I
C. G. VON FREDERSDORFF Institute of Gas Technology, Chicago, 111.
Process for Coal Hydrogasification Methane separation and hydrogen recycle minimize material requirements in manufacturing high-B.t.u. gas from coal T 2 . o processes for making high-B.t.u. gas are being investigated: hydrogasification of coal or oil shale under pressure, and catalytic conversion to methane of synthesis gas from coal. The catalytic method is more advanced technically, but coal hydrogenation is considered more acceptable. I t requires less oxygen, eliminates the need for sulfur purification except in the product gas, and improves thermal efficiency by reducing exothermic heats of reaction. When hydrogasification is used for making pipeline gas. cost of hydrogen is a major factor. A proposed modification of this process reduces the amount of hydrogen needed. Methane is separated by liquefaction, the recovered hydrogen is recycled, and, to achieve additional steam decomposition, steam is admitted into the hydrogasification zone. \Vith oxygen admission prior to hydrogasification, the process approaches thermal balance and becomes independent of an external hydrogen source.
fed at the top for fixed-bed, or at the bottom for fluid-bed operation (see diagram). After preliminary cleaning of dust, the raw gas (mostly methane, hydrogen, carbon monoxide, and dioxide, together with unreacted steam) is passed to the catalytic carbon monoxide shift, cooled in the hydrogen heat exchanger, and scrubbed with water. Prior to methane recovery, carbon dioxide and sulfur-bearing gases are removed, and after heat exchange the warm hydrogen is recompressed for recycle. This process may not require a carbon monoxide shift, in which case the recycled gas is a hydrogen-carbon monoxide mixture. Some carbon monoxide shift would presumably occur in the reactor because moderate excess steam would be present at the point where recycled gas is introduced. Thermodynamic Analysis
I n the first part of this analysis, the carbon-hydrogen-steam system, defined by: C C CO
Process
It is assumed that steam superheated to F. is use? and that hydrogen is recycled at 800 F. Solid fuel may be 1000'
+ 2Ha
+ H20
+
+
+ Hz0
CHI
CO
+
+ H?
COS
+ Ha
(11
(2)
(3)
is represented in Figure 1 in terms of equilibrium concentrations, hydrogen availability, and net heat effect, assuming
recycle gas as hydrogen only. The ratio X takes into account the hydrogen equivalent of carbon monoxide:
x = (1 - nH,O
x' = x + 0.336C/2nc~, (corrected for coal) (4a) The recycle process is independent of external hydrogen if X' 2 1.0. The net heat effect, AHR, in B.t.u. per pound-mole of methane represents the heat balance around the individual reactions : AHR = AH1 f
-
ncop
[l/nCH,]
-
ClHc
-
800'F
He RECYCLE
100OoE STEAM
(4)
where n is computed from available equilibrium constants (8). A corrected ratio, X',for actual fuels, based on data for bituminous coal containing (dry basis) about 75% (weight) carbon and 5% (weight) hydrogen, indicates that approximately 80% of the coal hydrogen content appears as mc-thane and higher hydrocarbons (2). Approximation of the moles of hydrogen available from coal for hydrogasification per mole of inlet steam is 0.8 CCf. With the above coal and 9570 carbon gasification this becomes 0.336 C. Then
[(I
High B.t.u. gas is made by separating methane from gases produced by coal g a s i f i c a t i o n with steam and recycled hydrogen. Either a fixed or fluid bed can b e used in the process
+ nCO)/2nCH,
X
nH,o)AHz f n c o , A H ~
+
+ ~ A H Hf, A H H , ~ ](5)
The ratio a: a = nHz
+ nH,o f
-
~ C H ,
1
(6)
may be corrected for hydrogen available from the above coal by the approximation : a'
= a
- 0.336C
(6a)
With coal a t 1700' F. (1200' K.), and at X' = 1.0 (Figure l), the process is endothermic by 35,000 B.t.u. per pound-mole of methane and at 1880" F. (1300' K.) by nearly 50,000 B.t.u. per pound-mole of methane. The endothermicity decreases with increasing hydrogen to steam inlet molar ratio; however, the X' ratio also decreases. Thus, a t high hydrogen to steam ratios the process approaches thermal balance but requires external hydrogen; a t low hydrogen to VOL. 52, NO. 7
0
JULY 1960
595
based on 4 standard cubic feet (SCF; a t 60" F., 30 inches of mercury absolute pressure, and saturated with water vapor) of carbon monoxide plus hydrogen in synthesis gas being equivalent to 1 SCF of methane by methanation. Data for Lurgi pressure gasification were adjusted for the methane content of the raw gas.
