Ind. Eng. Chem. Process Des. Dev. 1980, 79, 561-565
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Oxidation of Methanol in a Fluidized Bed. 1. Catalyst Attrition Resistance and Process Variable Study L. Calriatl,' L. DI Flore,' P. Forzattl,' I. Pasquon, and F. Trlflrb' Istituto CB Chimica Industriale del folitecnico, 20 133 Milano, Italy
A novel fluid bed process of methanol oxidation to formaldehyde is presented in this paper. It is shown that catalysts prepared by supporting Fe203-Mo03on silica with very low surface area are suitable for fluid bed operations. They combine good attrition characteristics and quite interesting formaldehyde yields (90-98 % selectivity X 98% conversion) at the laboratory scale. The effects of the key process variables are investigated and the following are presented as the best potential operating conditions: catalyst composition, around 1.7% by weight; gas pressure, varying from atmospheric to 6 atm; reaction temperature, from 270 to 340 OC depending on gas pressure; CH30H/02 molar riati0 equal to 1 and CH30H level slightly higher than 6 % . A discussion follows on the significance of fluidized bed data when taking into account that the laboratory reactor operates in the slug flow regime.
Introduction Formaldehyde is prepared for approximately half of its production by the air oxidation of methanol over iron molybdate based catalyst (Adkins and Peterson, 1931). Fe2(Mo04),is the true active component (Boreskov et al., 1966), but a large excess of Moo3 is always present in commercial catalysts (Pernicone, 1974). The MOO, excess is required to provide coherence and high surface area and to prevent the formation of molybdenum defective iron molybdate (Alessandrini et al., 1977; Le Page et al., 1978). The process is operated in fixed bed reactors around 300 "C with maximum of 400-410 "C. At excess temperatures MOO, segregates insidle the pellets. The process of segregation is responsible for the loss of the mechanical strength of the catalyst and for all related problems such as powder losses and increase of pressure drop along the bed. The segregation of Moos also produces molybdenum deficient iron molybdate which is believed to decompose under reaction condition, thus causing catalyst deactivation (Le Page et al., 1978). Usually the catalyst is replaced after 6-12 months of operation. Given the fact that the temperature control is critical for catalyst life and further considering that, due to the high exothermicity of the reaction, hot spots are always experienced in the fixed bed, the operation in a fluidized bed seems attractive. Indeed fluidized bed reactors achieve quite a good temperature control. They also offer further advantages such as ease of solid recirculation, catalyst regeneration, and replacement. On the other hand, their use imposes severe requirements on catalyst attrition resistance and product specficity. From the economical point of view, fluid bed techinology might be convenient at high capacities because a single reactor could be used. However, up to now the fluid bed oxidation of methanol to formaldehyde has nlot received much attention. To the authors' knowledge, n o commercial methanol oxidation process is accomplished in fluidized beds; neither are reports dealing with this; type of operation available in the scientific literature. A program designed to prepare a catalyst suitable for the fluidized bed oxidation of methanol to formaldehyde has been developed in our laboratories. After confirming that Fe,(MoO,),-MoO,, catalysts cannot be used due to the 'Euteco Spa, Via Gmzioli 11, 20100 Milano, Italy. Istituto di Tecnologie Chimiche Speciali, Universitk di Bologna, 40136 Bologna, It(a1y. 0196-4305/80/1119-0561$01.00/0
