MgAl2O4 as n-Butane Dehydrogenation Catalyst in a Two-Zone

May 29, 2009 - n-Butane dehydrogenation over a Pt-Sn/MgAl2O4 catalyst has been carried out in a two-zone fluidized-bed reactor (TZFBR), where oxidatio...
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Ind. Eng. Chem. Res. 2009, 48, 6573–6578

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Pt-Sn/MgAl2O4 as n-Butane Dehydrogenation Catalyst in a Two-Zone Fluidized-Bed Reactor M. P. Lobera, C. Te´llez, J. Herguido, and M. Mene´ndez* Aragon Institute for Engineering Research (I3A), UniVersity of Zaragoza, 50009 Zaragoza, Spain

n-Butane dehydrogenation over a Pt-Sn/MgAl2O4 catalyst has been carried out in a two-zone fluidized-bed reactor (TZFBR), where oxidation and reduction of this catalyst occurs in the same vessel but in separate zones. This allows working in a continuous mode with in situ catalyst regeneration, unlike traditional systems that require two different reactors or a single reactor with periodic operation. The effect of the main TZFBR operating variables was studied and its performance compared with that of a conventional fluidized-bed reactor loaded with the same catalyst and with other reactor + catalyst systems. Compared with a previous work in a similar reactor with another catalyst, the selectivity is improved because of the lower tendency to coke formation of the Pt-Sn/MgAl2O4 catalyst. 1. Introduction The increasing demand for olefins creates an opportunity for developing other modes of producing these basic chemicals, as opposed to their traditional sourcing as byproducts of steam cracking and catalytic cracking.1 An alternative to these processes is the dehydrogenation of short-chain alkanes. These highly endothermic reactions are carried out at high temperatures to increase the equilibrium conversion. Two types of industrial catalysts for n-butane dehydrogenation are used, those based on platinum supported on alumina and those based on chromia/ alumina. In both cases a series of competitive reactions occur simultaneously, such as the formation of short-chain hydrocarbons (cracking) and the deposition of coke on the surface of the catalyst (coking),2 and the fast catalyst deactivation produced by coke makes its regeneration necessary. The majority of industrial processes operate cyclically, with fixed-bed reactors and cofeeding of n-butane with hydrogen (to decrease coke formation), which decreases the equilibrium conversion. Another problem to be solved in the development of a suitable reactor for an industrial dehydrogenation process is how to supply the huge quantity of heat needed by the reaction, maintaining in the meantime careful control of the temperature to minimize the yield to cracking products.3 Thus, the main technological problems within the constraint of operating as a continuous process are as follows: (1) sufficient heat supply to the reactor; (2) careful temperature control to minimize cracking products and to maximize conversion; and (3) catalyst regeneration.4 In this context, the use of a two-zone fluidized-bed reactor (TZFBR) for the catalytic dehydrogenation of n-butane has several potential advantages in relation to the three problems mentioned above. In this type of reactor, oxygen and hydrocarbon are fed separately: oxygen to the bottom of the reactor and the hydrocarbon at an intermediate point of the bed. Figure 1 shows a scheme of a TZFBR. Under suitable operating conditions, it is possible to achieve the segregation of two different zones in a single reactor: at the bottom of the reactor the oxygen fed is consumed by reacting with coke deposited in the catalyst, whereas at the top of the reactor the desired chemical reaction takes place. If a good separation of both zones is achieved, the hydrocarbon reacts in the absence of gas-phase oxygen. The continuous internal circulation of solid, a feature * To whom correspondence should be addressed. E-mail: [email protected].

of fluidized beds, provides the solid transport between the two areas, allowing continuous operation of the system. As a consequence of this design, the use of the TZFBR provides in situ regeneration of the catalyst and can operate continuously using a single reactor.5 The advantages of the TZFBR have been investigated in two types of reactions: selective oxidation processes6 and dehydrogenation of hydrocarbons.7-9 When the TZFBR is employed for oxidation processes, the role of the oxygen in the lower section includes the replenishing of the catalyst lattice with oxygen, which is employed in the reaction in the upper section instead of gas-phase oxygen. Recent studies have been conducted on the dehydrogenation of n-butane over Pt/Al2O3 catalyst using sequential pulses of butane and oxygen to combine the reaction and regeneration of

Figure 1. Scheme of the two-zone fluidized-bed reactor.

