THE UNIVERSITY OF MICHIGAN INDUSTRY PROGRAM OF THE COLLEGE OF ENGINEERING THE MANUFACTURE OF HIGH BTU GAS FROM COAL Ralph H. Schatz This report has been prepared under the sponsorship of the Michigan Gas Association. November, 1957 IP-246

TABLE OF CONTENTS Page LIST OF TABLES......................................... iv LIST OF FIGURES........................................ v I. INTRODUCTION..................................... 1 II. GASIFICATION..................................... 3 A. Fixed-Bed Countercurrent Gasifiction...... 9 1. High Pressure......................... 9 2. Atimospheric Pressure.................. 13 B. Fixed-Bed Concurrent, Down-Draft Gasification..................................... 1 C. Gasification in Suspension................. 1 1. Fluid-Bed Gasification................ 18 2. En-trai- ed Solids Gasification......... 20 III. PURIFICATION OF SYNTHESIS GAS.................... 2 A. Dry-Box Process............................ 2 B. GirboLol Process........................... 27 C. Shell Phosphate Process.................... 29 D. Rectisol Process........................... 29 E. Other Processes............................ 30 F. Conclusions............................... 32 IV. METHANrE SYNTHESIS.....3.................... 35 A. Chemistry and Thermodynamics............... 32 B. Catalysts and Processes.................... 39 1. Catalysts............................. 40 (a) Iron Catalysts.................. 43 (b) Nickel and Cobalt Catalysts.... 4 (c) Carbon Deposition............... 50 (d) Fluid-Bed Catalysts............. 1 (e) Summary......................... -4 2. Processes.............................4 C. Conclusions................................ 56 ii

TABLE OF CONTENTS CONT'D Page V. ECONOMICS OF HIGH BTU GAS MANUFACTURE FROM COAL... 59 A. Basis for Economics........................ 59 B. Investment Costs........................... 61 C. Production Costs........................... 64 D. Effect of Fuel Cost........................ 56 E. Comparitive Cost of Natural Gas and High BTU Gas.................................... 66 F. Summary.................................... 68 VI. CONCLUSIONS........................................ 71 iii

LIST OF TABLES Table Page I Relative Cost of the Various Steps in Manufacturing High BTU Gas From Coal as a Percentage of the Total Cost........................... 3 II Calculated Results of Gasification of Char at Elevated Pressures and Temperatures....... 8 III Typical Composition of Synthesis Gas Produced in Lurgi Gasification Units..................... 12 IV Chemical Processes for Hydrogen Sulfide Removal......................................... 33 V Evaluations of Fixed-Bed Methanation Catalysts.. 42 VI Evaluations of Fluid-Bed Methanation Catalysts.. 32 VII Analysis of a Typical Product Gas Produced by Methanation with a Nickel Catalyst.............. 3 VIII Gas Compositions at Various Points in a High BTU Gas Plant.................................. 62 IX Investment Cost Summary.................... 63 X Production Cost Summary 6..................5.... iv

LIST OF FIGURES Figure Page 1 Flow Diagram of the Girbotol Process............. 28 2 Flow Diagram of the Rectisol Process............. 31 3 Effect of Temperature, Pressure, and Feed H2/CO Ratio on the Composition of the Product Gas Which Satisfies Both the Methane and the Water-Gas Equilibria....................................... 3 Gross Heating Value of the Scrubbed Product Gas as a Function of the Feed Gas Composition Using a Steel Catalyst................................ 45 Effect of Temperature and Feed Carbon Monoxide Concentration on the Amount of Carbon Deposition Which Occurs During Synthesis Using a Steel Catalyst......................................... 6 Effect of Temperature on the Product Gas Composition Using a Cobalt-Thoria Catalyst.......... 4 7 Schematic Diagram of a High BTU Gas Synthesis Plan............................................ 60 Effect of Fuel Cost on the Cost of Producing High BTU Gas.......................................... 67 v

I. INTRODUCTION Increasingly larger quantities of fuel gas are being used yearly in the United States for both domestic and industrial heating. by consumers who prefer the convenience, low cost, and cleanliness of gas to other sources of energy. Natural gas from gas wells located primarily in the Southwest accounts for over ninety percent of the fuel gas consumed yearly, and this percentage is increasing with each succeeding year. Like most natural resources, natural gas is not an inexhaustible commodity, and the gas industry has a vital interest in having a substitute for natural gas available if the need should arise. While the existence of large known reserves of natural gas and the continual discovery of new reserves insures the availability of this source of energy for many years, nevertheless, the dedication of large portions of these reserves to existing inter-state pipe lines and to large volume users at the point of production may in the near future make natural gas unavailable or in short supply in many areas of the country not served directly by pipe lines. It is becoming increasingly apparent that the possibility of obtaining fuel gas from other sources should be considered long before the day when natural gas reserves actually near exhaustion. The existence of extensive known coal reserves in the United States makes coal a logical choice as the raw material for fuel gas manufacture. Combustible gas mixtures can be manufactured by gasifying coal with air or oxygen and steam, using any one of several processes. The gas mixtures produced, however, are of low -1

-2heating value. At present there is not a one-step gasification process known which can deliver a gas with a heating value greater than about 500 BTU/SCF. Long distance transmission lines and utility equipment in use today have, for the most part, been designed for natural gas which has a heating value of about 1000 BTU/SCF. Because the transmissions of low BTU gas and also the conversion of utility equipment to accommodate a 450 to 500 BTU gas may be prohibitively expensive, low BTU fuel gas manufactured from coal must be upgraded to a high BTU gas comparable to natural gas if it is to serve as an attractive substitute. Upgrading can be accomplished by catalytically converting the hydrogen and carbon monoxide present in the gaseous products from coal gasification to methane. Processes and catalysts for methane synthesis have been studied extensively. An overall process for the conversion of coal to high BTU fuel gas would require five steps: (1) gasification of coal with steam and oxygen to produce synthesis gas; (2) removal of carbon dioxide and sulfur compounds from this synthesis gas; (3) catalytic conversion of synthesis gas to methane; (4) removal of carbon dioxide from the gas produced; and (5) compression of the purified high BTU gas to transmission pipe line pressure. It is the purpose of this report to review briefly the processes which may be used for gasifying coal and for purifying synthesis gas, to discuss in detail the catalysts which have been investigated and developed for the methanation reaction, and to present economics for the overall process of making high BTU gas from coal in the light of present-day costs.

II. GASIFICATION Because there is not a one-step process known which will convert coal or other solid fuel to a fuel gas with a heating value greater than about 500 BTU/SCF, any process to manufacture high BTU gas of 850 to 1000 BTU/SCF from coal must necessarily consist of a number of sub-processes. Gasification of the solid fuel to produce a mixture of CO and H2, generally referred to as "synthesis gas," is the first step of an integrated process (which includes purification and methanation steps) to produce high BTU gas. The relative importance of the gasification step compared to the subsequent processing steps is illustrated by Table I which shows the share which each sub-process contributes to the overall cost of manufacturing high BTU gas from coal. TABLE I. RELATIVE COST OF THE VARIOUS STEPS IN MANUFACTURING HIGH BTU GAS FROM COAL AS A PERCENTAGE OF THE TOTAL COST Gasification 70.0% (including raw materials cost) Initial Purification 11.4 Methanation 4.1 Final Purification 2.4 Final Compression 2.2 Auxiliaries 9.9 Total 100.0% Fuel Cost (coal) 29.2% Raw Material Cost 60.5% -3

-4The data of Table I are developed and discussed in greater detail in a later section on the Economics of High BTU Gas Manufacture, but serve at this point to show that gasification of the solid fuel plus purification of the synthesis gas prior to upgrading comprise about 80 of the total cost of producing high BTU gas. The need for developing efficient and inexpensive methods of gasification, as well as improved methods for mining coal, are obvious. A great deal of research is justified to devise means by which the cheapest of fuels can be converted to synthesis gas in large units with a minimum of operating expense. Gasification of coal is by no means a new development. There are numerous references in the literature to gasification processes, including foreign developments. A historical review of gasification developments will not be attempted, but rather this section will be devoted to discussing those processes which appear particularly promising for production of synthesis gas from American coals. These, to a large degree, are of the coking variety, and therefore, gasification units which may ultimately be developed for processing American coals are likely to differ from those developed in Europe where non-coking coals are available in great quantity. The reactions which take place and the products formed when solid carbonaceous fuel is gasified by oxygen and steam may be described by the following equations: C + H20 < CO + H2, AH = 58,340 Btu/lb mol C at 1300~ K. (1) C + 2H20 i CO2 + 2H2, AH = 44,635 Btu/lb mol C at 1300~ K. (2)

-5C + 1/2 02 < CO, AH = -48,982 Btu/lb mol C at 1300~ K. (3) C + 02 _ C02, AH = -170,021 Btu/lb mol C at 1300~ K. (4) The formation of CO, CO2 and H2 according to equations (1) and (2) requires a high reaction temperature. Moreover, the two reactions are endothermic in nature, making it difficult to maintain the required temperature. Early practice in water-gas manufacture was to blow a fixed-bed of fuel with air [equations (3) and (4)] in order to raise the temperature of the bed to a desired level and then to pass steam through the hot bed to produce a mixture of H2 and CO [equation (1)]. This type of operation was necessarily cyclic in nature, because the fuel bed was continuously cooled during the endothermic water-gas step and periodically had to be reheated by air-blowing. The gas produced during the water-gas operation had a gross heating value in the range of 250 to 350 BTU SCF after H20 and CO were removed. It was not economical practice to combine the pre-heating step with the water-gas step by simultaneously passing air and steam through the fuel bed, because the inert nitrogen present in the blast air ended up in the product, reducing its heating value. Synthesis gas which has been produced by gasification of coal with air and steam cannot be upgraded to high BTU gas by methane synthesis, if by high BTU gas we mean a product with a heating value above 350 BTU SCF, because of the presence of inert nitrogen which is introduced with the air blast and acts to dilute the heating value of the product.

