I n d . E n g . Chem. Res. 1991,30, 50-55
50
[alindene, 130797-29-4; fluoranthene, 206-44-0.
Literature Cited Allen, D. T.; Petrakis, L.; Grandy, W.; Gavalas, G. R.; Gates, B. C. Determination of Functional Groups of Coal-derived Liquids by n.m.r. and Elemental Analysis. Fuel 1984,63,803. Beveridge, G. S. C.; Schechter, R. S. Optimization: Theory and Practice; McGraw-Hill: New York, NY, 1970. Bhinde, M. V. Ph.D. Thesis, University of Delaware, 1979. Constant, W. D.; Price, G. L.; McGlaughlin, E. Hydrocracking of Model Coal-Derived Liquid Components over a Zeolite Catalyst. Fuel 1986,65,8. Gates, B. C.; Katzer, J. R.; Schuit, G. C. A. Chemistry o f Catalytic Processes; McGraw-Hill: New York, NY, 1979. Guin, J.; Tarrer, A,; Taylor, L.; Prather, J.; Green, S., Jr. Mechanisms of Coal Particle Dissolution. Ind. Eng. Chem. Process Des. Deu. 1976,15, 490. Haynes, H. W.; Parcher, J. F.; Heimer, N. E. Hydrocracking Polycyclic Hydrocarbons over a Dual-Function Zeolite (Faujasiteb Based Catalyst. Ind. Eng. Chem. Process Des. Deu. 1983,22, 401-409. Huang, C.; Wang, K.; Haynes, H. W. Hydrogenation of Phenanthrene over a Commercial Cobalt Molybdenum Sulfide Catalyst Under Severe Reaction Conditions. In Liquid Fuels f r o m Coal; Ellington, R. T., Ed.; Academic Press: New York, NY, 1977;p 63. Lapinas, A. T. Catalytic Hydrocracking of Fused-Ring Aromatic Compounds: Chemical Reaction Pathways, Kinetics and Mechanisms. Ph.D. Thesis, University of Delaware, 1989. Lapinas, A. T.; Klein, M. T.; Macris, A.; Lyons, J. E.; Gates, B. C. Catalytic Hydrogenation and Hydrocracking of Fluoranthene: Reaction Pathways and Kinetics. Ind. Eng. Chem. Res. 1987,26, 1026. Lemberton, J.; Guisnet, M. Phenanthrene Hydroconversion as a Potential Test Reaction for the Hydrogenating and Cracking Properties of Coal Hydroliquefaction Catalysts. Appl. Catal. 1984,13,181. Petrakis, L.; Allen, D. T.; Gavalas, G. R.; Gates, B. C. Analysis of Synthetic Fuels for Functional Group Determination. Anal. Chem. 1983a,55, 1557. Petrakis, L.; Ruberto, R. G.; Young, D. C.; Gates, B. C. Catalytic Hydroprocessing of SRC-I1 Heavy Distillate Fractions. 1. Ind. Eng. Chem. Process Des. Deu. 1983b,22,292-298.