10.1021/ie900381p CCC: $40.75  2009 American Chemical Society Published on Web 05/29/2009

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the catalyst by the combustion of coke formed in the former.2 The use of neutral supports such as MgAl2O4 (instead of acid supports such as alumina) implies an improvement in the properties of the catalysts for dehydrogenation reactions because they inhibit undesirable side reactions such as cracking and polymerization. In addition, PtSn catalysts supported on MgAl2O4 have shown good performance in the dehydrogenation of n-butane, maintaining good Pt dispersion after different oxidation-reaction cycles,10 and therefore this kind of catalyst was selected for this study. The objective of this work is to study n-butane dehydrogenation over a PtSn/MgAl2O4 catalyst in a TZFBR at laboratory scale. This work provides a more detailed study of the n-butane dehydrogenation over a PtSn/MgAl2O4 catalyst using a TZFBR. It includes a study of the effect of the main operating conditions on the performance of the TZFBR (temperature, relative gas velocity, oxygen feed, and relationship between the height of the reaction zone and the regeneration zone, (H - ha)/ha). In addition, the results with the TZFBR will be compared with results in the literature and the results obtained in a previous work8 with commercial Cr2O3/Al2O3 catalyst in a TZFBR for n-butane dehydrogenation. 2. Experimental Section 2.1. Catalyst Preparation. 2.1.1. Synthesis of Supports. Mg-Al-O mixed oxides were prepared by the coprecipitation method in accordance with the procedure described in the literature.11,12 Stoichiometric quantities of aqueous solutions (0.5 M) of Mg(NO3)2 · 6H2O (Sigma-Aldrich, 99% purity) and Al(NO3)3 · 9H2O (Sigma-Aldrich, 98% purity) were dropped into a flask with constant stirring and the pH value was maintained at about 9.5 with an aqueous solution of ammonium hydroxide (Sigma-Aldrich, 5.0 N). The temperature of the solution was maintained at 50 °C. The resulting gel was continuously stirred for 1 hr and then aged overnight at room temperature. It was then filtered and carefully washed with distilled water several times. The resulting slurry was dried at 120 °C for 15 h and finally calcined at 800 °C for 8 h in air. 2.1.2. Pt and Sn Impregnation. The Pt-Sn/MgAl2O4 catalyst was prepared by incipient wet impregnation10 and its nominal composition was 0.3 wt % in Pt and 0.3 wt % in Sn. The salts used were SnCl2 · 2H2O (Sigma-Aldrich, 98% purity) and H2PtCl6 · 6H2O (Sigma-Aldrich, 8 wt % solution in water). The catalyst was prepared by sequential impregnation; each component was impregnated separately. Sn was first impregnated with SnCl2 · 2H2O dissolved into HCl (Sigma-Aldrich, 1 N); the catalyst was aged at room temperature for 6 h and then dried in air at 120 °C for 12 h. Pt was subsequently impregnated using an aqueous solution of H2PtCl6 · 6H2O, and the same catalyst drying process was again repeated. Finally, the dried catalyst was heated (3 °C/min) to 600 °C in air and calcined for 3 h. Before the study was carried out, the catalyst was pretreated with six reduction-reaction-regeneration cycles to obtain an aged and stable catalyst.7,9,13 In the first step, the catalyst was reduced in situ in flowing H2:He (molar ratio) ) 1:2 for 120 min at 550 °C; in the second step a reaction was carried out in C4H10:Ar (molar ratio) ) 1:10 atm for 60 min at 530 °C, and in the third step the catalyst was regenerated in O2:He (molar ratio) ) 1:20 for 90 min at 540 °C. Finally, the catalyst was ground and sieved to a particle size of 160-250 µm. Before each reaction test the catalyst was regenerated in O2:He (molar ratio) ) 1:20 for 120 min at 540