-6 The development of economical processes to produce large quantities of relatively pure oxygen, e.g., the Linde-Frankel process, eliminated the need for cyclic operation in water-gas manufacture. Oxygen and steam could be simultaneously introduced to produce a product containing CO, H2, C02, H20 with only a small amount of inert nitrogen. The availability of pure oxygen is of fundamental importance, for without oxygen the production of high BTU gas is virtually impossible. The subject of oxygen production will not be discussed other than to say that commercially proven processes, which are based upon low temperature distillation of liquid air, are available. Unfortunately, however, oxygen is expensive; the cost of oxygen used in gasification accounts for better than 10% of the cost of the final high-BTU product. Any economies which can be made in oxygen production will, therefore, be a primary factor in reducing the cost of producing high BTU gas. The American Gas Association has sponsored research in which the effect of gasification temperature and pressure on the equilibrium composition of the products from gasification of char have been determined.1 These calculations show that an increase of reaction temperature brought about by increasing the oxygen to steam ratio of the blast gas decreases the H2 to CO ratio of the product gas. Oxygen consumption is increased and steam consumption is reduced. Total gas yield of net combustibles (CO + H2 + CH4) is virtually unchanged. Increasing the gasification pressure greatly reduces oxygen consumption at a given gasification temperature. The yield of net combustibles (CO, H2 and CH4) falls off slightly with

-7higher pressure, but the yield of methane increases appreciably. Increased pressure results in an increase in the product H2 to CO ratio, as shown by the following tabulation. GASIFICATION OF CHAR H2/CO Ratio at 700~C and Various Pressures1 Pressure, atm. H2/CO ratio 1 1.2 5 1.97 10 2.38 20 2.78 85 3.52 At pressures above 10 atm. the gas composition has reached the point where the H2 to CO ratio is out of the range for complete stoichiometric reaction, and the gas becomes more and more unsuitable for upgrading. Raising the operating temperature will remedy this situation, however, since higher temperatures decrease the H2 to CO ratio of the product gas. The calculated results of gasification of char at 85 atm. and 800~C are shown in Table II. At temperatures in the region of 8000C and above, ash formed by gasification can be above its softening point. In this event, ash must be removed from gasification reactors in the form of liquid slag. Another method of adjusting H2 to CO ratios of product gas to desired values can be utilized at temperatures below the ash softening point. Carbon dioxide can be recycled to the gasification reactors, thus shifting the equilibrium described by the equation

-8H2 + C02 < CO + H20 (5) to form more CO. Theoretically, gas with a suitable H2 to CO ratio for upgrading can be produced at any pressure and suitable gasification temperature by using proper ratios of oxygen to steam in the blast gas, and if necessary by recycling carbon dioxide. TABLE II. CALCULATED RESULTS OF GASIFICATION OF CHAR AT ELEVATED PRESSURES AND TEMPERATURES1 Pressure, atm. 85 Temperature, ~C 800 Blast Gas Composition, vol. % H20 91.3 02 8.3 N2 0.4 Product Gas Composition, o As produced H20 + C02 free CO 12.6 25.8 C02 17.8 H2 24.7 50.5 H20 33.3 CH4 10.7 21.9 H2S 0.2 0.4 N2 0.7 1.4 100.0 100.0 H2/CO ratio 1.96 Gas Yield SCF/lb. fuel Total 65.3 Net Combustible (CO + H2 + CH4) 31.4 Three main principles of gasification have shown promising results. These are: 1. Fixed-bed countercurrent gasification operating at either elevated or atmospheric pressure and employing oxygen-steam blast. 2. Down-draft gasifica-tion; e.g., the Flesh-DEMAG process. 3. Gsific LtioYr in suspension.

-9Any of the above could serve as the basis for a process to manufacture synthesis gas from American coking coals. A short discussion citing the advantages, disadvantages and future potentialities of each follows. A. Fixed-Bed Countercurrent Gasification The upward flow of blast gas (oxygen and steam) and the slow downward movement of the fuel bed are characteristic of the countercurrent gasification process. Countercurrent reactors have deep fuel beds with a large excess of carbon present. A sequence of zones exist in the fuel bed. In the direction of gas flow these are the ash zone, the oxidation zone, the reduction zone, and the preheating zone. Countercurrent processes, which are fed by gravity flow of the fuel, require a uniformly sized, non-agglomerating fuel with a minimum of fines to insure uniformity of flow. Classical gasification fuels are lump coke, double-screened anthracite, noncaking bituminous coals, and briquetted brown coal or lignite. Countercurrent processes have been operated at pressures from atmospheric up to 25 atmospheres. Gasification at higher pressures, i.e., in the range of pipe line pressures, may prove possible and desirable. The advantages of both high pressure and atmospheric pressure gasification are discussed below. 1. High Pressure High pressure fixed-bed gasification as exemplified by the Lurgi process appears at present to be one of the most promising methods for producing synthesis gas from coal. The fixed bed countercurrent system is high in efficiency. Furthermore, high pressure gasification increases the throughput of fuel and blast gas

-10considerably. The synthesis gas produced is at high pressure which aids in subsequent purification and methanation operations and also reduces gas compression load. If fixed-bed gasification pressures can ultimately be raised to a level comparable to longdistance pipe line pressures, which range from 800 to 1200 psig, the final product gas compression step can be eliminated entirely. Elimination of the compression step would reduce the cost of producing high BTU gas by about 2% as shown by the information of Table I, and decrease the required plant investment by about $1,000,000. Gasification of non-caking fuels utilizing solid ash discharge has been proven in practice at medium pressures, i.e., around 300 to 350 psi, using Lurgi gas generators. The Lurgi generator reached a commercial state of development between 1930 and 1932. The first installation, which manufactured city gas from lignite, was built in 1936 at Hirshfelde, Saxony, to supply the town of Zittau. Since that time, there has been a great deal of information reported about the Lurgi gasification process. A recent bibliography cites 2 many of the published reports. Essentially, the high-pressure Lurgi gasification process consists of continuously blowing a hot bed of solid fuel with a mixture of 95% purity oxygen and superheated steam. Oxygen is mixed with steam to give about 10 vol % of oxygen in the blast mixture entering the fuel bed. The temperature of the bed is controlled by the proportion of oxygen used, and is maintained above 1900~F but below the softening point of the ash formed. Reactor pressures range from 20 to 25 atmospheres. Raising the gasification

-11pressure increases the heating value of the synthesis gas produced and reduces oxygen consumption; however, heretofore on the basis of economic and structural considerations the maximum reactor pressure has been considered to be about 30 atm. The Lurgi reactor itself is a cylindrical vessel which is refractory-lined and water-jacketed. Blaw-Knox Construction Company has reported that the optimum size is the maximum which can be shipped in one piece, i.e., a reactor with an I.D. of 11.8 feet.1 Fuel is supplied to the fuel bed by gravity flow from a hopper positioned above the reactor at a rate corresponding to the rate of descent of the fuel bed. In operation, the blast of oxygen and steam results in an intense reaction zone at the bottom of the fuel bed. Below this zone, fine ash collects. The ash is supported on a mushroom-shaped grate which is rotated by a hollow shaft that extends through the bottom of the reactor and further serves to supply the steam-oxygen mixture to orifices positioned below the ash grate. The slow rotation of the grate moves the ash to the periphery, and plow blades underneath the bottom side of the mushroom head move the ash into an opening at the bottom of the reactor. The ash falls through an ash pipe and into a pressure vessel below the reactor. When the ash hopper becomes filled, a valve between the hopper and the reactor is closed, the hopper pressure is reduced, and the ash is dumped without interrupting the gasification operation in the reactor. Product gas is removed through the top of the generator. A water jacket surrounding the reactor serves to cool the shell of the reaction vessel. Sufficient steam is formed to maintain a pressure

-12on the outside of the reactor shell equivalent to the gasification pressure,and thus, eliminate the structural problem of containing a gaseous mixture at 2000~F and 20 atm. in a steel vessel. Synthesis gas produced at high pressure in Lurgi generators contains from 5 to 15 percent methane, and has a heating value ranging from 350 to 500 BTU/SCF. A typical analysis of a gas produced from Germian "Braunkohle" and coke briquet is presented in Table III. TABLE III. TYPICAL COMPOSITION OF SYNTHESIS GAS PRODUCED IN LURGI GASIFICATION UNITS3 Volume % Component Raw Gas Purified Gas CHH 14.9 20.6 H2 35.6 52.2 CO 12.9 18.3 CnH2n+2 0.8 0.7 H2S 2.0 CO0 32.1 6.8 02 0.2 0.2 N2 1.5 1.2 100.0 100.0 Heating Value, BTU/SCF 342 475 Before upgrading the crude synthesis gas by methanation, a purification step to remove C02, H2S and any other organic sulfur cornpounds is required. A typical composition of purified raw gas is shown in the second column of Table III.

-13In the past, the Lurgi process has primarily used predried lignite for solid fuel, but runs have also been made with coke, char from bituminous coal, semi-anthracite and bituminous coal. These all fall in the non-caking or only weakly agglomerating class. While successful operation in pilot plant scale equipment has been reported4'5 with caking bituminous coals by using a mechanical stirrer, successful operation with caking coals has yet to be demonstrated in commercial-size Lurgi generators. In summary, the advantages of high pressure gasification may be listed as: (a) High fuel rates. (b) Lower oxygen consumption. (c) Increased methane content in the product gas. (d) Lower gas compression requirements. Disadvantages are: (a) Inaccessability to equipment because of high pressure. (b) The need for mechanical agitation when processing caking coals. 2. Atmospheric Pressure While high pressure and medium pressure gasification processes have the advantages of large capacities, lower oxygen consumption, and the need for less gas compression in subsequent processing steps, there are advantages of atmospheric pressure gasification which should not be disregarded. The most important of these is the availability of suitable gas generators which have proven by years of commercial operation to be dependable. Moreover, atmospheric

-14pressure gasification permits greater accessibility to gasification equipment. Direct observation of the fuel bed is possible and corrective measures can be taken if visual observation indicates that gasification is not proceeding in the desired manner. Slagtapping operation, by which ash is drawn off from tne reactor in the form of liquid slag, is easily applied to atmospheric pressure gasification units. The fuel beds of updraft gasification processes are usually supported by grates. Rotating grates or fixed grates with rotating stirrers serve to continuously remove the ash produced during gasification. Such mechanical grates impose limits on the size of reactors to diameters of 10 to 12 feet. However, by raising the temperature of gasification so that the ash produced is above its melting point, it is possible to remove ash from gasifiers by tapping it off in the form of a liquid slag. The need for rotating grates or stirrers is, thus, eliminated, avoiding the high maintenance costs associated with such equipment. Basically, slagging ash gasifiers are simple in design and can be built in large units. A unit has been built and operated which can gasify 300 tons of coke per day.6 Moreover, based on experience with blast furnaces, which closely resemble slagging ash gasifiers but are several times larger, higher capacity units should be possible. The yield of methane from atmospheric-pressure gasifiers is considerably lower than that obtained from high pressure gasification. As a result, more oxygen is required to gasify coal at

