 , in which 34 parts of 
 ammonia by weight require 98 parts of sulphuric acid, thus 
 forming 34 + 98 = 132 parts of ammonium sulphate. As the 
 ton of coal produces 7.28 pounds of ammonia, we would have 
 
 /7.28x98\ 1fi QQ .„ Q 
 
 ( ^r — J 16 = 335.68 ounces. 
 
 equivalent to 33.568 gallons of 10 ounce liquor. 
 
 The amount of sulphate produced by this amount of 
 ammonia will be 7.28 x 3.88 = 28.25 pounds per ton of coal. 
 
 Raw ammonia liquor from the plant usually contains from 
 1 to 2 per cent of ammonia, and due to the cost of trans- 
 portation it would not in most cases be economical to sell this 
 liquor in its diluted state; the raw liquor is, therefore, usually 
 worked up for one of the following ammonia products: 
 
 1. Concentrated liquor, 
 
 2. Aqua ammonia, 
 
 3. Sulphate of ammonia, 
 
 but the product to be recommended for any particular gas 
 works must necessarily depend upon local market conditions. 
 
 Concentrated ammonia is a product in which the raw, liquor 
 has been freed of a large portion of water, and it is used for 
 producing other ammonia products. 
 
 Aqua ammonia is a solution of ammonia in distilled water, 
 a great deal of it, in a dilute solution, being used for cleansing 
 purposes. 
 
 Ammonium sulphate is produced by distilling the liquor, 
 and absorbing in sulphuric acid; by passing the gas through 
 an acid bath, thus absorbing the ammonia, or by combining 
 the ammonia in the gas directly with the sulphur radical, 
 and subsequent oxidation, without the use of sulphuric acid, 
 
AMMONIA 93 
 
 Sulphate finds its greatest use in the manufacture of fertilizer, 
 it being the nitrogen-carrying agent in this product. For this 
 latter use, sulphate has lately found an active competitor in 
 calcium cyanamide, a product of the electric furnace, where 
 calcium carbide absorbs nitrogen from the atmosphere to 
 form cyanamide with the liberation of carbon, this product 
 containing from 14 to 22 per cent of nitrogen which is lib- 
 erated as ammonia when treated with water, thus making it 
 suitable for agricultural fertilizer. It appears that with soils 
 which are not rich in humus or not deficient in lime, cyana- 
 mide is almost as good, presuming that they all have the 
 same nitrogen content, as ammonium sulphate or sodium 
 nitrate, but it is of doubtful value with peaty soils, or soils 
 containing little lime, and it is not nearly as good as either 
 of the above for a soil top-dressing. 
 
 Ammonium chloride (sal ammoniac) is obtained by using 
 hydrochloric acid as an absorbing agent, followed by evapora- 
 tion and crystallization. In the United States the principal 
 source of sal ammoniac is ammonium sulphate, this com- 
 pound being heated in conjunction with common salt. It is 
 used as an agent in galvanizing iron, for soldering, and in 
 electric batteries. 
 
 Liquid anhydrous ammonia is produced by freeing the 
 ammonia gas of all moisture and impurities, after which it is 
 liquefied by pressure; it is used for the production of low 
 temperatures by the absorption of heat, as in refrigeration. 
 
 Ammonium carbonate is produced by heating a mixture of 
 ammonium sulphate in conjunction with calcium carbonate, 
 and it is used for scouring purposes, especially wool. 
 
 Ammonium nitrate is produced by neutralizing nitric acid 
 with ammonia, and it is used in the manufacture of explosives. 
 
 There are many other ammonia compounds, but the ones 
 mentioned above are the principal products, and they can all 
 be prepared from ammonium sulphate, thus making this latter 
 compound one of the principal by-products of coal distillation. 
 
 The crude liquor coming from the works contains ammonia in 
 combination with various radicals, the principal ones of which 
 are ammonium carbonate, sulphide, chloride, and sulphate. 
 These salts may be grouped into two classes in accordance with 
 their conduct during the process of boiling the crude liquor. 
 
94 COAL GAS RESIDUALS 
 
 Carbonates of ammonia as well as sulphides of ammonia, 
 below the boiling point. of the liquor, decompose into carbonic 
 acid, hydrogen sulphide, and ammonia, and as these are com- 
 pletely freed at the boiling point, they are classed as "free" 
 ammonia compounds. On the other hand ammonia com- 
 bined as chloride and sulphate can only be liberated by the 
 addition of some stronger alkali, lime in the form of milk of lime 
 being usually introduced into the solution for this purpose. 
 
 Analysis will give the amount of free and fixed ammonia 
 in the liquor, from which the yield of ammonia as well as 
 the amount of lime to be used can easily be deduced; as a 
 rule the liquor contains from 70 to 80 per cent of free, and 
 20 to 30 per cent of fixed, ammonia. 
 
 The working up of the liquor is usually effected in continuous 
 stills (Fig. 18), operating with a supply of live steam, and as the 
 principle of design and operation is practically the same in the 
 usual stills on the market, only one method will be described, 
 the others differing only in detail, and not in principle. 
 
 The still usually consists of a series of superimposed 
 sections, one mounted upon the other, constructed of cast 
 iron, each section being provided with a steam passage in 
 the bottom covered with a hood or bell, and an internal 
 overflow for liquor. These sections should be liberally pro- 
 vided with cleaning holes, by means of which access can be 
 had to every part of the interior, so that the apparatus can 
 be thoroughly cleaned without dismantling it. The liquor, 
 having been previously heated, enters at the top of the still 
 and flows from section to section in a direction opposite to 
 that of the steam, which enters at the bottom, passes up 
 through the steam passages, and is caused to pass through 
 the liquor by means of the hoods or bells. 
 
 The crude liquor is thus gradually brought to the boiling 
 point, the free ammonia and other gases mixing with the 
 steam in the upper portion of the apparatus. The milk of 
 lime is introduced in small quantities into the lower, or 
 liming sections, the hoods of these sections having a some- 
 what deeper seal than the others, in order to permit of more 
 efficient mixing of lime and liquor. A final boiling of the 
 liquor in the lower sections completely liberates all of the 
 fixed ammonia, and the stills should be so operated that 
 
AMMONIA 
 
 95 
 
 Lime Milk 
 
 the waste liquors leaving the bottom should not contain more 
 than 0.005 per cent of ammonia. 
 
 Great precaution must be taken in the operation of the 
 ammonia stills, as it is at this point where a large amount 
 StoomOc NHjVfapor °f the ammonia produced may 
 be lost, due to improper or in- 
 efficient liming for the liberation 
 of the fixed am- 
 Anmcnio Liquor monia, the liming 
 apparatus being 
 the subject of a good deal of 
 attention if proper pecuniary 
 returns are expected, and, be- 
 sides maintaining a constant 
 flow of liquor to the stills, it is 
 essential to see that a constant 
 feed of milk of lime is secured, 
 and that the gravity of the 
 solution, or its alkali content, 
 is constantly maintained at pre- 
 determined figures based upon 
 the quantity of fixed 
 
 ammonia in the liquor, 
 staarn.. 
 
 Concentrated Ammonia: 
 
 A plant for concentrat- 
 ing the liquor 
 f ToWaste Pit. produced in 
 the works is 
 shown in Fig. 
 
 19, and it consists of a still, as described above, a reflux cooler, 
 automatic waste lime overflow, lime washer, sectional cooler, 
 collecting tank for the storage of concentrated liquor, automatic 
 steam lime pump with agitator tank, iron hand pump and an 
 overhead supply tank. 
 
 The crude liquor is delivered from the overhead supply 
 tank (1) to the sectional cooler (2), where the vapors coming 
 from the still are cooled, thus transferring their heat to the 
 incoming crude liquor, and thereby reducing the amount of 
 cooling water required. The crude liquor enters the bottom 
 of the cooler, leaving it at the top, and passing from thence 
 
 Fig. 18. — Ammonia Still. 
 
96 
 
 COAL GAS RESIDUALS 
 
 to the upper portion of the still (3). As described above, 
 the still consists of several lime and several water sections, 
 the free ammonia being volatilized in the water sections. 
 
 
 The lime enters the upper liming secnon, being pumped from 
 the agitator tank (4) by means of the automatic lime pump 
 (5); this pump is so controlled by means of an air brake 
 (compressed air being supplied by means of a small air pump 
 
AMMONIA 97 
 
 not shown) that the delivery of milk of lime can be varied 
 within very wide limits, or from one to ten double strokes per 
 minute. The lime milk is agitated in the tank by means of 
 an attachment connected by a lever to the piston rod of the 
 pump, thus ensuring a perfect mixture of lime and water. 
 
 In the first lime section the lime becomes thoroughly 
 mixed with the liquor, being then passed through the remain- 
 ing sections and leaving the still at the bottom through the 
 automatic waste lime overflow. This overflow valve is 
 mounted in the bottom of a cast-iron casing, a lead-covered 
 sheet-iron float being connected to the valve by means of a 
 stem. The float moves up and down inside the casing, and 
 the steam pressure acting on it is the same as that in the 
 lower part of the still, so that the water level at which the 
 valve opens always remains constant. 
 
 The volatile ammonia is expelled through prolonged boil- 
 ing of the liquor, but the fixed ammonia can only be ex- 
 pelled by decomposing the salts with lime. The vapors 
 rise through the tubes in the bottoms of the sections, being 
 baffled by means of the hoods or bells located over the tubes, 
 while the raw liquor, as well as that portion of the liquid 
 condensed by the cooling action of the air, gradually flows 
 through the internal overflow tubes from one section to the 
 other. The frequent repetition of this cycle in the different 
 sections of the still produces ammonia vapors of high strength. 
 
 In the bottom section of the still the waste liquor is freed 
 of practically the last traces of ammonia. The ammonia 
 vapors are now cooled in the water-cooled top section of the 
 still and they are then dephlegmated in the reflux cooler (7), 
 thence passing into the lime washer (8), and are finally 
 condensed in the sectional cooler (2). The reflux cooler is 
 arranged over the still and can be fed with more or less 
 water as required for the purpose of condensing a large amount 
 of vapors and sending them back into the still in order to 
 produce a concentrated liquor of high NH 3 content, and this 
 liquor usually contains from 15 to 25 per cent of ammonia, 
 dependent upon the manner of operating the plant. 
 
 As the formation of ammonium carbonate may obstruct 
 the condenser, or sectional cooler, the vapors from the still 
 
98 COAL GAS RESIDUALS 
 
 are first passed through the lime washer (8), in which suffi- 
 cient lime should be injected to prevent any excessive forma- 
 tion of carbonate. The lime milk is injected into the washer 
 by means of the automatic pump (5), the delivery pipe from 
 this pump being provided with a proportional distributing 
 valve system, so arranged that two successive strokes of 
 the pump will deliver the lime milk into the still, and the 
 following stroke will deliver to the lime washer. The lime 
 milk from the washer in turn flows into the first liming sec- 
 tion of the still, where it gives up the ammonia absorbed 
 in the washer and mixes with the lime direct from the pump. 
 
 The sectional cooler (2) is made of cast iron, the upper 
 portion of the cooler being supplied with the necessary 
 amount of cooling water, this amount being considerably 
 less than that used in the usual cooler, because the principal 
 cooling medium consists of the crude ammonia liquor coming 
 from the supply tank (1). 
 
 This entire process for producing concentrated liquor is 
 practically automatic, it being only necessary to carefully 
 attend to the filling of the crude liquor supply tank (1) so 
 as to ensure a proper feed to the still, and to see that lime 
 is delivered at proper fixed intervals. 
 
 Concentrating plants as erected and operated in America 
 vary a great deal in detail, but the method of operation as 
 given above will cover almost any system in use, as the 
 principle of concentration is a fixed one. 
 
 The heating of the ammoniacal liquor, and bringing it up 
 to the boiling point, requires the greater amount of steam 
 used, the remainder of the steam being required for heating 
 the water in the lime milk, for splitting up and expelling the 
 ammonia, for replacing the amount of heat lost by radia- 
 tion, and for supplying that portion condensed to carry the 
 ammonia in solution. 
 
 Unger (John S. Unger, Chicago) gives the amount of 
 steam required in his 30-inch still, having 186 square feet 
 of radiating surface, radiating 200 B.T.U. per square foot 
 per hour, with steam containing 960 B.T.U. per pound, as 
 follows: 
 
 Steam required per pound of ammonia to concentrate a 
 liquor containing 1 per cent of ammonia, 
 
AMMONIA 
 
 99 
 
 Item 
 
 B.T.U. 
 
 Steam 
 lbs. 
 
 Per cent 
 
 One pound of ammonia requires for dissociation . . . 
 To heat 100 pounds of liquor from 70° to 220° F. . 
 To heat 4.2 pounds water with the lime 
 
 1,880 
 
 15,000 
 
 630 
 
 1,860 
 
 1.97 
 15.63 
 0.66 
 1.94 
 3.33 
 
 8.37 
 
 66.43 
 
 2.80 
 
 Heat radiated in three minutes 
 
 Pounds of water in concentrate 
 
 8.25 
 14.15 
 
 Total steam required per lb. of NH» 
 
 or 1.96 lbs. of steam per gallon of liquor 
 
 23.53 
 
 100.00 
 
 Steam required per pound of ammonia to concentrate a 
 liquor containing 2 per cent of ammonia, 
 
 Item 
 
 B. T. U. 
 
 Steam 
 lbs. 
 
 Per cent 
 
 One pound of ammonia requires for dissociation . . 
 To heat 50 lbs. of liquor from 70° to 220° F 
 
 1,880 
 
 7,500 
 
 630 
 
 930 
 
 1.97 
 7.81 
 0.66 
 0.97 
 3.33 
 
 13.35 
 
 53.02 
 
 4.48 
 
 Heat radiated during one and one-half minutes . . . 
 
 6.58 
 22.57 
 
 
 
 Total steam required per pound of NH 8 
 
 or 2.52 pounds of steam per gallon of liquor. 
 
 14.74 
 
 100.00 
 
 The above figures indicate that it requires eight and 
 three-quarter pounds more steam, per pound of ammonia, 
 to concentrate a 1 per cent liquor than is required with a 
 2 per cent liquor. 
 
 Lime Required: 
 
 One and sixty-five hundredths pounds of lime are theo- 
 retically required to replace one pound of ammonia, but as 
 the usual commercial lime contains some impurities, this 
 amount is usually increased to 2.5 pounds, which latter also 
 allows for some excess. This lime is only required for the 
 fixed ammonia, and if five pounds of ammonia containing 
 25 per cent of fixed salts should be recovered per ton of 
 coal, we would require (5 X 0.25) 2.5 = 3.125 pounds of lime 
 per ton of coal. 
 
100 COAL GAS RESIDUALS 
 
 Cooling Water: 
 
 About nine gallons of cooling water at a temperature 
 of 60° F. are required as a maximum, but in well constructed 
 and operated plants not more than seven gallons of water 
 per pound of ammonia should be used. 
 
 EARNING CAPACITY OF A CONCENTRATING PLANT 
 In the following example a plant carbonizing 300 short 
 tons of coal per day will be used as a basis, or one producing 
 12,000 gallons of 8 oz., or 2.0728 per cent liquor per day, 
 which with 250 maximum working days will give 3,000,000 
 gallons of liquor per year, recovering 5 pounds of ammonia 
 per ton of coal. 
 
 As we will assume that the plant is constructed as 
 described above, where the crude liquor is preheated in 
 the sectional cooler by means of. the concentrate proper, 
 only about 20 pounds of steam per pound of ammonia in a 
 1 per cent liquor, and about 12 pounds of steam in a 2 per 
 cent liquor, will be required. 
 
 I. Labor: 
 
 300 operating days, 3 men at $2.00 per day $1800.00 
 
 20 days cleaning waste pits, 1 man at $1.50 30.00 
 
 Total labor $1830.00 
 
 II. Materials: 
 
 4500 M pounds of steam at 20 cents $ 900.00 
 
 117 tons of lime at $8.50 994.50 
 
 2.75 million gallons of water at $10.00 27.50 
 
 Total materials ' $1922.00 
 
 III. Miscellaneous: 
 
 Drayage on about 100 tons of sediment at 50ji $ 50.00 
 
 Illuminating the plant 60.00 
 
 Estimated repairs 200.00 
 
 Lubricating oil, waste, etc 150.00 
 
 Total Miscellaneous $460.00 
 
 The total cost per year will then be: 
 
 I. Labor $1830.00 
 
 II. Materials 1922.00 
 
 III. Miscellaneous 460.00 
 
 Total $4212.00 
 
AMMONIA 101 
 
 making the total cost of concentrating one pound of ammo- 
 nia, less interest and depreciation on plant, 1.123 cents, while 
 the gross revenue at eight cents per pound will be 
 
 375000 X $0.08 = $30,000.00, 
 
 giving a net revenue, less interest and depreciation, of 
 
 $30,000.00 - $4,212.00 = $25,788.00 per year. 
 
 If the weak liquor, as produced in the works, can be sold as 
 such, and if the ammonia in this liquor based on the above 
 price of concentrate is worth six cents a pound, the yearly 
 gross revenue for this ammonia would be $22,500, but from this 
 would have to be deducted the added freight due to shipping a 
 larger bulk, if the ammonia liquor is to be transported away 
 from the works, as well as pumping charges. In order to make 
 the sale of weak liquor a profitable undertaking as compared 
 with concentrate, it would be necessary that an ammonia works 
 which would take the entire output of liquor be established in 
 very close proximity to the gas-works. 
 
 Aqua Ammonia: 
 
 This plant, a diagram of which is shown in Fig. 20, consists 
 of an ammonia still (3), provided with a reflux cooler (7); 
 automatic lime overflow (6); lime washers (8); coolers (9), 
 the latter consisting of wooden vats containing lead-cooling 
 coils; overhead supply tank (1); liquor heater (2); lime 
 tank (4); automatic lime pump (5); charcoal filters (10); 
 oil washer (11); caustic soda washer (12); absorbers (13); 
 and lime trap (14). 
 
 Chemically pure ammonia is a concentrated solution of 
 caustic ammonia in water, and can be made directly from 
 the gas liquor; it is a colorless, clear liquid and should not 
 expel any empyreumatic odors, while a lead-paper test should 
 show the entire absence of sulphides. 
 
 Aqua ammonia is stored in carboys or drums, neither of 
 which should be filled to capacity, and they must be kept 
 in a cool place to avoid bursting due to excessive gas pres- 
 sure. If drums are used for storage or transportation, care 
 should be taken to caulk them tightly in order to avoid loss 
 of ammonia by evaporation. 
 
102 COAL GAS RESIDUALS 
 
 The manufacture of aqua ammonia in some localities is 
 very remunerative, as the demand for this product is very 
 large in the textile industry as well as in refrigeration, but 
 the plant is also at times subject to large losses. 
 
 The crude liquor from the supply tank (1) enters the 
 heater (2) by gravity, the crude liquor being heated here 
 through the medium of the waste liquors coming from the 
 still; from the heater the liquor enters the top of the still 
 (3), where it is treated with steam and lime as in the pre- 
 vious case. The offgoing vapors are cooled in the reflux 
 cooler (7) and then enter into the lime washers (8), pass- 
 ing from one to the other, and from thence into the lime 
 trap (14), the vapors being thus deprived of all lime 
 water. 
 
 After leaving the trap the vapors enter the coolers (9), 
 going from thence into the first charcoal .filter (10), where 
 the moisture is expelled, and then into the oil washer (11), 
 where tarry matters are removed. The vapors are now 
 passed through the remaining charcoal filters in order to 
 expel all empyreumatic substances, and from thence into 
 the caustic soda washer (12), this latter washer being filled 
 with a 10 per cent solution of caustic soda for the purpose 
 of arresting any sulphides which may be contained in the 
 vapors, and the purified vapors then finally enter the sat- 
 urators or absorbers (13). 
 
 This process of absorption produces heat, and the re- 
 ceivers of the absorbers are therefore kept cool by being 
 placed in wooden vats filled with water. 
 
 THE EARNING CAPACITY OF AN AQUA AMMONIA PLANT 
 
 Capacity of plant at 300 tons of coal per day and 250 
 maximum working days, giving 3,000,000 gallons of liquor 
 per year. 
 
 J. Labor: 
 
 300 operating days, 3 men at $2.00 $1800.00 
 
 20 days cleaning waste pits, 1 man at $1.50 30.00 
 
 300 days liming, cleaning drums, loading, etc., 4 men at $1.50 1800.00 
 
 Total labor $3630.00 
 
1 Overhead Supply Tank. 
 
 2 Liquor Heater. 
 8 Ammonia Still. 
 
 4 Milk of Lime Tank. 
 
 5 Automatic Steam Lime Pump 
 
 6 Automatic Waste Lime Valve 
 
 7 Reflux Cooler 
 
 8 Lime Washer. 
 
 9 Cooler. 
 
 10 Charcoal filters. 
 
 11 Oil WASHER. 
 
 12 caustic Soda Washer. 
 
 13 absorbers. 
 
 14 Lime Trap 
 
 15 Waste Pits. 
 
 Fig. 20. — AQUA AMMONIA PLANT 
 
 (Facing Page 101) 
 
AMMONIA 103 
 
 //. Materials: 
 
 17,500 M. pounds of steam at 20^ $3500.00 
 
 450 tons of lime at 88.50 3825.00 
 
 1500 bushels of charcoal at 15^ 225.00 
 
 6 million gallons of water at $10.00 6000 
 
 Total material $7610.00 
 
 III. Miscellaneous: 
 
 12 gallon carboys, 12,500 at $1.60 $20,000.00 
 
 Drayage, 550 tons of sediment at 50ji 275.00 
 
 Illuminating the plant 65.00 
 
 Estimated repairs 750.00 
 
 Oil, waste, etc 200.00 
 
 Total miscellaneous $21,290.00 
 
 The total cost per year will then be: 
 
 I. Labor $ 3,630.00 
 
 II. Materials 7,610.00 
 
 III. Miscellaneous 21,29000 
 
 Total $32,530.00 
 
 The gross revenue to be derived will amount to: 
 780 tons of Aqua of 26° Be\, equivalent 
 
 to about 27.2% NH 3 , at $102.00 $79,560.00 
 
 and the net revenue will be $79,560.00 - $32,530.00 = 
 $47,030.00 per year, less depreciation and interest. This 
 revenue should be increased after the first year, because a 
 number of the carboys will be returned to the plant, thus 
 making their cost less than originally, only a nominal allow- 
 ance being made to the purchaser for them. 
 
 INDIRECT SULPHATE PROCESS 
 
 This plant consists of an overhead supply tank (1); a 
 liquor heater (2); ammonia still (3); lime agitating tank 
 (4); automatic lime pump (5); automatic overflow (6); 
 water trap or separator (7); saturators (8); acid catch box 
 (9); and drainage table (10), all as shown in Fig. 21. 
 
 Here again, the principle of design remains the same, but 
 the various manufacturers have subjected the various parts of 
 the plant to variations complying with individual experience. 
 
 The liquor from the supply tank flows into the preheater, 
 where it is heated by the escaping waste liquors before they 
 enter the waste pits, the heated liquor then passing into the 
 upper portion of the still, where the volatile ammonia is 
 
1 104 
 
 COAL GAS RESIDUALS 
 
AMMONIA 105 
 
 driven off, the fixed ammonia being treated with lime in the 
 lower sections as described above. The water separator (7) 
 serves to remove any condensed water from the ammonia 
 gases, this condensate being returned to the still. The 
 ammonia vapors now enter the saturator, this apparatus 
 being lead-lined throughout, and are here brought in contact 
 with sulphuric acid, whereby the ammonium salts are formed. 
 These saturators are so constructed that they may be op- 
 erated intermittently, the process being thereby interrupted 
 while one saturator is being cleared of salt as the ammonia 
 is diverted into the other, which has been prepared for the 
 purpose in the meantime, or the process may be made con- 
 tinuous by entirely closing the saturator, feeding it regularly 
 with acid, and removing the salts onto the draining table 
 by means of a steam or air ejector. 
 
 The sulphuric acid employed is of 60° to 66° Be\, and in 
 large works it is usually transferred by means of compressed 
 air and a lead-lined monteju. Passing the sulphate from the 
 drainage table through centrifugals dries it so that the salt 
 contains practically no moisture, and it can be bagged for 
 immediate shipment. 
 
 The indirect plant of C. and W. Walker, London, embraces 
 a still column of square cross-section, the upper section being 
 for free ammonia, the intermediate section being the liming 
 chamber, while the lower sections take care of the fixed ammonia. 
 The crude liquor flows from an overhead supply tank through 
 the tubes of a vertical tubular heater, Fig. 21a, and there absorbs 
 heat from the waste gases coming from the saturator. The 
 tubes in the heater are provided with expansion joints at either 
 end. The lime is mixed with water in a mixing tank provided 
 with mechanical agitators, these agitators being driven from 
 the pump which injects the lime milk into the still. The lower 
 section of the still, as well as the liming chamber, is provided 
 with perforated steam coils, steam being supplied to these 
 coils through a reducing valve which maintains the desired 
 pressure. Steam is admitted into the still through the lower 
 coil, while the one in the liming chamber is used primarily to 
 agitate the lime and thus cause it to mix properly with the 
 liquor. The ammonia vapors coming from the still pass through 
 a baffle box, where condensed water is thrown out, and then 
 through an inverted U leg into the saturator. This saturator is 
 
106 
 
 COAL GAS RESIDUALS 
 
 cylindrical in form, made of cast iron, with cone bottom lined 
 throughout with lead, and is supported on four columns above 
 the drain board. The ammonia inlet and the waste gas outlet 
 are both located on the flat top of the saturator. The center 
 of the top is provided with a clean-out opening of a diameter 
 large enough to permit a man to enter the saturator to make 
 repairs, and a lead curtain pipe depends from this opening and 
 dips into the acid in the saturator, thus permitting access to the 
 latter at any time. The saturator discharges through a special 
 
 A suPPi-V TANtc 
 
 B HEATEC 
 
 C Still 
 
 D LIME MlXEia. 
 E SATUBATOO. 
 
 F Drain DoAisft [ 
 
 1 Fig. 21a. — Walker Sulphate Plant. 
 
 discharge valve made of copper, the valve consisting of a casing 
 in which is mounted a plunger capable of a forward and back- 
 ward motion through the medium of a lever. The discharge 
 sulphate and the mother liquor are received on the drain board, 
 the liquor being drained off to a cistern, while the sulphate is 
 carried to a centrifugal dryer. It is claimed that the waste 
 liquor from such a still contains less than 0.006 per cent of NH 3 . 
 In the Wilton apparatus, also of English design, we find the 
 raw liquor being run into preheaters, and from there to the 
 stills, where they enter at a temperature of about 187° F., and 
 at which point all volatile ammonia is liberated. The pre- 
 heaters are tubular, made of cast iron, and heated by the ex- 
 
AMMONIA 107 
 
 haust gases from the saturators, these gases being almost 
 entirely condensed by the time they leave the last heater. 
 The construction of the still permits of easy access for cleaning, 
 and the saturator, like that of Walker, is completely inclosed. 
 It is of interest to note the method adopted for the removal of 
 the salt, which is accomplished by means of a steam ejector. 
 The bottom of the saturator is perfectly flat, except for a 
 small sump in the center, in which the ejector is located. The 
 NH 3 vapors enter the saturator through two pipes, passing 
 vertically through the top, and entering the center of two 
 crescent shaped perforated pipes which surround the well. The 
 perforations are arranged in such manner that the vapor shoots 
 downward towards the well, thus blowing all the salt into the 
 well, and at the same time thoroughly mixing and saturating 
 the acid bath. The mother liquor enters from a mother liquor 
 pot and falls to the bottom of the saturator immediately be- 
 hind the perforated pipes. The saturator is operated in the 
 same manner as any other continuous saturator, the bath being 
 maintained at about 32°-33° Be\ by a constant stream of 
 acid flowing in, and the ejector may be kept in constant opera- 
 tion removing the salt, or the salt may be permitted to fill the 
 well up to the spray pipes, when the ejector is set in operation 
 intermittently. A bell and vent pipe are provided toward the 
 discharge end of the ejector pipe to carry off the steam in order 
 to prevent its escape into the house. After being ejected onto 
 the drain-board the salt is dried in the usual manner. 
 
 EARNING CAPACITY OF AN INDIRECT SULPHATE PLANT 
 
 Capacity of the plant at 300 tons of coal per day, and 250 
 maximum working days, which at 5 pounds of ammonin per ton 
 of coal will produce 727 tons of sulphate per year. 
 
 J. Labor: 
 
 300 working days, 3 men at $2.00 $1800.00 
 
 20 days cleaning pits, 1 man at $1.50 30.00 
 
 Total labor $1830.00 
 
 //. Materials: 
 
 4500 M. pounds of steam at 20(! $900.00 
 
 1 13 tons of lime at $8.50 960.50 
 
 727 tons of sulphuric acid at $11.50 8,360.50 
 
 Total material $10,221.00 
 
 ///. Miscellaneous: 
 
 Drayage on 100 tons of sediment at 50 cents $50.00 
 
 Illuminating the plant ^.'9^ 
 
 Estimated repairs 50 0.00 
 
 Oil, waste, etc 150-00 
 
 Total miscellaneous $760.00 
 
108 COAL GAS RESIDUALS 
 
 The total cost per year will then be: 
 
 I. Labor $ 1,830.00 
 
 II. Material 10,221.00 
 
 III. Miscellaneous 760.00 
 
 Total $12,811.00 
 
 making the cost of producing one ton of sulphate equal 
 $17.62, and the net yearly income with a sulphate value of 
 727 tons at $60.00, equal to $43,620.00, will be $43,620.00 - 
 $12,811.00 = $30,809.00, less depreciation and interest. 
 
 The unit charges given in the above three examples are 
 of course dependent upon local conditions, and the quantity 
 of materials used are the theoretical amounts; these condi- 
 tions may change in each individual case, somewhat altering 
 the net results. 
 
 The successful operation of a sulphate plant is dependent 
 upon close attention to the character of the crude liquor, to 
 the sulphuric acid, and to the lime used. While the chemi- 
 cal process is a very simple one, it is very sensitive to slight 
 variations from the normal conditions, and derangements 
 are therefore possible unless the required attention is given 
 to the raw materials. 
 
 Clean liquor is a most important requisite, and this is 
 especially true if closed saturators are used, the latter re- 
 quiring that the crude liquor be entirely free of light oils 
 and tar, or nearly so, this foreign admixture causing heavy 
 frothing inside the saturator with a consequent deposition 
 of a hard crust on the lead walls, which deposition leads to 
 stoppages necessitating a shut-down and a complete cleaning 
 of the plant. 
 
 In the purchase of sulphuric acid attention should be 
 drawn to its content of iron and arsenic, as the acid should 
 be as free of these foreign constituents as possible. The 
 acid under no condition should be accepted if it contains 
 more than 10 to 12 grains of iron per gallon, and the arsenic 
 should not exceed 15 parts per million; besides this, the acid 
 storage tanks should be subject to periodical cleaning. The 
 lime used should also be as free from foreign matter as 
 possible, and should analyze 95 per cent CaO. 
 
 The results of investigations made by K. Leo, and reported 
 to the German Coke Commission, on a series of tests made 
 
AMMONIA 
 
 109 
 
 on the sulphuric acid used in producing the corresponding 
 ammonium sulphates, is given in Table XIII, 60° Be\ acid 
 having been used. 
 
 TABLE XIII. — INFLUENCE OF FOREIGN MATTER ON COLOR 
 OF AMMONIUM SULPHATE 
 
 Content of 
 foreign matter 
 
 Color of salt 
 
 Tar oil + small quan- 
 tity of tar, about 
 1.2% 
 
 Wash oil, but no tar, 
 1.2% 
 
 Wash oil + 0.01 % . . 
 
 Dark gray-brown 
 White, with red- 
 dish tinge 
 
 White 
 
 Cu . . . . 
 0.01 % Cu. 
 
 0.03% Cu 
 
 0.05 %Cu 
 
 0.03%Cu+0.03%As 
 0.01% As 
 
 0.03 % As 
 
 0.05 % As 
 
 0.03% As + 0.01 %Cu 
 
 0.03% As + 0.03 %Cu 
 0.03% As + 0.03 %Pb 
 
 0.01% Cd. 
 
 0.03% Cd. 
 0.05 %Cd. 
 
 White 
 
 White, with light 
 grayish tinge . . 
 
 White, with gray 
 ish tinge 
 
 White, with pale 
 grayish tinge 
 
 White, with pale 
 light yellow 
 tinge 
 
 Content of 
 foreign matter 
 
 0.075% Cd 
 
 0.03 % Cd + 0.01 % 
 
 Cu 
 
 0.03 % Cd + 0.03 % 
 
 Cu 
 
 0.03 % Cd 4- 0.03 % 
 
 Pb 
 
 Light yellow . 
 Light yellow. 
 Red 
 
 Violet gray 
 Light yellow. 
 
 03% Pb 
 
 0.01% Cu (CuS at 
 point of oxidation) 
 
 0.03 % Cu (CuS at 
 point of oxidation) 
 
 0.03 %Cu (CuS part- 
 ly oxidized) 
 
 0.03% Cu from a 
 strong acid bath 
 
 Color of salt 
 
 White, with 
 weak pale 
 yellow tinge 
 
 Light, reddish 
 yellow 
 
 Light, reddish 
 yellow 
 
 Reddish yellow 
 
 Light gray 
 
 Light gray 
 Light, reddish 
 yellow 
 
 White 
 White with 
 
 pale gray 
 
 tinge 
 
 Light gray 
 White with 
 grayish tinge 
 
 Light gray 
 
 0.04% Pb from a 
 strong acid bath . . 
 
 White 
 
 As stated before, ammonia is one of the principal by- 
 products as well as source of revenue to the coal-gas pro- 
 ducer, but as a rule the usual course of procedure is to 
 manufacture concentrated liquor, only a few of the larger 
 works having taken up the manufacture of ammonium 
 sulphate. The competition of Chili saltpeter with sulphate 
 
110 COAL GAS RESIDUALS 
 
 cannot last much longer, as these beds are gradually becom- 
 ing depleted and the future will demand that this natural 
 product be displaced by ammonium compounds. 
 
 As regards ammonia in liquid form, it may be prophesied 
 that synthetic ammonia will in the near future be an active 
 competitor, and if Germany, which produces about 40 per 
 cent of the total ammonium sulphate made in England, Ger- 
 many, and the United States, be taken as an example, we 
 find that they work up about 22,000 tons of liquid ammonia 
 per year, all of which can readily be replaced by synthetic 
 ammonia at the present moment. 
 
 The drop in ammonia prices in Germany has been a cause 
 of uneasiness for some time, and as synthetic ammonia con- 
 tinues to make inroads into the market this drop will prob- 
 ably become more marked, and difficulty will be experienced 
 in finding a favorable market for this product; this condition 
 will in all probability lead to a greater production of sulphate, 
 the market for which is gradually increasing, as is witnessed 
 by the fact that we are still importing quite an appreciable 
 quantity of this product from Europe. Franz Buchner, Engi- 
 neer of the Economic Association of German Gas Companies, 
 advises his associates that synthetic ammonia production will 
 necessarily lead to an increased sulphate production in gas- 
 works practice, and recommends that smaller companies shall 
 pool their ammonia liquor at some central point where sul- 
 phate can be prepared at less expense than if this were done 
 by the individual. Table XIV shows the production and con- 
 sumption of ammonium sulphate in the United States from 
 1902 to 1913: the amounts being expressed in tons of 2000 
 pounds, while Table XVI gives the world's sulphate production 
 in tons. 1 
 
 These figures show an increased production of 158,876 
 tons in the year 1913 over the year 1902, or an increase of 
 440 per cent; the increase in imports was largest in 1911, 
 having been subject to a gradual increase, but this has 
 declined since then; the consumption shows an increase of 
 208,366 tons in the year 1913 over that of 1902, an increase 
 of 384 per cent. It is well to note here that in the figures 
 
 1 American Coal Products Company. 
 
AMMONIA 
 
 111 
 
 TABLE XIV.— PRODUCTION AND CONSUMPTION OF SULPHATE 
 IN THE UNITED STATES 
 
 
 Production 
 
 Imported 
 
 
 Year 
 
 
 
 Consumption 
 
 
 By product 
 
 Coal gas and 
 
 
 
 
 coke ovens 
 
 carbonization 
 
 
 
 1902 
 
 18,483 
 
 17,641 ' 
 
 18,146 
 
 54,270 
 
 1903 
 
 24,098 
 
 17,775 l 
 
 16,777 
 
 58,650 
 
 1904 
 
 32,653 
 
 22,011 » 
 
 16,667 
 
 71,331 
 
 1905 
 
 41,864 
 
 23,432 
 
 15,288 
 
 80,584 
 
 1906 
 
 75,000 ' 
 
 9,182 
 
 84,182 
 
 1907 
 
 62,700 
 
 36,609 
 
 32,669 
 
 131,978 
 
 1908 
 
 50,073 
 
 33,327 
 
 34,274 
 
 117,674 
 
 1909 
 
 75,000 
 
 31,500 ' 
 
 40,192 
 
 146,692 
 
 1910 
 
 86,000 
 
 30,000 i 
 
 63,178 
 
 179,178 
 
 1911 
 
 95,000 
 
 32,000 ' 
 
 103,743 
 
 230,743 
 
 1912 
 
 130,000 
 
 35,000 ' 
 
 81,188 
 
 246,188 
 
 1913 
 
 153,000 
 
 42,000 l 
 
 67,636 
 
 262,636 
 
 Estimated. 
 
 TABLE XV. — SULPHATE OF AMMONIA CONSUMPTION BY 
 
 COUNTRIES 
 
 Year 
 
 United 
 Kingdom. 
 Gross tons 
 
 Germany. 
 Metric tons 
 
 United States. 
 Net tons 
 
 France. 
 Metric tons 
 
 1900 
 
 65,000 
 
 127,700 
 
 36,011 
 
 49,000 
 
 1901 
 
 68,297 
 
 168,000 
 
 43,756 
 
 43,300 
 
 1902 
 
 63,750 
 
 175,000 
 
 54,270 
 
 53,300 
 
 1903 
 
 71,700 
 
 172,500 
 
 58,650 
 
 53,900 
 
 1904 
 
 68,500 
 
 200,000 
 
 71,331 
 
 53,800 
 
 1905 
 
 75,000 
 
 213,000 
 
 80,584 
 
 56,380 
 
 1906 
 
 82,000 
 
 229,500 
 
 84,182 
 
 67,000 
 
 1907 
 
 87,500 
 
 240,000 
 
 131,978 
 
 81,000 
 
 1908 
 
 90,000 
 
 260,000 
 
 117,674 
 
 84,000 
 
 1909 
 
 87,000 
 
 274,000 
 
 146,692 
 
 78,500 
 
 1910 
 
 87,000 
 
 350,000 
 
 179,178 
 
 73,086 
 
 1911 
 
 89,000 
 
 370,000 
 
 230,743 
 
 84,026 
 
 1912 
 
 90,000 
 
 425,000 
 
 246,188 
 
 
 1913 
 
 97,000 
 
 435,000 , 
 
 262,636 
 
 
112 
 
 COAL GAS RESIDUALS 
 
 covering importations, the sulphate equivalent of the am- 
 monia imported as chloride is included for the years 1912 
 and 1913. 
 
 TABLE XVI. — WORLD'S PRODUCTION OF SULPHATE IN TONS 
 
 Country 
 
 1905 
 
 1906 
 
 1907 
 
 1908 
 
 1909 
 
 1910 
 
 1911 
 
 1912 
 
 1913 
 
 England 
 
 Germany .... 
 United States . 
 
 France 
 
 Belgium- Hol- 
 land 
 
 Spain 
 
 Italy 
 
 Other countries 
 
 273,550 
 190,000 
 59,250 
 47,300 
 
 24,200 
 
 10,000 
 
 4,500 
 
 40,500 
 
 294,170 
 
 235,000 
 
 68,000 
 
 49,100 
 
 30,000 
 
 10,000 
 
 5,000 
 
 40,000 
 
 318,400 
 
 287,000 
 
 90,120 
 
 52,720 
 
 1 55,000 
 
 2 12,000 
 
 11,000 
 
 65,000 
 
 330,450 
 
 313,000 
 
 79,500 
 
 52,600 
 
 35,000 
 
 80,000 
 
 354,747 
 
 330,000 
 
 96,600 
 
 53,600 
 
 40,000 
 12,000 
 12,000 
 73,000 
 
 373,590 
 
 373,000 
 
 105,143 
 
 56,000 
 
 43,000 
 
 9,000 
 
 12,000 
 
 79,000 
 
 391,160 
 
 418,000 
 
 115,245 
 
 60,000 
 
 43,000 
 
 I 166,500 
 
 394,540 
 
 492,000 
 
 149,700 
 
 68,500 
 
 50,000 
 
 J 175,000 
 
 426,745 
 
 548,558 
 
 176,900 
 
 74,500 
 
 51,000 
 
 | 161,500 
 
 Including Norway and Sweden. 
 
 Including Portugal. 
 
 Table XV shows the sulphate of ammonia consumption 
 by countries, the amounts given including the equivalent of 
 sulphate in other forms of ammonia. 
 
 The importance of these figures to the American sulphate 
 producer can be seen when it is remembered that the total 
 area of Great Britain, Germany, France, Austria-Hungary, 
 Spain, Holland, and Belgium is about 992,180 square miles, 
 with about 672,550 square miles under cultivation, and in 
 the United States with its total of 3,026,800 square miles 
 only 746,012 square miles are under cultivation, and the 
 greater portion of the sulphate produced is used in fertilizer; 
 a great deal of the land under cultivation in the States is 
 still virgin soil, but this condition cannot continue, and 
 nitrogen will have to be returned to the soil to make it 
 productive. 
 
 In this connection it may also be well to state that in 
 1902 Japan imported 4000 tons of sulphate, while in 1907, 
 or five years later, this amount was increased to 64,000 tons. 
 
 Diagram II, prepared by the Barrett Company, shows the 
 price paid for ammonium sulphate from 1894 to the month of 
 May, 1916; from this it is seen that the average price of sulphate 
 during the twenty years prior to 1916 given in the diagram was 
 $58.60 per ton, for the last ten years of that period $61.28 per 
 ton, and for the last five years $61 per ton. 
 
Missing Page 
 
AMMONIA 113 
 
 During the last fifteen years the attention of investigators 
 has been turned towards an endeavor to simplify the methods 
 used to produce ammonium sulphate from the product due 
 to coal distillation, and instead of treating the gas liquor 
 itself, as in the old or indirect system just described, the 
 endeavor has been to pass the gas directly into and through 
 the saturator, which is a closed one in this case; this method 
 of procedure has been termed the "direct process" and, as 
 the name indicates, it is a reversal of the old method. Other 
 investigators have proceeded along somewhat different lines, 
 and while passing the gas through the closed saturator, they 
 at the same time vaporize the ammonia produced by previous 
 condensation and then pass these vapors into the saturator 
 in conjunction with the gas, this latter development being 
 known as the "semi-direct" process. 
 
 Other investigators, such as Burkheiser and Feld, make 
 use of the sulphur contained in the gas, thus doing away 
 with the saturator and its bath of acid, this being the only 
 real "direct" process, as it is entirely independent of the use 
 of sulphuric acid. While the Burkheiser process has been 
 tried in an experimental way at Tegel, Germany, it has so 
 far, to my knowledge, not been successful, and the experi- 
 ments at that plant have been abandoned. The principal 
 direct processes in active operation are those of Brunck, 
 Otto, and Feld, while those of the semi-direct procedure are 
 known as Koppers, Mont Cenis, Collin, and Still. 
 
