Production of Hydrogen and Synthesis Gas by Oxygen Gasification of

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592

INDUSTRIAL AND ENGINEERING CHEMISTRY

engineering standards i f a service life of several years is desired. At no time during the operation of the experimental equipment did it appear that the rate of heat transfer through the tube walls was the limiting factor which determined the rate of gasification. In order to permit operation a t a considerably higher and more favorable temperature level, it would appear possible to employ silicon carbide refractories for construction of the walls of the reaction chamber. This material has lower thermal conductivity than the alloy, but it is possible t h a t sufficient heat could be transferred a t a higher temperature, which the refractory could withstand, to give a significant increase in capacity. It appears possible by suitable modifications to the equipment to operate the alloy tubes under a pressure of several atmospheres, which would act to increase the residence time of the particles in the tube, and thus increase the capacity of a tube of given dimensions. The gasification characteristics of the pulverized coal at elevated pressures would need to be determined experimentally. If the chemical activity of the fuel is determined by the rate at which steam can be brought to the fuel surface for reaction, i t seems possible that the apparent activity might be increased t o imparting a relative motion t o the fuel particles with reference by the ambient gases. In the present apparatus, the particles

Vol. 40, No. 4

are carried through the tube by the flow of gases. If gasification were carried out in a cylindrical chamber with the steam admitted by jets directed tangentially around the periphery, the circular motion imparted to the particles would subject them to a centrifugal force which would tend to throw them outward toward the cylinder wall against the flow of gas and steam toward a central outlet. This could conceivably result in a marked increase in the rate of reaction. Acknowledgment

Acknowledgment is due F. E. Graves, formerly research engineer a t Battelle, for help and suggestions during some of the construction and testing phases of this work and to the Gasification Committee of Bituminous Coal Research, Inc., Eugene J. Kerr, chairman, for interest and helpful suggestions. Literature Cited (1) Barnes, C. A., Tech. Rept. V, Bituminous Coal Research, Inc., Washington, D. C., June 1939. (2) Institute of Gas Technology, “Gas Making Processes,” American Gas Association, New York, October 1945. RECEIVED October 13, 1947

Production of hydrogen and synthesis gas By the Oxygen Gasification of Solid Fuel C. C. WRIGHT AND K. M. BARCLAY The Pennsylvania State College, State College, Pa.

R. F. MITCHELL The Consolidated Mining and Smelting Company of Canada, Ltd., Trail, B. C.

TEST data on the gasification of rice and barley sizes of anthracite and of lump coke in a commercial producer gas plant, slightly modified to blast the bed continuously with oxygen-steam mixtures instead of the conventional airsteam mixtures, are presented. Data for plant scale tests on the catalytic conversion of the excess carbon monoxide to hydrogen form the basis for heat and material balances calculated for the production of “synthesis gas” having a hydrogen-carbon monoxide volume ratio of 2, and for commercial hydrogen suitable for ammonia synthesis. The relative efficiency of synthesis gas production by this and other commercial solid fuel gasification processes is discussed.

P

RODUCTION of synthesis gas for the Fischer-Tropsch and related syntheses has been the subject of considerable interest in recent years, and several excellent reviews have dealt with foreign and domestic developments utilizing gaseous hydrocarbons (7, l a ) and solid fuels (1, 4, 9, 11, 18, 16, 1 7 ) . Of the processes using solid fuels, current interest in America has centered largely on foreign methods for the gasification of pulverized fuel in fluid beds, or in suspension, because of the greater ease of handling, generally lower cost of fine sizes, and higher gasification rates attainable. There are, however, certain less desirable

features of the pulverized fuel processes, such as the poor carbon efficiency, the relatively poor quality of the raw gas, and the problem of removing the fines carried in the gas stream which up t o the present have not been entirely overcome. Whether the extensive research now in progress has or will overcome all these difficulties cannot be stated, but no such process has yet been successfully demonstrated commercially. Although less publicized in the technical literature, work has progressed on several fixed-bed processes utilizing oxygen for continuous gasification (5, 6, 8, 9, 10) and commercial plants have operated successfully. As early as 1937, Stewart (16) reported on the possibilities of. using oxygen for the gasification of solid fuels in a fixed-grate water gas machine. More recently Mitchell (8)and Wright and Newman (18)have presented reports on the commercial development of the steam-oxygen blast fixedbed process at the synthetic ammonia plant of the Consolidated Mining and Smelting Company of Canada, Ltd., at Trail, British Columbia. Although a t this particular plant the conversion is carried completely t o hydrogen for the synthetic ammonia process, the essential features of the plant and method of operation would be similar were the gas being produced for the FischerTropsch or related synthesis. This paper presents additional information regarding the production of hydrogen or synthesis gas by this process.

INDUSTRIAL AND ENGINEERING CHEMISTRY

April 1948

593

Equipment and Test Procedure

The gas producer and auxiliaries are identical t o those used in standard producer gas operations, except that the conventional air-steam blast is replaced by an oxygen-steam blast. Figure 1 shows a phantom cross-section view of one of the gas producers and Figure 2 shows a diagrammatic sketch of the producer and auxiliaries for raw gas production, together with the locations of test and control equipment. Figure 3 shows a similar diagrammatic sketch of the system for conversion of raw gas. The test procedures employed, methods of sampling and analysis, and special operating techniques have been described (8, 18). Briefly, the test procedure followed closely that recommended in the test code procedure for gas producers (g), modified only where necessary because of the use of oxygen instead of air. For the sampling and analysis of material and products the methods of the American Society for Testing Materials (3) were employed except for precision gas analyses where the procedure recommended by Shepherd (14)-fractional combustion of hydrogen and carbon monoxide using the copper oxide tubewas adopted. Experimental Results

