728
Ind. Eng. Chem. Process Des. Dev. 1982, 21, 728-750
absorption with temperature of the Fen(NTA)NO complex. An equilibrium constant for Fe"(NTA)NO of 2 X lo6 M-' at 25 "C was obtained by extrapolating the data illustrated in Figure 2. With an upper limit for the relaxation time of 10 p s and the lowest NO concentrations of 1.4 X M, lower limits for 12, = 7 X lo7M-I s-l and k-, = 35 s-l at 25 "C were obtained. The lower limits obtained for k, did not exhibit any pH dependence over the range of pH 5-7. The lower limit of k, determined for Fe"(NTA)NO is consistent with the value of kl obtained by indirect means for a similar complex Fe(EDTA)NO (Teramoto et al., 1978)-kl = 1X 10s M-' s-l-and is much faster than that observed for the Fe(Hz0)5N0complex (Kustin et al., 1966)-k1 = 6.2 X lo5 M-' s-l. The substitution of NTA for some of the water molecules in the Fen inner coordination sphere weakens the bond between the Fe" ion and the remaining water molecules. The higher rate constant observed for Fell(NTA)(HzO)zcompared to that of Fen(HzO), is in agreement with other cases of higher-order substitution kinetics.
Conclusion It has been suggested (Yaverbaum, 1979) that FenEDTA can be an effective binding reagent/additive for NO in wet simultaneous flue gas desulfurization and denitrification scrubbers. The equilibrium constant of the reaction of NO with FenEDTA to form Fen(EDTA)NO (Hishinuma et al., 1979) is about a factor of 2 larger than that with FenNTA to form Fe"(NTA)NO a t a typical wet scrubber temperature, T = 50 OC. However, NTA is less expensive than EDTA. The rate of binding of NO to form nitrosyl ferrous complexes by both Fe"NTA and FemEDTA is probably within the same order of magnitude and is much faster compared with the mass transfer rate at typical wet scrubber conditions. Determining the rate of regeneration of these metal chelates by investigating the kinetics of the reaction of Fe"(NTA)NO and Fe"(EDTA)NO with absorbed SOzin aqueous solutions is important for assessing whether NTA or EDTA is an optimum chelating agent.
Acknowledgment We thank Professors Leo Brewer, Robert Connick, and Scott Lynn for helpful discussions, and we appreciate the support and encouragement of Miss Stephanie Bialovak and Dr. Jack Halow. This work was supported by the Morgantown Energy Technology Center, Contract No. 81MC14002, through the Assistant Secretary of Fossil Energy of the U.S.Department of Energy under Contract NO. DE-AC 03-76SF00098. Literature Cited Armor, J. N. J. Chem. Eng. Data 1974, 19, 82. Chang, S . G . "Reactions of Sulfite and Nitrite Ions in Aqueous Solutions"; paper presented at the Technical Advisory Committw. DOEIAdvanced Envkonmental Control Technology Program, Morgantown, WV, Nov 6-7, 1980; Lawrence Berkeley Laboratory Report LBL-11800, 1980. Chang, S.-G. "Kinetics of Reactions in a Wet Flue Gas Simultaneous Desulfurization and Denitrlficatlon System"; 181st National Meeting of the American Chemlcal Society, Atlanta, GA, Mar 30, 1981; Lawrence Berkeley Laboratory Report LBL-13063, 1981. Chang, S.-G.; Toossi, R.; Novakov. T. Atmos. Environ. 1881, 15, 1287. 1959, 63, 652. Czerllnski, G.;Eigen, M. 2.€/ektf&m. Gomiscek, S.; Clem, R.; Novakov, T.; Chang, S.-G. J. Phys. Chem. 1881, 85, 2567. Hishlnuma, Y.; Kaji, R.; Akimoto, H.; Nakajima, F.; Mori, T.; Kamo, T.; Arikawa, Y.; Nozawa, S. Chem. SOC.Jpn. Bull. 1879, 52, 2863. Kustn, K.; Taub, I. A.; Weinstock, E. Inorg. Chem. 1866, 6, 1079. Llttlejohn, D.; Chang, S. G. J. Phys. Chem. 1982, 86, 537. Oblath, S. B.; Markowitz, S. S.; Novakov, T.; Chang, S A . J. Phys. Chem. 1981a, 85, 1017. Oblath, S. B.; Markowltz, S. S.; Novakov, T.; Chang, S.-G. "Kinetics of the Initlal Reaction of Nitrite Ion in Bisulfite Solutions"; Lawrence Berkeley Laboratory Report LBL-12581, 1981b; accepted for publication in J. Phys. Chem . Oblath, S. 8.; Markowitz, S. S.; Novakov, T.; Chang. S . G . "Reaction of Nitrite Ion with nydroxylamine-N-sulfonate in Aqueous Solution"; Lawrence Berkeley Laboratory Report LBL-13255, 1981c; submitted to Inorg. Chem Pietrzyk, D. J.; Frank, C. W. "Analytical Chemistry"; Academic: New York, 1974; p 453. Pltzer, K. S.; Brewer, L. "Thermodynamics"; McGraw-Hill: New York, 1961, p 298. Teramoto, M.; Hiramine, S.; Shlmada, Y.; Sugimoto, Y.; Teranishi, H. J. Chem. Eng . Jpn. ?S78, 1 1 , 450. Yaverbaum, L. H. Nitrogen Oxides Control and Removal, Recent Developments"; Noyes Data Corporatlon: Park Ridge. NJ, 1979.
.
Receiued f o r review April 27, 1982 Accepted June 1, 1982
Thermodynamic and Kinetic Constraints of Catalytic Synthetic Natural Gas Processes Reuei Shinnar,. Gilead Fortuna, and Dan Shapka Department of Chemical Engineering, The City College of New York, The City Universe of New York, New Yo&, New York 10031
The work contains a thorough analysis of the thermodynamic constraints of converting coal to synthetic natural gas (SNG). I t is shown that the thermodynamic constraints that limit the thermal efficiency are not inherent but the result of design decisions based on available technology, as weH as the kinetic properties of available catalysts. The latter limits the yield of methane to that obtainable at global equilibrium over carbon in the presence of CO and H,. The equitlbrlum composition is shown to be independent of the thermodynamic properties of the char or coal fed. These limitatlons give nonisothermai two-stage processes significant thermodynamic advantages. The analysis results in suggesting directions for modifying present processes to obtain higher thermal efficiencies, and it presents two-stage process schemes that would have significant advantages over present technology. As the methodology used for the thermodynamic analysis contains some novel elements, it should be of interest to the reaction engineer in general and should be applicable to a wlde range of catalytic and noncatalytic processes.
1. Introduction There has been considerable attention in the recent literature to the so-called second law or the availability analysis of coal gasification processes (IGT, 1978-1979, 0196-4305/82/1121-0728$01.25/0
Gaggioli and Petik, 1976) as a tool to evaluate their thermal efficiency. Thermal efficiency is a very important consideration in conversion of coal to either SNG, fuel gas, or syngas. The importance of thermal efficiency is far 0 1982 American Chemical Society
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 4, 1982
greater than utilization of coal. It determines the amount of C02released to the atmosphere. Furthermore, the required investment is strongly correlated to thermal efficiency (Shinnar and Kuo, 1978; Stone and Webster, 1977). Therefore, any clues for improvement in a given thermal efficiency are valuable. However, the constraints that limit thermal efficiency are not really thermodynamic but mainly kinetic, process, and material constraints, which interact in a complex way. If a process temperature is limited due to material constraints, then this may have thermodynamic implications which limit the thermal efficiency. Kinetic properties of catalysts also have thermodynamic implications. In order to draw important inferences for improved process schemes and reactor designs in SNG processes, it will be shown that it is sufficient to know that some reactions are faster than others. Thermodynamics is a powerful tool to utilize such knowledge. The methods presented here should be of general interest to those engaged in either reactor or process design. In the following, a complete analysis of the factors which determine the thermal efficiency of SNG processes will be presented. It will also be shown how present processes are compared with one another and how they can be improved. 2. Outline of the Analysis In this work an attempt is made to present the analysis of thermodynamic constraints of the SNG process by using a somewhat novel approach. For the sake of easy reading, the sequence of the paper is briefly described. In section 3, the relevant process efficiencies are defined. The analysis of chemical processes is compared to that of power plants and the differences are pointed out. The analogue of the power cycle in a chemical process is the idealized flowsheet in which all changes are reversible and no driving forces are needed. Chemical processes have no inherent constraints, but design decisions based on available materials or technology have thermodynamic implications that limit thermal efficiency. The above approach allows us to look at the impact of each decision separately. In section 4, the effect of the heat balance and equilibrium limitation on the process is analyzed. The analysis uses a completely idealized process, in which neither CO nor H2is formed, and ideal semipermeable membranes are allowed. It is shown that one of the main penalties limiting thermal efficiency is the degradation of unconverted steam. Section 5 shows the thermodynamic implications that result from the properties of presently available catalysts which form CH4 by first forming CO and H2 followed by the shift and the methanation reactions. The gasification of pure carbon is analyzed and it is shown how the law of detailed balancing can be used in conjunction with a knowledge of the relative magnitude of various reaction rates to predict the space of accessible compositions. This section presents a technique to analyze the interaction between properties of catalysts and thermodynamic constraints. Furthermore, it is shown why for presently known catalysts nonisothermal two-stage processes have inherent advantages. Section 6 discusses some idealized flowsheets based on properties of presently proposed processes and looks at the implications of the previous sections on these processes. The main factors limiting the thermal efficiency of several process schemes are analyzed and discussed. In section 7 the implications of gasifying carbon (ideal case) to gasification of coal and char is discussed. It is shown that the results of the previous sections are relevant due to two factors. Present Catalysts require temperatures at which coal devolatilizes fast to form char and the ki-
729
Table I. Availability Calculations" datum level: To= 298 K ; P , = 1 atm stable reference: air 0.21 0, 0.79 N, H,0(1), co, availability of gas mixture P t Amk(T,P) = J ~ o O X j [ C ~ pTOl C p jd ]T t RTo In P.
