586
Ind. Eng. Chem. Process Des. Dev. 1900, 19, 586-592
i = component designation index: 1,NH3; 2, N,; 3, H,; 4, Ar; 5, CH,
J = defined by eq 34 j = component designation index (same as i) n = total number of points, eq 25 ni,nj = number of moles of component i or j P = pressure, atm P, = critical pressure, atm P,,,,Pcj = critical pressure of component i or j , atm P,,, = critical pressure of vapor mixture, atm Pi' = saturation pressure of NH3, atm Pi = pressure of component i when occupying the total volume at the given temperature R = gas constant, (atm cm3)/(g-mol K) T = temperature, K T , = critical temperature, K T,,i,T,, = critical temperature of component i or j , K Tc,, = critical temperature of mixture, K T R ,=~ reduced temperature of NH3 u = molar volume, cm3/g-mol uc,+, = critical molar volume of component i or j , cm3/g-mol u1 = molar volume of saturated liquid NH3, cm3/g-mol u, = molar volume of the mixture, cm3/g-mol D l = partial molar volume of NH3, cm3/g-mol Dim = partial molar volume of component i for infinite dilution in NH3, cm3/g-mol D y = partial molar volume of component i for infinite dilution in NH3 at saturation temperature of NH3, cm3/g-mol x i , x j = mole fraction of component i or j in the liquid phase y, yj = mole fraction of component i or j in the vapor phase yyll,c;ypbl,, = calculated and measured equilibrium concentration of NH3 in the vapor phase, respectively, at pressure , P i at a chosen temperature, eq 24, mole % ~ ' 1 y'l,, , ~ ; = calculated and measured concentration of NH3 in the vapor phase, respectively, at different pressures and temperatures, eq 25, mole % 2 = compressibility factor
Greek Letters ~ u i i , ~ ;u i j = , ~ self-interaction
constant of molecule i or j in the environment of molecule 1 Pls = saturated liquid compressibility of NH3 pima= partial compressibility of solute i at infinite dilution y1 = activity coefficient of solvent NH3 y*i = activity coefficient of solute i ~ 1 - t 4 = prescribed constants $ ~ i = fugacity coefficient of component i R, = parameter in eq 12 R,, = value of Q, for NH3 fib = parameter in eq 13 a b , = value of for NH3 w = acentric factor
Literature Cited Alesandrini, C. G., Lynn, S., Prausnitz, J. M., Ind. Eng. Chem. Process Des. Dev.. 11. 253 (1972). Ammonk Casale S.A., Italy, Fig.No.TD 48. Chueh, P. L., Prausnitz, J. M., AIChE J., 15, 471 (1969). Dodge, 8. F., "Chemlcal Engineering Thermodynamics", p 197, McGraw-Hill Book Co., Inc., New York, 1944. Guerreri, G., Prausnitz, J. M., Process Technol. Int., 18(4/5), 209 (1973). Gunn, R. D., Prausnitz, J. M., AIChE J., 4, 494 (1958). Husaln, A., Gangiah, K., "Optimization Techniques for Chemical Engineers", pp 27-31, Macmillan Co. of India Ltd., Delhi, 1976. "International Critlcal Tables of Numerical Data, Physics, Chemistv and Technology", McGraw-Hill Book Co. Inc., New York, 1926. Joffe, J. Ind. Eng. Chem., 39, 837 (1947). Joffe, J., rnd. Eng. Chem., 40, 1738 (1948). Nelder, J. A., Mead, R., Comput. J., 7, 308 (1964). Nielsen, A., "An Investigation on Promoted Iron Catalysts for the Synthesis of Ammonia", pp 28-29, Jut. Gjellerup Forlag, Copenhagen, 1956. Prausnitz, J. M., Chueh, P. L., "Computer Calculations for High Pressure Vapor-Liquid Equilibria", pp 82-90, PrenticeHall, Englewood Cliffs, N.J., 1966. Prausnitz, J. M., "Molecular Thermodynamlcs of Fluid-Phase Equilibria", p 156, Prentice-Hall, Englewood Cliffs, N.J., 1969. Redlich, O., Kwong, J. N. S., Chem. Rev., 44, 233 (1949). Wada, Y., J . Phys. SOC. 4, 280 (1949).