 The first process on the "direct" principle was devised 
 by Franz Brunck, of Dortmund, Germany, and his success- 
 ful experiments began about fifteen years ago, making Brunck 
 the father of all of the "direct" or "semi-direct" systems 
 with the exception of the Feld. Brunck's first plant was 
 erected at the Kaiserstuhl works in Dortmund, being de- 
 signed to treat 2,750,000 cubic feet of coke-oven gas per 
 24 hours, the gas containing from 150 to 195 grains of am- 
 monia per 100 cubic feet at a temperature of from 80° to 
 95° F. Brunck passed this cooled gas through an acid washer, 
 the washer being filled with earthenware cubes measuring 
 about three inches to the side; these cubes were constantly 
 sprayed with sulphuric acid, the resultant liquor flowing 
 from the bottom of the washer being returned to the sprays 
 
114 COAL GAS RESIDUALS 
 
 until it possessed a certain degree of concentration, after 
 which the acid liquor was removed to the saturator where 
 the salt was formed. This plant was also provided with a 
 still for vaporizing the ammonia thrown down by previous 
 condensation, these ammonia vapors meeting the acid liquor 
 in the saturator. 
 
 This plant led to a further development by Brunck, and 
 his revised method of operation is as follows : — The hot, 
 raw gas is first passed through a temperature regulator, one 
 or more being used according to circumstances, and then 
 into a tar centrifugal; after the gas has thus been brought 
 to a temperature of from 255° to 265° F., and after about 
 80 per cent of its tar has been removed, it is passed into 
 the saturators of the ammonia plant. Here the ammonia 
 reacts with the sulphuric acid and a further 10 to 15 per 
 cent of the tar is extracted, after which the gas passes through 
 washer-scrubbers and then through volumetric condensers, 
 where the gas is sprayed for the removal of naphthalene, and 
 then on through tubular condensers and exhausters. While 
 this process seemed successful at the time, another develop- 
 ment was soon due, brought about by the introduction of 
 the closed saturator, as the open saturator permitted both 
 the gas and the liquor to cool too rapidly, this fact having 
 compelled Brunck to maintain the temperature of the gas 
 ahead of the saturator at about 255° F. 
 
 The application of the closed saturator permitted a lower 
 gas temperature, thereby bringing about a more efficient 
 extraction of tar, a temperature of only 185° F. ahead of the 
 saturator also simplifying the apparatus. 
 
 As stated above, Brunck's success soon led to develop- 
 ments by other investigators producing some direct and 
 some semi-direct processes, and one of the most successful 
 of the latter type is the 
 Koppers System: 
 
 The Koppers process (Fig. 22) is termed a semi-direct 
 one because the gases are freed of tar and some ammonia 
 by cooling before they enter the acid bath, the condensed 
 ammonia being afterwards vaporized. Koppers grounded 
 his theory on the hypothesis that the gas could not be freed 
 of tar by mechanical means only, and that cooling was 
 
AMMONIA 
 
 115 
 
 necessary in order to accomplish this, and having a tar- 
 free gas he was thus able to bring the dew point of the gas 
 to a rather low temperature. 
 
 AMMONA VAPOR 
 
 QMBT IQO-C 
 
 CoonwqWATKja 
 
 Coour*q wter 
 
 
 CU pp l a 1 - 
 
 Tfc« 0* AMMONIA TO Cl&TERMS 
 ~R-GOOL.tR.~T 
 
 ~) -EXHAUSTER 
 C-T(^e 5CP«R(=(IDR 
 
 D" fSt-HOITER.. 
 ""- SpCTUR/WOR. 
 
 f - DisTiuiMq Column. 
 
 Fig. 22. — Koppers Sulphate Plant. 
 
 In this process the gases coming from the hydraulic main 
 enter a cooler, where a portion of the tar, as well as con- 
 densed water, is removed; about one-half of the cooling 
 system is used to effect the temperature transfer between 
 the raw gas and the cooled, tar-free gas. After the gas has 
 passed through the last tubular water condenser it passes 
 into the temperature transfer apparatus, where it is pre- 
 heated and then removed by means of an exhauster and sent 
 into a Pelouze, where the last portion of the tar is removed 
 and the gas made practically tar free. 
 
 Koppers also includes a temperature regulator in his sys- 
 tem, locating it between the Pelouze and the saturator, this 
 apparatus also being a preheater, and its use is of great 
 benefit to the entire process. The gas is now admitted into 
 the saturator where it reacts with the acid in the usual 
 manner. The condensate from the coolers (tar and ammonia) 
 is run into a separator where, due to their respective specific 
 weights, the tar and liquor separate, the liquor being con- 
 ducted to a still where it is vaporized by means of steam 
 and lime in the usual manner. The ammonia vapors com- 
 ing from the still are not conducted directly into the sat- 
 urator, but to the gas main ahead of the coolers, as Koppers 
 considers this method of more value, and its adoption in 
 practice seems to have borne out this opinion. 
 
 The advantages of this process over the old or indirect one 
 
116 
 
 COAL GAS RESIDUALS 
 
 are self-evident; they consist in simplicity of operation and 
 ease of supervision, the saving of steam in the ammonia 
 plant, owing to the fact that it is not necessary to distil off 
 all of the ammonia contained in the gas, and the obviation 
 of all mechanical washers for removing ammonia from the 
 gas. 
 
 The saving of steam, however, only holds good for evap- 
 orative purposes, as steam or power of some kind is required 
 at the exhausters in addition to what would be required with 
 the indirect system; and this holds good for all of those 
 systems which admit the gas directly into closed saturators, 
 as they throw a back pressure of many more inches, due to 
 the depth of the bath, than would ordinarily be the case, 
 and to which reference will be made later. 1 
 
 The Collin Process: 
 
 The semi-direct process introduced by Collin of Dort- 
 mund, Germany, is shown in Fig. 23; this system is very 
 similar to that of Koppers, the specific difference between 
 the two being that Collin does not use a temperature trans- 
 fer apparatus nor a gas preheater between the tar extractor 
 and exhauster. 
 
 SEPWKATOB 
 
 Fig. 23. — Sulphate Plant — " Collin " System. 
 
 The gas coming from the hydraulic main is reduced in 
 temperature in the cooler to a degree which will permit of 
 efficient tar extraction in the Pelouze; the exhauster then 
 transfers the .tar-free gas to the saturator, where the am- 
 
 1 See "The Cleaning of Blast-Furnace Gases," McGraw-Hill Book 
 Co., Inc. 
 
AMMONIA 117 
 
 monia is absorbed by the acid bath. The condensate thrown 
 down in the cooler is distilled in the usual manner in vapor- 
 izing stills, and the ammonia vapor thus produced is sent 
 into the saturator in such manner that it cannot unite with 
 the gas coming from the hydraulic main. Due to the pres- 
 sure produced in the still, the waste gases formed there 
 pass through the acid bath by means of their own pressure 
 and are caught under the central bell of the saturator, from 
 whence they pass on to the stack, or join the fuel gases 
 going to the coke ovens in a coke plant. Collin claims that 
 this procedure prevents adding rulphur vapors to the gas. 
 Due to the separate admission of the vapors from the still 
 to the saturator bath it is possible to carry out the process 
 of sulphatation without heating the bath by external means. 
 
 Collin states that it is of distinct advantage not to carry 
 the troublesome waste gases away with the other gases, as 
 Koppers does, but to carry them directly to a stack, which 
 latter procedure is made possible by the construction of the 
 Collin saturator. 
 
 In another system devised by Collin, Fig. 23A, the gas 
 from the hydraulic main is caused to pass through tubular 
 water coolers where the greater portion of the tar and am- 
 monia is precipitated, after which it passes through a P. 
 and A. Condenser for the removal of the remainder of the 
 tar. After being thus freed of tar and ammonia, correspond- 
 ing to a temperature of 113° F., the gas being of 212° F. as 
 it enters the first coolers, the gas passes into a second set 
 of water coolers of similar construction to the first set, where 
 it is cooled to about 73° F., and from thence through the 
 exhausters to the saturators. 
 
 In this system Collin uses two saturators, set at different 
 levels in the building, the cleaned gas entering the upper 
 saturator at about 73° F. and leaving it at 100° F.; the 
 bath in this saturator is of about 25 per cent free acid, 
 the liquor from this bath being used as a saturating medium 
 in the lower saturator, the latter receiving the ammonia 
 vapors from the still. 
 
 Collin claims that this arrangement of saturators, or the 
 separation of the direct gas from the still vapors, eliminates 
 the obnoxious vapors of H 2 S, (CN) 2 , etc., from the system. 
 
118 
 
 COAL GAS RESIDUALS 
 
 The Still Process: 
 
 In this system the 
 gas passes the cool- 
 ing water in a coun- 
 ter current, coming 
 in direct contact 
 with it, thus being 
 cooled itself while 
 si m ul taneously 
 heating the cooling 
 water, the quantity 
 of the latter being 
 increased by the 
 amount of water 
 condensed from the 
 gas, this water being 
 used later to raise 
 the temperature of 
 the gas in the same 
 direct manner, and 
 Still claims that the 
 water condensed 
 from the gas is 
 thus vaporized back 
 into it. 
 
 This system is 
 shown in Fig. 24; 
 the gas from the 
 hydraulic main en- 
 ters the condenser 
 (A) at (a) with a 
 temperature of 
 176° F., where it 
 comes in direct con- 
 tact with a counter 
 stream of water 
 and is thus cooled. 
 This condenser is 
 of very large di- 
 mensions, in some 
 
AMMONIA 
 
 119 
 
 £t 
 
 1 
 
 31WSW3OWO0 OTOD 
 
120 COAL GAS RESIDUALS 
 
 cases being 10 feet in diameter and 82 feet high, the 
 lower portion (E) acting as a receiver and separator for 
 the products of condensation. The upper or cooling portion 
 of the tower is provided with a number of perforated plates 
 (6), the condensing water trickling through the holes and over 
 the edge of the plates, this water being circulated by means 
 of the pump (H) and pipe (c). The gas is thus cooled to 
 85° to 95° F., while the water is heated to approximately 
 the same temperature as that of the entering gas. Still 
 claims that this method of cooling removes the entire con- 
 tent of tar and naphthalene, as well as the greater portion of 
 the water vapor from the gas. 
 
 The warm water, tar, etc., collect in the receiver (E), 
 where they are separated due to their respective specific 
 gravities. The tar is run off by pipe (e) directly to the tar 
 storage, while the warm water passes through the syphon 
 (d) to the reheater (B). The cooled gases now pass through 
 pipe (/) to a second cooler (C), of the tubular type, where 
 they are further reduced in temperature to from 68° to 77° F. 
 This second cooler is for the purpose of maintaining the 
 final temperature of the gas at a constant point, and Still 
 claims that this second cooler is the regulator governing the 
 proper performance of the plant. 
 
 The gas leaves the second cooler through pipe (g) and is 
 forced by the exhauster (G) into the reheater (B); here the 
 gas is brought into direct contact with a most minute spray 
 of warm water, thus heating the gas to about 158° F., while 
 the water is cooled down to about 85° to 95° F., or approx- 
 imately to the temperature of the entering gas. The water 
 passes off into the storage tank (F), from whence it is again 
 pumped into the cooling tower (A), thus continuously ex- 
 tracting heat from the gas and in turn giving this heat back 
 again, a portion of the ammonia in the gas being thus ab- 
 sorbed also (Still claims 120 to 180 grains per gallon of 
 water), and as soon as this degree of concentration is reached 
 no more ammonia is absorbed, and the entire content of 
 ammonia remaining in the gas is carried through the re- 
 mainder of the process. A certain portion of the absorbed 
 ammonia is removed from time to time and evaporated by 
 means of steam and lime in the usual manner, the ammonia 
 
AMMONIA 
 
 121 
 
 vapors thus produced being mixed with the gas leaving the 
 reheater and passing on through pipe (K) into the saturator 
 (D), where the ammonia is absorbed by the sulphuric acid. 
 
 The temperature of the gas entering the saturator is 
 sufficiently high to prevent the precipitation of water vapor, 
 and thus permits of direct sulphatation without the addition 
 of external heat. The salt is ejected from the saturator 
 onto a draining table and then dried in the usual manner. 
 The gas leaves the saturator at a temperature of about 
 176° F. and then passes into the final coolers (K). 
 The Coppie Process: 
 
 In this process, shown diagrammatically in Fig. 25, the 
 
 q qcatoBenzdfttohor 
 LfromOvsna 
 
 'Serturotor 
 CjoSor 
 
 Fig. 25. — Copp6e Sulphate Plant. 
 
 A*SPlt 
 
 gas coming from the hydraulic main is cooled down in the 
 usual condensers to about 90° F., after which it is passed 
 through a Pelouze condenser for the extraction of the remain- 
 der of the tar. The tar-free gas is then passed directly into 
 the inclosed lead saturator (B) at (A), where the ammonia 
 combines with the acid. The ammonia liquor condensed 
 by the process of cooling is distilled in the still (C), the 
 ammonia vapors from the still being passed into the gas 
 main at (Z>) directly in front of the saturator, while the 
 heat of the reaction in the saturator prevents any further 
 condensation at that point. This process, like the others of 
 its class, is continuous and only requires attention when 
 fresh acid is needed and when sulphate is to be removed 
 and dried. The sulphate is ejected from the saturator in 
 the usual manner by either a steam or an air ejector at (E), 
 
122 
 
 COAL GAS RESIDUALS 
 
 thus delivering the salt to the draining table (F), from 
 whence it passes into the centrifugal and thence to storage. 
 This process claims that a reheating of the gas is not neces- 
 sary after it has once been cooled to 90° F. 
 
 The Mallet Process: 
 
 The Mallet system is one of the latest and it differs from 
 all other acid bath systems in several particulars, but closely 
 
 Fig. 26. — Mallet Sulphate Process. 
 
 resembles Brunck's early acid washer process. In its opera- 
 tion, the gas is cooled and then passed through a tar ex- 
 tractor in the usual manner, after which the tar-free gas is 
 conducted into an acid washer (see Fig. 26), this washer 
 being built somewhat like a Glover tower, in that it consists of 
 a lead shell the inside of which is supplied with earthenware 
 pipes or tubes, down which the cool acid is caused to trickle, 
 this acid being of 53° Be\ strength. 
 
AMMONIA 
 
 123 
 
 The tower is built. up of lead rings burned to each other, 
 the bottom resting on a cast-iron base, also lined with lead, 
 encased in the masonry foundation, the entire tower being 
 supported by means of timber framing. The tubes (Fig. 27), 
 are small pipes 4 inches high, 3 inches external and 2 inches 
 internal diameter, forming concentric rings supported upon 
 lead-covered iron bars, thus forming a floor located about 
 20 inches above the bottom of the washer, the washer usually 
 being about 9'9" in diameter and 19'6" high. 
 
 At the top of these stacks of tubes are placed three rows 
 of earthenware cupels (Figs. 28 and 29), which in turn support 
 
 •>»>>>!>> 
 
 Fig. 27. 
 
 {V"'J'/ "'( \ 
 
 Fig. 28. 
 
 ! 
 
 >, 
 
 ! 
 
 Fig. 29. 
 
 pots for the purpose of splitting up the gas; these pots are 
 also of earthenware and are supplied with serrated edges, 
 receiving the washing acid from the lead tubes (E) (Fig. 
 26), these latter tubes having an internal diameter of about 
 three-quarters of an inch, being burned to the lead cover of 
 the tower. 
 
 The flow of the acid is regulated in such manner that 
 each tube (E) receives the same amount of acid, thus caus- 
 ing the entire washing surface to be sprayed uniformly. 
 The acid is distributed through the medium of the channels 
 marked (D), from whence it is sprayed by means of plun- 
 gers in the (E) tubes over the whole internal surface of the 
 washers. This distribution of the acid causes it to flow over 
 the pipes, thus coming into intimate contact with the gas, 
 the latter entering the washer at (F) and leaving at (G). 
 
 The acid mother liquor leaves the bottom of the washers 
 and is collected in the lower tanks (A) and (A 2 ), from whence 
 it is pumped by pumps (1) and (2) into the elevated tanks 
 (B) and (B 2 ). When this acid mother liquor has received 
 a sufficient degree of ammonia concentration, a certain 
 portion is withdrawn and conducted to the intermediate 
 
124 
 
 COAL GAS RESIDUALS 
 
 bath (C 2 ), from which it drains to the saturator, where it 
 combines with the ammonia remaining in the gas and with 
 the acid. 
 
 The usual troubles with impure sulphuric acid in the 
 saturator are said to be avoided in this system, because 
 the acid is subjected to a complete purification by means 
 of the hydrogen sulphide in the gas, during its passage through 
 the tower washer, precipitating arsenic sulphide, which 
 floats on top of the acid in the bath (A) and(A 2 ) and from 
 whence it may be removed by means of a ladle. 
 
 The absorption of ammonia is said to be practically com- 
 plete up to 0.25 grain per 100 cubic feet of gas, and it is 
 claimed that the back pressure thrown by the tower washer 
 is only 1 to 1| inches of water. 
 
 The Mont-Cenis Process: 
 
 In the Mont-Cenis system, shown in Fig. 30, the gas is 
 cooled in the usual manner in a cooler (A), the temperature 
 
 e«> ftr ioo"e 
 
 Ammonia \fjppR. ^ 
 
 A^*o™aT6 Cistern 
 
 A-Coolejs, D-5«nj(e«ioR. 
 
 5-EXHBUSTEJl OCOUNTOS CUSKEm COOLER. 
 
 C- Tar Extractor.. F - DCTLUNq Column. 
 
 Fig. 30. — Mont-Cenis Sulphate Plant. 
 
 of the gas entering the exhauster (B) being reduced to about 
 95° F. The condensate from (A) is run off to a separating 
 well, where the tar and ammonia are separated by specific 
 gravity. From the exhauster the gas enters the Pelouze 
 (C), where the remainder of the tar is removed, and then 
 enters the saturator with a temperature of 95° F. 
 
 The ammonia from the well is pumped into a still, where 
 it is treated with steam and lime, the vapors passing through 
 a counter-current cooler (E) where the water is thrown down, 
 after which these vapors enter the saturator in conjunction 
 
AMMONIA 125 
 
 with the gas, the chemical reaction in the saturator increas- 
 ing the temperature of the off-going gas to about 169° F. 
 
 A peculiarity of this system is that a layer of well-distilled 
 tar, about one to two inches thick, is poured on top of the 
 acid bath in the saturator, this tar being for the purpose of 
 preventing a too violent agitation of the acid bath during 
 the formation of salt; it is claimed that this addition of tar 
 does not affect the quality or cleanliness of the salt, providing 
 the gas itself is free of tar. 
 
 The Otto Process: 
 
 In the Otto process all ammonia washing and ammonia 
 liquor distillation is obviated, and the advantages of this 
 process lie in the saving of steam and water, and in the 
 simplicity of supervision. 
 
 The greatest difficulty which had to be overcome in the 
 development of this process lay in separating the tar fog from 
 the hot gases, and many experiments in direct sulphatation 
 failed on this account. The adoption by Otto of the tar 
 washing spray, operating at temperatures which closely 
 approach the dew point of the gas, finally brought about 
 the desired results. 
 
 All substances which can be separated from the gas are 
 removed after the gas has passed the tar sprays, and only such 
 hydro-carbons as are due to the prevaling temperatures are 
 carried forward wrth the gas to the saturators and coolers. 
 Due to intensive cooling and consequent decrease in volume, 
 water vapors, light fatty oils, and naphthalene are condensed 
 and, due to their difference in specific gravity, these sub- 
 stances can be readily separated from each other. 
 
 The tar oils separated at this point are of great value as 
 wash oils for benzol recovery, and therefore have a value of 
 three to four times that of the remaining tar. The acid in 
 the saturator has no effect upon these condensates, and also 
 does not affect the benzol-carbon formation in the gas. 
 With a normal coal and proper cooling the tar oils dissolve 
 the naphthalene thrown down, and it is claimed that stop- 
 pages are thus avoided. 
 
 This system is shown in Fig. 31, and the method of opera- 
 tion is as follows: — The gases coming from the hydraulic 
 
126 
 
 COAL GAS RESIDUALS 
 
 main, and which have been subjected to atmospheric cooling 
 during their passage, have a temperature of from 212° to 
 248° F., and they thus pass through an envelope surround- 
 ing the saturator (Fig. 32), after 
 which they pass into the tar- 
 spraying device, the latter being 
 constantly supplied with tar by 
 means of a rotary pump. Due to 
 the intimate contact between tar, 
 or tar liquor, and the gas, a prac- 
 tically tar-free gas is produced. 
 
 From the tar-receiving tank a 
 portion of the separated tar flows 
 through an overflow to the tar- 
 storage pit, and another portion 
 passes on to the rotary pump, 
 which latter again pumps the hot 
 tar into the spraying device. The 
 pipe connections on the discharge 
 side of this pump are so arranged 
 as to permit the temperature of 
 the washing tar to be kept at any 
 desired degree, thus maintaining 
 the temperature at which the tar 
 is separated at the dew point. 
 Under normal conditions this dew 
 point is at about 169° F., while 
 the temperature of the tar, due to 
 atmospheric changes, varies be- 
 tween 122° and 176° F. 
 
 The tar-spraying apparatus is 
 provided with a number of sprays 
 so arranged that they can be cut 
 out at will, thus making the tar 
 extractor readily adaptable to 
 fluctuations in manufacture. 
 
 After the tar has been removed the hot gas, with its entire 
 content of water vapor and ammonia, is led into the closed 
 saturator through a submerged bell, being sprayed with 
 sulphuric acid in the downpipe leading to the saturator 
 
AMMONIA 
 
 127 
 
 (see Fig. 32). The lower edge of this bell is serrated, and 
 the saturator is only partially filled with sulphuric acid. In 
 the saturator, due to the heat of reaction as well as to the 
 heat transmitted by the raw gas in passing through the 
 envelope, the temperature of the incoming gas is increased, 
 this increase being sufficient to prevent dilution of the bath 
 by condensation. 
 
 Fig. 32. — Otto Sulphate Plant. 
 
 After proper saturation the salts fall to the bottom of the 
 saturator, being removed from there onto the draining table 
 by means of an ejector, going from thence to the centrifugal 
 and finally to storage. 
 
 The Strommenger Process: 
 
 Strommenger claims that many of the direct acid bath 
 systems have failed to measure up to expectations due to 
 inability to entirely expel the tar from the gas, and that in 
 order to produce a saleable sulphate it is absolutely necessary 
 to remove all traces of tar before the gas is admitted into 
 the acid bath; he also claims that the usual devices are only 
 partially successful in the removal of tar, which fact has led 
 him to devise a new form of extractor, stating that the use 
 of this device entirely eliminates the tar and produces a 
 perfectly clean salt. This extractor also makes use of hot 
 tar as a washing medium, the hot tar entering the lower 
 half of the extractor, where it is broken up into a number of 
 fine tar streams in such manner as to bring the gas into 
 intimate contact with the hot tar; the gas then ascends into 
 
128 
 
 COAL GAS RESIDUALS 
 
 the upper half of the extractor, where this process of contact 
 between gas and hot tar is repeated, the tar circulating in a 
 direction counter to that of the gas, the latter also being 
 broken up into fine streams in order to increase contact. 
 
 The upper portion of the extractor is provided with a 
 filter, the gas being compelled to pass through it, thus 
 preventing the carrying forward of any entrained tar. 
 
 e*=J 
 
 VTFW1 
 
 >liJte>l)&>>i&ll^>»mP> s -J/)tir//,-;//is-W i?//}s;)/=M=//l= J ''s 
 
 ZJ - p 
 
 nh rnrii r^ 
 
 r yy/Bfiic/»&*rW>&»&* 
 
 Fig. 33. — Strommenger Sulphate Plant. 
 
 It is claimed that at the Phoenix Colliery, Germany, the 
 tar was completely removed from the gas by means of this 
 extractor, but that with a temperature a little higher than 
 that for which the extractor was designed, the gas retained 
 about eleven grains of tar per 100 cubic feet, a degree 
 of cleanliness which can be equaled by almost any well- 
 designed machine. 
 
 In the sulphate process proper (Fig. 33), the gas from the 
 hydraulic main enters the temperature regulator (A), where 
 its temperature is brought to the desired degree, after which 
 
AMMONIA 129 
 
 it enters the tar extractor (B). The wash tar required in 
 the extractor is constantly replaced by the tar expelled from 
 the gas, and all surplus tar is conducted to storage. The 
 ammonia liquor used in the temperature regulator is trans- 
 ferred to the cooler (F), where it is cooled and such part as is 
 sent to the saturator is replaced by fresh liquor. The water 
 required in the cooler is later cooled on the open frame (<7), 
 or it may be used for other purposes. 
 
 The gas, freed of tar, now enters the saturator (C), where 
 ammonia is extracted by combination with the acid, the gas 
 being then transferred by the exhauster (E) through the 
 final cooler (D) for further treatment. 
 
 The Feld Process: 
 
 In all of the systems thus described we find the necessity 
 of using sulphuric acid, and as the production of one ton of 
 ammonium sulphate requires approximately one ton of sul- 
 phuric acid, the cost of this acid becomes an important 
 feature of manufacture, and it is therefore of general interest 
 to be able to produce sulphate in a manner which makes 
 the producer independent of (lie sulphuric acid manufacturer, 
 or, instead of using acid, to make use of the hydrogen sul- 
 phide contained in the gas and to combine it with the am- 
 monia in the formation of ammonium sulphate. 
 
 For many years various investigators attempted to com- 
 bine ammonia direct with sulphurous acid, and as far back 
 as 1852 a British patent was issued to Laming on this subject. 
 Claus, in England, about 50 years ago tried to remove the 
 hydrogen sulphide from the gas by washing the latter with 
 ammonia, and then to regenerate the resultant solution of 
 sulphur, ammonia, and ammonium carbonate by heating, 
 using the remaining ammonia solution for wash purposes; 
 the high carbonic acid content of the gas required, however, 
 that a volume of ammonia equivalent to about six times 
 that originally contained in the gas should be in constant 
 circulation, thus producing an ammonia loss which prevented 
 the successful introduction of this system. 
 
 Laming, Claus, and the others who were working along 
 similar lines overlooked the fact that the reaction by which 
 ammonia and sulphur dioxide combine will not proceed until 
 
130 COAL GAS RESIDUALS 
 
 the reaction is complete. Neutral ammonium sulphite, 
 (NH 4 ) 2 S0 3 .H 2 0, gives off ammonia, and the quantity of 
 ammonia given off increases with the temperature, the 
 residue being a wet acid salt, or (NH^zHSOs, which, in con- 
 tact with air, will give off sulphur dioxide, and only a small 
 portion will be changed into sulphate. 
 
 This behavior of the sulphites is due to the high vapor 
 tension of both ammonia and sulphur dioxide, and for this 
 reason it becomes impossible to completely absorb the 
 ammonia contained in the gases by means of an aqueous 
 sulphurous acid solution, or to completely absorb sulphur 
 dioxide contained in gases by means of an ammonia solu- 
 tion. 
 
 Even if these difficulties could be overcome in practice, 
 the resulting salt would be of very unstable nature. To 
 use ammonium sulphite directly as a fertilizer, as proposed 
 in Lachomette's British patent of 1887, is impracticable for 
 more than one reason, the principal one being its tendency 
 to give off sulphur dioxide and ammonia to the atmosphere, 
 and attempts to change ammonium sulphite into ammo- 
 nium sulphate by contact with air have proven to be 
 abortive. 
 
 For many years Feld experimented with the reactions 
 occurring between sulphur dioxide and hydrogen sulphide; 
 he found these reactions rather complicated, although the 
 equation usually and erroneously assumed for it, 
 
 2H 2 S + S0 2 = 2H 2 + 3S, 
 
 is simple enough. 
 
 In the course of these researches Feld established the fact 
 that certain tar oils are excellent solvents for sulphur dioxide 
 and sulphur; pure sulphur dioxide introduced into heavy 
 tar oils is eagerly absorbed, and causes a strong evolution 
 of heat. If this absorption is carried out in a closed, agi- 
 tated bottle, the reaction will be so energetic that the atmos- 
 pheric pressure above the liquid is considerably reduced, 
 while the temperature rises. 
 
 Now, if a tar oil has been saturated with sulphur dioxide 
 and pure hydrogen sulphide is then passed into it, the latter 
 will also be completely absorbed, while the temperature 
 
AMMONIA 131 
 
 increases and the pressure decreases. In this case the 
 sulphur dioxide and the hydrogen sulphide react, forming 
 sulphur and water, and the sulphur dissolves in the hot tar 
 oil. 
 
 If the tar oil is now treated alternately with sulphur 
 dioxide and hydrogen sulphide, a portion of the dissolved 
 sulphur will in time crystallize out of the saturated oil as a 
 crystalline grain. 
 
 While this process appears to be very simple, it involves 
 great difficulties when applied to gases containing sulphur 
 dioxide and hydrogen sulphide in very dilute form. 
 
 Further experiments showed that hydrogen sulphide de- 
 composes zinc thiosulphate in such manner that zinc sulphide 
 and elementary sulphur are formed, or 
 
 ZnS 2 8 + 3H 2 S = ZnS + 3H 2 + 4S. 
 
 If concentrated or dilute hydrogen sulphide is introduced 
 into a solution of ZnS 2 3 , the hydrogen sulphide will be 
 completely absorbed, and in experiments carried out on a 
 working scale with illuminating gas, which had been prac- 
 tically freed of ammonia, one washer was sufficient to absorb 
 from 80 to 90 per cent of the hydrogen sulphide in the gas 
 by means of a zinc thiosulphate solution. 
 
 Feld's idea was to regenerate the zinc thiosulphate from 
 the zinc sulphide by means of sulphur dioxide according to 
 the equation, 
 
 2ZnS + 3S0 2 = 2ZnS 2 3 + S, 
 
 but after repeated attempts this had to be abandoned be- 
 cause the process of regeneration could not be made complete. 
 The zinc sulphide dissolves very slowly in the presence of 
 sulphur dioxide, and the regenerated zinc solution loses most 
 of its ability to absorb hydrogen sulphide; the regeneration 
 process does not take place in the manner expected because 
 when zinc sulphide and sulphur dioxide react, polythionate, 
 or ZnS 4 6 is mainly formed instead of the expected thio- 
 sulphate ZnS 2 3 , and hydrogen sulphide acts strongly on 
 zinc thiosulphate, but only slightly on zinc polythionate. 
 
 In spite of these shortcomings this process indicated a 
 direction which ultimately led to success, and it also indi- 
 
132 COAL GAS RESIDUALS 
 
 cated the possibility of simultaneously absorbing hydrogen 
 sulphide and ammonia from the gas in accordance with the 
 equation: 
 
 ZnS 2 3 + 2NH 3 + H 2 S = ZnS + (NH 4 ) 2 S 2 3 , 
 
 and further experiments confirmed this surmise. 
 
 The difficulties met with in attempting to regenerate the 
 ZnS to ZnS 2 3 were, however, too severe, and recourse was 
 then had to an iron salt, FeS 2 3 , which was used in place of 
 the zinc salt for absorption purposes. 
 
 In this latter case the gases containing ammonia and 
 hydrogen sulphide were washed with a solution of FeS 2 3 , 
 or iron thiosulphate, whereby iron sulphides were precipi- 
 tated according to 
 
 FeS 2 3 + 2NHs + H 2 S = FeS + (NH 4 ) 2 S 2 3 . 
 
 This iron sulphide, FeS, was then dissolved in sulphurous 
 acid, thus again forming FeS 2 3 which was then returned to 
 circulation for a further treatment of the gas, or 
 
 2FeS + 3S0 2 = 2FeS 2 3 + S. 
 
 When this alternate treatment of the solution with gas 
 and then with sulphur dioxide had been repeated a number 
 of times, the content of ammonium salt increased to such 
 an extent that the recovery of ammonium sulphate from the 
 solution became profitable. 
 
 The treatment of thiosulphate with sulphurous acid pro- 
 duces polythionate, or 
 
 FeS 2 3 + (NH4) 2 S 2 3 + 3S0 2 = FeS 3 6 + (NHOAO,, 
 
 and this solution being heated gave ammonium sulphate, or 
 
 FeS 3 6 + (NH 4 ) 2 S 4 6 + A = FeS0 4 + (NH 4 ) 2 S0 4 + 2S0 2 + 3S. 
 
 The heating of the solution may be done simultaneously 
 with the treatment with sulphurous acid, so that the last 
 reaction takes place simultaneously with the preceding one. 
 
 The formation of ferrous sulphate in this process is accom- 
 panied by the formation of sulphur dioxide and free sulphur, 
 the latter being burned in a sulphur burner to again form 
 
(Facing Page 132) 
 
Missing Page 
 
AMMONIA 133 
 
 sulphur dioxide for regenerative purposes while the sulphate 
 solution is again treated with crude gas, or 
 
 FeS0 4 + 2(NH 4 ) 2 S0 4 + 2NH 3 + H 2 S - FeS + 3(NH 4 ) 2 S0 4 , 
 
 the solution thus increasing in sulphate content with each 
 treatment, and when the solution had reached a certain 
 degree of ammonia concentration the liquor was evaporated 
 with the consequent formation of salts. 
 
 The successful operation of this process both in Germany 
 and in the United States led to further investigations, par- 
 ticular attention being paid to sodium and ammonium salts. 
 Feld thus found that ammonium thiosuphlate has the char- 
 acteristic property of being very easily changed into poly- 
 thionate by treating it with sulphur dioxide, or 
 
 2(NH 4 ) 2 S 2 3 + 3S0 2 = (NHOAO, + (XH 4 ) 2 S 3 6 , 
 
 and this led to the adoption of the present process as shown 
 diagrammatically in Fig. 34. 
 
 This latter reaction is rapid and complete, so that sulphur 
 dioxide may be completely washed out of dilute gases by 
 treating these gases with ammonium thiosulphate. 
 
 Ammonium polythionate has proven a very effective 
 means for extracting hydrogen sulphide and ammonia, cither 
 in combination or separately, from coal gas. The inter- 
 mediate reactions are many, and therefore the final ones 
 only are given here. 
 
 The polythionate secured with the reaction given above 
 is now treated with gas containing hydrogen sulphide and 
 ammonia, 
 
 (NH 4 ) 2 S 4 6 + (NH 4 ) 2 S = 2(NH 4 ) 2 S 2 3 + S, 
 
 and treating with hydrogen sulphide, we have 
 
 (NH 4 ) 2 S 4 6 + 3H 2 S = (NH 4 ) 2 S 2 3 + 5S + 3H 2 0, 
 
 and then treating with ammonia, we have 
 
 (NH 4 ) 2 S 4 6 + 2NH 3 + H 2 = (NH 4 ) 2 S0 4 + (NH 4 ) 2 S 2 3 + S. 
 
 In each of these three reactions thiosulphate is formed, 
 and this compound is again changed into polythionate by 
 
134 COAL GAS RESIDUALS 
 
 treating the liquor with sulphur dioxide according to the 
 first reaction, or 
 
 2(NH 4 )2S 2 03 + 3S0 2 = (NH 4 ) 2 S 4 6 + (NH 4 ) 2 S 3 0„. 
 
 By an alternate treatment of the solution with coal gas 
 and then with sulphurous acid, the content of ammonium 
 salts gradually increases; finally by treatment with sulphur- 
 ous acid polythionates are formed which are in turn broken 
 down by the application of heat with the resultant formation 
 of sulphate, or 
 
 2(NH 4 ) 2 S 2 3 + 3S0 2 + S = 2(NH 4 ) 2 S 4 6 , 
 and 
 
 3(NH 4 ) 2 S 4 6 + A - 3(NH 4 ) 2 S0 4 + 3S0 2 + S, 
 
 and by adding these two equations we have 
 
 (NH 4 ) 2 S 4 6 + 2(NH 4 ) 2 S 2 6 = 3(NH 4 ) 2 S0 4 + 5S. 
 
 That is, when the content of polythionate has reached a 
 certain maximum, the sulphurous acid which has been set 
 free during sulphate formation is sufficient to change the 
 thiosulphate present into polythionate and finally into am- 
 monium sulphate. The precipitated granular sulphur can 
 now be separated from the sulphate solution and the latter 
 can be evaporated for the production of the ammonium 
 sulphate salts. 
 
 This polythionate process presents a simple and reliable 
 solution of the old problem of completely combining am- 
 monia and sulphur in illuminating or coke-oven gas for the 
 production of ammonium sulphate. As is seen from the 
 reactions given above, no other oxydizing agent except 
 atmospheric oxygen is necessary, and the latter is only used 
 in the sulphur stove where the sulphur is burned to form 
 sulphur dioxide, this sulphur stove thus displacing the sul- 
 phuric acid factory which has hitherto been a necessity. 
 
 The method of operation can best be understood by 
 following the diagram (Fig. 34). The plant is started by 
 filling the two regenerator tanks (B) and C) as well as the 
 wash liquor tank (D) with weak ammonia liquor from the 
 works, or with fresh water, this liquor or water being then 
 
AMMONIA 135 
 
 pumped into the supply tank ( U) from whence it flows into 
 the top of the washer (A). Ammonia extraction immedi- 
 ately follows, and the liquor flows from the bottom of the 
 washer as thiosulphate, entering regenerator (B), where it 
 is treated with sulphur dioxide coming from the sulphur 
 stove (F), the sulphur dioxide being forced through the 
 liquor due to pressure produced by the compressor ((?). 
 The overflow from regenerator (B) enters regenerator (C), 
 where a further treatment with sulphur dioxide is effected, 
 and then enters the wash liquor tank (D), from whence it is 
 again pumped to the washer. 
 
 Alternately treating the liquor with sulphur dioxide and 
 with crude gas forms polythionate in the first instance and 
 thiosulphate in the second, the polythionate being the active 
 washing medium; an equivalent amount of hydrogen sul- 
 phide is removed from the gas .in conjunction with the 
 ammonia in the washer. 
 
 When the liquor has obtained a strength equivalent to 
 35 or 40 per cent of ammonium sulphate, a portion of this 
 concentrated solution is pumped into the oxidation or finish- 
 ing tank (H), it being replaced in the circulating system 
 with weak liquor from prior condensation. In the finishing 
 tank (H) the polythionate liquor is treated with heat through 
 the medium of a steam coil, and some sulphur dioxide if this 
 should be necessary; by this treatment the thiosulphates 
 are transformed into polythionates, which are in turn decom- 
 posed into sulphate and free sulphur. 
 
 The liquor, when decomposed, is drawn off from the 
 finishing tank onto the sulphur drain board (J), and the 
 free sulphur, which is in the form of hard grains, is placed 
 in the centrifugal (K) where its moisture is expelled; this 
 sulphur is returned to the sulphur stove (F) for the further 
 production of sulphur dioxide. 
 
 The liquor from the drain board (J), as well as that from 
 the centrifugal (K), flows to the precipitating tank (Q), 
 where it is treated with some concentrated ammonia, or 
 (NH 4 ) 2 S from the condenser (T); this liquor is then trans- 
 ferred to the storage tank (L) and is pumped from there to 
 the vacuum boiler (M). The concentrated ammonia, or 
 (NH^S, from the still and condenser is sent into the precipi- 
 
136 COAL GAS RESIDUALS 
 
 tating tank in order that any iron held in suspension may be 
 thrown down, thus ensuring a clean, white salt; the amount of 
 ammonia used for this purpose is very small, as will be shown 
 later, and it is all recovered in the sulphate. The liquor is 
 evaporated under a vacuum in the boiler (an open evaporator 
 may be used if desired) , and the salts are drawn out onto the 
 drain board (N), and from thence to the centrifugal (0) and 
 rotary dryer (P), the salts thus produced being ready for 
 bagging and shipping. All mother liquor from the drain- 
 board and centrifugal returns to tank (L) and is from thence 
 pumped back into the boiler. 
 
 The still (R) is used for the purpose of evaporating any 
 ammonia liquor thrown down by previous condensation, these 
 vapors being returned to the gas at the inlet to the washer, 
 a small portion of the vapor being sent into the condenser 
 (T) to produce the (NH 4 ) 2 S used in the precipitating tank. 
 The use of a still is dependent upon the method of operation; 
 if Feld's system of hot-tar extraction is installed, a still will 
 not be necessary, as there will be no precipitation of ammonia 
 under these conditions; a still is also not necessary if the 
 production of ammonia liquor does not exceed the amount of 
 liquor required in circulation. 
 
 This process requires but very little expert control, the 
 usual mechanics attending the machinery being capable of 
 making the necessary liquor tests after some little instruction. 
 
 COMBINED BUEB CYANOGEN AND FELD SULPHATE PLANT 
 
 As stated before, ammonium sulphate is formed in the Bueb 
 process, and therefore these two systems in combination form 
 a very remunerative proposition. The combined plant is 
 shown in diagram in Fig. 35; the operation of both systems 
 is the same as described for the separate processes, but the 
 final liquor from both systems is run into a common sulphate 
 liquor storage tank from whence it is pumped to the vacuum 
 boiler. 
 
 Example. Three Million Cubic Foot Coal-Gas Plant: — 
 Ammonia assumed at five pounds per ton of coal carbonized, 
 or five pounds per 10,000 cubic feet of gas, equivalent to 0.05 
 pound per 100 cubic feet of gas. 
 
Missing Page 
 
AMMONIA 137 
 
 Assume 25 per cent of the total ammonia removed in con- 
 densation (this amount is probably large), and as 0.05 pound 
 is equivalent to 350 grains of ammonia per 100 cubic feet, we 
 will have: 
 In condensation 350 X 0.25 = say 88 grains NH 3 per 100 
 
 cu. ft. 
 In Plant 350 x 0.75 = say 262 grains NH 3 per 100 cu. ft. 
 
 The total ammonia sent to the Cyanogen and Ammonia 
 plant will then be: 
 
 262 x 30,000 = 7,860,000 grains, or 
 
 7,860,000 , , XTXT 
 
 _ nnn — =1123 pounds of NH 3 per day. 
 
 The total ammonia in the condensate will be 
 
 88 X 30,000 = 2,640,000 grains per day, or 
 
 2,640,000 „„ 
 70Q0 377 pounds per day. 
 
 If the coal produces 0.07 pound of water per pound of coal, 
 or 2000 X 0.07 = 140 pounds of water per ton of coal, equiva- 
 lent to 140 X 300 = 42,000 pounds of water per day, equal to 
 5042 gallons, and all of this water is precipitated with a con- 
 tent of 377 pounds of ammonia (an extreme case) this water 
 would contain 377 X 16 = 6032 ounces of ammonia, and the 
 
 ammonia per gallon of water would be r^r = 1.19 ounces," 
 
 and the ounce strength of this liquor, with 
 
 2 (NH,) = 34, andH 2 S0 4 = 98 
 
 would be 
 
 qs 
 1.19 X £ = 3.42, 
 
 or there will be 5042 gallons of 3.42 ounce liquor in condensate 
 to be treated each day. All of this water does not pass off 
 with the ammonia, some of it remaining with the tar, but for 
 the purposes of this example, and to explain an extreme case, 
 this fact has not been considered. 
 
 We will also assume that the gas carries 600 grains of hy- 
 drogen sulphide and 120 grains of cyanogen per 100 cubic feet. 
 