steam ratios the process has sufficient hydrogen but lacks heat. This imbalance between thermal and hydrogen quantities prevails at all possible gasification temperatures. The heat deficiency can be supplied by any of three methods: greater preheat of reactants than employed here, internal heat exchange with a high-temperature fluid, or partial combustion of coal within the reactor. The third alternative is the most attractive because of the large deficiency and the need for rapid heat transfer. I t provides a high-temperature zone where a part of the inlet steam is rapidly decomposed-steam decomposition is relatively slow at temperatures which favor high rates of methane production. Because equilibrium steam decomposition cannot be achieved in practice, the system with oxygen admission is treated in the second part of this analysis (Figure 2) on the more realistic basis of equilibrium in Reactions 1 and 3, with the reaction c 0 2 + CO? (7)
Rough cost estimates indicate that the saving in oxygen and fuel over the Lurgi process about equals the incremental cost of methane separation. Reactor Size Estimation of the hydrogasification fuel bed depth is based on differential
1 7 0 0 O F(1200O K )
1880'
HEAT EFFECT FOR PURE CARBON WITH 8 0 0 . ~ H I , 1 0 0 0 . ~S T E A M
B
F ( 1300" K )
HEAT EFFECT FOR PURE CARBON W I T H B00.F l i t . 1 0 0 0 . F S T E A M
--
100
1 -
1 FOR PURE CARBON
40
40
30
30
20
20
IO
IO
+
going to completion and with arbitrary assignment of percentage steam decomposition. T h e hydrogen availability test, Equation 4 or 4a, applies. The process becomes thermally balanced or slightly exothermic to compensate for heat losses by selection of the proper oxygen to steam inlet molar ratio. T h e net heat effect is now:
0
0
0
05
15 20 25 30 I N L E T MOLAR RATIO
IO
H,/HzO
35
0
05
I O
nz/nzo
I 5 20 2 5 3 0 I N L E T MOLAR R A T I O
3 5
Figure 1. Increasing the hydrogen to steam inlet molar ratio in hydrogasification at 200 atm. increases equilibrium methane and hydrogen concentration, decreases hydrogen availability to consumption ratio, and decreases heat requirements
AHR
+ [ l / n c a , ] [ ( l - nco, - ~ H ~ AoH)z + nco2AH8 4-bAH4-k CAHc +
AH1
aA",
-k bAHo, f AHH,O] (8)
At conditions noted in Figure 2 and at low hydrogen to steam inlet molar ratios, X or X' decreases from values greater than 1.0, and the net heat effect is moderately exothermic. At 30% steam decomposition (Figure 2 , A ) the process becomes independent of external hydrogen if the hydrogen to steam inlet molar ratio is no greater than 1.8, the point where X' approaches unity. At this point AHR is practically zero, the equilibrium wet gas methane content is about 15Tc (21Yc dry basis), and the process requires 3.5 moles of hydrogen recycle per mole of methane. At 40% steam decomposition (Figure 2,B) the process is hydrogen sufficient at 2.5 hydrogen to steam inlet ratio, AH, is about zero, and the needed hydrogen recycle is 3.4 moles per mole of methane. This demonstrates operating conditions where the proposed process with oxygen admission appears thermodynamically feasible. By virtue of potentially greater thermal efficiency: oxygen and coal requirements per thousand cubic feet (MCF) of methane (Table I) are less than those of other gasification processes (Table 11). The data in Table I1 are
596
S T E A M DECOMP.:
40%:
Oz/HzO~0.15
HEAT EFFECT F O R PURE CARBON WITH
,
I
-2001 4
3
I I
I
I
I
I
'
I
I
ALLOWANCE FOR /
2
I
1
ne I N C O A L
!