Physical and chemical data for the fuels tested are presented in Table I. Operating data on raw gas production are shown in Table I1 for a few selected tests on rice and barley anthracite and on coke a t various gasification rates. In Figure 3 are shown, for a specific 1-hour test period, average data for temperatures and flow of materials in the converter system. Figure 4 shows a schematic diagram of the conversion and purification system, together with data on the composition of the gas a t various points, and calculated values for the quantities of gas based upon 1000 cubic feet of entering carbon monoxide plus hydrogen. Discussion of Results

The operation of the oxygen-steam producer differs but little from normal air-steam operation and is characterized principally by the fact that with oxygen very much higher gasification rates are possible, appreciably higher quantities of steam must be used, and fuel bed conditions change much more rapidly. The difference between operation with large sizes of coke and small sizes of anthracite is characterized principally by the fact that with the small anthracite, ash bed condition's are such as to require different grate spacings and eccentricity from either air-anthracite or oxygen-coke operations, and a larger quantity of steam is required because of the more concentrated combustion zone in the producer. These items are discussed in more detail in subsequent sections.

TABLE I. PHYSICAL AND CHEMICAL DATAON FUELS TESTED Kind of fuel= Nominal sizec

Anthracite Anthracite Coke! Rice Barley 2 X '/a inch

Proximate analysisd, %

HzO

Volatile matter Fixed carbon Ash Ultimate analysis, % C

H

0 N

s

Ash Heating value, B.t.u./lb. Ash fusion temperature, Initial deformation Softening temperature Fluid temperature

F. .2,100

2,060

,...

2,810

2,810

2,200

2,940

2,940

b Typical analysis for one shipment of coke used in tests. 0 Anthracite, rice 6/18 X 8/18 inch. barlev. a / l a Y ind ecreens. Coke, equivalent round-hole'scre d All analyses reported on as-

....

Figure 1.

Schematic cross section of producer

Gasification Rates. COKE. The gasificatiw rates of 4000 to 5000 pounds per hour regularly employed in oxygen-coke operations a t Trail are appreciably higher than the 3000 pounds per hour normally obtained in air-coke operations in the same equipment. Under test, a maximum gasification rate of 5750 pounds per hour (73 pounds per hour per square foot of grate area) has been attained but because of the additioml labor involved in feeding the large-size coke through the intermediate storage bin and feeder pipe system and maintaining uniform bed depths, this rate is not practical for continuous use with existing equipment. On the basis of test results, however, there is reasonable justification for assuming that rates higher than 5750 pounds per hour could be readily maintained after minor modifications to the feeder system and to the offtake and scrubber system t o permit more rapid movement of coke and of products. ANTHRACITE.For anthracite operations, oxygen gasification rates of 3200 and 1900 pounds per hour (41 and 24 pounds per hour per square foot of grate area) were attained in tests with rice and barley sizes of anthracite, respectively. These rates compare with 1600 and about 600 t o 700 pounds per hour attainable in the same producer using an air-steam blast. The limiting factor in these tests was ability to control ash bed conditions satisfactorily. Because of the more concentrated combustion zone resulting from the use of smaller sizes and oxygen, the anthracite ash tended t o surface-sinter and was very much more compact than the ash from either oxygen-coke or air-anthracite operations. This caused the ashes t o sift through the grates much faster than desired and necessitated stopping all grate motion a t times during the tests in order to prevent possible loss of the protecting ash layer on the grates. Slowing or stop-

594

INDUSTRIAL AND ENGINEERING CHEMISTRY

GOLD WATER

TEMF!,

1

[

Vol. 40, No. 4

Figure 2. Diagrammatic sketch of producer and auxiliaries

AS SAMPLING

TEMP: BLAST BEFORE S Q-Q I

POINT

-

Y

LOW P R E S S U R E STEBM HOT WATER

METERM

ping the grates prevented normal agitation of ash bed and caused ash crusts t o form which led t o subsequent difficulties v i t h ashlevel control and fostered clinker formation. Suitable niodifications t o the grate spacings and eccentricity would no doubt assist materially in overcoming these conditions and should increase appreciably the permissible gasification rates with these sizes of anthracite. The quantities of anthracite available for test purposes a t Trail viere insufficient, however, t o permit extensive experimentation on grate spacing or eccentricity. The use of larger sizes of anthracite should permit the attainment of gasification rates a t least as high as, if not appreciably higher than, those already attained with the larger sizes of coke. Gas Composition and Quality. COKE. As may be seen from the data in Table 11, the hydrogen plus carbon monoxide contents of the raw gas from the coke tests approach those normally attained in blue gas operations. The nitrogen content is relatively low and is largely a function of the quantity present in the oxygen used. Methane content is satisfactorily low. Tars and hydrocarbons higher than methane are produced in small quantities, but appear t o be a function of the quality of the coke used rather than the result of chemical reactions in the producer. These impurities are objectionable because of subsequent problems in the gas handling, purification, and conversion system. ANTHRACITE. The raw gas from anthracite tests shows a somewhat lower content of carbon monoxide plus hydrogen, but has a more favorable ratio of hydrogen t o carbon monoxideapproximately 1 to 1-than that from coke. As far as could be detected during 4 weeks of operation, no tars or higher hydrocarbons were produced, although the methane content of the raw gas was possibly a little higher than for similar coke operations. This latter point is open t o question, however, as the accurate determination of small quantities of methane in the presence of high percentages of carbon monoxide and hydrogen is extremely difficult by standard gas analysis methods. The problem of dust carry-over was not serious with either coke or anthracite. Some carry-over occurred, but the dust was removed in the regular scrubber normally used with this class of producer.

