INDUSTRIAL AND ENGINEERING CHEMISTRY
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Vol. 40, No. 4
eter analysis shows that a minute but definite quantity of mcthTABLE111. G s s AXALYSISOF PRODUCT FROXI Gas GENERATOR ane and of hydrocarbons of higher molecular weight are actually Run 15 synthesized in the generator. This is presumably due to partial - -< ;\lass Average hydrogenation of the carbon monoxide in the presence of either Spectrometer Orsat Orsat the nickel catalyst itself or possibly iron contaminants. Hz 65.35 68.1 67.1
co
cor
Cila
C2H6 C6H6 CIH8 C1H6
CaHio H?/CO
33.74 0.30
30.8 0.3
0.11
, . . . , . . .
0.21 0 03 0.16 0.05
0.8
....
....
0.05 100.00
lhO.0 __
1.93
2.21
30.9 0.2
.... .... .
1.8 , .
...,
....
100.0 -
2.17
a reactor temperature of 260 to 320 C. is adequate for the dehydrogenation reaction. Operating practice has been to regenerate when the catalyst temperature began to exceed the higher figure. Mass spectrometer and Orsat analyses run on a gas sample obtained during run 15 are presented in Table 111, together with a n averaged Orsat analysis for other runs. No explanation is offered for the discrepancy between the Orsat and mass spectrometer analysis, but it is believed t h a t the Orsat values for hydrogen and carbon monoxide are the more reliable. Only a small fraction of the product gas, usually about 2%, consists of products other than hydrogen and carbon monoxide. The mass spectromO
Conclusions
No difficulties in the operation of the unit have been encountered. The unit can be started in 3 hours from a complete shutdown and shut down in about a half hour. By leaving partial heat on the reactor and preheater, the starting time may be halved. The method is considered satisfactory and is highly recommended for producing moderate quantities of reasonably high purity hydrogen, and carbon monoxide in a ratio of 2 .
O
Acknowledgment
The authors appreciate the important contributions which W. B. Plummer, E. L. D’Ouville, B. L. Evering, J. Zisson, and F. Kalina made to various phases of the work. Literature Cited
S.Patent 2,010,427 (assigned t o Carbide and Carbon Chemicals Corp., August 6, 1936).
( 1 ) Eversole, J. F., U.
R E C E I V EOctober D 4, 1947.
Production of water gas from pulverized coa A Continuous Process JOHN F. FOSTER Battelle Rlemorial Institute, Columbus, Ohio
M
ANY processes for the continuous production of water gas from pulverized coal have been proposed in the past, and descriptions of a number of these have been given by the Institute of Gas Technology in a comprehensive review of gas-making processes ( 2 ) . A common basis for all the proposals for the use of pulverized coal in the continuous production of water gas has been the belief that the conventional water-gas generator is subject to improvement on two counts. (1) The conventional generator requires solid fuel of large and relatively uniform size in order that the flow of air or steam may be evenly distributed through the fuel bed, and the bed may thus be maintained in good condition for gas making over long periods. The necessity for using a n optimum size of coke or coal prevents the use of smaller, cheaper fuels. If a generator can be developed to utilize pulverized fuel, an attractive saving in fuel costs appears possible. (2) The production of water gas in a conventional generator is a n intermittent process, in which air is first blown through the fuel bed to release by combustion, and to store, the heat necessary for the subsequent endothermic reaction between carbon and steam, and to raise the fuel bed temperature to a level a t which the reaction will proceed at a reasonable rate. Next, the air is cut off and steam is passed through the bed; in this period water gas is produced, and the temperature falls because of the heat absorbed by the reaction. As soon as the temperature reaches a minimum below which the reaction will not proceed a t an eco-
nomical rate, the steam must be cut off and air again admitted t o raise the temperature of the bed before further gas can be produced. If the heat for the carbon-steam reaction could be supplied through the walls of the reaction chamber, and the process thus made continuous, the necessity for precise control of the timing of the gas-making cycle would be avoided, and possibly labor and investment costs for a generator of given capacity would also be reduced. As pointed out by Barnes ( I ) , there have been two obstacles to the development of successful water-gas processes using indirect heating of the reaction chamber to maintain continuous gas production: (1) When refractory walls were used, the comparatively low thermal conductivity of the refractory limited the rate a t which heat could be supplied, and this in turn limited the capacity of the apparatus. (2) With metal walls, the maximum temperature which could be used without rapid deterioration of the metal was comparatively low, and the rate of gasification was correspondingly small. Barnes suggested that the then recent development of temperature-resistant alloys, capable of withstanding temperatures several hundred degrees higher than the alloys previously available, justified an investigation of the possibilities of gasification of pulverized coal by steam in a n externally heated chamber constructed of one of these alloys. As a result of Barnes’ recommendations, the present investigalion was initiated under the sponsorship of Bituminous Coal
April 1948
I N D U S T R I A L A N D E N G I N E E R I N G CHEMISTRY
587
THIS paper presents results obtained in an experimental investigation on water gas sponsored by Bituminous Coal Research, Inc. The rate of production of water gas by reaction between pulverized coal and steam, flowing continuously through a vertical, externally heated alloy tube, is shown to depend upon the chemical activity of the fuel, the rate of fuel feed, the rate of steam supply, the temperature, and the dimensions of the tube. A n empirical equation is presented, describing the effect of these variables on gas production. The use of the equation permits an estimate of the capacity of equipment of commercial size. Comparison of investment costs for generators employing this process with conventional water-gas generators of equivalent capacity indicates that the conventional process is enough lower in investment cost to preclude the commercial development of the continuous process under present economic conditions.