-1)atmospheric pressure since the solid carbon is converted to gas primarily by reaction with oxygen and to a lesser extent by combination with hydrogen to form methane. To summarize, atmospheric pressure countercurrent gasifiers are of two major types: rotary-grate generators.and slagging ash generators. While atmospheric pressure rotary-grate generators are now available, they do not appear promising for future use in processes to manufacture high BTU gas particularly because of high maintenance costs and size limitations. Slagging ash gasifiers, however, are worthy of further consideration for they possess a number of desirable characteristics: (a) They can be built in large units of high capacity. (b) Operation at atmospheric pressure eliminates problems associated with operating and maintaining high pressure equipment. (c) Low-grade, high-ash caking coals can be gasified, and valuable by-products, mainly pig iron and slag for building materials can be recovered. (d) High gas yields can be obtained per unit of fuel gasified. Important disadvantages of atmospheric pressure slagging ash gasifiers, however, are increased oxygen consumption and the necessity of compressing product gas from atmospheric to pipe-line pressure. B. Fixed-Bed Concurrent, Down-Draft Gasification The use of concurrent flow of blast gas in gasification has found application primarily (1) to reduce carry-over of fines

-loat high rates of gasification and (2) to destroy tar and other volatile materials resulting from gasification of lowr-grade fuels, such as wood and peat, by passing the volatiles through the oxidation zone of the fuel bed. With down-flow of blast gas, the fuel bed acts like a filter, retaining small fuel particles and permitting the use of gas rates higher than are possible with any other type of fuel bed. Blast rates have been achieved that would have fluidized an updraft fuel bed of small-sized fuel. Unfortunately, down-draft systems are subject to limited rates of ignition; however, this difficulty can be overcome by preignition of the gasification fuel. A unique method for ignition is employed in the Flesh-DEMAG gasifier in which the entire inventory of the generator is periodically fluidized in order to distribute ignition nuclei throughout the fuel bed and produce practically an unlimited ignition rate. A large variety of fuels, including caking coals can be gasified in a down-draft fixed bed if the bed is periodically renewed by up-draft fluidization for declinkering and mixing in fresh feed. The Flesh-DEMAG process is a continuous gasification process employing two (or more) reactors which are interconnected at their upper ends. Blast gas is admitted at the top of both shafts and product gas is removed through the rotary grate which supports the fuel bed. Periodically, the flow of gas in one reactor is reversed and the reactor charge is fluidized using either blast gas or steam. Clinker which has formed during gasification breaks up and falls down to the rotary grate from which it is continuously removed. During the fluidization, fresh feed is introduced and is thoroughly

-17mixed with the reactor inventory. The off-gas from the fluidization Tperiod flows through the connecting pipework between the two reactors and is passed in down-flow through the hot bed of the second reactor. When the fluidizing operation has been completed, the bed is allowed to settle, and introduction of the blast gas at the top of the reactor is resumed. The second shaft is then fluidized in a like manner. To date, the process has been dem)-_istrated only on a semi-comnercial scale in two 5-1/4 ft. I.D. reactors operated at Ludwigshafen, Germany, by DEMAG, which can gasify 100 tons of fuel per day at pressures up to 10 atmospheres.7 To summarize the advantages of down-draft gasification processes and, in particular, the Flesh-DEMAG process: (a) Down-draft gasification processes can gasify with extremely high flow rates of oxygen and steam without appreciable carry-over of fine fuel particles. (b) The Flesh-DEMAG process is able to handle a large variety of fuels which include caking coals. (c) The Flesh-DEMAG process has the advantage over all other types of fixed-bed processes that the fuel bed is replenished and rebuilt before the beginning of each gasification cycle; i.e., about every 15 to 20 minutes. C. Gasification in Suspension Processes which gasify coal in suspension fall into two categories, those which use a fluidized bed of finely divided fuel, and those in which pulverized coal particles are entrained in the

-18moving blast gas stream. Both techniques can be carried out at elevated pressures. Gasification in suspension has the advantage of widely separating individual fuel particles which eliminates many of the shortcomings of fixed fuel beds. Particles of fuel can swell, soften and even slag without interfering with adjacent particles, and without changing the characteristics of the fuel bed. Separation of individual particles makes gasification in suspension much less sensitive to the type of fuel used than is the case with fixed-bed gasification processes. There is much research being devoted to suspension gasification processes because of their great potential as means for handling the many varieties of strongly caking American coals. 1. Fluid-Bed Gasification The success of fluid-bed processes for catalytic cracking and reforming has led to the application of fluid techniques to gasification of pulverized coal. Fluid-bed gasification processes operate by passing blast gas upward through a bed of finely divided fuel particles at velocities such that the solid particles are suspended in the rising gas and are maintained in continual random movement in a dense phase, wherein the solid fuel and blast gas react to form a product gas comprised of CO, H2, C02, CH4, CnH2n+2, and H20. A characteristic of fluid-bed operations in general is the absence of an appreciable temperature gradient within the dense bed, since the motion of the solid particles tends to equalize the temperature. Gasification reactions in the fluid bed take place at lower maximum particle surface temperatures than those observed in fluid-beds.

-19Chemical reactivity of the fuel tends to become important at lower surface temperatures. However, this effect can be minimized by operating fluid bed gasifiers at temperatures close to the softening point of the ash formed. Large fluid-bed gasifiers are possible. Reactors are essentially empty vessels with auxiliary equipment for introducing pulverized fuel and blast gas. Sizes are limited primarily by the ability to maintain good fluidization. Uniform distribution of blast gas to the fuel bed is of the utmost importance since by-passage of oxygen through the bed in the form of bubbles can lead to the formation of explosive product gas mixtres. A number of oil companies in the United States have conducted extensive research programs on fluid-bed gasification processes as sources of carbon monoxide-hydrogen mixtures for liquid fuels synthesis. Unfortunately, the resalts of these research programs, for the most part, have not been reported in the literature. This fact does not necessarily indicate that the fluid-bed gasification process lacks potential, but rather reflects a preoccupation on the part of oil companies in processes of more current interest and also a reluctance to release information which ~,ight be of interest to competitors. Actually, development of fluid-bed gasification processes has reached an advanced stage. Problems such as nonuniform blast gas distribution and stalactite formation of partially slagged ash particles at points in the reactor remain to be solved, but it is felt that with sufficient incentive a commercial process could be perfected.

-202. Entrained Solids Gasification A number of pilot plants have been built both in Europe and in the United States to investigate pulverized coal gasifica8,9 tion processes. The Bureau of Mines has conducted an extensive research program, and has demonstrated in pilot plant scale equipment that it is possible to gasify any type of American coal for short periods of time at pressures up to 300 psi with carbon conversions of 90 percent or higher.8 However, the equipment used has not been satisfactory for long periods of operation. Gasification of entrained solid fuel requires a high temperature level which can be provided by using either high oxygen concentrations or highly superheated steam in the blast gas. Temperatures as high as 1800~F are common. Ash removal in the form of a liquid slag is preferred. Elevated pressure operation is desirable in order to reduce the size of the reactor and auxiliary equipment as well as to reduce the amount of final product gas compression required. Small pilot plant units which have been used thus far suffer from high external heat losses and from erosion of reactor walls. Data obtained from these units are not easily extrapolated nor are they sufficient for the design of commercial size installations. Runs in larger pilot plants or in full-size experimental units are needed to fully demonstrate the gasification of American caking coals in entrained solids type equipment. D. Summary and Conclusions Three principle types of gasification processes, countercurrent, concurrent, and suspension gasification show promising

-21results. Extensive research programs both in the United States and in Europe have led to the development of a number of processes which are capable of producing synthesis gas from coal. Some of these are well advanced, and the equipment has been tested and perfected in many years of commercial operation. Still other processes have not passed out of the pilot plant stage, and much research will be required before successful commercial operation can be realized. Those processes which have already reached commercial status have used fuels that are not comparable to the types of coal which must be used in this country, and while they provide a "jumping-off point" for future research, they do not by any means provide all the information needed to design a process for use with American caking coals. Further research is needed to develop the optimum gasification process, which not only will be able to handle caking coals, but which will efficiently and economically produce synthesis gas at pressures equal to or above pipeline pressures. The three process which appear to offer the firmest foundation for future efforts to develop better equipment and improved operating procedures are: 1. The Lurgi countercurrent updraft fixed-bed gasification process. 2. Downdraft two-shaft gasification of the Flesh-DEMAG type at elevated pressure. 3. Suspension gasification either in a fluid-bed or as an entrained solid at elevated pressures.

-22The Lurgi gasifier, using a fixed fuel bed blown countercurrently with oxygen and steam, appears to be one of the most promising gasification processes. Modifications which would make the process suitable for producing synthesis gas from American coals for high BTU gas manufacture are: 1. The development of mechanical methods for breaking up caked ash to permit use of the process with caking coals. 2. Redesign of the equipment to operate at or above pipeline pressures. 3. Incorporation of slagging ash operation into the high pressure Lurgi design. The Flesh-DEMAG process and suspension gasification processes offer great promise, but need to be advanced from the pilot-plant stage to semi-commercial or commercial size units to gain experience necessary to improve designs. However, large size units are costly and based on present-day economics for highsBTU gas manufacture it is doubtful that commercial size units can be justified for research purposes. To summarize the present status on gasification processes, it can be said that while operable commercial processes do not now exist, basic processes are available which can be adapted to produce synthesis gas from American coals. Gasification and raw materials comprise the major cost of producing high BTU gas from coal, and any economies in gasification costs will do much to speed the day when high BTU gas can economically be made from coal. In turn,

-23favorable economics for the integrated high BTU manufacturing process will offer great incentive to develop improved gasification processes. Factors other than technological improvements in gasification processes should not be overlooked. The cost of raw materials is an important part of the total gasification cost. Improvements in mining methods and in oxygen manufacture would lower raw material costs and result in more economical gasification. Research along these lines could be of benefit not only to gasification, but to a number of other non-related processes.