Production of Hydrogen and Synthesis Gas. COKE. h material balance, raw gas t o hydrogen, for the ammonia synthesis is shown in Figure 4, based upon gas analyses a t various points throughout, the system as indicated. This balance has been calculated in terms of 1000 cubic feet of carbon monoxide plus hydrogen entering in the raw gas and upon complete conversion of carbon monoxide t o hydrogen for the ammonia synthesis. The calculated quantities of gas are based upon carbon monoxide plus hydrogen present rather than upon nitrogen contents, because the latter include all cumulative errors in analysis and are believed t o be less accurate. I n this balance, losses of gas other than those occurring in scrubbing operations have been neglected. This assumption appears t o be justified, as pump slippage losses are recovered and recycled, while other losses are believed t o represent but a small fraction of 1y0 a t this particular plant, As may be seen from an examination of this balance, the loss of carbon monoxide and hydrogen in the scrubber system is a significant item. This loss amounts t o 5.75% of total carbon monoxide and hydrogen entering. Although not practiced at this plant, scrubber losses could be reduced by a two-stage pressure let-down on the scrubber water and by recycling the gases liberated when the pressure is reduced t o 30 t o 40 pounds per square inch. Detailed information on the method and recovery possible has been published by Seifert and Ogilvie ( I S ) . Although the exact recovery and analysis of recycle gas will vary with a number of factors, plant scale results reported by these authors suggest that the loss can be reduced t o between 0.5 and 1yo. As an alternative the use of ethanolamines instead of water for the scrubbing operation is reported (19) t o reducc this loss t o negligible proportions. A similar material balance has been calculated for the paitial Conversion of carbon monoxide t o give a synthesis gas having a final hydrogen t o carbon monoxide ratio of approximately 2 t o 1. This balance depends, of course, upon t h e purity requirements of the process for which the synthesis gas is intended and upon the particular conditions selected. For t h e standard Fischer-Tropsch synthesis, inert contents up to 10 or 12y0are considered permis-

,

April 1948 RAW GAS

SAT. AT 20'C.

INDUSTRIAL AND ENGINEERING CHEMISTRY

Low STEAM,

AI

PRESSURE 19,00O*/HR.

t

595

Figure 3. Diagrammatic sketch of converter system, showing average operating data for a 1-hour test period BURNER TO

CONVERTED

AMPLE POINT

FROM ALTERNATE UNIT

balance shown in Figure 5 has been calculated, assuming a raw gas of the same composition as that used for the balance shown in Figure 4, a 2 t o 1 ratio of steam t o gas at converter inlet, a converter temperature of 450' C., a hydrogen t o carbon monoxide ratio in t h e converted gas of 2 t o 1, substantially complete removal of carbon dioxide, and no recycling of let-down gas from the scrubber. This latter assumption causes a change in the hydrogen t o carbon monoxide ratio of the finished gas but is of interest because it indicates the extent of the losses that may be incurred. Assuming a higher or lower scrubber DATAAND RESULTS FOR SELECTED ANTHRACITE TABLE 11. OPERATING efficiency or the use of recycling, the * AND COKE TESTSa quantity and composition of the waste A i 3 L scrubber gas would change and the Barley D E F G Rice Rice anthra- 2 Inch 2 Inch 2 Inch 2 Inch anthra- anthraratio of hydrogen to carbon monoxide cite coke coke coke coke cite cite Test Fuel in the finished gas would be altered 17 25 11.62 11.92 29.0 23.6 18.5 Duration, hours 69,280 36,160 15,920 72,100 72,200 76,300 68,560 Fuel consumed, Ib. slightly. 454 625 630 400 435 200 Pure oxygen flow, cu. ft./min,b 378 ANTHRACITE. A material balance 98.6 97.5 96.8 97.5 97.5 Oxygen purity % 95.8 98.0 3820 4840 1825 2140 3125 3260 1467 Steam additioAs auxiliary, lb./hour similar t o that shown.in Figure 4 was 3875 4663 1714 5630 2520 3365 3990 Total steam used Ib./hour 879 765 730 508 585 678 Offtake temperatLre F. 814 made during operation with anthracite 20 7-10 7-10 5-6 16 19 Combustion zone debth est., inches ... as fuel and although it differed some1.1 2.6 3.0 3.0 2.0 2.0 2.0 Combustible in dry refuse, % 2771 3112 3060 4125 2490 4020 1335 Fuel gasified, Ib./hour what in specific quantities because of 52.5 31.7 51.4 17.0 33.0 39.6 39.0 Fuel gasified, lb./hour/s ft. 9.00 8.39 9.50 p9.20 9.70 9.40 Oxygen used, cu. ft. Oz/ft;. fuel 8.67 the different initial gas composition, 1.28 1.40 1.51 0.97 1.01 1.10 1.40 Steam used lb. steam/lb. fuel the over-all results were substantially Steam dedmposed, Ib. steam/lb. 0.73 0.61 0.61 0 75 0.57 0.62 fuel 0.73 the same and need not be presented 23.3 28.4 24.6 24.3 24.6 23.4 , 23.6 Total fuel, Ib./M cu. ft. 35.2 33.0 Total steam, Ib./M cu. f t . 32.5 24.8 26.8 23.9 32.7 in detail. During the short period 228.6 227 233.5 234 220 Total oxygen, cu. ft./M cu. ft. 204.2 195.5 when only the anthracite producer was Gas analysis, % 17.70 19.10 17.60 11.3 12.8 18.1 10.7 COZ in operation and the synthesis plant 0.15 0.10 0.20 0.6 1.8 0.3 0.4 0 2 40.55 41.35 37.50 31.0 30.2 30.9 36 2 Hz operated on anthracite gas alone, the 3 9 . 7 5 3 7 , 5 5 4 2 . 3 5 54.1 54.2 53.3 44.3 co 0.85 0.90 0.75 0.4 C Ha 0.4 0.4 0.3 ratio of steam t o raw gas was not 1 1 . 0 0 .oo 1 . 6 0 3 . 6 2.6 2.3 0.7 N2 changed. Because of the initially 80.2 76.2 85.8 86.8 84.5 80.0 Cold gas efficiency. 80.7 _ . ._ . higher ratio of hydrogen t o carbon monoxide 0 All reoorted anthracite data corrected for small loss of carbon dioxide resulting from absorDtion in moisture t h a t condensed in mercury-sealed gasometer used for collecting compoiite hourly s'amples. in the raw gas, the shift reaction was Although similar corrections were not made for coke tests because necessary data were not secured correction probably would be negligible owing to use of lower quantities of excess steam and generail; carried more nearly t o completion and lower carbon dioxide concentrations. b All gas data reported a 60° F. and 30 inches dry. resulted in a converter offtake gas of Cold gas Bff. = gross B.t.u. value of gas from 1 lb. of coal as fired the percentage B.t.u. value of 1 lb. of coal as fired + p e a t content above 60" F. bf auxiliary steam 0.75 carbon dioxide, 38.80; oxygen, 0.10; hydrogen, 55.60; carbon monoxide, Bible, while compression t o 250 pounds per square inch for scrubbing purposes is impractical. Low pressure scrubbing with lower carbon dioxide removal and a subsequent decrease in hydrogen and carbon monoxide loss is preferred. For methanol synthesis, on t h e other hand, carbon dioxide is reduced t o a minimum and either pressure scrubbing with increased synthesis gas loss or the more expensive ethanolamine scrubbing is used. As an indication of the general character of t h e results, the