138 COAL GAS RESIDUALS 
 
 The amount of cyanogen present in the gas will then be 
 
 120X30,000 _,, , , 
 ^^ = 514 pounds per day. 
 
 The amount of iron sulphate (FeS0 4 ) required per day for 
 the absorption of this amount of cyanogen will be given by 
 the reaction 
 
 2FeS0 4 .7H 2 + 2H 2 S + 4NH 3 = 2FeS + 2(NH 4 ) 2 S0 4 , or 
 
 556: 156:: x: 514, and 
 
 x = 1832 pounds of FeS0 4 , 
 
 or 76.33 pounds per hour, and with a 30 per cent solution, 
 
 having a specific gravity of 1.174 and weighing 9.78 pounds 
 
 per gallon, 3.565 pounds being required per pound of cyano- 
 
 3 565 
 gen, ' = 1.22 gallons of solution will be required for each 
 
 A.UO 
 
 pound of cyanogen in the gas. 
 
 The amount of ammonia removed from the gas to precipi- 
 tate the FeS0 4 as FeS will also be given by the above reac- 
 tion, or 
 
 556:68:: 1832: x, and 
 x = 224 pounds per day. 
 
 The (NH 4 ) 2 Fe 2 (CN) 6 contains 11.84 per cent of ammonia, 
 and as 514 pounds of cyanogen per day is equivalent to 
 
 514 x 1.95 = 1002 pounds of (NH 4 ) 2 Fe 2 (CN) 6 , 
 
 we will find 1002 x 0.1184 = 119 pounds of ammonia present 
 in the "blue," therefore a total of 224+ 119 = 343 pounds of 
 ammonia are removed from the gas each day in the cyanogen 
 plant, of which 224 pounds are available for sulphate manu- 
 facture, being equivalent to 224 X 3.88 = 869 pounds of am- 
 monium sulphate per day. 
 
 In addition to this amount of ammonium sulphate, some 
 sulphate is also produced in the neutralizing tank by the 
 union of (NH 4 ) 2 S with the FeS0 4 and the acid of this liquor, 
 but the amount is indeterminate. 
 
 The amount of H 2 S removed from the gas is equivalent to 
 about 80 per cent of the NH 3 in practice, or 224 x 0.80 = 179.2 
 pounds, plus that due to the increased alkalinity of the sludge, 
 which latter is variable and due to operating conditions. 
 
AMMONIA 139 
 
 The total H 2 S in the gas is 30,000 x 600 = 18,000,000 grains, 
 and 179.2 X 7000 = 1,254,400 grains are removed as shown 
 above, leaving 16,745,600 grains still to be expelled. 
 
 The amount of sulphuric acid used per day in the cyanogen 
 neutralizing tank is a variable quantity, it depending entirely 
 upon the alkalinity of the liquor, but on an average it 
 amounts to about 0.9 pound per ton of coal carbonized, or in 
 this case to 300 X 0.9 = 270 pounds per day. 
 
 The amount of ammonia remaining in the gas going to the 
 Feld plant is 1123 - 343 = 790 pounds per day, equivalent to 
 790 X 3.88 = 3065 pounds of sulphate, and the amount of 
 hydrogen sulphide removed here will be 790 X 0.80 = 632 
 pounds per day, equivalent to or containing 632 x 0.94 = 594 
 pounds of sulphur. 
 
 The amount of condensate ahead of the plant contains 377 
 pounds of ammonia per day, equivalent to 377 X 3.88 = 1463 
 pounds of sulphate. 
 
 The additional hydrogen sulphide removed by this quantity 
 of ammonia is 377 X 0.80 = 302 pounds, containing 302 x 
 0.94 = 284 pounds of sulphur, making the total sulphur re- 
 covered from the gas 594 + 284 = 878 pounds per day. 
 
 The amount of sulphur required per day to enter into 
 combination is based upon the sulphur in the sulphate, or 
 0.2424 pound per pound of sulphate, and as the total sulphate 
 produced in the Feld plant will be 3065 + 1463 = 4528 pounds 
 per day, 4528 X 0.2424 = 1098 pounds of sulphur will be re- 
 quired, necessitating the purchase of 1098 - 878 = 220 pounds 
 of sulphur per day. 
 
 This combined plant saves quite an amount of iron in the 
 oxide contained in the purifiers, as seen below: 
 
 18 000 000 
 Total H 2 S in gas = ' ' — = 2572 pounds per day, while 
 7UUU 
 
 the total H 2 S removed is 179.2 + 632 + 302= 1113.2 pounds 
 per day, and the cyanogen removed amounts to 514 pounds. 
 
 Fe required in the purifiers for the total elimination of (CN) 2 
 and H 2 S contained in the gas is 
 
 For H 2 S = 2572 x 1.56 = 4012 pounds 
 For (CN) 2 = 514 x 1.19 = _612 " 
 
 Total Fe = 4624 pounds per day. 
 
140 COAL GAS RESIDUALS 
 
 Fe saved by the removal of H 2 S and (CN) 2 in the combined 
 plant is 
 
 For H 2 S = 1113.2 x 1.56 = 1737 pounds 
 For (CN)» = 514 x 1.19 = _612 " 
 
 Total Fe Saved 2349 pounds per day and 
 
 2349 
 the purifier saving will amount to = ttxt = 50.8 per cent. 
 
 4624 
 
 It is safe to predict that the future development of the Feld 
 process will soon present the means of entirely removing the 
 hydrogen sulphide from the gas, thus obviating the use of 
 purifiers with a consequent saving in labor as well as in 
 ground space occupied. 
 
 Summarizing the above, the net income to be expected from 
 this combined plant will be given by: 
 
 Gross Income 
 
 Cyanogen per day 514 lbs. at 13.25^ = $ 68.10 
 
 Sulphate " " 5397 " " 3.00j* = 161.91 
 
 NH, in (CN), Cake per day 119 " " 7.00^ 8.33 
 
 Total Gross Earnings $238.34 
 
 Operating Expenses 
 
 1832 lbs. at $ 9.50 per ton = $ 8.70 
 
 270 " " 11.50 " " = 1.55 
 
 220 " " 22.00 " " = 2.53 
 
 33.2 M. lbs. at 20f£ " M. = 6.64 
 
 " 1075 K.W.H. at ljf = 10.75 
 
 470 lbs. at $8.50 " ton = 2.00 
 
 Oil, Waste, Miscellaneous, Repairs 4.00 
 
 Labor 10.00 
 
 Total Operating Expense = $46.27 
 
 Gross Income per day $238.34 
 
 Operating Expense per day 46.17 
 
 Net Income per day $192.17 
 
 or with 250 maximum working days, $192.17 x 250 = 
 $48,042.50 per year. 
 
 While these figures can be used as a guide, the units should 
 be corrected to suit locations and cost of supplies. 
 
 FeS0 4 
 
 per da 
 
 H 2 S0 4 
 
 a a 
 
 Sulphur 
 
 tc it 
 
 Steam 
 
 a ti 
 
 Power 
 
 it u 
 
 Lime 
 
 it it 
 
AMMONIA 141 
 
 As stated before, the content of cyanogen and ammonia in 
 coal gas is dependent upon the kind of coal used, the type of 
 carbonizing chamber, and the conditions under which car- 
 bonization is effected, especially with reference to heats, 
 length of carbonization period, size of charge, and the amount 
 of seal carried in the hydraulic main, and it should be the 
 object of every operator to remove the gas from the retort 
 as quickly as possible in order to produce a large yield of 
 ammonia; the usual temperatures of carbonization so closely 
 approach the dissociation temperature of ammonia, that pro- 
 tracted contact with the incandescent coke in the retort is 
 liable to break down some of the ammonia already formed 
 and thus reduce the maximum yield. 
 
 The following tests were made to determine the cyanogen 
 and ammonia content at various temperatures in a plant 
 where cyanogen is extracted by the Bueb process and where 
 the Feld sulphate recovery system is used, this plant being 
 equipped with both horizontal and inclined retorts, operating 
 with a little more than one inch seal on the hydraulic main. 
 
 An average of ten tests on the gas leaving the hydraulic 
 main showed that at 143° F. there were 352.4 grains of am- 
 monia per 100 cubic feet of gas from the inclined retorts, and 
 the gas from the horizontal retorts contained 358 grains per 
 100 cubic feet at 135° F. 
 
 Simultaneous tests made at the outlet of the first cooler, 
 for the purpose of determining the amount of ammonia re- 
 moved by cooling for each degree Fahrenheit showed that at 
 104° F. the same gas contained 284 grains per 100 cubic feet, 
 or it had sustained a loss of 1.85 grains per 100 cubic feet for 
 each degree in the reduction of temperature, and the am- 
 monia removed with the tar at these points ranged from 1 to 
 3 per cent of the total content of the gas. 
 
 The factor of 1.85 was used on gas between 100° F. and 
 150° F., and by correction the gas at 150° F. contained 365 
 grains of ammonia per 100 cubic feet from the inclined re- 
 torts, and 385 grains from the horizontal retorts. 
 
 The cyanogen content of the gas at 146° F. was 129 grains 
 per 100 cubic feet, and there was practically no reduction in 
 this content due to a reduction in temperature. . 
 
 In examining into the operating conditions of the combined 
 
142 COAL GAS RESIDUALS 
 
 Bueb and Feld plant just described, 365 grains of ammonia 
 and 130 grains of cyanogen will therefore be assumed in place 
 of the amounts previously mentioned. There is absolutely no 
 difficulty in removing 100 per cent of ammonia, but the hydro- 
 gen sulphide extraction will depend upon the excess of poly- 
 thionates over thiosulphates in the wash liquor, or on the 
 excess of polythionates over what is required to entirely re- 
 move the ammonia. 
 
 Under these conditions we will have 
 
 I. NH 3 in gas at inlet to tar washer: 
 
 30,000 X 365 
 
 7000 
 
 = 1564.3 pounds at 150° F. 
 
 II. NH 3 in gas at inlet to naphthalene washer, assuming 
 that a naphthalene washer is interposed between the tar 
 and cyanogen washer, and assuming that 5 per cent of am- 
 monia would be removed with the tar, there would be at this 
 point : 
 
 1564.3 - (1564.3 X 0.05) = 1486 pounds at 147° F. 
 
 III. NH 3 in gas at inlet to cyanogen washer: Allowing a 
 drop in temperature of 5 degrees in the pipe connections be- 
 tween No. I and No. Ill, the ammonia in the gas at this 
 point would be 1446 pounds at 145° F. 
 
 At 145° F., and with an average cyanogen content of 130 
 
 grains per 100 cubic feet, the amount of cyanogen in the gas 
 
 would be 
 
 130 X 30,000 ccr7 , , 
 ^^ = 557 pounds per day, 
 
 and as one pound of cyanogen will remove 0.665 pound of 
 
 ammonia, 557 X 0.665 = 370 pounds of ammonia per day, or 
 
 370 
 
 ., ... = 25.5 per cent of the total ammonia will be removed 
 1.446 
 
 in the cyanogen washer. 
 
 About one-third, or 123 pounds of this ammonia exists as 
 combined ammonia in the " blue," the remaining 247 pounds 
 being available as ammonium sulphate, providing the cyano- 
 gen cake is thoroughly washed in the cake wash tanks, thus 
 removing all sulphate liquor from the combination. A sum- 
 mary of the above contents is given in Table XVII. 
 
AMMONIA 
 TABLE XVII.— AMMONIA CONTENT 
 
 143 
 
 
 Temp, of 
 
 gas in 
 
 degrees 
 
 F. 
 
 Ammonia 
 
 Location 
 
 Gr. per 
 100 
 
 Lbs. per 
 day 
 
 I. Inlet to tar washer 
 
 150 
 147 
 145 
 145 
 
 365 
 347 
 337 
 251 
 
 1564 3 
 
 
 1486 
 
 III. Inlet to cyanogen washer 
 
 IV. Inlet to polythionate washer 
 
 1446.0 
 1076.0 
 
 
 
 I As shown by the reactions, one pound of cyanogen will 
 require 3.565 pounds of copperas for its complete removal, or 
 there will be required 
 
 557 x 3.565 = 1986 pounds of FeS0 4 .7H 2 0, 
 
 and in terms of a 30 per cent solution, with a specific gravity 
 of 1.174, there will be required 
 
 1986 
 
 2.93 X 24 
 
 = 28 gallons per hour. 
 
 Calculating to a 25 per cent cyanogen press cake, there will 
 
 2228 
 be 557 X 4 = 2228 pounds of press cake per day, or -—-- =7.43 
 
 o(JU 
 
 pounds of cake per ton of coal. 
 
 Figuring the liquor as 70 per cent of the total used, there 
 
 would be 677 X 0.70 = 474 gallons of press liquor per day, this 
 
 958 
 liquor containing -r=r = 2 pounds of sulphate per gallon, and 
 
 the amount of ammonia at the outlet of the cyanogen washer 
 would be 1446 - 370 = 1076 pounds per day, or 251 grains 
 per 100 cubic feet. 
 
 In actual practice the amount of ammonia removed in the 
 cyanogen washer is higher than the theoretical 25.5 per cent 
 as calculated, for owing to variations in the volume of gas 
 produced at various times, and also the possibility of a change 
 in the cyanogen content of the gas, it is next to impossible to 
 
144 COAL GAS RESIDUALS 
 
 regulate the supply of wash liquor to take care of these 
 changes. 
 
 For instance, if the liquor supply per hour to the cyanogen 
 washer is regulated to treat 125,000 cubic feet of gas with a 
 cyanogen content of 130 grains per 100 cubic feet, and only 
 115,000 cubic feet of gas is actually treated, there would be 
 excess ammonia removed from the gas, as each pound of cop- 
 peras in excess of what is required to combine with the cyano- 
 gen will remove 0.44 pound of ammonia, but this ammonia is 
 available as sulphate in the resultant liquor. 
 
 The amount of sulphate recovered from the cyanogen sludge 
 will depend entirely upon the amount of washing to which 
 the press cake is subjected, and while the ammonia content 
 of the cake is paid for as such, it is worth more as sulphate 
 than as ammonia. 
 
 The amount of sulphuric acid required in the neutralizer 
 will depend upon the free ammonia and the iron sulphide 
 present in the cyanogen sludge in excess of what is required 
 to combine with the cyanogen. 
 
 However, these questions of excess ammonia removal and 
 increased sulphuric acid consumption depend entirely upon 
 liquor regulation, and by working to a fixed percentage of 
 cyanogen removal, say 95 per cent, instead of complete ex- 
 traction, there should be no difficulty in solving these 
 problems. 
 
 In the polythionate process, based on sulphate removal 
 and in terms of a 35 per cent liquor, there will, with 1076 
 pounds of ammonia, or 251 grains per 100 cubic feet, be pro- 
 duced each day 
 
 1076 X 3.88 = 4175 pounds of sulphate, or in terms of a 35 
 per cent liquor with a specific gravity of 1.20, 
 
 4175 
 
 „ _ =1193 gallons per day, 
 o.o 
 
 to which must be added the 474 gallons of press liquor con- 
 taining two pounds of sulphate per gallon. 
 
 The amount of liquor in circulation depends upon the 
 nature of the compounds present in the liquor, and these in 
 turn upon the degree of regeneration of thiosulphate into 
 polythionate, because ammonium thiosulphate, (NH^SaOs, 
 will remove ammonia but no hydrogen sulphide, and the 
 
AMMONIA 145 
 
 purification efficiency of thiosulphate and polythionate can 
 be obtained from the following equations: 
 
 I. (NH 4 ) 2 S 2 3 + 2NH, + 2H 2 = (NH 4 ) 2 S0 4 + (NH 4 ) 2 S + 
 
 H 2 0, or one pound of (NH 4 ) 2 S 2 3 will remove 0.23 pound of 
 ammonia. 
 
 II. (NH 4 ) 2 S 4 6 + 2NH 3 + H 2 S = 2(NH 4 ) 2 S 2 3 + S (A) 
 (NH 4 ) 2 S 2 3 + 2NH 3 + 2H 2 = (NH 4 ) 2 S + (NH 4 ) 2 S0 4 + 
 
 H 2 (B) 
 
 Due to this double decomposition, one pound of (NH 4 ) 2 S 4 0s 
 will remove 0.39 pound of ammonia. 
 
 At a European Plant, one where the Feld system was first 
 installed, the liquor with a specific gravity of 1.25 contained: 
 (NH 4 ) 2 S 2 3 . . .7.5 molecules, or 5.55 per cent, or 0.57 pound 
 of ammonium thiosulphate per gallon, 
 
 (NH 4 ) 2 S 4 6 . . .4.0 molecules, or 2.60 per cent, or 0.27 pound of 
 ammonium tetrathionate per gallon, or the thiosulphate was 
 to the polythionate as 7.5:4, and one gallon of this liquor 
 would remove 0.57 x 0.23 = 0.131 pound of NH 3 , and 
 0.27 X 0.39 = 0.105 " " " or 
 a total of 0.131 + 0.105 = 0.236 pound of ammonia, and one 
 pound of ammonia under these conditions would require 
 
 = 4.24 gallons of liquor; therefore, for the amount of 
 0.236 
 
 ammonia present in the above described plant, 
 1076 X 4.24 = 4562 gallons, or 
 
 190 gallons per hour, 
 3.17 gallons per minute,, 
 equivalent to 1.52 gallons of liquor of this composition in cir- 
 culation per 1000 cubic feet of gas. 
 
 If the above liquor, which contains 31.08 per cent 
 (NH 4 ) 2 S0 4 , or 8.02 per cent NH 3 , and 20 per cent of which is 
 composed of thiosulphate and tetrathionate compounds, had 
 this 20 per cent existing as all tetrathionate, then one gallon- 
 of 1.25 specific gravity liquor would contain 12.3 per cent 
 (NH 4 ) 2 S 4 6 , equivalent to 1.28 pounds of tetrathionate, and it 
 would remove 0.5 pound of ammonia, thus only two gallons 
 of liquor of this composition would be required in circulation 
 per pound of ammonia. 
 
146 COAL GAS RESIDUALS 
 
 The ratio of polythionate to thiosulphate is dependent upon 
 regeneration, or upon an emulsion with sulphur dioxide pro- 
 duced by the burning of sulphur, and a better ratio than 7.5 
 to 4 will ensure better extraction, and if the tankage is 
 arranged for the 7.5 to 4 ratio and the liquor is treated with 
 more S0 2 than this ratio requires, the regenerating reactions 
 will progress much slower, owing to the increased capacity 
 of the vessels, and they will therefore be more complete. 
 
 The amount of sulphur to be burned is obtained from the 
 following regeneration reactions: 
 
 III. 2(NH 4 ) 2 S + 3S0 2 = 2(NH 4 ) 2 S 2 3 + S. 
 
 IV. 2(NH 4 ) 2 S 2 3 + 3S0 2 + S = 2(NH 4 ) 2 S 4 06, 
 
 and calculating from this we find that one pound of thiosul- 
 phate requires 0.32 pound of sulphur; then 4562 gallons of 
 liquor containing 0.57 pound of thiosulphate per gallon, and 
 0.27 pound of polythionate per gallon, will require the com- 
 bustion of 
 
 (4562 x 0.57) 0.32 = 832 pounds of sulphur, 
 
 (4562 x 0.27) 0.73 = _899 " 
 
 or a total of 1731 pounds of sulphur per day 
 
 1731 
 for the regeneration of this amount of liquor, or = 1.61 
 
 pounds of sulphur per pound of ammonia in the gas. 
 
 If the liquor had been so regenerated that the entire 20 
 per cent of thiosulphate and polythionate compounds had been 
 converted into polythionate requiring two gallons of liquor 
 per pound of ammonia, or 1076 X 2 = 2152 gallons per day 
 containing 2152 x 1.28 = 2755 pounds of (NH 4 ) 2 S 4 6 , and 
 when reduced as per the above reactions, there would be 
 required to regenerate 2755 x 0.73 = 2011 pounds of sulphur; 
 therefore to obtain the same extraction results, as regards 
 ammonia, with this liquor it would be necessary to burn 
 280 pounds, or slightly over' 16 per cent more sulphur than 
 would be required to regenerate the 7.5 to 4 liquor. 
 
 From the efficiency ratio of polythionate to thiosulphate 
 which is 
 
 O4O6 '. S2O3 
 
 For ammonia 1.7 : 1.0 
 
 For hydrogen sulphide 1.0 : 0.0, 
 
AMMONIA 147 
 
 it is readily seen that a higher hydrogen sulphide extraction 
 will be obtained when the liquor contains a higher percentage 
 of (NH 4 ) 2 S40 6 than is given in the 7.5 to 4 ratio, and this can 
 readily be secured by burning the proper quantity of sulphur. 
 
 As stated before, the only oxidizing agent used during the 
 process is air, and this is used to maintain combustion in the 
 sulphur burner. One pound of sulphur requires one pound of 
 oxygen, or 4.315 pounds of air, or 4.315 X 13.06 = 56.3 cubic 
 feet of air per pound of sulphur; this would require in the 
 first case 1731 X 56.3 = 97,455 cubic feet of air per day, or 
 68 cubic feet per minute, and in the second case 2011 X 56.3 = 
 113,219 cubic feet per day, or 80 cubic feet per minute, and 
 for the production of a 10 per cent S0 2 double this quantity, 
 or 136 and 160 cubic feet respectively will be required per 
 minute, and an excess of air is an undoubted advantage and 
 aid to regeneration due to greater liquor agitation. 
 
 The sulphur thus burned is not lost, but it is all, with the 
 exception of the small amount which may be required to 
 complete the sulphate combination, recovered in the finishing 
 tank, from whence it is returned to the sulphur stove. 
 
 The amount of ammonia required to precipitate the iron in 
 the 1667 gallons of liquor is small. If this finished liquor 
 weighs nine pounds per gallon, giving a total of 15,003 pounds 
 of liquor, and assuming the maximum Fe content of this 
 liquor to be one-quarter of 1 per cent, the amount of Fe 
 to be precipitated will be 15,003 X 0.0025 = 37.5 pounds. 
 
 One pound of Fe requires 0.1223 pound of ammonia for its 
 precipitation, or the amount of ammonia required will be 
 37.5 X 0.1225 = 4.59 pounds per day, and if liquor containing 
 1.54 per cent of NH 3 is used for this purpose, or liquor con- 
 taining 8.33 X 0.0154 = 0.13 pound of NH 3 , the amount to be 
 
 4.59 
 evaporated per day will be ^pr- = 35.3 gallons. 
 
 Some of the Feld plants being built at present are designed 
 to operate under a vacuum instead of a pressure, in which 
 case the compressor is replaced by an exhauster and the sul- 
 phur dioxide is drawn through the liquor instead of being 
 pressed through as explained above. 
 
 As the Feld system of purification is but little known in 
 the United States, it may be well to mention that two plants, 
 each of 3,000,000 cubic feet capacity, are being built in the 
 
148 COAL GAS RESIDUALS 
 
 United States at present, one of them being for the produc- 
 tion of ammonium sulphate only, by the Feld process, the 
 other being a combination between the Bueb Cyanogen and 
 the Feld Sulphate systems, as described in the preceding 
 pages. 
 
 Among the larger Feld plants being erected in Europe at the 
 present time may be mentioned: 
 
 One in France for the direct recovery of by-products from 
 1250 tons of coal per day. 
 
 One in Germany for 1600 tons of coal per day. 
 
 One for the production of 5.5 tons of sulphate per day from 
 producer gas. 
 
 One for the purification of 5,500,000 cubic feet of coal gas with 
 sulphate production. 
 
 One for the production of pitch and oils from 800 tons of coal 
 per day. 
 
 The Feld Direct Sulphate Plant as built in the United States 
 differs radically from that previously described, in that the 
 European system is operated under pressure, while that in 
 America operates under a partial vacuum, as shown in diagram 
 on Fig. 35a. 
 
 The gas is passed through two washers in series, the fresh 
 wash liquor, from the regulating tank being admitted to the top 
 of the second washer, overflowing from there to the process 
 tank located on the lower floor; a centrifugal pump takes this 
 liquor from the process tank and pumps it up to the supply 
 tank over the first washer, this washer being fed from this 
 supply tank. 
 
 The liquor from the process tank under the first washer is 
 pumped up into the V-shaped supply tank, running from there 
 through a sight-overflow into the bottom of the mixing tank; 
 the liquor now flows from the top of this mixing tank through 
 the second and first acidifiers in the direction indicated. 
 
 The sulphur is burned to sulphur dioxide in the rotary sulphur 
 burner, this gas passing through a sulphur melter into an all 
 cast-iron condenser and from there up to the first acidifier, 
 that portion of the S0 2 gas not absorbed in the first acidifier 
 is taken up in the second, the two acidifiers and the mix- 
 ing tank being kept under a partial vacuum by means of the 
 exhauster located on the second floor. 
 
Lime Miunq 
 
 :an feld d 
 
 (Facing Page 14S) 
 
Missing Page 
 
AMMOXIA 149 
 
 After the liquor contains its burden of ammonium sulphate, 
 25 per cent is withdrawn from this train and run into the boil- 
 ing tank, where it is heated by indirect steam, and where some 
 SO2 is added. After the polythionates have thus been broken 
 down, the ammonium sulphate liquor, with granular sulphur 
 in suspension, is drawn off and enters the settling tank on the 
 second floor; any sulphur which does not drop out in this tank 
 will be removed in the sulphur centrifugal. This sulphur is 
 dried and then passed to the sulphur melter, where it is melted, 
 cast into sticks, and again charged into the rotary sulphur burner. 
 
 The sulphate liquor from the centrifugal is run into the 
 precipitating tank, where any iron which might discolor the salt 
 is precipitated by strong ammonia liquor produced in the am- 
 monia still. The sulphate liquor is now cleared and decanted 
 in the three clear liquor tanks, and then pumped up to the 
 evaporator where the salt is formed, this finished salt being 
 then dried in the sulphate centrifugal, all mother liquor from the 
 evaporator and centrifugal returning to the precipitating tank. 
 
 The sludge which finally collects in the precipitating tank is 
 filter-pressed, the cake being added to the oxide in the purifiers, 
 the clear liquor returning to the precipitator. 
 
 The plant is provided with an ammonia still to take care of 
 any ammonia which may come down in condensation prior to 
 the process proper, the strong liquor produced in this still being 
 run into the first acidifier, where it joins the polythionate liquor 
 in circulation and makes up for the liquor removed to the boil- 
 ing tank. 
 
 The chemistry of the process is the same as that previously 
 described, but the apparatus differs in many essential particu- 
 lars. The difficulty of perfect operation could heretofore be 
 traced back to the regeneration train, as the regeneration of the 
 liquor with S0 2 gas was difficult to secure with the apparatus 
 previously used, but the new design of vacuum acidifiers and 
 mixers has removed this difficulty, and the contact between 
 liquor and S0 2 gas is now practically perfect. 
 
 At the large scale experimental station, treating 1,500,000 
 cubic feet of gas per day, as high as 132 grains of hydrogen 
 sulphide per 100 grains of ammonia were removed, while only 
 an equivalent of 100 grains of H«S per 100 grains of XH 3 are 
 required for the sulphate combination; the ammonia was en- 
 
150 COAL GAS RESIDUALS 
 
 tirely removed from the gas, this removal averaging 100 per 
 cent during the entire operating period of one year. The 
 sulphate produced was clean and white, and contained 25.7 
 per cent of NH 3 on the dry basis. 
 
 The improved plant recently completed at another point will 
 shortly be placed in operation, and the results expected from 
 there should exceed anything done at the experimental plant. 
 
 CONCLUSIONS 
 
 As seen from the system just described, the old, or indirect 
 process required a priori cooling of the gas with practically a 
 complete absorption of ammonia in water, this latter liquor 
 being treated with steam and lime in a still and the ammonia 
 vapors thus produced sent into a saturator, while the so-called 
 "direct" or "semi-direct" systems, besides calling for a com- 
 plete elimination of the tar, require that the temperature in the 
 saturator be sufficiently high to prevent condensation with a 
 consequent dilution of the acid bath. 
 
 In the Brunck system the tar-free gas was supposed to have 
 a temperature sufficiently high to pass through the saturator 
 without condensation, but this temperature, in spite of the fact 
 that it was above the critical point, and notwithstanding the 
 additional heat produced by the reaction in the acid bath, dropped 
 with consequent condensation. Owing to the fact that an ar- 
 tificial increase in temperature above that of the entering gas 
 had a bad effect upon tar recovery, this method, which at first 
 had seemed so promising, found but little application. 
 
 The Otto process proved to be a successful endeavor to im- 
 prove upon what Brunck had accomplished; Otto replaced 
 Brunck's tar washer with liquid tar injectors of the atomizing 
 type followed by a Pelouze drum, the latter apparatus being 
 placed on the suction side of the exhauster. The Koppers sys- 
 tem starts with radical condensation, thus precluding con- 
 densation in the saturator, the gas being admitted into the acid 
 bath at a lower temperature than in the Otto, this gas being, 
 however, heated to the required temperature by means of hot 
 raw gas coming from the hydraulic main. Koppers distills the 
 condensed ammonia, but the vapors are not permitted to enter 
 the saturator immediately, as in the Mont-Cenis system, they 
 being mixed with the gases prior to entering the saturator. 
 
AMMONIA 151 
 
 All of these processes, except the indirect, require the pass- 
 ing of the gas through the acid bath, producing a back pres- 
 sure which is very high, as it depends upon the depth of the 
 liquid bath, the report made by C. Heck to the German Coke 
 Commission stating that these back pressures are 900 mm. = 
 35-J inches in the Otto system, and 750 mm. = 29| inches in the 
 Koppers system, but greater back pressures than these have 
 been known, and it is a known fact that the power required to 
 force a thousand cubic feet of gas through a given depth of liquid 
 rapidly increases with this depth. 
 
 It should also be remembered that the gas in rising through 
 the liquid of a certain height causes the useful effect of con- 
 tact to decrease quickly, owing to the fact that a number of 
 the smaller gas bubbles rapidly unite to form larger ones, thus 
 reducing surface contact. 
 
 Charles Berthelot, in the "Revue de Metallurgie," states 
 that it is difficult to obtain complete absorption by this acid 
 bath, and that it is necessary to renew the surfaces of contact 
 between the gas and the acid as often as possible; he also states 
 that Otto does this by violent agitation of the liquid by means 
 of a pump, thus making it possible for him to adopt a lesser 
 travel of the gas through the acid; in contradistinction to this, 
 it is seen that in the Mont-Cenis system a layer of tar is used on 
 the surface of the acid in the saturator in order to prevent this 
 very agitation. 
 
 Berthelot also states that "from experiments made it would 
 seem that the ammonia lost due to direct sulphatation in the 
 acid bath would be about 77 grains per ton of coal, but that 
 this is not found to be the case in actual practice." He states 
 that in the semi-direct process 6.8 per cent of the total am- 
 monia is retained in the tar, and that this loss cannot occur in 
 the indirect system because there the water which is condensed 
 with the tar in the mains dissolves out almost the whole of the 
 fixed salts, and that the loss in ammonia in both the direct and 
 semi-direct acid bath systems is practically the same. 
 
 He states that a serious objection to the adoption of these 
 direct systems is due to the presence of ammonium chloride 
 in the gas, the difficulties produced by this constituent being 
 due to its method of travel. A portion of this salt remains in 
 the tar and not only occasions ammonia loss, but also causes 
 
152 COAL GAS RESIDUALS 
 
 serious trouble due to its decomposition when distilled. The 
 only way in which this loss can be avoided is to wash the tar 
 with water and then to distill this water to recover the ammonia. 
 
 Further, he claims that the ammonium chloride also enters 
 the saturator and is there decomposed by the sulphuric acid, 
 causing a loss of heat and producing hydrochloric acid, which 
 seriously attacks all iron work in the apparatus following the 
 saturator unless the gas is kept alkaline, in which latter case 
 more ammonia is lost. A portion of the ammonium chloride 
 is always dissolved without decomposition in the bath, and 
 the salt thus produced loses some of its fertilizer properties. 
 
 Berthelot states that the loss of ammonia per ton of coal 
 amounts to: 
 
 Old, or indirect system 385 grains 
 
 Semi-direct system 231 
 
 Direct system '. 185 
 
 and that if one ton of coal produces 3000 grammes, equivalent 
 to 6 pounds of ammonia, these losses will be 
 
 Old, or indirect system 0.84 per cent 
 
 Semi-direct system 0.50 
 
 Direct system 0.40 
 
 The Feld system was offered with the endeavor to overcome 
 all of the objections given against the acid bath process, and 
 as the former is one of chemical absorption without the use of 
 external acid, it has successfully met a great many of the de- 
 mands which would remove the troubles due to the saturator. 
 If Feld's system of hot-tar extraction is included in the process 
 there will be no condensation, and the gas reaches the poly- 
 thionate washers containing all of its ammonia. Besides this, 
 the cost of production plays an important role in the manufac- 
 ture of any product, and in order to show what may be expected 
 from the various systems reference should be made to Table 
 XVIII, which shows the cost of sulphate manufacture by the In- 
 direct, Semi-direct, and Direct methods as reported to the Ger- 
 man Coke Commission by C. Heck, 1 and to which has been 
 added the corresponding figures for the Feld system. This 
 table shows the differences in the cost of producing one ton of 
 sulphate by the four systems in use, and, as stated before, syn- 
 
 1 See " Stahl und Eisen," Vol. 33, 1913. 
 
TABLE XVIII. — ECONOMIC BALANCE FOR THE PRODUCTION 
 OF (NHOsSO, FROM 100 COKE-OVENS 
 
 Items 
 
 Indirect 
 
 Semi- 
 Direct 
 
 Direct 
 
 Feld 
 
 
 33.3 
 
 33.3 
 
 33.3 
 
 33 3 
 
 
 
 
 12 
 500 
 
 6 
 
 0.3 
 
 0.3 
 
 
 
 C 1 Exhauster power back pressure in mm. of water, 
 
 750 
 
 900 
 
 150 
 
 
 39 
 
 58 
 
 70 
 
 12 
 
 
 
 
 12 
 
 7 
 3 
 
 12 
 
 3 
 3 
 
 12 
 
 5 
 2 
 3 
 
 12 
 
 3 " " spray cleaners, washers, circulating 
 
 40 
 
 
 2 
 
 5 *' " centrifugals, average per hr 
 
 3 
 
 
 61 
 
 76 
 
 92 
 
 69 
 
 
 
 D 1 Steam for NH< Still (250 kg/cbm) tons p. hr. ... 
 
 3 
 0.05 
 
 1.5 
 0.1 
 0.05 
 
 0.075 
 0.025 
 0.050 
 
 0.075 
 
 
 0.050 
 
 
 
 
 3.05 
 
 1.65 
 
 0.150 
 
 0.125 
 
 
 
 Value of power and steam consumed. 
 E Case I — Electric motive power and fresh steam for 
 heating and distilling; for comparison, and in view 
 o,f losses, take 1 H.P. at point of consumption =* 
 1 K.W. at switchboard, or 7 kg. steam at turbo- 
 generator, then 
 
 0.427 
 3.050 
 
 0.532 
 1.650 
 
 0.644 
 0.150 
 
 483 
 
 
 1.500 
 
 
 
 
 3.477 
 6.954 
 
 2.182 
 4.364 
 
 0.794 
 1.588 
 
 1.983 
 
 4 Value of steam at 2 M per ton, M per hour 
 
 3.966 
 
 
 60,900 
 
 38,100 
 
 13,900 
 
 34,800 
 
 
 
 F Caso II — Using exhaust steam for power produc- 
 tion in steam turbines. 
 
 3.05 
 120 
 
 1.65 
 
 66 
 10 
 0.07 
 
 0.150 
 
 6 
 
 86 
 
 0.602 
 
 0.125 
 
 2 With 25 kg. steam p. H. P. the power to be derived 
 
 5 
 
 
 
 
 448 
 
 
 
 
 3.050 
 
 1.720 
 
 0.752 
 
 573 
 
 
 
 Value of ateam at 2 M per ton, M per hour 
 
 6.100 
 
 3.440 
 
 1.501 
 
 1.146 
 
 
 53,400 
 
 31,000 
 
 13,150 
 
 10,050 
 
 
 
 G 1 Cooling water required (1 cbm = 0.08M) 
 
 year 
 
 2 Sulphuric acid required (1 cbm = 30.00M) . . 
 
 8,400 
 
 105,000 
 2,200 
 
 3,500 
 
 4,200 
 
 105,000 
 1,100 
 
 3,500 
 
 105,000 
 500 
 
 3,500 
 
 500 
 
 4 Sulphur required (1 ton = 92.00M) " 
 
 11,600 
 3,500 
 
 
 
 
 119,100 
 
 35,000 
 
 215,000 
 
 207,500 
 
 3,320 
 
 $16.25 
 
 $15.65 
 
 113,800 
 35,000 
 186,900 
 179,800 
 3,320 
 $14.08 
 $13.54 
 
 109,000 
 35,000 
 157,900 
 157,150 
 3,320 
 $11.89 
 $11.84 
 
 15,600 
 35,000 
 
 
 
 85,400 
 60,650 
 
 
 L 1 Sulphate produced, 25 lbs. per ton, tons per year. 
 
 3,320 
 $6.50 
 
 N 1 Cost per ton of sulphate, case II 
 
 $4.57 
 
154 COAL GAS RESIDUALS 
 
 thetic ammonia will affect the market wherever power is cheap 
 enough to permit of its production; therefore the gas producer 
 should closely examine into all existing conditions before de- 
 termining upon which system is best to install. While the cost 
 figures given are all theoretical, the same error will probably 
 apply in practice to all four systems, but the elimination of pur- 
 chased sulphuric acid in the Feld process reduces the ultimate 
 cost at once by the value of this constituent of sulphate. 
 
 Diagram III is a graphic representation of the principal acid 
 bath systems, showing the usual operating temperatures for 
 comparison. 
 
 Before closing the chapter on ammonia it will be well, under 
 present circumstances, to call attention to the manufacture of* 
 nitric acid from ammonia, the conversion process being com- 
 monly known as the "Ostwald," although there are quite a 
 number of such processes which operate on the same principle, 
 but which differ in the design of the apparatus as well as in 
 the character of the catalyzer used. 
 
 The conduct -of the present war will require a vast amount 
 of nitric acid, and although export from the natural saltpeter 
 fields of Chile is still open, political changes may close this 
 field to us at any time, and we would then perforce be obliged 
 to turn to atmospheric fixation of nitrogen, or some process 
 similar to the Ostwald to supply our pressing demands. 
 
 The fixation of atmospheric nitrogen is limited to such lo- 
 calities where cheap electric power is available, and therefore, 
 the Ostwald process lends itself to operations in various fields 
 where sufficient electric power at low cost is not obtainable. 
 Extensive data on the Ostwald process cannot be secured at 
 present, and although it has been reported that this process is 
 in extensive use in Germany, the operators are very careful as 
 to any information which might point to its success. 
 
 In 1900 Ostwald took up the study of the oxidation of am- 
 monia, and he soon discovered the conditions under which the 
 reaction proceeded; in the laboratory he was able to obtain a 
 yield of over 85 per cent of HNO on the ammonia. He es- 
 tablished the fact that platinized platinum, or smooth platinum 
 roughened by depositing a thin layer of platinum sponge on 
 its surface, was the best catalyzer, and that a thin layer of the 
 catalyst could be worked to better advantage than a thick one; 
 
AMMONIA 
 
 155 
 
 -»|003 
 
 cpoi i T j . a 
 
 8?" 
 j ceeuEofl 
 
 § $ 8 § § § 8 8 8 ? o 
 
156 COAL GAS RESIDUALS 
 
 also that the gas mixture must be caused to pass over the 
 catalyst at a high velocity. 
 
 A factory was built at Gerthe, near Bochum in 1909, and it is 
 stated that this plant had an annual production of 2400 tons of 
 53 per cent nitric acid, using ammonia derived from the car- 
 bonization of coal. Another report covering this plant states 
 that it produces 2000 tons of ammonium nitrate per year. 
 
 This process was purchased by the Nitrates Products Com- 
 pany, Ltd., London, in 1910, the new company taking over 
 the Gerthe plant, and erecting a new plant at Vilvorde, Bel- 
 gium, the latter plant producing 40 to 50 per cent nitric acid. 
 
 The original converters 1 were constructed in such manner 
 that an exchange of heat occurred between the incoming and 
 the outgoing gas, thus not only keeping the catalyzer at the 
 proper temperature of 572° F., because of the strongly ex- 
 othermic character of the reaction NH 3 + 20 2 = HN0 3 + H 2 0, 
 in which 385 B. T. U. are evolved, but also serving to regulate 
 this temperature, although considerable variation in the rate 
 of gas flow might exist. The thickness of the catalyzer element, 
 as well as the velocity of the gas, is so chosen that the time 
 of contact between the two does not exceed one-hundredth of 
 a second. 
 
 The proportion between ammonia and air can be varied 
 considerably; an excess of ammonia gives ammonium nitrate 
 as the end product, while if the ammonia and the air are so 
 proportioned that the equation NH 3 + 20 2 = HN0 3 + H 2 is 
 fulfilled, and no water is dripped on the absorption towers, the 
 maximum strength of nitric acid theoretically obtainable will 
 be 77.8 per cent. Ostwald states that no less air should be 
 used than represented by the equation 2NH 3 + 70 = 2N0 2 + 
 3H 2 0, but in this case a very rapid passage of the gas will 
 result in obtaining a mixture of ammonium nitrate and am- 
 monium nitrite in the exit gases. 
 
 Camille Matignon states that the above reactions are not 
 entirely correct, as at the temperature of the reaction neither 
 N0 2 nor HNO3 can exist, and that the equation should be 
 written 2NH 3 + 50 = 2NO + 3H 2 0. It is of course under- 
 stood that if an excess of air is used the NO will quickly oxidize 
 
 1 F. C. Zeisberg. "Met. Chem. Engrg." Vol. XV. Page 300. 
 
AMMONIA 157 
 
 to N0 2 as soon as the temperature is sufficiently reduced. The 
 process of oxidation begins as soon as the temperature drops 
 below 1112° F., but it cannot become complete until the tem- 
 perature drops to 284° F. 
 
 At this lower temperature, moreover, the reaction velocity 
 is very low, so that an appreciable time must elapse before 
 the reaction becomes complete. 
 
 "Metall und Erz," 1916, Vol. XIII, page 21, published a dia- 
 gram of a converter, the catalyzer in this apparatus being 
 platinum gauze electrically heated to 1292° F. The gases enter 
 at the bottom and pass upward through the gauze, no heat 
 exchange being effected. At this temperature the conversion 
 is almost quantitative, steam and nitric oxide (NO) being the 
 end products, which would indicate that the equation of re- 
 action is 4NH 3 + 50 2 = 4NO + 6H 2 0. 
 
 To keep the gauze heated, electricity is supplied at 25 volts 
 and 135 amp., or 3.375 kw. The article mentioned does not 
 give the capacity of a converter, but from the description of 
 the ammonia pas purifying apparatus, Zeisberg assumes that 
 three converters handle from 275 to 2750 pounds of NH 3 per 
 day, hence one converter can handle about 528 pounds per day, 
 or 22 pounds per hour. At a conversion efficiency of 85 per 
 cent, 10 pounds of NH 3 will produce 31.5 pounds of HN0 3 , and, 
 100 pounds of HNO3 will therefore require 4.9 K.W.H. to 
 maintain the temperature in the catalyzer. 
 
 The amount of platinum required to effect the conversion is 
 about 329 grains per 100 pounds of HNO3 per 24 hours, and 
 the elements are renewed once a month. 
 