- 0
z X j [ h j ( T o ) - Tosj(ToJjPo)- ~ j o a P j o = chemical potential of species j in the reference environment; hj, Sj = partial molar enthalpy and entropy of speciesj.
I
netics of char gasification are such that the equilibrium constraints are similar to that of carbon and independent of the thermodynamic properties of the char. The latter conclusion is derived from the law of detailed balancing and kinetic properties of known catalysts. Therefore, it allows us to extend the results of section 6 to coal gasification. Also discussed are process schemes which will have significant advantages on thermal efficiency over processes presently under development. In section 8 the idealized flowsheets are compared to the present state of technology and suggestions for process improvements and development are derived from the preceding analysis. 3. Definitions of Thermal Efficiency There are several ways in which thermal efficiency of a synthetic fuel plant can be defined. Based on heat balance, a thermal efficiency qH can be defined as qH
= [higher heating value (HHV) of useful products]/[(HHV) of total coal feed to plant] (1)
The total coal feed to plant includes both process coal and coal used for power generation. Another way is to base the definition of the efficiency on the available free energy (A) and define ?A
= [A(useful products)] / [A(total coal feed)] (2)
The usefulness of QA is that it represents the only rigid thermodynamic constraint on a chemical process, namely, qA 5 1.0. For fuels and chemicals (feedstocks or products) the availability is defined at ambient conditions with respect to the enviroment (see Table I) (Gaggioli and Petik, 1976). In SNG plants the only feeds to the process are coal and water, and therefore energy requirements such as electricity and process heat are assumed to be derived from coal. For most fuels the availability is very close to the lower heating value (LHV) (Flower and Linnhoff, 1979). Therefore, a third definition for efficiency (qL) in which LHV is substituted for HHV in eq 1 will be used. qL = [LHV (useful products)]/[LHV (total coal feed)]
(3) for SNG processes is very close to qA and has the advantage that it is easier to be calculated. In this article we are interested in deriving limits on vL and evaluating the role of thermodynamic analysis in process design. We are especially interested in understanding how catalyst and gasifier design affect the inherent thermodynamic constraints of the process. The latter are a very important tool in economic analysis of processes (Shinnar et al., 1981). However, one has to be careful with the term thermodynamic process constraint. In principle, all chemical processes can be run in a reversible way. qL
730
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 4, 1982
Table 11. Thermodynamics Constants for Reactions Used AGO298
AH" 298
AS0298
6.99 -94.26 23.9 9.56 17.83 -12.14 28.64
24.65 -94.05 41.88 21.29 32.95 -17.89 41.22
59.27 0.7 60.3 39.36 50.76 -19.3 42.21
-6.81 -27.15 -33.96 -40.78
-9.84 -39.41 -49.26 -59.10
-10.16 -41.16 -51.32 -61.49
-8.42 17.54 12.78 4.43
7.4 38.07 38.75 28.03
53.08 68.89 87.1 79.19
A. Graphite
carbon-gas 2C + 2H20(1)-+ CH, + CO, c + 0, co, C + HiO(1) +'CO + H, 0.5C + H,0(1) 0.5C0, + H, 1.5C + H,0(1) 0.5CH, + CO C + 2H, CH, c + co, -+ 2 c o -+
-+
-+
-+
gas-phase reactions
CO t H,O CO, + H, CO? + 4H, - + C H , + 2H20 CO t 3H, -+ CH, t H,O 2CO + 2H, +CH, + CO, -+
B. Coal 1.738CH,,,OO., + 1.304H20(1) CH, t 0.738C0, CH,,,O,., + 0.9H20(1) CO t 1.3H, CH,,,O,,, + 1.9H20(1)-+ CO, + 2.3H, 2.307CH,.,OO,, + 1.0763H20(1) CH, + 1.307CO -+
--f
-+
HEAT FROM ENVIROMENT
1300 OF or more, but investment as well as maintenance cost will go up. The Carnot principle is useful here for the analysis of the thermodynamic effect of a design decision. It is different from applying it to the analysis of power plants based on temperature gradient in the ocean where the Carnot principle is a hard constraint. Thermodynamic analysis can be used in a more detailed way in designing better power plants. Power plants do not operate by the Carnot cycle, a fact which reduces the efficiency even further. Representing the actual process by an idealized cycle (such as the Rankine cycle) allows a fast estimate as to the penalty introduced by a specific design approach. In a chemical plant the equivalent of a power plant cycle is an idealized flowsheet in which it is assumed that each step is carried out at reversible conditions, limited only by the first and second law of thermodynamics which in return limits the conversion in the reactor. Such analysis is useful as long as the efficiency of the real process is close to that of the ideal one, which applies in many cases. Exceptions are separation processes for which real efficiencies are often very low ( 1200 >600 S 600
> 300
output streams
factor
temp, OF
1.5 1.0
>700
1.0
400
0.5
0.5
factor
to deal with the most common option which is the feeding of oxygen. Feeding oxygen has a small additional penalty. It increases the COz content in the gasifier, which in return reduces methane yield and steam conversion. This penalty is shown in Figure 18. It increases with temperature, but in the temperature range of interest, the penalty is small. The fact that we put the reaction heat in the gasifier at high temperature and regain it at low temperature is a severe penalty, which under present design constraints is much larger than one would predict from a straight thermodynamic analysis. However, once we have data for this penalty, we can use thermodynamic analysis to compare different process options. We noted before a similar problem in the steam balance. There was a distinct difference between heat obtained above the boiling point of high pressure steam and below it. Similarly, heat at high temperatures has a much higher penalty than equivalent heating value of coal. Availability analysis predicts correctly that the cost of reaction heat should increase with temperature. However, there is one difference that design constraints impose on the heat balance of plants. The cost of high-temperature heat input is not a continuous function of temperature but rather a step function. Furthermore, the cost is about 50% more than the availability of the heat used from combusting coal and could be triple the availability value of the reaction heat required. For approximate analysis it might therefore be preferable to represent the value of any heat flow in the plant in terms of step functions, such as given in Table IV. Both the temperature of the step and the penalties depend on the plant design and on the design constraints accepted by the designer. Table IV is a strong oversimplification and accurate values can only be obtained by detailed flowsheets. However, methods like those in Table IV give a better approximation of design tradeoffs than looking at the available free energy of the heat flows. 6.2.3. Recycle Penalty. The two versions of the Exxon process have another penalty, which dominates their thermal efficiency and their design. The two versions require cryogenic separation of CO and H2 from methane and recompression of the gases. The cryogenic separation is energy-wise very inefficient and provides a heavy energy
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 4, 1982 741
-PENALTY FOR TEMPERATURE HEAT HIGH
NO PENALTY
I - TWO STAGE II . EXXON IU- MODIFIED EXXON
___ 1 a: 1 l
800
900
1000
1100
EQUILIBRIUM TEMPERATUREI'K)
1200
:
800
L
L
900
-
a-
.