I

-

I

INDUSTRIAL AND ENGINEERING CHEMISTRY

596

FT.

71 CU.

821 CU. FT. % 92.2

% 14.0

C02

O1 H2

0.7 4.9 69.6

co

CH+

1.6 9.0

COa

co

Ne

< M ( ,I

HZ

0. I 5.8

CH4 N2

0.1 0.6

0,

WASTE WASTE GAS TO ATMOSPHERE

I .e

-

REG€ NERATOR

I

I,

I

=T

Vol. 40, No. 4 gas installations where it was possible t o measure the blownover fuel by direct methods suggests losses up t o ten times this figure especially in view of the somewhat increased gas velocities in the upper part’ of the bed at the high gasification rates. 3. Other Possibilities for

PRODUCER

of t h e ash a t the start and end of the test; incorrectly establish9 6 7 CU. FT. 1874 CU. FT. 1190 CU. FT. ing the level of the fuel in the % % producer a t the start and end CO, 0.4 CO, 13.0 0, 0.7 0, 0.3 of the test; oxygen loss due to 9 7.5 H2 89.1 33.7 51 6 co 5.5 CO 4.6 C> 50.3 oxidation of mineral matter; CH4 0.1 CH4 0 3 CHI 1.0 Nz 4.2 N, 1.7 21 errors in gas analysis determination, especially carbon monFigure 4. Schematic diagram of conversion and scrubber system oxide and hydrogen; and erGas quantities and compositions for H2 production rors in meter readings due to Basis, 1000 cubic feet of pure CO + H2 entering fluctuations in load and pressure. 2.65; methane, 0.70; nitrogen, 2.15. As a result of this change I n the element balances, the major source of “unaccounted-for” in converter gas composition the relative compositions of the material is oxygen and the probable source of this loss is discussed in the preceding paragraphs. On a percentage basis, scrubber “offtake gas” and scrubber “waste gas” changed slightly. Total quantity and composition of the finished gas for the greatest discrepancy between input and output appears in the nitrogen balance. This, however, is t o be expected in such a ammonia synthesis, however, were but slightly altered. The raw gas from the oxygen gasification of anthracite has the balance, because the nitrogen which is present in only small desired ratio of hydrogen t o carbon monoxide for direct use in quantities in the gas, is determined by difference and includes the isosynthesis process. A light scrubbing t o lower the carbon all the cumulative errors in the analysis of other constituents. dioxide content would be necessary and in this case the use of In the material balance for the coke test, the input and output figures are substantially the same. As with the anthracite the ethanolamine process might be advantageous. since sulfide removal would also be desired. The raw gas from the oxygen tests, however, the element balances show a deficiencv of both oxygen and hydrogen. Although the oxygen deficiency is not so gasification of coke a t the highest steam-oxygen ratio also approaches a composition suitable for the isosynthesis process and great as for the anthracite test, the hydrogen deficiency is almost the same. Although less conclusive, these data again suggest could probably be used i\-ithout change in ratio. Material and Element Balances. Raw GAS. For the proloss of steam as the probable reason for errors in the element duction of raw gas, material and element balances have been balances. Converter System. An over-all material balance for the concalculated for a n anthracite and a coke test and are shonn in Tables 111 and IV, respectively. These balances are based upon verter system has been calculated. for one specific te?t period and is presented in Table V. This balance is based upon the measa carbon balance because actual gas produced was not measured ured output of hydrogen t o the svnthetic ammonia system and directly. I n the material balance for rice size anthracite, “unaccounted-for” losses amounted t o 2.6Y0,. The probable sources of loss in Tvhat is believed t o be IN POEXDS FOR TABLE 111. I\/IATERIALAND ELEMEXT BALAYCE t h e most likely order of magnitude are: ANTHRACITE TEST 1. Loss of Steam because of Condensation in Input Outpnt t h e Ash Pit. Condensation of some of the Lb. 70 Lb. 70 steam in the ash pit undoubtedly occurs. DurOver-all material balance Coal 69,280 31.7 Ashes 7,267 3.3 ing periods immediately follon-ing ash discharge 50,823 23.3 Dry gas 157,330 72.1 Oxsgen Water in oxygen 1,154 0.5 Water vapor a n d gas 47,980 22.0 the opportunity exists for the steam-oxygen Total steam 95,975 __44.5 Unaccounted-for losses 5,665 2.9 blast t o come into direct contact with the relaTotal 218,232 100.0 Total 218,232 100 .O tively cold n-alls of the ash cone. For an hour Oxygen balance Free oxygen 50,823 35.5 Total in gas leaving 93,837 65,6 or so after ash discharge, ivater does drip from Total in water and steam 91,066 63.7 Total in mater vapor 42,600 29.8 Combined in coal 1,167 0 . 8 Unaccounted-for losses 6,619 4.6 t h e ash-cone slide valve, but it is not possible t o -~ Total 143,056 100.0 Total 143,056 100.0 state definitely how niuch of this is the result Hydrogen balance 6.630 62.4 of condensation because water is used t o flush Total in mater and steam 11,497 9 0 . 8 Total in pas leaving Combined in coal 1,167 9 . 2 Total in water vapor 5,380 4 2.5 out the refuse. The fact that the oxygen and 12,864 1oO.o Unaccounted-for losses 654 6,l Total hydrogen balances both show a loss and that the Total 12,664 1 0 0 . 0 deficiency is roughly in the proport’ions of the Nitrogen balance 126.7 Total in gas leaving 2,175 26.4 Combined in coal 454 oxygen and hydrogen content of water suggests -26.7 Unaccounted-for losses -458 Free in oxygen entering 1,263 73.6 that steam condensation is a definite possibility. 100.0 Total 1,717 1,717 100.0 Total 2. Loss of Blovn-Over Fuel in Excess of Carbon balance Combined in coal 64,536 1 0 0 . 0 Total in gas leaving 54,347 99.7 That Measured by Impingement Method. Test 189 .3 in ashes __ _ _ _0_ -- TotalTotal data from t h e impingement method showed a Total 54,536 1 0 0 . 0 54,536 100.0 loss of only 0.05%, whereas experience in producer