 The converters are constructed of iron up to the contact 
 element, and beyond this of iron lined with aluminum; at a 
 point far enough away from the converter for condensation to 
 occur, the gas lines must be made of an acid-resisting material, 
 such as chemical ware. Ostwald states that nickel steel is not 
 attacked by hot nitrous gases, and can therefore be used in 
 place of the aluminum-lined iron. 
 
 One design of plant embracing the Ostwald process is shown 
 diagramatically in Fig. 35b, the catalyzer being of a type de- 
 vised by Frank and Caro. The ammonia is purified by dis- 
 tillation in the still (7) and washed in the lime washers (8), the 
 gas passing from these washers to the ammonia and air 
 
158 COAL GAS RESIDUALS 
 
 mixer (12), a container for the ammonia gas, not shown in the 
 figure, being interposed between the ammonia-purifying plant 
 and the mixer. Air is supplied to the ammonia through the 
 blower (11), the properly proportioned mixture of air and am- 
 monia then entering the catalyzers (13). From the catalyzers 
 the nitrous oxide is led to the absorbing towers (16), where it 
 meets with a large cooling surface covered with nitric acid, 
 which continuously trickles down and through the stone filling 
 of the tower. The nitric acid and the water formed during the 
 reaction are both condensed in the tower to liquid nitric acid, 
 and any other of the higher oxides of nitrogen which may have 
 been produced are thus given an opportunity to further oxidize 
 and to come in contact with sufficient water to form nitric acid. 
 The nitric acid in the towers is used repeatedly, until the proper 
 degree of concentration is reached, all acid flowing from the 
 towers being cooled in the condensers (17) and returned from 
 the storage (18) through the blow tanks (19) to the top of the 
 towers. 
 
 A coal-gas plant carbonizing 500 tons of coal per day will 
 produce 500 X 5 = 2500 pounds of NH 3 per day, or 75,000 pounds 
 per month. As one pound of NH 3 contains 0.824 pound of N, 
 this amount of ammonia will give 61,800 pounds of N per month, 
 and at an 85 per cent conversion we will secure 
 
 61,800x0.85x4.5 ... 
 
 = 118 tons of absolute nitric acid per 
 
 month. A conversion of 85 per cent has been questioned, and 
 in order to be conservative a conversion of 75 per cent will be 
 used in this estimate, this giving 104 tons of absolute acid, or 
 say 196 tons of 53 %, or 36° Be. nitric acid per month. 
 
 Owing to the rapid fluctuation of the prices of all the materials 
 entering into a construction of this character at the present 
 time, an estimate covering the cost of a plant for producing 
 196 tons of acid per month can only be indicative, and it is 
 given as such in this statement. 
 
 Assuming that one catalyzer can convert 500 pounds of am- 
 monia per day, a plant for the above purpose would require 
 eight catalyzers, five of which would be in constant use, and 
 three in reserve, each catalyzer requiring 4067 grains of plati- 
 num, or 4067 X 8 = 32,536 grains total. 
 
Tfc> STOOAQE 
 
 »*H«1 
 
 (Facing Page 158) 
 
Missing Page 
 
AMMONIA 159 
 
 Under these conditions, the estimated cost of a plant for the 
 above production would be 
 
 Buildings, foundations and supports $20,000.00 
 
 Ammonia apparatus 12,000.00 
 
 Catalyzer " 10,000.00 
 
 Condensing and absorbing apparatus 25,000.00 
 
 Storage tanks 2,000.00 
 
 Air compressor and fan 4,000.00 
 
 $73,000.00 
 Contingencies, 7} % 5,475.00 
 
 S7S.475.0O 
 Platinum contacts, original investment, 32,536 grains, or 2113 
 grammes at $3.20 $6,761.60 
 
 Total cost $85,236.60 
 
 OPERATING EXPENSE 
 
 1. Ammonia. The raw ammonia liquor, containing 75,000 pounds of 
 NH S per month, can probably be purchased at 6 cents per pound of NHa, 
 or $4500.00 per month. 
 
 2. Steam. With a 1 per cent ammonia liquor, there will be required 
 75,000 x 23.53 = 1,764, 750 pounds of steam per month, which at 20 cents 
 per M., will cost $352.95. 
 
 3. Lime. Allow 0.8 pound of lime per pound of ammonia or 75,000 x 
 0.8 = 60,000 pounds, or 30 tons per month, which, at $8.50 per ton, will 
 cost $255.00. 
 
 4. Cooling Water. Allow 55,000 gallons per day, or 1,650,000 gallons 
 per month, equivalent to 220,000 cubic feet at 65 cents per M., or $143.00 
 per month. 
 
 5. Power. A plant of this size will require about 25 HP. for its opera- 
 tion, or say 13,500 K.W.H. per month, which, at 2 cents per K.W.H., 
 will cost $270.00. 
 
 6. Current for Catalyzer. The plant will produce 104 tons of absolute 
 HNOs per month, and as each 100 pounds requires 4.9 K.W.H. to main- 
 tain the temperature in the catalyzer, or say 5 K.W.H., the current per 
 month will be 10,400 K.W.H. at 2 cents, or $208.00. 
 
 7. Platinum. If the platinum contact is to be renewed each month, 
 each contact weighing 4067 grains, or say 264 grammes, five of these, or 
 1320 grammes, will have to be renewed monthly, and allowing 10 cents 
 
160 COAL GAS RESIDUALS 
 
 per gramme for the difference between scrap and new platinum, the re- ' 
 place platinum per month will cost $132.00. 
 
 8. Wages. Two men will be required per shift to operate valves and 
 cocks, and one additional helper per day, or 
 
 6 men at $2.50 = $15.00 
 1 man at 2.00 = 2.00 
 
 $17.00 per day, 
 
 or $510.00 per month, to which should be added $125.00 for superintend- 
 ance, making a total of $635.00 per month. 
 
 9. Repairs. The repairs on a plant of this size are estimated at about 
 $175.00 per month. 
 
 Summary of Cost of Manufacture. 
 
 1. Ammonia $4500.00 
 
 2. Steam 352.95 
 
 3. Lime 255.00 
 
 4. Water 143.00 
 
 5. Power : 270.00 
 
 6. Electric current 208.00 
 
 7. Platinum 132.00 
 
 8. Wages 635.00 
 
 9. Repairs 175.00 
 
 12% for interest and depreciation on $78,475.00, 
 
 per month 784.75 
 
 Total cost per month $7455.70 
 
 or say 2.30 cents per pound of 36° Be. acid. 
 
 Production — 324,000 lbs. of 36° Be. HN0 3 @ 6f! $19,440.00 
 
 Cost of manufacture 7,455.70 
 
 Estimated profit per month 11,984.30 
 
 The above values are of course only comparative, as the costs 
 of raw material, the efficiency of the plant, and the selling 
 value of the nitric acid are dependent upon local conditions and 
 market for the product. 
 
 If ammonuim nitrate, instead of nitric acid, is to be the end 
 product, the nitric acid as produced above must be neutralized 
 with an additional quantity of ammonia, this neutralization 
 being conducted in wooden vats, the resultant liquor being then 
 evaporated so as to cause the salts to crystallize out. The salts 
 are dried in hydro-extractors, and are ready for sale, while the 
 small amount remaining in the mother liquor is also dried and 
 disposed of as a second grade product. 
 
AMMONIA 161 
 
 An installation for the production of ammonium nitrate, besides 
 the apparatus specified above, requires in addition another am- 
 monia purifying plant, vacuum evaporators, hydro-extractors, 
 and other appurtenances, it being estimated that this additional 
 plant including buildings will cost approximately $45,000.00. 
 
 This will also require 63,000 additional pounds of NH 3 for 
 neutralization, and the total ammonia used will produce 270,000 
 pounds, or 135 tons of ammonium nitrate. 
 
 Cost of Manufacture. 
 
 Cost as given for nitric acid $7,455.70 
 
 63,000 lbs. NH a @ U 3,780.00 
 
 Additional labor @ $25.00 per day 750.00 
 
 Purifying additional NH 3 , evaporating and drying the salt 1890.00 
 
 12 % for interest and depreciation on $45,000.00, per month 450.00 
 
 Total cost per month $14,325.70 
 
 or say 5.30 cents per pound. 
 
 Production — 270,000 pounds ammonium nitrate at 15)!. . . . $40,500.00 
 Cost of Manufacture 14,325.70 
 
 Estimated profit per month $26,174.30 
 
 The profits here indicated are of course very high, but war 
 conditions haVe greatly increased the value of both nitric acid 
 and ammonium nitrate beyond normal figures; these prices 
 cannot continue, but as the installation can be quickly amortized 
 at present, the plant should be paid for when normal conditions 
 return, and a normal manufacturer's profit may then be expected. 
 
 In neither of the cases cited above has a boiler plant been 
 included, as it is presumed that the works contemplating such 
 an installation would have the steam to spare from its present 
 output. 
 
 Nitric Acid from Coal Gas. The Hausser process for the 
 production of nitric acid from coal gas lends itself excellently 
 to such coke-oven plants as have an excess of gas, or in such 
 localities where nitric acid may have more commercial value 
 than would ammonium sulphate, and consists of converting the 
 nitrogen in the gas into nitric acid by explosion. A large scale 
 trial plant operating under this process was built at the Nurem- 
 berg Gas Works, Germany, before the war, but it is understood 
 that this plant has been considerably modified since then. 
 
162 COAL GAS EESIDUALS 
 
 The explosion of the gas is effected in steel explosion-proof 
 vessels, these vessels being surrounded by a water-jacket, and 
 to which the air and gas were admitted under pressure through 
 separate valves in the original plant. Ignition of the charge 
 was accomplished by means of a high-tension electric spark 
 derived from a magneto attached to one side of the vessel, 
 this magneto being connected with, and driven from, the same 
 mechanism which operates the air and gas valves. 
 
 The chemical reactions involved in this process are of an ex- 
 ceedingly simple nature, the nitrogen combines with the oxygen 
 to form nitric oxide, or 
 
 N 2 + 2 = 2NO, 
 
 and these nitric-oxide mixtures are cooled and passed into a 
 gas holder, where sufficient time of storage is allowed for oxida- 
 tion to take place, or 
 
 2NO + 2 = 2N0 2 . 
 
 The nitrogen peroxide thus formed by oxidation is con- 
 ducted to counter-absorption vessels where, due to reaction with 
 water, it forms a mixture of nitric and nitrous acid, or 
 
 2N0 2 + H 2 = HNO3 + HN0 2 . 
 
 The nitrous acid thus formed is then oxidized to nitric acid 
 by means of an excess of oxygen. The acid obtained in this 
 manner varies in strength from between 30 and 50 per cent, 
 and the weak product is usually evaporated and re-distilled in 
 order to produce a concentrated solution. The waste gases 
 from the plant contain some nitric oxide, and they are therefore 
 treated in special washing towers, thus recovering the greatest 
 possible quantity of nitrous products. It is estimated that 
 between 2 and 3 per cent of the original nitric oxide gas is 
 discharged with the waste products. 
 
 It is understood that the original plant built at Nuremberg 
 has been considerably modified, this modification relating 
 particularly to the explosion chamber, these changes having 
 increased the efficiency of the process beyond expectations. 
 
 The later design of plant embraces a battery of chambers, or 
 bombs, the entire battery being connected to two compressors, 
 and instead of the gas and air being compressed together, as 
 in the old plant, they are now dealt with separately. High 
 
AMMONIA 163 
 
 temperature is essential during explosion, and therefore the air, 
 after compression, is passed through a specially designed pre- 
 heater operated by gas; owing to temperature conditions, it is 
 also essential that a small proportion of pure oxygen be regularly 
 mixed with the air in the compressor. 
 
 The latest form of container is driven through gearing from 
 an electric motor, and it is supplied with four valves, the speed 
 of the motor determining the number of explosions per minute. 
 It is also essential that the waste products be cooled rapidly, as 
 otherwise there would be a tendency for the nitric acid to 
 undergo decomposition, and a special exhaust-valve is therefore 
 provided for the purpose of facilitating a drop in temperature. 
 
 A condenser coil, immersed in water, is located immediately 
 after the outlet from the bomb, or explosion chamber, the water 
 resulting from the explosion being thrown down in this coil 
 and then drawn off. Practice requires that the explosions be 
 so regulated that they occur in each bomb at the rate of about 
 sixty per minute, and it is stated that, within limits, the greater 
 the capacity of the explosion chamber the larger will be the yield 
 per unit of gas. 
 
 This process can be applied to producer gas as well as to all 
 low-grade gas, and as a general statement it may be said that 
 the yield in nitric acid, employing coke-oven gas, will be about 
 6J- pounds of acid per 1000 cubic feet of gas exploded, but this 
 is probably a maximum condition. 
 
CHAPTER V 
 BENZOL 
 
 While the removal of benzol from gas is not of remuner- 
 ative interest to the coal-gas producer under the present 
 antiquated standard of gas sale by candle-power, it is of 
 special interest to the by-product coke-oven operator, and 
 as all of the preceding data applies to both of these pro- 
 ducers, the method of removing and purifying benzol is 
 herewith included. 
 
 Benzol and its homologues were first used as dissolving, 
 purifying and cleaning agents, and for many years their most 
 important function lay in the production of aniline dyes by 
 conversion into CeHsNOj and then into C 6 H 5 NH2, thus 
 forming a raw basic material which is used in the preparation 
 of artificial organic paints and dyes. 
 
 Benzol also forms the basis for the production of artificial 
 perfumes and in the preparation of pharmaceutic compounds, 
 as well as in the production of saccharine, a substitute for 
 sugar. 
 
 A very important industrial purpose is served by benzol 
 and its homologues in the production of explosives, such as 
 nitro-benzol and trinitro-toluol. 
 
 Benzol is also used for carburetting illuminating gas, an 
 important function where straight coal gas is sold on the 
 candle-power basis; but its most modern as well as impor- 
 tant use is its substitution for gasoline as motor fuel, and it 
 is claimed that if benzol is mixed with toluol all carbonizing 
 and sooting of cylinders and sparkers is avoided; this mixed 
 fuel also being non-freezing under ordinary conditions. 
 
 All of these uses have created a constant demand for 
 benzol, and the market price is such that its removal from 
 the gas is a very profitable undertaking. 
 
 F. Puening, of the Koppers Company, states that the 
 average American coal will produce about two gallons of 
 
BENZOL 165 
 
 benzol per ton carbonized, and if it is compared with gaso- 
 line as motor fuel the producer should receive 15 cents per 
 gallon, while the cost of production in a fair sized plant 
 should not exceed 6 cents per gallon; Puening also states 
 that if the benzol in the 14,800,000 tons of coal coked in 
 the 6200 by-product coke ovens in the United States per 
 year is extracted, 29,600,000 gallons having a value of $4,400,- 
 000 would be produced, thus making benzol recovery a con- 
 servation of resources which would add greatly in reducing 
 the cost of a ton of coke. 
 
 Benzol is one of the hydro-carbon compounds which 
 escapes with the gas produced during the dry distillation 
 of coal, and owing to the fact that it is a low volatile it is 
 retained in the gas at temperatures as low as 60° F., the 
 other hydro-carbons being higher volatiles and therefore 
 condensing at the various intermediate temperatures above 
 60°; the condensation of the other high volatile oils, how- 
 ever, dissolves and absorbs about 5 per cent of the total 
 benzol contained in the gas. 
 
 Toluol, xylol, and some other high boiling compounds, 
 such as solvent naphtha, are known as homologues of benzol, 
 and these are found in the volatile portion of the coal and 
 consequently in the resultant gas. 
 
 Benzol, CeHo, is a colorless fluid and it greatly resembles 
 gasoline; its specific gravity is 0.88, its boiling point 191° F., 
 and its flash point 17.6° F., solidifying at 32° F. into large 
 rhombic crystals which melt at 37.5° F. 
 
 Toluol, C 6 H 6 CH3, so closely approaches the appearance of 
 benzol that it is difficult to distinguish one from the other by 
 ordinary means; its specific gravity is 0.87, boiling point 230° 
 F., and freezing point -130° F. 
 
 Xylol, C-eH^CHs)", is a combination of three products, closely 
 resembling benzol and toluol and having a specific gravity of 
 0.87, the three boiling points being at 278.6° F., 282.2° F., and 
 287.6° F. 
 
 Solvent naphtha is a compound made up of several hydro- 
 carbons, but their exact combination is only partially known. 
 The name "solvent naphtha" is due to its very high solvent 
 effect on rubber. It has a specific gravity of 0.87, and a 
 boiling point of from 302° F. to 356° F. 
 
166 COAL GAS RESIDUALS 
 
 These characteristics of benzol and its homologues only- 
 hold true for these products in chemically pure form, and 
 considerable differences are found in the products as received 
 from coal gas or tar, due to the fact that as such by-products 
 they exist as more or less of a mixture, and repeated dis- 
 tillations are necessary to even approach the pure article, 
 but as far as the by-product producer is concerned, the 
 methods of recovery and purification to be explained later 
 will produce the usual commercial product, the rectifica- 
 tion for purity being left to the chemical manufacturer. 
 
 Germany is the largest producer of benzol, and there are 
 but few coke-oven plants without their benzol recovery 
 system; therefore, if it has been found to be so remunerative 
 in Germany to recover this by-product the example should 
 be followed in the United States, where some coke-ovens are 
 still operated without efficient by-product recovery. 
 
 Puening states in the "Gas Record" that the profits 
 accruing from benzol plants in Europe are very high, having 
 ranged from 50 to 100 per cent upon the investment during 
 the last few years. He also states that, due to several 
 circumstances, the profits in the United States have not 
 been so large, but as seen from the following tabulation 
 these expected profits are sufficient to make the investment 
 give very handsome returns. He bases his calculation on a 
 plant carbonizing 2000 net tons of dry coal per day, with 
 the recovery of two gallons of benzol and its homologues 
 per net ton of coal, or 
 
 Yearly Recovery 
 
 Benzol, 67 per cent 978,000 gallons 
 
 Toluol, 16 per cent 234,000 " 
 
 Xylol, 8 per cent 117,000 " 
 
 Solvent naphtha, 9 per cent 131,000 » 
 
 Total 1,460,000 gallons 
 
 With a gross income of: 
 
 1,460,000 gallons at 15 cents $219,000.00 
 
 Crude naphthalene, 330 net tons at $5.00 1,650.00 
 
 Regenerated acid of 40° Be\, 360 net tons at $6.00 2,160.00 
 
 Total gross income $222,810.00 
 
 The operating expense per year being: 
 
BEXZOL 167 
 
 Raw material, consisting of wash oil, sulphuric acid, caustic soda . .$15,000.00 
 Steam for distillation, air compressor, loading pumps, acid regen- 
 eration and cooling water 15,000.00 
 
 Electric power for water pumps, oil pumps, agitator, illumination. 6,200.00 
 
 Wages for three distillers and two helpers 5,000.00 
 
 Overhead expenses, fire insurance, maintenance and depreciation, 
 
 assuming the cost of the complete plant to be about $300,000.00, $30,000.00 
 
 Calorific loss of the gas 13,000.00 
 
 Total operating expense $84,200.00 
 
 giving a net profit of $222,810.00 - $84,200.00 = $138,610.00 
 or 46 per cent on the investment capital, and making the 
 cost of producing one gallon 
 
 $84,200.00 . 
 
 T&om = 5 - 8 cents - 
 
 This profit is of course dependent upon the selling price 
 of benzol and the cost of raw material, but as it seems to 
 be an established fact that benzol will soon be an active 
 competitor of gasoline for motor fuel, the selling market 
 will probably be governed by the price paid for gasoline; 
 the latter price is increasing daily, and benzol should be subject 
 to the same conditions. The prices paid for benzol from 
 1881 to 1913 are shown in Diagram IV. 
 
 The first commercial benzol plant was built by F. Brunck 
 in Germany in 1887, and since then the development of this 
 system has been very rapid; in spite of the fact that several 
 builders are engaged in erecting benzol recovery plants, the 
 system as laid down by Brunck still maintains, the differ- 
 ences lying only in details, therefore a complete description 
 of operation covering any one system will practically hold 
 good for all others. 
 
 THE KOPPERS SYSTEM OF BENZOL RECOVERY AND 
 PURIFICATION i 
 
 Benzol is recovered from coal gas after the tar and am- 
 monia have been removed; in plants where the ammonia is 
 removed from the gas by the water absorption process the 
 gas usually has a temperature of from 68° to 100° F., while 
 where the direct sulphatation in acid baths is practised the 
 
 1 F. Puening in "Gas Record." 
 
168 
 
 COAL GAS RESIDUALS 
 
 gas temperature is usually from 108° to 170° F., but, as the 
 treatment for benzol should be carried out at about 77° F., 
 the gas must be cooled before entering the benzol scrubbers 
 if its temperature exceeds this figure. After the gas has 
 been cooled to the required temperature it is admitted into 
 the benzol washers (A) (Fig. 36), where it is brought into 
 
 
 x> 
 
 Price of Benzol. 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 290 
 
 200 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 1 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 260 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 240 
 230 
 220 
 
 210 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 190 
 180 
 l?= 
 160 
 ISO 
 140 
 130 
 120 
 110 
 100 
 90 
 SO 
 70 
 60 
 BO 
 40 
 SO 
 20 
 IO 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 h 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 8 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 z> 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 k 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 f 
 
 3 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 ft 
 
 
 
 \ 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 \ 
 
 
 
 
 
 
 
 
 
 
 
 1 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 \ 
 
 
 
 
 
 
 / 
 
 \ 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 \ 
 
 
 
 
 
 / 
 
 s 
 
 \ 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 \ 
 
 
 s 
 
 •" 
 
 
 
 \ 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 \ 
 
 S 
 
 
 
 
 
 
 
 
 
 
 1 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 1/ 
 
 
 
 
 1 
 
 L 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 A 
 
 
 
 
 r 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 V 
 
 
 
 
 I 
 
 
 
 
 
 
 
 
 r* 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 l 
 
 ^ 
 
 V 
 
 V 
 
 
 
 
 " 
 
 
 ^ 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 8 
 
 N t 
 O o 
 
 33 
 
 3 <0 
 
 i 
 
 g 
 
 B 
 
 o 
 
 § 
 
 s 
 
 o 
 
 (O 
 
 o « 
 
 si 
 
 s 
 
 
 8 
 
 c- 
 
 
 o 
 
 8 
 
 0) 
 
 s 
 
 (T> 
 
 8 
 
 8 
 en 
 
 8 
 
 8 
 
 0> 
 
 8 
 
 8 
 
 0) 
 
 8 
 
 o 
 
 01 
 
 0> 
 
 50 
 
 0) 
 
 
 Diagram IV. — Prices of Benzol. 
 
 intimate contact with wash oils which at low temperatures 
 are capable of dissolving the benzol constituents and which, 
 after being heated to about 266° F., will again release these 
 constituents. 
 
 The principal oil used in Europe for this purpose consists 
 of tar oil of which not more than 5 per cent will distil 
 below 390° F., 90 per cent will distil between 390° and 
 572° F.; and the naphthalene content must not exceed 7 per 
 
BENZOL 
 
 169 
 
 -rail IT] 
 
 cent, while the water content 
 shall not be more than 1 per 
 cent, and the oil shall contain 
 only traces of anthracene. 
 
 This wash oil is taken from 
 the tanks (B) and (5 2 ) by the 
 pumps (C), and is pumped suc- 
 cessively over the benzol washer 
 (A), the current of the wash oil 
 being counter to that of the gas. 
 The wash oil is spread out over 
 the hurdles or trays contained in 
 the towers (A), and it absorbs 
 the benzol constituents up to 
 . within very small quantities, 
 g finally being delivered to the tank 
 ^ (Bi) in a saturated state. From 
 
 3 tank (Bi) the saturated wash oil 
 
 a 
 
 pq is delivered to the benzol re- 
 2 covery plant, where it is heated 
 a for the purpose of driving off the 
 W absorbed benzol constituents. 
 
 In order to use as small an 
 m amount of steam as possible for 
 gP heating the enriched wash oil, 
 Koppers causes all of the heat 
 used to distil the wash oil to be 
 transferred to the cold enriched 
 oil on its passage to the still. 
 This cold wash oil enters the 
 heat exchanger (D), where it is 
 heated to about 175° F. by benzol 
 and steam vapors issuing from 
 the still (E). This oil now passes 
 through a second heat exchanger 
 in which it is preheated to a 
 temperature of about 212° F. by 
 means of the hot de'benzolized 
 wash oil issuing from the still 
 (£). It is then heated to about 
 
170 COAL GAS RESIDUALS 
 
 266° F. by means of live steam in one of the superheaters 
 (G) for the purpose of driving off the water contained in it 
 and thus rendering it possible to separate the naphthalene. 
 
 This water is transferred to the oil from the wet gases, 
 the latter having been cooled to below their point of satura- 
 tion; again, it may happen that the oils are found to be 
 cooler than the saturated gas, in which case water is pre- 
 cipitated. The naphthalene is washed from the gas in a 
 manner similar to that of washing out benzol and its homo- 
 logues. Puening states that 1000 cubic feet of gas contains 
 0.05 pound of naphthalene at a temperature of 80° F., and 
 it therefore becomes necessary to liberate the naphthalene 
 absorbed by the wash oil, as its presence increases the vis- 
 cosity and decreases the absorbing power of the oil. 
 
 At a temperature of 266° F. in the superheater ((?), all 
 benzol, toluol, and water is expelled, and the oil with its 
 burden of xylol, solvent naphtha and naphthalene now enters 
 the still (E). The oil flows through the lower portion of 
 this apparatus and also through its individual chambers, the 
 steam being blown directly into the lowest part of the still, 
 thence traveling in a direction counter to that of the oil, 
 thus driving off all xylol, solvent-naphtha, and naphthalene 
 from the wash oil. The resultant mixture of benzol and 
 water vapors passes through the upper portion of the still, 
 where crude rectification is performed and where all liquid 
 entrained wash oil particles are separated from the vapors. 
 The benzol and water vapors issuing from the superheater 
 ((?) enter the upper portion of the still where they are recti- 
 fied in conjunction with the other vapors. 
 
 These vapors now issue from the top of the still at a 
 temperature of about 220° F. and enter the heat exchanger 
 (D), where they are nearly all condensed and, as stated before, 
 the enriched wash oil is preheated by this exchange of heat. 
 The remaining vapors and the condensates formed in the 
 heat exchanger no\/ enter the water cooler (H), where the 
 vapors are completely condensed and all condensates are 
 cooled down to about atmospheric temperature. These 
 condensates, water and light oils, enter the separator (J), 
 where they are separated due to their difference in specific 
 gravity. 
 
BENZOL 171 
 
 This so-called light oil is a mixture of crude and impure 
 benzol, toluol, xylol, solvent naphtha, naphthalene, and a por- 
 tion of the wash oil, and it is collected in the collecting 
 tank (K), where it is stored for further fractionation and 
 purification, while the water from the separator is led to 
 the waste liquor tank, or, if it contains any ammonia, it may 
 be sent to the ammonia plant. The wash oil, which has 
 thus been freed of benzol and naphthalene, leaves the still at 
 a temperature of about 257° F. and enters the heat exchanger, 
 where it transfers a portion of its heat to the enriched wash 
 oil, after which it enters the oil coolers (F), where it is cooled 
 by water to a temperature of about 77° F., being then 
 delivered to the wash oil tank (Bi), from whence it is again 
 pumped into the scrubbers. 
 
 The light oil is now transferred from the storage tank 
 (K) to the still tank (M), where it is fractionated into its 
 components, the still tank having a capacity of from 6000 
 to 12,000 gallons, as this larger capacity permits of more 
 decided fractionations than would a smaller one. The oil 
 is heated here by means of two internal-heating coils and 
 one steam spray; due to this heat, the distillation of the 
 light oil begins at a temperature of about 176° F., at which 
 point benzol is driven off, the vapors passing through the 
 rectifying column ( N) and then through the dephlegmator 
 (0). The construction of this rectifying column is similar 
 to that of the usual ammonia or spirit still. The dephleg- 
 mator is a water-cooled tubular cooler arranged to cool the 
 vapor mixture, the heavier portions condensing and draining 
 back through the rectifying column, while the frosh rising 
 vapors evaporate the lighter portions of this condensate. 
 
 From the dephlegmator the vapors thus rectified pass 
 into the cooler (P), which is provided with vertical water- 
 cooled tubes in the upper section, where the vapors condense 
 and the resultant condensate is cooled to about atmospheric 
 temperature. The lower portion of this cooler acts as a 
 separator, separating the benzols from water in those cases 
 where direct steam is used for distillation. The various 
 benzol products then drain from the cooler and separator 
 (P) into their respective storage tanks (Q). 
 
 After benzol, toluol, xylol, and solvent naphtha have been 
 
172 
 
 COAL GAS RESIDUALS 
 
 distilled off, any wash oil with its naphthalene burden remains 
 in the still tank (M). The products thus secured from 1000 
 gallons of light oil vary, this variation being not only due to 
 the character of the coal carbonized, but also to the method 
 of operating the light oil plant, but Puening states that the 
 following will represent the average result: 
 
 500 gallons of crude benzol, 
 
 120 gallons of crude toluol, 
 60 gallons of crude xylol, 
 70 gallons of solvent naphtha, 
 
 250 gallons of wash oil containing naphthalene remaining in 
 the still tank (M). 
 
 The boiling points and the amounts of products from this 
 distillation are shown in Diagram V, prepared by Puening; 
 
 356T 
 
 ~ 
 
 
 
 
 r. 
 
 KCT 
 
 / 
 
 
 
 / 
 
 
 U6* 
 
 
 
 
 
 ■) 
 
 
 
 
 
 
 OUST 
 
 
 
 
 
 
 zxrr 
 
 2I2T 
 IMT 
 I76T 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 
 XOQoWcra of O-ud© eenzot 
 
 i20Qollo»T3 
 
 6O0ol 
 
 TOG3I 
 
 2SOQ31 fJOsh Oil cortann-to 
 
 CrOdo Crude 
 Xylol 3o^-NOp(h. 
 
 floplhofenc 
 
 Diagram V. — F. Puening. Boiling Points of, and Products from the 
 Distillation of Light Oil. 
 
 it is readily seen that the boiling points differ from those of 
 the pure products, due to the fact that these products are 
 distilled from a mixture of different benzols, the difference 
 being especially noticeable near the points where one kind 
 of benzol should finish and the next one begin. 
 
 The temperature up to which indirect steam for distilling 
 can be used is dependent upon the pressure of the live steam 
 available. Generally speaking, it is necessary to use both 
 direct and indirect steam heating for temperatures at and 
 
BENZOL 173 
 
 above 280° F. By referring to Diagram V, it is seen that 
 the distillation of solvent naphtha and a portion of the xylol 
 follows the curve A-B at 280° F., while if the temperature 
 is raised above 280° and without the use of indirect steam, 
 distillation will follow the curve A-C. 
 
 The wash oil retained in the still tank (M) is drained into 
 cooling pans, where it is cooled by the atmosphere to crystal- 
 lize out the naphthalene, after which the oil is drained off and 
 returned to the oil circulating tank (B 2 ), thus again entering 
 the washing system. A small cooler, for the purpose of 
 cooling the oil to about 175° F., receives the wash oil and 
 naphthalene just before it enters the cooling pans, in order 
 to prevent the giving off of obnoxious odors. The naphtha- 
 lene thus deposited in the pans may be dried in a centri- 
 fugal, thus separating the small amount of oil remaining in 
 the crystals, this naphthalene being sold as a crude product, 
 or it may be mixed with the tar. 
 
 The crude benzols in the storage tanks (Q) are now further 
 purified by washing them separately in the washer (R) with 
 sulphuric acid, caustic soda, and water. This washer con- 
 sists of a large lead-lined vessel provided with an agitator 
 or centrifugal mixing arrangement; here the benzols are 
 washed with 66° B6, or 1.84 specific gravity sulphuric acid, 
 the amount used varying for benzol and toluol from S to 
 14 per cent by weight, and for xylol or solvent naphtha from 
 20 to 24 per cent by weight. The acid is supplied from the 
 storage tank (S), it passing through a meter ((Si) before 
 entering the washer. The mixer is located at a height which 
 permits of discharging the acid in the bottom of the washer, 
 near the level of the benzol, and to distribute it in such 
 manner as to obtain a thorough mixture. This mixing or 
 washing is performed in about thirty minutes, after which 
 the agitator is shut down and twenty minutes allowed for 
 the used acid to settle to the bottom of the washer, the acid 
 being used to separate the unsaturated hydro-carbons, prin- 
 cipally phenols and defines. The majority of these im- 
 purities in conjunction with the acid form thick resinous 
 compounds, which are insoluble in benzol and which, due to 
 their high specific gravity, settle out in the bottom of the 
 washer, but a small portion of these resinous bodies, es- 
 
174 COAL GAS RESIDUALS 
 
 pecially the polymeric olefines, dissolve in the benzol and thus 
 give it a brown color; redistillation will separate these dis- 
 solved olefines from the benzol and thereby cause the brown 
 color to disappear, the acid used being later regenerated. 
 
 The benzol is now washed with water in order to remove 
 the remaining acid, after which a 16° Be\ caustic soda solu- 
 tion is run into the washer for the purpose of neutralizing 
 any remaining traces of the acid in the benzol, this latter 
 solution being prepared in tank (T), and the proper quantity 
 required measured in tank (7\) before admittance into the 
 washer. After being properly mixed with the benzol, the 
 caustic soda is permitted to settle, and is then drained off, 
 the benzol being again washed with water in order to remove 
 any traces of caustic soda. The acid usually reduces the 
 volume of the benzol, due to the extraction of impurities, 
 this reduction amounting to from 5 to 7 per cent of benzol 
 and toluol, and it may be as high as 20 to 30 per cent of 
 xylol and solvent naphtha. 
 
 The benzol from the washer (R) is now delivered to the 
 still tank (£/), this latter vessel being of similar construction 
 to tank (M), but it is advisable to interpose storage tanks 
 for washed benzol between the washer (R) and the still 
 tank (U) in order to permit the washer to operate independ- 
 ently of the still. The still tank (U) is provided with heat- 
 ing coils, an especially high rectifying column (V), and a 
 dephlegmator (W). After passing through the rectifier and 
 dephlegmator, the vapors pass into the cooler (X), where 
 they are condensed and cooled to about atmospheric tem- 
 perature; the condensates from the cooler drain into the 
 receivers (Y), where the quality of the product is tested 
 before being admitted into the storage tanks (Z). 
 
 This latter distillation is dependent upon the character 
 of the final product desired. Puening states that distilla- 
 tion carried out as described above improves the limits of 
 the boiling points and that the finished product more closely 
 approaches the chemical characteristics of the pure benzols. 
 However, due to the fact that the largest demand for benzol 
 in the present market will probably require products of 
 rather inexact boiling points, the last distillation will only 
 have to separate the benzol from the resinous olefines held 
 
BENZOL 175 
 
 in suspension, and therefore but little attention need be paid 
 to the final boiling points. 
 
 Under these conditions this last distillation will be ac- 
 complished very quickly, and the vapors thus produced are 
 sent directly into the cooler (X), the rectifier and dephelg- 
 mator being by-passed. The benzols produced in this man- 
 ner are known as "purified benzols," and they will have the 
 same general characteristics as regards their boiling points 
 as have the "crude benzols" shown in Diagram V. 
 
 The water and caustic soda used in the washer are drained 
 to the waste sewer, but the acid is conducted to small tanks 
 where it is regenerated to 40° Be\ by mixing it with steam, 
 which condenses while heating the acid, the lighter parts of 
 the resinous compounds being evaporated at the same time; 
 after the steam is shut off and the contents of the tanks 
 cooled, the remaining portions of the resinous compounds 
 solidify, and can then be removed, the 40° Be\ acid being 
 sent to the ammonia plant. These acid boilers must be so 
 arranged as to prevent the emission of any of the resulting 
 vapors, the greater portion of the latter being conducted to 
 condensers, a smaller quantity being usually discharged into 
 the atmosphere. 
 
 This last distillation for the production of benzols approach- 
 ing the pure article in chemical characteristics requires about 
 five times the distilling period required for crude benzol. 
 The still tank (U) in the Koppers system is filled with about 
 10,000 gallons of washed benzol having boiling points as 
 shown in Diagram V, while the distillation is accomplished 
 at the boiling points shown in Diagram VI; in the latter 
 diagram the portion from A to B shows the forerunnings, 
 which consists of a few per cent of impurities, such as carbon 
 bisulphide, while the portion from B to C shows "pure 
 benzol," and that from C to D a mixture of benzol and 
 toluol, the temperature curve of the latter following the 
 line from C to D. 
 
 The amount of "pure benzol" recovered is claimed by 
 Puening to be about 55 per cent of the "crude benzol" run 
 into the washers, and while distilling 10,000 gallons of washed 
 crude toluol, in order to recover "pure toluol," the distilla- 
 tion follows the course of the curve shown in Diagram VII. 
 
176 
 
 COAL GAS RESIDUALS 
 
 264V 
 
 Ze6V 
 
 240V 
 
 230T 
 
 az'r 
 
 194V 
 
 175V 
 
 -v— 
 
 Mngs 55%Pore Benzol 
 
 Mixture of 
 Benzol and Toluol. 
 
 Diagram VI. — Boiling Points and Product — Pure Benzol. 
 
 2B4°F 
 266T 
 f46°F 
 Z20°r 
 2I2°F 
 194V 
 U6> 
 
 1 
 
 1 
 
 B , 
 
 C ./ 
 
 f 
 
 
 
 / 
 
 
 
 J A 
 
 
 
 . —v / 
 
 ^ v ^ 
 
 » 
 
 &%rAxi. *****■****. -sres&l 
 
 Diagram VII. — Boiling Points and Product — Pure Toluol. 
 
BENZOL 177 
 
 The distillation in this latter instance begins at A, a mixture 
 of benzol and toluol being evaporated up to the point B, 
 while "pure toluol" results from B to C. After the toluol 
 has been expelled the distillation follows the curve from C 
 to D, and produces a mixture of toluol and xylol; the pro- 
 duction of "pure xylol" will produce a curve similar to the 
 one shown in Diagram VII. Diagrams VI and VII for "pure 
 benzol" and for "pure toluol" both show a slight increase in 
 temperature, amounting to 1.1° F. for benzol, and to 1.8° F. 
 for toluol between the points B and C, an indication that 
 although these products are termed "pure," they are not ab- 
 solutely chemically so, since if they were, the curve would not 
 rise at all. Usually the mixtures produced between A and B, 
 and C and D in Diagrams VI and VII, are mixed and sold 
 as such. Table XVIIIA shows the benzol specification as pre- 
 pared by the "Deutsche Benzol Vereinigung," which company 
 controls the benzol trade of Germany; column 4 of this 
 table relates to the chemical purity of the product. 
 
 The Feld System: 
 
 In the Feld system (Fig. 37) the gas is washed in the 
 washers (A) with light oil in a manner similar to that of 
 Koppers, the saturated oil entering the receivers (l) and (2), 
 finally flowing from (1) to the wash oil tank (4), the fresh 
 oil to the washers being pumped from tank (5). From the 
 wash oil tank (4) the saturated oil is pumped into the regu- 
 lator (B), flowing from thence into the outflow heater (C), 
 rising up through the still preheater (D) and stoam pre- 
 heater (E), and finally entering the column still (F), where 
 it is distilled by means of both direct and indirect steam. 
 The wash oil flows from the bottom of the still and enters 
 the outflow preheater (C), where it serves to heat the wash 
 oil coming from the regulator (£), after which it flows 
 through the final cooler (G), where it is cooled by means of 
 cold water, then entering fresh oil tank (5), from whence it 
 is sent back into the washing system. 
 
 Light oil vapors mixed with water vapors leave the column 
 still (F) and pass through the still preheater (D), where 
 they serve to impart additional heat to the saturated wash 
 oil on the way to the still, the vapors being thus condensed 
 
178 
 
 COAL GAS RESIDUALS 
 
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BENZOL 179 
 
 in (D); this condensate enters the separator (H), where the 
 the water is separated from the oils, the water being led 
 away while the light oils flow into the oil receiver (J), the 
 latter being provided with a water pocket in the bottom for 
 the purpose of collecting any entrained water. 
 
 At stated intervals the still receiver (K) is charged with 
 light oils from tank (J), where they are fractionately sepa- 
 rated through distillation by means of direct and indirect 
 steam. The vapors from the still pass through the rectifier 
 (L), where the vapors of high boiling points, such as toluol 
 and xylol, are condensed, and the partially purified benzol 
 enters the dephlegmator (M). The temperature of the cool- 
 ing water in the dephlegmator is kept at about 80° C; the 
 benzol vapors pass through the pipe system of the dephleg- 
 mator without condensation, being later cooled and con- 
 densed by means of cold water in the final cooler (N), 
 passing from thence into the separator (0), and then into 
 the receiver (P), one of the latter being provided for each 
 of the distillates, as after the benzol has been removed the 
 toluol, xylol and solvent naphtha are driven off. 
 
 After distillation has been completed, the wash oils re- 
 maining in the still tank (K) are run into the cooling pans 
 (3), where the naphthalene crystallizes out. 
 
 The raw benzol is now led to the purifier (S), where it is 
 treated with sulphuric acid, caustic soda, and water as in 
 the Koppers system, after which it is again distilled in still 
 (U) and rectified, dephlegmated, and separated, the final 
 products being found in tanks (Z). 
 
 In the extraction of benzol direct contact water coolers 
 should be used for cooling the gas; usually they consist of 
 tall, cylindrical shells filled with wooden trays, the gas enter- 
 ing at the bottom and flowing in a direction counter to that 
 of the cooling water, which latter enters at the top. 
 
 Puening states that when tubular coolers were formerly 
 used, the gas being separated from the water through the 
 walls of the tubes, it was impossible to cool the gas to a 
 temperature as low as that in direct contact cooling, the 
 lower temperature being required for efficient benzol ex- 
 traction, as its absorption is best at low temperatures; 
 besides this, the naphthalene deposits prevented the cooling 
 
180 COAL GAS RESIDUALS 
 
 of the gas by checking the transfer of heat through the 
 tubes, while in the direct contact coolers the separated 
 naphthalene is washed out by the water passing through the 
 cooler, and thus occasional deposits will not affect the ex- 
 traction efficiency. 
 
 The naphthalene carried by the water flowing from the 
 coolers is easily separated out by crystallization. Instead of 
 contact coolers such as described above, Feld uses his vertical 
 centrifugal washers for cooling the gas by direct contact 
 with water. 
 
 Puening also states that hurdle washers are used to a 
 great extent in Germany and England for benzol absorption, 
 the hurdles in these towers being so arranged as to cause 
 the gas to take a zigzag direction in passing through, the 
 hurdles consisting of wooden planks about tI" wide spaced 
 ri" apart, their lower edges being serrated to assist the 
 dripping oil to distribute itself over the next lower lying 
 layer. 
 
 Hirzel uses three different kinds of washers, the first one 
 being known as a tuyre washer and consisting of a steel shell 
 provided with from 200 to 400 spray nozzles so arranged as to 
 fill the entire space on the inside with a fine spray of oil; the 
 second washer is built in sections, similar to an ammonia still, 
 the bottom of each section being provided with numerous 
 tubes covered with seal caps, the latter dipping into the oil 
 in each section, thus causing the gas, which enters the bottom 
 of the washer, to pass up and break the seal at each hood, 
 being thus mixed with the wash oil; the third washer is of 
 the hurdle type as explained above. 
 