Figure 2. XRD patterns of the support (MgAl2O4) and aged Pt-Sn/MgAl2O4 catalyst. Table 1. Specific Surface of Synthesized Support and Catalyst BET surface areas (m2 · g-1) MgAl2O4 support Sn/MgAl2O4 dry Pt-Sn/MgAl2O4 dry Pt-Sn/MgAl2O4 calcined Pt-Sn/MgAl2O4 stable Sn/MgAl2O4 after 150 h of operation

114 86 89 76 78 75

°C and then reduced in situ in flowing H2:He (molar ratio) ) 1:2 for 120 min at 550 °C. 2.2. Catalyst Characterization. The surface area was obtained by the Brunauer-Emmett-Teller (BET) method with an ASAP 2020 Micromeritics automatic analyzer. The powder X-ray diffraction (XRD) patterns were measured in a Rygaku/ Max diffractometer equipped with a graphite monochromator operating at 40 kV and 80 mA and using nickel-filtered Cu KR radiation 1.2 (2θ ) 10° a 80°). The JCPDS-International Centre for Diffraction Data-2000 database has been used for the determination of the phases. As can be seen in Figure 2, after calcination at 800 °C, the support sample showed the XRD pattern characteristic of the MgAl2O4 spinel structure, and this structure was maintained after Pt and Sn impregnation and pretreatment of the catalyst, when it was aged and stable. This fact is in accordance with the results obtained for the BET areas, which do not show significant differences between various stages of the catalyst life (Table 1). Morphology and particle shape were examined by transmission electron microscopy (TEM). TEM data were recorded with a JEOL 2000 FXII transmission electron microscope. Energy-dispersive X-ray analysis was performed using an eXL-10-LINK ANALYTICAL that can determine the chemical composition of the microzone that focuses on the electron beam. Three separate samples were prepared for calcined, stable, and used catalyst (after 150 h of operation). Many areas were examined for each TEM specimen to ensure that the reported results were truly representative of the samples. No catalyst sintering was observed during the different steps of the study (results not shown). 2.3. Catalytic Tests. In this work, a TZFBR was used whose most important characteristics have been described previously.5 The reactor was a 30 mm i.d., 300 mm long quartz tube equipped with a quartz distributor plate. It was placed inside an electrical furnace and the temperature was measured by a thermocouple; a PID controller maintained temperature variations within (0.5 °C of the set point. A movable axial quartz

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Figure 3. Evolution of n-butane conversion with time for different O2 flows fed in the reactor. T ) 530 °C, Wr/FC4H10 ) 430 s · kg · mol-1, ur1 ) 1.7, 10% C4H10, ha/H ) 0.5.