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INDUSTRIAL AND ENGINEERING CHEMISTRY

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Vol. 40, No. 4 -

TABLE VI.

MATERIAL BALAKCES OVER SATURATOR, CONVERTER, AND WATERHEATER

Water Ib. Water'vapor, Ib. Dry gas, lb. Unaccounted for Total Steam and/or water vapor, lb. D r y gas, lb. Unaccounted for Total Dry gas, Ib. Water vapor with gas, lb. Water lb. Unaccbunted for Total

Dry gas, Ib. Water vapor with gas, lb. Water as liauid. lb. Unaccounted for Total

TABLE1711.

(Basis, 1 hour) Input Lb. % Saturator 519,000 96.9 250 0.0 16,570 3.1

output Lb. % 509,035 10,215 16,570

... lo0.0

.....

535,820

.....

95.0 1.9 3.1 0.0

j35,820

io0.0

22,500 24,089 -774 45,785'

49.2 52.5 -1 7

loo.0

24,059 14,480 617,072 -17

4.3 2.6 93.1 0.0

Converter 29,215 16,570

63.8 36.2

45,785

100.0

- .....

...

Water Heater 24,059 4.3 22,500 4.0 509,035 91.7

...

-

c

555,594

_

...

100.0

Condenser 24,059 7.8 14,480 4.7 270.000 87.5

..... -

. . I

308,539

100.0

555,594

ioo.0

24,059 415 284.074 -9

7.8 0 1 92.1 0.0

808,j39

l0o.b

HEATBALANCEFOR OXYGENPRODUCER TESTS WITH RICEANTHRACITE AND COKE

(Basis, 1 pound of fuel as fired.

Reference temperature, 60' F.) Anthracite Test Coke Test % B.t.u. B.t.u. Innu -t 12,250 89.8 12,000 92.9 Oa 0 05 0 n 05 0 0 1,376 10.1 898 7.0 4 3 0 0

TABLEVIII. VERTER

HEaT BALANCES OVER MAJOR UNITS O F CONSYSTEMFOR CONVERSION OF RAWGAS TO SYNTHESIS GAS OF COMPOSITIONS SHOWNIN TABLE V

(Basis, 1 hour.