 Feld uses his vertical centrifugal washer for benzol extrac- 
 tion, and it has been found exceedingly efficient on this duty, 
 throwing far less back pressure than any of the other devices 
 used for this purpose. 
 
 Koppers advises that steel be used in the construction of 
 oil heating, distilling and cooling devices, except in the column 
 rectifier, where cast iron is more suitable; steel is a suitable 
 material where the temperatures do not exceed 210° F., but 
 at this temperature and above cast iron should be used on 
 account of the corrosive effect of the hot oils containing pos- 
 sible slight traces of ammonia. 
 
BENZOL 181 
 
 Oil superheaters arc subjected to the most severe corrosive 
 conditions, and spare heaters should be provided, so arranged 
 that they can easily be shunted into or out of the system. 
 Koppers recommends that oil superheaters with spiral tubes 
 be used only in small plants, because repairs are generally very 
 high in cost due to entirely replacing the coils. Koppers also 
 arranges his apparatus for heating and cooling the wash oil 
 in such manner that each individual unit can easily be taken 
 out without disturbing the other apparatus. 
 
 As the cast iron stills do not corrode, it is not necessary to 
 provide spare stills, and Koppers states that stills have been in 
 operation for from 10 to 15 years without repairs. The con- 
 struction of oil heaters and coolers requires special attention, 
 because hot oils and benzol escape through the smallest open- 
 ing produced by expansion and contraction during changes in 
 temperature. 
 
 Hirzel does not agree with the other designers as regards 
 the method of distilling the oil and driving off the benzol 
 vapors, and has made his views in this particular a subject 
 of U. S. patent No. 991,205, issued May 2, 1911/ 
 
 He states that the ordinary methods of fractional distilla- 
 tion, or the separation of the components of a mixed liquid, 
 are usually based on the differences existing in the boiling 
 points of the components and, as explained above, the usual 
 method consists of a series of regulated and methodical boil- 
 ings at different temperatures, thus boiling away the compo- 
 nent of a lower boiling point from the one of a higher boiling 
 point in a column still. Hirzel states that the balancing effect 
 between the base heating and the cooling due to the skin 
 contact of the shell in the usual still, as well as the introduc- 
 tion of a relatively cool liquid at the top, causes a regular 
 diminution of heat upward, and by introducing a mixed 
 liquid at the top of the column and causing it to travel suc- 
 cessively through the liquid holding devices of the still, it is 
 boiled at a series of successively increasing temperatures; 
 thus the vapors due to the boiling liquid in each liquid hold- 
 ing device, or tray of the still, serve to boil the liquid in the 
 device next above, condensing therein the greater portion of 
 their high-boiling components, and adding the vapor of their 
 low-boiling components to that which is there evolved. 
 
182 COAL GAS RESIDUALS 
 
 Hirzel claims that this method, which is suitable for some 
 purposes, is not well adapted for boiling oils containing 
 many constituents. 
 
 As explained before, the boiling point of a mixture of mutu- 
 ally soluble liquids is perfectly definite with a given mixture 
 under given conditions, and depends not only upon the respec- 
 tive boiling points of its constituents and their relative pro- 
 portions, but also upon many other factors, such as vapor 
 tensions at different temperatures, mutual affinity and solu- 
 bility in liquid and vapor states, pressure, etc. 
 
 Hirzel states that when these conditions change, these 
 factors also change quite independently of each other, and that 
 with complex mixtures there is no simple relation between the 
 compositions of the vapors evolved in boiling at different 
 temperatures; or, in other words, there is no certainty as to 
 the result which will be obtained in running such a mixture 
 through the usual column still. 
 
 The subject of the above patent calls for simplified con- 
 ditions of distillation, and Hirzel claims that the results are 
 correspondingly improved by departing from the usual prac- 
 tice of relying solely upon differences in boiling points, but 
 rather depending upon differences in vapor tensions. 
 
 The liquid to be treated is converted into a large area film, 
 and, by a suitable arrangement of heating elements, the film 
 thus formed is uniformly heated to a temperature approach- 
 ing, but below the boiling point of the constituent to be re- 
 moved, or near its point of maximum tension. This film is 
 treated with a heated current of gas or vapor into which the 
 constituent in question volatilizes and from which it may be 
 later removed by condensation, or in other words, the opera- 
 tion consists of evaporation instead of the boiling of ordinary 
 fractional distillation methods, the still being so arranged as 
 to practically maintain a uniform temperature throughout the 
 entire height. 
 
 Hirzel's scheme is shown in Fig. 38, where the washers are 
 represented as (A), two or more being required, dependent 
 upon conditions, with the usual gas connections, the wash oil 
 being supplied to the first washer from tank (B) by means of 
 pump (Ci), the wash oil flowing from the bottom of the 
 washer into tank (D), and being forwarded from there by 
 
BENZOL 
 
 183 
 
 pump (C 2 ) to the top of the second washer (A x ), flowing from 
 thence into the saturated oil tank (E). 
 
 The saturated oil flows from (E) into the heat interchanger 
 (F), where it is heated by the vapors coming from the still 
 (G), and then flowing into the top of the still. In the still 
 
 Fig. 38. — Hirzel Benzol Plant. 
 
 the heated oil is deprived of its volatile constituents, or crude 
 benzol, the oil leaving the base of the still and flowing into 
 the cooler (H), and chiller (J), both being devices of the 
 tubular condenser type, the cooler being provided with an 
 inlet and an outlet for cooling water, while the chiller is sup- 
 plied with similar connections for carrying a chilling fluid. 
 
184 COAL GAS RESIDUALS 
 
 The wash oil is now ready for absorption purposes in the 
 washer again, and it is therefore pumped away from the 
 chiller by pump (C 3 ) and sent to the wash oil tank (B). 
 
 The Hirzel still is of peculiar construction, as it is not 
 intended that it shall be subjected to a differential heat from 
 the bottom to the top, as is usual in the ordinary column still, 
 but it is uniformly heated throughout, because Hirzel does 
 not wish to fractionate the liquid treated, as is usually done 
 by a series of successive boilings, but only to remove a single 
 desired constituent by exposing the liquid in a large film area 
 uniformly heated throughout to the proper temperature. 
 
 The still is provided with a series of annular trays, each 
 tray being supplied with a steam coil, all coils communi- 
 cating with a header on the outside of the still. As Hirzel 
 does not desire either skin cooling or differential heating, he 
 lags the entire still with insulating material to assist in main- 
 taining a uniform temperature. Each annular tray is supplied 
 with a hood over the central aperture, a current of heated 
 indifferent gas being sent upward through these apertures 
 against the descending filmed liquid to strip the liquid of its 
 volatile constituents. Hirzel uses either steam for this pur- 
 pose, or gas from the mains. 
 
 The gas (at the temperature at which the benzol still is 
 ordinarily run, about 266° F., steam may be considered a 
 gas) carrying the volatilized benzol leaves the still at the 
 top, passes around the tubes in the heat interchanger (F), 
 thus warming the inflowing wash oil by transferring some of 
 its heat, and then passes down into the condenser (K), where 
 it is cooled by means of water, the water leaving this con- 
 denser and entering the wash oil cooler (H). From the con- 
 denser (K) the condensed liquids enter the collecting tank 
 (L), and if steam has been used as the carrier, the condensed 
 water accumulates in the bottom of this tank from whence it 
 is drawn off, the condensed benzol flowing out of the top of 
 the tank to the storage vessel (M). 
 
 If gas from the mains has been used as the carrier, the gas 
 is removed from the tank (L) by means of the exhauster (N) 
 and returned to the gas system; this gas as it leaves the 
 mains and enters the still is saturated with benzol at the 
 temperature at which it enters, but by being heated in the 
 
BENZOL 185 
 
 still it is capable, according to Hirzel, of absorbing more benzol 
 and thus serves to remove this constituent from the wash oil, 
 depositing this excess in the condenser. 
 
 The benzol still used in the Gasser system is a radical departure 
 from all of the others, and consists of a series of long horizontal cells, 
 separated from each other, each cell being supplied with an in- 
 dependent source of heat. It is claimed that this method of con- 
 struction permits the saturated wash oil to flow to the cells at 
 rather low temperature, and then gradually to heat the various 
 cells to a degree corresponding to the decreased amount of benzol 
 carried by the oil. 
 
 The extended form of the cell also makes it possible to force 
 the oil through in a thin film, constantly bringing this film in 
 contact with fresh steam, it being claimed that this procedure is 
 very effective in liberating the benzol. Besides this, the method 
 of constructing this still is such that it can be arranged for any 
 capacity, the additional capacity being secured by adding indi- 
 vidual cells. 
 
 This same construction is also used for distilling off the par- 
 tially purified, as well as the purified, benzol, the construction of 
 the cells being such that continuous distillation is claimed, only 
 such an amount of benzol as can be distilled in a given time 
 being admitted into the still. 
 
 Figure 38a is a diagram of a light-oil plant for medium sized 
 works, and contemplates the removal of benzol and its homologues 
 from the gas and the driving off of the wash oil from the benzolized 
 oil, returning this wash oil to the system again. 
 
 The proper location for a plant of this character is undoubtedly 
 after the purifiers, or especially after the ammonia, has been re- 
 moved, thus avoiding an ammonia loss which might occur by 
 the deposition of salts due to the low temperature at which the 
 benzol plant should be operated. If located after the purifiers 
 an added advantage is secured in the removal of corrosion danger 
 due to the action of sulphur compounds. If located ahead of the 
 purifiers, there is a possibility of coating the purifying material 
 with oils which may be entrained and carried over in the gas, thus 
 rendering a portion of the purifiers inactive. 
 
 The plant shown in Fig. 38a contemplates the use of a con- 
 denser ahead of the benzol washer, in order to reduce the tem- 
 perature of the gas to as low a point as possible, 70° F. being 
 
Missing Page 
 
BENZOL 187 
 
 The wash oil used in America is usually a petroleum dis- 
 tillate, known as "straw oil," having a specific gravity of 0.88, 
 or 30° Be\, and of which at least 90 per cent boils between 
 450° and 630° F. 
 
 The average oil used will absorb between 3 and 4 per cent 
 of its own volume or weight in benzol and its homologues, and 
 between 75 and 100 gallons of oil are kept in circulation per 
 ton of coal carbonized, the volume of oil used of course de- 
 pending upon its absorbing power. A small portion of this oil 
 will have to be replaced from time to time, as some of it will 
 be carried over with the light-oil vapors from the still, but the 
 amount is small. 
 
 The principal difficulties encountered in Great Britain, owing 
 to oil washing, have been due to: water in the oil; emul- 
 sion of wash oil and water; thickening of oil; naphthalene; 
 lack of fractionation; corrosion, and air locks in the system. 
 
 Water in the oil is due to the deposition of water coming 
 from the gas, or it may come from the condensation of steam 
 in the still. Water separators should be applied ahead of the 
 benzol washers, or very large storage tanks should be used to 
 permit of a separation of water from oil by difference in specific 
 gravity. Very wet steam also accounts for this trouble, and 
 superheating the steam will avoid it. 
 
 If the wash oil is much cooler than the gas, emulsification 
 may occur, and it is therefore advisable to maintain the tem- 
 perature of the wash oil a few degrees F. higher than the gas. 
 
 Emulsion of wash oil and water may be due to the water 
 coming from the gas being thoroughly mixed with the wash oil 
 in the washers, or, if the benzolized oil is admitted into the still 
 at too low a temperature, condensation may take place. 
 
 Thickening of the oil may be due to the evaporation of lighter 
 fractions of the wash oil by steam distillation, this being es- 
 pecially so if the gas is free of tar-fog. If this thickening 
 occurs, it can be remedied by adding a quantity of especially 
 light creosote oil. 
 
 Naphthalene troubles may be due to the gas picking up 
 naphthalene from the oil, or to the extraction of the naphthalene 
 carriers from the gas. If due to the first cause, it will be well 
 to look into the naphthalene content of the wash oil, as if the 
 latter does not contain more than 7 per cent, this trouble 
 
188 COAL GAS RESIDUALS 
 
 should not occur. The use of open steam greatly reduces the 
 quantity of naphthalene in both the gas and the oil, and while 
 its use tends slightly to reduce the quality of the distillate, the 
 best results for both benzol and naphthalene are usually se- 
 cured when the total distillate consists of half water and half 
 crude benzol. 
 
 If the trouble is due to the removal of naphthalene carriers 
 from the gas, it can be corrected by adding solvents to the gas, 
 as explained in Chapter II. 
 
 Lack of fractionation was usually found in plants where tar 
 dehydration apparatus, without any modification, was used for 
 debenzolizing the oil. The resultant benzol was found to be 
 of poor quality, and priming of the wash-oil often took place. 
 
 Corrosion was found to be quite serious on the parts in con- 
 tact with hot benzolized oil, providing much ammonia was 
 present, and cast iron was found best adapted to withstand 
 this action. 
 
 Trouble with air-locks may be expected by a faulty arrange- 
 ment of heat exchangers, coolers, etc., in which proper allowance 
 has not been made for the escape of air, this condition requiring 
 that the oil always enter the exchangers at the lowest point 
 and pass out at the highest, otherwise it will be necessaiy to 
 apply air vents. 
 
 Mr. J. W. Shaeffer, of Milwaukee, in his paper read before 
 the American Gas Institute, gives the results of his studies in the 
 loss of heat value due to the removal of the light oils from 
 the gas. He determined the heating value of the light oil to 
 be 17,400 B.T.U. per pound of liquid light oil, and to this he 
 added the latent heat of vaporization of benzol, or 167 B.T.U. 
 per pound, thus making 17,567 B.T.U. the gross heating value 
 of one pound of light-oil vapor; this value is somewhat lower 
 than the corresponding value for pure benzol, the latter being 
 18,793 B.T.U. Shaeffer accounts for this by the presence of 
 carbon disulphide and other impurities of a lower heat value 
 in the light oil. The latent heat of vaporization was added in 
 order to make the results comparable with calorimeter tests in 
 which the light oil enters the calorimeter in the shape of 
 vapor. 
 
 Calorimeter tests were made on the gas before and after oil 
 scrubbing, and the difference in B.T.U. divided by the total 
 
BENZOL 189 
 
 B.T.U. of the gas before benzol scrubbing gave the percentage 
 of actual loss. The B.T.U. per cubic foot of gas after the 
 scrubbers was calculated from 
 
 a - 121,8276 
 1 -356 
 
 in which a = B.T.U. per cubic foot of gas before scrubbing, 
 
 b = gallons of light oil removed per cubic foot of gas, 
 35 = cubic feet of vapor per gallon of light oil. 
 
 The constant 121,827 is the product of 17,567, the B.T.U. 
 per pound of light oil, 7.3 the pounds per gallon of light oil, 
 and 0.95, a correction factor used for correcting the amount 
 of wash oil in the light oil, this amount of wash oil being assumed 
 at 5 per cent. 
 
 The average of all calculations showed a theoretical loss of 
 4.45 per cent, while the actual loss as determined by calorimeter 
 tests at the same plants was 4.62 per cent. The above averages 
 are low, because they include plants where the scrubbing re- 
 sulted in a comparatively low efficiency, and for a plant where 
 the scrubbing is accomplished with a fair degree of efficiency, 
 the loss in heat value will be higher, probably 5.8 per cent. 
 The conclusions reached by Shacffer were: 
 
 1. The actual loss in the gross calorific value of the gas due 
 to the removal of light oil is about 5.8 per cent at plants where 
 the removal takes place with a fair degree of completeness. 
 
 2. The theoretical loss is practically the same as the actual 
 loss. This is an average determination. 
 
 3. Some plants show a greater actual loss than the theo- 
 retical loss and some vice versa; all agree within the limits of 
 probable error in determination. 
 
 4. The observed loss due to the removal of the light oil is 
 considerably greater than the actual loss. This is accounted 
 for by the difference between the gross and net heating values 
 and an improper air supply to the burners. 
 
 5. The observed loss has led many gas engineers to believe 
 that the actual loss is greater than it really is. 
 
 6. The true value of fight oil in a gas lies in its effectiveness 
 per unit volume. 1.1 per cent of light oil by volume furnished 
 7.4 per cent of the total heat, which is formed when the gas is 
 burned. 
 
190 COAL GAS RESIDUALS 
 
 The removal of light oils from the gas therefore involves the 
 burning of more gas to secure a given heating effect, and if 
 10,500 cubic feet of gas are produced per ton of coal carbonized, 
 the loss of 5.8 per cent in heat value will be equivalent to the 
 loss of 609 cubic feet of gas. 
 
 The Public Service Gas Company of New Jersey recently 
 erected a light-oil washing plant at one of its stations, this 
 plant handling about 4,000,000 cubic feet of mixed coal and 
 water gas in 24 hours, between February 14 and April 1, 1917, 
 this plant treated 183,468,000 cubic feet of gas, of which amount 
 72,817,000 cubic feet was coal gas and 110,651,000 cubic feet was 
 water gas. The light-oil washing plant removed 87,589 gallons 
 of light oil from this volume of gas, thus averaging 0.477 
 gallon per 1000 cubic feet of gas washed. After deducting 
 waste and residuals, about 0.35 gallon of benzol products per 
 1000 cubic feet of gas resulted, this large amount being due to 
 the water gas, which contains more benzol than does coal gas. 
 
 The average heating value of the total mixed gas dropped 
 from 620 to 605 B.T.U. and as all the gas made was not 
 washed, this indicated a heating value of about 575 B.T.U. for 
 the washed gas. The illuminating power of the total gas, 
 washed and unwashed, dropped from 20 to 15.88 candles, this 
 indicating a loss of about 7.70 candles of the washed gas. 
 
 The quantity of artificial gas used for illuminating purposes, 
 and only a small percentage of this in open-flame burners, is 
 but a small proportion of that used for heating and power 
 operations. The caloric value is the main consideration in the 
 latter case, and this also is true of incandescent mantle lighting, 
 as the lighting effect is not increased with an increase in the 
 candle value of the gas used. 
 
 In view of the very small reduction in caloric value resulting 
 from the removal of the benzol homologues, and the future 
 developments in the industrial and chemical fields for these 
 products, the conservation of such products is a telling argument 
 in favor of the replacement of the obsolete candle value standard 
 for gas sales by the more logical calorific standard. 
 
CHAPTER VI 
 SULPHURIC ACID 
 
 Owing to the possible difficulty of securing acid for sulphate 
 manufacture, and the trouble and inconvenience this lack of 
 supply would entail, it may become necessary for the gas-works 
 or coke-oven operator to produce the acid on the plant, and a 
 description of the process of manufacture is therefore given 
 below. Some works may not be large enough to take up acid 
 manufacture individually, but it is possible for several neighbor- 
 ing works to combine, ship their spent oxide, with its sulphur 
 burden, to some central point, and there produce the necessary 
 acid for the benefit of the participants. 
 
 Sulphuric acid is made by either the "Contact" process, or 
 the "Chamber" process, either of which is primarily dependent 
 upon the sulphur gases generated by the burning of sulphur- 
 bearing material. 
 
 In the "contact" process, the spent oxide, bearing its sulphur 
 burden, is burned in a furnace with air, the sulphur dioxide and 
 other gaseous products of combustion thus produced being 
 later cooled and freed from dust and the greater portion of the 
 moisture content. These purified gases are then mixed with 
 air and passed through a catalyzer, the catalytic agent being 
 either ferric oxide or finely divided platinum. It is absolutely 
 necessary that the impurities be removed from the gas before 
 entering the catalyzer, as otherwise they would act detrimentally 
 on the contact material, and it would soon lose its catalytic 
 power. 
 
 The greater portion of the sulphuric acid made, however, is 
 provided by the "chamber" process, the usual product being 
 commercial oil of vitriol, which is not a pure or concentrated 
 article, as it contains only from 60 to 70 per cent of sulphuric 
 acid. The advantage of the "contact" process over the 
 "chamber" process is that in the former a pure, concentrated 
 acid is directly formed, but for purposes of ammonia absorption 
 the ''chamber" acid will, as a rule, answer all purposes. 
 
192 COAL GAS RESIDUALS 
 
 The apparatus required in the "chamber" process embraces 
 the sulphur-burning furnace, combined with a dust separator, 
 or purifier, and niter oven; the Glover tower for cooling the 
 gas, concentrating the chamber acid, and denitrating the nitrous 
 vitriol; the lead reaction chambers; and the Gay-Lussac tower 
 for absorbing the oxides of nitrogen. 
 
 In this process, the spent oxide is burned in a furnace, and 
 the gas thus produced in this sulphur furnace contains between 
 7 and 8 per cent of S0 2 , the balance being excess nitrogen and 
 air. The sulphur furnace may be of any type suitable for the 
 purpose, the one shown in the diagram, Fig. 38b, combining the 
 furnace proper, the dust chamber, and the niter oven in one 
 setting. The furnace selected should be of a type which will 
 continuously supply a sulphur gas of constant quality and 
 quantity, and it should produce as little dust as possible; it 
 should be provided with accurate means for regulating the air 
 supply to the burning sulphur, as otherwise some of the sulphur 
 might sublime, and it should also be capable of maintaining a 
 constant temperature at the inlet to the Glover tower. Per- 
 haps the best furnace to meet these requirements is one of the 
 Wedge type, although a large number of furnaces of the general 
 design shown in the diagram are used. The Wedge, or mechani- 
 cal type of furnace, will give a regular supply of gas, and the 
 composition of the gas will be more constant than that produced 
 in the hand-fired furnace. 
 
 The gas leaving the sulphur furnace (B) , Fig. 38b, contains dust, 
 and the greater portion of this dust is deposited in the dust 
 chamber (C) , under the nitre pots (D) , as the gas is caused to take 
 a tortuous path through the chamber. The gases now pass 
 over the niter pots, where the nitric acid required is introduced 
 in the form of a vapor, being generated by the action of sul- 
 phuric acid on sodium nitrate. In some plants liquid nitric 
 acid is introduced into the Glover tower, trickling through the 
 filling of the tower together with the nitrous vitriol. The re- 
 action in the niter pots may be expressed as 
 
 NaN0 3 + H 2 S0 4 = HN0 3 + NaHSO*, 
 
 or the action of sulphuric acid on sodium nitrate, produces 
 nitric acid and sodium acid sulphate, or niter cake. 
 This mixture of gases is now passed into the Glover tower (F) , 
 
SULPHURIC ACID 193 
 
 through the flue (E), an acid-resisting fan, not shown in the dia- 
 gram, being installed ahead of this tower to give the required 
 draught on the furnace. The gas enters the bottom of the tower 
 and passes up through the packing or filling, the hot gases thus 
 being brought into intimate contact with a down-flowing stream 
 of aqueous sulphuric acid, evaporating the water from this acid, 
 by giving up a portion of its heat, and thus also supplying a 
 portion of the steam or water necessary to produce the required 
 reactions in the lead chambers (G), (H), and (J). This moisture 
 and heat also assist to drive off the lower oxides of nitrogen from 
 the nitrous vitriol fed into the top of the tower, this denitration 
 process being assisted by the simultaneous concentration and 
 action of SO2, as is shown in the equation 
 
 2S0 2 (OH)(ONO) + S0 2 + 2H 2 = 3H 2 S0 4 + 2NO. 
 
 The Glover tower is usually erected upon an elevation which 
 will permit the acid to flow from the bottom of the tower 
 through the lead-cooling apparatus and to storage. This tower 
 is usually a lead cylinder about fl feet in diameter and from 
 25 to 30 feet high, the lead shell being supported by an external 
 frame of wood, and protected inside for about two-thirds of its 
 height by a lining of acid-resisting bricks. The interior is filled 
 with bricks, or small cylinders of clay to give adequate contact 
 surface to the ascending gases. The upper part of the tower is 
 provided with two lead vessels, one containing nitrous sul- 
 phuric acid, about two-thirds of the entire acid used, and the 
 other sulphuric acid from the chambers. 
 
 These two acids are distributed by sprays over the entire 
 area of the tower, where they mix and flow down through the 
 filling against the current of hot SO2 gas coming from the 
 burner. The nitrous vapors are liberated and escape into the top 
 of the tower from where, together with SO2 and air, they pass 
 out into the first lead chamber. The concentrated sulphuric 
 acid formed here, 60° Be\, collects at the bottom of the tower 
 and passes out to the cooling apparatus and storage. 
 
 The reactions in this tower, according to M. Neumann, are 
 of two kinds, that in the lower portion of the tower, where the 
 temperature is at its maximum, being a reducing action, and 
 in the upper portion, where the temperature is at a minimum, 
 an oxidizing action. 
 
03 
 
 o 
 
 B 
 
 cS 
 
 of 
 
 s 
 
 00 
 CO 
 
 s 
 
 $ d S £39 
 
 ' ■ A Q Z |L(5 
 
 <CiOOQlUU.UI^y-iSZO 
 
SULPHURIC ACID 195 
 
 The first of these reactions may be expressed as follows: 
 ONO 
 2S0 2 <^ + H 2 - 2H 2 S0 4 + N 2 3 , 
 
 OH 
 and 
 
 N 2 O s + S0 2 + H 2 = H 2 S0 4 + 2NO, 
 or, adding these together, 
 ONO 
 
 2S0 2 <^ + S0 2 + 2H 2 = 3H 2 S0 4 + 2NO. 
 
 OH 
 
 The first portion of this reaction causes heat to be evolved, 
 while the second portion, where N 2 3 is reduced to NO, causes 
 the absorption of heat, the formation of sulphuric acid in the 
 tower therefore requiring that a temperature be maintained in 
 the lower portion of from 150° to 160° C, 302° to 320° F., this 
 temperature being necessary to decompose the nitrous sulphuric 
 acid. The reactions in the upper portion of the tower, where 
 the gases are liberated together with NO, may be expressed as 
 
 and 2NO + = N 2 3 , 
 
 OH 
 
 2S0 2 + N 2 3 + 2 + H 2 = 2S0 2 <^ 
 
 ONO 
 or together, 
 
 OH 
 
 2S0 2 + 2NO + 3 + H 2 = 2S0 2 ^ 
 
 ONO 
 
 These two reactions are exothermic, and they are greatly 
 assisted by the gradual abstraction of the evolved heat, thus 
 causing the formation of some additional nitrosyl-sulphuric 
 acid, yielding sulphuric acid in the lower portion of the tower. 
 
 The vapors leave the Glover tower and enter the first lead 
 chamber, being so regulated in their flow that the acid formed 
 here has a concentration of 50° to 53° Be\; in the last chamber, 
 the train usually consisting of three, where no steam or water 
 is injected, the gases are poor, and the acid may be as low as 
 
196 COAL GAS RESIDUALS 
 
 40° Be\ About 75 per cent of the S0 2 coming from the Glover 
 tower is converted into sulphuric acid in the first chamber, 20 
 per cent in the second, and 4.5 per cent in the third, the latter 
 mainly acting to cool and dry the gas. 
 
 Sufficient steam or water spray must be admitted into the 
 lead chambers to transform practically all of the S0 2 into 
 sulphuric acid, this reaction concerning all the constituents 
 except nitrogen. 
 
 The reactions in the lead chamber are very complex, and are 
 not very well understood, but the equilibrium between the 
 various reactions seems to depend upon the quantitative re- 
 lation between the reacting substances, the temperature, and 
 the volume of the lead chamber. It may be advanced, however, 
 that nitrogen peroxide (N0 2 ), due to the loss of one atom of 
 oxygen, will be reduced to nitric oxide (NO), and the latter 
 in turn will unite with atmospheric oxygen and thus be re- 
 converted into nitrogen peroxide, therefore, when sulphur 
 dioxide, oxygen and nitrogen peroxide are mixed in the presence 
 of water, the series of reactions which take place will lead to the 
 formation of dilute sulphuric acid. The series of reactions which 
 thus take place may be expressed as 
 
 S + 2 = S0 2 , 
 3S0 2 -I- 2HN0 3 + 2H 2 = 3H 2 S0 4 + 2NO, 
 
 2NO + 2 = 2N0 2 , 
 2N0 2 + 2S0 2 + 2H 2 = 2H 2 S0 4 + 2NO, 
 
 but the theoretical interpretation of these reactions is still 
 under continuous discussion. 
 
 The lead chambers are erected on an elevation above the 
 ground, the floor of the chamber resting on a platform of wood. 
 The walls and roof are also supported on a wood frame in order 
 to supply a solid backing to the lead plates. The volume of 
 the lead chambers must be large, because the yield will greatly 
 depend upon properly mixing the gases; the first chamber is 
 the largest, the other two decreasing in size. 
 
 Since nitrogen comprises about four-fifths of the air, it be- 
 comes necessary to provide some means for the escape of this 
 nitrogen, and at the same time to prevent the escape of the 
 oxides of nitrogen as far as possible, and for this purpose the 
 chamber gases are caused to pass through the Gay-Lussac 
 
SULPHURIC ACID 197 
 
 towers. These towers are constructed entirely of lead, similar to 
 the Glover tower, supported on a wooden frame; the bottom of 
 the tower is set about 7 to 8 feet above the ground, the internal 
 diameter being from 7 to 8 feet, with a height of from 25 to 
 30 feet. (Both the Glover and the Gay-Lussac towers are also 
 built of a square cross-section.) The interior of the tower is 
 filled with acid-resisting surface material, supported on lead 
 gratings. The nitrous gases from the lead chamber enter the 
 bottom of the tower, and are brought into intimate contact 
 with sulphuric acid of 60° Be\ during the upward passage to 
 the outlet, the acid being fed in continuous streams from a 
 distribution system located under the roof at the top of the 
 tower. This acid spray dissolves the nitrogen oxides con- 
 tained in the gases, and forms the so-called nitrous sulphuric 
 acid, this being a solution of nitrosyl-sulphuric acid, or 
 
 ONO 
 S02^ in sulphuric acid. 
 
 OH 
 
 The acid used in the towers is collected in receivers of lead, 
 stoneware, or cast iron, pressure pots, or Montejus, being used 
 to raise the acid to the top of the towers by means of air 
 pressure. 
 
 After the gases exit from the Gay-Lussac towers, they are 
 practically purified from both S0 2 and oxides of nitrogen, and 
 they are then finally exhausted through a stack into the air, a 
 fan being sometimes located at this point to assist the draught. 
 
 From the foregoing it will be seen that sulphuric acid is 
 formed by combining sulphur dioxide, oxygen, and water by 
 means of the oxides of nitrogen, this acid being formed in the 
 lead chambers, precipitated on the walls, and collected on the 
 floor of the chamber. The acid in the last chamber is the purest, 
 but also the weakest (48° to 50° Be\) made in the system, and 
 it can be taken directly from here and concentrated in a sepa- 
 rate plant, whenever a purer and stronger acid than ordinary 
 chamber acid is desired. 
 
 Under ordinary conditions, however, this acid flows forward 
 into the second, or middle chamber, where it increases slightly 
 in strength by union with the stronger acid made in this warmer 
 chamber; due to the dust settling in this chamber the acid is, 
 
198 COAL GAS RESIDUALS 
 
 however, slightly contaminated. From the second chamber the 
 acid is fed to the first, where it reaches a strength of between 
 52° and 54° Be. From here the product may be partly with- 
 drawn for sale or use, or for concentration in a separate plant, 
 or it can be taken direct to storage. 
 
 If the above procedure is not followed the acid may be ele- 
 vated, by air pressure from Montejus, to the tank in the top 
 of the Glover tower, where it is mixed with nitrous vitriol, thus 
 concentrating it, but it will soon be diluted again from the steam 
 condensed by the two cold acids. It also loses strength due to 
 the conversion of any excess oxides of nitrogen or sulphur 
 dioxide contained in it, and assists to liberate the lower oxides 
 from their solution in the nitrous vitriol. 
 
 The nitric and nitrosyl-sulphuric acids are decomposed by 
 coming in contact with the sulphur dioxide in the hot gases, 
 and the resulting oxides of nitrogen are carried away with the 
 gases leaving the Glover towers. The acid, as it trickles down 
 through the filling material, receives further heat from the hot 
 ascending gases, and it soon begins to give up water, thus 
 becoming concentrated, and flowing from the tower into the 
 coolers, reaching a strength of between 59° and 61° B6. After 
 being cooled, a portion of this strong acid is fed to the top of a 
 second Gay-Lussac tower, where it descends and meets the 
 outgoing gases, dissolving and recovering the final portions of 
 nitrous gas not absorbed in the first Gay-Lussac tower. The 
 acid flows from the bottom of this second Gay-Lussac and enters 
 storage, and the excess of this circulating acid thus formed by 
 the continuous addition of weaker nitrous vitriol, becomes the 
 strong nitrous vitriol that is fed to the top of the Glover tower, 
 and thus returns the recovered nitrous gases to the process. 
 
 The acid coolers consist of a wooden tank, divided into com- 
 partments, lined with lead, each compartment containing a 
 series of lead pipe coils, the acid being fed into these coils by 
 means of a silica funnel; the outlet end of the coil is brought 
 to the top edge of the tank and ends in a trough leading to 
 storage. The tank is kept full of moderately cold water, and 
 the acid should not be admitted into the coils until they are 
 cooled by the water. 
 
 The use of nitric acid in the Glover tower has worked out 
 in practice to better advantage than when niter pots are used 
 
SULPHURIC ACID 199 
 
 in the oven, as the action of the niter pots is not as regular as 
 feeding nitric acid to the tower, the latter method being under 
 the complete control of the operator. 
 
 The 60° B6. acid is usually stored in steel tanks, or tanks 
 built of wood and lined with lead, and it can be shipped in 
 tank cars, tank wagons, carboys, or drums. 
 
 If the acid is to be used for benzol washing, it will have to 
 be concentrated, and it will therefore have to undergo a process 
 of evaporation. The usual method of procedure in this case is 
 to feed the acid into a cascade concentrating plant, this cascade 
 arrangement consisting of a series of fused silica trays or basins, 
 so set that the acid fed into the topmost basin overflows into 
 the one directly below, and so on down the entire series. The 
 heat applied under these basins slowly and gradually evaporates 
 the water from the acid, and the latter thus becomes concen- 
 trated more and more as it descends the series of trays, finally 
 entering pot-type coolers, from whence it flows into a lead 
 cooler, where the temperature is reduced to that of the atmos- 
 phere, and then on into storage. The hot gases are drawn from 
 the furnace to under the basins, and then pass out to a stack 
 communicating with the atmosphere, while the acid fumes, due 
 to evaporation, are drawn off in a separate flue and passed to a 
 coke, or similar scrubber. 
 
CHAPTER VII 
 TESTS 
 
 Test fob Tar (Dry) in Coal Gas 
 
 To test for tar, estimated as dry tar, since no account is 
 taken of the contained water in this instance, a tube such as 
 shown in Fig. 39 is inserted in the gas main; this tube should 
 be of such length as to permit it to project a sufficient dis- 
 tance into the gas main 
 to avoid aspirating the 
 skin of "dead" gas near 
 the outer circumference 
 of the main, this projec- 
 tion being usually from 
 four to six inches, de- 
 pending upon the diam- 
 eter of the main. 
 
 Fig. 39. — Tar Test. 
 
 The clean, dry tube, having a bore of about one-half inch, 
 is packed with glass wool as shown, and after having been 
 placed in a desiccator for from one to two hours, the tube 
 with its contents is accurately weighed and the weight noted. 
 In placing the tube in the main care must be taken to see 
 that the one-eighth inch orifice at the end is directly in the 
 current of gas, the outer end of the tube being connected to 
 the two wash bottles and meter as shown, the one bottle con- 
 taining H 2 S0 4 for the absorption of NH 3 , and the other con- 
 taining NaOH for the absorption of H 2 S. 
 
 After five or ten cubic feet of gas, the average temperature 
 of which has been observed and noted at the meter, have 
 passed through the apparatus, the glass test tube is connected 
 to a Chapman or similar filter pump and the air, which 
 is drawn through the H 2 S0 4 , is aspirated by means of the 
 pump until a constant weight is secured, at least three weigh- 
 ings being necessary to determine this weight. Then 
 
 Increase in Weight 
 
 ~ , , „ — =—. X 100 = grammes of tar. 
 
 Corrected Gas Volume ' 
 
TESTS 201 
 
 and this product multiplied by 15.432 will give the grains of 
 tar per 100 cubic feet. 
 
 The CioH 8 content of this tar may be determined by placing 
 two (2) wash bottles containing picric acid solution between 
 the tar tube and the pump, the estimation of naphthalene 
 being then carried out according to the method described 
 for the naphthalene test. 
 
 This method is one of usual practice, but Feld (see "Journal 
 fur Gasbeleuchtung," Jan., 1911) states that a gas saturated 
 with fluid vapors at the temperature of saturation cannot 
 absorb any more of these same vapors, and he based his 
 method of tar determination on this principle, and proceeds 
 as follows: 
 
 The gas, the tar content of which is to be determined, is 
 led through a dry "U" tube filled with absorbent cotton, the 
 exact weight of which is known, this*"U" tube being kept at 
 the same temperature as the temperature of the gas while the 
 test is being run, so that the absorbent takes up as little water 
 with the tar as possible. The drying of the tar in the "U" 
 tube is accomplished through the heat contained in gas which 
 has been freed of tar fog and water, at the same or at a 
 slightly lower temperature; as this gas is saturated with the 
 vapors of tar constituents, but is water free, it can absorb 
 water vapors, but no vapors of tar constituents. 
 
 Fig. 40 shows the arrangement used by Feld; this consists 
 of a rectangular, insulated vessel (a), the top being provided 
 with the necessary plug holes in which three "U" tubes (b), 
 (c), and (d) are inserted. The tubes (6) and (d) contain ab- 
 sorbent cotton (or glass wool), while (c) contains calcium 
 chloride; the vessel (a) is filled with a liquid having a tem- 
 perature equivalent to the one at which the tar test is to be 
 run. The tube (rf) is first dried until it has a constant weight, 
 the gas used for drying leaving the gas main at tube (e) and 
 entering the "U" tube (6), where it deposits all condensed 
 tar constituents, then giving up its water content to the 
 calcium chloride in (c), while in (d) the moisture contained in 
 the cotton only is found. 
 
 After tube (d) has been dried to a constant weight, it is 
 connected directly to the gas main by means of tube (e) and 
 by by-passing tubes (b) and (c) a measured quantity of gas 
 
202 
 
 COAL GAS RESIDUALS 
 
 is led through it, after which tube (d) is again connected to 
 and behind tubes (6) and (c) and dry, tar-fog free gas is 
 again drawn through it. During the test, as well as during 
 the drying period, tube (d) is kept in the vessel (a) filled with 
 warm water, and if the vessel (a) is well insulated the water 
 temperature will remain constant during the test, but if 
 necessary the warm water may be renewed from time to time 
 through the medium of inlet (/) and outlet (g). 
 
 Owing to the fact that the gas retains its temperature while 
 being passed through tube (d), the absorbed tar will contain 
 but little water, and the drying process is therefore very rapid. 
 Usually constant weight is secured in (d) as soon as a little 
 
 Tff Warm Water Intel 
 
 lb Meter 
 
 iNater Overflow 
 
 XaClz 
 
 Fig. 40. 
 
 Cbtfon 
 Feld Tar Test. 
 
 more gas is passed through for drying than was used for the 
 analysis, and tubes (6) and (c) are refilled from time to time 
 as required. 
 
 A tar test should either show the amount of tar fog re- 
 maining in the gas at a determined temperature and at a 
 predetermined point, or the amount of tar constituents which 
 will be separated from the gas by cooling to a certain tem- 
 perature at a point behind, or after the point where the test 
 is made. In the first case the temperature of the bath is 
 maintained at the temperature of the gas to be tested, and in 
 the latter case at the temperature to which the gas is to be 
 cooled. 
 
 If, for example, the tar content of the gas is to be deter- 
 mined directly ahead of the Pelouze, and if the temperature 
 
TESTS 203 
 
 of the gas at this point is 20° C. (68° F.), the temperature of 
 the bath should be maintained at 20° C, but if the test is 
 to determine how much tar will be separated by cooling the 
 gas from 60° to 30° C. (140° to 86° F.), the test should be 
 made at a point where the temperature of the gas is 60° C, 
 and the temperature of the bath should then be maintained 
 at 30° C, but in the latter case the gas should be passed very 
 slowly through the tubes, so that the temperatures may be 
 equalized. 
 
 If it should be desired to determine the amount of tar fog 
 present at 60° C, and how much tar oil or naphthalene will 
 be separated during a drop in temperature from 60° to 20° C. 
 (140° to 68° F.), two test tubes are placed behind each other, 
 the first being maintained in a bath of 60° C. and the second 
 in a bath of 20° C, while the drying process can be carried 
 out without any difficulty at the lower temperature, because 
 the gas at a lower temperature is always saturated with the 
 tar constituents which would be separated at a higher tem- 
 perature, but in no case should the drying be carried out at a 
 higher temperature than the one at which the test is to be 
 run. 
 
 Drying with gas consumes much less time than with air, 
 and consequently Feld's method has been adopted to quite 
 an extent in Europe. 
 Test for Naphthalene in Coal Gas: 
 
 In making a test for naphthalene in crude gas containing 
 NH 3 and H 2 S, wash bottle (A) containing a saturated solu- 
 tion of oxalic acid, and bottle (E) containing an NaOH 
 solution, Fig. 41, are placed before and after the two bottles 
 (B) and (C) containing picric acid as shown. Twenty-five cc. 
 of picric acid solution is placed in each of the bottles (B) 
 and (C), and the 250 cc. flask (D) should contain 75 cc. of 
 the same solution, while the NaOH equivalent of this 
 picric acid solution should be determined prior to each test, 
 unless more than one test per day is made. 
 
 The glass tube, which in the case of crude gas should be 
 filled with glass wool, is inserted into the gas main as shown, 
 and from two to ten cubic feet of gas should be passed 
 through the test, this volume depending upon the Ci H 8 
 content of the gas. After this volume of gas has been passed, 
 
204 
 
 COAL GAS RESIDUALS 
 
 the test is disconnected and the bottles. (B) and (C) washed 
 into flask (D), sufficient picric acid being added from a pipette 
 to make about 200 cc. of solution, after which the tube (F) 
 is placed in the flask and connected to a Chapman pump, 
 exhausting until all air bubbles cease to appear, after 
 which the flask is sealed by pulling the glass tube up into 
 the cork until the hole in the side of the tube no longer 
 communicates with the flask. The flask is then placed in a 
 water bath and boiled until all of the yellow crystals are 
 dissolved. 
 
 The flask should then stand for from six to eight hours, 
 or until the flask is cold, after which the volume is made 
 up to 250 cc. with distilled water; the contents are then 
 
 To Meter 
 
 A e C D E 
 
 Fig. 41. — Naphthalene Test. 
 
 agitated by shaking the flask, after which the contents are 
 filtered through a dry filter into a dry beaker. Take 
 25 cc. portions of the filtrate, using lacmoid as an indi- 
 cator, and titrate with N/i NaOH until a final drop produces 
 a green coloration. 
 
 One cc. of N/i NaOH is equivalent to 0.197 grain of Ci H 8 , 
 and if 175 cc. of the picric acid solution were used in the test, 
 and if this is equivalent to "A" cc. of N/i NaOH, and if the 
 250 cc. of filtrate is equivalent to "B" cc. of N/i NaOH 
 then 
 
 (A - B) 0.197 
 corrected gas volume Xl00=grams of Cl ° Hs P er 10 ° cubic 
 feet of gas. 
 
 It is very important that this test should be run and kept 
 at the same temperature as that of the gas being tested, and 
 this can best be done in the same manner as described in 
 Feld's tar test. 
 