Received f o r review September 25, 1979 Accepted June 18, 1980
Global Model of Countercurrent Coal Gasifiers Phlllp G. Kosky' and Joachlm K. Floess General Electric Research and Development Center, Schenectsdy, New York 1230 1
This is a model of a fixed-bed coal gasifier in which CO, C02, HO , and H, are assumed to be in thermodynamic shift equilibrium over a zone in which the primary gasification reactions occur. Exiting temperatures from this zone are in excess of 550 O C and the shift reaction is readily catalyzed by gas-borne impurities. Fresh coal is pyrolyzed in this gas stream and its gaseous products are added quantitatively to the shift gases. The final raw product gases thus calculated are close to experimentaldata from several sources for oxygen- and air-blown gasifiers. The model, which is simple conceptually and mathematically, correctly predicts the effect of heat leak in establishing the composition of the raw coal gas from a fixed bed gasifier. This important variable has not had the visibility that its significance demands.
Introduction It is possible to model coal gasifiers by assuming detailed knowledge of local chemistry and kinetics (Arri and Amundson, 1878; Amundson and Arri, 1978; Yoon et al., 1976,1978; Biba et al., 1978). The gross picture of a fixed bed (i.e., slowly descending) gasifier and reactant gases is as Figure 1. The reactor is presumed to have the approdimate zones shown. The coal which enters the reactor is rapidly heated by the upflowing gases. The coal is devolatilized in the process-which is to say that the chemical bonds of the original coal macerals are broken producing 0196-4305/80/1119-0586$01.00/0
gases as H2, CO, COz, CH4, H2S, NH,, and water, trace gases, some oils, and tars (Figure 2). A residue of coal char remains. This material is mostly carbon ("fixed carbon"). At the temperatures achieved in a fixed bed gasifier the devolatilization is near completion but the possibility remains that some small mass of hydrogen is retained in the char (Juntgen and van Heek, 1968). On a molar basis the hydrogen may not be quite negligible compared to the fixed carbon. The devolatilization process is ill understood and is is unwise to be dogmatic about the distribution of devolatilized gases from this zone. Their concentration 0 1980 American Chemical Society
Ind. Eng. Chem. Process Des. Dev., Vol. 19, No. 4, 1980 587
Table I. Volatile Matter from Devolatilization‘
COAL
4
wt % water wt % tar and oil (as CH) wt % coal gas
ZONES DEVOLATI L l Z A T l O N
23 20 57
Composition of Devolatilization Gas volume % CO 20.6 6.1 volume % CO, 13.1 volume 5% H, volume % CH, 50.3 volume % H,S 9.9 volume % NH,
REDUCTION (GASIFICATION)
t
COMBUSTION GRATE
I
ASH
a IBLAST
TEMPERATURE
heterogeneous catalysis due to ash or coal mineral particles (Desai and Wen, 1978; Walker et al., 1937). In any case, it is usually much faster than any of reactions 1, 2, or 3 (Ergun and Mentser, 1965). This above discussion is not to imply that the elementary reactions obey simple rate laws or that their rates are not dependent on undisclosed parameters-for example, one elementary step in the homogeneous water-gas shift reaction, reaction 4, involves the concentration of [OH] species (Bamford and Tipper, 1977). The evidence is, however, that the shift reaction is readily catalyzed so that in industrial equipment at “sufficient” temperature it is at thermodynamic equilibrium. Below the gasification zone one presumes that oxidation occurs to CO or COz
GASES
UNBURNT CARBON
Figure 1. Physical model of gasifier.
HzO, Hz, CO,
1
GOz, HzS. NH3, CHI, TAR. OIL
I
rul
COAL
II I I I I I
RAW GAS
i
Loison and Chauvin (1964); Yoon e t al. (1978).