Reference temperature 60' F.) Input output B.t.u. % B.t.u. % Saturator 64,875,000 98.8 50,802,000 77.4

Total heat in water Total heat in water vapor in gas 278,000 0.4 11,514,000 17.5 Sensible heat in gas 496,000 0.8 686,000 1.0 Radiation and unaccounted for ...... 0 .O 2,647,000 4. Total 65,649,000 100.0 65,649,000 100.0 Converter and Heat Exchanger Total heat in steam and water vapor 33,677,000 2 9 . 1 28,638,000 24.8 Sensible heat in gas 1,005,000 3,318,000 2.9 0.9 Potential heat in CO .48,016,000 4 1 . 5 6,362,000 5.5 Potential heat in Hz 30,074,000 2 6 . 0 72,051,000 62.3 2.5 1,324,000 Potential heat in CHa 2.930.000 1.1 4,009,000 Radiation and unaccounted for ....... 0.0 3.4 Total 115,702,000 100.0 115,702,000 100.0 Water Heater Total heat in steam and water 28,023,000 34.1 16,292,000 19.8 vapor Sensible heat gas 3,489,000 4.2 1,035,000 1.3 Total heat in water 50,776,000 6 1 . 7 66,496,000 80.8 Radiation, unaccounted f o r , ....... 0.0 -1,535,000 -1.9 and errors -___ Total 82,258,030 100.0 82,288,000 100.0 Condenser Total heat in water 216,000 1.2 19,658,000 112.1 Total heat in water vapor 16,292,000 92.9 416,000 2.5 Sensible heat in gas 1,036,000 5.9 23,000 0.1 Radiation. unaccounted for, ....... 0.0 -2,584,000 3 7 and errors Total 17,543,000 100.0 17,543,000 100.0

-

~

--

---A-

Calorific value of fuel Sensible heat of fuel Total heat of water supplied Total heat of steam entering Sensible heat of oxygen entering Total heat of water in oxygen entering Total

- - - 18 13,651

0.1 100.0

13 12,914

0.1 100.0

output 1 Calorific value of gas 11,376 83.4 11,173 86.5 2 Sensible heat of dry gas 715 5.2 492 3 8 3 Total heat of water vapor in gas 1,006 7.4 487 3 8 Q 4 Sensible heat of ashes 4 0 0 1 5 Heat lost in jacket overflow water 468 3.4 75: 5 8 6 Calorific value of ashes 40 0.3 0 3 39 7 Radiation and convection 31 0 2 21 0.2 8 Errora and unaccounted for 11 0.1 -0 5 58 Total 12,914 13,651 100.0 1ZiT3 a N o actual data. Assumption made t h a t temperature of fuel and water supplied was 60' F.

- -

~

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such calculation would add little to the information already presented. Comparisons between Oxygen Producer Process and Other Synthesis Gas Processes

Regardless of the process or source of raw material, the ultimate objective of all synthesis gas processes is to produce a finished gas of satisfactory hydrogen to carbon monoxide ratio containing the lowest content of impurities commensurate with process requirements and cost. The desired ratio of hydrogen to carbon monoxide will, of course, vary with the process and purpose for which the gas is intended. In general, however, a ratio between 1 t o 1 and 2 to 1 will be desired for the Fischer-Tropsch and related synthetic fuel processes. Similarly, the permissible inert content may vary from as high as 10 t o 12% for the fuel synthesis processes to as low as 1 or 2% for methanol synthesis. Because of these variations in requirements and because thc relative values of raw materials, equipment, and labor vary in different localities, no single figure encompassing these variables can be used to compare the relative merits of the various proposed processes. Therefore, the results for the oxygen producer tests together with published data on other processes have been

calculated and are presented in Table X, in terms of carbon, oxygen, and steam requirements for the production of raw gas containing 1000 cubic feet of pure hydrogen plus carbon monoxide. Because the raw gases produced from solid fuels by any of the existing processes fail to meet the 2 to 1 ratio of hydrogen t o carbon monoxide desired for some synthesis processes, and in many cases have considerably less than a 1to 1ratio, additional steam may be required to perform the shift reaction. The ratios of hydrogen to carbon monoxide are also reported in Table X, therefore, together with calculated values for additional steam requirements to produce the 2 to 1ratio. These calculated values are based upon the assumption that 3 volumes of steam will be required per volume of carbon monoxide converted. The use of the threefold $xcess in the calculations may be open to question but is a compromise between the total steam required, which itself is a variable depending upon raw gas composition and conversion temperature, and the steam produced in the process due to heat produced in the shift reaction. For comparative purposes the inert contents are also reported in terms of cubic feet per 1000 cubic feet of pure hydrogen-carbon monoxide mixture. D a t a used in the calculations for processes other than the oxygen producer were obtained from recent reports on the Winkler, Koppers, Thpsen-Galocsy, and Leuna Slagging producers b y Newman (&?,IO)and on the water-gas process by Young (19). As may be seen from a n examination of these data, the oxygen gasification results for both coke and anthracite compare very favorably with data for other processes. I n general the results with anthracite are better than comparable data for coke because about 3.5 cubic feet of volatile matter are released per pound of coal, and this volatile matter will normally average better than 80% hydrogen by volume. Carbon consumptions, per 1000 cubic feet of pure hydrogen plus carbon monoxide produced, are as low as or lower than for any of reported commercial processes. Oxygen consumption is as low as or lower than for any process except the regular water-gas process, where air rather than oxygen is used. I n this intermittent process an additional 10 to 12 pounds of carbon is consumed in place of oxygen. Stcam consumption, except in the case of the water-gas process,

April 1948

INDUSTRIAL AND ENGINEERING CHEMISTRY

599

Synthesis gas plant of Consolidated Mining and Smelting Company of Canada, Ltd., Trail, B. C.