TESTS 
 
 205 
 
 Test for Ammonia in Coal Gas: 
 
 The standard solutions used in this test consist of standard 
 HC1, Hydrochloric acid — 1 cc. = 0.5 grain NH 3 , and standard 
 NaOH, Sodium-hydroxide, 1 cc. = 1 cc. HC1 = 0.5 grain NH 3 ; 
 it is extremely important that these two solutions should 
 exactly check each other, or that 1 cc. of NaOH should 
 neutralize 1 cc. of HC1. 
 
 Ten cc. of the standard- HC1 solution, with sufficient 
 distilled water to provide a seal of from one-half to 
 three-quarters of an inch, should be placed in each of two 
 gas-washing bottles, these bottles being connected to the sam- 
 pling tube which should be inserted into the gas main for a 
 distance of about four inches, depending upon the diameter 
 of the main, using rubber tubing for sleeve connections but 
 bringing the ends of the glass tubes together in the sleeves. 
 A bottle containing a 20 per cent NaOH solution should be 
 placed before the wet experimental meter to absorb any H2S, 
 as shown in Fig. 42, after which gas is passed through the 
 
 solutions at the rate of 
 one-half to three-quarters 
 of a cubic foot per hour, 
 the meter reading and 
 the temperature of the 
 gas leaving the meter 
 being accurately ascer- 
 tained and noted. 
 Prior to the ammonia scrubber, the passing of one cubic 
 foot of gas through the solutions will be sufficient for the 
 test, but after the scrubber about five cubic feet should be 
 run through the bottles, the latter test being so regulated that 
 both tests will be run at the same time, these tests being for 
 the purpose of determining the amount of ammonia removed 
 from the gas in the scrubber. 
 
 If much tar is present in the gas, the sampling tube leading 
 to the wash bottles should be packed with glass wool in 
 order to prevent the tar entering the bottles. 
 
 After sufficient gas has been passed, the bottles are dis- 
 connected, the contents washed into a clean beaker of at 
 least 250 cc. capacity, a few drops of lacmoid indicator added, 
 and enough of the standard NaOH solution run in from a 
 
 TO Mater 
 
 HCl HC1 AtaOH 
 Fig. 42. — Ammonia Test. 
 
206 
 
 COAL GAS RESIDUALS 
 
 burette, so that a final drop of this solution causes a faint blue 
 coloration of the liquid, indicating that the end point is 
 reached, the solution being constantly stirred during titration. 
 The following calculation will give the quantity of ammonia 
 in the gas thus tested : 
 
 Amount of standard HC1 used 20 cc. 
 
 " " NaOH for titration, say 14 cc . 
 
 " HC1 combined 6 cc. 
 
 Gas registered by test meter 1.00 cu. ft. 
 
 Temperature of gas at outlet of meter 75° F. 
 
 Corrected volume: 1.0 x 0.960 0.96 cu. ft. 
 
 then 
 
 6^2 
 
 0.96 
 
 Excess 
 
 X 100 = 312.5 grains per 100 cubic feet, or 
 2 
 
 X 100 = grains per 100 cubic feet. 
 
 Corrected gas volume 
 
 In order to make a standard sodium carbonate solution, 
 take 55 grammes of dry Na^COs, C.P., and place it in a clean 
 porcelain or platinum crucible and heat gently, weighing 
 from time to time until the weight becomes constant, or until 
 no further loss in weight is noted. This heating is best ac- 
 complished by placing the crucible on a bent triangle and 
 placing the latter on wire gauze (Fig. 43), as the crucible 
 should only show a dull glow on the bottom. 
 
 When the weight has become constant take exactly 50.330 
 grammes of the carbonate and place it 
 in a 500 cc. stoppered volumetric \ | 
 
 flask, then add distilled water to the 
 mark on the flask. 
 
 To prepare the standard HC1 or 
 hydrochloric acid solution in which 1 
 cc. = 0.5 grain of NH 3 , place about 
 160 cc. of C.P. hydrochloric acid of 1.19 
 to 1.20 specific gravity and containing 
 37.5 per cent HC1 (absolute) in a 
 liter flask, and add enough distilled water to fill to the 
 mark on the flask. 
 
 After being well shaken, a clean burette is filled with this 
 solution and enough of it is run into a beaker, the latter con- 
 taining 25 cc. of the sodium carbonate solution and a few 
 
 wireGouza 
 
 Fig. 43. — Ammonia 
 
 Test. 
 
TESTS 207 
 
 drops of methylorange, the latter as an indicator, until a 
 final drop from the burette causes a red coloration which 
 remains permanent for one-half minute. 
 
 If it requires 24.5 cc. of HC1 for 25 cc. of NajCOa, then 
 each 980 cc. of the HC1 must be diluted to 1000 cc, 
 as 24.5 X 40 = 980 cc. equal to 1000 cc. NasCOa, and the 
 dilution is necessary as 1 cc. HC1 should exactly equal 1 cc. 
 of Na^COs solution. 
 
 After dilution, clean out burette and fill with the diluted 
 solution, titrating as before until 25 cc. of HC1 neutralizes 
 25 cc. of Na 2 C0 3 solution. 
 
 To prepare a standard NaOH solution, dissolve 80 grammes 
 of C.P. caustic soda in sufficient distilled water to make 
 1000 cc. With, a pipette place 25 cc. of this solution in a 
 beaker and run in the standard HC1 solution, using lacmoid 
 as an indicator, until a final drop produces the end point; 
 as it will probably require more of the HC1 to neutralize, 
 it will be necessary to dilute, as in the former case. 
 
 Then the number of cc. of HC1 x 40 = number of cc. to 
 which 1000 cc. NaOH solution must be diluted, or 25.6 X 
 40 = 1024 cc, or each 1000 cc. must be diluted to 1024 cc. 
 with distilled water. 
 
 After a thorough mixing, the solutions should be checked 
 until they both agree in strength, and as the accuracy of all 
 tests depends upon the solutions, a great deal of care should 
 be exercised in their manufacture; the sodium carbonate 
 (Na2C0 3 ) is the starting point, and if that is not just what 
 it should be the other two solutions will also fail. The car- 
 bonate should be well dried and not heated too high, other- 
 wise the salt will lose C0 2 , and the HC1 and NaOH must 
 exactly check. 
 Test for Cyanogen in Coal Gas : 
 
 In testing for cyanogen, as practised in the United States, 
 the following solutions are used: 
 
 Ferrous sulphate (FeS0 4 .7H 2 0) 10% Solution 
 
 Sodium hydroxide (NaOH) or Potassium 
 
 hydroxide 20% Solution. 
 
 Zinc sulphate (ZnS0 4 ), standard solution, the value of each 
 
 cc. in terms of (CN) 2 must be known. 
 Ferric alum in a 2 per cent solution is used as an indicator. 
 
208 COAL GAS RESIDUALS 
 
 Into each of three gas-washing bottles place 15 c.c. of the 
 FeS0 4 solution and 15 c.c. of the KOH solution, and shake 
 well to prevent the formation of lumps; then add sufficient dis- 
 tilled water to provide a one-inch seal for the gas inlet tubes. 
 
 Pass from 3 to 5 cubic feet of gas through the bottles, and 
 then wash the contents of the bottles into a suitable casserole, 
 being careful to remove all insoluble matter, and boil until no 
 more NH 3 is evolved, testing for this result with turmeric paper; 
 wash the hot solution into a graduated volumetric flask, cool, 
 and make up to the mark on the flask with distilled water, 
 shake well and then filter through a dry filter. 
 
 Place 50 or 100 c.c. portions of the filtrate in a beaker, and 
 acidify to slightly acid reaction with dilute H2SO4; then add 
 an excess of 10 c.c. of dilute acid, and titrate with the standard 
 zinc solution until a final drop placed in proximity to a drop of 
 the ferric alum solution on Schleicher and Schul drop reaction 
 paper No. 601 shows no blue zone at the point of contact, 
 care being taken that the iron solution does not touch the inner 
 circle of the test drop. Then 
 
 Total c.c. zinc used x equivalent (CN) 2 per c.c. X 100 _ 
 corrected gas volume ° 
 
 per 100 cubic feet. 
 
 To standardize the zinc solution, which may be made up 
 by dissolving say 16.5 grammes of zinc sulphate C.P. in 1000 
 c.c. of distilled water, a 1 per cent solution of potassium ferro- 
 cyanide, K 4 Fe(CN) 6 , C.P., made with clear yellow crystals, 
 is utilized. Twenty-five c.c. portions of this solution are made 
 up to 100 c.c. with distilled water and acidified with dilute 
 H2SO4, and after the addition of 10 c.c. more dilute acid, the 
 solution is titrated with the zinc solution until a final drop 
 does not produce the blue zone in contact with the drop of 
 indicator. It is essential that the same conditions as to vol- 
 ume and acidity be observed on both standardization and 
 test titrations. Dividing the equivalent of the 25 c.c. in 
 terms of the potassium ferro-cyanide by the number of c.c. of 
 zinc solution used, will give the value of each c.c. of zinc solu- 
 tion in terms of potassium ferro-cyanide, and multiplying this 
 result by 0.3695 and again by 15.432 will give the equivalent 
 of 1 c.c. ZnSC>4 solution in grains of cyanogen. 
 
TESTS 209 
 
 Determination of prussian blue in Cyanogen: 
 
 The method devised by Feld to determine the "blue" 
 content of cyanogen seems to give a simple, sure, and rapid 
 analytic process for the use of those chemists who are com- 
 pelled to analyze either spent oxide or cyanogen sludge and 
 to determine its value based on "blue" content. 
 
 Feld showed (see "Journal fur Gasbeleuchtung," 1904) by 
 a great many tests that the loss in cyanogen through the 
 formation of sulphocyanides is very large when the methods 
 of analyses devised by Moldenhauer and Lybold, as well 
 as Knublauch are used, due to the decomposition of the 
 cyanogen in contact with caustic potash. He further claims 
 that the method of Drehschmidt does not entail these losses, 
 but that the conversion of cyanogen combinations into 
 mercuric cyanide is a very slow process, and therefore leads 
 to discrepancies. 
 
 Table XIX gives the results of an examination of twelve 
 samples of spent oxide according to Feld's method, and 
 to the combined method of Knublauch and Drehschmidt 
 as recommended by Burschell and Lubberger. This combi- 
 nation analysis — breaking up cyanogen combinations with 
 caustic potash, separating the "blue" therefrom, decom- 
 posing the latter with mercuric oxide, and transposition into 
 ammonium cyanide and sodium cyanide, and then titrating 
 the latter according to Volhard — has to a great extent 
 displaced the individual methods of both Knublauch and 
 Drehschmidt. This combination method gives results which 
 generally coincide with the Knublauch method, although the 
 values are somewhat higher, but it is also subject to the 
 danger of sulphocyanide formation by long contact with 
 the caustic potash, while on the other hand the conversion 
 of the pure, fine-particled "blue" into mercuric cyanide 
 through boiling for fifteen minutes with mercuric oxide can 
 be thoroughly accomplished without any loss. 
 
 It is however, of extreme importance, after the "blue" 
 has been washed out, and during the remainder of the 
 analysis, to use only such reagents as are free from chlorine, 
 otherwise the result of the titration will be too high, and the 
 chemist is cautioned that the zinc dust used to decompose 
 the mercuric cyanide often contains chlorine which is difficult 
 
210 COAL GAS RESIDUALS 
 
 TABLE XIX. — "BLUE" CONTENT OF SPENT OXIDE 
 
 According to Feld 
 
 According to Knublauch- 
 Drehschmidt 
 
 No. 
 
 Value 
 
 % 
 
 Differ- 
 ence 
 
 Average 
 value % 
 
 Value 
 % 
 
 Differ- 
 ence 
 
 Average 
 value % 
 
 Difference 
 as com- 
 pared with 
 Feld." % 
 
 I 
 
 9.106 
 9.106 
 
 0.000 
 
 9.106 
 
 9.09 
 
 8.88 
 
 0.21 
 
 8.985 
 
 -0.121 
 
 Without 
 
 Cyanogen 
 
 Alkalies 
 
 8.869 
 8.715 
 
 0.154 
 
 8.792 
 
 
 
 
 
 II 
 
 9.164 
 9.221 
 
 0.057 
 
 9.193 
 
 8.89 
 8.72 
 
 0.17 
 
 8.805 
 
 -0.388 
 
 Without 
 
 Cyanogen 
 
 Alkalies 
 
 8.978 
 8.954 
 
 0.024 
 
 8.966 
 
 
 
 
 
 III 
 
 10.062 
 10.120 
 
 0.058 
 
 10.091 
 
 10.08 
 
 9.88 
 
 0.20 
 
 9.980 
 
 -0.111 
 
 Without 
 
 Cyanogen 
 
 Alkalies 
 
 10.043 
 10.081 
 
 0.038 
 
 10.062 
 
 
 
 
 
 IV 
 
 7.931 
 7.931 
 
 0.000 
 
 7.931 
 
 7.38 
 7.33 
 
 0.05 
 
 7.355 
 
 -0.576 
 
 V 
 
 8.959 
 8.897 
 
 0.062 
 
 8.928 
 
 8.00 
 8.00 
 
 0.00 
 
 8.000 
 
 -0.928 
 
 VI 
 
 8.812 
 8.812 
 
 0.000 
 
 8.812 
 
 8.20 
 7.92 
 
 0.28 
 
 8.060 
 
 -0.752 
 
 VII 
 
 8.100 
 8.188 
 
 0.088 
 
 8.144 
 
 7.66 
 7.67 
 
 0.01 
 
 7.665 
 
 -0.479 
 
 VIII 
 
 8.615 
 8.591 
 
 0.024 
 
 8.603 
 
 7.98 
 7.90 
 
 0.08 
 
 7.94 
 
 -0.663 
 
 IX 
 
 8.725 
 8.772 
 
 0.047 
 
 8.748 
 
 8.23 
 8.25 
 
 0.02 
 
 8.24 
 
 -0.508 
 
 X 
 
 9.040 
 8.963 
 
 0.077 
 
 9.002 
 
 8.43 
 8.43 
 
 0.00 
 
 8.43 
 
 -0.572 
 
 XI 
 
 9.470 
 9.527 
 
 0.057 
 
 9.499 
 
 9.29 
 9.29 
 
 0.00 
 
 9.29 
 
 -0.209 
 
 XII 
 
 8.925 
 8.925 
 
 0.000 
 
 8.925 
 
 8.68 
 8.68 
 
 0.00 
 
 8.68 
 
 -0.245 
 
TESTS 211 
 
 to remove. As a whole, this method requires the utmost 
 care, is rather difficult to carry out, requires a great deal of 
 time, and, as stated before, does not give results which cannot 
 be disputed. 
 
 The simplicity and rapidity of Feld's method has therefore 
 found much favor in Europe, especially as a method for 
 forming the basis of all analysis for the sale of the "blue," 
 where disputes so often arise between the chemist of the 
 producer and the one employed by the purchaser. 
 
 The principle upon which this method of Feld's is based 
 consists in separating the cyanogen in the form of hydro- 
 cyanic acid (HCN) from its various combinations by dis- 
 tillation, its absorption in a sodium solution, and the direct 
 determination of the resultant sodium cyanide with a nitrate 
 of silver (AgN0 3 ) solution. 
 
 In order to prevent the iron cyanide combinations being 
 transformed into sulphocyanides during the analyses, the 
 separation of the cyanogen (CN) 2 with caustic potash must 
 be done in a cold atmosphere and in the shortest possible 
 time, and the further action of the caustic must be neu- 
 tralized by adding magnesium chloride (MgCl 2 ) before heat- 
 ing. The formation of sulphocyanides is thus totally 
 obviated. Due to an excess of magnesium chloride solution 
 all free caustic is converted into chloride with the separation 
 of magnesium hydroxide, and ammonia and hydrogen sul- 
 phide are soon liberated by boiling: 
 
 K 2 S + MgCl 2 + 2H 2 = Mg(OH) 2 + 2KC1 + H 2 S, 
 or the addition of magnesium chloride and water to the 
 potassium monosulphide, forms magnesium hydroxide and 
 potassium chloride with the liberation of hydrogen sulphide. 
 
 Thereupon the iron cyanide combinations are decomposed 
 with a mercuric chloride solution, from which, due to the 
 excess of magnesium chloride and in spite of the alkalinity 
 of the solution, no mercuric oxide is separated, but mercuric 
 cyanide is produced with a separation of ferric hydroxide, or 
 2K 4 Fe(CN) 6 + 8HgCl 2 + 3Mg(OH) 2 = 6Hg(CN) 2 + Hg 2 Cl 2 
 + 2Fe(OH) 3 + 8KC1 + 3MgCl». 
 
 The complete decomposition is indicated by a light 
 brownish-red precipitate in the clear liquid. The excess of 
 
212 COAL GAS RESIDUALS 
 
 the amount of magnesium chloride over that equivalent to the 
 alkali should be such that from 3 to 4 molecules of MgCl 2 
 should be present for each molecule of HgCl 2 , because with 
 a less amount mercuric oxide, or basic mercury salts, is sepa- 
 rated, retarding the transposition into mercuric cyanide. 
 
 The solution of mercuric chloride must be added to the 
 boiling iron cyanide solution in a boiling condition, since if 
 one of the solutions should not be hot enough, mercuric 
 ferro-cyanide, which is difficult to decompose, will be formed. 
 
 The HCN is liberated from the mercuric cyanide by dis- 
 tilling with dilute sulphuric acid, absorbing it in a dilute 
 sodium solution. 
 
 As an indication of incomplete decomposition while boil- 
 ing with the mercuric chloride solution, a blue discoloration 
 1 will appear after the acid has been added, due to the action 
 of the iron sulphate on the mercuric ferro-cyanide. Traces, 
 which always escape during any process of decomposition, 
 will give the distillation residual a dirty-green color. The 
 excess of acid, after the removal of the amount equivalent 
 to the alkali present, must be equivalent to at least four 
 to five times the added mercuric chloride. 
 
 During distillation sulphur, evolved from free sulphur or 
 from polysulphides, also passes over and dissolves in the 
 sodium solution forming sulphide or thiosulphate, and thus 
 disturbs the titration with the silver solution, but the addi- 
 tion of about 0.5 gramme of lead carbonate solves this 
 difficulty. 
 
 According to whether the total cyanogen, or only that part 
 which is in combination with iron, is to be determined, the 
 free alkaline cyanide is either transformed into iron cyanide 
 by the addition of iron sulphate, 
 
 6NH4CN + 6NaOH + FeS0 4 = Na 4 Fe(CN) 6 + Na^SO* + 
 6NH3 + 6H 2 0, 
 
 or the cyanogen contained in the potassium cyanide is 
 liberated by heating with magnesium chloride: 
 
 2KCN + MgCl 2 + 2H 2 = Mg(OH) 2 + 2KC1 + 2HCN. 
 
 Feld recommends that a few c.c. of potassium iodide be 
 added for the titration of the sodium cyanide with the silver 
 
TESTS 
 
 213 
 
 solution, and states that the titration will be correct even 
 though chlorides are present. 
 
 The apparatus recommended by Witzeck for this analysis 
 is shown in Fig. 44. 
 
 The solution to be distilled is placed in the flask (A) 
 having a capacity of 700 c.c. A dropper (B) combined with 
 the Leibig condenser (C) is passed into the flask through a 
 rubber stopper on the one side, and a funnel (D) with a 
 glass cock on the other side, this funnel serving to admit 
 the sulphuric acid. The condenser terminates in a flask 
 (E) containing the sodium solution, this flask also being 
 
 Fig. 44. — Feld "Blue" Test. 
 
 connected to a safety tube (F) comprising three bulbs, the 
 latter also containing sodium solution. 
 
 The total cyanogen in spent oxide, or in cyanogen sludge, 
 is then determined as follows: 
 
 Two grammes of the mass (or 0.5 gramme of sludge) are 
 placed in a. mortar and 1 c.c. of 2/ N FeS0 4 .7H 2 solution 
 (278 grammes of iron vitriol per liter) is added together 
 with 5 c.c. of 8/ N NaOH (320 grammes per liter). This is 
 pestled for five minutes, and then 50 c.c. of 6/ N MgCl2 solution 
 (610 grammes per liter) is slowly added, the mixture being 
 constantly stirred, after which the entire mass is washed 
 with hot water into the distilling flask (A), sufficient hot 
 water being added to make up about 200 c.c; after boiling 
 about five minutes, 100 c.c. of boiling N/ 6 HgCl2 solution (27.1 
 grammes of sublimate per liter) is added to the mixture, and 
 then boiled again for ten more minutes, after which flask (A) 
 
214 COAL GAS RESIDUALS 
 
 is connected to the condenser, 30 c.c. of 8/ N H 2 S0 4 (392 grammes 
 per liter) are added, and then distilled for from 20 to 30 min- 
 utes; flask (E) is supplied with 20 c.c. of 2/ N NaOH (80 grammes 
 per liter). If the distillate is turbid, 0.5 gramme of lead car- 
 bonate is added to the measuring flask to which it has been 
 transferred. 
 
 Aliquot portions of the solution, after the addition of 5 c.c. 
 of N/ 4 KI (41.5 grammes per liter) are titrated with N/ioAgN0 3 , 
 the end point being indicated by the formation of a yellowish- 
 milky turbidity. One c.c. of N/i AgNO 3 (17.0 grammes per 
 liter) is equivalent to 0.009548 gramme of "prussian blue," 
 Fe 7 (CN) 18 . 
 
 DETERMINATION OF BENZENE, TOLUENE, AND SOLVENT 
 NAPHTHA IN LIGHT OILS 
 
 Method of D. Wilson and I. Roberts, Laclede Gas Light 
 Coyipany, St. Louis, Mo. 
 
 2100 c.c. of the crude oil is placed in a 2500 c.c. separatory 
 funnel, and 100 c.c. of sulphuric acid (1.84 specific gravity) is 
 slowly added. If much heating should take place, the acid 
 should be added a little at a time, and thoroughly mixed with 
 the oil by shaking for about one minute after each addition. 
 The mixture of oil and acid should not be cooled unless the 
 container becomes uncomfortably warm, and in that case it 
 should be cooled only to about 50° C. (122° F.), as a warm 
 temperature is especially conducive to good and quick separa- 
 tions when working with oils which are liable to heat very 
 much; oils of this character also will create considerable pressure 
 in the vessel and this condition should be relieved every few 
 seconds by inverting the container and opening the stopcock. 
 
 After all the acid has been added, the mixture is agitated 
 thoroughly for five minutes and then allowed to stand for 
 fifteen minutes, in order that the acid tar may settle out. At 
 the end of this period the sludge is carefully drawn off into a 
 beaker of water, and if it partially dissolves in the water, 
 forming a muddy solution, sufficient acid will have been added 
 to destroy the defines; if the sludge sinks to the bottom without 
 mixing with the water, a second 100 c.c. of acid must be added, 
 and the above procedure repeated. This acid treatment must 
 be repeated until the sludge partially dissolves in the water. If 
 
TESTS 215 
 
 any oil should follow the sludge, it can be blown off into a small 
 separatory funnel, the water can be drawn off, and the oil 
 then returned to the main sample. After the final acid treat- 
 ment, a little water is poured down the side of the container, 
 allowed to settle, and is then drawn off, thus removing the small 
 amount of sludge remaining in the bottom. 
 
 200 c.c. of warm sodium hydroxide solution (1 to 10) is now 
 added, and the mixture is thoroughly agitated for five minutes. 
 This should change the oil to a lighter color, but if this change 
 fails to appear, a little more of the hydroxide should be added. 
 The mixture is now allowed to settle until a clean-cut separa- 
 tion between the oil and the hydroxide appears, after which the 
 hydroxide is carefully drawn off. 
 
 Fig. 44a. — Apparatus for Live Steam Distillation. 
 
 This washed oil is now placed in the one-gallon copper still, 
 fitted for live steam distillation, Fig. 44a; this still is connected 
 to a Liebig condenser having a 30-inch jacket, ending in an 
 adapter on the lower end of the condensing tube. The washed 
 oil is heated in the still to the boiling point by means of a small 
 burner underneath, and a slow current of steam is then turned 
 into the still. (One half-gallon still, fitted with a large riser 
 to permit of draining back any condensation, will furnish suf- 
 ficient steam for this distillation.) Distillation is continued 
 until the flow of oil practically ceases, and the temperature 
 of the vapors leaving the still will gradually rise to 100° C. 
 (212° F.) before the distillation is finished. If a small flame is 
 maintained under the steam still during distillation, it will 
 prevent the condensation of much steam in the still, and also 
 increase the rate of distillation. The distillate is collected in 
 the same vessel as was used for the acid treatment, and the 
 water is drawn off when necessary. When the distillation is 
 
216 
 
 COAL GAS RESIDUALS 
 
 completed, the remaining water is carefully drawn off, and the 
 oil is measured at the original temperatures. 
 
 The oil is now dried by passing it through the calcium chlo- 
 ride dryer (Fig. 44b), and ff of its volume (measured at the 
 same temperature) is placed in the copper flask of the fractional 
 distillation apparatus. The calcium chloride dryer consists of a 
 1-inch by 30-inch tube, 
 the bottom of the tube 
 containing glass wool, the 
 remainder of the tube 
 being filled with fresh 
 granular calcium chloride. 
 
 caj=aiX=I 
 
 Fig. 44b. — Calcium Chloride 
 Dryer. 
 
 Fig. 44c. —Fractional Distillation Apparatus. 
 
 The fractional distillation apparatus for benzene, toluene, and 
 solvent naphtha is shown in Fig. 44c. A 2200-c.c. copper flask 
 (H) is fitted with a 40-inch Hempel column (G) having an in- 
 side diameter of 1 inch and filled with J-inch glass balls for 
 36 inches of its length. Attached vertically to the top of the 
 Hempel column, is a 21-inch Liebig condenser tube (C), with 
 the enlarged end up, this tube having an inside diameter of 
 
TESTS 217 
 
 I inch. The condenser tube is supplied with an open-top 
 Liebig condenser or dephlegmator jacket, 14 inches long (E), 
 and is fitted inside with a sealed tube (F), | inch by 16 inches 
 outside, set concentrically by means of vertical wires or by 
 protuberances on the tube itself. The top of the condenser 
 tube is fitted with a side-neck connecting tube (B), the de- 
 livery tube of which is bent down, and connected to a Liebig 
 condenser (J) with a 30-inch jacket set almost vertical. An 
 extra long condenser tube, with an adapter on its lower end, 
 is used, thus bringing the distillate down to a convenient place 
 for changing the receivers. The top of the apparatus is sup- 
 plied with a £-inch thermometer (A), and the condenser jacket 
 is provided with a dephlegmator thermometer (D). 
 
 All vapor temperatures are measured with a certified gas- 
 filled Centigrade thermometer, which is scaled for 3-inch 
 immersion, and has a range of 75° to 150° C. (167° to 302° F.) 
 in one-fifth degrees. The graduations start about one inch 
 from the immersion line, so as not to be covered by the cork, and 
 run about forty to the inch. All thermometers used for any of 
 the distillations are set with the immersion line at the bottom of 
 the cork, and the top of the bulb not higher than the bottom 
 of the side-neck opening. All volumetric measurements are made 
 in stoppered cylinders at a temperature of 20° C. (68° F.). 
 
 A Jena side-neck flask is used for distilling off the desired 
 portion of the naphtha remaining in the copper flask at the end 
 of the distillation period, a 12-inch Liebig condenser and a 3-inch 
 immersion certified thermometer, graduated in single degrees, 
 being used for this purpose. 
 
 The intermediates and known mixtures are tested in a 50-c.c. 
 Jena side-neck flask connected to a Liebig condenser having 
 a 12-inch jacket, and set at an angle of about 30° from the 
 horizontal, the lower end of the condenser tube being provided 
 with an adapter, while the temperatures are taken with the one- 
 fifth degree thermometer. The measuring cylinder is graduated 
 in 0.2 c.c. 
 
 The oil is purified in a 2500-c.c. separatory funnel, or in a 
 half-gallon aspirator bottle with a glass stop-cock inserted in 
 the upper stopper. 
 
 In order to prevent excessive distillation losses, the following 
 precautions should be taken: (1) While making tests, all con- 
 
218 COAL GAS RESIDUALS 
 
 densers and receivers in use should be cooled to at least 15° C. 
 (59° F.). (2) Extra select corks should be used for connections. 
 (3) While the distillate is being collected, a damp cloth should 
 be wrapped around the adapter and the receiver opening in 
 order to prevent the circulation of air. 
 
 If the fractional distillation apparatus is new, it will be best 
 to boil some solvent naphtha in the copper flask, until the whole 
 column is filled with vapors, then cool the flask with water and 
 allow the column to drain for 15 minutes, thus practically- 
 eliminating the error that would otherwise be due to oil re- 
 maining in the column. 
 
 After the washed oil has been placed in the copper flask, the 
 latter is connected tightly to the column, the dephlegmator 
 jacket is filled with water, and the flask is slowly heated. After 
 the oil begins to boil, the condensation in the column will be 
 rapid, and it may fill up with liquid, in which case the burner 
 should be withdrawn just long enough to allow the column to 
 drain. As soon as the distillate begins to come over, the rate 
 should be regulated to one drop each second. 
 
 The first runnings are taken off up to 79° C. (corrected) or 
 174° F. As benzene will at times come over below this tem- 
 perature, due to a small amount of water in the oil, it is best 
 to test this fraction for paraffins. The amount absorbed then 
 can be added to the benzene percentage. The rate is now 
 increased to 5 c.c. per minute, and the temperature will soon 
 become stationary (80°-80.5° C. or 176°- 177° F.) and re- 
 main so until most of the benzene is over. As soon as the tem- 
 perature begins to rise again, the rate is again decreased to one 
 drop a second until 81° C. (corrected) (178° F.) is reached, 
 when the benzene fraction is taken off. Up to this point, the 
 water leaving the dephlegmator should be kept at a temperature 
 well below 80° C. (176° F.), but high enough to permit the 
 distillation to take place at the stated rate. 
 
 The next fraction, which Wilson and Roberts call the B-T 
 intermediate, is distilled at the rate of one drop a second, the 
 dephlegmator water supply is shut off, and as soon as the 
 water boils it is quickly drawn off and replaced with a hot 
 solution of calcium chloride having a boiling point of 105° C. 
 (221° F.). This fraction is taken off when the temperature 
 reaches 109° C. (corrected) (228° F.). 
 
TESTS 219 
 
 Toluene now comes over, and while it is being collected the 
 rate is increased to 5 c.c. a minute after having wrapped the 
 Hemple column in asbestos paper. The temperature will soon 
 become constant around 110.5°-111° C. (231°-232° F.), and 
 when it starts to rise again the rate is decreased to one drop 
 a second, and the fraction is taken off when the temperature 
 reaches 111.5° C. (corrected) (233° F.). 
 
 The calcium chloride solution is now removed from the de- 
 phlegmator, the remainder of the column is wrapped in asbestos 
 paper, and the distillation is continued at the rate of one drop 
 a second until 137° C. (corrected) (247° F.) is reached, when the 
 distillation is stopped. The condenser is allowed to drain and 
 the T-S intermediate is taken off. 
 
 The asbestos wrappings are now removed, the copper flask 
 is cooled with water, and the column is allowed to drain for 
 15 minutes. The naphtha remaining in the flask may be 
 measured or distilled off to any desired temperature in the Jena 
 side-neck flask. The distillation loss should not exceed one or 
 two per cent, and this can be proportionately added to the vari- 
 ous constitutents if the content percentage of the oil is desired. 
 
 The intermediates can be taken at the average figures of 
 two-thirds of benzene to one-third of toluene, and two-thirds 
 of toluene to one-third of xylene, or their composition can be 
 determined as described below, and the amounts of benzene, 
 toluene, and solvent thus found is added to their respective 
 fractions. 
 
 To test the intermediates, 20 c.c. are distilled in the 50 c.c. 
 side-neck flask at the rate of one drop each second. The 
 corrected temperature is noted when the first drop falls into 
 the measuring cylinder, and also the percentage off at the even 
 degrees (corrected), i.e., 86, 88, 90, etc., and 114, 116, 118, etc. 
 Known mixtures are also distilled in the same way, and a per- 
 manent table is thus prepared. The composition of the in- 
 termediates can then be determined very closely by comparison. 
 The following benzene-toluene and toluene-xylene mixtures will 
 cover most intermediates: 60/40, 65/35, 70/30, 70/25. 
 
 The oils for preparing these mixtures can be obtained by dis- 
 tilling some of the pure oils in the fractionation apparatus; the 
 benzene and toluene should be taken off while the temperature 
 remains constant, and the xylene between 137° and 140° C. 
 
220 COAL GAS RESIDUALS 
 
 (279° and 284° F.). If the intermediate should be less than 
 20 c.c, toluene (prepared as above) is added to make 20 c.c, 
 and due allowance must later be made for this addition. The 
 50-c.c. flask should be supported by a 6-inch square of asbestos 
 board about \ inch thick, this asbestos being provided with a 
 1-inch diameter hole in the center. 
 
 The steam distillation may be omitted if desired, in which 
 case the oil should be measured after the washing with sodium 
 hydroxide. The naphtha residue from the fractional distilla- 
 tion is distilled to the desired temperature, and the oil remain- 
 ing in the flask is measured. 
 
 A low rate of distillation is desirable as the temperature rises 
 to 81° and 111.5° C. (178° and 236° F.), thus assisting to pro- 
 duce small intermediates. 
 
 In order to correct for the effect of barometric pressure, the 
 following fractional degrees should be added or subtracted for 
 each mm. of pressure above or below 760 mm. respectively: 
 
 81° C. (178° F.) 0.043 
 
 109° and 111.5° C. (228° and 236° F.) 0.047 
 
 137° C. (247° F.) 0.052 
 
 Benzene-toluene intermediate 0.045 
 
 Toluene-xylene intermediate 0.050 
 
 By starting with less than 2200 c.c. and using a smaller 
 copper flask, the time of distillation can be cut proportionately, 
 but the fractionating efficiency will remain the same in either 
 case; however, unless the oil contains at least 8 per cent of 
 solvent naphtha, it will be best to start with more than 1000 c.c. 
 of the crude oil, as about 70 c.c. of naphtha are required to send 
 the T-S intermediate over. 
 
 Sample Distillation. p er cent 
 
 Oil after steam distillation 93.3 
 
 Up to 79° C. (174° F.) 0.2 
 
 (due to water) 
 
 81° C. (178° F.) 44.7 
 
 109° C. (228° F.) 2.4 85.15 
 
 (exceptional) 
 
 111.5° C. (236° F.) 37.1 
 
 137° C. (247° F.) 2.0 65.35 
 
 Residue (solvent naphtha) 6.5 
 
 Distillation loss ...,,,,.,,.,,....,.,,.,.., 0,4 
 
TESTS 221 
 
 Summing up, and eliminating the loss, we have: 
 
 
 
 Per cent 
 
 
 Benzene 
 
 
 
 
 
 
 44.7 
 
 
 
 
 2.0 
 
 
 
 
 0.2 (loss) 
 
 
 
 
 47.1 
 
 ■ 47.1% 
 
 Toluene. 
 
 
 
 
 
 
 37.1 
 
 
 
 
 1.3 
 
 
 
 
 02 (loss) 
 
 
 
 
 39.0 
 
 39.0% 
 
 Solvent 
 
 
 6.5 
 
 
 
 
 7.2 
 
 7.2% 
 
 defines, 
 
 
 
 6.7% 
 
 100.0% 
 
 As the specific gravity of benzene at 15° C. (59° F). is very 
 close to 0.88, and that of toluene and solvent naphtha close to 
 0.87, specific gravities lower than these figures would indicate 
 paraffins. These can be tested for as follows: 
 
 10 c.c. of oil to be tested are carefully measured into a 110-c.c. 
 flask (with 10 c.c. graduations in tenths on the neck) and 25 c.c. 
 of a mixture containing two parts by weight of sulphuric acid 
 (sp. gr. 1.84), with one part of fuming sulphuric acid (20 per 
 cent), are added; this mixture is agitated thoroughly for several 
 minutes, and then allowed to stand for half an hour. The 
 volume is then made up to nearly 110 c.c. by pouring concen- 
 trated sulphuric acid down the side of the flask. After being 
 allowed to stand for an hour, the paraffins can be read off. 
 
 The advantages to be secured with this method, and with 
 the apparatus developed for the purpose, are: 
 
 (1) The benzene and toluene produced will distill within 
 1.5° C. when tested in a side-neck distillation flask. 
 
 (2) The intermediates will average 25 c.c, and should not 
 exceed 40 c.c. 
 
 (3) One fractional distillation will give the percentage of ben- 
 zene, toluene, and solvent naphtha within a few tenths of 1 per 
 cent. 
 
 (4) If the distillation is carried straight through at a rate 
 not exceeding one drop a second, the benzene and toluene 
 
222 COAL GAS RESIDUALS 
 
 produced will distill 97 per cent within one degree, the inter- 
 mediates remaining the same as above. 
 
 C. P. GRADES 
 
 Benzol Specifications. No standard specifications for benzol 
 and its homologues have as yet been adopted in this country, 
 but the following are used by the larger dealers in these products : 
 
 Pure Benzol. Under this specification 100 per cent of the 
 material shall distill entirely within a range of 2° C, and in 
 which range shall be included the true boiling point of benzol. 
 When three parts of the benzol are shaken with one part of 
 C.P. sulphuric acid and allowed to stand for fifteen minutes, 
 the benzol shall remain colorless and the acid may be slightly 
 colored. 
 
 Pure Toluol. 100 per cent of the material shall distill en- 
 tirely within a range of 2° C, in which range shall be included 
 the true boiling point of toluol. When three parts of the toluol 
 are shaken with one part of C.P. sulphuric acid and allowed 
 to stand for fifteen minutes, the toluol shall remain colorVss 
 and the acid may be slightly colored. 
 
 COMMERCIAL GRADES 
 
 90 Per Cent Benzol. 90 to 93 per cent shall distill below 
 100° C; specific gravity at 15.5° C. about 0.88. Color, water 
 white. In sulphuric acid wash test only very slight color shall 
 be imparted to acid (1.5 bichromate scale). 
 
 50 Per Cent Benzol. 50 per cent shall distill below 100°, 
 90 per cent below 120° C; specific gravity about 0.88. Color, 
 water white. In sulphuric acid wash test only slight yellow 
 color shall be imparted to acid. 
 
 Toluol. 90 per cent shall distill between 100 and 120° C; 
 specific gravity about 0.87. Color, water white. In sulphuric 
 acid wash test only slight yellow color shall be imparted to acid. 
 
 Solvent Naphtha. Distillation does not begin below 120° C. 
 and 90 per cent distills below 160°; specific gravity about 0.87 
 to 0.88. Color, water white to slightly yellow. Sulphuric acid 
 colored slightly yellow by wash test. 
 
APPENDIX 
 
 TABLE XX. — COEFFICIENTS FOR VOLUME CORRECTION 
 
 TO 60° F. OF GAS MEASURED AT DIFFERENT 
 
 TEMPERATURES 
 
 Temperature 
 
 Coefficient 
 
 Temperature 
 
 Coefficient 
 
 40 
 
 1.050 
 
 69 
 
 0.977 
 
 41 
 
 1.049 
 
 70 
 
 0.974 
 
 42 
 
 1.045 
 
 71 
 
 0.971 
 
 43 
 
 1.042 
 
 72 
 
 0.968 
 
 44 
 
 1.040 
 
 73 
 
 0.966 
 
 45 
 
 1.038 
 
 74 
 
 0.963 
 
 46 
 
 1.035 
 
 75 
 
 0.960 
 
 47 
 
 1.032 
 
 76 
 
 0.958 
 
 48 
 
 1.030 
 
 77 
 
 0.955 
 
 49 
 
 1.028 
 
 78 
 
 0.952 
 
 50 
 
 1.025 
 
 79 
 
 0.949 
 
 51 
 
 1.022 
 
 80 
 
 0.946 
 
 52 
 
 1.020 
 
 81 
 
 0.944 
 
 53 
 
 1.018 
 
 82 
 
 0.941 
 
 54 
 
 1.015 
 
 83 
 
 0.938 
 
 55 
 
 1.013 
 
 84 
 
 0.935 
 
 56 
 
 1.010 
 
 85 
 
 0.932 
 
 57 
 
 1.008 
 
 86 
 
 0.930 
 
 58 
 
 1.005 
 
 87 
 
 0.927 
 
 • 59 
 
 1.003 
 
 88 
 
 0.924 
 
 60 
 
 1.000 
 
 89 
 
 0.922 
 
 61 
 
 0.997 
 
 90 
 
 0.919 
 
 62 
 
 0.995 
 
 91 
 
 0.916 
 
 63 
 
 0.993 
 
 92 
 
 0.913 
 
 64 
 
 0.990 
 
 93 
 
 0.910 
 
 65 
 
 0.987 
 
 94 
 
 0.907 
 
 66 
 
 0.985 
 
 95 
 
 0.905 
 
 67 
 
 0.982 
 
 96 
 
 0.902 
 
 68 
 
 0.979 
 
 97 
 
 0.900 
 
224 
 
 APPENDIX 
 
 TABLE XXI. — SPECIFIC GRAVITY OF AMMONIUM SUL- 
 PHATE SOLUTIONS AT 15° C. (59° F.) (SCHIFF) 
 
 Specific 
 
 Percentage 
 
 Specific 
 
 Percentage 
 
 Specific 
 
 Percentage 
 
 Gravity 
 
 (NH 4 ) 2 S0 4 
 
 gravity 
 
 (NH 4 ) 2 S0 4 
 
 Gravity 
 
 (NH 4 ) 2 SO< 
 
 1.0057 
 
 1 
 
 1.1035 
 
 18 
 
 1.2004 
 
 35 
 
 1.0115 
 
 2 
 
 1.1092 
 
 19 
 
 1.2060 
 
 36 
 
 1.0172 
 
 3 
 
 1.1149 
 
 20 
 
 1.2116 
 
 37 
 
 1.0230 
 
 4 
 
 1.1207 
 
 21 
 
 1.2172 
 
 38 
 
 1.0287 
 
 5 
 
 1.1265 
 
 22 
 
 1.2228 
 
 39 
 
 1.0345 
 
 6 
 
 1.1323 
 
 23 
 
 1.2284 
 
 40 
 
 1.0403 
 
 7 
 
 1.1381 
 
 24 
 
 1.2343 
 
 41 
 
 1.0460 
 
 8 
 
 1.1439 
 
 25 
 
 1.2402 
 
 42 
 
 1.0518 
 
 9 
 
 1.1496 
 
 26 
 
 1.2462 
 
 43 
 
 1.0575 
 
 10 
 
 1.1554 
 
 27 
 
 1.2522 
 
 44 
 
 1.0632 
 
 11 
 
 1.1612 
 
 28 
 
 1.2583 
 
 45 
 
 1.0690 
 
 12 
 
 1.1670 
 
 29 
 
 1.2644 
 
 46 
 
 1.0747 
 
 13 
 
 1.1724 
 
 30 
 
 1.2705 
 
 47 
 
 1.0805 
 
 14 
 
 1.1780 
 
 31 
 
 1.2766 
 
 48 
 
 1.0862 
 
 15 
 
 1.1836 
 
 32 
 
 1.2828 
 
 49 
 
 1.0920 
 
 16 
 
 1.1892 
 
 33 
 
 1.2890 
 
 50 
 
 1.0977 
 
 17 
 
 1.1948 
 
 34 
 
 
 
 TABLE XXII. — SPECIFIC GRAVITY OF SAL-AMMONIAC 
 SOLUTIONS AT 15° C. (59° F.) (GERLACH) 
 
 Specific 
 
 Percentage 
 
 Specific 
 
 Percentage 
 
 Specific 
 
 Percentage 
 
 gravity 
 
 NH 4 C1 
 
 gravity 
 
 NH 4 C1 
 
 gravity 
 
 NH 4 C1 
 
 1.00316 
 
 1 
 
 1.03081 
 
 10 
 
 1.05648 
 
 19 
 
 1.00632 
 
 2 
 
 1.03370 
 
 11 
 
 1.05929 
 
 20 
 
 1.00948 
 
 3 
 
 1.03658 
 
 12 
 
 1.06204 
 
 21 
 
 1.01264 
 
 4 
 
 1.03947 
 
 13 
 
 1.06479 
 
 22 
 
 1.01580 
 
 5 
 
 1.04325 
 
 14 
 
 1.06754 
 
 23 
 
 1.01880 
 
 6 
 
 1.04524 
 
 15 
 
 1.07029 
 
 24 
 
 1.02180 
 
 7 
 
 1.04805 
 
 16 
 
 1.07304 
 
 25 
 
 1.02481 
 
 8 
 
 1.05086 
 
 17 
 
 1.07575 
 
 26 
 
 1.02781 
 
 9 
 
 1.05367 
 
 18 
 
 1.07658 
 
 26.297 
 
APPENDIX 
 
 225 
 
 TABLE XXIII. — SPECIFIC GRAVITY OF CAUSTIC AMMONIA 
 
 CONTAINING DIFFERENT QUANTITIES OF NH,. TEMP. 
 