probe was used to introduce the n-butane at the desired reactor height. All the streams were mass flow controlled (Brooks). The reaction products were analyzed with an online gas chromatograph (Agilent 3000 µ-CG). Concentration profiles of the different species along the bed length were measured by means of a movable quartz probe connected to an online gas chromatograph. The carbon balances were always better than (3%. The C4* products group includes 1-butene, isobutene, 2-butene, and 1,3-butadiene; and the cracking products group includes ethane, ethylene, and methane. The coke content in the catalyst along the bed was determined. Once steady state was reached under reaction conditions, the flow of gas was abruptly cut off and a small flow of argon was then fed (insufficient to fluidize the catalyst) and the reactor temperature reduced to room temperature. Different samples of the catalyst were then taken at several points with a movable quartz vacuum probe. These samples were analyzed by thermogravimetric analysis in a Mettler-Toledo TGA/SDTA 851e microthermobalance by means of combustion tests in nonisothermal conditions under an air flow (50 cm3 · (STP)/min). The temperature was programmed from 30 to 600 °C with a heating rate of 5 °C/min. 3. Results and Discussion Several sets of experiments were performed in the two-zone fluidized-bed reactor, varying the bed temperature (T), the relative superficial gas velocity (ur ) u/umf), the oxygen fed in the reactor as a percentage of the total gas feed flow at the bottom of the bed, and the relative height of the regeneration and reaction zones (H - ha)/ha. The range of operating conditions was as follows: temperature, 470-550 °C; relative gas velocity (ur ) u/umf), 1.7-3.7; % O2, 0-5%; relationship between the height of the reaction and the regeneration zones ((H - ha)/ha), 0.33-3.00. The size fraction of catalyst between 160 and 250 µm had a minimum fluidization velocity (obtained with He, at 500 °C) of 36 mm/s (equivalent to 1.28 cm3 · (STP) · cm-2 · s-1). Coke formed during n-butane dehydrogenation is burnt in the regeneration zone. Thus, the amount of oxygen fed is a main variable in the operation of a TZFBR. Different experiments were performed with a constant input of n-butane, while the percentage of oxygen fed at the bottom of the reactor was varied to keep a constant relative gas velocity (ur2) in the regeneration zone. In addition, the (H - ha)/ha relationship, Wr/FC4H10 ratio, and reaction temperature were maintained constant. Figure 3 shows the evolution over time of n-butane conversion for experiments performed in a conventional fluidized-bed reactor without feeding of oxygen (FBR), and in the TZFBR

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Figure 4. Evolution of C4* (1-butene, isobutene, 2-butene, and 1,3butadiene) yield with time for different O2 flow fed in the reactor. T ) 530 °C, Wr/FC4H10 ) 430 s · kg · mol-1, ur ) 1.7, 10% C4H10, ha/H ) 0.5.

Figure 5. Yield to main reaction products obtained in TZFBR for different percentages of O2 fed at steady state. T ) 530 °C, Wr/FC4H10 ) 430 s · kg · mol-1, ur ) 1.7, 10% C4H10, ha/H ) 0.5.

with different percentages of oxygen in the reactor feed. It can be seen that it was possible to reach a steady state in the TZFBR when an appropriate amount of oxygen was used. This is due to the stabilization of the coke content of the catalyst along the reactor for a given set of experimental conditions. That is, when equilibrium was reached between coke formation (in the reaction zone, ha < h < H) and coke burning (in the regeneration zone, 0 < h < ha), a constant exit flow rate was achieved. A decrease in the amount of oxygen fed into the reactor increases the time needed to reach steady state and also the butane conversion decreases. These results can be explained because the higher the amount of oxygen fed, the better the coke burning; therefore, the equilibrium between the coke formation and coke combustion is achieved with a higher activity of the catalyst. A second (and undesired) factor that increases the conversion is the transformation of butane to COx. Some COx comes from increased coke combustion since a cleaner catalyst will also be more active in coke formation, but also if the oxygen flow is too large (i.e., more than what is needed to burn the coke formed in the upper zone) or the residence time is not long enough in the lower zone, some of it can reach the upper zone, where it will oxidize butane, butene, and butadiene to COx. Moreover, as can be seen in Figure 4, the C4* yield (1-butene, isobutene, 2-butene, and 1,3-butadiene) decreases sharply with time in a FBR as a result of deactivation by coke deposition, whereas a constant C4* yield is reached in the TZFBR. An initial transient behavior was observed for about 30 min in the TZFBR, after which the yield was stable. In the studied range, an increase in the amount of oxygen fed into the reactor decreases the C4* yield and increases the yield to cracking products and COx (Figure 5). The increase of COx formation has been previously explained. The increase in cracking products can be related to the higher activity of the catalyst. It can be observed that the maximum C4* yield was obtained for a percentage of oxygen

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Figure 6. Effect of temperature upon n-butane conversion and selectivity to C4* (1-butene, isobutene, 2-butene, and 1,3-butadiene) at steady state. Wr/FC4H10 ) 430 s · kg · mol-1, ur ) 1.7, 10% C4H10, ha/H ) 0.5, 1% O2.