low mechanical strength of this system, the decision was made to increase the catalyst strength by supporting the active components onto suitable commercial materials. Microspheroidal silica was chosen as Fez03-MOO, carrier because of its widespread use in fluidized bed applications. The precaution to reduce the surface area of this commercial material by high temperature calcination was adopted since it has been found that silica with high surface area decomposes methanol mainly to CO, methyl formate, and COP. Supported FePO-MoOs catalysts suitable for the fluidized bed oxidation of methanol to formaldehyde have been prepared (Cairati and Di Fiore, 1977; Cairati and Trifirb, 1977; Cairati et al., 1979) along these lines. Previous works by Carbucicchio (1979) and Carbucicchio and Trifirb (1980) clarify the nature of the active components, the interactions of these components with the support, as well as the redox properties of these catalysts. This paper describes the results of an investigation carried out to determine the attrition characteristics of the catalysts and the effects of the key process variables on conversion and selectivity for the defining of the potential operating conditions of the fluidized bed methanol oxidation process. Experimental Section Materials. Fresh microspheroidal silica (Grace 951) was used as received. Calcined silica was obtained from this by heat treatment at 1200 "C for 2 h; the resulting surface area, determined by the BET method, was less than 1 m2 g-1. Silica-supported FeMo samples were prepared by impregnating calcined silica with aqueous solutions of Fe(N03)3'9H20and (NH4),Mo,O2,.4H20 in the presence of citric acid, as reported previously by Cairati et al. (1979). Samples with 0.7, 1, 1.25, and 1.7% FeMo by weight (as Fe2O3-MoO3) were prepared. The Fe/Mo molar ratio in all samples was 2. The procedure followed for drying and calcining has been reported elsewhere (Cairati et al., 1979). Reference to the catalysts is made by indicating the symbol of the support @io2) followed by the FeMo 70by weight. The characteristics of the solids are given in Table I. Attrition Apparatus and Test Procedure. The attrition resistance characteristics of the samples were investigated by means of the test described by Forsythe and Hertwig (1949). In this laboratory accelerated attrition test a catalyst sample is subjected to a jet attrition action for 1 h. The resulting change in the distribution of catalyst particle size, determined by dry screen analysis, is taken as a measure of attrition characteristics and compared with
e 1980 American Chemical Society
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Ind. Eng. Chem. Process Des. Dev., Vol. 19, No. 4, 1980
Table I. Solids Characteristics
materials microspheroidal silica (Grace 951) calcined silica Si0,-0.7, Si0,-1, Si0,-1.25, Si0,-1.7
apparent particle density, g/cm3 0.47 1.15 1.15
Size Distribution for Si0,-0.7, Si0,-1, Si0,-1.25, Si0,-1.7 Charged t o the Fluidized Beda
a
size. um
cumulative % undersize Ibv wt)
44 53 62 74 125 +125
28.96 49.44 49.61 93.70 99.75 100.00
Geometric means size, 55.6 pm.
that of other catalysts under the same conditions. Although the comparisons of catalysts are relative and do not permit direct quantitative prediction of actual plant losses, the test reproduces in a reasonable way the mechanism of particle breakdown in commerical units so that the results can be representative of commercial operation (Forsythe and Hertwig, 1949). The attrition apparatus and the test procedure used were perfectly similar to those reported by Forsythe and Hertwig (1949); only the canvas filter a t the top of the glass pipe was replaced by a fines collection vessel provided with filter bags. Fluidized Bed Apparatus and Testing Procedure. The fluidized bed reactor consisted of a 350 x 3.9 cm i.d. stainless steel tube heated by a three-zone electrical resistance furnace and provided with a sintered metal porous plate distributor and with a vessel, 1 2 cm in diameter and 25 cm in height, at the top. A sintered metal porous plate was placed at the top of the vessel to prevent any loss of catalysts entrained in the exit gas stream. The reaction temperature was measured by an iron-constantan thermocouple placed at the center of the bed height. The fluidizing gas consisted of CH30H-air or CH30HNz-O2 mixtures. The air and the Nz-O2 streams were obtained from compressed gas cylinders through a pressure regulator and metered on a rotameter. Liquid CH30H was supplied through a metering micropump Type CE 150. The two streams were mixed and heated for methanol evaporation up to the reactor. Gas samples from the reactor inlet and outlet gas streams were analyzed in two gas chromatographs: a Perkin-Elmer Model 154 with a molecular sieve 5A column (1 m long) at 50 OC for Oz, Nz, and CO analysis and a Hewlett-Packard Model 5700 A with a Porapack N column (1.8 m long) a t 140 OC for COz, CH20, HzO, CH30H, and other oxidized products analysis. Results and Discussion (1) Attrition Characteristics of the Catalysts. The Forsythe test adopted in this study for the evaluation of the catalyst attrition resistance involves comparison with a commercially acceptable catalyst. Given the fact that no commercial catalysts for the fluidized bed oxidation of methanol were available, fresh microspheroidal silica was assumed as a commercially acceptable catalyst due to its large use in industrial fluidized bed applications. The results obtained in the laboratory test for sioz-1.7 and for fresh microspheroidal silica are given in Table 11. Calcined silica behaves similarly to sio2-1.7 and thus the table points out the effect of high-temperature calcination
Table 11. Effect of High-Temperature Calcination on Attrition
screen analysis
silica Grace 951 and calcined silica before attrition
+ 1 2 0 mesh,
0.25
0.3
0.34
6.30
0.6
5.8
43.84
17.0
6.2
0.17
1.8
20.48
42.2
19.6
28.96
38.2
66.42
calcined silica Grace silica 951 after after attrition attrition
%
-120 + 2 0 0 mesh, % -200 + 230 mesh, % -230 + 270 mesh, % -270 t 3 2 5 mesh, % - 3 2 5 mesh,
1.64
%
increase in % of -325 mesh
9.24
37.56
Table 111. Effect of High Temperature Calcination on Size Distribution cumulative % undersize (by wt) size, pm
silica Grace 951, fresh
silica Grace 951, calcined
44 53 62 74 125 + 125
26.05 81.05 83.17 92.86 99.62 100.00
53.00 88.61 90.20 98.51 99.04 100.00
on the attrition resistance behavior. The test resulted in a 9.24% increase in -325 mesh fines in the case of fresh microspheroidal silica and in a 37.56% increase in the case of sio2-1.7. The difference in the increase in -325 mesh fines indicates that the high-temperature calcination we were forced to adopt in order to minimize the negative role of silica in the oxidation of methanol gives rise to a less abrasion resistant carrier. We note that the attrition in the test is presumably caused by the collisions of particles rapidly accelerated by the air jet with slower moving particles, which simulates the actual situation in the grid zone of a fluidized bed, as well as by the collisions of the particles in the zone of proper fluidization. The hightemperature calcination produces a remarkable increase in particle density (see Table I) which presumably results in higher production of fines due to the action of the air jet. The appearance of larger cavities in the catalyst as well as in calcined silica, relative to fresh microspheroidal silica, was also observed by means of scanning electron microscopy. This observation is in line with the change in silica size distribution upon calcination reported in Table 111, viz. an increase in the content of the smallest size fraction. We deduce that both the larger cavities and the higher particle density are responsible for the lower attrition resistance of sioz-1.7. The behavior in the Forsythe test of sio2-1.7 and of fresh microspheroidal silica were further investigated as functions of the time jet action. In Figure 1the -325 mesh contents for the two samples and the percent by weight of FeZO3-MoO3in the +325 mesh fraction for sio2-1.7 are given against 20,40,60,80-min intervals of jet action. The -325 mesh fines were screened from the samples at every test interval. The figure shows that most of the catalyst breakdown occurred within the first 20-min interval while later on the attrition rates (increase in the -325 mesh fines/+325 mesh fraction charged) were the same for the two systems. No preferential loss of the active phase occurred in the case of sioz-1.7as shown by the constancy
Ind. Eng. Chern. Process Des. Dev., Vol. 19, No. 4, 1980 563 Fe Mo% b w
-325 mesh content ,%
t
Aeaction temperature C;
CH,O selectivity,%
f
i'
loot
\
2601
0
T i m e , mtn
Figure 1. Effect of time j , t action on attrition resistance (0, fresh microspheroidal silica (Grace 951); ., sio2-1.7) and on FeMo % by weight (A). The size distribution for the two samples at the start of the test was: +120 mesh %, 5.57, -120 +200 mesh %, 11.46, -200 +230 mesh %, 41.90, -230 .t270 mesh %, 7.67, -270 +325 mesh 70, 33.40.