Research, Inc., t o determine on a pilot plant scale the gasification capacity of a n externally heated alloy tube and to evaluate the commercial feasibility of a full-scale plant consisting of multiple tubes externally heated by a single furnace. Experimental Apparatus
4 n exterior view of the gasification furnace is given in Figure I, and the experimental apparatus is shown diagrammatically in Figure 2. Gasification takes place in a reaction tube, which is suspended by its upper end from the top of the furnace, passes completely through the furnace, and has its lower end set in a water seal at the inlet to the ash receiver. Pulverized coal is fed to the top of the reaction tube by means of a rotary table feeder, with an attached hopper that can be seen in the upper part of Figure 1. Connection is made between the feeder and reaction tube by means of a n adapter with a built-in Venturi section. The steam for the process is admitted through a nozzle at the Venturi throat, which serves t o reduce the pressure a t the feeder to slightly below atmospheric, and thus prevent excessive steam condensation in the hopper, which would interfere with proper feeding. The furnace has a steel shell with a 9-inch lining of insulating brick. Internal measurements of t h e furnace are 18 X 18 inches and it is 9 feet 3 inches high. Four gas burners are set into each of two opposite sides of the furnace for heating the reaction tube externally. Reaction tubes of two different diameters both 11 feet long, were used during the experimental work. The smaller tube had an inside diameter of 5 inches, with a 0.375-inch wall, and was cast from a 28-12 chromium-nickel alloy. The larger was 8.25 inches in inside diameter, with a 0.5-inch wall, and was cast from a 25-12 chromium-nickel alloy. Thermocouples were attached a t intervals t o the outer surface of the reaction tube, with leads brought out through the wall of the furnace and thence to a selector switch and indicating potentiometer. The temperature of the furnace was controlled by manual adjustment of the gas burners. Once steady conditions were attained, it was possible to control temperatures along the entire length of the tube with a variation of less than 50' F. Gas and entrained solids passed from the reaction tube through a series of baffles in the ash receiver, where the bulk of the solids was removed from the gas stream, and thence through a wash box, a spray tower paeked with glass fiber, and finally an oilfilled automotive-type dust filter. The clean gas passed through a n exhausting pump and gas meter and was sampled before being finally released t o the atmosphere and burned. The exhausting pump was a Roots-Connersville Type XA-33 positive pressure blower, which was run at constant speed with the pumping rate controlled by adjusting a throttling valve in a by-pass connection (not shown) between the inlet and outlet. The throttling valve was set so that the pressure at the mouth of the coal feeder was approximately 0.5 inch of water below atmospheric, and adjusted as required to compensate for variations in the rate of gas production. Pressures at various points in the system were indicated by eight manometers, whose locations are
Figure 1.