Petrakis, L.; Young, D. C.; Ruberto, R. G.; Gates, B. C. Catalytic Hydroprocessing of SRC-I1 Heavy Distillate Fractions. 2. Detailed Structural Characterizations of the Fractions. Ind. Eng. Chem. Process Des. Dev. 1 9 8 3 ~22, 298. Pines, H. The Chemistry of Catalytic Hydrocarbon Conversion; Academic Press: New York, NY, 1981;p 177. Qader, S. A. Hydrocracking of Polynuclear Aromatic Hydrocarbons Over Silica-Alumina Based Dual Functional Catalysts. J . Inst. Pet. 1973,59, 178-187. Qader, S. A.; Hill, G. R. Development of Catalysts for the Hydrocracking of Polynuclear Aromatic Hydrocarbons. Prepr. Pap.Am. Chem. Soc., Diu. Fuel Chem. 1972,16,93-106. Qader, S. A.; McOmber, D. B.; Wiser, W. H. Evaluation of Mordenite Catalysts for Phenanthrene Hydrocracking. Prepr. Pap.-Am. Chem. Soc., Diu. Fuel Chem. 1973,18, 127-137. Salim, S. S.; Bell, A. T. Effect of Lewis Acid Catalysts on the Hydrogenation and Cracking of Three-Ring. Fuel 1984,63,469-476. Shabtai, J.; Veluswamy, L.; Oblad, A. G. Steric Effects in Phenanthrene and Pyrene Hydrogenation Catalyzed by Sulfided Ni-W/ alumina. Prepr. Pap.-Am. Chem. Soc., Diu. Fuel Chem. 1978, 23, 107-1 12. Stephens, H. P.; Chapman, R. N. T h e Kinetics of Catalytic Hydrogenation of Pyrene: Implications for Direct Coal Liquefaction Processing. Prepr. Pap-Am. Chem. Soc., Diu. Fuel Chem. 1983, 28(5),161. Sullivan, R. F.; Scott, J. W. The Development of Hydrocracking. In Heterogeneous Catalysis: Selected American Histories; Davis, B. H., Hettinger, W. P., Eds.; American Chemical Society: Washington, DC, 1983;p 293. Unzelman, G. H.; Gerber, N. H. HydrocrackingToday and Tomorrow. Pet. Chem. Eng. 1965,Oct, 32. Veluswamy, L. R. Catalytic Hydroprocessing of Coal-Derived Liquids. Ph.D. Thesis, University of Utah, Salt Lake City, UT, 1977. Ward, J. W. The Varieties of Hydrocracking. Hydrocarbon Hydroprocess. 1975,S e p t , 101-106. Wu, W.; Haynes, H. W. Hydrocracking Condensed-Ring Aromatics Over Nonacid Catalysts. ACS Symp. Ser. 1975,20,65-81. Wuu, S. K. Zeolite Catalysts for Hydrocracking Polynuclear Aromatics: Phenanthrene Kinetics. Ph.D. Thesis, University of Alabama, 1983. Received for reuieur March 12,1990 Accepted August 6, 1990
Modeling Catalytic Gauze Reactors: Ammonia Oxidation Daniel
A. Hickman a n d Lanny D. Schmidt*
Department of Chemical Engineering and Materials Science, University of Minnesota, Minneapolis, Minnesota 55455
The performance of the Ostwald process for NO synthesis from NH, and O2on a Pt gauze catalyst is simulated by using rate equations obtained from well-defined surface experiments. Variations in the selectivity with composition, total pressure, and gauze temperature are calculated. The dependences of the adiabatic gauze temperature on the feed composition and flow rate are also examined. At atmospheric pressure with 10% NH, in air and a reactor temperature of 1100 K, the model predicts a maximum NO selectivity of 95%, which is close to performance observed. The model also predicts that the NO selectivity increases slightly with decreasing temperature and with increasing pressure. For a fixed residence time, the adiabatic temperature rise has a maximum with respect to the ammonia composition. This maximum shifts to higher fractions of ammonia as the residence time increases. For a given feed composition, the adiabatic temperature rise decreases as the flow rate increases. Introduction Gauze catalysts are used commercially for HCN synthesis from CH4,NH3,and O2 (the Andruasow process) and for NH3 oxidation followed by homogeneous oxidation and
* Author to whom inquiries should be addressed.