is t o some extent a function of the ratio of hydrogen to carbon monoxide produced and of the character of the fuel used. I n the Koppers and WinMer pulverized fuel processes the steam production from the waste heat of the gasification processes amounts t o 44 and 54 pounds, respectively, so that although the actual consumption for gasification is, respectively, 29.4 and 47.3 pounds, the net result is a negative consumption. Several other processes including the oxygen producer operation could develop a substantial percentage-if not all-of the steam requirements by utilizing waste heat in the offtake gases. Whether such re covery would be profitable to the oxygen producer is open to question because of relatively low offtake temperatures. Over-all carbon efficiencies, defined as pounds of carbon appearing as carbon monoxide in the finished synthesis gas per pound of carbon in the original fuel, are of considerable interest. I n general, the carbon efficiencies for all the fixed-bed complete gasification processes range in the vicinity of 45%. The watergas process and the Winkler process both show relatively low carbon efficiencies. I n the former i t is due t o the intermittent character of the process and in the latter to the high percentage of fly material. The high carbon efficiency of the Koppers pulverized fuel process is surprising, especially in view of the high offtake temperatures reported. More detailed information on plant-scale operations will be of interest. It is not the purpose of this paper t o discuss the relative economics of the various processes, but in addition to conversion efficiency-solid fuel t o finished synthesis gaa-other major factors must be considered in comparing the merits of various processes for synthesis gas production: type of fuel required, over-all cost of gasification equipment, and labor required. With respect t o quality of fuel required, all fixed bed processes are at a disadvantage as compared to the pulverized fuel processes because; in general, sized material commands a higher price. F o r the oxygen producer process the fuel size requirements are not

so rigid as for the conventional water-gas process, but in comparison with pulverized bituminous coal or lignite neither sized coke nor anthracite is competitive. Some possibilities for application of the oxygen producer process may exist, however, using material of coarse size such as high-carbon mine refuse, pelletized silt, coke breeze, or other sources of relatively low-cost carbon. The over-ai1 cost of the gasification equipment should be compared only in terms of equipment cost per unit of synthesis gas produced, but since most of the foreign figures for equipment costs are not directly translatable into costs in this country, comparison is difficult. The oxygen producer is, however, a relatively low-cost piece of equipment and experience has shown that maintenance costs are comparatively low. The best available estimates ( I ) indicate that equipment and maintenance costs per unit of gas produced would be lower for the oxygen producer than for any process thus far developed. The gasification rates for the oxygen producer in terms of quantity of carbon monoxide plus hydrogen produced per square foot of generator area per hour were, respectively, 1340 and 1800 cubic feet for tests B and F. This is substantially lower than the 3630 and 8630 cubic feet per square foot per hour reported for the Winkler process plants at Bohlen (9) and a t Zeitz (10). The gasification rates in terms of cubic feet of generator volumea much more significant figure re5ecting over-all size and probable cost of equipment-show, however, t h a t the oxygen producer has a n appreciably higher capacity than the two Winkler units mentioned. The rates for the oxygen producer were, respectively, 155 and 205 cubic feet per hour per cubic foot of generator volume, while the Winkler units produced only 60 and 145 cubic feet per hour per cubic foot of volume. The general belief that the fluidbed type of gasification is capable of giving gasification rates markedly superior t o those attained in other types of equipment appears t o be totally unfounded if consideration is given t o the over-all size rather than just the diameter of the generator.

INDUSTRIAL AND E N G I N E E R I N G CHEMISTRY

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TABLE IX. OVER-ALLHEATBALANCE FOR CONVERTER SYSTEM SATURATOR INLET TO WATERHEATEROUTLET (Basis, 1 hour. Reference temperature B O 0 F.) Input B.t.u. Sensible heat in gas entering 496,000 Total in water vapor in gas 278,000 Potential heat in gas entering co 48,016,000 30,074,000 H2 2,930,000 CH4 64,875,000 Total heat in water entering saturator 22,711.000 Total heat in steam entering after saturator 169,380,000 Total

% 0.3 0.2 28.4 17.7 1.7 38.3 13.4 ~

hydrogen. This loss may be reduced appreciably by two-stage pressure letdown and recycling part of the gas. Material and heat balances over the producer and over the converter system are shown for the production of hydrogen, and a calculated material balance is shown for the production of synthesis gas of 2 to 1 ratio of hydrogen t o carbon monoxide. These balances indicate some of the factors involved in the overall problem of synthesis gas production from solid fuels. The economics of the oxygen gasification of solid fuels in fixed beds hinges largely upon the cost of the solid fuel suitable for the process and upon the cost of oxvgen, the relative efficiency as compared with other processes, and the capital charges.

100.0

Aclmowledgments

output Sensible heat in gas leaving Total heat in water vapor in gas Potential heat in gas leaving

co

H2 C H4 Total heat in water leaving water heater Radiation and unaccounted for Total

1,035,000 16,292,000

0.6 9.6

6,362,000 72,051,000 1,324,000 66,496,000 5,820,000 169,380,000

3.8 42.6 0.8 39.2 3.4 __ 100.0

With respect t o labor involved in the gasification processes, little information is available as t o requirements for foreign processes. For the oxygen-producer, labor requirements per unit of gas produced should run about the same as or a little higher than for the standard water-gas process. On a large scale operation advantage could be taken of mechanical devices t o save labor, but these materially increase the equipment cost and might not justify the increase in overhead and maintenance. Conclusions

The oxygen gasification of solid fuels in fixed-bed producer gas equipment has been demonstrated a s a practical commercial process. The construction and operation of the equipment are simple, its performance is well known, and its cost is relatively low. The results of plant scale operations indicate t h a t the over-all carbon, oxygen, and steam consumption for synthesis gas production compare favorably with any fixed-bed, fluid-bed, or coal-suspension proceBs thus far developed t o commercial operation. Gasification efficiencies for synthesis gas production are high. The gasification rates per unit of generator area are 50 to 15070 higher than those normally attained in air-blown producer operations and higher than those attained in the standard watergas generator when the same size fuel is used, but appreciably lower than the rates attained for the pulverized fuel processes. The gasification rates per unit of generator volume are, however, 2 t o 4 times those of the Winkler process. The quality of the raw gas from oxygen gasification of anthracite was such t h a t with a slight reduction in carbon dioxide content it could be used directly for the isosynthesis process. Raw gas from the coke at the highest steam-oxygen ratio approached this composition. For conversion t o a gas of 2 t o 1ratio of hydrogen t o carbon monoxide the shift reaction must be employed. Removal of carbon dioxide from the gas by pressure scrubbing with water results in a substantial loss of carbon monoxide and

TABLE X.