1000
'
1100
-
1200
E O U I L I B R I U M TEMPERATURE I'io
Figure 19. (a) Recycle requirements for routes I1 and I11 as a function of gasifier temperature; (b) energy requirements for cryogenic separation and recompreesion for routes I1 and I11 as a function of gasifier temperature (based on heat input to boiler in power plant).
penalty (Air Products, 1980). Both the number of moles gas recycled and the energy requirements for the recyde are shown in Figure 19. In both routes the energy requirements are far larger than the requirements indicated by thermodynamics and we therefore use energy requirements based on practical experience which are large a t high temperatures. We note, here, that the modified Exxon route has a very significant advantage at high temperatures. Figure 19 shows one important feature of the separation process in the two Exxon routes. For low recycles real energy requirements are almost independent of recycle or of the CO and H2 content, which is a typical feature of many practical separation processes. 6.2.4. Comparing the routes-the Practical Advantage of the Two-Stage Process. The penalty for unconverted steam decreases with increasing temperature. The penalty due to supplying reaction heat in the twostage process and due to recycling requirements in the two Exxon routes increases with increasing temperature. Therefore, we should expect an optimum temperature for each of the three processes. Figure 20a shows the minimum total energy requirements of the three routes as a function of temperature. The total energy requirements include the basic requirements for all processes, namely, 2 moles of steam per mole of CHI, the heat of reaction, and the three main penalties of the processes: (1)penalty due to incomplete steam conversion, (2) penalty due to the heat of reaction, and (3) penalty due to the cryogenic separation process. Since the penalties for recycle and oxygen separation are taken from practical experience, the same is done for the other penalties. Therefore, the penalty for excess steam includes the efficieny of the steam generation (85%). Two lines are shown for all processes: one for temperature range, where indirect heat transfer might be possible, and the other line which includes the penalty of indirect heat transfer at high temperatures. We note that the two-stage process is superior compared with the two Exxon processes. Since recycle separation has an inherently large minimum penalty, it is preferable to operate the reactor at low temperatures, where the heat requirements as well as CO and H2content, are small. If a catalyst which operates at low temperatures can be found, the two-stage process will be clearly superior. There is no present process under development which uses indirect heat transfer to the reactor by burning coal. It will only make sense at temperatures below 900 K, and there are no catalysts which operate well a t that temperature. The main advantage of the two-stage and the modified Exxon processes over the original Exxon process is not in the small differences in energy requirements, but in the ability to operate at higher temperatures without a severe
I 800
1000
(UJ
_______.. --
40 1200 800 GASIFIER TEMPERATURE,
1000
1201
O K
tb)
Figure 20. Minimum total process energy losses for the routes mentioned: (a) ideal flowsheet; (b) actual flowsheet.
penalty. Operating at higher temperatures can either allow use of cheaper catalysts or use of smaller reactors due to higher reaction rates. It will be shown later that these conclusions also hold for reactions with coal. The catalytic two-stage process as well as the modified Exxon process allow, with existing catalysts, significantly higher efficiencies than existing processes, which, due to process constraints, operate far from the optimum conditions. However, all such processes involve similar compromises between the penalty, due to feed of oxygen (or other ways of supplying the heat of reaction) and efficient steam utilization. Therefore, our approach is applicable to them, too. Noncatalytic processes require higher temperatures in the fiist stage and in most cases have lower methane yield than the yield at equilibrium. The penalty of lower methane formation can be estimated by looking a t the limiting case with no methane formation in the first reactor. The penalty for supplying the heat of reaction is 40 kcal/g-mol of methane vs. 10-20 kcal for the catalytic two-stage gasification methanation reador with maximum methane formation in the first stage. Figure 20a is based on a comparison of the main penalties, excess steam and separation requirements (oxygen for two-stage and cryogenic separation for the recycle). Figure 20b shows similar overall results based on actual flowsheeting. 6.2.5. Hydrogasification. Till now hydrogasification was excluded from the discussion. It could have been the best route if a catalyst for hydrogasification, operating at 650 K, could have been found. Yet, present catalysts (or noncatalytic hydrogasification) require high temperatures (above 1000 K and preferably above 1100 K) to achieve reasonable gasification rates. At high temperatures the equilibrium is unfavorable, as can be seen from Figure 21, where the conversion of Hz over carbon as a function of temperature is plotted. If the gas already contains CH,, as it does for any gasifier operating below 1300 K, then conversion is further limited. To get a high hydrogen conversion, a cryogenic separator is needed. A possible process scheme is given in Figure 22. In this case an incremental 0.069 mole of CHI per mole of steam fed is achieved compared with the two-stage process at the same temperature (see Figure 14). The higher yield of C H I per mole of HzO also reduces the penalty of supplying heat to the first reactor. However, this advantage is more than
742
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 4, 1982 loo
1
lb00.K' (NO OXYGEN)
1
EXXON MODIFIED EXXON
h i\ 0
840
t
c-
,
20 400
600
SO0
1000
1200
EQUILIBRIUM TEMPERATURE
1400
1600
Figure 21. Conversion of H, over carbon in an hydrogasifier. if
n
,
c*.
4 I
I
30
IO
( O K )
20
30
40 50 60 70 GASIFIER PRESSURE (ATM)
80
90
100
Figure 23. Minimum total energy losses as a function of pressure for the three routes (product at 68 atm).
CO,
,-
Figure 22. Schematic flowsheet for hydrogasification process.
nullified by the large energy requirementa of the cryogenic separator. In the two Exxon processes the need for an oxygen plant was eliminated and a cryogenic separator substituted. Here, we need both. To illustrate that, Figure 20a gives the energy requirements for two points, one for the hypothetical case where the hydrogasifier operates at 650 K, and a second where both the steam gasifier and the hydrogen gasifier operate at 1100 K. We note that the second case is worse than the two-stage route. For certain coals, hydrogasification can still be attractive due to other process considerations, especially since many noncatalytic gasifiers operate far from the optimum conditions. 6.2.6. Effect of Pressure. Till now all processes were considered at a fmed pressure, 68 atm. It was noted before that the main parameters determining thermal efficiency for the three routes under consideration were the steam requirements, heat of reaction, and recycle ratio for the two Exxon processes. Pressure has one further effect. If the CH, is deaired at 68 atm, lower operating pressures will require compression of the product. The equivalent of Figure 20 is shown in Figure 23, where the total minimum energy requirements for the three processes are compared as a function of pressure at lo00 K. The total minimum energy requirements include the steam requirements, the heat of reaction, and the recycle losses. For the Exxon process the heat of reaction and the net steam requirements are independent of pressure. For the other routes, the heat of reaction decreases with pressure, but net steam requirement increases with increasing pressure. The advantage of low steam requirement a t low pressures is negated by higher reaction heats, higher recycle ratios, and the requirement of product compression. 7. Reactions with Coal 7.1. Hypothetical Direct Reaction with Coal (Low Temperatures). Till now we have concentrated on re-
actions of carbon. However, the previous analysis will not be useful if the results for coal and coal chars will not be similar. In practice they are, and the main reason for this similarity is that in all gasifiers coal decomposes to char and volatiles faster than any direct reactions with steam or hydrogen. As will be shown later, char has similar properties to thoae of carbon, which justifies our emphasis on carbon. Consider a hypothetical coal with the composition CH,,.800.1 (typical Eastern coal. Real coal will also contain sulfur and nitrogen but we will neglect them here.) The overall reaction of this coal is given in reaction XI1 (Table 11). We have no accurate data for AGO, and AHoB8,but a reasonable estimate would be (see Appendix) 1.738cHo&o.~ + 1.304H,O(g) CH4 0.738C02 (13)
+
AGO298
= -10.46 - 1.738 X AGo2g8(cod) = -1 1.53 kcal/g-mol
m0298
= -11.93 - 1.738 X mo,9a(coal) = -6.7 1 kcal/g-mol
It is noted that the reaction has a negative AGO,,. Although the heat of vaporization for water is still needed, only 1.3 moles of steam per mole of C H I is required (instead of 2 moles of steam per mole of CH4 in the case of the reaction with carbon). In reality, despite the negative &Gozse, the process is only slightly better than for carbon and the reason for that is that the reaction CHo.aO0.l char + volatiles (14)
-
is much faster than all other direct reactions of steam with coal. Thus, char and not coal is gasified. Since coal decomposes to char rather fast above 700 K (800 OF), one would need a catalyst operating below 650 K (600 O F ) to prevent char formation. Such a catalyst would be a major achievement, but this paper is mainly concerned with known catalysts or known reactions. Therefore, it will be taken into account that coal devolatilizes first and the char left is gasified. The properties of this char are very close to that of graphite and are independent of the properties of the coal gasified. The reader interested mainly in process design can therefore skip the next section and proceed to section 7.3. However, many claims have been published that methane equilibrium over char is strongly dependent on char properties. As this would significantly affect our analysis, a thorough discussion of the evidence and the thermodynamic constraints involved is given in the next section. The analysis presented in this chapter should also
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 4, 1982 743 BASIS: 1 MOLE H20 FEED 68ATM
600
700
800 900 1000 1100 1200 EPUlLlBRlUM TEMPERATURE ( O K )
1300
Figure 24. Estimate of product yield at equilibrium for steam-char reaction.