III. PURIFICATION OF SYNTHESIS GAS When coal is gasified, much of the sulfur present is converted to gaseous sulfur compounds. Since all of the catalysts presently known to be active for methanation are susceptible to poisoning by sulfur, the purification of synthesis gas is an important step in the manufacture of high BTU gas. Sulfur is present in raw synthesis gas primarily in the form of hydrogen sulfide, carbonyl sulfide and carbon disulfide, and can range in concentration from about 200 to 500 grains per 100 SCF of synthesis gas. In order to obtain reasonable catalyst life with both nickel and iron methanation catalysts, it has been found necessary to reduce the sulfur content of synthesis gas to 0.1 grain of sulfur per 100 SCF or lower. A large number of processes have been developed and used commercially for the removal of sulfur compounds from gaseous streams. A detailed bibliography of literature references is available.10 Several of the sulfur removal processes commonly used in the refining and the manufactured gas industry will be briefly described. A. Dry-Box Process The dry-box process has long been used in the gas industry for removing relatively low concentrations of H2S from industrial gases. Hydrated iron oxide is coated on wood shavings or other supporting material and spread in thin layers on trays in rectangular boxes. Raw gas to which sufficient air or oxygen has been added to provide an oxygen concentration of 0.6 to 1.0 volume % is passed over the iron oxide. Hydrogen sulfide present reacts to form iron -23

-26sulfide, which in turn is oxidized by oxygen added to the gas stream back to iron oxide and sulfur. After continued operation it becomes necessary to thoroughly revivify the iron oxide sponge, either by passing air or oxygen through the beds or by removing the iron oxide from the boxes and exposing it to the atmosphere for a prolonged period. The iron oxide sponge is finally discarded when the total sulfur content reaches 50 to 60%. The dry box process has the disadvantage of not removing organic sulfur compounds. Consequently, for purifying synthesis gas a two-step process is required. The Germans employed such a combination process for the purification of the feed gas used in the synthesis of liquid fuels. Raw synthesis gas was passed over a mixture of hydrated iron oxide and sodium carbonate at 200 to 3000C. Organic sulfur components present were catalytically converted to hydrogen sulfide which was subsequently removed in dry-boxes. Major advantages of the dry-box technique are the completeness of hydrogen sulfide removal and the simplicity of the process. However, the dry-box process does not appear practicable for purifying synthesis gas in sufficient quantity for economic high BTU gas manufacture because of the prohibitively large amount of iron oxide sponge required. For example, a high BTU gas plant producing 100 million SCF per day of 850 BTU gas would consume 364 million SCF per day of synthesis gas. Since about six cubic feet of iron oxide sponge are used to remove the hydrogen sulfide from 1000 SCF per day of gas, in excess of two million cubic feet of iron oxide sponge would be required to purify the synthesis gas for a 100 million SCF per day high BTU gas plant.

-27B. Girbotol Process The Girbotol process, developed by the Girdler Company, has become the most widely used process for the removal of hydrogen sulfide from natural and refinery gases.ll112 The process utilizes one of the ethanol amines (mono, di, or tri) to absorb hydrogen sulfide from gaseous streams in conventional bubble-tray towers. Diethanolamine is most commonly used in the treatment of refinery gases because of its inertness toward carbonyl sulfide. Diethanolamine exhibits basic properties and absorbs acidic components (H2S and C02) as shown by the following reactions: 2R2NH + H2S < (R2NH)2 H2S (6) 2R2NH + H20 + C02 4 (R2NH)2 H2C03 (7) where R represents a (CH2CH2OH +) group. Regeneration of the absorbent is possible, since at the boiling point of the solution the amine becomes practically non-alkaline, releasing the H2S and C02 from solution. The process is carried out in a two-column system as shown in Figure 1. The relatively high cost of the ethanolamines necessitates special precautions to prevent excessive losses of the absorbent. Monoethanolamine, which has a high absorption capacity, reacts with carbonyl sulfide to form the relatively heat-stabel compound diethanolurea, CO(NHCH2CH20H), and thus, cannot be used economically with gas streams containing carbonyl sulfide, as is the case with raw synthesis gas produced from coal. In addition, impurities such as tars, cyanogen, and oxygen, which are often present in manufactured

-23C, w 00 \ o w O I - --- o Q4~~~O~~ ~ 0 H1 U) CC W <I:f C jI 0 IL cr I —--- 0 42 z aH o ~LLC~) LL) ~~~~~~xO 0 D C _, 3W:o w 44Q O Q~~~~~~~~~~~~~U

-29gases react to form non-regeneratable compounds with the ethanolamines, Thus, while the Girbotol process would be satisfactory process-wise for the purification of synthesis gas, the presence of reactive impurities in the raw gas is likely to result in excessive absorbent losses. C. Shell Phosphate Process The Shell phosphate process13 is essentially similar to the Girbotol process, but uses an aqueous solution of tripotassium phosphate as absorbent. The solution used normally contains from 40 to 45 wt. % tripotassium phosphate. Hydrogen sulfide is absorbed according to the reaction: K3P04 + H2S > K2HP04 + KHS (8) Regeneration is accomplished by steam stripping. In general, the Shell phosphate process has higher operating costs than the Girbotol process. D. Rectisol Process Currently, the most promising process for the removal of sulfur from synthesis gas is the Rectisol process, developed cooperatively by Linde and Lurgi in Germany. The removal of moisture, carbon dioxide, all forms of sulfur, and gum-forming materials are combined into a single operation, thus, simplifying the purification step considerably and providing the Rectisol process with a distinct advantage over all other purification processes. The first Rectisol installation has recently been constructed as a part of the South African Fischer-Tropsch plant.

-30A simplified flow diagram for the Rectisol process is shown in Figure 2. Raw synthesis gas is contacted with methanol in an absorber which operates at -30~F and 500 psig. Purified synthesis gas leaves the top of the absorber, and after heat exchange with the raw gas is ready for the synthesis step. Methanol rich in sulfur compounds and carbon dioxide leaves the bottom of the absorber and is partially flashed. A portion of the methanol from the flash tank is pumped back to an intermediate plate of the absorber and the balance is sent to the stripper. The flash gas, rich in gaseous sulfur compounds and carbon dioxide, is fed to the bottom of the reabsorber which operates at -30~F and atmospheric pressure. Sulfur is reabsorbed in methanol, and a carbon dioxide stream containing negligible quantities of sulfur compounds leaves the top of the reabsorber and is vented from the plant. The methanol stream from the bottom of the reabsorber and methanol from the flash tank are combined and pumped to the top of the stripper. Gaseous sulfur compounds and the remaining carbon dioxide pass overhead from the stripper. Clean methanol leaving the bottom of the stripper is split into two streams and returned to the top of the absorber and the reabsorber towers. In addition to the absorption and stripping towers, refrigeration equipment and an extensive heat exchanger system is required. E. Other Processes The removal of hydrogen sulfide from gaseous streams can be accomplished by other processes than those discussed. Several

-31cu 0.J ~~~0 0.( 0, (0.-o0~ -'I.V,, H38UOS8V3I IU 0 i__ X c>NVI. HSVI. 0 U. 8i38HiOSSV Li CD LL (/) a::E-U t9 tJ38tlOS8~~~~~3t? (n 0. p~~~~~~~~~Lr < <~~~~ 0: 0~~~~~~~L

-32additional processes are briefly described in Table IV. The chemical reaction by which hydrogen sulfide is removed as well as the method of regeneration of the absorbent are given. Most of the processes listed are suitable only for removing hydrogen sulfide, though several will also remove carbon dioxide. Purification of raw synthesis gas produced from coal, which would contain hydrogen sulfide and carbon dioxide as well as organic sulfur compounds by any of the methods given in Table IV would necessitate a twostep and in some cases a three-step process. Organic sulfur compounds would first have to be catalytically converted to hydrogen sulfide. The hydrogen sulfide formed could then be removed by any one of the ten processes cited. However, if the hydrogen sulfide step does not also remove carbon dioxide, an additional step is required for this purpose. Carbon dioxide is commonly removed by water scrubbing. F. Conclusions High operating costs, expensive and large equipment installations, high absorbent losses, as well as the need for multiple processes, all serve as serious disadvantages to the above-described schemes for sulfur removal. The Rectisol process which can remove moisture, carbon dioxide, all forms of sulfur and gum-forming materials in a single operation, therefore, appears to be an outstanding process for use in purifying synthesis gas for high BTU gas manufacture. Reportedly, Rectisol gas purification cost has been cut to half the cost of purification with conventional methods. A good deal more information on the Rectisol process should be available when operations of the South African Fischer-Tropsch plant are reported.

-33CH 0 H r-I IJ (D r C _ d r edco cx i O *,H h ) 0 a a) 4- X.*H - -P -P -P - 0.q ~ > >.q > > >.< d- O o p Q C2 rU o Q r 4 0 Ci C - C CO Co o *HFI= ^ > - ~ 0 * (C tD.rcx O OC OO OC c i U) U) q cL\ U) U ) U X 4 Ca3 O 4) U) + C^ C H C] CH w 0P C 0 ~i U (\ 0 4 n 0 -4i 0;0 0 d H CD C r *H C d Cd, c q 0 i -r4 CO rQ rCd lMrr-lC J Q H o cCo - d0 M t0 x O m x 0 o rn O O U ioj + + cQ (X) H a.p + oix h c + + d + o 0 Cd ll (S c rgo p Fr - t + O O o C) U C I CU Co CC CQ i +, ) C) X C0 + Ut I U] U U t VW xM + 4 o ^ + + L~\ 0 + U J ) co + + 0 W O-U 1c O) J o2 -P N CM X cxi U) O + U)3~ CM U) U ) XCM F CM CM C\ rZ v 11o oI+ 0 0 0 Uo + + o 1f + N 0 U) + + V o CU C UR C) C - Z PO H. o o 0 Co o F H Co (x P C H F U) > 0 0 < HPL U CM C; PL \D0 t CO 0 H OJ O Ij- U YO 1>- 00 0\

-34Even with the Rectisol process, gas purification costs represent 14 percent of the total cost of manufacturing high BTU gas from coal. In view of the large potential savings which might be made by improving existing processes or by developing new processes for purifying synthesis gas, further research in this area is clearly warranted.