CONTROL VOLUME DEVOLATlLlZATlON
c + ‘/202 I
I
t
ASH
(5)
(6) with CO favored a t high temperatures (Arthur, 1951). It seems also plausible that if any hydrogen is formed in the combustion zone by decomposition of steam with carbon, it may combine with oxygen. H2
I BLAST GASES
is strongly dependent on local conditions during devolatilization (Anthony and Howard, 1976). The char thus produced is then slowly consumed in a series of heterogeneous reactions with steam, hydrogen, and carbon oxides. This zone, probably the largest zone in the reactor, is known as the gasification (reduction) zone. Heat for this zone is generated by combustion of the char with the feed oxygen in an intense combustion zone. The speed of this reaction restricts it to a narrow zone, the heat from which is gradually consumed by the endothermic reactions in the gasification zone. While there is no unique set of reactions which may be implied by considering the exiting gas from a coal gasifier, there seems to be general agreement amongst several investigatprs (Desai and Wen, 1978; Arri and Amundson, 1978; Biba et al., 1978; Yoon et al., 1976, 1978) as to the primary reactions. These are
CO
co
c + 0 2 = c02
VOLUME SHIFT REACTION
Figure 2. Global model of gasification.
C
=
+ HzO s CO + H2 c + coz s 2 c o
(1) (2)
C! + 2H2 e CH,
(3)
+- Hz0 e C02 + H2
(4)
which are presumed to occur in the gasification zone. Reaction 3 is generally quite slow even in comparison to the other heterogeneous reactions 1 and 2. It may be presumed that reaction 4 is pseudohomogeneous in that it proceeds between gaseous components but probably with
+ ‘/202
=HZW
(7) Below the combustion zone an ash (or slag) layer exchanges some heat with the incoming blast gases before exiting at the grate. Both the Lurgi gasifier (von Fredersdorff and Elliot, 1963) and the Morgantown Energy Research Center gasifier (Desai and Wen, 1978) have a water jacket. In the commercial Lurgi gasifier some of the steam for the blast gases is generated in this jacket. In the pilot-scale, advanced fixed bed GEGAS-D gasifier (Woodmansee and Palmer, 1977) the vessel is insulated and the requisite steam is entirely produced elsewhere. This points to an interesting question: What is the effect of the radial heat leak on the exit conditions from the gasifier? Different coals ranging from lignites to anthracites produce quite different quantities of volatile matter (Juntgen and van Heek, 1965) and depending on pyrolysis conditions quite different coal gas from the thermal decomposition of the coal (Friedman et al., 1968; Anthony and Howard, 1976). The complexity of the pyrolysis processes can be qualitatively determined from the Anthony and Howard (1976) review. Rather than realistically attempt complete volatiles prediction for all coals a typical devolatilization gas analysis was selected. The procedure emulates Yoon et al. (1976, 1978), who basically used the coal gas analyses of Loison and Chauvin (1964). The data are reproduced in Table I. The effects of such an assumption are relatively important and will be discussed in a later section. Moreover, the standard heat of reaction(s) of coal char pyrolysis products
-
+
will be assumed to be zero, following Yoon et al. (1976,
588
Ind. Eng. Chem. Process Des. Dev., Vol. 19, No. 4, 1980
1978). This affects only the temperature of the exiting, “raw” gas. One important aspect of the operation of a coal gasifier needs to be considered. The coal feed rate is dependent on the chosen blast gas composition, rate, enthalpy, and pressure as well as its intrinsic reaction rate with the coal. More specifically the coal consumption rate for a given coal is determined not by the feed mechanism, but by these chosen blast gas conditions. Sufficient coal can be added only to maintain the coal bed height for steady-state operation. In practice, however, the operator of a gasifier does have one additional control strategy-to adjust the grating rate for removal of the bed from the gasifier shaft; in a slagging gasifier this option is automatically eliminated. Where such an option does exist, the constraint on grating is to prevent excessive carbon loss in the ash and to maintain reducing conditions in the gas. Coal feed rate is therefore mostly dependent, as it must be, on the rate or reaction of the coal with the reactant gases. This is exemplified by the data reported by Kydd (1975), in which the effect of inlet blast gas enthalpy is shown to increase the heating value of the coal gas while maintaining all other conditions constant except for additional coal feed. It is there reasoned that the consumption of coal in the gasification zone was controlled by the heat generated in the combustion zone. Therefore it appeared reasonable to assert proportionality between the coal feed and enthalpy of the blast gases of a given composition which produce the combustion products. This concept will be recalled later. There