hydration to HN03 (the Ostwald process) (Satterfield, 1980; Twigg, 1989). In this study, we have developed a simple model for ammonia oxidation that predicts the behavior of the industrial process with reasonable accuracy. The ammonia oxidation reaction is typically operated with 10% NH, in air over a Pt-10% Rh gauze catalyst in an adiabatic atmospheric pressure reactor. The catalyst
0888-5885191/ 2630-0050$02.50/0 0 1991 American Chemical Society
Ind. Eng. Chem. Res., Vol. 30, No. 1, 1991 51 consists of 10-50 layers of woven Pt-10% Rh forming a gauze several millimeters thick and up to several meters in diameter. Linear velocities of the gases are on the order of 100 cm/s, giving a total contact time of about 1 ms with a negligible pressure drop. Generally, the gases are preheated to about 400K, with the exothermicity of the reactions heating the gauze to about 1100K. In a typical reactor, all of the NH3 is converted with over 90% conversion of NH3 to NO. The overall reactions that produce NO and N2 are assumed to be NH3 + V 4 0 2 NO + 3/2H20 (AH = -54 kcal/mol) (1) and
species. The Andrussow process also operates at a higher temperature (1100 K) and uses excess fuel, while ammonia oxidation uses excess OF Nevertheless, these two processes are very similar, and the results from the two models should have comparable interpretations. In this paper, we simulate the behavior of the Ostwald process. The rate expressions that we use have been obtained experimentally on polycrystalline Pt at low pressures, where no mass-transfer limitations exist. Our objectives are (1) to be able to predict the performance of the process under conditions far away from those currently used, (2) to test the validity of the rate equations in a multicomponent, high-pressure situation, and (3) to test the validity of a simple model of flow and temperature in the gauze.
Y2Nz+ 3/2H20 (AH = -75 kcal/mol) (2) However, these overall reactions do not describe the mechanism of the surface reactions. From many studies of the unimolecular and bimolecular reactions in this system (Pignet and Schmidt, 1974,1975; Hasenberg and Schmidt, 1985, 1986, 1987; Loffler and Schmidt, 1976; Mummey and Schmidt, 1981; Takoudis and Schmidt, 1983; Blieszner, 1979), it has been suggested that the following reaction steps might occur:
Model The gases follow a tortuous path through the gauze, so a reasonable model of the process is a plug flow tubular reactor with only axial concentration gradients. Because of the high rate of heat conduction throughout the gauze and radiation within the gauze, we model the gauze a t a uniform temperature throughout. Therefore, assuming steady-state and neglecting diffusion, the following mass balance applies:
-
NH3 + 3/4O2
-
-
NH3 + 7 4 0 2
-
NH3 + $$NO NH3
NO
-
+ 3/2H20 5/4N2 + 3/H20
NO
(3) (4)
YZNZ + Y2H2
(5)
1/N2 + 1 / 0 2
(6)
(7) Hz + 7 2 0 2 HzO These are the major reactions observed when one or two of the individual reactants are passed over Rh and Pt surfaces, and the products in the above reactions are the major species observed in industrial reactors. Relatively small amounts of N20and NO2are observed in laboratory experiments; N20 is only important at temperatures below 700 K (Pignet and Schmidt, 1975; Heck et al., 1982), while NOz is formed by successive homogeneous oxidation of NO. Many investigators have attempted to model this and similar oxidation processes (Trimm, 1983). Most of these models have been concerned with the flow pattern and temperature gradients in the reactor. Likewise, many investigators have postulated mechanisms for this process, but most have limited their arguments to qualitative explanations. Over the past several years, we have been studying the kinetics of these reactions over wide ranges of temperature, total pressure, and gas composition on Pt and Rh surfaces (Pignet and Schmidt, 1974, 1975; Hasenberg and Schmidt, 1985, 1986, 1987; Loffler and Schmidt, 1976; Mummey and Schmidt, 1981; Takoudis and Schmidt, 1983; Blieszner, 1979). We recently used these rate expressions to predict the behavior of the Andrussow process (Waletzko and Schmidt, 1988), which is described by the following overall reactions: +
CH4 + NH3 + 3 / 2 0 2
and
-
-
HCN + 3H20
(8)
CHI + NH3 + Y4O2 f/2N2+ CO + y2H20 (9) The model required for this reaction system is more complex than for ammonia oxidation since approximately 13 reactions are involved and the feed consists of 3 reactant
where ci is the molar concentration of component i (S total components),u is the linear velocity of the gas in the gauze, A / V is the surface area of catalyst per unit volume of gas, No is Avogadro's number, rj is the rate of the jth reaction (in molecules per catalyst area per time, R total reactions), and vi, is the stoichiometric coefficient of component i in reaction j. A t this point, we define a new variable t = z/u
(11)
where t is the time the gas has been in the plug flow reactor. The total residence time is defined as 7
= L/u
(12)
where L is the length of the plug flow reactor. Then, with pi being the partial pressure of component i and T the gas temperature, the mass balances become
Partial pressures are used in the model instead of concentrations because of the form of the rate expressions. Thus, we neglect the small effects of changes in the total number of moles. In fact, since the feed consists mostly of N2 in reactant air, the total number of moles typically . assume that the gas temperachanges by only ~ 3 % We ture is identical with the gauze temperature. The total rate of heat generation Q (per unit volume) is then
We solve eq 13 with rate data from Table I for a given feed composition and reactor temperature. Reactions that form NO2and N20 have been ignored. By use of the operating conditions typically found in the Ostwald process, a value of 7 that gives a particular conversion of NH3 can be calculated. By fixing this value, the reactor length is fixed. For these flow conditions, the effects of variations in the composition, temperature, and pressure can be evaluated.