be of general interest to the reaction engineer as it demonstrates the value of thermodynamic analysis in interpreting kinetic data in complex multiphase systems. 7.2. The Effect of Char Propoerties on Global Equilibrium. An estimation based on different assumptions for the structure of char is presented in the Appendix. Since the estimates are approximate, a range is given for the heat of reaction and for the free energy of formation. The overall reaction of steam with char can be written as l.86CHo.3+ 1.72H20(g) CHI + 0.86COz (15) = 0.76 - 1.86 AG0298(char)= -2.03 to -0.82 kcal/g-mol
-
AHoZg8= -12.52 - 1.86 AHzg8(char)= -10.66 to -12.74 kcal/g-mol AGO298 is negative a t low temperatures, but at the temperature where reaction rates are appreciable, AG(V is slightly positive. The equivalent equilibrium yields of Figure 6 are given in Figure 24, assuming again that char is the only solid compound and that CHI, HzO, CO, COz, and Hz are the only compounds in the gas phase. At a temperature of lo00 K we get a much higher C H I yield at global equilibrium than the CHI yield at global equilibrium over carbon, but steam conversion is incomplete. Thus, conversion is still equilibrium constrained. However, close to equilibrium the assumption that only one type of char exists in the solid phase is not realistic. It implies that any reverse reaction in the gas phase forms exactly the same char. The gas-solid phase reactions known to have reasonably large reaction rates are the Boudouart reaction 2 c o + c + c02 (16) and methane decomposition reaction (17) CHI + C 2H2 (see reactions VI1 and VI in Table 11). The conditions of carbon formation from synthesis gas and from CHI have been studied extensively in connection with steam reforming and shift reaction (Rostrup-Nielsen, 1972; Gruber, 1975; Dent, 1945). A t temperatures below lo00 K the gas can sustain a higher supersaturation (gas phase concentrations which have higher ratios of [CH4]/ [HJ2 or [CO]2/[COz]than permissible at equilibrium over graphitic carbon) for long times. The permissible supersaturation depends on the catalyst. The carbon formed has a higher free energy of formation compared with graphite. At temperatures above 1000 K the free energy of the carbon formed is independent of the chemical
+
composition of the catalyst and similar to that of graphite. The gas phase cannot sustain a supersaturation since carbon forms easily on any surface. The few available data on equilibrium over char, such as those of Exxon (1978), indicate that the properties of this char are fairly close to those of graphite. There is no evidence that either thermal or catalytic reactions can synthesize a complex char from CO, Hz, CHI, and C02in the time scale of the gas residence time in a gasifier. This leads to two interesting conclusions: (1)If a complex char is brought to equilibrium with steam or hydrogen, the equilibrium mixture will contain carbons or simple chars which result from the reactions such as methane and CO decomposition. (2) The reaction of a complex chemical structure of char with steam or hydrogen, near equilibrium, is a two-step reaction involving an irreversible decomposition to carbon followed by reaction with steam. The first conclusion results from the fact that CH4 and CO decomposition have f i i t e reaction rates forming simple chars a t temperatures and at pressures of interest. The second conclusion results from the requirement of detailed balancing. We note from Figure 24 that steam conversion is incomplete. Therefore, the overall reaction is reversible and the forward reaction rate, at the begining, is larger than the backward reaction by just one order of magnitude. If the formation of the char from the gas is very slow, the forward reaction near equilibrium involving steam and char must be slow too. The fact that coal and complex chars are not formed at reasonable reaction rates from either CO, Hz, or CHI implies that the forward reaction must have some irreversible steps in order to be fast, irreversible in the sense that conversion in these steps is far from equilibrium or that equilibrium conversion is so high that the ratio between backward and forward reaction is very large. In the latter case even a negligible rate for the backward reaction is consistent with a fast forward reaction. Consider, for example, the case of reaction VI with the same char as in eq 15. CHo.3 + 1.85H2 CH4 (18)
-
Let us assume that this char is significantly different from graphite. At 1100 K and 68 atm the conversion of Hz is 83% at equilibrium. It is also known that at 1100 K, the backward reaction of 18 or formation of this specific char from CH4 is slow compared to formation of a simple char from methane with properties similar to graphite. At equilibrium the forward and backward reactions have equal reaction rates and therefore near equilibrium the forward reaction of reaction 18 must be slow compared to the hydrogasification of a graphitic carbon. Let us look now at a composition far from equilibrium. For pure hydrogen at 68 atm the hydrogen partial pressure will only be 4 times larger than at equilibrium. As the overall reaction rates of eq 18 and reaction VI are both approximately same order in pressure, reaction 18 should still be very slow compared to hydrogasification of the char formed from decomposition of CHI. If hydrogasification of this char is fast far from equilibrium then it must involve an irreversible step different from reaction 18. What do we mean by an irreversible step? Consider, for example, devolatilization of coal. Let us take a hypothetical overall reaction CHo.800,1 0.7C + 0.2CH4 + 0.1CO (19) At 1000 K the equilibrium pressure is 1O'O atm. This is 80 much higher than standard pressures of gasification that the reaction can be truly irreversible and the coal completely disappears. Therefore, at normal pressures (less than 1000 atm) one can never observe a net backward
-
744
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 4, 1982
Table V. Equilibrium Constant for Reactions of Char with H, and Steam' equilibrium constant C + 2H, f CH, CH, 3 + 1.85H22 CH, lzck0.3 + 0.2H2fCH, + 11C 6.66CH0., + H, Z C H , + 5.66C C + H , O f C O + H, CH,., + H , D f C O + 1.15H2 6.66CHO., + H,O t 'CH, + CO + 4.66C
1000 K
1100 K
0.089 0.522 1.4 x lo* 1.1x l o 4 2.5 15.2 3 x 104
0.03 5 0.217 9.4 x 10' 6.0 x lo3 11.0 68.9 6.8 X lo4
a The numbers given in the table are upper bounds estimated on the basis that the free energy of formation of the char is A G , ~ 3.5; ~ = AGlioo= 3.95 kcal/gmol.