1
I
FOALREA >;
50
40
30
20
IO
-
0
2
COS
01 0
0.5
1
1
1
1.0
1.5
2.0
0 2.5
n 2 / ~ 2 0 I N L E T MOLAR R A T I O
3.0
3.5
a
0 5
IO
Hz/HpO
I 5 20 25 I N L E T M O L A R RATIO
30
3 5
Figure 2. In hydrogasification a t 50 atm. and 1700" F., the effect of increasing the hydrogen to steam inlet molar ratio is to approach a balance in thermal and hydrogen requirements a t equilibrium
INDUSTRIAL AND ENGINEERING CHEMISTRY
HIGH B.T.U. fluid-bed rate data for lowtemperature Disco char, 65 to 100 mesh, with hydrogen-steam mixtures at 1 to 30 atm. and 1700’ F. (70). Methane production and total carbon gasification rates, moles of carbon gasified per mole of carbonminute, irere determined for pure hydrogen and for hydrogen to steam inlet molar ratios from 0.1 to 2.0 at carbon burnoffs of 0 to loo$&. The total carbon gasification rate is the sum of the rates of methane production and steam reactions with carbon. These reported data indicate that at the same hydrogen partial pressure, steam effectively increases the methane formation rate; the increase, being pressure dependent, ranges from 100 to 2 times the rate in a pure hydrogen atmosphere at 1700’ F. and 1 to 30 atm. total pressure. The methane rates with Disco char are not sufficiently high at pressures below30 atm. and temperatures below 1700’ F. to be of practical interest. Rate plots for 50 atm. (Figure 3) were prepared by plotting the reported rates against pressure (1 to 30 atm.) on
logarithmic scales, for each hydrogen to steam inlet molar ratio investigated (0.32, 1.O, and 2.0) at 0 to 100% carbon burnoff. The resulting nearly straight lines were extrapolated to 50 atm. and
Table I.
Theoretical Material Requirements (Reactor Only) Are Quite Low for Coal Gasification by Oxygen-Steam and Hydrogen Recycle 1000’ F. inlet steam; 800’ F. hydrogen, 60’ F. oxygen and coal, outlet temperature 1700’ F. Pressure, Atm. 50
O?/H20 inlet molar ratio H 2 / H 2 0 inlet molar ratio Inlet steam decomp., % RequirementsIMCF CHI : Pure C, Ib. Coal, at 95% gasification, Ib. Steam, Ib. 02, loo%, SCF Theoretical C appearing as CH4, 0/7”*
Hz recycle, moles/mole CHa a
0.15 2.4 40
0.1 1.8 30 67.5” 95“ 111 25Cia
.200
68‘ 96n 79 273‘
43 3.5
0.1 0.75 30
0.1 1.0 40
63 88 110 230
64.5 90.5 151 318 49 1.8
43 3.4
GAS
the extrapolated points replotted in the desired form. The linear methane rate plot (Figure 3,A) was extended (dashed lines) beyond the range of hydrogen to steam inlet ratios investigated. The total carbon gasification rate plot (Figure 3,B) was also extended by continuing the trend of the curves (dashed lines) to a 4.0 hydrogen to steam inlet molar ratio. The relationship for computation is the material balance over a differential bed height, dh, in which incremental methane, dn,, is formed: dn,
= rm dh =
-qqFndB
(9)
Assuming uniform temperature and piston flow, this relation is integrated graphically over the limits of the hydrogenation zone, h = 0, where burnoff is Bo, and h = Zwhere B = 0:
51 1.7
Increased by 10% over theoretical values.
Table It. Material Requirements (Coal Gasifier Only) for Methane Production from Synthesis Gas in Current Processes Are Largely Dependent upon Operating Conditions and Methods of Fuel Contacting Process‘ USBM ( 5 ) PAN Coal type
Bit.
PROD
(4)
Bit.
L (5)
ITGS ( I )
Bit.
Char
Coke
7.5 2500 0.45 1500 100 31 92
21 1100 0.77 500 500 34 97
K-T ( 6 ) IGT (7) Bit.