IV. METHANE SYNTHESIS Since the beginning of the century when Sabatier and Senderens first produced methane by passing hydrogen and carbon monoxide over reduced nickel and cobalt catalysts, there has been a sizeable amount of research on the synthesis of methane, motivated primarily by a need for suitable catalysts and processes to up-grade the low heating values of gases manufactured from coal. Reports of early investigations and processes can be found in the literature.4 Water-gas, comprised principally of carbon monoxide and hydrogen in approximately equal proportions, is one of several gases manufactured from coal. The heating value of water-gas is low, ranging from about 280 to 330 BTU per cubic foot. European countries, lacking natural gas, but with ample supplies of coal, have for years manufactured large quantities of water-gas by the steam-carbon reaction. The literature contains references to many early investigations on the synthesis of methane from water-gas as a method of increasing its heating value. 2225 More recently, the British Gas Research Board and the Fuel Research Board have conducted an extensive experimental program on the synthesis of methane from water-gas and other industrial gases using nickel catalysts as a means of enriching town gas26-31 gas. Kinetic studies of the synthesis of methane from both carbon monoxide32-35 and carbon dioxide-hydrogen36,37 mixtures on nickel catalysts have been made. The mechanism of the methanization reaction has been considered.38'39 There has been much research -35

-36on methanization catalysts,443 particularly with regard to supported nickel catalysts.44-56 While most of the catalyst development effort has been devoted to supported catalysts, some work has been done on non-porous metal catalysts. A comparison of a number of methanation catalysts will be presented later. A. Chemistry and Thermodynamics The predominant reaction in the synthesis of methane from carbon monoxide and hydrogen is usually represented by the equation CO + 3H2 2 CH4 + H20 (9) and is rapid in the presence of active catalysts at temperatures above 250~C. Catalysts used for the synthesis also promote the water-gas shift reaction. CO + H20 _- CO2 + H2 (10) which produces carbon dioxide by interaction of water vapor, formed according to equation (9), and unconverted carbon monoxide. Both reactions are reversible, and from known equilibrium constants it is possible to calculate the composition of the gas which would be produced if various mixtures of carbon monoxide and hydrogen were allowed to come to equilibrium at constant pressure and temperature. The results of such calculations30 are shown in Figure 3 to illustrate the effect of temperature, pressure, and H2/CO ratio of the feed gas on the composition of the product which satisfies both the methane and the water-gas equilibrium. Figure 3 shows that the formation of methane from carbon monoxide and hydrogen is favored by low temperatures and high pressures. This would be expected in view of the decrease in volume and the

-37H2/CO ratio of feed gas=3 H2 H2/CO ratio of feed gas =3 Pressure: I atmosphere Pressure: 25 atmospheres 60 - - 60 H HO~~~~ H2 o o 50 O 5 0 0 o4C o-040 -- H- Hz z Lw w 0 30 - 30 120 I 10 C i oCO rao to n fe gs > > CO 0 2 0 00 300 400 500 600 700 400 500 600 700 800 TEMPERATURE,~C TEMPERATURE, ~C LH22/CO ratio of feed gas I H2/CO ratio in feed gas I Pressure: I atmospherePressure: 25 atmospheres 60 60 50 4050,, a. /CO o0 0o a- / / ^\ Hz aw w 320 3 320 00 H 1o 00.0H20 -. H_-_- - 300 400 500 600 700 400 500 600 700 800 TEMPERATURE, ~C TEMPERATURE, ~C

-38evolution of heat which occurs in the methane reaction. As the H2/CO ratio of the feed gas is reduced, much larger percentages of carbon dioxide appear in the product gas, while the amount of unused hydrogen is reduced. During synthesis at high temperatures and low H2/CO ratios, free carbon can be deposited on the catalyst according to the equation 2CO ~ CO2 + C (11) [Other equations may be used to describe the appearance of carbon, but all of them can be obtained by various combinations of equations (9), (10) and (11).] For synthesis conditions at which carbon deposition would be expected, it is possible to determine the theoretical product gas composition which simultaneously satisfies the methane, the water-gas shift and the carbon equilibria43 provided equilibrium constants for equation (11) are known. Values for the carbon equilibrium constants depend upon the nature of the carbon formed, and it makes a considerable difference in the calculated equilibrium product composition whether the equilibrium form of the carbon is graphite, a more active metastable carbon, or perhaps a metal carbide. Equilibrium calculations predict only the composition of the product gas which is theoretically possible from the standpoint of thermodynamics. The composition of the product which actually would be formed under a given set of reaction conditions is dependent upon many factors: the rates of reaction on the catalyst surface, the temperature distribution through the catalyst bed, the rate of mass transfer to and from the catalyst surface; and must necessarily be determined by experiment.

-39B. Catalysts and Processes The objectives of research on methanation have been twofold: first, to develop processes which can be utilized in commercial size installations to upgrade product gases obtained from coal gasification to fuel gases with heating values of over 850 BTU/SCF; and secondly, to develop catalysts which not only are active for methanation, but which have process lives of such length that the cost of the catalyst is small compared to the total cost of the process. A number of problems which have been encountered in developing methanation processes and catalysts are as follows: 1. The reactions associated with the production of methane are extremely exothermic and methods for removing heat must be incorporated into the methanation process equipment. The excessive temperature rise over the catalyst bed which would occur if the heat of reaction were not adequately removed, could result in catalyst deactivation, in less favorable equilibrium conditions for the synthesis of methane, in excessive carbon deposition, and in extreme situations in damage to process equipment. 2. When using synthesis gas with ratios of hydrogen to carbon monoxide lower than three to one, carbon deposition can, and often does, occur in regions of temperature and pressure where high yields of methane are obtained. The conversion of carbon monoxide to carbon and carbon dioxide rather than to methane results in a loss in potential heating value since carbon dioxide, not being combustible, would be scrubbed from the product gas before use as a fuel gas.

-40Excessive carbon deposition can result in blockage of fixed-bed reactors, in loss of fluidization in fluidized-bed reactors and, in the case of porous-type catalysts, in fragmentation of the catalyst particles. 3. Catalysts for the conversion of carbon monoxide and hydrogen to methane are susceptible to poisoning by sulfur compounds present in the feed gas. If synthesis gas is not purified of sulfur compounds by processes which have been described earlier, the active life of methanation catalysts is greatly reduced and catalyst costs may become excessive. The development of processes and catalysts suitable for commercial methanation of carbon monoxide and hydrogen mixtures must necessarily deal successfully with these problems. In the following sections, developments in the field of methanation, catalysts and processes will be discussed to provide an indication as to the knowledge available on methanation and to the areas where further research would be profitable. 1. Catalysts A number of metals have been found to be catalytic for the conversion of carbon monoxide and hydrogen to methane. These include ruthenium, iridium, rhodium, nickel, cobalt, osmium, platinum, iron, molybdenum, palladium, and silver.41 Catalysts containing these active metals can be divided into two general classes: those which are porous in nature, and those which have non-porous structures. Porous type catalysts are generally of the supportedmetal type and are prepared by employing impregnation or precipitation

-41techniques on supports such as asbestos, alumina, kaolin, and kieselguhr. A typical method for impregnating a catalyst support with an active metal consists of adding the nitrate salt of the metal plus any catalyst promoters used to a quantity of support to form a slurry. The slurry is first dried and then heated to convert the nitrate salts to oxides. Precipitated catalysts can typically be prepared by treating the slurry of nitrate salts and support with a carbonate solution. The mixture is boiled with constant stirring and filtered. After washing with hot water, the filtered solid material is dried. Dried residue obtained from either technique is crushed, sized and reduced in hydrogen before use as a catalyst. Iron filings, nickel pellets, and steel shot are examples of non-porous catalysts. The catalytic area of such catalysts is limited to the external surface of the particles, since these materials lack the extensive internal pore structure characteristic of highly active supported catalysts. By virtue of the absence of pores, however, non-porous catalysts possess much higher mechanical strength than do porous catalysts and are not susceptible to fragmentation by carbon deposition. Extensive experimental programs to evaluate various porous and non-porous methanation catalysts have been conducted and reported by the United States Bureau of Mines, the British Gas Research Board, the British Fuel Research Board, and the Michigan Gas Association Laboratory at the University of Michigan. A number of nickel, cobalt, and iron-base catalysts have been investigated both in laboratory-scale fixed-bed and fluidized-bed reactors. The results of experiments in fixed-bed reactors have been tabulated and are shown in Table V.

-42co il ae a a t o m I M co I & Al 3 \ I II 4' ae! Sa t! \S & I2 o 8 2 d - d o@A d! +H H H, HH H 0X; -ti-( (ne mmcvi0'ro ti4i4ixmtm or( o-t mHO K u. d. T-Io~~~~~~~~~ I II Cl\. * LIN m i. -- L Lr C\ m MH 000: \ c \oII-I OV OLd en C \ C \ fr jC Ie^~~~~~C - ~C H C~ 4,4..I.3..3 3 0 64 d d d d d oo o Q o888888 8 888888888888 88888888" V 8f co f7 M M MM0000000-t000000t -CM _ -t n 000M 0 l " P 888COr8 888 88 8 8 8 C' I,fi I HH ^ YHHH?eH II I 11 8 8 88 888H' 8888888888888888888888 c888 / co.a| myi OHI IH mi m I i \ L m Ci I i I V\ OcO -E o, o r\ U \ =n O QI (4 u o< c Hfi^ i ^oc v~ r NjcMcn oo otc O N fi <a 28 0 Jl) 0 0 Lt\ aTOO A, 04 0.. O I \' t.,:., -, o x\'-'U >. 0~~,~.~.~.~.~.~.~. o,~.~o,~,~ o~d,~~,..o.U..o V 0 S H r-L u f\ H ^ 1 0 H -t ^-t O tt H t/ Or lO t OQ O \ O 0l O: O Ni H 1' ~ L0 0 ~ n O n NO, C t O O o m W 0 * N I' \o U S'du o |t8*.(U ff -' ~4o r4 Ce jU Ce t; | 0; r; t d t15 ] 0;; co. 4', e 8~~S!f 0 —' 4 d dd ~ 3 co 43 0 hd4' fl ^ ^ < t Od -- a Nd0e O R~ x 0 o 0 H I Id ~F * o l -d ~Id O V dtt ~ d.4 W I d t g:o o I d 9d pd k t eO I *j o * PI 4 oo- - -5e l rJI g8^ ^ * 8 Id 2 +3 O d 4) 4 S 3.3 ~ t o r~)3~ I E 5 0a s1 O a 8 6 a^11 S4 * Nt %0 N 4O O l O 0'-l.mE''d d EH 4n n0np Ou a Ac dtd >;-' < < 8 -^s'sr a 3a n t S^\a 1 > ^t~t AS " s Sg^l8 [| HSH ~ ~ ~ ~ ~ B r ~ " aId coa 4 " 4 -c s 8' 0 40a H S... ^.cHA AAu a...S o0 I Id Id & r ~l 0 ~OCI ifl-lacu M ~~ 8rCIP O fl 1 0ffI ff1 - H ir 0 i j d 0PM *+ 4CJ 4 -3 CM P * 0 *s 0 C + 4 o V HH c- 4Se *'t- r 0 P r0a a0~t a" a) w a ~ M^KpS- ^&<p ^ d t c\ CO Z d t0 0. r __A __a l4 - -A. P It,9 E-l CQ a I^ f iK^ ^ rlri-n6r^Z E -.i( <c c lC' u ^rirl11Hrlrlr(H ^M*H (; i4-t~~:* ~ z:Z u> ^ > > 0 ~ ~ ~ ~ ~ s ^acO O UU OB) K^ o Sz;s lU 67ill) " IJ UU P *l s r 3 (^C\J ~ in m ~ ~4 ~ 48^ 0. "6^ ^ O r ap ^ ufPpppppp cr u'.^ -~-'-Titib~fi<i Tt(df 0b U U x (<( ( OCOC O OC C O ( m Hn l^ra- l^ H a. fc ~ r K ~ 4d ~, 5 aft f ( 3 3 ^~j,9=11; MI ~ "HuB ~ ~ brU r