reaction, which will require a pressure greater than 1O'O atm to be observable. A very fast forward reaction is therefore consistent with the knowledge that the backward reaction is not observed. Reaction 19 at normal gasifier pressures will go to completion, just as devolatilization of coal. It will be misleading to consider this an equilibrium as for a solid-gas reaction complete conversion can still be far from equilibrium. Any reaction that has a very high conversion can occur fast, despite the fact that the backward reaction is not observed. This is not true for reactions where conversion is strongly limited by equilibrium considerations. Irreversible reaction can also occur in gasification of devolatilized char. Consider for example the reaction 12CHo.3 + 0.2H2
-
CH4 + 11C
(20)
At 1100 K and 68 atm Kp = 108 and potential conversion is very high. By reacting 0.2 mole of H2per atom of carbon and maintaining the pressure of 68 atm, the equilibrium partial pressure for H2 for this reaction is atm. The ratio between this pressure and initial pressure is 10%. If reaction 20 is first order the reaction rate at equilibrium will be 10%times slower than the initial rate. The rate of the forward reaction at equilibrium gives a good estimate for the backward reaction. A fast reaction rate of reaction 20 at 68 atm will therefore be consistent with the fact that the backward reaction is too slow to be observed, as the formation of other chars from CHI is 10 times faster. Equation 20 is just a hypothetical overall reaction. A large number of such a reaction is possible between char and hydrogen (or char and steam). (See Table V). Such a mechanism can explain the results of Wen and Huebler (1965). In this reference, hydrogen and devolatilized char were reacted in a bomb at 975 K and 137 atm. Initially, the CHI concentration was high and the ratio of [CH4]/[H2]2exceeded equilibrium. The CHI concentration later decreased with time and reached a final equilibrium value equal to that over graphite. Char conversion was low in these experiments and high char conversion could only be reached with large excess of H2 At high char conversion the ratio [CH4]/[H2I2did not exceed Kp for graphitic carbon. Initial high CH4 yields have been observed by other investigators and it has been suggested that freshly formed char is more active. However, no case has been reported with almost complete char conversion and high ratio of [CH4]/[H2]2in either hydrogen or steam gasification. A reaction mechanism similar to eq 20 followed by gasification of char would explain the experiments of Wen and Huebler and be consistent with the overall knowledge of char gasification. The above example illustrates another application of thermodynamic analysis in reaction engineering that merits
wider application. The knowledge that a backward reaction is slow coupled with an estimate of the equilibrium constant allows us to estimate which forward reactions are likely to occur. The high irreversibility of reactions 19 and 20 was due to the fact that the primary coal or char formed another char with a lower free energy together with the gaseous products. This is a feature of many irreversible decomposition processes. The exact nature of the process may be complex but there is no need to know it for estimating the likelihood of the overall reaction. Very little is known about the nature and extent of irreversible reactions of different chars. All is known is that in hydrogasification, it is possible to achieve much higher yields of CH, per mole of H2 fed at low char conversion than at high char conversion (Wen and Heubler 1965) and that at high char conversion the equilibrium properties of the char are close to that of graphite for both steam and hydrogasification (Blackwood and McGrory, 1957; Johnson, 1974). Less is known about irreversible reactions between char and steam, but again there is no evidence that high steam or char conversions can exceed that of graphite. It is also known that properties of the char can change during gasification (Blackwood and McGrory, 1957). Regrettably, these effects have not been taken into account in kinetic studies of gasification rates, most of which have been carried out at low char conversion. The data are therefore questionable for design at high char and steam conversion. For our purposes this is no obstacle, but for real design proper data are highly desirable. Much more is known about the devolatilization of coal (Anthony and Howard, 1976; Solomon, 1979; Zahradnik and Grace, 1974). It is generally assumed that when coal devolatilizes, it first forms volatiles and char. The volatiles either decompose to char, gases and liquid products or react with steam or hydrogen to form methane, CO, and H2 As the volatiles have a higher free energy compared to coal, they can undergo truly reversible reactions with both steam and hydrogen. The products distribution of the devolatilization of coal is therefore a result of competing reactions (decomposition of volatiles vs. their reaction with steam or hydrogen). It therefore, depends on pressure, temperature, rate of heating, and the composition of the gas in which the coal is devolatilized (Wen and Tone, 1978). These conditions can affect the overall efficiency of the process. Thermodynamics cannot give any information on relative rates. It can only provide bounds on likely product distributions. But once it is realized that gasification of coal is a two-step process, thermodynamic analysis provides guidelines for reactor design. 7.3. Advantage of Prior Devolatilization. The advantage of prior devolatilization can be understood if it is realized that equilibrium is just a function of mass balance and temperatures. The same results will be obtained if instead of coal, the reactor will be fed with char, CHI, and COP It is obvious that feeding the product to the gasifier will reduce conversion of the steam. Rememberring the guidelines obtained from the membrane case, it is preferable to remove products selectively as the devolatilization stage does. At global equilibrium over char, feeding coal without prior devolatilization is completely equivalent to feeding carbon, hydrogen, and oxygen separately (or carbon, CHI, and C02). The only way the coal's chemical structure affects the reactor is in the AH required. Incomplete conversion simply increases the ratios of H/C and O/C fed to the reactor. If conversion is limited and unconverted
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 4, 1982 745 0.77 CH, 0.7 COI
BASIS: 1 M O L E He0 FEE0 6 8 ATM
1.92 CHO.,OO,l
9 0 % CARBON CONVERSION
j F /
0.21co' 0.65 1 ' n, 2
CPTALYTIC
0.73 Cop 4
68otm ~OOO*K
1.0
H
2
S0H I FM T AN0 T
1
1.0 CH,
METHANATION COOLING
STEAM CHo
H2O
O
w>
-
w J 0 0.8 1.92 ~
~
~
.
~
cn,
0.62 COz 0 0.3~CO .
0.5 n2 0 . 9 8 HzO DEVOLlTlLlZER
~
0.73 C o p SHIFT AN0 METHANATION
1.0
cn,
0.58 c o p
BOO 1000 EOUlLlBRlUM TEMPERATURE
600
0.2 co
1200
Figure 25. Real equilibrium yields for steam gasification of Eastem coal. 2.0
1 1.5
w
1
68otm 1OOO'K
I BASIS:
1 MOLE Hz0 F E E D
68 ATM
(bl
Figure 27. Schematic flowsheet for a two stage process: (a) without
9 0 % CARBON CONVERSION
prior devolatilization; (b) with prior devolatilization.
1.0
600
107 H f l
(OK)
800 1000 EOUlLlERlUM TEMPERATURE
1200 (OK)
Figure 26. Real equilibrium yields for steam gasification of Weatem coal.
char is removed from the reactor, the composition of the char and the amount of char removed are both required to compute steam conversion at equilibrium. If char is gasified, it is enough to assume that the overall char composition does not change during the reaction and that the removed char has the same composition as the char in the reactor. For computing the equilibrium yields in this case, only one gasaolid phase reaction is needed. It is assumed to be reaction I11 (the water gas reaction) in Table II. The results for two types of coals are given in Figures 25 and 26. The yields per mole of steam feed are insensitive to conversion if conversion exceeds 90 % . In all the examples with coal, an ultimate conversion of 90% is used. It is assumed that the unconverted char is burned in a fluidized bed boiler to raise steam. Thermodynamically, it is better to burn char instead of coal because we do not want to burn hydrogen, yet it requires the development of boilers which are able to handle the char. This does not mean that the CHI yield cannot exceed global equilibrium. The CHI yield from devolatilization (or reasonable initial reactions) exceeds the global equilibrium value a t high temperatures. The problem is to achieve a high char conversion by the steam-carbon reaction without decomposing the CHI by either reaction VI or X (steam reforming of methane). The latter is catalyzed by all catalysts that promote char gasification and also by many minerals contained in the ash. Designing a reactor to achieve a kinetic optimum between steam conversion and methane yield would be very difficult as the reaction rates vary from coal to coal due to the difference in ash composition and catalytic activity.