Furnace for gasification of pulverized coal
shown by M in Figure 2. Manometer readings were recorded periodically throughout a test in order to detect any obstruction to the gas flow, within the system, before it could build up to serious proportions. The gas produced was measured by means of a No. 2 Emco meter with a rated capacity of 30 cubic feet per minute. Steam was measured by means of a sharp-edge orifice, and natural gas for firing the furnace by similar means. Steam was taken from the service lines, and its temperature was raised to approximately 500" F. in a gas-fired superheater before it was introduced a t the top of the reaction tube through the nozzle, which was constructed of a drilled pipe plug. Experimental Procedure
Before each test the hopper of the coal feeder was filled with pulverized coal, and the feeder was weighed on a beam balance, then lifted with a n overhead crane to the top of the furnace, and bolted to the reaction tube. At the end of the test, the feeder
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INDUSTRIAL AND ENGINEERING CHEMISTRY
Vol. 40, No. 4
MI
ST EAW
2
WASH B O X
ASH RECEIVER
F i g u r e 2.
was again weighed, and the total coal fed determined by difference. After the furnace was brought t o the desired temperature, which usually required less than one hour, the coal feed and steam were started simultaneously. Periodic readings were taken of tube temperatures, steam rate and temperature, gas production, manometer pressures, and gas used for firing. Fifteen-minute composite samples were withdrawn a t half-hourly or hourly intervals, and analyzed immediately in an Orsat apparatus. Fuels S t u d i e d
Table I shows the source and analyses of the five different lots of coal, including two bituminous coals and three lignites, which have been investigated.
TABLE I. ANALYSESOF COALSUSEDIX GASIFICATION STUDIES Lot n-0. Rank Seam State County Moisture Volatile matter Fixed carbon Ash B.t.u. Carbon Hydrogen Oxygen Nitrogen Sulfur Ash Total
182 167 196 Bituminous A Lignite Lignite Elkhorn .... .... S. D. S . D. Kentucky .... .... Letcher Proximate Analyses 13.0 12.0 0.4 38.7 39.2 33.9 41.6 40.5 62.5 8.8 3.2 6.2 14420 9390 9500 Ultimate Analyses 55.7 57.4 80.7 5.2 4.9 5.2 29.3 29.9 8.7 0.9 0.9 1.5 0.4 0.4 0.7 6 . 2 8.8 3 .2 100.0 100.0 100.o
--__
201 Lignite
210 Bituminous A Pittsburgh i . ' D , * Pennsylvania Henry Allegheny 11.5 38.0 43.8 6.7 9540
57.2 5.1 29.7 0.9 0.4 6.7 100.0
MI- Me= M A N O M E T E R S
A r r a n g e m e n t of gasification a p p a r a t u s
3.6 32.4 56.8
7.2 11570 67.6 4.4
18.6 1.4 0.8 7.2 1
W
All lots were pulverized in a ball mill t o 93 t o 95% through 200mesh. The Elkhorn coal had medium coking properties, while the others, including the Pittsburgh, which was a n outcrop coal, were noncoking; the Pittsburgh outcrop coal of high oxygen content was tested because of its known low reactivity. A few tests were made on this coal after the addition of 2% by weight of commercial soda ash, containing 587, sodium carbonate; in subsequent discussions Lot 210 coal with soda ash is designated a s 2108. Lot 167 lignite, used in some of the first tests, had been held in storage a t approximately 100' F. for 20 months, with a resulting reduction in moisture content to 13Cr,, from approximately 387, as mined. Lot 201 lignite was air-dried in a rotary dryer a t 220" F. t o 11.5% moisture before pulverization. Lot 196 was a steam-dried lignite of approximately the same moisture content, supplied through the courtesy of L. C. Harrington, Collegeof Engineering, University of North Dakota. Experimental Results The effect of each of the five independent variables-rate of coal feed, steam rate, tube temperature, tube diameter, and added soda ash-on the volume of gas produced has been determined experimentally with one or more of the fuels described above. Preliminary tests made with the Elkhorn coal, Lot 182, were unsuccessful in that the particles became plastic in passing from the feeder to the tube entrance and coked a t the top of the tube into large masses which interfered with steam and gas flow. Much effort was expended in attempting to prevent agglomeration of the particles by modifications of the design of the entrance to the reaction tube. Several designs were tried which employed various types of orifices between the top of the tube and the base of the diffusion cone to shield the entering particles from radiation on the upstream side and then to bring them through the plastic range very rapidly as they entered the orifice. Other designs varied the method of injection of steam, in an effort to protect the internal surfaces of the apparatus from contact with plastic
t
TABLE 11. EFFECT OF FEED RATEAND TUBE DIAMETER ON GAS PRODUCTION FROM LOT 201 LIQNITE
Run No. Tube diameter f t . Tube temperaiure O F. Feed rate lb./hou; Steam &e lb./hour Excess steim, % Gas production, cu. ft./hour a t std. cond. Gas comp., % by volume
cot 01
co
C H4 HP Nz
F.