are extant “equilibrium” models of gasification in the literature of which the prototype is that of Gumz (1950). Gumz assumed heterogeneous equilibria between carbon and steam, carbon and carbon dioxide, carbon and hydrogen (reactions 1-3, respectively) with some empirically specified “approaches” to equilibrium in those reactions which do not actually achieve “full” equilibrium. Woodmansee (1976) described a similar thermodynamic model as did Desai and Wen (1978). It is hard to imagine the slower heterogeneous reactions achieving equilibrium and harder yet to allow physical significance to the “reaction temperature” a t which these reactions yield equilibrium. Such models do have an advantage, however. They are capable of computing a feed rate of coal to be consumed at a specified blast condition. In recent simultaneous oral presentations both Denn (1978) and Kosky (1978) have recognized the insensitivity of exit raw gases from a coal gasifier to the kinetic conditions assumed in the heart of the gasifier. Denn et al. (1979) have further amplified this observation in a kinetics-free model of coal gasification and have showed that it agrees closely with a kinetics model except where the direct hydrogenation reaction, eq 3, is important. This kinetics-free model is multifaceted with specific assumptions made for modeling the combustion zone, shift equilibrium and in consideration of gasification reactions 1 and 2. For these reasons the model, which follows, does not include consideration of any heterogeneous reaction and has only the capability of calculating coal feed rate based on the proportionality suggested by Kydd (1975). Since this is only proportionality, without at least one experimental point, it cannot predict coal consumption rates but merely codify changes will occur with different blast enthalpies.
Model Basis The whole of the chemistry in the reactor is ignored in this model except as follows: (1)The gases which emerge on the top of the coal char bed are in water-gas shift
equilibrium (eq 4) at the exit temperature from the gasification zone. (2) Added to this gas are the products of devolatilization of the coal. The model is exemplified in Figure 2, in which the whole basis of the model is described. The oxidation and gasification zones are considered together; it is assumed that char from the devolatilization zone enters the reactor a t some unknown temperature T*, this char to be of known composition in carbon and hydrogen. This can be determined from the ultimate (Le., elemental), analysis and proximate analysis of the coal less the assumed products of devolatilization. It is further assumed that the exiting gases are at r* and that the CO, C 0 2 ,H2, and H 2 0 are in the ratio as determined by the thermodynamics of the shift reaction, i.e.
where
{E:/
K(T*) = 0.0265 exp -
There are thus five unknowns in the system, T* and NCi, i = CO, C02, H2 and HzO. The equations which have to be solved are the carbon balance (assuming the coal feed can be specified for a given blast gas rate), the oxygen and hydrogen balances, the overall energy equation plus eq 8 representing the shift equilibrium assumption. These equations are (9) Nc = Nco + Nco2 NHz = “20’
- “20
(10)
+ ‘”2
in which all of the oxygen is consumed in the combustion zone. The trivial amount of C02 ingested with the original blast gases is ignored; the inerts, if present (N, and Ar), are included in the energy balance but of course take no active part in the shift reaction. The overall energy equation can be expressed as CNG/’CPGi(T* - TB)- MsCs(T*- TB) Q = W H-~b”,)AH, - Nco2AH6- NcoAHb (12) in which AH, is the combustion heat of the reaction of eq j a t the temperature T*. At this point it is noted that a small error is incurred in assuming that the residue a t a mass rate M s exits at the incoming blast temperature TB. It may be significantly hotter, but it represents only a small overall enthalpy term. Since N H z , Nco2, and Nco are all unknowns it is convenient to rewrite eq 12 using eq 9, 10, and 11. By simple algebraic rearrangement the energy equation can be expressed in terms of the reaction heats at T* of the shift, Boudouard, and carbon monoxide combustion reactions CNGi”CpGi(T*- T B ) - MsCs(T* - TB) + Q = -NH200(l- X)AH4 - NcAH2 - 2No,OAH14 (13)
+
where X = NHp/ NH200,the fraction undecomposed steam. Also co + ‘ / 2 0 2 = c02 (14) Note now, given Nc, eq 13 contains just two unknowns, T* and X. Proceeding in a like manner eq 8a can be expressed as X2NH2o0(l- K ) + X[Nc - 2N0,O - “,0°(2 + 6y) + (NH200+ 2No: 2Nc)K] + (NHZo0 + 2No,0 - Nc)(l + by) = 0 (15)
Ind. Eng. Chem. Process Des. Dev., Vol. 19, No. 4, 1980
589
Table I1 blast graphical gas symbol
source
coal
Desai and Wen (1978) Lacey 1967) Gumz 11950) Woodmansee and Palmer (1977) Woodall-Duckham, Ltd. (1974)
Artwrigh t, bituminous coke anthracite Pittsburgh No, 8, bituminous R.O.M.Illinois No. 6, bituminous