Anyway, a two-stage reactor with a separate devolatilization stage can have higher potential steam conversion and better overall thermal efficiency compared to such a kinetic optimum, even if it is achievable, since methane formation in a char gasifier always reduces steam and heat requirements (seeFigure 16). All that is needed to evaluate the potential of an optimal countercurrent gasifier with a devolatilization stage is a knowledge of equilibrium over carbon and an estimate for the products of the devolatiliir. These depend on the conditions in the devolatilizer as well as on the coal, but for our purposes an approximate estimate is enough. Typical devolatilization products for two coals are given by (Yoon, Wei, and Denn, 1977) Eastern coal CHo.sOo.l 0.8CHo,3+ 0.1296CH4 + 0.0176c02 + 0.0528CO + 0.012H20 + O.OO88H2 4
-
Western coal CHo.sOo.2 0.71CHo.3 + 0.1396CH4 + 0.0316C02 0.1128CO + 0.012H20 + O.OO88H2
+
What are the implications of these assumptions? Consider, for example, the gasification-methanation route. In Figure 27 two versions of this process are given. In the first (27a), the coal is introduced directly into reactor 1 and the products are fed to reactor 2. In 27(b), the coal is fist devolatilized and the products are fed to reactor 2. A clear advantage is noted for prior devolatilization in terms of steam requirements. Since devolatilization is fast, it will have very little effect on the composition of the product gas, and it can be achieved by contacting the product gas from reactor 1 with the fresh coal. Prior devolatilization has an additional advantage, which is neglected here. In both cases the coal is heat exchanged with the product gas. In practice this is impossible without devolatilization to occur,and so, to prevent devolatilization, the coal has to be fed cold. Feeding the coal cold into the reactor increases the heat penalty by 3-10 kcal/g-mol of CH,. This is a process constraint that affects the limiting efficiency. If, for process reasons, a higher temperature is required in reactor 1, the advantage of prior devolatilization increases significantly. The equilibrium yield of methane decreases with temperature (see Figures 25 and 26) and
748
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 4, 1982
at high temperatures it is even less than the one obtained from devolatilization alone. Devolatilization therefore has a significant advantage for thermal efficiency. Its disadvantage is that the coal pyrolysis generates tars and agglomerated compounds which may not completely decompose. This not only makes heat recovery difficult but introduces byproducts which in some cases may be undesirable. One way to overcome this is to increase the residence time in the devolatilization stage and choose the temperature such that all undesirable byproducts such as tars and phenols are completely cracked, while the other gasification reactions are slow. While this has not been demonstrated in a working pilot plant, there are catalysts that can achieve this at 900 K. One will do this either in a separate catalytic reactor or in a properly designed devolatilization zone adding the catalyst to the coal. (There are no data to show if the mineral components of the ash are sufficient, but a zeolite as well as many other cracking catalysts will achieve the goal.) This is therefore neither a thermodynamic nor a kinetic constraint but relates to process development and cost considerations. We included here under the overall term devolatilization all the irreversible processes that ultimately result in gaseous products and a graphitic char. In practice the result of these irreversible processes depend on the composition (hydrogen and steam content) temRerature and pressure of the gas used for devolatilization. Our knowledge of these effects is rather incomplete for any optimization. The overall trends and advantages are not sensitive to these effects in the range of conditions discussed here. The quantitative advantage of the prior devolatilization for specific process schemes will be shown in the next section Table VI). 7.4. Comparison of Idealized Limiting Processes. In the previous section it was shown that the lowest steam and heat requirements are obtained in a countercurrent gasifier in which the coal is devolatilized in a separate section with the hot gas from the char gasifier, which receives the hot char from the devolatilization. Therefore, the overall heat requirements strongly depend on the properties of the char gasification section, and therefore, the results of Figure 20 are directly applicable to coal gasification as one can add the devolatilization zone to any of the four processes mentioned. While now each process may have added a separate devolatilization reactor, we will still use the basic nomenclature. Ideal process requirements for the twestage process with or without prior devolatilization are given in Table VI. It is noted that the subbituminous coal has a small advantage over the bituminous coal with prior devolatilization and a disadvantage without it. In noncatalytic gasifiers subbituminous coals have an additional advantage due to higher reactivity, which allows them to approach equilibrium more closely. The oxygen content of the coal is utilized by prior devolatilization in a beneficial way (forming CO), whereas direct feed to the gasification zone wastes the free energy of the coal compared to char. In principle, the same approach can be used for the two Exxon processes. We can devolatilize the feed, send the combined product from both the reactor and the devolatilization to a cryogenic separator, and recycle the CO and H2to the reactor. Figure 28 shows simplified flow sheets for one typical coal. There, too, we note an advantage for the modified Exxon process. If coal is fed without prior devolatilization, we get a significant penalty, as can be seen from Table VI. Devolatilization generates CO and H2.In the Exxon process given in Figure 28 (and also in Table VI), the
I I
-;a1
"u 'u
0
zE l
d l
I
???
Y".?
v m a amarl t- t - P - O D
P-P-t-
2m
2 P-
LD.
rl
@I
m
N
P-
oo*
Id
0 0
0 0
1c?
CJ
1N
m a
N?
r3 r l r 3
x 000
m o o
222
000 8 " 0 0 r3
222 c a
3
cc U
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 4, 1982 747
Table VII. Process Requirements (mol/mol of CH, Final Product)a Western coal ideal two stage Lurgi-dry ash 1100 K 1100 K Lurgi-dry ash 68 25 pressure, atm 25 25 steam
a
3.5
1.57
1.8
5.7
oxygen 0.45 0.27 0.2 0.67 Steam for practical shift requirement to a H,/CO ratio of 1:l included. 0.73 COS 0.35CO 0.65 H z 1.24 H#
CONOEN SAT ION AND ACID GAS REMOVbL
DEVOLITILIZER
0.75 CH4
COOLIhG
0.63H, CRYOGENIC SEPARATION
GASIFICATION
IO)
'."
Cn~.(lO~.i
1 CH4
1 CH4
0.52 COP 0.29t o
0.73Cop
0.6Hz
COOLING
GASIFICATION 6 8 0 l m 1a)O.K
fl
CONDENSATION AND b C l D GAS REMOVAL
SHIFT
DEVOLITILIZER
coz
0.08CO 0.81 n *
COOLING
CRYOGENIC SEPARATION
I
Figure 28. (a) Schematic flowsheet of Exxon process with prior devolatilization; (b) schematic flowsheet of Modified Exxon process with prior devolatilization.
incremental CO and H2formed during devolatilization is recycled to the reactor where it is methanated. The Exxon processes with devolatilization are therefore sensitive to the exact product distribution of devolatilization. The obvious choice is to methanate it in a separate reactor which can be done by methanating a side stream of the product after gas cleaning. One can then reduce the recycle by 30%. While thiscan affect design and cost significantly it has only a small impact on our ideal and limiting efficiency. The heat of methanation evolved in the reactor, at total recycle, reduces the amount of high-temperature heat that has to be supplied to the feed stream which compensates for the losses due to the larger recycle. The two-stage process has a small significant advantage over the two Exxon routes at all temperatures. This small advantage can disappear if other process penalties are taken into account, but what is significant is that this advantage persists even if the temperature of the two-stage process is increased to 1100 K, which allows much smaller reactor sizes. A smaller reactor size can also be achieved by the proposed modification of the Exxon process (shifting the exit gases before recycling). Table VI clearly indicates the strong advantage of the proposed modified Exxon process. It allows the increase of reactor temperature without any penalty in thermal efficiency, maintaining the same recycle and improving steam conversion. (This will reduce one of the main present disadvantages of the Exxon process which is the very large reador size, and merits closer checking in terms
Eastern coal ideal two stage slagger
30 2.8 0.58
1100 K 25 1.74 0.29
1100 K 68 2.0 0.22
of actual process performance). 8. Comparison of Ideal and Real Process Requirements In all processes, high thermal efficiency requires that all feeds should enter the reaction zone as hot as they can be heated without penalty. Steam can indirectly be heated up to 950-1000 K by combusting coal. The coal can be heat exchanged with product gas in a countercurrent moving bed (or in a countercurrent multistage fluid bed). Both steps are essential to minimize heat requirements at high temperatures. Table VI gives an indication of the theoretical limiting efficiencieswhich are close to 80%. In practice 65% is about the highest thermal efficiency achieved. The limiting efficiency does not include the energy requirements for COz removal nor the process requirements and the losses for heating and cooling nor does it include the heat required for catalyst recovery. However, there is still a strong potential for improvement as it can be seen by comparing the process penalty of present gasifiers with the ideal process in Table VI. Table VI1 shows the total steam and oxygen requirements per mole of CHI final product for several gasifiers. We note that present noncatalytic processes have almost double the steam and oxygen requirements. Each mole excess steam has a penalty close to 10 kcal, and each mole of excess oxygen used has a penalty of 40 kcal. There is, therefore, considerable room for improvement. We note that in a Lurgi gasifier, Eastern coal has a significant penalty compared to a typical Western coal. This is not due to any thermodynamic or kinetic reasons of the type described before but due to an additional process constraint specific to moving bed gasifiers. Since the reaction of char with oxygen is much faster than the reaction with steam, it cannot operate with optimal steam to oxygen ratios or highly preheated steam. Operating at these conditions will lead to maximum temperature in the range of ash agglomeration. To overcome this problem, the Lurgi gasifier uses excess steam at lower temperature. The amount of excess steam required to cool the combustion zone depends on the selectivity of the coal, as the endothermic gasification contributes to the cooling. Western coal has here a significant advantage due to its higher reactivity. This advantage applies to all noncatalytic gasifiers operating below 1300 K. A slagger gasifier overcomes this difficulty by using excess oxygen which leads to molten slag. The reactivity of the coal is therefore less important. In principle, fluid beds should be able to overcome this penalty but none of the presently reported results for SNG production has yet achieved this. Our results indicate that a well-designed catalytic gasifier can lead to significant improvement in thermal efficiency, provided the catalyst itself does not involve a significant thermal penalty. Discussion of these design problems and the effect of the conclusions of this paper on design will be discussed in a future paper. 9. Summary and Discussion The article presents a thorough analysis of the thermodynamic constraints which face the developer in de-