39 41 42 44 45 46 48 50 49 0.407 0.407 0.407 0.687 0.687 0.687 0.687 0.687 0.687 2098 2092 2089 2064 2047 2064 2077 2051 2025 16.0 19.6 26.0 20.2 21.8 28.2 29.4 38.4 49.4 21.6 21.9 28.1 24.7 26.0 30.8 37.2 48.0 50.2 58 30 26 42 39 27 48 45 18 552
749
793
9.4 7.3 8.6 6.4 1.5 1.5 1.4 0.8 31.5 34.7 32.9 36.1 3.2 2.9 3.2 3.0 50.2 48.0 49.3 49.8 4.2 5.6 4.6 3.9 100.0 100.0 100.o 100.o
649
760
7.8 0.7 33.8 3.7 49.6 4.4
938
949
1150
1264
7.1 10.0 11.5 10.8 1.2 0.9 0.5 0.7 24.4 30.8 29.7 29.3 3.7 3.3 4.0 5.4 49.8 50.4 3.8 4.6 100.0 100.0 100.0 100.0
- - - - ioo.o -----
Total
Gross heating value, B.t.u./ cu. f t . Reaction efficiency yo Gas temperature a't tube exit, Gasification constant
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INDUSTRIAL AND ENGINEERING CHEMISTRY
April 1948
x
10-6
297 78.7
...
90.8
297 76.4
299 67.7
io?
108
,,
309 88.5
.
,
..
94.9
308 86.8
..
,
99.8
311 80.0
.,.
101
coal particles, by directing part of the steam flow across these surfaces by means of slots, tangential jets, vertical jets, or a short coaxial cylinder extending a few inches into the top of the reaction tube, with steam and coal flowing through the center and steam only through the annular space. It was possible to minimize coking t o the extent that continuous operation was feasible, but not to eliminate it completely. Inasmuch as the primary purpose of the investigation was t o determine the gas production capacity of the apparatus, attention was turned to noncoking fuels, and further attempts at solution of coking difficulties were indefinitely postponed. Effect of Feed Rate on Gas Production. Table I1 presents data obtained in a number of experimental runs with Lot 201 lignite at different feed rates in reaction tubes of two diameters. The tube temperatures represent the arithmetic mean of temperatures read a t 15-minute intervals throughout the run by means of njne thermocouples attached t o the outer wall of the reaction tube and spaced along its length. Operation was maintained for periods of approximately 4 hours in the majority of t @ runs. Excess steam was calculated from the rate of coal feed and rate of steam supply. As each pound of carbon requires 1.5 pounds of steam for the water-gas reaction, C H20 = CO H?, the reported value for excess steam represents the steam supplied above the theoretical requirements for complete reaction with the carbon in the coal. The steam rate recorded in Table I1 represents the sum of the moisture in the coal and the steam introduced through the jet a t the top of the tube. Gas production is reported in cubic feet per hour at standard conditions of 30 inches of mercury pressure, 60" F., dry gas. Temperature and moisture content of the gas a t the meter were measured a t intervals throughout each run t o provide the necessary correction factors. The reported gas compositions are averages of a number of 15-minute samples taken at 0.5- to 1-hour intervals throughout each run. The composition of individual samples did not vary significantly from the average composition reported. The reported calorific value of the gas is calculated from the composition. No experimental determinations of calorific value were made. The reaction efficiency is an approximate indication of the percentage gasification of the lignite. It is the ratio of total calorific value of the gas produced to the sum of the calorific value of the lignite fed and the endothermic heat of reaction of the total carbon with steam, calculated for reactants and products both at 60' F. The latter sum represents approximately the maximum potential heat in gas which could be obtained from complete reaction betyeen the solid fuel and steam. It can be seen from Table I1 t h a t the reaction efficiency decreased as the rate of feed of the lignite increased.