-43(a) Iron Catalysts Table V shows that iron catalysts, in general, are less active for methane synthesis than nickel catalysts. In fact, with iron catalysts it is difficult to obtain product gases with heating values over 800 BTU/SCF. Most porous iron catalysts produce large quantities of C3 + hydrocarbons in addition to methane and ethane over the range of temperatures where they are most active for converting carbon monoxide and hydrogen to products. Hydrocarbons containing more than three carbon atoms tend to condense in highpressure pipe lines and for this reason do not constitute suitable components for inclusion in high BTU gas streams. Because of the conversion of synthesis gas to liquid products and the resulting reduction in the yield of gaseous products of high heating value, porous iron catalysts in general are not suitable for methane synthesis at temperatures below about 800~F. At' temperatures above 800~F, thermodynamic equilibrium conditions become unfavorable for the formation of liquid hydrocarbons. Synthesis gas is converted primarily to methane and, to a much lesser extent, to ethane. While higher temperatures yield a more favorable distribution of hydrocarbon products for high BTU gas manufacture, two other factors combine to make porous iron catalysts generally unsatisfactory for methanation at these temperatures. Carbon formation becomes excessive, and its deposition in the pore structure of porous catalysts can eventually cause fragmentation of the catalyst pellets. In addition, supported type porous catalysts are often susceptible to sintering at temperatures above 800oF which can reduce catalyst surface area and decrease catalyst activity.

-44Non-porous iron catalysts are not subject to fragmentation by carbon deposition nor to loss of activity by sintering, and thus, would appear to be superior to porous iron catalysts for methanation. Smith57 and Schatz58 have rather thoroughly studied the hydrogenation of carbon monoxide on 1/8-inch diameter S.A.E. 1013 steel balls and have found that the catalyst exhibits a large increase in its ability to catalyze the reaction between carbon monoxide and hydrogen to form methane, water and carbon dioxide as the reaction temperature is raised from 900~ to 1040~F. The catalyst shows a maximum activity at 1100~F to 11500F, depending upon synthesis gas composition, and then declines with increase in temperature. The effect of temperature and synthesis gas composition on the gross heating value of the product gas is shown in Figure 4. This figure indicates that the optimum temperature and feed gas composition for the synthesis of a product of the highest heating value at 31 atmospheres pressure and a space velocity of 0.10 mols/hr/sq.ft. fall into the following ranges: Temperature: 1090~ - 1150~F Percent CO in feed gas: 47% - 52% (remainder H2) For these conditions, the heating value of the product gas, scrubbed free of carbon dioxide and water, ranges from 735 to 745 BTU/SCF. Carbon deposition on the steel catalyst increases rapidly as the reaction temperature is raised from 900~ to 10000F, but decreases substantially at higher temperatures. This trend is clearly shown by Figure 5, in which the effect of temperature and synthesis gas composition on the amount of carbon deposited during synthesis is presented. Carbon deposition, though decreasing with

-45Pressure: 31 atmospheres 800- Feed rate: 5 SCFH Space velocity: 0.10 8 0. 16 mols/hr-sq ft 50% CO in feed.gas I700 1050 14100 1150 1200 OF 40%CO W S L C LYS ELI CO )35%CO. 600 )0% CO 500 1050 1 100' 1150 1200 TEMERATURE OF FIGURE 4. EFFECT OF TEMPERATURE AND FEED GAS COMPOSITION ON THE GROSS HEATING VALUE OF THE SCRUBBED PRODUCT GAS FROM METHANATION WITH A STEEL CATALYST

-46(/) oc~~~~O (0 1 zz 0 o._ Q I0 ( >.- ~ ~- 00 E 0 C c_ o 0 EE:2 0 0 0 C:o: II C.) o < U + E0 0 -.^-e ^| c< < CL 4 - ~ 1h 0^ >^'o <*-? 0 J / CZ..... 0' w 0~ h 0 I& Ir'U) -- ^- U M 0L I.)0. 0 U) cB\ -- z ">0 0 o o LL ___0 =ro, d -.i U) uw W=I L ~ ~ ~ ~ ~ ~ L - IC) 5I 5'' _____ __' a. LL 0 NO8'VOD 33]J Sv 0311sOd30 a33d 3HI JO B0O +00 NI 0 JO IN3B3id IH913M

-47with increasing temperature, is appreciable at temperatures where maximum conversions of carbon monoxide to methane are obtained. However, carbon can be removed from the catalyst by air oxidation without deleterious effect on catalyst activity. It is significant to note from Figure 4 that the heating value of the product gas reaches a maximum when the synthesis gas used contains approximately 50% carbon monoxide. At higher carbon monoxide concentrations in the synthesis gas, a decrease in product heating value results from a decreased conversion of carbon monoxide to methane and ethane and an increased conversion to carbon dioxide. It is apparent that with the non-porous steel catalyst studied, an upper limit on the heating value of the product exists which is below the 850 BTU/SCF value necessary for economical high BTU gas manufacture. On the basis of thermodynamic equilibrium considerations, it has been pointed out previously that low temperatures and high pressures are most favorable for the conversion of carbon monoxide and hydrogen to methane. Unfortunately, most non-porous steel catalysts are not active for methane synthesis at temperatures much below 1000~F. Thus, while steel possesses a number of attributes that make it an attractive methanation catalyst, it is necessary to use reaction temperatures that are not the most favorable for methane synthesis. As a result, it is not possible to approach complete conversion of synthesis gas to methane. Even at optimum reaction conditions for non-porous steel catalysts, appreciable quantities of carbon monoxide and hydrogen remain unreacted, thus, producing product gases with heating values lower than those necessary for

-48economic high BTU gas manufacture. Steel catalysts may be attractive in Europe, however, where the required heating values for commercial gases fall in the range of 500 to 600 BTU/SCF. (b) Nickel and Cobalt Catalysts Supported (porous) nickel and cobalt base catalysts, are active for methane synthesis at temperatures at which almost complete conversion of carbon monoxide and hydrogen to methane is thermodynamically possible. While such porous type catalysts are subject to sintering and to fragmentation by carbon deposition, making them less desirable from a structural standpoint, potentially they are capable of producing a product of higher heating value than is obtainable with either porous or non-porous iron catalysts. Table V shows that nickel catalysts formed by precipitation techniques are more active than varieties of impregnated nickel catalysts. There is little difference in activity, however, between precipitated nickel catalyst and the extracted Raney nickel catalyst. Both are capable of producing product gases with heating values in excess of 900 BTU/SCF. Likewise, cobalt-thorium oxides supported on kieselguhr and also extracted Raney cobalt alloys are active methanation catalysts which can produce gases with heating values above 900 BTU/SCF. The composition of the product gas which can be synthesized from a mixture of 29% CO and 71% H2 using a cobalt-thoria catalyst, is shown in Figure 6 as a function of synthesis temperature. At temperatures ranging from 650~ to 750~F carbon monoxide is almost completely converted to products. Some hydrogen remains unreacted, but a more judicious selection of synthesis gas composition could

-49Pressure: 30 atmospheres Feed rate: 5 SCFH Feed composition: 29% CO- 71% H2 Catalyst: cobalt- thoria (200 pellets- dia. x 8 long) 50 6H4 40 0 U0 Z 3O LU I 0 20 H2 10 TEMPERATURE OF FIGURE 6. EFFECT OF TEMPERATURE ON THE PRODUCT GAS COMPOSITION USING A COBALT- THORIA CATALYST.

-50have resulted in virtually complete conversion to methane, water, and carbon dioxide. After removal of C02 and H20, the product stream would be close to 100% methane and have a heating value in excess of 1000 BTU/SCF. (c) Carbon Deposition Deposition of carbon during synthesis is generally considered to be detrimental to efficient methanation. Conversion of carbon monoxide to carbon constitutes a loss of potential heating value since yields of methane are reduced. Carbon deposition can fragment porous catalysts, plug fixed-bed reactors, or cause loss of fluidization in fluid-bed reactors. Nevertheless, because of the nature of the catalyst used or the composition of synthesis gas available, carbon deposition often must be tolerated even though it may be the most significant factor reducing catalyst life. Carbon deposition is encountered with all iron catalysts, and becomes excessive as the H2 to CO ratio of the synthesis gas is reduced below three. In contrast, very little carbon is deposited on nickel catalysts when H2 to CO ratios of 2.5 or higher are used. At lower H2 to CO ratios, e.g., a ratio of 1.0, excessive carbon deposition is encountered with all of the active catalysts listed in Table V. Carbon can be removed from some catalysts, notably nonporous iron catalysts, without deleterious effect on catalyst activity by passing air over the hot catalyst to convert the deposited carbon to CO and C02. With porous catalysts, this technique is not always successful for the high surface temperatures obtained during oxidation may cause sintering of the catalyst and subsequent reduction in catalytic activity.