52 Ind. Eng. Chem. Res., Vol. 30, No. 1, 1991 Table I. Reaction Rate Expressions
rate expression (molecules/(cm*s); P, Torr; T , K)
reaction
NH3
+ -54 0 2
NO H2
-+
+
+
NO
1 - N, 2 1 -2 0,
-C
Pignet and Schmidt, 1974, 1975; Hasenberg and Schmidt, 1985, 1986, 1987
+ -32 HZO
+ -21 02
ref
5.53 x 10l6 e x p ( 2 6 2 5 / T ) P ~ o 1
+ 6.96 X lo4 exp(4125/T)P~o+ 1.56 exp(4775/T)Po2 1.5 x 101gPOzpH,
HZO
Mummey and Schmidt, 1981
The integral energy balance (14) was used to compute the adiabatic temperature rise for specified feed conditions, reactor geometry, and feed velocity. Since the reactor temperature is a constant, its value for a given inlet temperature can be found by computing the composition at an assumed T , determining T from eq 14, and iterating until the temperature is found. For the adiabatic temperature rise calculations, we have assumed that both the gas density, p , and heat capacity, C are constant throughout the reactor. Since the gas pcase is typically about 70% N, and the reactor is at a uniform temperature, both of these assumptions are reasonable. We have also assumed that the heat capacity of the reacting gases is independent of temperature. Typically, the heat capacity changes by only about 10% between the 400 K feed and the 1100 K product. The model calculates conversions as a function of 8, which is related to the residence time by the expression
Assuming ideal gas behavior and a constant molar flow rate (F)and pressure (P)across the gauze, the gas velocity in the gauze is URT
(16) where A, is the cross-sectional area of flow through the catalyst. Likewise, the gas velocity uo before the gauze is
where A. is the cross-sectional area of flow before the gauze. Substituting these equations into (15), we see that
Hence, the model variable 8 is linearly related to the reactor length for a specified feed temperature, feed velocity, and reactor geometry. Rate Expressions The rate expressions shown in Table I are based on experimental data from this laboratory. All experiments were performed on polycrystalline Pt wires or foils for reactant pressures between lo-, and 10 Torr with substrate temperatures between 500 and 1500 K.