+ -
+
297 73.9
301 69.8
308 61.1
1687 94.9
1589 103
1602 92.8
b2T2
The gasification constant given in Table I1 is a measure of the inherent activity of the solid fuel toward reaction with steam. It is useful for the On common basis Of the gasification characteristics of different fuels at one temperature, or one fuel at different temperatures, as it practically eliminates the effect of tube dimensions, fuel feed rate, and minor varia+ions in tge proportion of steam t o each Of which may influence the volume of gas produced experimentally. Values tabulated for the gasification constant were obtained by solving for k after substitution of the experimental data in the semi-empirical equation: = 4 log,
1 r2 f-q- 2
-3r
where k = gasification constant L = heated length of reaction tube, feet d = inside diameter of reaction tube feet c = weight of carbon supplied t o the apparatus, pounds per hour b = total weight of steam supplied t o apparatus, pounds per hour, including t h a t introduced as moisture with fuel T = absolute temperature of steam and gases inside reaction tube, degrees Rankine r = 0.0238 V / b volume of air-free gas produced by apparatus, cubic feet per hour, measured at 60" F., 30 inches of mercury, dry
v=
The form of the equation was determined by making certain arbitrary assumptions as t o the manner in which gasification occurs, and, on the basis of these assumptions, deriving the relation existing between the process variables and the volume of gas produced. These assumptions may be stated briefly as follows: 1. The only materials present in the gasification zone are carbon, water, carbon monoxide, and hydrogen. The reaction H2. occurring is C.4- H20 +CO It is recognized t h a t other reactions are occurring simultaneously t o produce carbon dioxide and methane, but to simplify the mathematical treatment, the principal reaction was assumed to be the only one. 2. The residence time of the particles within the tube is determined by the gas velocity, and the effect of gravity is insignificant. 3. The solid carbon occupies a negligible volume in the tube compared t o the other three materials, which are gaseous. 4. The rate of reaction of carbon with steam is proportional to the total surface of the carbon in contact with the gases, and is also proportional to the partial pressure of the steam present in the mixture of the three gases. 5. The total exposed surface of the carbon in the tube is proportional t o the feed rate.
+
I n applying the experimental data to the equation, the heated length of the reaction tube was taken as 9 feet. The temperature of the gases inside the reaction tube could not be determined precisely. A velocity thermocouple was used in some runs to measure the temperature of the gases within the tube at a location approximately 11 inches below the heated portion of the tube. This temperature varied between 1600" and 1700" F., and was employed in the equation when available. I n cases where the gas temperature was not determined, the temperature of 1650' F. was arbitrarily selected. An uncertainty of 100" in T is relatively unimportant, when the gasification equation is used for comparisons at one tube temperature. Gas compositions given in Table I1 ghow t h a t there was some air present in the samples, possibly from small leaks in the
INDUSTRIAL AND ENGINEERING CHEMISTRY '
550
of lignite 201, as indicated by the gasification constant, is more than 25Yc below that of the other two samples. ,Effect of Excess Steam on Gasification. Three runs were made with lignite 167 a t a constant feed rate of 19 pounds per hour and tube temperature of 2100' F., but with a large variation in the amount of steam injected with the fuel. Table IV shows the gasification constants calculated from the experimental data by means of the gasification equation for each value of excess steam. The apparent activity of the lignite varied with the amount of excess steam supplied. The gas composition also changed considerably, with the percentage of carbon dioxide increasing as the proportion of steam increased. This large variation in the gasification constant reveals the necessity for placing some limitation on the use of the gasification constant for predicting gas production a t various feed rates. The constant has been found to be a measure of the activity of the solid fuel, as determined from the volume of gas produced, but there is an implicit assumption that the composition of the gas does not change materially for the conditions under consideration. With a large excess of steam the secondary, gas-phase reaction between carbon monoxide and steam, CO H20 = CO2 HB,is favored, which is confirmed by the increase in the percentage of carbon dioxide shown in Table 1V. The volume of noneondensable gas is thus increased independently of the inherent activity of the solid fuel. Therefore, the use of the gasification constant is limited to prediction of gas production for a moderate variation of coal-steam ratio from that used in the experimental measurements by means of which the constant is determined. I n all tests, except those on the effect of excess steam reported in this section, the nominal value for excess steam has been 40Yc, although there has been considerable experimental variation from this target value. It has been iound that variations from 20 t o 66% excess steam have relatively little influence on the gasification constant determined from the rvperimental measurements.
+
FEED R A T E , LB. PER HR.