-51(d) Fluid-Bed Catalysts Catalysts which are found to be active for methanation of CO and H2 in fixed-bed testing units are also active when used as fluidized catalysts. However, from a structural standpoint they do not always make satisfactory fluid catalysts. Catalysts must be physically stronger for service in fluid-bed reactors where maintenance of a constant particle size distribution is necessary for good fluidization. Data on a number of catalysts which have been tested in laboratory scale fluid-bed units are tabulated in Table VI. From an activity standpoint, fluid-bed catalysts are similar to fixed-bed catalysts. Iron catalysts are less active than nickel catalysts and are unable to produce high BTU fuel gas. High yields of C3+ hydrocarbons are obtained at temperatures below 800~F. At higher temperatures, equilibrium conditions become less favorable for the formation of methane. Carbon deposition, which can fragment porous catalysts and can ultimately result in plugging of fixed-bed reactors, is even a more serious problem in fluid catalyst operations. Changes in the size distribution of a charge of fluid catalyst because of particle fragmentation or carbon buildup can result in loss of fluidization at total carbon levels much lower than those which cause trouble in fixed-bed operations. Carbon deposition on the fluid iron catalysts described in Table VI is excessive even at synthesis gas CO to H2 ratios as high as three over the range of temperatures covered. Thus, low catalytic activity, high yields

-52co M M MMo o 0 \O \ QLf\ 0000000 A \D t~- Lm L\\O OV \ WI ( R J <oQ mrQ r~ A A A A A t VoI o I, n H I o iU # o 000 U.i o o UN o 8.oo 88 8% 88o co.4 00 10 0 o00 888008 *^ H 00 0 f 00 Q Q 0 Ef~c - O O OD M O O L 0 M M o itC pi C HO H +U To l.. I 0 _o o C lrs 0 e Ui -4 4( ooohq oOQ000 0 0000000 00 0 ooo00000 8fl a ~-r-4 0: 0 0. o o ~J, rX(, r- (n z (g H M:t co r-4 t — tl- a- t-d t H 03~ 4 0 0 1 I E-53> 6 0 0 0 0 0 0 0 C3j IO t< 00 tD CO l o o ~ oo0 o o ooo i| 1 p IO O O OJl~~hr3i~t i (D (' { ('0 (,' (' ( N 0 co N 0 t ) $ U - t \ t \ t z a. -. 4.)C\J 0 L- - 4 0 o0 a~(! I(~ H,;'P i-I 0... do ~: *8 4)... UNI II-i -t /j -.,t - -t Q, U 0 I o8'\ d o( lco O + 0:3 4) l\ 4 I < co -,0o o o"o o 0 r{ - )d * i 00000 of to. -.-... 0, ms) m~ ol\ in m 8 co d lao r0202o 0 2 POOOOO~oCdCCO US5BBKS5 S25S

-53of C3+ hydrocarbons, and excessive carbon deposition combine to make iron catalysts generally unsatisfactory for methanation of synthesis gas. Nickel base fluid catalysts, on the other hand, are highly active for methanation. The nickel catalysts listed in Table VI produce methane yields high enough to give gases having heating values of over 900 BTU/SCF. Catalyst process lives as high as 6500 pounds of methane per pound of nickel have been obtained. A typical product gas composition is shown in Table VII. TABLE VII. ANALYSIS OF A TYPICAL PRODUCT GAS PRODUCED BY METHANATION WITH A NICKEL CATALYST43 Catalyst Composition: NiO-A1203-Ka olin (100: 159:97) Average reactor temperature 579 F Reactor Pressure, psig 300 Fresh-feed space velocity, V/V/hr. 5530 Recycle ratio 1:1 H2/CO feed ratio 2.5 Catalyst age, hrs. 65 Product Composition, volume % CH4 81.8 C2+ 0.0 C02 14.2 H2 4.8 CO 1.2 Gross heating value, BTU/SCF 960 Carbon deposition has not proven to be a problem with nickel fluid catalyst as long as feed gas H2/CO ratios are maintained above 2.5.

(e) Summary The development of methanation catalysts has progressed to the stage where active catalysts are now available for methanation on a commercial scale in either fixed or fluidized bed processes. It has generally been established that nickel is superior to iron as a methanation catalyst for processes designed to produce only high heating value fuel gas. However, it is conceivable that in a combined operation to produce both gaseous and liquid products, iron might prove to be a more suitable catalytic material. It is doubtless that further research would result in improved methanation catalysts. However, in view of the present advanced state of catalyst development, and recognizing the need for intensive research effort in the fields of solid fuels gasification and gas purification, further research on methanation catalysts can best be deferred in favor of work in fields of more immediate need. 2. Processes The formation of methane from CO and H2 is a highly exothermic reaction, and the heat of reaction must be quickly and effectively removed in order to prevent excessively high reaction temperatures which could result in catalyst deactivation, unfavorable equilibrium conditions, carbon deposition or in da.mage to process equipment. Two primary means are available for temperature control: 1) Cooling by boiling a liquid (usually water) contained either in tubes incorporated into the methanation reactor or in a jacket which surrounds tubes containing catalyst. Equipment of this sort has been used in the Fischer-Tropsch process for synthesis of liquid hydro59 carbons.

-552) Recycling a large quantity of cooled product gas. Since both methods are applicable to either fixed or fluidized-bed operation, four general types of methanation operations are possible. To summarize, these include: (a) Use of fixed-beds of catalyst in which reaction heat is removed by boiling a liquid. (b) Use of fixed-beds of catalyst in which reaction heat is removed by recycling cooled product gas. (c) Use of fluid-beds of catalyst in which reaction heat is removed by a boiling liquid contained in tubes inmmersed in the catalyst bed. (d) Use of fluid-beds of catalyst in which reaction heat is removed by recycling cooled product gas. Control of reactor temperature by recycle of cooled product gas has the added advantage of reducing or preventing carbon deposition in the catalyst bed. The presence of carbon dioxide in the combined feed to a methanation catalyst serves to displace the equilibrium between carbon monoxide, carbon dioxide, and solid carbon, described by the equation 2CO = C02 + C (11) so that conversion of carbon monoxide to carbon is reduced. The heat of reaction of the methane synthesis is so great, however, that internal cooling of catalyst beds with boiling liquids may be required in addition to recycle of cooled product gas in order to maintain desired temperature levels. Fixed-bed reactor systems have received the major attention

-56for use in methanation processes primarily because they have proven to be ideal units for both laboratory and pilot plant scale investigations. Fixed-bed reactors have the advantage over fluid-bed reactors that for a given diameter vessel, higher linear gas velocities are possible, allowing higher synthesis gas throughputs. On the other hand, fluid catalyst units have more uniform catalyst temperatures and improved heat transfer to cooling surfaces. Circulation of catalyst to a regenerative vessel in which carbon could be burned from the catalyst, is possible in a fluid catalyst system. Both fixed and fluidized-bed pilot plant units have been described in the literature.43 Operation of the methanation step at elevated pressure is desirable both from a thermodynamic and an engineering point of view. Pressure operation allows higher flow rates and reduces the size of equipment. Elevated pressures favor the formation of methane as shown earlier in Figure 3. Furthermore, methanation at high pressure reduces the amount of work required to compress the final high BTU gas to pipe line pressures. Operation of the gasification, methanation, and purification steps above pipe line pressures (800 to 1200 psig), would eliminate the need for a final compression step, and is one goal toward which research in these fields should be directed. C. Conclusions Research on both methanation processes and catalysts has reached an advanced stage. Methanation has been demonstrated on a pilot plant scale, and no particular problems are expected in building

-57and operating commercial size units as long as the synthesis gas is purified of sulfur compounds and adequate means of removing reaction heat are incorporated into methanation reactors. To be sure, further research will be helpful in arriving at the optimum reactor design and the most effective methanation catalyst. However, this research is not urgent, and can best be postponed until the economics for high BTU gas manufacture become favorable. Until that time, research effort is more urgently needed in the field of gasification and gas purification, where the economic incentives for developing improved processes are greater.

V. ECONOMICS OF HIGH BTU GAS MANUFACTURE FROM COAL Previous sections of this report have dealt wi'th the process steps required to manufacture high BTU gas from coal. Specific processes have been cited which have been developed, some only on the pilot plant scale, to accomplish the various steps which constitute the integrated process of making high BTU gas. These are of obvious importance, for a workable process is a fundamental requisite. Nevertheless, the ability to produce high BTU gas is not the ultimate story. Unless the price of the product, delivered to the consumer, can compete with the price of natural gas and of other forms of energy, coal will not be used as a raw material for high BTU gas production. It is the purpose of this section to present economics for the overall process of making high BTU gas from coal in the light of present day costs. A. Basis for Economics In 1951, Blaw-Knox Construction Company prepared an economic 60 study of the production of methane from coal and this work has served as the basis for the economics presented herein. Investment and production costs have been escalated from 1951 to January 1957 levels. A simplified flow plan for a typical plant to produce 100 million SCF/day of high BTU gas is shown in Figure 7. Non-caking coal is reacted with oxygen and superheated steam in Lurgi pressure gasifiers. The gas produced is completely purified of sulfur compounds in an initial purification step employing the Rectisol process. Simultaneously, the CO2 content of the gas is reduced. Following purification, the gas is passed over a -59

-6ow -J LiJ C/)ULJ < C/U~ C/) < C/. XI (, 1' C0 0 ~ 4 0 0I) ------- a) 4 c/ o IL< I —L0 U 0 w m 8-~~~~~ 0 z IU a ~~~~~~~~~~~C/)~~ ~ ~ ~ ~ ~ (/U 4w(0~~~~~~ 0 zis c~ ~ro z, I<~ ~ J C 3) 4 0E (1 H I aZ0a. " — tI L a Z I Z O V)~a V) V)~~~~~~~~~~~~~~~~~~~1 o o~~~~~~~~3 z b.I~ ~ ~~~~~V > - 3 ooW -r Q I~~~ O Q ~~~L

-61supported nickel methanation catalyst which converts the synthesis gas to CH4, C02 and H20. A final purification of the methanation product with the Rectisol process removes C02 and H20. The resulting gas stream has a heating value in excess of 900 BTU/SCF, and is ready for compression to transmission pipe line pressure. In the Blaw-Knox study, fuel consumption is based upon a hypothetical non-caking coal having a heating value of 13,000 BTU/lb. which would not produce tars or oils in the gasifiers. To produce 100 million SCF/day of high BTU gas, daily consumption of fuel would amount to 4,500 tons. Oxygen requirements would be 2,070 tons of 95% oxygen supplied at 375 psig. In addition, 13,380 tons/day of 375 psig, 850~F steam would be consumed. Representative compositions of the gas stream at various points in the overall process to convert coal to high BTU gas as outlined by the flow diagram in Figure 7, are shown in Table VIII. B. Investment Costs The investment costs for the process units and the utilities involved in the production of high BTU gas are presented in Table IX. Investments have been escalated from the 1951 level to January 1957, using a 1.39 escalation factor. These costs include the design and procurement of equipment, and also the labor and overhead charges incurred during construction. The designation "auxiliaries" include site preparation, grading, storage tanks, loading facilities, roads, rail sidings, administration buildings, warehouses, shops, etc. Interest charged on the capital required during the construction period has been added to these costs. Working capital has also been