748
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 4, 1982
signing an SNG process with present technology and suggests directions for further research and development. The results can be summarized as follows. There are three factors that dominate the thermal efficiency of SNG processes: (a) the conversion of steam in the gasifier. Recovery of steam from a steam-gas mixture is inefficient and expensive and therefore steam utilization is a dominant factor; (b) transferring heat to a reactor at high temperature (>900 K) by combustion of coal involves a significant penalty (0.5 kcal/kcal supplied). The common way of using oxygen inside the reactor is as efficient as any of the other routes suggested till now. It is therefore important to minimize the high temperature heat requirement of the reactor; (c) separation processes in a gas phase have very low efficiencies as compared to their theoretical requirements and have to be minimized. These constraints are not inherent results of the first or second law of thermodynamics, but rather thermodynamic consequences of available technology. AU chemical processes involve irreversible losses in available energy inside the reactor. In SNG production these losses are small. However, the thermodynamic efficiency of the process is a consequence of the kinetic processes as well as process constraints that dominate reactor performance. The best hypothetical process, which will be consistent with the overall state of technology, will be the development of a cheap catalyst that can directly react coal with steam at a temperature of 600 K at a reasonably high pressure. No such catalysts are presently known, and the best process with catalysts presently available is the one described in Figure 27b. Gasification occurs in a two-stage countercurrent reactor in which coal is devolatilized by the product gases from the primary gasification zone. Steam and oxygen are reacted with the char (formed in the first zone) and brought to equilibrium. The gases are then methanated and cleaned (or cleaned and methanated). Such a system can have a significantly better thermal efficiency than present processes (over 70% (LHV) vs. 62% for present processes). To achieve maximum efficiency, the following items are important. (a) Methane formation is maximized in the char gasification zone. Direct methane formation from char and steam has a low heat of reaction, whereas CO and H2 formation, which is the primary reaction for present catalysts, is highly endothermic. Catalysts which form methane by a direct route can have significant advantages. With present catalysts global equilibrium gives the highest achievable methane yield per mole steam feed. This yield decreases with temperature, whereas overall steam conversion increases with temperature. Therefore, reactor heat requirements increase with temperature and the penalty more than negates the benefits of increased steam conversion. The optimum conditions require the reduction of reactor temperature as low as possible. In that sense SNG production is different from syngas production (Shinnar et al., 1978) for which higher temperatures are advantageous. However, the penalty of increasing the temperature by 100 "C is still relatively small. (b) In order to minimize heat requirements in the reactor, all feeds should be preheated as high as can possibly be achieved by heat transfer from coal combustion. ( c ) It is important that the coal is devolatilized, prior to its feeding to the gasification zone, under conditions that maximize CH4 and CO formation. Tars should be preferably cracked in the devolatilizer. If the coal is fed directly, as in a one stage fluid bed or cocurrent gasifier, the
free energy of the coal is not utilized. CH4 formed by methanation in the gasifier limits further methane formation due to equilibrium constraints. Proper devolatilization is therefore essential for all gasifiers, unless a real low-temperature catalyst can be developed. (d) Methanation of CO and H2 should be carried out without additional steam. The unconverted steam from the gasifier can be used to shift the product gas, but the use of additonal steam will involve a penalty in thermal efficiency. The methanator should operate at conditions of H2/C0 ratio close to 1. The present Exxon process, the only catalytic process under large-scale development in the US.,prevents CO and H2 formation by separating the CO and H, cryogenicaly from the methane and recycling them to the reactor. This reduces the heat requirement of the reactor and eliminates the need for oxygen. Here, instead of the penalty for cryogenic separation of air, a similar penalty is required for the cryogenic separation of CHI fron CO and HF The energy requirement for the separation of the recycled gas increases with temperature much steeper (Figure 19) than the oxygen requirements in the two-stage process. These facts force the Exxon process to operate at the lowest feasible temperature. The ability to operate at a 100 OC higher temperature, with the same thermal efficiency, is a significant advantage for a two-stage catalytic process (using oxygen and steam in the first stage and methanation in the second) and leads to smaller reactors or can allow the use of cheaper catalysts. The efficiency of the Exxon catalytic SNG process can be significantlyimproved if the product gas is shifted with its own unconverted steam before the separation. This process modification reduces both recycle and steam requirements by preferentially removing COz from the system and can lead to a significant improvement of the present catalytic gasification process. This modified process permits operation of the gasifier a t a higher temperature (an increase of about 30 to 50 "C)without any penalty in recycle requirements, and leads to higher steam conversion. Operating a t a higher temperature can significantly reduce the large size reactor of the present gasification process which is one of the main disadvantages of this route. The analysis in this article deals with catalytic gasifiers using known catalysts. Catalytic gasification gives the highest thermal efficiencies, at least for idealized processes. It also indicates how much the present processes can be improved by either optimizing their process conditions or by using catalysts. The analysis provides a guideline for such improvements. The literature on SNG production contains some claims that the equilibrium methane yield is a strong function of char properties. However, no such data exist to back up these claims. The preceding article contains a thorough discussion of the thermodynamic constraints of methane formation and presents the argument that the equilibrium constant of methane formation over char is very close to that over graphitic carbon. This argument is backed by experimental data. The experimental results used to indicate higher equilibrium constants for chars are explained by our hypothesis. The article demonstrates that thermodynamic analysis can be a powerful tool in process development. It is especially useful in understanding the effect of kinetic properties of a catalyst (or a noncatalytic process) on the maximum conversion that can be achieved. The main tool of the analysis is the principle of detailed balancing, which is shown to lead to useful and interesting results. The
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 4, 1982 749
Table VI11
C
~~
char properties, kcal/ gmol G AH^^^^ ~ ~ TAS",, ~ ~
assumption for char structure A graphite (reference) n-hexane cyclohexane char I (benzene) char I1 ( m -xylene)
gasification (A-7) A G aHa2* ~ ~ 23.9
41.9
-0.02 t0.16 t1.55
-1.0 -0.73 1.0
-0.98 -1.09 -0.55
23.9 23.7 22.3
42.9 42.8 40.9
t0.77
-0.18
-0.95
23.1
42.0
~
Acknowledgment The work reported here was supported by DOE Contract No. DE AC01-79 ET14811, and the authors want to express their appreciation for the cooperation extended by the division of coal gasification at DOE. The opinions expressed in the paper are solely those of the authors and do not reflect any position of DOE.
Appendix. Estimation of Thermodynamic Properties of Coal and Char A. Coal at Room Temperature. No accurate or even approximate measurements of the thermodynamic properties of coal are available in the literature (IGT, 1978-1979). Some approximations have been suggested (Krikosian, 1972) as used by IGT (1978-1979). In the theoretical approximation some assumptions about the coal structure have to be made. The data for coal for TASomof a structure CH,O,NpS, were given by Krikosian (1972) as = -0.05689 - 3.6909n - 6.109m - 5.627~- 1.079q (A-1)
TAS02g8
where TASo298 is in kcal/g-mol. For a typical Eastern coal (Detman, 1976) the structure of the dry coal is equivalent to CHo.800.1No.ozSo.023. If we neglect the N and S, or substitute them for an extra oxygen, a typical Eastern coal will be considered as CHo,800,1. For the coal, TAS0298 from eq A-1 is -3.62 kcal/g-mol. AHo can be determined by the heat of combustion. For the above coal (Detman, 1976) the heat of combustion is 14800 Btu/lb or 8.22 kcal/g, or 8.22 X 14.4 = 118.4 kcal/g-mol of coal. The Moms of the coal is obtained by using the combustion data
mocoal
= -94.054
=
AHoC02
-
COP + OAH,O(l)
+ 0*4AH0HzO(l)-
(A-2)
AHcomb
+ 0.4 X (-68.3) + 118.4 = -3.0
kcal/g-mol
AGO can be calculated using the relation AGO298
= AH0298 -
TASo2g8
= -3.0 -(3.62) =
0.62 kcal/g-mol Those are used as typical values for the Eastern coal when compared to graphite AH0298
= -3.0 kcal/g-mol
= +0.62 kcal/g-mol
B. Effect of Temperature. The correction for temperature was done by using the data of (HT- Hm8) for the formation reaction
-
CH,O,
the molecular weight of the coal is MWd = 12 + n m
application of the analysis to existing gasifiers will be presented in a separate article.