Blieszner, 1979
Rates cannot exceed the fluxes of reactants to the surface; thus, a flux limit is imposed on all rates. This flux limit is approximated as
where si and M iare the sticking coefficient and the molecular weight of component i. In the program, each calculated rate is compared to the flux limit of each of its reactants and the smaller of these two values is used as the rate. A t the flux limit, this approximation merely transforms the rate of a given reaction from having Arrhenius temperature dependences and strong pressure dependences to being first order in the limiting reactant, zero order in the other reactant, and inversely proportional to the square root of the temperature. A sticking coefficient of 0.1 was assumed for all components except 02,for which a sticking coefficient of 0.04 was used. Although exact values for the sticking coefficients have not been measured directly for these conditions, these values agree qualitatively with sticking coefficients measured on Pt under ultrahigh vacuum conditions (Waletzko and Schmidt, 1988). The sticking coefficient of oxygen was used as an adjustable parameter; its value was chosen so that the NO selectivity was optimized at 10% NH3 for industrial conditions. The rate expression used for the ammonia oxidation reaction (Table I) was modified slightly from the previously reported expression (Waletzko and Schmidt, 1988) by making the rate first order in oxygen rather than half order. This change was prompted by evidence in other similar reaction systems in which the oxidation reaction is first order with respect to oxygen, even though it adsorbs dissociatively (Hasenberg and Schmidt, 1987). The modified expression still gives a reasonable fit to the original experimental data. In addition, the results support the hypothesis that this reaction is actually flux-limited at typical operating conditions for the Ostwald process. Results Figure 1 shows a typical calculation of the component partial pressures and the individual reaction rates as a function of reaction time for a feed of 10% NH, in air at 1100 K and atmospheric pressure. Complete NH3 con(0 is defined by eq version is obtained for 0 ? 10 x 18). Thus, 8 = 10 X is used to set the reactor length for later calculations. As shown in Figure lb, NO production is initially much faster than the formation of N2. However, as the NO
Ind. Eng. Chem. Res., Vol. 30,No. 1, 1991 53 100
R
I
HP
No
I
98
96
94
92
0
5
10 90
0
4
12
8
16
20
% Ammonia Figure 3. Selectivity (without NH, recovery) of NO formation versus percent NH, in the feed at 1100 K and 8 = 10 X
is evident that these two reactions are the most significant ones throughout the reactor for these operating conditions. Figure 2 plots the NO selectivities at a fixed total pressure of 1 atm and a fixed residence time (e = 10 X We define the selectivity of NO production as 5
0
10
SNO
0x Figure 1. (a) Component partial pressures as a function of residence time (proportional to position) in the plug flow reactor. (b) Reaction rates as a function of residence time in the plug flow reactor. The NO decomposition rate is much slower and is not shown on this plot. All three plots are for 10% NH, in air at a reaction temperature of 1100 K and atmospheric pressure. loo
84
,
I
-
0
4
8
12
16
20
% Ammonia Figure 2. Selectivity of NO formation versus percent NH3 in the feed for a fixed reaction time 8 10 X and atmospheric pressure operation. For each reaction temperature, the two selectivities (with and without NH3 recovery) diverge at higher levels of NH,. The upper branch of each divergent curve is the selectivity with NH3 recovery. The experimental data pointa are for operation at 7.7 atm with gauze exit temperatures between 1190 and 1220 K (Heck et al.,
-
1982).