of FeMo percentage by weight with time jet action. The behavior is consistent with a uniform distribution of the active phase within the catalyst particle. Indeed the high-temperature calcination of silica is known to produce macropores (Bossi e t al., 1973) and thus ensures effectiveness in impregnatiing internal catalyst surface. On the basis of the above results we conclude that the attrition characteristics of sio2-1.7 can be compared to those of a commercially acceptable catalyst (i.e., microspheroidal silica) provided sio2-1.7has first been subjected to the Forsythe test for 20 min. For this reason the catalyst samples charged to the fluidized bed reactor for activity measurements were screened from the samples obtained after a 20-min Forsythe test. (2) Process Variable Study. A study was conducted to determine the effects of the key process variables such as the FeMo percentage by weight in the catalyst, the reaction temperature, the operating pressure and the oxygen partial pressure on CH,OH conversion, and CH20 selectivity with the aim at defining the potential operating conditions of the process. The CH30H concentration in the inlet was 6% by volume, a level allowing the highest CH20 productivity and still consistent with CH30H-air explosion limits. The superficial gas velocity U and the bed height a t incipient fluidization Hmf,unless otherwise specified, were U = 6.26 cm/s and Hmf= 104 cm. The investigation was carried out a t almost complete CH,OH conversions in order to study the behavior of the catalyst under the most interesting and severe experimental conditions. In Figure 2 the temperature of 98% CH30H conversion and the CH20 selectiviky are shown as functions of the FeMo percentage by weight in the catalyst. The catalysts present quite nice selectivities to formaldehyde ranging from 90 to 98%. These values are comparable to those reported for commercial fixed bed catalysts a t the same methanol conversion levels (viz.,-around 95%). The increase in the FeMo percentage by weight results in a markedly higher activity, due to the larger quantity of the active phase a t the surface, and in a slightly higher selectivity. The trend of CH20 selectivity can be related to the negative role of the carrier since silica with very low
1
05
1
10 FeMo% b w
1
15
Io
Figure 2. Effect of FeMo % by weight on the temperature of 98% and on the CH20 selectivity (0).ExperiCH,OH conversion (0) mental conditions: P = 1 atm; CHBOH= 6% ; 0 2 = 20% ; Ho = 104 cm; U = 6.26 cm/s.
surface area can still decompose CH30H, although to a small extent, mainly to CO, C02, and methyl formate. Inspection of the figure also suggests that small increases in activity are to be expected for FeMo percentage by weight higher than 1.7 probably due to the almost complete coverage of the catalyst surface by the active components a t this level. On the basis of these results further investigation was confined to sio2-1.7. The effect of successive increases in the operating pressure were studied. High-pressure fluidized beds are potentially very attractive because of the following factors. High-pressure units lead to smaller beds since the productivity is increased for a given superficial gas velocity and a given bed cross section. They also lead to higher heat density, kcal/h m3, and so allow for the best utilization of the heat transfer capacity of the cooling tubes. In addition, increase in pressure aids fluidization, since fluidization is a function of gas density. High pressures also result in considerably smaller bubbles and hence in more favorable gas-solid contacting and in a reduction of the ejection and elutriation of particles at the bed surface. Investigation of the effect of total gas pressure may also allow for greater flexibility in defining the operating conditions to meet the time depending demands of the production. In Figure 3 the temperature of 98% CH30H conversion and the CH20 selectivity are given as functions of gas pressure. High pressures result in higher reaction temperatures for the same CH30H conversion to be secured and in losses of few points in CHzO selectivity. Considering that Fe2(Mo04):,has been identified by Carbucicchio and Trifirb (1980) as the active component in silica-supported Fe203-Mo03 catalysts, the effects can be explained with the different kinetic dependences for the CHBOHand CHzO oxidations over Fe2(Mo04):,and with the greater relevance of the former oxidation. Indeed, the trend in activity is consistent with a CH30H pseudo-order dependence lower than 1for the alcohol consumption rate (Jiru et al., 1964; Dente et al., 1964) while the decrease in selectivity with pressure is consistent with a higher CH20 pseudo-order dependence (equal to 1)in the case of the oxidation of the aldehyde (Dente and Collina, 1965). In view of these data and despite the advantages of highpressure operation the choice of the proper pressure level appears to be still uncertain. Finally, the effect of oxygen partial pressure was investigated. No changes in activity and selectivity from
564
Ind. Eng. Chem. Process Des. Dev., Vol. 19, No. 4, 1980 tH,O s e l e c t i v i t y , OO/
temperature, 'C
100 -
80
-
60 -
10-
20-
I
2
1
1
1 Gas pressure ,ata
Figure 3. Effect of gas pressure on the temperature of 98% CH,OH conversion (0) and on the CH20selectivity (0). Experimental conCH30H = 6%;O2 = 20%; Ho= 104 cm; ditions: catalyst, sio2-1.7; U = 6.26 cm/s.