Pressure, atm. 21 1 1 Outlet temp., F. 2500 2350 2300 0 ~ / H 2 0inlet molar ratio 1.5 0.65 0.81 Inlet steam temp., ’ F. 625 930 250 625 930 100 Inlet 02 temp., O F. Inlet steam decomp., yo 25 25 25 90 84 85 C gasified, % Requirements/MCF CHg : C, Ib. 96.5 119 112 Coal (or coke), Ib. 137 185 168 Steam, lb. 42 85 77 0 2 , loo%, SCF 1332 1060 1444 C appearing as CHI by methanation, 7‘ 33 26.7 28
93.5 125 131 1248
84 112 360 584
(9)
Barley anth. 1 1 1200 880 0.38 0.29 210 285 210 285 54 52 99.5 99.4 100 111
152 1144
At fixed temperature and pressure the quantities rm,r I ! and q depend upon carbon burnoff and the composition of the reacting gases. The effect of gas composition in rate hindrance is approximated by the ratio of the actual partialpressure product of the reacting components to the equilibrium constant of the particular reaction. By use of this concept, the recommended forms (70) of the rate equations are:
87 110 154 951
34 38 32 36.6 USBM = E. S.Bureau of Mines: P-lK = Panindco process; K-T = Koppers-Totzek gasifier; IGT = Institute of Gas Technology; L = Lurgi unit: WGS = water gas set; PROD = producer.
where r and p and y : t,
rt
X ’ are related through = 60 p v R , ’ / l Z
60 p y R , ’ / I Z
VOL. 52, NO. 7
JULY 1960
(13)
597
For practical purposes the product py is considered constant through the hydrogenation zone, for fuels of normal ash content, and Equation 10 becomes:
Fixed-Bed Reactor. The integral in Equation 14 is the area under a curve of 1 / R f Lus. B. The outlet gas composition is taken from Figure 2,B at X’ = 1.O and a = 2.5, or 47.2% hydrogen. 17.9% undecomposed steam, 13.57” carbon monoxide, 3.5y0 carbon dioxide. and 17.9% methane. hlaterial balances are incorporated to relate the variation of gas composition with burnoff (Figure 4) based on 80% of the hydrogen content of coal appearing as methane, weighted linearly with carbon burnoff; maintenance of carbon monoxide shift equilibrium; and 10% incremental steam decomposition in the hydrogenation zone. Here the exothermicity of the oxidation zone is approximately balanced by an initial 30% steam decomposition. Beyond the hydrogen recycle point the exothermicity of coal hydrogenation is approximately balanced by incremental steam decomposition and sensible heat transfer to coal and hydrogen. LVith y ’ = 1.0 and y = 1.1 [which is justified on the basis that, given sufficient reaction time, the equilibrium methane content based on carbon as @-graphite may be exceeded (2) with carbon as coal], graphical integration yielded 86.25 Fo/5pJ. With practical values of Fo = 20 pound-moles of carbon per hour-square foot, p = 40 pounds per cubic foot, and y = 0.75, this becomes Z = 11.5 feet of hydrogenation zone at 1700’ F. and 50 atm. This is applicable for 65 to 100 mesh fuel particles. A fixed bed would operate with perhaps 0.25-inch average size fuel. Some correction is required for the effect of fuel size, but necessary data are lacking. The assumption of 10% steam decom-
u
position in the hydrogenation zone checked out at 11.5% by graphical integration of Equation 9. Fluid-Bed Reactor. With assuniption of the reactor completely stirred with respect to coal, and other conditions the same as above, Z = 102.9 Fo/5p!, which: with practical values of Fo = 11.5. p = 35, and J = 0.75, becomes Z = 9 feet. This again is applicable only for 65 to 100 mesh fuel particles. The assumption of 10% steam decomposition in the hydrogenation zone checked out at 12.1%. Afore than 9 feet in height \rould be required for assumption of complete mixing of the gas stream. Under these assigned conditions the hydrogenation zone fuel bed depth for fixed- and fluid-bed operation appears reasonable for a practical system. Hoivever, the fixed-bed estimate must be viewed with reservation. Information is insufficient to correct for effects of particle size and floiv characteristics. together with gas-contacting efficiencies between fluid and fixed beds. Discussion
kfethane separation and hydrogen recycle offers an attractive approach for high-B.t.u. gas production, provided there are no excessive counterbalancing costs. If operated as a fixed-bed process, it presents an adaption of the Lurgi pressure gasifier modified for hydrogen recycle to the upper portions of the fuel bed. If operated as a fluidized process, it does not differ in principle from fluidized-hydrogenation, residualfuel gasification schemes proposed or under investigation. ’4single-vessel fluid bed, where both oxygen and hydrogen are introduced, is disadvantageous by virture of rapid fuel mixing, resulting in some volatile matter being burned by oxygen. Because the volatile portion of the fuel is the most readily hydrogenated component, an effective fluid-bed system would require a separate reactor to complete the volatile-matter hydrogenating reactions and a second vessel. immediately below the first: to gasify residual carbon with oxygen and steam. The hot gases from the second vessel would be admitted to the hydrogenation reactor. An alternative to methane separation in this proposed process is catalytic methanation of carbon monoxide and hydrogen in the raw gas, yielding gases of high methane content but requiring an external source of hydrogen.