 14° C. (57° F.) (CARIUS) 
 
 Specific 
 
 Percentage 
 
 Specific 
 
 Percentage 
 
 Specific 
 
 Percentage 
 
 gravity 
 
 NH, 
 
 gravity 
 
 NH a 
 
 gravity 
 
 NH, 
 
 0.8844 
 
 36.0 
 
 0.9021 
 
 28.2 
 
 0.9239 
 
 20.4 
 
 0.8848 
 
 35.8 
 
 0.9026 
 
 28.0 
 
 0.9245 
 
 20.2 
 
 0.8852 
 
 35.6 
 
 0.9031 
 
 27.8 
 
 0.9251 . 
 
 20.0 
 
 0.8856 
 
 35.4 
 
 0.9036 
 
 27.6 
 
 0.9257 : 
 
 19.8 
 
 0.8860 
 
 35.2 
 
 0.9041 
 
 27.4 
 
 0.9264 
 
 19.6 
 
 0.8864 
 
 35.0 
 
 0.9047 
 
 27.2 
 
 0.9271 
 
 19.4 
 
 0.8868 
 
 34.8 
 
 0.9052 
 
 27.0 
 
 0.9277 
 
 19.2 
 
 0.8872 
 
 34.6 
 
 0.9057 
 
 26.8 
 
 0.9283 
 
 19.0 
 
 0.8877 
 
 34.4 
 
 0.9063 
 
 26.6 
 
 0.9289 
 
 18.8 
 
 0.8881 
 
 34.2 
 
 0.9068 
 
 26.4 
 
 0.9296 
 
 18.6 
 
 0.8885 
 
 34.0 
 
 0.9073 
 
 26.2 
 
 0.9302 
 
 18.4 
 
 0.8889 
 
 33.8 
 
 0.9078 
 
 26.0 
 
 0.9308 
 
 18.2 
 
 0.8894 
 
 33.6 
 
 0.9083 
 
 25.8 
 
 0.9314 
 
 18.0 
 
 0.8898 
 
 33.4 
 
 0.9089 
 
 25.6 
 
 0.9321 
 
 17.8 
 
 0.8903 
 
 33.2 
 
 0.9094 
 
 25.4 
 
 0.9327 
 
 17.6 
 
 0.8907 
 
 33.0 
 
 0.9100 
 
 25.2 
 
 0.9333 
 
 17.4 
 
 0.8911 
 
 32.8 
 
 0.9106 
 
 25.0 
 
 0.9340 
 
 17.2 
 
 0.8916 
 
 32.6 
 
 0.9111 
 
 24.8 
 
 0.9347 
 
 17.0 
 
 0.8920 
 
 32.4 
 
 0.9116 
 
 24.6 
 
 0.9353 
 
 16.8 
 
 0.8925 
 
 32.2 
 
 0.9122 
 
 24.4 
 
 0.9360 
 
 16.6 
 
 0.8929 
 
 32.0 
 
 0.9127 
 
 24.2 
 
 0.9366 
 
 16.4 
 
 0.8934 
 
 31.8 
 
 0.9133 
 
 24.0 
 
 0.9373 
 
 16.2 
 
 0.8938 
 
 31.6 
 
 0.9139 
 
 23.8 
 
 0.9380 
 
 16.0 
 
 0.8943 
 
 31.4 
 
 0.9145 
 
 23.6 
 
 0.9386 
 
 15.8 
 
 0.8948 
 
 31.2 
 
 0.9150 
 
 23.4 
 
 0.9393 
 
 15.6 
 
 0.8953 
 
 31.0 
 
 0.9156 
 
 23.2 
 
 0.9400 
 
 15.4 
 
 0.8957 
 
 30.8 
 
 0.9162 
 
 23.0 
 
 0.9407 
 
 15.2 
 
 0.8962 
 
 30.6 
 
 0.9168 
 
 22.8 
 
 0.9414 
 
 150 
 
 0.8967 
 
 30.4 
 
 0.9174 
 
 22.6 
 
 0.9420 
 
 14.8 
 
 0.8971 
 
 30.2 
 
 0.9180 
 
 22.4 
 
 0.9427 
 
 14.6 
 
 0.8976 
 
 30.0 
 
 0.9185 
 
 22.2 
 
 0.9434 
 
 14.4 
 
 0.8981 
 
 29.8 
 
 0.9191 
 
 22.0 
 
 0.9441 
 
 14.2 
 
 0.8986 
 
 29.6 
 
 0.9197 
 
 21.8 
 
 0.9449 
 
 14.0 
 
 0.8991 
 
 29.4 
 
 0.9203 
 
 21.6 
 
 0.9456 
 
 13.8 
 
 0.8966 
 
 29.2 
 
 0.9209 
 
 21.4 
 
 0.9463 
 
 13.6 
 
 0.9001 
 
 29.0 
 
 0.9215 
 
 21.2 
 
 0.9470 
 
 13.4 
 
 0.9006 
 
 28.8 
 
 0.9221 
 
 21.0 
 
 0.9477 
 
 13.2 
 
 0.9011 
 
 28.6 
 
 0.9227 
 
 20.8 
 
 0.9484 
 
 13.0 
 
 0.9016 
 
 28.4 
 
 0.9233 
 
 20.6 
 
 0.9491 
 
 12.8 
 
226 
 
 APPENDIX 
 
 TABLE XXIII.— SPECIFIC GRAVITY OF CAUSTIC AMMONIA 
 
 CONTAINING DIFFERENT QUANTITIES OF NH 3 . TEMP. 
 
 14° (C. 57° F.) (CARIUS) — Continued 
 
 Specific 
 
 Percentage 
 
 Specific 
 
 Percentage 
 
 Specific 
 
 Percentage 
 
 gravity 
 
 NH 3 
 
 gravity 
 
 NH 3 
 
 gravity 
 
 NH 3 
 
 0.9498 
 
 12.6 
 
 0.9654 
 
 8.4 
 
 0.9823 ' 
 
 4.2 
 
 0.9505 
 
 12.4 
 
 0.9662 
 
 8.2 
 
 0.9831 
 
 4.0 
 
 0.9512 
 
 12.2 
 
 0.9670 
 
 8.0 
 
 0.9839 
 
 3.8 
 
 0.9520 
 
 12.0 
 
 0.9677 
 
 7.8 
 
 0.9847 
 
 3.6 
 
 0.9527 
 
 11.8 
 
 0.9685 
 
 7.6 
 
 0.9855 
 
 3.4 
 
 0.9534 
 
 11.6 
 
 0.9693 
 
 7.4 
 
 0.9863 
 
 3.2 
 
 0.9542 
 
 11.4 
 
 0.9701 
 
 7.2 
 
 0.9873 
 
 3.0 
 
 0.9549 
 
 11.2 
 
 0.9709 
 
 7.0 
 
 0.9882 
 
 2.8 
 
 0.9556 
 
 11.0 
 
 0.9717 
 
 6.8 
 
 0.9890 
 
 2.6 
 
 0.9563 
 
 10.8 
 
 0.9725 
 
 6.6 
 
 0.9899 
 
 2.4 
 
 0.9571 
 
 10.6 
 
 0.9733 
 
 6.4 
 
 0.9907 
 
 2.2 
 
 0.9578 
 
 10.4 
 
 0.9741 
 
 6.2 
 
 0.9915 
 
 2.0 
 
 0.9586 
 
 10.2 
 
 0.9749 
 
 6.0 
 
 0.9924 
 
 1.8 
 
 0.9593 
 
 10.0 
 
 0.9757 
 
 5.8 
 
 0.9932 
 
 1.6 
 
 0.9601 
 
 9.8 
 
 0.9765 
 
 5.6 
 
 0.9941 
 
 1.4 
 
 0.9608 
 
 9.6 
 
 0.9773 
 
 5.4 
 
 0.9950 
 
 1.2 
 
 0.9616 
 
 9.4 
 
 0.9781 
 
 5.2 
 
 0.9959 
 
 1.0 
 
 0.9623 
 
 9.2 
 
 0.9790 
 
 5.0 
 
 0.9967 
 
 0.8 
 
 0.9631 
 
 9.0 
 
 0.9799 
 
 4.8 
 
 0.9975 
 
 0.6 
 
 0.9639 
 
 8.8 
 
 0.9807 
 
 4.6 
 
 0.9983 
 
 0.4 
 
 0.9647 
 
 8.6 
 
 0.9815 
 
 4.4 
 
 0.9991 
 
 0.2 
 
APPENDIX 
 
 227 
 
 TABLE XXIV. — SPECIFIC GRAVITY OF SULPHURIC ACID 
 AT 15° C. (59° F.) (KOLB) 
 
 
 
 
 100 parts by weight 
 
 
 1 liter contains 
 
 
 
 
 
 
 correspond to 
 
 
 
 kilos 
 
 
 
 
 
 Per cent 
 
 Per cent acid 
 
 Chemically 
 
 
 
 
 
 
 
 
 
 
 pure 
 
 acid 
 
 
 
 Sp. Gr. 
 
 Dog. 
 Be 
 
 Deg. 
 Twd. 
 
 
 
 
 
 
 
 Acid of 
 60° Be 
 
 Acid of 
 
 
 
 
 
 
 
 50° Be 
 
 
 
 
 SOi 
 
 HjSO 
 
 of 60° 
 Be 
 
 of 50' 
 Be 
 
 SO, 
 
 H,SO« 
 
 
 
 1.000 
 
 
 
 
 0.7 
 
 0.9 
 
 1.2 
 
 1.4 
 
 0.007 
 
 0.009 
 
 0.012 
 
 0.014 
 
 1.007 
 
 1 
 
 1.4 
 
 1.5 
 
 1.9 
 
 2.4 
 
 3.0 
 
 0.015 
 
 0.019 
 
 0.024 
 
 0.030 
 
 Mil 1 
 
 2 
 
 2.8 
 
 2.3 
 
 2.8 
 
 3.6 
 
 4.5 
 
 0.023 
 
 0.028 
 
 0.036 
 
 0.045 
 
 1.022 
 
 3 
 
 4.4 
 
 3.1 
 
 3.8 
 
 4.9 
 
 6.1 
 
 0.032 
 
 0.039 
 
 0.050 
 
 0.062 
 
 1.029 
 
 4 
 
 5.8 
 
 3.9 
 
 4.8 
 
 6.1 
 
 7.7 
 
 0.040 
 
 0.049 
 
 0.063 
 
 0.078 
 
 1.037 
 
 5 
 
 7.4 
 
 4.7 
 
 5.8 
 
 7.4 
 
 9.3 
 
 0.049 
 
 0.060 
 
 0.077 
 
 0.096 
 
 1.046 
 
 6 
 
 9.0 
 
 5.6 
 
 6.8 
 
 8.7 
 
 10.9 
 
 0.059 
 
 0.071 
 
 0.091 
 
 0.114 
 
 1.0S2 
 
 7 
 
 10.4 
 
 6.4 
 
 7.8 
 
 10.0 
 
 12.5 
 
 0.067 
 
 0.082 
 
 0.105 
 
 0.131 
 
 1.060 
 
 8 
 
 12.0 
 
 7.2 
 
 8.8 
 
 11.3 
 
 14.0 
 
 0.076 
 
 0.093 
 
 0.120 
 
 0.149 
 
 1.007 
 
 9 
 
 13.4 
 
 8.0 
 
 9.8 
 
 12.6 
 
 15.7 
 
 0.085 
 
 0.105 
 
 0.134 
 
 0.168 
 
 1.075 
 
 10 
 
 16.0 
 
 8.8 
 
 10.8 
 
 13.8 
 
 17.3 
 
 0.095 
 
 0.116 
 
 0.148 
 
 0.186 
 
 1.083 
 
 11 
 
 16.6 
 
 9.7 
 
 11.9 
 
 15.3 
 
 19.0 
 
 0.105 
 
 0.129 
 
 0.164 
 
 0.206 
 
 1.091 
 
 12 
 
 18.2 
 
 10.6 
 
 13.0 
 
 16.7 
 
 20.8 
 
 0.116 
 
 0.142 
 
 0.182 
 
 0.227 
 
 1.100 
 
 13 
 
 20.0' 
 
 11.6 
 
 14.1 
 
 18.1 
 
 22.6 
 
 0.126 
 
 0.155 
 
 0.199 
 
 0.248 
 
 1.108 
 
 14 
 
 21.6 
 
 12.4 
 
 15.2 
 
 19.5 
 
 24.3 
 
 0.137 
 
 0.168 
 
 0.216 
 
 0.268 
 
 1.116 
 
 15 
 
 23.2 
 
 13.2 
 
 16.2 
 
 20.7 
 
 25.9 
 
 0.147 
 
 0.181 
 
 0.231 
 
 0.290 
 
 1.125 
 
 16 
 
 25.0 
 
 14.1 
 
 17.3 
 
 22.2 
 
 27.1 
 
 0.159 
 
 0.195 
 
 0.250 
 
 0.312 
 
 1.134 
 
 17 
 
 26.8 
 
 15.1 
 
 18.5 
 
 23.7 
 
 29.6 
 
 0.172 
 
 0.210 
 
 0.269 
 
 0.336 
 
 1.142 
 
 18 
 
 28.4 
 
 16.0 
 
 19.6 
 
 25.1 
 
 31.4 
 
 0.133 
 
 0.224 
 
 0.287 
 
 0.359 
 
 1.152 
 
 19 
 
 30.4 
 
 17.0 
 
 20.8 
 
 26.6 
 
 33.3 
 
 0.196 
 
 0.239 
 
 0.306 
 
 0.383 
 
 1.162 
 
 20 
 
 32.4 
 
 18.0 
 
 22.2 
 
 28.4 
 
 35.3 
 
 0.209 
 
 0.258 
 
 0.330 
 
 0.413 
 
 1.171 
 
 21 
 
 34.2 
 
 19.0 
 
 23.3 
 
 29.8 
 
 37.3 
 
 0.222 
 
 0.273 
 
 0.349 
 
 0.437 
 
 1.180 
 
 22 
 
 36.0 
 
 20.0 
 
 24.5 
 
 31.4 
 
 39.3 
 
 0.236 
 
 0.289 
 
 0.370 
 
 0.463 
 
 1.190 
 
 23 
 
 38.0 
 
 21.1 
 
 25.8 
 
 33.0 
 
 41:3 
 
 0.251 
 
 0.307 
 
 0.393 
 
 0.491 
 
 1.200 
 
 24 
 
 40.0 
 
 22.1 
 
 27.1 
 
 34.7 
 
 43.4 
 
 0.265 
 
 0.325 
 
 0.416 
 
 0.520 
 
 1.210 
 
 25 
 
 42.0 
 
 23.2 
 
 28.4 
 
 36.4 
 
 45.4 
 
 0.281 
 
 0.344 
 
 0.440 
 
 0.550 
 
 1.220 
 
 26 
 
 44.0 
 
 24.2 
 
 29.6 
 
 37.9 
 
 47.4 
 
 0.295 
 
 0.361 
 
 0.463 
 
 0.579 
 
 1.231 
 
 27 
 
 46.2 
 
 25.3 
 
 31.0 
 
 39.7 
 
 49.5 
 
 0.311 
 
 0.382 
 
 0.489 
 
 0.610 
 
 1.241 
 
 28 
 
 48.2 
 
 26.3 
 
 32.2 
 
 41.2 
 
 51.5 
 
 0.326 
 
 0.400 
 
 0.511 
 
 0.639 
 
 1.252 
 
 29 
 
 60.4 
 
 27.3 
 
 33.4 
 
 42.8 
 
 53.5 
 
 0.342 
 
 0.418 
 
 0.536 
 
 0.670 
 
 1.263 
 
 30 
 
 52.6 
 
 28.3 
 
 34.7 
 
 44.4 
 
 55.5 
 
 0.357 
 
 0.438 
 
 0.561 
 
 0.702 
 
 1.274 
 
 31 
 
 54.8 
 
 29.4 
 
 36.0 
 
 46.1 
 
 57.6 
 
 0.374 
 
 0.459 
 
 0.587 
 
 0.735 
 
 1.285 
 
 32 
 
 57.0 
 
 30.5 
 
 37.4 
 
 47.9 
 
 59.9 
 
 0.392 
 
 0.481 
 
 0.616 
 
 0.769 
 
 1.297 
 
 33 
 
 59.4 
 
 31.7 
 
 38.8 
 
 49.7 
 
 62.1 
 
 0.411 
 
 0.503 
 
 0.645 
 
 0.805 
 
 1.308 
 
 34 
 
 61.6 
 
 32.8 
 
 40.2 
 
 51.5 
 
 64.3 
 
 0.429 
 
 0.526 
 
 0.674 
 
 0.841 
 
 1.320 
 
 35 
 
 64.0 
 
 33.9 
 
 41.6 
 
 53.3 
 
 66.6 
 
 0.447 
 
 0.549 
 
 0.704 
 
 0.878 
 
 1.332 
 
 36 
 
 66.4 
 
 35.1 
 
 43.0 
 
 55.1 
 
 68.8 
 
 468 
 
 0.573 
 
 0.734 
 
 0.917 
 
 1.345 
 
 37 
 
 69.0 
 
 36.2 
 
 44.4 
 
 56.9 
 
 71.0 
 
 0.487 
 
 0.597 
 
 0.765 
 
 0.955 
 
 1.357 
 
 38 
 
 71.4 
 
 ■37.2 
 
 45.5 
 
 58.3 
 
 72.8 
 
 0.505 
 
 0.617 
 
 0.791 
 
 0.987 
 
 1.370 
 
 39 
 
 74.0 
 
 38.3 
 
 46.9 
 
 60.0 
 
 75.0 
 
 0.525 
 
 0.642 
 
 0.822 
 
 1.027 
 
 1.383 
 
 40 
 
 76.6 
 
 39.5 
 
 48.3 
 
 61.9 
 
 77.3 
 
 0.546 
 
 0.668 
 
 0.856 
 
 1.069 
 
 1.397 
 
 41 
 
 79.4 
 
 40.7 
 
 49.8 
 
 63.8 
 
 79.7 
 
 0.569 
 
 0.696 
 
 0.891 
 
 1.117 
 
 1.410 
 
 42 
 
 82.0 
 
 41.8 
 
 51.2 
 
 65.6 
 
 82.0 
 
 0.589 
 
 0.722 
 
 0.925 
 
 1.155 
 
228 
 
 APPENDIX 
 
 TABLE XXIV. — SPECIFIC GRAVITY OF SULPHURIC ACID 
 AT 15° C. (59° F.) (KOLB) — Continued 
 
 
 
 
 100 parts 
 
 by weight 
 
 
 1 liter contains 
 
 
 
 
 
 
 correspond to 
 
 
 
 kilos 
 
 
 
 
 
 Per cent 
 
 
 Chemically 
 
 
 
 
 
 
 
 
 
 
 pure 
 
 acid 
 
 
 
 Sp. Gr. 
 
 Deg. 
 Be 
 
 Deg. 
 Twd. 
 
 
 
 
 
 
 
 Acid of 
 60° Be 
 
 Acid of 
 
 
 
 
 
 
 
 50° Be 
 
 
 
 
 SOa 
 
 H 2 SO 
 
 of 60° 
 Be 
 
 of 50° 
 Be 
 
 S0 2 
 
 H2SO4 
 
 
 
 1.424 
 
 43 
 
 84.8 
 
 42.9 
 
 52.6 
 
 67.4 
 
 84.2 
 
 0.611 
 
 0.749 
 
 0.960 
 
 1.198 
 
 1.438 
 
 44 
 
 87.6 
 
 44.1 
 
 54.0 
 
 69.1 
 
 86.4 
 
 0.634 
 
 0.777 
 
 0.994 
 
 1.243 
 
 1.453 
 
 45 
 
 90.6 
 
 45.2 
 
 55.4 
 
 70.9 
 
 88.6 
 
 0.657 
 
 0.805 
 
 1.030 
 
 1.288 
 
 1.468 
 
 46 
 
 93.6 
 
 46.4 
 
 56.9 
 
 72.9 
 
 91.0 
 
 0.681 
 
 0.835 
 
 1.070 
 
 1.336 
 
 1.483 
 
 47 
 
 96.6 
 
 47.6 
 
 58.3 
 
 74.7 
 
 93.3 
 
 0.706 
 
 0.864 
 
 1.108 
 
 1.382 
 
 1.498 
 
 48 
 
 99.6 
 
 48.7 
 
 59.6 
 
 76.3 
 
 95.4 
 
 0.730 
 
 0.893 
 
 1.143 
 
 1.429 
 
 1.514 
 
 49 
 
 102.8 
 
 49.8 
 
 61.0 
 
 78.1 
 
 97.6 
 
 0.754 
 
 0.923 
 
 1.182 
 
 1.477 
 
 1.530 
 
 50 
 
 106.0 
 
 51.0 
 
 62.5 
 
 80.0 
 
 100.0 
 
 0.780 
 
 0.956 
 
 1.224 
 
 1.530 
 
 1.540 
 
 51 
 
 108.0 
 
 52.2 
 
 64.0 
 
 82.0 
 
 102.4 
 
 0.807 
 
 0.990 
 
 1.268 
 
 1.584 
 
 1.563 
 
 52 
 
 112.6 
 
 53.5 
 
 65.5 
 
 83.9 
 
 104.8 
 
 0.836 
 
 1.024 
 
 1.311 
 
 1.638 
 
 1.580 
 
 53 
 
 116.0 
 
 54.9 
 
 67.0 
 
 85.8 
 
 107.2 
 
 0.867 
 
 1.059 
 
 1.355 
 
 1.694 
 
 1.597 
 
 54 
 
 119.4 
 
 56.0 
 
 68.6 
 
 87.8 
 
 109.7 
 
 0.894 
 
 1.095 
 
 1.402 
 
 1.752 
 
 1.615 
 
 55 
 
 123.0 
 
 57.1 
 
 70.0 
 
 89.6 
 
 112.0 
 
 0.922 
 
 1.131 
 
 1.447 
 
 1.809 
 
 1.634 ' 
 
 56 
 
 126.8 
 
 58.4 
 
 71.6 
 
 91.7 
 
 114.6 
 
 0.954 
 
 1.170 
 
 1.499 
 
 1.872 
 
 1.652 
 
 57 
 
 130.4 
 
 59.7 
 
 73.2 
 
 93.7 
 
 117.1 
 
 0.986 
 
 .1.210 
 
 1.548 
 
 1.936 
 
 1.671 
 
 58 
 
 134.2 
 
 61.0 
 
 74.7 
 
 95.7 
 
 119.5 
 
 1.019 
 
 1.248 
 
 1.599 
 
 1.996 
 
 1.691 
 
 59 
 
 139.2 
 
 62.4 
 
 76.4 
 
 97.8 
 
 122.2 
 
 1.055 
 
 1.292 
 
 1.654 
 
 2.037 
 
 1.711 
 
 60 
 
 142.2 
 
 63.8 
 
 78.1 
 
 100.0 
 
 125.0 
 
 1.092 
 
 1.336 
 
 1.711 
 
 2.118 
 
 1.732 
 
 61 
 
 146.4 
 
 65.2 
 
 79.9 
 
 102.6 
 
 127.8 
 
 1.129 
 
 1.384 
 
 1.772 
 
 2.284 
 
 1.753 
 
 62 
 
 150.6 
 
 66.7 
 
 81.7 
 
 104.3 
 
 130.7 
 
 1.169 
 
 1.434 
 
 1.838 
 
 2.294 
 
 1.774 
 
 63 
 
 154.8 
 
 68.7 
 
 84.1 
 
 107.7 
 
 134.0 
 
 1.219 
 
 1.492 
 
 1.911 
 
 2.387 
 
 1.796 
 
 64 
 
 159.2 
 
 70.6 
 
 86.5 
 
 110.8 
 
 138.0 
 
 1.268 
 
 1.554 
 
 1.990 
 
 2.416 
 
 1.819 
 
 65 
 
 163.8 
 
 73.2 
 
 89.7 
 
 114.8 
 
 143.5 
 
 1.332 
 
 1.632 
 
 2.088 
 
 2.671 
 
 1.842 
 
 66 
 
 168.6 
 
 81.6 
 
 100.0 
 
 128.0 
 
 149.4 
 
 1.523 
 
 1.842 
 
 2.358 
 
 2.872 
 
 TABLE XXV. — STRENGTH OF CAUSTIC SODA SOLUTION 
 AT 15° C. (59° F.) 
 
 
 
 
 ) 
 
 
 1 cubic meter con- 
 
 Sp. Gr. 
 
 Baum6 
 
 Twaddell 
 
 Per cent 
 NajO 
 
 Per cent 
 NaOH 
 
 tains in kilos 
 
 
 
 
 
 
 
 
 
 NA 2 
 
 NaOH 
 
 1.007 
 
 1 
 
 1.4 
 
 0.47 
 
 0.61 
 
 4 
 
 6 
 
 1.014 
 
 2 
 
 2.8 
 
 0.93 
 
 1.20 
 
 9 
 
 12 
 
 1.022 
 
 3 
 
 4.4 
 
 1.55 
 
 2.00 
 
 16 
 
 21 
 
 1.029 
 
 4 
 
 5.8 
 
 2.10 
 
 2.71 
 
 22 
 
 28 
 
 1.036 
 
 5 
 
 7.2 
 
 2.60 
 
 3.35 
 
 27 
 
 35 
 
 1.045 
 
 6 
 
 9.0 
 
 3.10 
 
 4.00 
 
 32 
 
 42 
 
 1.052 
 
 7 
 
 10.4 
 
 3.60 
 
 4.64 
 
 38 
 
 49 
 
APPENDIX 
 
 229 
 
 TABLE XXV.— STRENGTH OF CAUSTIC SODA SOLUTION 
 AT 15° C. (59° F.) — Continued 
 
 
 
 
 
 
 1 cubic meter con- 
 
 Sp. Gr. 
 
 Baum6 
 
 TwaddeU 
 
 Per cent 
 NsfcO 
 
 Per cent 
 NaOH 
 
 tains in 
 
 kilos 
 
 
 
 
 NajO 
 
 NaOH 
 
 1.060 
 
 8 
 
 12.0 
 
 4.10 
 
 5.29 
 
 43 
 
 56 
 
 1.067 
 
 9 
 
 13.4 
 
 4.55 
 
 5.87 
 
 49 
 
 63 
 
 1.075 
 
 10 
 
 15.0 
 
 5.08 
 
 6.55 
 
 55 
 
 70 
 
 1.083 
 
 11 
 
 16.6 
 
 5.67 
 
 7.31 
 
 61 
 
 79 
 
 1.091 
 
 12 
 
 18.2 
 
 6.20 
 
 ■ 8.00 
 
 68 
 
 87 
 
 1.100 
 
 13 
 
 20.0 
 
 6.73 
 
 8.68 
 
 74 
 
 95 
 
 1.108 
 
 14 
 
 21.6 
 
 7.30 
 
 9.42 
 
 81 
 
 104 
 
 1.116 
 
 15 
 
 23.2 
 
 7.80 
 
 10.06 
 
 87 
 
 112 
 
 1.125 
 
 16 
 
 25.0 
 
 8.50 
 
 10.97 
 
 96 
 
 123 
 
 1.134 
 
 17 
 
 26.8 
 
 9.18 
 
 11.84 
 
 104 
 
 134 
 
 1.142 
 
 18 
 
 28.4 
 
 9.80 
 
 12.64 
 
 112 
 
 144 
 
 1.152 
 
 19 
 
 30.4 
 
 10.50 
 
 13.55 
 
 121 
 
 156 
 
 1.162 
 
 20 
 
 32.4 
 
 11.14 
 
 14.37 
 
 129 
 
 167 
 
 1.171 
 
 21 
 
 34.2 
 
 11.73 
 
 15.13 
 
 137 
 
 177 
 
 1.180 
 
 22 
 
 36.0 
 
 12.33 
 
 15.91 
 
 146 
 
 188 
 
 1.190 
 
 23 
 
 38.0 
 
 13.00 
 
 16.77 
 
 155 
 
 200 
 
 1.200 
 
 24 
 
 40.0 
 
 13.70 
 
 17.67 
 
 164 
 
 212 
 
 1.210 
 
 25 
 
 42.0 
 
 14.40 
 
 18.58 
 
 174 
 
 225 
 
 1.220 
 
 26 
 
 44.0 
 
 15.18 
 
 19.58 
 
 185 
 
 239 
 
 1.231 
 
 27 
 
 46.2 
 
 15.96 
 
 20.59 
 
 196 
 
 253 
 
 1.241 
 
 28 
 
 48.2 
 
 16.76 
 
 21.42 
 
 208 
 
 266 
 
 1.252 
 
 29 
 
 50.4 
 
 17.55 
 
 22.64 
 
 220 
 
 283 
 
 1.263 
 
 30 
 
 52.6 
 
 18.35 
 
 23.G7 
 
 232 
 
 299 
 
 1.274 
 
 31 
 
 54.8 
 
 19.23 
 
 24.81 
 
 245 
 
 316 
 
 1.285 
 
 32 
 
 57.0 
 
 20.00 
 
 25.80 
 
 257 
 
 332 
 
 1.297 
 
 33 
 
 59.4 
 
 20.80 
 
 26.83 
 
 270 
 
 348 
 
 1.308 
 
 34 
 
 61.6 
 
 21.55 
 
 27. SO 
 
 282 
 
 364 
 
 1.320 
 
 35 
 
 64.0 
 
 22.35 
 
 28.83 
 
 295 
 
 381 
 
 1.332 
 
 36 
 
 66.4 
 
 23.20 
 
 29.93 
 
 309 
 
 399 
 
 1.345 
 
 37 
 
 69.0 
 
 24.20 
 
 31.32 
 
 326 
 
 420 
 
 1.357 
 
 38 
 
 71.4 
 
 25.17 
 
 32.47 
 
 342 
 
 441 
 
 1.370 
 
 39 
 
 74.0 
 
 26.12 
 
 33.69 
 
 359 
 
 462 
 
 1.383 
 
 40 
 
 76.6 
 
 27.10 
 
 34.96 
 
 375 
 
 483 
 
 1.397 
 
 41 
 
 79.4 
 
 28.10 
 
 36.25 
 
 392 
 
 506 
 
 1.410 
 
 42 
 
 82.0 
 
 29.05 
 
 37.47 
 
 410 
 
 528 
 
 1.424 
 
 43 
 
 84.8 
 
 30.08 
 
 38.80 
 
 428 
 
 553 
 
 1.438 
 
 44 
 
 87.6 
 
 31.00 
 
 39.99 
 
 446 
 
 575 
 
 1.453 
 
 45 
 
 90.6 
 
 32.10 
 
 41.41 
 
 466 
 
 602 
 
 1.468 
 
 46 
 
 93.6 
 
 33.20 
 
 42.83 
 
 487 
 
 629 
 
 1.483 
 
 47 
 
 96.6 
 
 34.40 
 
 44.38 
 
 510 
 
 658 
 
 1.498 
 
 48 
 
 99.6 
 
 35.70 
 
 46.15 
 
 535 
 
 691 
 
 1.514 
 
 49 
 
 102.8 
 
 36.90 
 
 47.60 
 
 559 
 
 721 
 
 1.530 
 
 1 50 
 
 106.0 
 
 38.00 
 
 49.02 
 
 581 
 
 750 
 
230 
 
 APPENDIX 
 
 TABLE XXVI. — REDUCTION OF GAS. VOLUMES TO 0° 
 AND 760 MM. 
 
 Volume at 0° and 760 mm. - , ( 76Q (1 | ^ J and (P - p) 
 
 v = observed volume of gas 
 
 t = observed temperature of gas in degrees Centrigade 
 P = observed barometric pressure, corrected, in millimeters 
 p = tension of aqueous vapor in millimeters 
 The logarithm of the volume at 0° and 760 mm. is obtained by adding the 
 
 i of v and I ) and (P — p). 
 
 \760(1 +.003670/ 
 
 
 Logarithm of 
 
 Tension 
 
 
 Logarithm of 
 
 Tension 
 
 °c. 
 
 1 
 
 aqueous vapor 
 mm. 
 
 °C. 
 
 1 
 
 aqueous vapor 
 
 
 760(1 + .00367t) 
 
 760(1. +00367t) 
 
 mm. 
 
 0. 
 
 3.11919 
 
 4.60 
 
 5.8 
 
 3.11004 
 
 6.90 
 
 0.2 
 
 3.11887 
 
 4.65 
 
 6.0 
 
 3.10973 
 
 7.00 
 
 0.4 
 
 3.11855 
 
 4.71 
 
 6.2 
 
 3.10942 
 
 7.09 
 
 0.6 
 
 3.11824 
 
 4.78 
 
 6.4 
 
 3.10911 
 
 7.19 
 
 0.8 
 
 3.11792 
 
 4.85 
 
 6.6 
 
 3.10880 
 
 7.29 
 
 1.0' 
 
 3.11760 
 
 4.92 
 
 6.8 
 
 3.10848 
 
 7.39 
 
 1.2 
 
 3.11728 
 
 4.99 
 
 7.0 
 
 3.10818 
 
 7.49 
 
 1.4 
 
 3.11696 
 
 5.06 
 
 7.2 
 
 3.10786 
 
 7.60 
 
 1.6 
 
 3.11665 
 
 5.14 
 
 7.4 
 
 3.10755 
 
 7.70 
 
 1.8 
 
 3.11633 
 
 5.21 
 
 7.6 
 
 3.10724 
 
 7.81 
 
 2.0 
 
 3.11601 
 
 5.29 
 
 7.8 
 
 3.10693 
 
 7.91 
 
 2.2 
 
 3.11570 
 
 5.36 
 
 8.0 
 
 3.10662 
 
 8.02 
 
 2.4 
 
 3.11538 
 
 5.44 
 
 8.2 
 
 3.10631 
 
 8.13 
 
 2.6 
 
 3.11507 
 
 5.52 
 
 8.4 
 
 3.10600 
 
 8.24 
 
 2.8 
 
 3.11475 
 
 5.60 
 
 8.6 
 
 3.10570 
 
 8.36 
 
 3.0 
 
 3.11443 
 
 5.68 
 
 8.8 
 
 3.10538 
 
 8.47 
 
 3.2 
 
 3.11412 
 
 5.76 
 
 9.0 
 
 3.10508 
 
 8.58 
 
 3.4 
 
 3.11380 
 
 5.84 
 
 9.2 
 
 3.10477 
 
 8.70 
 
 3.6 
 
 3.11349 
 
 5.92 
 
 9.4 
 
 3.10446 
 
 8.82 
 
 3.8 
 
 3.11317 
 
 6.00 
 
 9.6 
 
 3.10415 
 
 8.94 
 
 4.0 
 
 3.11286 
 
 6.09 
 
 9.8 
 
 3.10384 
 
 9.06 
 
 4.2 
 
 3.11255 
 
 6.17 
 
 10.0 
 
 3.10354 
 
 9.18 
 
 4.4 
 
 3.11223 
 
 6.26 
 
 10.2 
 
 3.10323 
 
 9.30 
 
 4.6 
 
 3.11192 
 
 6.35 
 
 10.4 
 
 3.10292 
 
 9.43 
 
 4.8 
 
 3.11160 
 
 6.44 
 
 10.6 
 
 3.10262 
 
 9.55 
 
 5.0 
 
 3.11129 
 
 6.53 
 
 10.8 
 
 3.10231 
 
 9.68 
 
 5.2 
 
 3.11098 
 
 6.62 
 
 11.0 
 
 3.10200 
 
 9.81 
 
 5.4 
 
 3.11067 
 
 6.71 
 
 11.2 
 
 3.10170 
 
 9.94 
 
 5.6 
 
 3.11031 
 
 6.81 
 
 11.4 
 
 3.10139 
 
 10.07 
 
APPENDIX 
 
 231 
 
 TABLE XXVI. — REDUCTION OF GAS. VOLUMES TO 0° 
 AND 760 MM. — Continued 
 
 
 Logarithm of 
 
 Tension 
 
 
 Logarithm of 
 
 Tension 
 
 °c. 
 
 1 
 
 aqueous vapor 
 mm. 
 
 °C. 
 
 1 
 
 aqueous vapor 
 
 
 760(1 + .00367t) 
 
 760(1 +.00367t) 
 
 mm. 
 
 11.6 
 
 3.10108 
 
 10.21 
 
 19.4 
 
 3.08932 
 
 16.78 
 
 11.8 
 
 3.10078 
 
 10.34 
 
 19.6 
 
 3.08902 
 
 16.98 
 
 12.0 
 
 3.10047 
 
 10.48 
 
 19.8 
 
 3.08873 
 
 17.19 
 
 12.2 
 
 3.10017 
 
 10.62 
 
 20.0 
 
 3.08843 
 
 17.41 
 
 12.4 
 
 3.09986 
 
 10.76 
 
 20.2 
 
 3.08813 
 
 17.62 
 
 12.6 
 
 3.09956 
 
 10.90 
 
 20.4 
 
 3.08783 
 
 17.84 
 
 12.8 
 
 3.09925 
 
 11.04 
 
 20.6 
 
 3.08754 
 
 18.06 
 
 13.0 
 
 3.09895 
 
 11.19 
 
 20.8 
 
 3.08724 
 
 18.28 
 
 13.2 
 
 3.09864 
 
 11.33 
 
 21.0 
 
 3.08695 
 
 18.50 
 
 13.4 
 
 3.09834 
 
 11.48 
 
 21.2 
 
 3.08665 
 
 18.73 
 
 13.6 
 
 3.09804 
 
 11.63 
 
 21.4 
 
 3.08635 
 
 18.96 
 
 13.8 
 
 3.09773 
 
 11.78 
 
 21.6 
 
 3.08606 
 
 19.19 
 
 14.0 
 
 3.09743 
 
 11.94 
 
 21.8 
 
 3.08576 
 
 19.42 
 
 14.2 
 
 3.09713 
 
 12.09 
 
 22.0 
 
 3.08547 
 
 19.66 
 
 14.4 
 
 3.09682 
 
 12.25 
 
 22.2 
 
 3.08517 
 
 19.90 
 
 14.6 
 
 3.09652 
 
 12.41 
 
 22.4 
 
 3.08488 
 
 20.14 
 
 14.8 
 
 3.09622 
 
 12.57 
 
 22.6 
 
 3.08458 
 
 20.39 
 
 15.0 
 
 3.09592 
 
 12.73 
 
 22.8 
 
 3.08429 
 
 20.63 
 
 15.2 
 
 3.09561 
 
 12.89 
 
 23.0 
 
 3.08400 
 
 20.88 
 
 15.4 
 
 3.09531 
 
 13.06 
 
 23.2 
 
 3.08370 
 
 21.14 
 
 15.6 
 
 3.09501 
 
 13.23 
 
 23.4 
 
 3.08341 
 
 21.39 
 
 15.8 
 
 3.09471 
 
 13.39 
 
 23.6 
 
 3.08312 
 
 2165 
 
 16.0 
 
 3.09441 
 
 13.57 
 
 23.8 
 
 3.08282 
 
 21.91 
 
 16.2 
 
 3.09411 
 
 13.74 
 
 24.0 
 
 3.08253 
 
 22.18 
 
 16.4 
 
 3.09381 
 
 13.91 
 
 24.2 
 
 3.08224 
 
 22.45 
 
 16.6 
 
 3.09351 
 
 14.09 
 
 24.4 
 
 3.08194 
 
 22.72 
 
 16.8 
 
 3.09321 
 
 14.27 
 
 24.6 
 
 3.08165 
 
 22.99 
 
 17.0 
 
 3.09291 
 
 14.45 
 
 24.8 
 
 3.08136 
 
 23.27 
 
 17.2 
 
 3.09261 
 
 14.63 
 
 25.0 
 
 3.08107 
 
 23.55 
 
 17.4 
 
 3.09231 
 
 14.82 
 
 25.2 
 
 3.08078 
 
 23.83 
 
 17.6 
 
 3.09201 
 
 15.00 
 
 25.4 
 
 3.08048 
 
 24.11 
 
 17.8 
 
 3.09171 
 
 15.19 
 
 25.6 
 
 3.08019 
 
 24.40 
 
 18.0 
 
 3.09141 
 
 15.38 
 
 25.8 
 
 3.07990 
 
 24.69 
 
 18.2 
 
 3.09111 
 
 15.58 
 
 26.0 
 
 3.07961 
 
 24.99 
 
 18.4 
 
 3.09081 
 
 15.77 
 
 26.2 
 
 3.07932 
 
 25.28 
 
 18.6 
 
 3.09051 
 
 15.97 
 
 26.4 
 
 3.07903 
 
 25.58 
 
 18.8 
 
 3.09021 
 
 16.17 
 
 26.6 
 
 3.07874 
 
 25.89 
 
 19.0 
 
 3.08992 
 
 16.37 
 
 26.8 
 
 3.07844 
 
 26.19 
 
 19.2 
 
 3.08962 
 
 16.57 
 
 27.0 
 
 3.07816 
 
 26.50 
 
232 
 
 APPENDIX 
 
 TABLE XXVI. — REDUCTION OF GAS. VOLUMES TO 0° 
 AND 760 MM. — ContinvM 
 
 
 Logarithm of 
 
 Tension 
 
 
 Logarithm of 
 
 Tension 
 
 °c. 
 
 1 
 
 aqueous vapor 
 mm. 
 
 °C. 
 
 1 
 
 aqueous vapor 
 
 
 760(1 +.00367t) 
 
 760(1 + .00367t) 
 
 mm. 
 