20% down to 5% oxygen levels were observed when operating at high conversions (195%). A decrease in activity and again no significant change in selectivity were observed at lower conversion levels by decreasing the oxygen content in the gas phase. The effects are perfectly in line with the kinetics of CH30H and CH20 oxidation since the former oxidation depends on oxygen partial pressure while the successive oxidation is practically of negligible importance (Jiru et al., 1964; Dente et al., 1964). Consequently, oxygen partial pressure affects CH30H conversion and not CH20 selectivity. It is worth noting that operation at low oxygen levels around 5% and a t C H 3 0 H / 0 2molar ratio around 1 did not produce catalyst deactivation even for almost complete conversions. Thus CH30H/02molar ratio equal to 1 together with CH30H levels higher than 6% may be chosen for commercial operation, provided that we worry about the CH30H-02-N2 explosion limits. In view of the high heat transfer capacity and the gas-solid contacting in fluidized beds, even these limits can be partially overcome, as in some commercial processes. The data so far obtained can be summarized by indicating the following as the best potential operating conditions for the fluidized bed oxidation of methanol to formaldehyde: catalyst composition, around 1.7% by weight; gas pressure, varying from atmospheric to 6 atm; reaction temperature, from 270 to 340 "C depending on the gas pressure; C H 3 0 H / 0 2molar ratio equal to 1,and CH30H levels slightly higher than 6%. (3) The Significance of Laboratory Fluidized Bed Data. An important characteristic required for a correct interpretation of fluidized bed behavior and for successful scaling-up is the regime of gas flow governing gas solid contacting. Our laboratory reactor, as in many laboratory units, operated in the slug flow regime. Stewart's criterion for the onset of slugging (Stewart, 1965) is fulfilled since
( U - U m f ) / ( 0 . 3 5 m )= 0.277 > 0.2 However, the results obtained by decreasing the height of the bed Ho showed a striking effect of the grid zone on conversion. Figure 4 gives R, the fraction of CH30H converted, as a function of Ho at 335 "C; it comes out that a great deal of the reaction occurs near the vicinity of the distributor. By replotting the data as -log (1- R ) against Ho and extrapolating to zero bed height on a straight line
oL 0
I
20
I
1
60
LO
I
80
n o , cm
Figure 4. Effect of the bed height Hoon CH30H conversion R (0). Experimental conditions: catalyst, sio2-1.7;P = 6 atm; T = 335 "C; CH30H = 6 % ;O2 = 20%; U = 6.15 cm/s.