pressure increases, the rate of NO reduction to N2increases until this reaction is the major cause of decreasing selectivity with increasing ammonia conversion. Eventually, both NO and N2 production rates are limited by the flux of NH3 to the surface (0 > 7 X From this plot, it
=
pNO
if NH3 is not recovered and as SNO
=
pNO ~'ONH~
- PNH~
if NH3 is recovered. These selectivities of course converge when all of the NH, is consumed. The NO selectivity decreases as the temperature increases for a given feed composition. However, the maximum selectivity only shifts from about 11% NH3 at 900 K to about 9% NH3 at 1300 K. The selectivity curves diverge when ammonia conversion is incomplete for a given 8; therefore, the choice of 0 is not important in determining the optimal selectivity as long as it is chosen large enough for complete conversion. In addition, because the rate of the NO decomposition reaction is negligible and there is no other mechanism in the model for NO loss after all NH3 has been consumed, the upper selectivity curves correspond to an infinite residence time. The effect of total pressure on NO selectivity at 1100 K is shown in Figure 3. As the pressure increases from 1 to 5 atm, the maximum selectivity increases from 95% at 10% NH, to 99% at 11% NH,. For a given feed composition, higher pressures are expected to improve the NO selectivity. Calculations of the adiabatic reactor temperature for a preheat temperature of 400 K are presented in Figures 4 and 5. As mentioned earlier, C, is assumed to be a constant, independent of composition and temperature. This constant was set by calculating the value of C needed for a 400 K feed to heat to 1100 K for 10% Nfi, in air at atmospheric pressure. The heat capacity was found to be 8.1 cal/(mol K) for this case. Between 400 and 1100 K, the actual value for N2 ranges from 7.1 to 7.8 cal/(mol K). For large residence times, Figure 4 demonstrates that an upper limit in the temperature rise exists for small enough levels of ammonia. For feed concentrations less than lo%, the temperature rise is linear with respect to
54 Ind. Eng. Chem. Res., Vol. 30, No. 1, 1991
Applicability of the Model
h
% 41
2
1203
800
400
n
0
8
4
12
16
20
% Ammonia Figure 4. Adiabatic temperature rise for various residence times (xlO-%)as a function of percent NH, in the feed. All curves were calculated for a feed inlet temperature of 400 K. The data points are from experiments at pressures ranging from 1 to 10 atm (Twigg, 1989). I-
800
-
h
% 41
2 2
600-
Q)
c1
kE
400-
F
200
-
0.0
0.2
0.4
flow rate,
0.6
1/
0.8
I .O
( e x iozo)
Figure 5. Adiabatic temperature rise for an inlet temperature of 400 K as a function of flow rate (which is proportional to the inverse of the residence time) for various NH3 levels in the feed.
ammonia concentration. Above 1070, the slope of the limiting curve gradually increases with increasing levels of ammonia. The NO formation reaction, eq 3 (AH = -54 kcal/mol), dominates in the smaller concentration region, while N2 formation, eq 4 (AH = -108 kcal/mol), becomes significant at higher levels of NH, in the feed. For each choice of 0, the adiabatic temperature rise reaches a maximum and decreases with a relatively small slope. In this region beyond the maximum, the endothermic ammonia decomposition reaction begins to have a significant effect on the system. Figure 5 shows the effect of the gas flow rate (which is inversely proportional to 0) on the adiabatic temperature for various feed compositions at 400 K. At low enough flow rates, the ammonia reacts completely. Apparently, only the rates with NH3 as a reactant are significant for these conditions, so the temperature reaches a maximum as l/0 approaches zero. (The 15% and 20% curves are off-scale for very low flow rates.) As the flow rate increases, NH, consumption is incomplete, resulting in decreasing reactor temperatures.