Vol. 40, No. 4

For permission t o publish this paper thanks are extended t o the management of the Consolidated Mining and Smelting Company of Canada, Ltd., whose operating and research staffs assisted in obtaining the data contained herein. The authors wish also t o extend thanks and appreciation to the anthracite Institute, which financed the anthracite tests, t o the several individuals and organizations who contributed time and effort to the performance of the several tests, and t o the staff of the Division of Fuel Technology of The Pennsylvania State College for help in calculating and analyzing the data and for helpful criticism and comments during preparation of the manuscript. Literature Cited (1) Am. Gas Assoc., New York, “Gas Making Processes,” October 1945.

(2) Am. Soc. Mech. Engrs., New Y o r k , “Test Code for Gas Pro(3) (4)

(5) (6) (7)

(8) (9)

ducers,” Kovember 1928. Am. Soc. Testing Materials, Philadelphia, “Standards on Coal and Coke,” October 1938. Anon., Coke and Smokeless-Fuel A g e , 4, 103. 127, 141 (1942). Gradzovsky, M . K., and Choukhanoff, Fuel, 25, 321 ( 1 9 3 6 ) . Holroyd, R . , Combined Intelligence Objectives Sub-Committee, Report XXXII-107, Item 30 ( 1 9 4 6 ) . Lane, J. C., and Weil, B. H., Petroleum Re.finer, 25, 355 ( 1 9 4 6 ) . Mitchell, R . F., Can. Chem. Process I n d . , 30, 34 (Auguet 1946). Newman, L . L.,Am. Gas Assoc.. Joint Conference of Production Chemical Committees, Preprint (June 1946) ; Am. Inst. Mining Met. Engrs., Coal Tech., Tech. Pub. 2116 (November 1946).

(10) Newman, L. L., IND. EXG.CHEM.. 40, 559 ( 1 9 4 8 ) . (11) Powell, A . R., Am. Gas Assoc., Joint Conference of Production

Chemical Committees (June 1947).

(12) Reed, R. (13)

M.,Petroleum Refiner, 25, 367

(1946).

Seifert, H. L., and Ogilvie, J. D. B., Can. Chem. Process Ind., 28, 665 (1944).

Shepherd, M . , Natl. Bur. Standards, Research Paper 1704 (March 1946). (15) Stewart, A. T., Proc. Can. Gas Assoc. (1937) ; Can. Chem. Met., (14)

21, 283 (1937) : Gas J . , 219, 343 ( 1 9 3 7 ) .

Von Fredersdorff, G., Kats, S., and Pettyjohn, E. S.,Am. Gas Assoc., Joint Conference of Production Chemical Committees (June 1947). (17) Weir, H . M.,IND. EXG.CHEX.,39, 48 (1947). (18) Wright. C. C.. and Newman, L. L., Am. Gas Assoc., Joint Conference of Production Chemical Committees (June 1947). (19) Young, H. B., Proc. Am. Gas Assoc., 901 (1932). (16)

RECEIVED OCTOBER 16, 1947.

SYNTHESIS GASPRODUCTION EFFICIENCIES FOR VARIOUS PROCESSES AND FUELS

Lb. Required t o Produce 1000 CU. F t *of Pure co * Inert Content, Total, Lb. Carbonb Pure Hz/CO Cu. Ft./M Cu. Ft. Steam/v Cu. Ft. Efficiency, Proceas Fuel Size Test Carbon oxygen Steam Ratio Pure Synthesis Gas Synthesis Gas5 7% 45.8 57.6 2.1 34.5 1.02 Oxygen producer Anthracite 21.5 Rice 4 5.6 5 6 . 9 20.9 36.5 1.10 2.4 Rice Oxygen producer Anthracite 44.6 62.9 2.9 24.2 34.8 0.89 Barley Oxygen producer Anthracite 4 5 .2 6 3 . 8 3 . 5 20.4 0.57 22.6 2 inch Oxygen producer Coke 45.4 65.0 4.7 23.4 20.7 0.56 2 inch Oxygen producer Coke 4 4 . 7 6 6 . 8 2 . 2 2 2 . 0 0 . 5 8 23.5 2 inch Oxveen Droducer Coke 45 0 64.4 1.2 33.2 0.82 23.1 2 inch Ox?ien producer Coke 2 6 . 7 3 . 8 3 . 0 6 . i a 1 . 4 5 2 6 . 8 1/1 inch Crude ,coke Winkler 46.8 9.4 -14.6d 1.00 25.2 1.1 Bituminous ooa Dust Koppers 43.7 72.7 14.0 0.34 4.5 27.3 11/2-21/2 inch Thyssen-Galocs y Coke 43.0 67.4 0.9 20.1 0.50 22.2 Lump Coke Leuna-Siagging 29.8 72.6 7.3 Air 56.5 1.24 Lump Coke Water gas a Assuming conversion t o 2H2:lCO and 3 volumes of steam used per volume of carbon monoxide converted. b Defined as pounds of carbon appearing in pure synthesis gas (2Hz:lCO) per pound of carbon in original fuel. Steam used 47.3, steam produced 54.0. d Steam used: 29.4; steam produced: 44.0.

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