air
A
0 2
0 E!
0
* w
2
air 0 2
yc0 - EXPTL
Figure 4. Comparison of dry mole fraction of CO between theory and experiment.
s
0" I
0"
>
1 0
10 yc0,-
I
1
30 EXPT~
40
/a
Figure 3. Comparison of dry mole fraction of C02 between theory and experiment.
in which 6y = 8" /ivH200. Equation 15 contains two unknowns, X and through K , T*. Solution of (13) and (15) is trivially accomplished by the method of Newton-Raphson (Hildebrand, 1956). To express the reaction heab as a function of T* it is convenient to use Kirchoff s equation
lo/
/
*
/ 20 50% 10
30
40
yHzEXPTL.
Figure 5. Comparison of dry mole fraction of H2 between theory and experiment.
Tt
AHj(T*) = Hj(298)
+ J298 A C ~ Gd, T
(16)
where ACpc, expresses the difference in heat capacity of the compounds which enter the reaction written as eq j. The standard heats of reaction a t 298 K are widely available; pressure effects on these enthalpy items have been ignored in view of the high temperatures of the system. Solution of eq 13 and 15 yields X and P.Application of the elemental balances, eq 9, 10, and 11, yields the exit gases in shift equilibrium at the top of the gasification zone. To this is added the components of the devolatilization zone in the amounts determined via proximate volatile analysis and given in Table I. An energy balance over this section then simply yields the exiting raw gas temperature with the assumption of an athermal standard heat of devolatilization. Results (a) Comparison t o Experiment. The predictions of this model have been compared to several sets of data corresponding to those entered in Table 11. These data include a range of bituminous coals and coke over a range of conditions from air blast to oxygen blast at a variety of steam to oxygen feed ratios. Significant differences exist in the gasification pressure, coal feed rates, inlet blast enthalpies, and reactlor heat loss as well as in the stoichiometry and proximate analyses of the coals. The last point is most importantly illustrated in the cases of the reported coke and anthracite gasification data since this has only a few weight percent volatiles. While these are accounted for as per Table I, their effect is quite small on
Y~~,EXPTL
Figure 6. Dry mole fraction of CH, proportioned from data of Loison and Chauvin (1964) compared to measured dry mole fraction CHI at gasifier exit.
the observed concentration of species and hence represent a clearer test of the primary shift gas equilibrium thesis assumed below the devolatilization zone. Insofar as the data and the computed results produce a vector of solutions for comparison purposes, a graphical display of the results has been chosen. The trends of the data are clearly displayed in Figures 3-6 as the computed (dry) mole fraction of the indicated component against its corresponding experimental value. The legend is listed as a column of Table 11. (b) Predictions Using the Model. As explained in the introductory sections, a gasifier operates by internally adjusting its coal feed rate according to consumption. Kydd (1975), on grounds previously discussed, has plau-
590
Ind. Eng. Chem. Process Des. Dev., Vol. 19, No. 4, 1980
probably depending on the adventitious concentration of gas-borne catalyst.) There is also a sharp dependence on temperature of the products of pyrolysis in the devolatilization zone (Juntgen and van Heek, 1968). These real effects will limit the desirable range of exit gas temperature from the reduction zone. As the heat leak rises, there is a corresponding decrease in the gaseous fuel value because of compositional changes in shift gases. In the case discussed, approximately 0.5% decrease in (lower) combustion heat is suffered for each 1% increase in heat leak.