This model gives results consistent with published data on the Ostwald process: a 95-99% yield of NO with the maximum selectivity occurring near 10% NH3 at a reactor temperature of 1100 K (Twigg, 1989). We also observed a broad maximum in the NO selectivity, although we predict that increasing the NH, level much beyond the optimal 10% should drastically reduce NO production. The model predicts slightly improved performance at higher operating pressures. Unfortunately, the dependence of the commercial Ostwald process on pressure is not clear. Apparently contradictory claims of higher selectivity (Handforth and Tilley, 1934) and lower selectivity (Satterfield, 1980; Twigg, 1989) at higher pressures exist. However, different operating pressures in commercial plants are accompanied by other changes, such as changes in the mass or volumetric flow rate. Hence, it is difficult to determine the effect of pressure independent of other process variables. The model prediction of increasing selectivitywith lower temperatures and lower NH3 levels cannot be easily tested. Industrial reactors operate adiabatically, so the temperature is fixed once the feed conditions are established. In addition, increasing the percentage of NH3 will increase the adiabatic reactor temperature. A reactor could be cooled by providing some type of heat conduction path from the gauze or by eliminating any existing radiation shields. However, cooling the reactor may not be advantageous since the flow rate must then be reduced to attain the same NH3 conversion. Likewise, preheating the reactants or enriching the feed with O2should not improve the selectivity since these approaches would increase the adiabatic temperature. A comparison of this model with the Andrugsow process model reveals some interesting similarities and differences. For that system, an optimal HCN selectivity was predicted (Waletzko and Schmidt, 1988) with respect to both the CH4/02and NH3/O2 ratios, as is observed experimentally. The HCN model predicted higher selectivities, and the selectivity increased slightly with increasing total pressure. However, the maximum selectivity for HCN synthesis was less than 6070,which is significantly lower than that for NH, oxidation. The NH, oxidation reactions are undesired side reactions for HCN synthesis, but since the Andruasow process operates at higher temperatures (1400 versus 1100 K),the rate of the NH, oxidation reaction decreases enough to give the needed HCN selectivity. In addition, we have shown previously that CN adsorbed on the catalyst surface blocks the NH3 decomposition reaction (Hasenberg and Schmidt, 1987; Hwang and Schmidt, 1987). The crucial issues for this model are (1)the validity of the rate expressions and (2) the validity of the simple plug flow model. The rates in Table I were obtained at lower pressures in the absence of other competing reactions and may not be accurate at atmospheric pressure with multiple components. As discussed in the previous paper (Waletzko and Schmidt, 1988), the decoupling of multiple surface reactions presents a fundamental problem in all catalytic processes because all other species can act as poisons or promoters, and the reactants must partition between various reaction channels. Hence, this model is valid only if the species coverages are low enough to validate the approximation of independent reactions. The absence of carbon-containing species and the high operating temperatures should make this approximation valid. Neglect,ing radial composition variations or masstransfer effects in the simple plug flow model presents another possible problem. However, the model may still
55
I n d . Eng. Chem. Res. 1991,30, 55-60
give proper product distributions since the main effect of radial variations and mass transfer is to increase the effective length of the reactor. As shown in these calculations, the reactor length is not an important variable, so the results are not particularly sensitive to these problems. Summary This model gives results that appear to agree well with known properties of the Ostwald process. Thus, the assumptions on which the model is based either accurately describe the real system or at least do not introduce large errors into the calculations. Apparently, the high temperature and relatively small number of reactions in the gauze reactors justify these simpliciations, which make these reactors the first to yield to quantitative prediction. These calculations may therefore have utility in designing other types of reactors. Similar models of other catalytic reactors with low area catalysts such as catalytic combustors may also be successful in predicting behavior. Acknowledgment This material is based upon work supported under a National Science Foundation Graduate Fellowship. This research was also partially supported by DOE under Grant DE-FG02-88ER13878-AO2. Registry No. NH3, 7664-41-7; NO, 10102-43-9.
Literature Cited Blieszner, J. Ph.D. Thesis, University of Minnesota, Minneapolis, 1979.