we obtain a 20% grid conversion. Strictly speaking, a straight line relationship holds for first order kinetics and fast reactions (Botton, 1970) under the assumption of the May's reactor model (May, 1959). In our case, the situation appears to be more complex, but nevertheless the trend is consistent and the extrapolation provides an approximate value for the grid conversion. The data in Figure 4 confirm that the flow characteristics near the vicinity of the grid are of importance for CH30H conversion so that the grid design may be critical in scaling-up. Finally, it is useful to stress that due to the different flow regimes (slug flow and free bubble respectively) laboratory fluidized beds provide much higher conversions for a given bed height and a given superficial gas velocity than large commercial beds (Matsen and Tarmy, 1970). In a largescale reactor, bubbles will grow to a large diameter and these large bubbles will provide a path for gas bypassing. Therefore, information on the flow characteristics in the commercial unit or in the proper pilot plant unit along with information on the kinetics is required for correct scaling-up. Conclusions The present investigation has demonstrated that catalysts prepared by supporting Fe203-Mo03onto low surface area silica (S < 1m2 g-') are suitable for the fluidized bed oxidation of CH30H to CH20. These catalysts provide 90-98% CH20 selectivity at almost complete methanol conversion at the laboratory scale. These values are comparable to those secured by commercial fixed bed catalysts at the same conversion level. The results support the commercial interest of Si02-1.7% (Cairati and Di Fiore, 1977; Cairati and Trifirb, 1977) while economical considerations indicate that operation in a fluidized bed may be convenient for capacities over 100000 lb/year of 100% CH20. The study has also offered useful information on the choice of the most favorable potential operating conditions for the process. The fluid bed technology for the oxidation of methanol to formaldehyde is being advanced in another work which involves a kinetic study and a pilot plant investigation. Nomenclature D = bed diameter, cm
Ind. Eng.
Chem. Process Des. Dev. 1980,
g = acceleration of gravity, cm sW2 Ho = bed height, cm H,f = bed height at incipient fluidization, cm U = superficial gas velocity, cm s-l U d = superficial gas vellocity at incipient fluidization, cm s-l
Literature Cited Adkins, H., Peterson, W. P., J . Am. Chem. Soc., 53, 1512 (1931). Alessandrini, G., Cairati, L., Foczatti, P., Viiia, P. L., Trifir6, F., J. Less-Common Met., 54, 373 (1977). Boreskov, G. K., Kolovertnov, 13. D., Kefeli, L. M., Plyasova, L. M., Karachiev, L. G., Matikhin, V. N., Popov, M. I., Dzis'kd, V. A,, Tarasova. D. V., Kinet. Ketal., 7, 144 (1966). Bossi, A., Leofanti, G., Moretti, E., Giordano, N., J . Mater. Sci., 8, 1101 (1973). Botton, R. J., Chem. Eng. Prcy. Symp. Ser. No. 101, 66, 8 (1970). Cairati, L., Di Fiore, L. (to Euteco Spa), Italian Patent 21426 (Mar 21, 1977). Cairati, L., TrifirB, F. (to Euteoo Spa), Italian Patent 27409 (Sept 9, 1977). Cairati, L., Carbucicchio, M., Fluggeri, O., Trifir6, F., Stud. Surf. Sci. Cafal., 3, 279 (1979). Carbucicchio, M., J . Chem. Phys., 70(2), 784 (1979). Carbucicchio, M., Trifir6, F., J . Catal., 62, 13 (1980).
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Dente, M., Poppi, R., Pasquon, I., Chim. Ind. (Milan), 46, 1326 (1964). Dente, M., Coliina, A., Chim. Ind. (Milan), 47, 821 (1965). Forsythe, W. L., Hertwig, W. R., Ind. Eng. Chem., 41, 1200 (1949). Jiru, P., Wichteriova, B., Tichy, J., Proc. 3rdInt. Congr. Catal. Amsterdam, 1, 199 (1964). Le Page, J. F., Cosyns, J., Courty, P., Freund, E., Frank, J. P., Jacquin, Y., Juguin, E., Marclily, C., Martino, G., Miquel, J., Montarnai, R., Sugier, A,, Van Landeghem, H., "Catalyse de Contact", Technip, Paris, 1978. Matsen, J. M.. Tarmy, B. L., Chem. Eng. Prog. Symp. Ser. No. 101, 66. 1 (1970). May, W. J., Chem. Eng. Prog., 55, 49 (1959). Pernicone, N., J . Less-Common Met., 36, 289 (1974). Stewart, P. S. B., Ph.D. Thesis, University of Cambridge, Cambridge, England, 1965.