CHo.8Oo.l + 1.1502
+ (n/2)H2 + (m/2)02
(A-3)
+ 16 X
~
The ST- s298 and H T - Hzg8are available from standard JANAF tables or can be calculated by the correlation for Cp of hydrogen, graphite, and oxygen. The AGO is calculated from AGoT = A H O T - TASoT. C. Char. The char in most cases is composed of very little oxygen and has about 0.2-0.4 molar ratio of hydrogen to coal. For our demonstration a typical char would be assumed to be CHo.3. The char structure is not known. Following different experimental information, it could be assumed to be aliphatic, cycloaliphatic, or aromatic. The activity in the gasification reaction is expected to be greater as we move from aliphatic to cycloaliphatic, to aromatic. The calculation would be according to the formation reaction, assuming that the char is a physical composition of graphite and the structure which is assumed. The three options are shown in 0.874C + 0.021C6H14 CHo.3 (A-4) 0.85C 0.025C&12 CH.3 (A-5)
--
+
0.70C + 0.05C& CHo.3 (A-6) Following eq A-3 the m 0 2 9 8 of the char are AGO298 = 0.021 X ASGon-hexme(l),298 = 0.021 X (-0.91) = -0.02 kcal/g-mol AH0298
= 0.021 x
Ak?on4,exme(l),298
= 0.021 x (-47.52) =
-1.02 kcal/g-mol The results are summarized in Table VIII. Also are given the corresponding value for the gasification reaction (A-7) CHo.3 + H20(1) CO + 1.15H2 As expected, the gasification reaction is enhanced as the char structure to aliphatic and aromatic. The AGO, value for the n-hexane is almost the same as for graphite. Literature Cited +
Air Products and Chemical Corp. "Cryogenic Methane Separation-Catalytic Hydrogaslflcatlon Process Analysis", Final Report, Allentown, PA, 1980. Ark, R. "Elementary Chemical Reactor Analysis"; Prentice-Hall, Englewood Cliffs, NJ, 1969. Blackwood, J. D.;McGrory, F. Aust. J. Chem. 1957, 1 1 , 16-33. Cabrera, A. L.; Heinemann, H.; Somoryai, G. A. "Methane Production from Catalyzed Reaction of Graphtte and Water Vapor at Low Temperatures (500-600 K)"; LBL-12812, 1981. Dent, J. F., et al. "An Investigation Into the Catalytic Synthesis of Methane by Town Gas Manufacture"; 49th Report of the Joint Research Committee, University of Leeds, 1945. Detman, R. "Factored Estlmates for Western Coal Commercial Concepts"; FE-2240-5, 1976. Exxon Research & Englneerlng Co. "Exxon Catalytic Coal Gasification Process"; FE-2369-24, 1978. Faith, E. L.; Vermeulen, T. AIChE J. 1961, 13, 936. Flower, J. R.; Linnhoff, B. "Thermodynamic Analysis in the Design of Process Networks"; paper presented at conference C.A.C.E. 79, 214th event of EFCE, Montreux, Switzerland 1979. Gaggioli, R. A,; Petik, P. J. ACS Repr. 1976, 27, 56. Gruber, 0. A&. Chem. Ser. 1975, No. 146, 31-46. IGT "Coal Gasification Pilot Plant Support Studies"; FE-2806-1,2,3,4, 1978-1979. Johnson, J. L. A&. Chem. Ser. 1974, No. 131, 145. Krambeck, F. J. ARMA. 1970, 38, 317.
Ind. Eng. Chem. Process Des. Dev. 1982, 21, 750-760
750
Krikosian. 0. M. “Thermodynamics Propetties of Wyoming Coals”; Technical Note 72-18, UCID-16587, Livermore, CA, Lawrence Livermore Laboratory, 1972. May, W. G.; Mueiler, R. H.; Sweetser, S. B. Ind. Eng. Chem. 1058, 5 0 , 1289. Rostrup-Nlelsen, J. R. J . Catal. 1972, 27, 343. Othmer, H. G. (2”.Eng. Scl. 1078, 37,993. Shlnnar, R. CHEMTECH 1978. 8, S86. Shinnar, R.; Kuo, J. C. W. “Gasifier Study for Mobil Coal to Gasoline Processes”; FE-2766-13, 1978. Shinnar, R.; Shapka, D.; Zakai, S. Ind. Eng. Chem. Process Des. D e v . 1081, 20, 581. Shinnar, R.; Feng, C. “Thermodynamic Constraints of Catalytic Processes”; 1981, to be published.
Squires, A. M. Trans. Inst. Chem. Eng. 1061, 39,3. Stom, and Webster Engr. Co. “Comparative Evaluation of I#&I and Low T e n peratwe Gas Cleaning for Coal Gasiflcation-Combined Cycle Power Systems”; EPRI-AF-416, 1977. Wei, J.; Prater, C. D. A&. &tal. 1082, 13. 204. Wen, C. Y.; Huebler. J. I d . Eng. Chem. Process Des. D e v . 1985, 4 , 147. Wen, C. Y.; Tone, S. ACS Symp. Ssr. 1078, 72, 56. Yoon, H.;Wei. J.; Denn. M. M. “Modelling and Analysis of Moving Bed Coal Gasifiers”; Final Report, EPRI AF-590, 1977. Zaharadnik, R. L.; Grace, R. J. A&. Chem. Ser. 1974, No. 131, 127.
Received for review July 14, 1981 Accepted May 11, 1982
Hydroisomerization of n -C5 and n -C6 Mixtures on Zeolite Catalysts James J. Spivey’ Research Triangle Institute, Research Triangle Park, North Carolha 27709
Phllllp A. Bryant’ Chemical Engineering Department, Louisiana State University, Baton Rouge, Louisiana 70803
Hydroisomerization of n C , and n-C, mixtures was carried out on a Pt-H-mordenlte and a Pd-H-faujasite catalyst. A different type of anomalous behavior was observed on each catalyst in the conversion of mixtures relative to the conversion of pure components. Statistical analysis of various Langmuir-Hinshelwood type models for the hydroisomerization rate constant show that this behavior can be related to the adsorption parameters of each system.
Introduction Hydroisomerization of low molecular weight paraffins has become an increasingly important conversion process in recent years. This is due to the phase-down of antiknock concentrations in motor gasoline and the resulting need for higher octane, Le., more branched hydrocarbon constituents. The research octane number of streams containing C5 and (& mostly straight-chain hydrocarbons can be boosted from about 70 to over 80 with the extent of improvement depending on the isomerization temperature and C 5 / C 6ratio (Kouwenhoven and Langhout, 1971). Zeolite catalysts have been employed almost exclusively in new efforts in this area. Paraffin isomerization requires a stable catalyst with high activity to take advantage of the higher equilibrium conversions to branched products at lower temperatures (Chick et al., 1977). Zeolites have been shown to be highly active for this type of reaction and have additional advantages such as relative insensitivity to moisture, sulfur, and nitrogen (Minachev et al., 1972). Little has been reported in the literature on systematic studies of mixed n-Cs and n-C6 feeds although it is this type of feed that is typical of hydroisomerization feed stocks for octane enhancement. Any anomalies observed in hydroisomerization of such mixed feeds are important in the design of commercial systems and in the investigation of the mechanism and dynamics on a molecular level. Exxon Research and Development Labs, Baton Rouge, LA 70821. 0196-4305/82/1121-0750$01.25/0
This study examines the kinetics of n-C5/n-C,hydroisomerization on a mordenite and a faujasite typical of developmental catalysts. The objectives are to describe the effect of reactant partial pressures on a simplified rate constant, explain any anomalies in terms of fundamental kinetic quantities, and compare the performance of the two catalysts in terms of kinetic parameters.
Zeolite Catalysts Zeolites are a special class of crystalline aluminosilicate compounds which have specific pore dimensions and pore structure. These pores may be tailored to exclude certain reactants and thus obtain high selectivity. In addition, the special acidic and adsorptive properties of the zeolite surface promote specific conversion reactions such as hydroisomerization. By introducing both noble metal and acid sites into a zeolite using special techniques, a dual function catalyst is formed. Such a catalyst is active for isomerization. It is thought that hydroisomerization of a normal paraffin on such a catalyst proceeds through a mechanism whereby olefiis are formed at the metallic site by dehydrogenation of the normal paraffin feed and are then adsorbed at an acidic site or the catalyst surface. A t this acidic site, a carbonium ion is formed which undergoes skeletal rearrangement. The reeulting isocarbonium ion is converted to an isoolefin which is then hydrogenated at a metallic site and desorbed (Kouwenhoven, 1973). Mordenite. Mordenite is unique among zeolites in that the structure consists of parallel elliptical channels which do not intersect. Mordenite is the most active catalyst 0 1982 American Chemical Society