-62TABLE VIII. GAS COMPOSITIONS AT VARIOUS POINTS IN A HIGH BTU GAS PLANT Gas Leaving Gas Raw Lurgi Initial Leaving High BTU Gas Purification Methanization Gas H2 40 54.8 ) ) ) 8.6 ) 10.9 CO 20 27.4 ) ) COo 30 4.1 22.1 1.0 CHR, 8 10.9 63.6 80.9 C2H6 & ill. 1 1.4 2.85 3.5 N2 1 1.4 2.85 3.7 100 100.0 100.0 100.0 Volumes 100 73 35.0 27.5 Heating Value (BTU/SCF) 293 401 721 915 I —-'i I _~~~ ~~ ~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~

-63TABLE IX. INVESTMENT COST SUMMARY Process Units Labor, Material & Overhead Gasification $ 8,150,000 Initial Purification 5,160,000 Methanation 3,560,000 Final Purification 1,510,000 Final Gas Compression 1,070,000 Auxiliaries 9,150,000 Utilities High Pressure Steam 10,520,000 Low Pressure Steam 1,320,000 Electric Power 2,700,000 Cooling Water 4,070,000 Boiler Feed Water 3,280,000 Oxygen Production 14,380,000 Oxygen Compression 1,270,000 Total Construction Cost $66,140,000 Interest on Capital During Construction (assumed) 3,480,000 Working Capital (assumed) 3,320,000 TOTAL INVESTMENT $72,940,000

-64provided. Investment costs total $73,000,000 and have increased about $20,000,000 since 1950. C. Production Costs A summary of costs involved in producing 100 million SCF/day of high BTU gas is presented in Table X. The costs shown are based upon coal at $4.00 per ton. The sale of 100 million SCF/day of high BTU gas involves a public utility which would be limited by law to a maximum return on capital investment of 6%. Depreciation costs average about 6% on capital investment. An additional 2.5% on invested capital has been included in the cost of producing high BTU gas to cover items such as real estate and personal property taxes, annuities, insurance, etc. Maintenance and repair costs have been assessed on the basis of a fixed percentage of the capital investment. Two percent has been used for electric power generation, cooling water supply, oxygen production and compression, and final gas compression. For all other units, maintenance and repair costs have been assumed to be 4%. Direct operating labor costs have been escalated to account for current average wage levels. Overhead and administrative expenses have been taken as 60% of the direct labor costs. The total cost of producing high BTU gas as shown in Table X amounts to 61.5 cents/MSCF, based on coal at $4.00 per ton. Inspection of Table X indicates that gasification represents 70% of the total cost of producing high BTU gas. The cost of initial purification is 11.5% of the total, methanation is 4%, final purification is 2.5%, final gas compression to 1200 psig is 2%, and auxiliaries

-65IH oo og 0 O o:ts CO 0\ 4 ri cX., H'oc c t s vo H i ^ H n,H a 0HHC t'N'0 6 0 O IM O\ Ch OR Q 0 CY UN IO * F-4 Ho 4 H (0 H O | 00 | I H m 0 0 o!! 1! 6 I t It o | 0: i o 66 olo I 6 o 6 |6 6 o o 66 6 i E ~ CQC o B ~ 6 H6|r!o]l 6 6 0:I 6,, ooo 0 6H6 6 c ~3 \9 8 CY 00 r-< 0 H o E-4 M - 0 H o g t. N g 1,,.,!. ( Co M I! I M CY m CY COI, 8 o o -o o 4 0 ^ 5^ O^ )1It O OO 0 0 0 2 O* * 3too0 0 >4 k i 4 4* i d* * d j 8 a t5 h U *0 0410 1 m H t 8 H 0 0 O a

-66amount to 10%. Production of the purified synthesis gas accounts for fully 84% of the total cost of producing high BTU gas (when coal is charged at $4.00 per ton) which clearly points out the areas where the greatest economies can be made, namely, in the processes of mining coal (lower fuel costs), gasification of coal, and purification of synthesis gas. D. Effect of Fuel Cost The effect of variation in fuel cost on the production cost of high BTU gas is shown in Figure 8. This figure indicates that the cost of high BTU gas increases linearly with the cost of coal. The 1951 Blaw-Knox study was based on coal priced at $4.00 per ton, and showed that high BTU gas could be produced for 55.4 cents/MSCF. The corresponding figure for January 1957 production costs is 61.5 cents/MSCF, an increase of 6.1 cents/MSCF over 1951 costs. This increase is the result of much higher investment costs, slightly higher catalyst cost, and increased wage rates. Since 1951, the wholesale price of bituminous coal has increased only very slightly. It is reasonable to assume that coal which cost $4.00 per ton in 1950, is available for approximately the same price today. Thus, comparison of the revised production costs for high BTU gas shown in Table IX with the 1951 Blaw-Knox study is a valid comparison, which shows that with coal at $4.00 per ton the true increase in the cost of high BTU gas over 1951 costs amounts to about 6.1 cents/MSCF. E. Comparitive Cost of Natural Gas and High BTU Gas In the interval 1951 to 1957, the average wholesale price of natural gas increased by 13%.61 During this same interval, the

-67100 90 80 70 REVISED COSTS,, (JAN 1957) 60 eo 50,, 50 i / ^ —- BLAW-KNOX STUDY COSTS (1951) m 401-^,, 30 0 0 20 ~ 20 10 0 1 2 3 4 5 6 7 8 9 COST OF COAL, S/TON FIGURE 8. EFFECT OF FUEL COST ON THE COST OF PRODUCING HIGH BTU GAS.

-68cost of producing high BTU gas has risen about 11% from 55.4 to 61.5 cents/MSCF. These statistics point out that while the price of natural gas has risen considerably over the last six years, the cost of producing high BTU gas has increased at a nearly comparable rate. The sizeable gap between the price of natural gas and that of high BTU gas produced from coal has not been narrowed an appreciable extent during this period. F. Summary It has been estimated, that the investment for a plant designed to produce 100 million SCF/day of high BTU gas would be approximately $73,000,000. The cost of gas production has been estimated to be 61.5 cents/MSCF when charging coal at $4.00 per ton. The cost of producing high BTU gas is closely related to the cost of coal. For example, with coal costing $4.00 per ton, fuel represents about 40% of the cost of high BTU gas, while at $8.00 per ton, fuel is roughly 55% of the gas cost. By far the greater portion of the total cost of producing high BTU gas lies in the production of the purified synthesis gas. Therefore, efforts to reduce costs should be directed toward reducing the cost of the synthesis gas. The escalation of the 1951 Blaw-Knox economics to present cost levels, shows that production costs of high BTU gas are rising at a rate which is nearly equal to the rate of increase of natural gas prices. This fact raises a very valid question as to whether high BTU gas manufactured from coal will become competitive with natural gas before the day when large supplies of inexpensive nuclear energy are available for general commercial and household use.

-69The economics presented in this section are based on an overall process as outlined in Figure 7. Selection of other subprocesses for gasification methanation purification, etc., than those used in this study could change these economics to some extent. However, the numbers presented herein are believed to be typical of the cost of producing high BTU gas using the most advanced of the commercially proven processes available.

VI. CONCLUSIONS Before the day when natural gas reserves near depletion, a replacement must be ready. High BTU gas manufactured from coal can serve as that replacement, and much research has been devoted to developing an efficient and economical process. Coal is a logical choice as the basic raw material from which to manufacture high BTU gas because of the ample reserves of coal which lie within easy reach of the highly industrialized and densely populated areas of the United States now supplied with natural gas. Present engineering know-how is adequate to design a workable plant to produce high BTU gas from non-caking coal. Such a plant would be comprised of five sections wherein 1) coal is gasified with steam and oxygen to produce a mixture of carbon monoxide, hydrogen, and carbon dioxide; 2) the raw synthesis gas is purified of carbon dioxide plus sulfur compounds which poison methanation catalysts; 3) the mixture of carbon monoxide and hydrogen is catalytically converted to methane, water and carbon dioxide using a supported nickel catalyst; 4) carbon dioxide and water are removed from the gas produced; and 5) the product high BTU gas is compressed to pipe line pressure. Possible alternative subprocesses for accomplishing these operations have been discussed in previous sections of the report. Research devoted to the development of processes to gasify American caking coals has been noted. Areas where additional research may lead to improved processing and to large economies in the production of high BTU gas, have been pointed out. In general, these lie in the manufacture of the purified synthesis gas. -71

-72The development of methanation catalysts and processes has reached an advanced state. Further research on methanation can best be deferred in favor of work on gasification and purification processes, where a greater return on research effort can be realized. Estimated economics for the production of high BTU gas from noncaking coal have been presented and indicate that with coal priced at $4.00 per ton, production costs would amount to 61.5 cents/MSCF of gas. Production costs have increased by 11% since 1951. Moreover, the total capital investment required to build a high BTU gas plant has risen by roughly $20,000,000. During the period 1951 to 1957, the average wholesale price of natural gas increased by 13%. This rate of increase has been only slightly more rapid than the increase in the estimated production cost of high BTU gas. This fact clearly points out that the gap between the cost of natural gas and high BTU gas manufactured from coal, has not been closing at a very rapid rate. The day when manufactured high BTU gas will be competitive with natural gas appears to be still quite far in the future. The potential of obtaining tremendous quantities of inexpensive useful atomic energy from fission or fusion cannot be overlooked. The advent of atomic power complicates the problem of predicting not only the future of high BTU gas from coal, but of all our present supplies of energy as well. However, no matter how uncertain the future may be, the possibility of utilizing gasification and methanation to produce and upgrade low BTU gas cannot be discounted. While at present, the construction and operation of a large plant to produce high BTU gas from coal would be prohibitively expensive, plans for the future should be made. There is a need for

-73much pilot plant scale work to improve and reduce the cost of presently proposed processes, and further, to answer many questions concerning the optimum design of a commercial size plant. When a shortage of natural gas becomes imminent and a replacement is needed, the design of a plant to manufacture high BTU gas from coal should be ready.

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