 27.2 
 
 3.07787 
 
 26.82 
 
 31.2 
 
 3.07211 
 
 33.80 
 
 27.4 
 
 3.07758 
 
 27.13 
 
 31.4 
 
 3.07182 
 
 34.19 
 
 27.6 
 
 3.07729 
 
 * 27.45 
 
 31.6 
 
 3.07154 
 
 34.58 
 
 27.8 
 
 3.07700 
 
 27.78 
 
 31.8 
 
 3.07125 
 
 34.97 
 
 28.0 
 
 3.07671 
 
 28.10 
 
 32.0 
 
 3.07097 
 
 35.37 
 
 28.2 
 
 3.07642 
 
 28.43 
 
 32.2 
 
 3.07068 
 
 35.77 
 
 28.4 
 
 3.07613 
 
 28.77 
 
 32.4 
 
 3.07039 
 
 36.18 
 
 28.6 
 
 3.07584 
 
 .29.10 
 
 32.6 
 
 3.07011 
 
 36.59 
 
 28.8 
 
 3.07555 
 
 29.44 
 
 32.8 
 
 3.06983 
 
 37.01 
 
 29.0 
 
 3.07527 
 
 29.78 
 
 33.0 
 
 3.06954 
 
 37.43 
 
 29.2 
 
 3.07498 
 
 30.13 
 
 33.2 
 
 3.06926 
 
 37.85 
 
 29.4 
 
 3.07469 
 
 30.48 
 
 33.4 
 
 3.06897 
 
 38.28 
 
 29.6 
 
 3.07440 
 
 30.84 
 
 33.6 
 
 3.06869 
 
 38.71 
 
 29.8 
 
 3.07411 
 
 31.19 
 
 33.8 
 
 3.06841 
 
 39.15 
 
 30.0 
 
 3.07383 
 
 31.56 
 
 34.0 
 
 3.06812 
 
 39.59 
 
 30.2 
 
 3.07354 
 
 31.92 
 
 34.2 
 
 3.06784 
 
 40.03 
 
 30.4 
 
 3.07325 
 
 32.29 
 
 34.4 
 
 3.06756 
 
 40.48 
 
 30.6 
 
 3.07297 
 
 32.66 
 
 34.6 
 
 3.06727 
 
 40.93 
 
 30.8 
 
 3.07268 
 
 33.04 
 
 34.8 
 
 3.06699 
 
 41.39 
 
 31.0 
 
 3.07239 
 
 33.42 
 
 35.0 
 
 3.06671 
 
 41.85 
 
 TABLE XXVII. 
 
 ■FERROUS SULPHATE AT 15° 
 Gerlach 
 
 Specific 
 gravity 
 
 Per 
 
 cent 
 
 FeS0 4 
 
 Per 
 
 cent 
 
 FeS0 4 
 
 .7H 2 
 
 Specific 
 gravity 
 
 Per 
 
 cent 
 
 FeS0 4 
 
 Per 
 
 cent 
 FeSO., 
 .7H 2 
 
 Specific 
 gravity 
 
 Per 
 
 cent 
 FeS0 4 
 
 Per 
 
 cent 
 
 FeSO, 
 
 .7H 2 
 
 1.005 
 1.011 
 1.016 
 1.021 
 
 0.565 
 1.130 
 1.694 
 2.258 
 
 1 
 2 
 3 
 4 
 
 1.0267 
 1.0537 
 1.0823 
 1.1124 
 
 2.811 
 
 5.784 
 
 8.934 
 
 12.277 
 
 5 
 
 10 
 15 
 20 
 
 1.1430 
 1.1738 
 1.2063 
 1.2391 
 
 15.834 
 19.622 
 23.672 
 27.995 
 
 25 
 30 
 35 
 40 
 
APPENDIX 
 
 233 
 
 TABLE XXVIII. — AQUA AMMONIA 
 W. C. Ferguson 
 
 Degrees 
 
 Sp. Gr. 
 60° 
 
 Per cent 
 
 Degrees 
 
 Sp. Gr. 
 60° 
 
 Per cent 
 
 Baum6 
 
 — F. 
 60° 
 
 NH, 
 
 Baum6 
 
 — F 
 60° 
 
 NH, 
 
 10.00 
 
 1.0000 
 
 .00 
 
 19.75 
 
 .9349 
 
 17.28 
 
 10.25 
 
 .9982 
 
 .40 
 
 20.00 
 
 .9333 
 
 17.76 
 
 10.50 
 
 .9964 
 
 .80 
 
 20.25 
 
 .9318 
 
 18.24 
 
 10.75 
 
 .9947 
 
 1.21 
 
 20.50 
 
 .9302 
 
 18.72 
 
 11.00 
 
 .9929 
 
 1.62 
 
 20.75 
 
 .9287 
 
 19.20 
 
 11.25 
 
 .9912 
 
 2.04 
 
 21.00 
 
 .9272 
 
 19.68 
 
 11.50 
 
 .9894 
 
 2.46 
 
 21.25 
 
 .9256 
 
 20.16 
 
 11.75 
 
 .9876 
 
 2.88 
 
 21.50 
 
 .9241 
 
 20.64 
 
 12.00 
 
 .9859 
 
 3.30 
 
 21.75 
 
 .9226 
 
 21.12 
 
 12.25 
 
 .9842 
 
 3.73 
 
 22.00 
 
 .9211 
 
 21,60 
 
 12.50 
 
 .9825 
 
 4.16 
 
 22.25 
 
 .9195 
 
 22.08 
 
 12.75 
 
 .9807 
 
 4.59 
 
 22.50 
 
 .9180 
 
 22.56 
 
 13.00 
 
 .9790 
 
 5.02 
 
 22.75 
 
 .9165 
 
 23.04 
 
 13.25 
 
 .9773 
 
 5.45 
 
 23.00 
 
 .9150 
 
 23.52 
 
 13.50 
 
 .9756 
 
 5.88 
 
 23.25 
 
 .9135 
 
 24.01 
 
 13.75 
 
 .9739 
 
 6.31 
 
 23.50 
 
 .9121 
 
 24.50 
 
 14.00 
 
 .9722 
 
 6.74 
 
 23.75 
 
 .9106 
 
 24.99 
 
 14.25 
 
 .9705 
 
 7.17 
 
 24.00 
 
 .9091 
 
 25.48 
 
 14.50 
 
 .9689 
 
 7.61 
 
 24.25 
 
 .9076 
 
 25.97 
 
 14.75 
 
 .9672 
 
 8.05 
 
 24.50 
 
 .9061 
 
 26.46 
 
 15.00 
 
 .9655 
 
 8.49 
 
 24.75 
 
 .9047 
 
 26.95 
 
 15.25 
 
 .9639 
 
 8.93 
 
 25.00 
 
 .9032 
 
 27.44 
 
 15.50 
 
 .9622 
 
 9.38 
 
 25.25 
 
 .9018 
 
 27.93 
 
 15.75 
 
 .9605 
 
 9.83 
 
 25.50 
 
 .9003 
 
 28.42 
 
 16.00 
 
 .9589 
 
 10.28 
 
 25.75 
 
 .8989 
 
 28.91 
 
 16.25 
 
 .9573 
 
 10.73 
 
 26.00 
 
 .8974 
 
 29.40 
 
 16.50 
 
 .9556 
 
 11.18 
 
 26.25 
 
 .8960 
 
 29.89 
 
 16.75 
 
 .9540 
 
 11.64 
 
 26.50 
 
 .8946 
 
 30.38 
 
 17.00 
 
 .9524 
 
 12.10 
 
 26.75 
 
 .8931 
 
 30.87 
 
 17.25 
 
 .9508 
 
 12.56 
 
 27.00 
 
 .8917 
 
 31.36 
 
 17.50 
 
 .9492 
 
 13.02 
 
 27.25 
 
 .8903 
 
 31.85 
 
 17.75 
 
 .9475 
 
 13.49 
 
 27.50 
 
 .8889 
 
 32.34 
 
 18.00 
 
 .9459 
 
 13.96 
 
 27.75 
 
 .8875 
 
 32.83 
 
 18.25 
 
 .9444 
 
 14.43 
 
 28.00 
 
 .8861 
 
 33.32 
 
 18.50 
 
 .9428 
 
 14.90 
 
 28.25 
 
 .8847 
 
 33.81 
 
 18.75 
 
 .9412 
 
 15.37 
 
 28.50 
 
 .8833 
 
 34.30 
 
 19.00 
 
 .9396 
 
 15.84 
 
 28.75 
 
 .8819 
 
 34.79 
 
 19.25 
 
 .9380 
 
 16.32 
 
 29.00 
 
 .8805 
 
 35.28 
 
 19.50 
 
 .9365 
 
 16.80 
 
 
 
 
234 
 
 APPENDIX 
 
 Specific Gravity determinations were made at 60° F., compared with 
 water at 60° F. 
 
 From the Specific Gravities .the corresponding degrees Baume' were calcu- 
 lated by the following formula: 
 
 140 
 
 Baume = 
 
 Sp. Gr. 
 
 130. 
 
 Baume Hydrometers for use with this table must be graduated by the above 
 formula, which formula should always be printed on the scale. 
 Atomic weights from F. W. Clarke's table of 1901. O = 16. 
 
 TABLE XXIX. — ALLOWANCE FOR TEMPERATURE 
 
 The coefficient of expansion correction for ammonia solutions, varying with 
 the temperature, must be applied according to the following table: 
 
 Corrections to be added for 
 
 Corrections to be subtracted for each 
 
 each degree below 60° F. 
 
 degree above 60° F. 
 
 Degrees 
 Baume 1 
 
 40° F. 
 
 50° F. 
 
 70° F. 
 
 80° F. 
 
 90° F. 
 
 100° F. 
 
 14° B<5 
 
 .015° B6 
 
 .017° B6 
 
 .020° Be 
 
 .022° Be 
 
 .024° Be 
 
 .026° B6 
 
 16 " 
 
 .021 " 
 
 .023 " 
 
 .026 " 
 
 .028 " 
 
 .030 " 
 
 .032 " 
 
 18 " 
 
 .027 " 
 
 .029 " 
 
 .031 " 
 
 .033 " 
 
 .035 " 
 
 .037 " 
 
 20 " 
 
 .033 " 
 
 .036 " 
 
 .037 " 
 
 .038 " 
 
 .040 " 
 
 .042 " 
 
 22 " 
 
 .039 " 
 
 .042 " 
 
 .043 " 
 
 .045 " 
 
 .047 " 
 
 
 26 " 
 
 .053 " 
 
 .057 " 
 
 .057 " 
 
 .059 " 
 
 
 
'he Temperature of Products of Combustion is reduced to 18" 
 
 C. = 64.4° P. 
 
 
 XX 
 
 XXI 
 
 XXII 
 
 XXIII 
 
 XXIV 
 
 XXV 
 
 XXVI 
 
 ' pound of combustible 
 
 Heat of formation at const, pres. 
 
 
 Products of 
 
 Calories 
 
 B. T. U. 
 
 Name of 
 
 combustion 
 
 per molec- 
 ular wt. in 
 
 
 gas or vapor 
 
 
 
 
 
 
 
 COi 
 
 HiO 
 
 
 grammes 
 
 Per cu. ft. 
 
 Per pound 
 
 
 
 
 CO-2.333+ 
 
 
 
 
 Carbon, to CO 
 
 3.066 + 
 
 
 
 
 
 
 
 Carbon, to COi 
 
 1.571 
 
 9.000 
 
 
 
 
 + 138.4 
 
 + 1869.2 
 
 Carbonic oxide 
 Hydrogen 
 
 2.750 
 
 2.250 
 
 
 
 + 21,750 
 
 + 103.1 
 
 + 2435.6 
 
 Methane 
 
 2.933 
 
 1.800 
 
 
 
 + 28,560 
 
 + 136.0 
 
 + 1713.6 
 
 Ethane 
 
 3.000 
 
 1.636 
 
 
 
 + 35,110 
 
 + 167.2 
 
 + 1436.3 
 
 Propane 
 
 3.034 
 
 1.552 
 
 
 
 + 42,450 
 
 + 202.2 
 
 + 1317.3 
 
 Butane 
 
 3.055 
 
 1.500 
 
 
 
 + 47,850 
 
 + 227.9 
 
 + 1196.2 
 
 Pentane 
 
 3.069 
 
 1.465 
 
 
 
 + 61,080 
 
 + 290.9 
 
 + 1278.4 
 
 Hexane 
 
 3.142 
 
 1.286 
 
 
 
 - 2,710 
 
 - 12.9 
 
 - 174.2 
 
 Ethylene 
 
 3.142 
 
 1.286 
 
 
 
 + 3,220 
 
 + 15.3 
 
 + 138.0 
 
 Propylene 
 
 3.142 
 
 1.286 
 
 
 
 + 10,660 
 
 + 50.7 
 
 + 342.6 
 
 Butylene 
 
 3.142 
 
 1.286 
 
 
 
 + 18,970 
 
 + 113.7 
 
 + 614.1 
 
 Amylene 
 
 3.384 
 
 0.692 
 
 
 
 - 47,770 
 
 - 227.5 
 
 - 3300.7 
 
 Acetylene 
 
 3.300 
 
 0.900 
 
 
 
 - 39,650 
 
 - 188.8 
 
 - 1784.2 
 
 Allylene 
 
 3.259 
 
 1.300 
 
 
 
 
 
 
 Crotonylene 
 
 3.384 
 
 0.692 
 
 
 
 - 12,610 
 
 - 47.3 
 
 - 229.3 
 
 Benzene 
 
 3.348 
 
 0.782 
 
 
 
 - 3,520 
 
 - 16.7 
 
 - 68.8 
 
 Toluene 
 
 3.311 
 
 0.849 
 
 
 
 
 
 
 Xylene 
 
 3.300 
 
 0.900 
 
 
 
 + 490 
 
 + 2.3 
 
 + 7.3 
 
 Mesitylene 
 
 3.437 
 
 0.563 
 
 
 
 
 
 
 Naphthalene 
 
 
 0.529 
 
 SOa- 1.883 
 
 + 4,740 
 
 + 22.6 
 
 + 250.9 
 
 Hydrogen sulphide 
 
 
 1.588 
 
 N - 0.823 
 
 + 11,890 
 
 + 56.7 
 
 + 1259.0 
 
 Ammonia 
 
 1.630 
 
 0.333 
 
 N - 0.518 
 
 - 27,480 
 
 - 131.1 
 
 - 1832.0 
 
 Hydrocyanic acid 
 
 1.692 
 
 
 N - 0.538 
 
 - 65,700 
 
 - 313.2 
 
 - 2273.9 
 
 Cyanogen 
 
 0.579 
 
 
 SO>- 1.684 
 
 - 26,010 
 
 - 124.0 
 
 - 616.0 
 
 Carbon bi-sulphide 
 
 1.375 
 
 1.125 
 
 
 + 51,450 
 
 + 246.4 
 
 + 2894.0 
 
 Methyl alcohol 
 
 1.913 
 
 1.174 
 
 
 
 + 58,470 
 
 + 278.5 
 + 463.1 
 + 327.1 
 + 337.3 
 
 + 2288.0 
 + 3979.1 
 + 6870.4 
 + 1999.1 
 
 Ethyl alcohol 
 
 Carbonic acid 
 
 Water 
 
 Sulphur dioxide 
 
 Oxygen 
 
 Nitrogen 
 
 Air 
 
 Column VII i 
 
 9 from Ganot's "Physics," edition 
 
 1896, page 445. 
 
 olumns X 
 
 and XXIII are from Julius Tl 
 
 lomsen's "Thermo- 
 
 nicul Invest 
 
 igations," and his results are tra 
 
 Dslated into English 
 
 s in column 
 
 s XI-XII and XXIV-XXV. 
 
 
 'olumns XI 
 
 I and XVIII are calculated on t 
 
 he assumption that 
 
 air = 2 
 
 0.9% oxygen + 79.1% nitrogen 
 
 by volume. 
 
 air = '. 
 
 3.13% oxygen + 76.87% nitroge 
 
 i by weight. 
 
 
 
 
 
 
 
 (Feu 
 
 ring Page 2U) 
 
Missing Page 
 
APPENDIX 
 
 235 
 
 a 
 
 1 
 
 Is 
 as 
 
 U 
 
 o o 
 
 h 
 
 © 
 
 ■-" » — -* h- 
 
 O W H >* I- H O 
 
 ■-< oo oo >~ cc i< od 
 
 05 ■* 
 
 Ol ffl »f N 
 COOiOH 
 
 dodo 
 
 oo £■ 
 os o 
 oo -i< 
 
 CM. 
 
 O 115 ^ N M -f 00 
 
 H (O (6 d H »t pi 
 
 y-4 »-H CI i-H CO <N 
 
 O '.O co O 
 
 to n n ^* 
 
 hoho 
 
 13 
 
 I' 
 
 TS'o io 
 
 <N O O CO CO 
 
 .-< O ~V CD CO CN ■* 
 
 i-H 00 t^ O O CO *H 
 
 a lO (N » N O 
 
 CN ^ 
 
 O -t< 
 d CO 
 
 ooffig 
 
 CD a £ 
 
 3- 
 
 T3 
 O 
 O 
 
 be 
 
 OS s 
 
 H 
 
 •3.2 
 
 O "5 
 r^ CO CJ 
 
 CO 
 
 OS CO CO 
 
 (N °. 1 
 
 00 o 
 CN d 
 
 CO ^ 
 ■>)< cN 
 
 "5 CN 
 
 O s 
 
 W 
 
 o <u 
 
 ■3 
 
 Oh h -t nc5h 
 
 OS OS 
 
 H (O IO H 113 !0 H 
 CO CO I- CO i-H ■* 
 
 -H O 
 
 >- CN O 
 
 CN 00 
 
 on 
 
 b/l 
 
 OS 
 
 
 
 >, 
 
 
 
 
 
 
 > 
 
 
 fc. 
 
 
 o 
 
 
 o 
 
 _ 
 
 a 
 
 CO 
 
 fc 
 
 Ej 
 
 CO net-J 
 
 O cN os r- co co -h 
 
 H !C 113 ^i N 00 N 
 
 »-H i— ( CO »-H ^ »-l 
 
 CO CO O 
 ^ O 1^ 
 
 H H 
 
 1 °°. 
 
 CO 
 
 1 .2 
 
 a °S3 
 
 tH OS "^ OS CO 
 
 OODJNHNH 
 
 h loci to ri nh 
 
 t^ t^ os co 
 
 H M ■* rt 
 
 -h d d d 
 
 h- 6? 
 
 a 
 & 
 
 a « 
 
 OSjaX ,-* O O »0 »0 
 
 >,-aii-idcNiocNodcN 
 
 i-H o o o 
 
 tOWH 
 
 i-i oci 
 
 CN 
 
 OS ,;, 
 
 a *> 2 a 
 
 £ a s'~ 
 
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236 
 
 APPENDIX 
 
 TABLE XXXII. — COMPOSITION OF THE PRINCIPAL 
 FRACTIONS OF COAL TAR 
 
 Crude Naphtha. Light 
 
 Oils 
 (Sp./Gr. less than 1.000) 
 (Distilling below 180° C.) 
 
 Methanes (Paraffins) 
 Tetrane to Decane 
 
 Ethenes (Olefines) 
 Pentene to Heptene 
 
 Ethines (Acetylenes) 
 
 Benzenes 
 Benzene to Durene 
 
 Benzene Hexhydrides 
 
 Naphthalene Hydrides 
 
 Phenol (Carbolic Acid) 
 
 Methylphenols (Cresols) 
 
 Pyridines 
 
 Sulphur Comp'ds (Thio- 
 phenes) 
 
 Amine Derivatives 
 
 Dead Oils. Creosote 
 
 Oils 
 (Sp./Gr. more than 
 1.000) 
 (Distilling between 
 
 180° — 270° C.) 
 
 Phenol 
 
 Methylphenols 
 
 Naphthalenes 
 Naphthalene Hy- 
 drides 
 Anilines 
 Leucolines 
 
 Pyridines 
 
 Quinolines 
 
 Green or Anthracene 
 Oils 
 
 (Distilling above 
 270° C.) 
 
 Methyl-naphthalene 
 Phenanthrene 
 
 Anthracene 
 
 Pyrene 
 Chrysene 
 
 Benzerythrene 
 Solid Paraffin 
 Benzonitrile 
 Carbazol 
 Acridine 
 
 Pitch 
 
INDEX 
 
 A 
 
 Absorption by anthracene oils. . . 58 
 
 by creosote oils 58 
 
 by water gas tar 58 
 
 of ammonia 89 
 
 Acid, nitric, conversion of am- 
 monia into 154 
 
 sulphuric 191 
 
 Advantages of cyanogen removal 67 
 
 Koppers sulphate process ... 115 
 
 Alkylation process 44 
 
 Allowance for temperature 234 
 
 American Feld sulphate plant . . . 148 
 American Gas Institute report 
 
 on naphthalene 54 
 
 Ammonia 89 
 
 absorption 89 
 
 anhydrous 93 
 
 aqua 92 
 
 concentrated 92-95 
 
 content 143 
 
 conversion of cyanogen into . 71 
 conversion of, into nitric 
 
 acid 154 
 
 free and fixed 94 
 
 in condensate 137 
 
 in gas 137-141 
 
 lost in direct sulphatation ... 150 
 
 precipitation 90 
 
 production 91 
 
 reactions, Feld 133 
 
 removal 4 
 
 removed in cyanogen extrac- 
 tion 143 
 
 required to precipitate iron . 147 
 
 still 94 
 
 synthetic 110 
 
 table of strength of 90 
 
 test 205 
 
 Ammonia testing apparatus .... 205 
 
 Ammonium carbonate 93 
 
 chloride 93 
 
 nitrate 93 
 
 sulphate 92 
 
 sulphate, specific gravity of . 224 
 
 Analysis of gas liquor, Table 
 
 XXXI 235 
 
 Anhydrous ammonia 93 
 
 Anthracene oil as absorber 58 
 
 oils 18 
 
 Apparatus for ammonia test 205 
 
 Feld's "blue" test 213 
 
 naphthalene test 204 
 
 tar test 200-202 
 
 Aqua ammonia 92-101-224 
 
 table of 233 
 
 B 
 
 Bayer's experiments 45 
 
 Benzol 164 
 
 and its homologues 166 
 
 apparatus, materials of con- 
 struction 180 
 
 boiling point of 165 
 
 extraction, loss in heat value 188 
 extraction, loss in candle 
 
 power 190 
 
 extraction, wash oil required 
 
 for 187 
 
 Feld process 177 
 
 in Germany 166 
 
 Koppers process 169 
 
 plant, Hirzel 182 
 
 plant, location of 185 
 
 prices 168 
 
 production 165 
 
 recovery, earnings 166 
 
 recovery, operating cost .... 167 
 
 specific gravity of 165 
 
238 
 
 INDEX 
 
 Benzol, specifications 178, 222 
 
 still, Gasser 185 
 
 still, Hirzel 181 
 
 test 214 
 
 value 165 
 
 washers 180 
 
 Black salt 83 
 
 Boiling constituents, high 28 
 
 low 28 
 
 Boiling points, benzol 165 
 
 light oil 172 
 
 naphthalene 52 
 
 pure benzol 176 
 
 pure toluol 165-176 
 
 solvent naphtha 165 
 
 xylol 165 
 
 British Cyanides Co. process ... 69 
 
 Brown coal products 30 
 
 Brunck sulphate process 113 
 
 Bueb cyanogen plant 73 
 
 plant, cost of operation .... 76 
 
 process, chemistry of 74 
 
 process of cyanogen extrac- 
 tion 69 
 
 product 75 
 
 Bueb revenue 76 
 
 Bunte, experiments on tar pro- 
 duction 7 
 
 Calcium cyanamide 93 
 
 Caloric compensating point 19 
 
 Candle power loss, due to benzol 
 
 extraction 190 
 
 Carboxylation process 44 
 
 Caustic ammonia, specific gravity 
 
 of 225 
 
 soda, strength of 228 
 
 Chamber acid process 191 
 
 Changes in condition of gas .... 19 
 
 Chemistry of Bueb process 74 
 
 Feld cyanogen process 76 
 
 prussiate of potash process . 82 
 
 Chili saltpeter 109 
 
 Chlorination process 43 
 
 Circulating wash tar 25 
 
 • Coal gas tar 5, 7 
 
 Collin sulphate process 116 
 
 Color of salt 108 
 
 Combined cyanogen and sul- 
 phate process 136 
 
 Compensating point, caloric 19 
 
 Composition of free carbon 8 
 
 gas 2 
 
 tar 6-8 
 
 tar, fractions, Table 
 
 XXXII 236 
 
 weak liquor 93 
 
 Concentrated ammonia . . 92-95, 100 
 
 Concentrated earnings 100 
 
 Conclusions covering sulphate 
 
 production 150 
 
 Condensation process 44 
 
 theory of 4 
 
 Condensing system, Feld 2 
 
 usual ' 2 
 
 Condition of gas 19 
 
 Consumption of artificial dyes ... 41 
 
 of sulphate Ill 
 
 Contact acid process 191 
 
 Continuous distillation of tar. ... 40 
 Conversion of ammonia into 
 
 nitric acid 154 
 
 cyanogen into ammonia. ... 71 
 
 Cooling the gas 21 
 
 Coppee sulphate process 121 
 
 Corrosion due to cyanogen 67 
 
 Cost of concentrating 100 
 
 operation, aqua ammonia . . . 103 
 operation, indirect sulphate. 107 
 
 producing sulphate 153 
 
 Coupling and diazotizing ....... 44 
 
 Creosote oil 17 
 
 as absorbent 58 
 
 Crude tar substances 42 
 
 Cyanogen 66-141 
 
 a corrosive 67 
 
 and sulphate earnings 140 
 
 from spent oxide 84 
 
 and sulphate operating 
 
 costs 146 
 
 in gas 137 
 
 in purifiers 67 
 
 plant, Bueb, operating cost . 76 
 process, British Cyanides 
 Co 69 
 
INDEX 
 
 239 
 
 Cyanogen process, Feld 76 
 
 process, Davis-Neill 80 
 
 removal, advantages of ... . 67 
 
 revenue from 68-76 
 
 test 207 
 
 D 
 
 Dehydration of tar 35 
 
 Dehydrated tar, value of 16 
 
 Deposition of naphthalene 52 
 
 Description of Bueb cyanogen 
 
 plant 73 
 
 Feld cyanogen plant 76 
 
 Deutz naphthalene engine 63 
 
 Davis-Neill cyanogen process. . 80 
 
 Dew points 21 
 
 Diagram A, Feld condensing 
 
 system 3 
 
 I, Tar fractions 10 
 
 II, Price of sulphate 112 
 
 III, Temperatures of sul- 
 phatation 155 
 
 IV, Benzol prices 168 
 
 V, Boiling point, light oil . . 172 
 
 VI, Boiling point, pure ben- 
 zol 176 
 
 VII, Boiling point, pure 
 toluol 176 
 
 Diazotizing and coupling 41 
 
 Direct sulphatation, ammonia 
 
 lost in 151 
 
 sulphate process 113 
 
 Distillation of light oils 28 
 
 tar 37^0 
 
 Distribution of nitrogen, Table 
 
 XI 67 
 
 Drugs, synthetic 49 
 
 Dye, Indigo 45 
 
 E 
 
 Earnings, aqua ammonia 102 
 
 benzol 166 
 
 concentrated ammonia 100 
 
 cyanogen and sulphate 140 
 
 indirect sulphate 107 
 
 Efficiency of purifiers 140 
 
 vertical washers 62 
 
 Electrical precipitation of tar. . . 33 
 
 Feld, American, sulphate plant . . 148 
 
 condensing system 3 
 
 cyanogen plant, descrip- 
 tion of 76 
 
 cyanogen process 76 
 
 hot tar washing 26 
 
 iron reactions 132 
 
 plants under construction. . . 148 
 
 process, chemistry of 78 
 
 process, liquor in circulation 145 
 
 process, oxidation 147 
 
 process, products from 78 
 
 process, sulphur required. . . 146 
 
 sulphate process 129 
 
 system of fractional separa- 
 tion 17 
 
 tar extracting system 21 
 
 tar test 201 
 
 test for prussian blue 208 
 
 vertical washer 60 
 
 zinc reactions 131 
 
 Feld's ammonia reactions 133 
 
 benzol process 177 
 
 early experiments 130 
 
 Ferrous sulphate 138 
 
 Figure 1, P. and A. condenser. . . 11 
 
 2, Injector tar extractor. ... 12 
 
 3, Cyclone tar extractor 12 
 
 4, Centrifugal tar extractor. 13 
 
 5, Rotary tar extractor 13 
 
 6, Static centrifugal ex- 
 tractor 14 
 
 7, Tar dehydrating plant ... 16 
 
 8, Feld tar washer 25 
 
 9, Static tar extractor 33 
 
 9a, Vertical tar still 36 
 
 9b, Horizontal tar still 37 
 
 9c, Phenol diagram 50 
 
 10, Oil regeneration plant . . 58 
 
 11, Horizontal "Standard" 
 washer 59 
 
 12, Feld vertical washer. . . 60 
 
 13, Vertical "Standard" 
 washer 61 
 
240 
 
 INDEX 
 
 Figure 14, Bueb cyanogen plant. 74 
 
 15, Feld cyanogen plant .... 77 
 
 16, Potassium ferro-cyanide 
 plant, Bueb 82 
 
 17, Potassium ferro-cyanide 
 plant, Feld 84 
 
 18, Ammonia still 95 
 
 19, Concentrating plant .... 96 
 
 20, Aqua plant 102 
 
 21, Indirect sulphate plant . . 104 
 21a, Walker sulphate plant. 106 
 
 22, Koppers sulphate plant . 115 
 
 23, Collin sulphate plant 116 
 
 23a, Collin sulphate plant. . 118 
 
 24, Still sulphate plant 119 
 
 25, Coppee sulphate plant . . 121 
 
 26, Mallet sulphate plant ... 122 
 
 27, Mallet sulphate plant ... 123 
 
 28, Mallet sulphate plant ... 123 
 
 29, Mallet sulphate plant ... 123 
 
 30, Mont-Cenis sulphate 
 plant 124 
 
 31, Otto sulphate plant 126 
 
 32, Otto sulphate plant 127 
 
 33, Strommenger sulphate 
 plant 128 
 
 34, Feld sulphate plant 132 
 
 35, Bueb cyanogen and Feld 
 sulphate plant 136 
 
 35a, American Feld plant . . 148 
 35b, Ostwald process plant. 158 
 
 36, Koppers benzol plant. . . 169 
 
 37, Feld benzol plant 178 
 
 38, Hirzel benzol plant 183 
 
 38a, Benzol plant 186 
 
 38b, Sulphuric acid plant. . . 194 
 
 39, Tar testing apparatus . . . 200 
 
 40, Feld tar testing apparatus 202 
 
 41, Naphthalene testing 
 apparatus 204 
 
 42, Ammonia testing appa- 
 ratus 205 
 
 43, Ammonia testing appa- 
 ratus 206 
 
 44, " Blue " Test 213 
 
 44a, Apparatus for live 
 
 steam distillation 215 
 
 Figure 44b, Calcium chloride 
 
 dryer 216 
 
 44c, Fractional distillation 
 
 column 216 
 
 Fixed ammonia 94 
 
 Flash point of benzol 165 
 
 Foreign matter, influence of 108 
 
 Formation of naphthalene 52 
 
 Free ammonia 94 
 
 carbon, composition of 8 
 
 carbon in Feld tar 18 
 
 Freezing point of toluol 165 
 
 Fusion process 43 
 
 G 
 
 Gas, composition of 2 
 
 constants, Table XXX, fac- 
 ing page 235 
 
 from first Feld washer 22 
 
 from second Feld washer. . . 23 
 liquor analysis, Table 
 
 XXXI 235 
 
 treatment of 2 
 
 volume, reduction of 230 
 
 Gasser benzol still 185 
 
 H 
 
 Heavy oil, removal of 4 
 
 oils 18 
 
 High boiling constituents 28 
 
 temperature tar 8 
 
 Hirzel benzol plant 182 
 
 process 182 
 
 still 181 
 
 system of tar distillation .... 40 
 
 Horizontal "Standard" washer. . 59 
 
 tar still 37 
 
 Hot tar washing 26-27 
 
 results, Bohemia 29 
 
 results, Ilsider 31 
 
 results, Moncieu 32 
 
 results, Silesia 29 
 
 Hydrogen sulphide, removal of . . 4 
 
 removed 139 
 
 I 
 
 Ilsider results 31 
 
 Income, benzol recovery 166 
 
INDEX 
 
 241 
 
 Indigo dye 45 
 
 Indirect sulphate, cost of opera- 
 tion 107 
 
 earnings 107 
 
 plant 104 
 
 process 103 
 
 Influence of foreign matter on 
 
 sulphate 108 
 
 nature of coal on tar 7 
 
 oxygen on tar production ... 7 
 
 temperature on tar 7 
 
 Iron reactions, Feld 132 
 
 sulphate required 138 
 
 K 
 
 Koppers benzol plant 167-169 
 
 sulphate plant 114 
 
 sulphate plant, advantages of 116 
 
 Kubierschky and Bormann sys- 
 tem of tar distillation .... 40 
 
 Light oil, boiling point 172 
 
 oil, removal 4 
 
 oils 4,17 
 
 Lime required 99 
 
 Liming process 44 
 
 Liquor in circulation, Feld 
 
 process 145 
 
 strength 137 
 
 Location of benzol plant 185 
 
 Loss in candle power of gas 190 
 
 in heat value of gas 188 
 
 Low boiling constituents 28 
 
 M 
 
 Mallet sulphate process 122 
 
 Manufacture of picric acid 47 
 
 synthetic drugs 49 
 
 synthetic perfumes 47 
 
 synthetic phenol 49 
 
 trinitrotoluene 47 
 
 Materials used, benzol apparatus 180 
 
 Melting point of naphthalene 52 
 
 Middle oil removal 4 
 
 Moncieu results, hot tar washing 32 
 
 Mont-Cenis sulphate process .... 124 
 
 N 
 
 Naphthalene 56 
 
 deposition of 52 
 
 engine, Deutz 63 
 
 extraction 62 
 
 for engine use 63 
 
 formation of 52 
 
 removal of 4, 62 
 
 report, A. G. 1 54 
 
 saturation 53 
 
 solvents 56 
 
 synthesis of 52 
 
 test 203 
 
 testing apparatus 204 
 
 Nature of tar products 18 
 
 Nitration process 43 
 
 Nitric acid, conversion of am- 
 monia into 154 
 
 from coal gas 161 
 
 O 
 
 Oils, anthracene 18 
 
 creosote 17 
 
 heavy 18 
 
 light 17 
 
 medium 17 
 
 regeneration plant 58 
 
 Operating cost, aqua ammonia . . 103 
 
 benzol plant 167 
 
 Bueb cyanogen and Feld 
 
 sulphate 140 
 
 Bueb cyanogen plant 76 
 
 concentrated ammonia 100 
 
 indirect sulphate 107 
 
 Ostwald process 154 
 
 Otto sulphate process 125 
 
 Oxidation process 43 
 
 Oxidizing agent, Feld process. . . 147 
 
 Paris Gas Company's experi- 
 ments 7 
 
 P. and A. condenser 11 
 
 Phenol, synthetic 49 
 
 Picric acid manufacture 47 
 
 Pitch removal 4 
 
 Polythionate liquor 145 
 
242 
 
 INDEX 
 
 Potassium ferro-cyanide 66 
 
 chemistry of 82 
 
 plant, Bueb 81 
 
 plant, Feld 83 
 
 Precipitation of ammonia 90 
 
 Principal residuals 1 
 
 Process, British Cyanides Co. ... 69 
 
 Brunck 113 
 
 Collin 116 
 
 Coppee 121 
 
 direct sulphate 113 
 
 Feld 129-177 
 
 Hirzel 182 
 
 Koppers 114-169 
 
 Mallet 122 
 
 Mont-Cenis 124 
 
 Otto 125 
 
 potassium ferro-cyanide ... 81 
 
 semi-direct sulphate 113 
 
 Still 118 
 
 Strommenger 127 
 
 Product, Feld process 78 
 
 Products, tar 35 
 
 Production of ammonia 91 
 
 benzol 165 
 
 Prussian blue 138 
 
 sulphate in TJ. S Ill 
 
 Products of distillation 9-30 
 
 tar 5 
 
 Prussian blue produced 138 
 
 test, comparison 210 
 
 test, Feld 208 
 
 Pure benzol boiling point 176 
 
 toluol boiling point 176 
 
 R 
 Raschig system of tar distilla- 
 tion '. 41 
 
 Ratio of polythionate to thiosul- 
 
 phate 146 
 
 Recovery of benzol and its homo- 
 
 logues 166 
 
 residuals 1 
 
 Reduction of gas volume 230 
 
 process 43 
 
 Removal of ammonia 4 
 
 hydrogen sulphide 4, 139 
 
 light oil 4 
 
 Removal of naphthalene 62 
 
 tar 10 
 
 Residual recovery 1 
 
 Results of Feld separation 31 
 
 hot water washing 27 
 
 Revenue from cyanogen 68 
 
 S 
 Sadewasser system of tar dis- 
 tillation 40 
 
 Sal ammoniac 93 
 
 specific gravity of 224 
 
 Selling prices, sulphate, Diagram 
 
 II 113 
 
 Semi-direct process 113 
 
 Solvent naphtha 165 
 
 naphthalene 56 
 
 naphtha specifications 222 
 
 naphtha tests 214 
 
 Specific gravity, ammonium sul- 
 phate 224 
 
 benzol 165 
 
 caustic ammonia 225 
 
 sal ammoniac 224 
 
 solvent naphtha 165 
 
 sulphuric acid 227 
 
 toluol 165 
 
 xylol 165 
 
 Specifications for benzol 222 
 
 solvent naphtha 222 
 
 toluol 222 
 
 Spent oxide, cyanogen from .... 84 
 
 sulphuric acid from 191 
 
 Static tar extractor 33 
 
 Steam required in concentrating . 99 
 
 Still, ammonia 94 
 
 Gasser 185 
 
 Hirzel 181 
 
 tar, horizontal 37 
 
 tar, vertical 36 
 
 Still sulphate process 118 
 
 Stoppages in pipe 55 
 
 Strength of caustic soda 228 
 
 liquor 137 
 
 Strommenger process 127 
 
 Sulphate, ammonia 92 
 
 Brunck process 113 
 
 Collin process 116 
 
INDEX 
 
 243 
 
 Sulphate Coppee proems . . 121 
 
 Feld process 129 
 
 from cyanogen sludge 144 
 
 in Japan 112 
 
 indirect process 103 
 
 Koppers process 114 
 
 Mallet process 122 
 
 Mont-Cenis process 124 
 
 Otto process 125 
 
 Still process 118 
 
 Strommenger process 127 
 
 Sulphate plant, Walker 105 
 
 Wilton 107 
 
 Sulphonation process 43 
 
 Sulphur required, Feld process . . 146 
 
 Sulphuric acid 227 
 
 from spent oxide 191 
 
 specific gravity 227 
 
 used 108-139 
 
 Synthesis of naphthalene 52 
 
 Synthetic ammonia 110 
 
 drugs 49 
 
 perfumes 47 
 
 phenol 49 
 
 Table I, Tar yields 9 
 
 II, Composition of tars .... 9 
 
 III, Tar distillation 10 
 
 IV, Feld hot tar washing. .-. 26 
 
 V, Brown coal products. ... 30 
 
 VI, Products of distillation . 30 
 
 VII, Ilsider results 31 
 
 VIII, Moncieu results 32 
 
 IX, Naphthalene saturation 53 
 
 X, Naphthalene extraction . 62 
 
 XI, Distribution of nitrogen 67 
 
 XII, Ammonia strength .... 90 
 
 XIII, Influence of foreign 
 matter 109 
 
 XIV, Production and con- 
 sumption of sulphate Ill 
 
 XV, Sulphate consumption . Ill 
 
 XVI, World's sulphate pro- 
 duction 112 
 
 XVII, Ammonia content .. . 143 
 
 XVIII, Cost of producing 
 sulphate 153 
 
 Table XVIIIA, Benzol specifica- 
 tions 178 
 
 XIX, Prussian blue tests . . . 210 
 
 XX, Volume correction 223 
 
 XXI, Specific gravity, am- 
 monium sulphate 224 
 
 XXII, Specific gravity, sal 
 ammoniac 224 
 
 XXIII, Specific gravity, 
 caustic ammonia 225 
 
 XXIV, Specific gravity, 
 sulphuric acid 227 
 
 XXV, Strength of caustic 
 soda 228 
 
 XXVI, Reduction of gas 
 volume 230 
 
 XXVII, Ferrous sulphate... 232 
 
 XXVIII, Aqua ammonia ... 233 
 
 XXIX, Allowance for tem- 
 perature 234 
 
 XXX, Gas constants, facing 
 page 235 
 
 XXXI, Analysis of gas liquor 234 
 
 XXXII, Composition of tar 
 fractions 236 
 
 Tar composition 7 
 
 crude substances 42 
 
 dehydrating plant 15 
 
 dehydration of 35 
 
 distillation 10 
 
 distillation of 37^0 
 
 electrical precipitation of . . . 33 
 
 extractor, centrifugal 13 
 
 extractor, cyclone 12 
 
 extractor, injector 12 
 
 extractor, P. and A 11 
 
 extractor, rotary 13 
 
 extractor, static 33 
 
 extractor, static centrifugal . 14 
 fractions, composition of . . . 236 
 
 from coal gas 5 
 
 high temperature S 
 
 horizontal still 37 
 
 in saturators 151 
 
 low temperature 8 
 
 production, Bunte 7 
 
 production, influence of 
 oxygen 7 
 
244 
 
 INDEX 
 
 Tar production, Paris experi- 
 ments 7 
 
 products 5, 35 
 
 products, Feld system 17 
 
 products, nature of 18 
 
 removal of 10 
 
 tests 200-201 
 
 vertical retort 8 
 
 vertical still 36 
 
 washer, Feld 25 
 
 yields, Table 1 9 
 
 Temperature for ammonia re- 
 moval 4 
 
 for heavy oil removal 4 
 
 for hydrogen sulphide re- 
 moval 4 
 
 for middle oil removal 4 
 
 for naphthalene removal .... 4 
 
 for pitch removal 4 
 
 of sulphatation 155 
 
 Test for ammonia 205 
 
 cyanogen 207 
 
 naphthalene 203 
 
 Prussian blue 208 
 
 tar 200 
 
 Tests, benzol 214 
 
 solvent naphtha 214 
 
 toluol 214 
 
 Theory of condensation 4 
 
 Thiosulphate liquor 145 
 
 Toluol 165 
 
 specifications 222 
 
 tests 214 
 
 Treatment of gas 2 
 
 Trinitrotoluene manufacture .... 47 
 
 U 
 Usual method of condensation. . . 
 
 Value of benzol 165 
 
 constituents 2 
 
 dehydrated tar 16 
 
 Vertical "Feld" washer 60 
 
 retort tar 8 
 
 "Standard" washer 61 
 
 tar still 36 
 
 Volume correction 223 
 
 W 
 
 Walker sulphate plant 105 
 
 Wash oil for benzol extraction . . 168 
 
 required 187 
 
 Washer insulation 23 
 
 Washers, benzol 178 
 
 Washing mediums 23 
 
 Water gas tar as absorhent 58 
 
 required in concentrating . . . 100 
 
 Weak liquor, composition of ... . 93 
 
 Wilton sulphate plant 107 
 
 Wright's experiments 8 
 
 Wyne-Roberts experiments 9 
 
 X 
 
 Xylol 165 
 
 boiling point 165 
 
 specific gravity 165 
 
 Zinc reactions, Feld 131 
 

 
 
 
 
 
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