Figure 7. Coke concentration along the catalyst bed for different temperatures at steady state. Wr/FC4H10 ) 430 s · kg · mol-1, ur ) 1.7, 10% C4H10, ha/H ) 0.5, 1% O2.

fed of 1%. This percentage is lower than the optimum found in the same reactor9 for propane dehydrogenation using a chromia/ alumina catalyst (around 5%), probably because of the lower tendency of the PtSn/MgAl2O4 to coke formation. From this point on, this percentage is used and all the results reported correspond to the steady state, measured after 5 h of reaction. The effect of temperature upon the n-butane conversion and C4* selectivity is shown in Figure 6. These experiments were carried out with a constant (H -ha)/ha ratio, relative gas velocity, ur, and catalyst mass and flow rates of n-butane, argon, and oxygen. It can be observed that C4* yields show a maximum at 490 °C. Dehydrogenation and cracking are both endothermic reactions; therefore, equilibrium conversion increases with the reaction temperature. However, the coke formation rate also increases with temperature,13 so the catalyst will be increasingly more deactivated; thus, the C4* yield has an optimum temperature. Figure 7 shows the coke content in catalyst particles along the bed for two different temperatures. The coke content in the regeneration zone (0 < h < ha) is lower than that in the reaction zone (ha < h < H). The amount of coke increases along the bed, most of this change appearing in the regeneration zone, whereas the coke content is quite constant in the reaction zone. Probably the higher gas velocity in the reaction zone favors the solid mixing. The coke content in the catalyst increases with the reaction temperature, in agreement with the findings of other authors.8,13,14 As would be expected, the temperature is a very important variable for controlling the coke formation, which is consistent with the trends in the C4* yield (Figure 6). Figure 8 shows another important variable in the operation of the TZFBR, the relationship between the height of the reaction and regeneration zones, (H - ha)/ha. These experiments were carried out with the same catalyst weight in the reactor (WT) but different n-butane feeding points (ha). The variation of the (H - ha)/ha relationship affects the Wr/FC4H10 ratio and

Figure 8. n-Butane conversion and yield to main reaction products obtained in TZFBR for different ratios of length of the reaction zone/length of the regeneration zone at steady state. T ) 490 °C, WT ) 32 g, ur ) 1.7, 10% C 4H10, 1% O2.

Figure 9. Selectivity to C4* (1-butene, isobutene, 2-butene, and 1,3butadiene) vs n-butane conversion for experiments performed with different relative gas velocities at steady state. T ) 490 °C, WT ) 32 g, ha/H ) 0.5, 10% C4H10, 1% O2.

the size of the regeneration zone, so an increase in (H - ha)/ha involves a higher Wr/FC4H10 ratio and a smaller regeneration zone. It can be seen that the n-butane conversion and butene yield initially increase with (H - ha)/ha, but decrease for higher values of (H - ha)/ha. So an optimum of n-butane conversion and butene yield was found when (H - ha)/ha ) 1; i.e., the reaction zone and the regeneration zone were the same size. However, whereas 1,3-butadiene and cracking yields increased slightly with (H - ha)/ha, the COx yield decreased. The changes of conversion with the height of the reaction zones can be explained by considering two different effects: (a) an increase in the amount of catalyst, if the activity is constant, should increase the conversion; (b) the larger the reaction zone, the larger the amount of coke formed per unit of time, and since the same flow of oxygen is fed, the coke concentration in the equilibrium is larger; i.e., the catalyst is more deactivated. The opposing effects of these two factors cause the appearance of an optimum, which in this case corresponds to approximately the same height in both zones. The changes in yield to butene follow the same trends as the conversion, but here there is a different evolution of the yield to butadiene and cracking products. This implies that the catalyst deactivation affects butene formation in a different way than cracking or butadiene formation. A small effect of coke on cracking was also found in propane dehydrogenation over PtSn/ Al2O3.15 Figure 9 shows the selectivity to C4* versus n-butane conversion for experiments performed with different relative gas velocity, ur. It can be observed that both n-butane conversion and selectivity to C4* sharply decrease with the relative gas velocity, ur. This is because the Wr/FC4H10 decreases and also because the amount of gas in the bubble phase increases,