Handforth, S. L.; Tilley, J. N. Catalysts for Oxidation of Ammonia to Oxides of Nitrogen. Ind. Eng. Chem. 1934,26 (12), 1287-1292. Hasenberg, D.; Schmidt, L. D. HCN Synthesis from CH, and NH3 on Clean Rh. J. Catal. 1985,91, 116-131. Hasenberg, D.; Schmidt, L. D. HCN Synthesis from CHI and NH3 on Platinum. J. Catal. 1986, 97, 156-168. Hasenberg, D.; Schmidt, L. D. HCN Synthesis from CHI, NH3, and O2 on Clean Pt. J . Catal. 1987,104, 441-453. Heck, R. M.; Bonacci, J. C.; Hatfield, W. R.; Hsiung, T. H. A New Research Pilot Plant Unit for Ammonia Oxidation Processes and Some Gauze Data Comparisons for Nitric Acid Process. Ind. Eng. Chem. Process Des. Dev. 1982,21, 73-79. Hwang, S. Y.; Schmidt, L. D. Decomposition of CH3NH2on Pt(ll1). Surf. Sci. 1987, 188, 219-234. Loffler, D. G.; Schmidt, L. D. Kinetics of NH3 Decomposition on Polycrystalline Pt. J. Catal. 1976,41, 440-454. Mummey, M. J.; Schmidt, L. D. Decomposition of NO on Clean Pt Near Atmospheric Pressures. Surf. Sci. 1981, 109, 29-59. Pignet, T.; Schmidt, L. D. Selectivity of NH3 Oxidation on Pt. Chem. Eng. Sci. 1974,29, 1123. Pignet, T.; Schmidt, L. D. Kinetics of NH3 Oxidation on Pt, Rh, and Pd. J. Catal. 1975,40, 212-225. Satterfield, C. N. Heterogeneous Catalysis in Practice; McGraw-Hill: New York, 1980; pp 214-221. Takoudis, C. G.; Schmidt, L. D. Reaction between Nitric Oxide and Ammonia on Polycrystalline Platinum. 1. Steady-State Kinetics. J. Phys. Chem. 1983,87, 958-963. Trimm, D. L. Catalytic Combustion (Review). Appl. Catal. 1983, 7, 249-282. Twigg, M. V. Catalyst Handbook; Wolfe Publishing Ltd.: London, 1989; pp 470-489. Waletzko, N.; Schmidt, L. D. Modeling Catalytic Gauze Reactors: HCN Synthesis. AIChE J. 1988,34 (7), 1146-1156.
Received for review March 1, 1990 Revised manuscript received June 25, 1990 Accepted July 5, 1990
Characterization and Long-Range Reactivity of Zinc Ferrite in High-Temperature Desulfurization Processed Raul
E.Ayala* and Donald W . M a r s h
C E Corporate Research and Development, P.O.Box 8, Schenectady, New York 12301
The chemical reactivity of zinc ferrite was studied experimentally to demonstrate the potential use of zinc ferrite as a sorbent in high-temperature desulfurization of coal gases. Fifty cycles of H2S absorption from simulated coal gas and regeneration under 4.5% oxygen were conducted in a laboratory-scale, packed-bed reactor system simulating gas compositions of a fixed-bed, air-blown gasifier and a regeneration scheme typical of a moving-bed process. Approximately 70% of the theoretical fractional conversion, as determined by thermogravimetric analysis (TGA), was maintained by the sorbent. Less than 1% residual total sulfur and total carbon were measured in the sorbent. Undesired solid phases, (e.g., metal carbides, sulfates, and elemental iron) were absent in the samples as determined by powder X-ray diffraction. Introduction A major cost in the integrated gasification combined cycle (IGCC) for generation of electricity from is the gas 'leanup to remove contaminants such as solid particulates, species (e*g*tH2S and COS),and tars. Conventional cleanup systems use energy-inefficient, low-temperature scrubbing processes where costly heat exchanger systems are necessary to cool off the gas from the gasifier, the water 'Ondensate, and reheat the gases to match the required turbine inlet tern'Presented at the Symposium on Adsorption and Reaction on Oxide Surfaces of the 198th ACS National Meeting, Miami Beach, FL, Sept 10-15, 1989. 0888-5885/91/2630-0055$02.50/0
peratures (Corman, 1986). A novel approach is the use of high-temperature desulfurization utilizing mixed-metal oxides as sulfur sorbents (Grindley and Steinfeld, 1982, 1983, 1984; ~ ~ ~ z a n ~ ~ ~ ~ e p ~eta n o1987; p o u~~~~l ~os et 1989). Sulfur absorption is accomplished at ternperaturesand pressures matching those of gasifier and turbine components in the IGCC system, thus minimizing energy and cost expenditures in the process, Zinc ferrite, ZnFe20r,is a leading mixed-metal sorbent for high-temperature desulfurization of coal gas having a crystalline spinel structure formed by two metal oxides: ZnO and Fe;O3 (Kolta et al., 1980). Zinc oxide is able to reduce H2S and COS levels below 10 parts per million by volume (ppmv) in coal gas while iron oxide has twice the sulfur absorption capacity of zinc oxide on a molar basis; 0 1991 American Chemical Society