Relation of gas production to coal feed in 8.25-inch tube
Figure 3.
sampling lines or from the tap wat'er used for Jyashing the gas. The measured volume of gas produced was corrected to an air-free basis before being used in the gasification equat,ion. Although the gasification constant varies between the limits 90.8 and 108 X 10-6 for the nine runs of Table 11, it. mas found that gas production calculated by the use of an average gasification constant of 99 x 10-6 agreed n-it11 the corresponding experimental values of gas production within a maximuni error of 5.5% and with a mean error of only 1.0% for the nine runs. This indicates that the average gasification constant is applicable to t'ubes of two diameters over a considerable range of feed rates, and that in all probability some extrapolation to o t'her feed rates and tube dianiet,ers is justified. It is believed that extrapolation to longer tube lengths is also justified. This point is discussed below, when the capacity of a gasification plant of commercial size is considered. Effect of Drying Method on Gasification Characteristics of Lignite. Three lignites of similar composition, as shown in Table I, were studied. Lot 201 was rapidly dried in air a t 200" F., while Lots 167 and 196 were subjected to less dmstic treatment'. Lot 167 was allowed to dry naturally Trhile in storage over a period of months a t an ambient temperature of approximately 100" F., while Lot 196 was dried in an atmosphere of &am.
TABLE 111. GASIFICATION COXSTANTS OF LIGNITES DRIEDBY THREE METHODS Drying Lot lMethod 167 Air-dried, looo F. 196 Steam-dried 201 Air-dried, 220' F.
Gasification Constant X 10-8 Individual tests Average 137 137 124, 133, 137, 131 131 91, 107, 108, 95, 100, 101, 95, 103, 93 99
Table I11 compares the gasification constants as determined for Lot 167 from one run a t 19 pounds per hour, for Lot 196 from four runs at feed rates of 19 to 29 pounds per hour, and for Lot 201 as shown in Table 11. All runs were made with approximately 40% excess steam, at nominal tube temperatures of 2100' F. Similar gas compositions were obtained in all runs. It is immediately apparent that the gasification characteristics are appreciably affected by the drying method, for the activity
Vol. 40, No. 4
TABLE IV.
+
EFFECTOF EXCESSSTEAMON GASIFICATIOX CONSTAKT OF LIGNITE
(Lignite 167: feed rate, 19 Ib. per hour; tube temp., 2100' F.) Excess Steam, Gasification COz in Gas, % Constant X 10-6 % 8 89 4.9 38 137 8,0 142 169 14.6
Effects of Temperature, Type of Fuel, and Catalyst on Gas Production. Table V coniparcs the average gasification constants calculated from runs n i t h three lots of coal a t two different teinperatuies. The gas temperature a t the tube exit in the runs a t 1860" F. averaged approximatell 1310" F., and this ternpeiature was used in the gasification equation for calculation of the constant. The gasification constant a t 1850" F. tube tenipeiature is only 25 t o 35% of the correiponding constant a t 2100" F. for the three fuels tested a t both temperatured. This establishes the lon-er rate of reaction of the fuels a t the lower temperature. It should not be inferred that there is a proportionate decrease in the actual volume of gas produced. Figure 3 shows the gas production plotted against feed rate for all the test runs completed on these fuels, data from which s e r e used in calculating the gasification constants of Table V. The particles of the fuels of low activity remain in the tube for a longer period than those of high activity, because of the smaller amount of gas produced per unit time and the resultant decrease in velocity of flow through the tube. Comparison of the average constants of Table V for Lot 201 lignite and Lot 210 high-oxygen bituminous coal reveals the relatively low activity of this coal in its reaction with steam.
April 1948
INDUSTRIAL AND ENGINEERING CHEMISTRY
TABLE V. VARIATIONOF GASIFICATION CONSTANT WITH TEMPERATURE, TYPEOF FUEL,AND ADDEDCATALYST ~ ~ Temp.,
Lot
Description
O F .