Received f o r review April 18, 1979 Accepted June 18, 1980 This work has been supported by Euteco Spa and by Consiglio Nazionale delle Ricerche (Roma).
Reaction of Sulfur Dioxide and Hydrogen Sulfide with Porous Calcined Limestone Girard A. Slmons' and Wilson T. Rawlins Physical Sciences Incorporated, Woburn, Massachusetts 0 180 7
A simple theory is developed to describe the mass transport and heterogeneous chemistry which occurs when either SO2 or HzS reacts with calcined limestone. The reactant gas diffuses into the porous calcine and is consumed on the interior surface. The basic heterogeneous rate constants for HzS and SOz with CaO have been inferred from existing laboratory data. Our results indicate that the reaction of HS , with CaO proceeds almost as fast as that of :SOzwith CaO. Hence, any procedure which utilizes limestone removal of SOz is potentially capable of removing HzS at approximately the same rate.
I. Introduction The high-temperature removal of HzS by calcium-based sorbents such as limestone and dolomite may provide a useful cleanup technique for fluidized bed coal gasification processes (Keairns et al., 1976). A similar method for SOz removal in fluidized bed coal combustion processes has already been shown to be a promising technique (Case et al., 1978). In both processes, high-temperature capability of the sorbent permits in situ cleanup of the producer gas and the resulting long gas-solid contact time in the fluidized bed allows high sorbent utilization. Limestone is primarily calcium carbonate (CaCOJ, or calcite. Upon heating, the calcite decomposes into calcium oxide (CaO) and C02. The CaO is referred to as calcined limestone or simply calcine. The HzS reacts with calcine to form calcium sulfide (Cas) whereas SOz reacts with calcine to form calcium sulfate (CaS04). The ultimate level to which sulfur can be removed depends upon the equilibrium properties of the gas mixture; however, the efficiency of sorbent utilization is determined by the overall rate of the gas-solid readion. This work is concerned with evaluating the kinetic rates with which SO2 and H2Sreact with calcine. The practical utility (of the limestone technique cannot be evaluated on the baisis of the chemical kinetics alone. In the case of SOz removal, there is clear evidence (Borgwardt and Harvey, 1972; Hartman and Coughlin, 1974) that the diffusion of SO2 through porous CaO may be rate controlling. In addition, it has been proposed oi96-43051a0111 i9-0565$oi.o010
(Hartman and Coughlin, 1976,1978) that the diffusion of the reactant gas through the solid sulfur deposits within the porous structure is, under some circumstances, rate limiting. The whole process is complicated by the fact that the porous structure itself is dependent upon the degree of calcination and sulfation. These complicated processes must be described in order to predict the sulfur removal rate in a diffusion-controlled environment. Transport theories have been developed to describe diffusion-limited processes in SO2 cleanup. Hartman and Coughlin (1976, 1978) represent the porous CaO by a pattern of small spherical grains surrounded by pores. The SOz diffuses through the pores to the grain where it reacts with the CaO. As the CaO reacts, CaS04 is formed on the outside of the grain and further reaction requires that the SOz must diffuse through the solid phase CaS04. The net reaction is then limited by both the diffusion of SOz through solid CaS04 and by the diffusion of SOz through the pores. The reaction ultimately terminates because the porosity of the stone is reduced to zero and prevents SOz from reaching the grain. The grain theory possesses four parameters: the grain size, the gas phase diffusion coefficient, the activated diffusion coefficient, and the heterogeneous rate constant. The grain radius was obtained by X-ray diffraction analysis and the gas phase diffusion coefficient (D) was taken from the experimental results of Campbell et al. (1970). The heterogeneous rate constant ( K ) was evaluated from experimental data in the limit of low fractional conversion 0 1980 American
Chemical Society