I
1wO
300
5w
400
Mx)
7w Te ‘K
Bw
9w
loo0
11w
Figure 7. Effect of blast temperature by Kydd’s (1975) method. REPORTED HEAT LOSS WOODMANSEEa PALMER ~ 7 7 1
,
I
I
,
I
I
1
t
I
2
4
6
Y
HEAT LOSS, O h OF GROSS COAL COMBUSTION HEAT
Figure 8. Effect of gasifier heat loss on gasifier shift gas composition.
sibly suggested that an increment of the coal rate is determined by the inlet blast gas temperature. In the notation of this paper this result may be mathematically expressed as
G(Nc/No,O) 6hB =Nc/No,O AharCL
(17)
where ahBis the increment in the specific enthalpy rate of the fixed blast gas, AhcL is the (lower) heating value of the coal and rcLis the stoichiometric mass ratio between the inlet blast gas and the coal. Figure 7 shows the effect of increasing the inlet blast temperature in the range 300-1100 K for conditions of coal feed rate, heat leak, etc., reported for the GEGAS-D gasifier by Woodmansee and Palmer (1975), perturbed about their stated blast condition of 450 K. Over a 800 K inlet blast temperature range the (lower) heating value of the coal gas has risen -40%. It is noteworthy that the temperature at the top of the reduction zone rises from only 1180 K to 1320 K in this example. This implies a relatively poor conversion of blast sensible heat to product sensible heat (which is advantageous). On the other hand, heat loss from the gasifier vessel manifests itself primarily as a decrease in the sensible heat of the reduced gases. For example, in Figure 8 the effect of a variable heat leak on the otherwise nominal Woodmansee and Palmer (1977) gasification products is shown. The heat leak, expressed as a percentage of the (lower) combustion heating value of the coal feed, has a dramatic effect on the value of P. (In fact, essentially 65% of the heat leak has appeared as a direct reduction in the sensible heat of the reduced gases.) Commensurate with this effect, the shift constant, eq 8, changes and with it the composition of the gases assumed to be in shift equilibrium. (Actually the rate of t h e shift reaction will become sensibly finite “substantially” above 800 K (Ergun and Mentser, 1965)
Discussion The comparison of model results with experiment seems to confirm that assumption of shift gas equilibrium is the basis of a rational theory for a substantial amount of fixed bed coal gasification data. The plausibility of the assumption is enhanced in the case of fixed bed gasifiers because of their countercurrent operation which increases the contact time for the shift conversion to occur. It is also plausible to assume, as we have, and as proposed by Yoon et al. (1976), that the shift gas is quenched by the cold feed coal in a very short distance thus freezing the shift equilibrium. To this gas are added the products of devolatilization. The literature vouches for the complexity of this step. The estimate of these gases presented here is the estimate of Yoon et al. (1978), Table I. Fortunately, the fractional addition of gaseous moles to any component is small, except in the case of methane, which was assumed entirely due to coal distillation. This may be egregiously in error if mechanism 3 operates to any extent. Yoon et al. (1978) analyzed the data of Elgin and Perks (1974) (apparently abstracted from Woodall-Duckham, Ltd., 1974) and assumed additional coal tar cracking to account for the observed methane content of 8-9%. This was not done here. In view of the high heat content of methane, this point is clearly worthy of future study. The global model presented here has assumed that the char has an elemental composition of carbon, hydrogen, and minerals where the hydrogen is determined from the difference between the hydrogen based on the ultimate analyses and hydrogen present in the volatile fraction products, the total volatile content of the coal being based on the proximate analysis. Two points are worth noting. Firstly, the mass of hydrogen in the char is quite small, a few percent, but its volumetric influence is nonnegligible on the shift gases equilibrium. Secondly, no kinetic model of coal or char gasification includes the kinetics of this residual hydrogen. The global model is essentially independent of kinetic parameters, parameters which are poorly understood at best (Desai and Wen, 1978, Figures 15 and 16). Clearly a comparison to raw gas data is no guarantee of the applicability