201
Lignite
2100
201 210
Lignite High-oxygen bituminous coal
1850 2100
210 210A 210A
High-oxygen bituminous coal Lot 210, plus soda ash Lot 210, plus soda ash
1850 2100 1850
Gasification d ~ Constant, ~ l x 10Individual runs Average 90.8, 107, 108, 94.9, 101. 94.9. 103. 99.8, 92.8 . 99 38.4, 35.4, 32.9 35 18 7 24 8 34.4 i f . 9 , 24.5, 32.3: 21 s 28 8.f-'8 40.0, 47.7 44 11 11.0
The activity of the outcrop coal is significantly increased by addition of 2% soda ash, as shown by comparison of the constants given in Table V for Lots 210 and 210A. This finding is in agreement with numerous reports in the literature of the effectiveness of sodium carbonate in increasing the activity of solid fuels. Feasibility of Commercial Development 'of Continuous Gasification Process
Estimate of Capacity of Full-scale Equipment. I n considering a n increase in the scale of the eiperiment equipment from pilot plant to a larger size, it was envisaged that a commercial plant might consist of multiple tubes of approximately the same diameter, but of greater length, all suspended in a common furnace, which could be heated by pulverized coal burners. Physical limitations on the construction of the furnace and fabrication of the tubes suggested that a maximum heated length of tube of 27 feet could be provided in the commercial plant. It is believed t h a t the gasification equation can be employed with reasonable certainty for the determination of the gasification capacity of each of these longer tubes, because no regular variation has been found in the experimental values of the gasification constant with a variation in feed rates from 20 t o 49 pounds per hour mith lignite 201, as shown in Table I1 for runs 44 to 50 in the 8.25-inch tube. The absence of any definite trend in the constant indicates t h a t gasification is proceeding according to the equation over the whole range of feed rates, and t h a t the reduction in the percentage of gasification, as shown by decrease of the "reaction efficiency" from 89 t o Sl%, is primarily a result of a decrease in the time of residence of the particles within the tube, as the feed rate is increased. It is therefore inferred that an increase in tube length would have only the effect of increasing the residence time of the particle, without changing the course of the reaction described by the gasification equation, provided the proportion of fuel t o steam is maintained constant. The equation has been used t o estimate the gas production at 2100 F. from %'-foot tubes of 5- and 8.25-inch diameters over a range of feed rates up t o 140 pounds per hour with lignite 201, using 40% excess steam. Figure 4 shows calculated gas production plotted against feed rate. Curves for 9-foot tubes of the same diameters are also shown on the same figure. Estimate of Investment Cost of Full-scale Equipment. As a basis for estimation of plant investment costs, i t was assumed that 27-foot tubes of 8.25-inch diameter would be used, and t h a t each tube would produce water gas of the approximate composition obtained in the pilot plant under the conditions outlined above. From Figure 4 i t can be seen that gas would be produced at a rate of approximately 2300 cubic feet per hour, with 90% gasification of the fuel. It was further assumed that the make gas would be coldenriched with propane to legal standards for city distribution, and the costs of propane handling and storage have been included in the estimate. It was found t h a t the total cost of a plant for production of approximately 12,000,000 cubic feet per day of 554 B.t.u. gas would be approximately $1,100,000 a t the
59 1
time the estimate was made. This estimate was prepared with the cooperation of the engineers of the Surface Combustion Company, Toledo, Ohio. Space limitations preclude a detailed presentation in this paper of the assumptions used in arriving at the estimated cost. T)le significant result is, however, that the investment cost per thousand cubic feet per day capacity is of the order of $82 for the enriched gas, and $65 for the unenriched water gas. These costs compare unfavorably with the author's estimate of less than $40 for conventional carbureted water gas generators, based on authoritative information on two recent installations. The curves of Figure 4 used for estimating the gas production from each tube were founded on the experimental production of gas from lignite, which is known to be highly reactive compared to many coals. The estimate is thus more favorable than could be justified if the gasification of a variety of bituminous coals were contemplated. It was concluded from the unfavorable comparisons based on lignite t h a t further work t o determine the gasification characteristics of additional coals would not be justified, as it seemed unlikely t h a t more highly reactive fuels than lignite would be generally available for use in the process. Estimates of the cost of gas in the holder showed t h a t it was not far out of line with published costs for carbureted water gas, but detailed studies were not carried out because of the unfavorable investment costs. Possible Improvements by Modifications to the Process
There appear t o be three possible avenues of approach to increasing the capacity of equipment for gasification of pulverized coal in a n externally heated chamber. The maximum temperature a t which the alloy chamber can be operated is limited by the physical properties of the alloy. A temperature as high as 2100" F. is not permissible by present
*0 4
O
-
FEED R A T E , L B . P E R H R .
Figure 4.
Calculated effect of feed rate on gas production
A t 2100' F. for varying tube dimensions with lignite 201 and 4U% excess steam
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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
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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.
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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.