20
rn
2
Nomenclature
10 0 100
90 80 70 6 0 5 0 40 THEORETICAL CARBON
30 20 10 C BURNOFF, %
Figure 4. The variation of gas composition through a fixed fuel bed operated with oxygen-steam and hydrogen recycle is traced by calculation
598
b
= hydrogen-steam inlet molar ratio = oxygen-steam inlet molar ratio
B
= fractional carbon burnoff
c
= hydrogen-carbon
a
C
INDUSTRIAL AND ENGINEERING CHEMISTRY
molar ratio in coal = equilibrium carbon gasified, mole/mole of inlet steam
f F
= fraction of carbon gasified
= carbon input, 1b.-mo1e’hr.q. ft. = bed height, ft.
h AH
= enthalpy of reactant, B.t.u. Ib.-
mole, between inlet and outlet temperatures AHi, AHz. AH3 = heat of Reactions 1, 2, 3 AH = net heat effect, B t.u /lb. ’mole methane KIJY? = equilibrium constants for Reactions 1, 2 I) = equilibrium product, rnolrs! mole inlet steam p = partial pressure, arm. 4 = rm r = reaction rate. lb -mole hr.-cu. f t . reactor = reaction rate, lb. mole,,’lb.-mole R carbon-min. R’ = reaction rate corrected for hindrance, Ib. mole/lb.-mole carbon-min. .Y = ratio of hydrogen available to that consumed S’ = X corrected for actual fuels I = weight fraction of carbon in fuel 2’ = reaction zone height, ft. y = activity of carbon in hydrogenation reaction y’ = activity of carbon in steam reaction p = fuel bulk density
Subscripts m = methane formation t = total carbon gasification Acknowledgment
Adaptation of the hydrogen recycle in coal hydrogenation investigations was suggested by E . L . Tornquist, Xorthern Illinois Gas Co. Literature Cited (1) , - , Blatchford. J. LV., A m . Gas. Assoc. Proc. 1950, pp. 652. ~
~
(2, Channabasappa, K. C., Linden, H. R.: IND. ENG.CHEM. 48, 900 (1956). (3) Cooperman, J., Davis, 3. D., Seymour, W.? Ruckles, W. L., U. S. Bur. Mines Bull. No. 498 (1951). (4) Foch, P., Loison, R., “International Conference on Complete Gasification of Mined Coal,” p. 224, R. Louis, Brussels, 1954. (5) Strimbeck, G. R., Cordiner, J. B., Jr., Taylor, H. G., Plants, K. D., Schmidt, L. D., U. S. Bur. Mines RI 4971 (1953). (6) Totzek, F., Chem. Eng. Progr. 50, 182 (1954). (7) von Fredersdorff, C. G., Pyrcioch, E. J., Pettyjohn, E. S., Inst. Gas Technol. Research Bull. No. 7 (1957). (8) Wagman, D. D., Kilpatrick, J. E., Taylor, W. J., Pitzer, K. S., Rossini, F. D., J . Research Natl. Bur. Standards 34, 143 (1945). (9) Wright, C. C., Newman, L. L., Am. G a s . Assoc. Proc. 1947, pp. 701. (10) Zielke, C. W., Gorin, E., IND. ENG. CHEM.49, 396 (1957). for review December 21, 1959 RECEIVED ACCEPTED May 9, 1960
Division of Gas and Fuel Chemistry, 136th Meeting, ACS, Atlantic City, N. J., September 1959. Work sponsored by American Gas Association.