of the model. Desai and Wen (1978) use an “equilibrium model” in which a fraction approach to equilibrium is allowed for several reactions. Provided shift equilibrium is obtained, there is no effect of kinetic parameters at the outlet condition unless the reactor was too small to process the coal feed or that methane conversion is significant. If unreacted carbon is lost at the bottom grate the exit product will be poorer in carbon (and perhaps hydrogen); the extent of the loss will be reflected in a “leaner” coal gas. The current global model is blind to all heterogeneous reactions and requires no pseudo-equilibria assumptions. It is instructive at this point to quantify exactly what is lost vis-&vis the simple equilibrium model presented here and the more realistic rate model of Yoon et al. (1978), which assumed kinetic parameters for all of the reactions
Ind. Eng. Chem. Process Des. Dev., Vol. 19, No. 4, 1980 591
IERTED
Table 111. Dry Volume % Calculated by This Model and Yoon et al. (1978) fixed C HZ Yoon et al. in char char
co CO, H2
CH,
(others)
co CO, HZ CH,
(others)
c
Illinois Coal !12.1 !16.6 43.9 6.4 1.1
LOSS
W
21.9 25.1 45.4 5.8 1.8
22.0 25.9 44.4 5.9 1.8
Wyoming Coal 20.4 !11.5 !16.8 25.1 42.5 47.0 8.0 5.8 1.2 1.7
20.5 25.9 46.0 5.9 1.8
; z
V
k
Ee
z
1 to 6 except that, as here, the shift reaction was in equilibrium. In Table 111, a comparison is made between the results of the modlel of Yoon et al. for an Illinois and a Wyoming coal, and those of the global model calculated with and without hydrogen attachment to the char, for an adiabatic reactor. Feed conditions are those assumed by Yoon et al. except that, additional compositional data were required for the C and H analyses of the coal (obtained from Walters et al., 1967) for the HP/char case. Hydrogen and methane show the largest discrepancies. In particular the Wyoming coal data show higher methane yields than the global model predicts. These coals are more reactive than Eastern bituminous varieties, and it is reasonable to assume that direct methanation proceeds in such reactive coals at a nonnegligible rate. This is probably the prime advantage of the rate model. Indeed, this point is well made by Denn et al. (1979) in their comparative study of kinetics-included and kinetics-free gasification models. It is also possible that more methane is produced during the complex devolatilization process than the devolatilization data base used here indicated. However, the advantage of this present model is its simplicity and its insensitivity to kinetic parameters with the agreement levels as indicated in Table 111. We further note that our “predictive” model requires a t least one experimental point of coal feed to oxidant ratio. We have found that all of the experimental results whether air-blown or oxygen-blown have a range of char carbon moles to oxygen in the approximate range 2.0 to 2.5. Clearly the coal reactivity enters the calculation in the variables discussed. However, if we accept Kydd’s (1975) model (eq 17) just one point is required for a given reactor and reactivity coal. The model of Yoon et al. (1976, 1978) and the results of Hebden (1975) confirm the general results that the combustion and devolatilization zones are physically small compared to the gasification or reduction zone. Presumably then the limiting kinetics is in the reduction zone and it should be a promising line of inquiry to establish a relation between its kinetics and the “standard” coal processing point required for eq 17. Even without the exact coal feed there is much useful comparative data that can be obtained. From Figure 8, for example, the dramatic effect on the exit reduction zone temperature has been noted. Most significant is the variation of reduction zone composition with temperature via heat leak. The heat leak question is most critical to the data of Desai and Wen (1978), who report Morgantown Energy Research Center data. The Morgantown reactor is about 1 m diameter and has a water jacket. Heat leak is profoundly felt. Dlesai and Wen report heat losses as approximately 23% of the coal combustion heat. It was not possible to predict reasonable exit raw gas tempera-
I 0
I 10O h
I
I
20%
HEAT LOSS AS % OF COAL HEATING VALUE
Figure 9. Composition and temperature of raw gas measured at the Morgantown gasifier as reported by Desai and Wen (1978).