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Figure 10. Variation of the n-butane and oxygen concentration profiles at steady state along the reactor for ur1 ) 1.7, T ) 490 °C, Wr/FC4H10 ) 430 s · kg · mol-1, ha/H ) 0.5, 10% C4H10, 1% O2.

worsening the gas-solid contact. Both facts decrease the n-butane conversion to butenes and butadiene, but the COx yield is always approximately the same because oxygen is totally depleted so that its selectivity increases at the expense of C4 selectivity. Although no large gas backmixing is expected with the ur values employed in this work, it could occur to some extent at the highest ur. Gas backmixing implies that some butane and reaction products flow to the oxidizing zone and could react to COx products, lowering the C4* yield, as has been observed previously with propane dehydrogenation.5 In a fluidized bed, backmixing of the gas occurs through a mechanism that is associated with the mixing of solids. Solid material forming the bubble wake is transported upward with the bubbles; the same solid flow, but in the downward direction, exists in the emulsion phase that surrounds the bubbles. If the velocity of the solid moving downward is higher than the minimum fluidization velocity (which is roughly the gasssolid velocity in the emulsion phase), then the solid in the emulsion phase draws gas downward, and the undesired gas backmixing is produced.5 However, as shown in Figure 10, a clear separation between the reaction and regeneration zones can be obtained if the operating conditions are chosen carefully. It can be seen that the oxygen concentration drops rapidly as a consequence of the combustion of carbonaceous deposit on the catalyst, and practically all the oxygen is consumed below the n-butane inlet.

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Figure 11 shows selectivity to C4* (1-butene, isobutene, 2-butene, and 1,3-butadiene) vs n-butane conversion for experiments performed during this work in the TZFBR. Also, some of the best results extracted from a survey of the literature are shown. These works were performed using different catalysts for n-butane dehydrogenation in several experimental systems. McNamara et al.2 used a Pt/γ-Al2O3 catalyst in a pulse-flow microreactor system. Callejas et al.8 used a commercial Cr2O3/ Al2O3 catalyst in a TZFBR. Results presented by Bocanegra et al.10 correspond to Pt-Sn/MgAl2O4 and Pt-Sn/ZnAl2O4 catalysts in a fixed-bed reactor, n-butane being co-fed with H2. Harlin et al.14 developed their work using a MoO3/SiC catalyst in a fixed-bed reactor, n-butane being co-fed with H2. Garcı´a et al.16 used a VOx/USY catalyst in a fixed-bed microreactor. Results shown by Jackson and Rugmini17 correspond to VxOy/ θ-Al2O3 in a fixed bed. Bocanegra et al.18 used an In-Pt-Sn/ MgAl2O4 catalyst in a fixed-bed reactor, where n-butane was co-fed with H2. In the above articles the catalyst was deactivated by coke (except in the results with TZFBR by Callejas et al.8), its regeneration being necessary in another process phase. In many of the commercial processes available for olefin dehydrogenation, hydrogen is needed to reduce coke formation. In several of these works n-butane was co-fed with hydrogen to reduce coke formation. However, in the TZFBR no hydrogen is required. The lower C4* selectivity in the TZFBR is due to COx, which appears as a product of coke burning in the regeneration zone. Also, the TZFBR has other advantages over conventional reactors. The coke combustion becomes an important source of energy for the endothermic reaction of dehydrogenation. Thus, reaction and catalyst regeneration take place in the same vessel. Furthermore, with carefully chosen operating conditions, a C4* yield of 45% could be reached in a TZFBR and this system allows in situ regeneration of the catalyst so that it can work in steady state. When the results in this work are compared with those obtained in a similar reactor with a Cr2O3/Al2O3 catalyst, a clearly higher selectivity to olefins is obtained with the Pt-Sn/MgAl2O4 catalyst. This higher selectivity for a given butane conversion and the lower flow of oxygen required are related to the slower rate of coke formation in the catalyst selected in this work. This raises a significant issue that should

Figure 11. Selectivity to C4* (1-butene, isobutene, 2-butene, and 1,3-butadiene) vs n-butane conversion for experiments performed during this work in a TZFBR configuration and comparison with results in the literature. Diverse conditions.

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be considered in TZFBR: Although the loss of catalyst activity by coke formation is counteracted by the continuous regeneration in the lower zone, the rate of coke formation may affect significantly the selectivity to the desired products. Otherwise, some yield to coke may be desirable since its combustion provides the heat for the (endothermic) dehydrogenation reaction. In TZFBRs the n-butane and oxygen feeds are segregated. It is possible to select operating conditions to maintain the oxygen concentration in the reaction zone below the minimum value required to avoid a risk of explosion. The percentage values of oxygen in the reactor feed used in this study were much lower than 20%, which according to the flammability diagram for butane-oxygen-inert mixtures,19 is the level of oxygen necessary to reach the explosion region when 10% of n-butane is fed into the reactor. 4. Conclusions A stable operation can be achieved in n-butane dehydrogenation over a Pt-Sn/MgAl2O4 catalyst in a two-zone fluidizedbed reactor. Under appropriate conditions, coke formed during n-butane dehydrogenation is burnt in the regeneration zone so that it can work in continuous mode. The control of the amount of oxygen fed into the reactor is very important. An increase in oxygen fed produces a higher quantity of coke burnt and the catalyst activity is higher, so the conversion of n-butane increases. However, the yield to C4* (1-butene, isobutene, 2-butene, and 1,3-butadiene) decreases as the amount of carbon oxides increases. Temperature, relationship between height of the reaction and the regeneration zones, and relative gas velocity are also important operating conditions for achieving an optimum equilibrium between coke regeneration and C4* yield. Under the conditions tested, the best results were achieved with 1% of oxygen in the feed, 490 °C, and ur ) 1.7 and when the reaction zone and the regeneration zone were of the same size. In addition, the results were compared with data from the literature2,8,10,14,16-18 for different reaction systems, and it was found that the best C4* yield (1-butene, isobutene, 2-butene, and 1,3-butadiene) was achieved in this study and with the great advantage of working in steady state. The use of Pt-Sn/ MgAl2O4 instead of Cr2O3/Al2O3, as was used in a previous work,8 improves the selectivity to olefins and reduces yield to coke and thus the amount of oxygen needed for the regeneration of the catalyst. Nomenclature FC4H10) molar flow of n-butane (mol · s-1) FBR ) fluidized-bed reactor h ) bed height (cm) ha ) n-butane feeding height (cm) H ) total bed length (cm) QT ) total gases flow rate (cm3 · (STP) · min-1) T ) reaction temperature (°C) TZFBR ) two-zone fluidized-bed reactor uo ) superficial gas velocity (cm3 · (STP) · cm-2 · s-1) umf ) minimum fluidization velocity (cm3 · (STP) · cm-2 · s-1) ur ) relative gas velocity (uo/umf) ur1 ) relative gas velocity in reaction zone (uo_reaction_zone/umf) ur2 ) relative gas velocity in regeneration zone (uo_regeneration_zone/ umf) WT ) mass of catalyst in the reactor (g)

Wr ) mass of catalyst in the reaction zone (ha < h < H) (g) Xn-butane ) n-butane conversion Yi ) yield to compound i

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ReceiVed for reView March 9, 2009 ReVised manuscript receiVed May 6, 2009 Accepted May 12, 2009 IE900381P