tures with such heat losses. In order to achieve the reported raw gas temperatures we were forced to assume lower heat losses as per Figure 9. A t the reported heat loss the computed raw gas temperature was below 400 K compared to 950 K measured. Figure 9 also shows the computed fixed gases which vary significantly with heat loss. It is expected that the heat loss is lower than the 23% reported; indeed Desai and Wen computed 15.7% which agrees with all but the experimental raw gas temperature. The latter calculation is influenced by the assumption of a zero heat of devolatilization and may not be meaningful. Conclusions It is possible to model rather accurately the output conditions of a fixed bed gasifier irrespective of whether it is air or oxygen blown, whether high steam/oxygen or low. The assumption of exit shift equilibrium from the reduction zone appears surprisingly good considering the number of probable gasification reactions which are being modeled. The accuracy of prediction of coal gas components is worst in the case of methane since it may be produced by direct hydrogasification (not included in the model) or by devolatilization and volatilization processes which are characterized. The induced errors in other components will be smaller. As a predictive model the global model shows the effect of both blast conditions and of heat leak. The heat leak calculation shows a first-order effect of the sensible heat of the reduced gases and hence on their composition and through composition on the heat of combustion. The latter point has not been previously stressed in the literature. Acknowledgment Valuable conversations and insights were afforded us by interaction with Dr. D. E. Woodmansee. We gratefully acknowledge these contributions. Nomenclature CpGi = molar heat capacity at constant pressure of gas i Cs = specific heat of ash material 6hB = increment in blast gas enthalpy A h c ~= specific lower heating value of coal AHj = molar heat of reaction j as written M s = solid ash flow rate Nci, Nc = molar flowrate of i and of carbon, respectively NG/’ = molar feed rate of blast gases at feed grate level 6NH2 = molar rate of hydrogen trapped in char to gasification zone Q = reactor heat loss rcL = stoichiometric blast gas to coal mass ratio
592
Ind. Eng. Chem. Process Des. Dev. 1980, 19, 592-599
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Received for review October 12, 1979 Accepted May 21, 1980
Behavior of Gas Bubbles in Bubble Columns Korekaru Ueyama Department of Chemical Engineering, University of Tokyo, Bunkyo-ku, Tokyo, 113, Japan
Shlgeharu Morooka Department of Applied Chemistry, Kyushu University#Higashi-ku Fukuoka, 8 12, Japan
Koro Kolde Department of Chemical Engineering, Shizuoka University. Johoku, Hamamatsuahi, Shizuoka, 432, Japan
Hlsatsugu Kajl Nishiki Research Laboratory, Kureha Chemical Industry Company, Ltd., Nishlki-machi, Iwaki-shi, Fukushima, 974, Japan
Terukatsu Mlyauchl' Department of Chemical Engineering, University of Tokyo, Bunkyo-ku, Tokyo, 113, Japan
The behavior of gas bubbles in a bubble column of 0.6 m i.d. was experimentally studied by changing gas velocity, liquid depth, and the gas distributor used. The resutts were compared with those obtained using a large-scale bubble column of 5.5 m i.d. There was no clear effect of column diameter on the average gas holdup corresponding to bubbling bed height. However, the axial and lateral distribution of local gas holdup which was detected by an electric resistivity probe showed strong dependence on the absolute value of liquid depth and the type of gas distributor used. The axial and lateral distributions of bubble velocity and bubble size were also obtained using twin electric resistivity probes. When the superficial gas velocity was 0.02-0.04 m s-', the mean diameter of bubbles for 5.5 m i.d. was almost twice as large as that for 0.6 m i.d. On the basis of these data, the flow characteristics of the bubble column were discussed.
Introduction
The bubble column is a typical reactor used for gasliquid systems as it can be easily constructed owing to simplicity of design and an absence of moving parts. To date, many studies have been undertaken on the flow characteristics of the bubble column (Maeda, 1963; Kato, 1963; Ptergaard, 1968; Sada, 1969; Akita, 1973). Most of these studies, however, are confined to the bubble-flow
regime where the gas flow rate is small and individual bubbles ascend homogeneously after uniform generation by a gas sparger. Recently, special interest has been directed toward operating a bubble column in the higher gas feed regime where the majority of bubbles concentrate in the central region of the column and ascend by alternately coalescing and breaking up (Pavlov, 1965; Towel1 et al., 1965; Yosh-
0196-4305/80/1119-0592$01.00/00 1980 American Chemical Society
