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SiO2 catalyst
Fuel Processing Technology 90 (2009) 1319–1325
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Fuel Processing Technology j o u r n a l h o m e p a g e : w w w. e l s ev i e r. c o m / l o c a t e / f u p r o c
Comparative study of the two-zone fluidized-bed reactor and the fluidized-bed reactor for oxidative coupling of methane over Mn/Na2WO4/SiO2 catalyst A. Talebizadeh, Y. Mortazavi ⁎, A.A. Khodadadi Oil and Gas Center of Excellence, School of Chemical Engineering, University College of Engineering, University of Tehran, P.O. BOX 11155-4563, Tehran, Iran
a r t i c l e
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Article history: Received 25 January 2009 Received in revised form 3 June 2009 Accepted 23 June 2009 Keywords: Oxidative coupling Methane Two-zone Fluidized Mn/Na2WO4/SiO2 Redox
a b s t r a c t In this work, oxidative coupling of methane over Mn/Na2WO4/SiO2 catalyst is studied in a two-zone fluidized-bed reactor (TZFBR) and its performance is compared with a fluidized-bed reactor (FBR). Diluted oxygen in argon was introduced into the bottom of the TZFBR through a quartz ferrite and methane was entered at higher positions along the fluidized bed. The catalyst circulated between the oxygen-rich and methane-rich zones in the TZFBR reactor. The effects of the main operating variables including bed temperature, the methane/oxygen ratio (Rmo), and the height at which methane was introduced into the reactor (Hm) were investigated. It is found that under some operating conditions the TZFBR gives a higher C2 selectivity than that obtained in the FBR reactor. Reaction of methane with lattice oxygen of the Mn/ Na2WO4/SiO2 redox catalyst in the methane-rich zone may have led to the higher selectivity. © 2009 Elsevier B.V. All rights reserved.
1. Introduction A great deal of attention has focused on the production of C2+ hydrocarbons, mainly C2H6 and C2H4, from oxidative coupling of methane (OCM) since the early work of Keller and Bhasin [1]. It is noticeable that in this process CO2 and CO are undesirable byproducts which cause decrease of the C2 selectivity. A lot of investigation has been done for increasing C2 selectivity and yield by means of different catalyst formulations and/or by reaction engineering in order to make the process commercially feasible. One of the most effective catalysts for the OCM is believed to be Mn/Na2WO4/SiO2, which was first introduced by Fang et al. [2]. It was reported that with this catalyst in the fixed-bed reactor, it is possible to reach a C2+ selectivity up to 80% at a CH4 conversion of about 20% for periods longer than 90 h [3,4]. The OCM reactions are highly exothermic and predominantly carried out in the presence of a heterogeneous catalyst. In the fixedbed reactors, especially in the industrial scale ones, due to the large amount of heat released in the course of reaction, the operation of the reactor is accident prone. Pak and Lunsford [4] described the thermal effects and magnitude of the hot spots for Mn/Na2WO4/SiO2 and Mn/ Na2WO4/MgO catalysts. They found that the exothermic reactions that occur during the OCM over these catalysts resulted in a small region in the catalyst bed which had a temperature that was 150 °C higher than the reactor wall. Also, Follmer et al. [5] observed steep axial and radial temperature gradients in the catalytic fixed-bed
⁎ Corresponding author. Tel.: +98 21 61112193; fax: +98 21 66967793. E-mail address: [email protected] (Y. Mortazavi). 0378-3820/$ – see front matter © 2009 Elsevier B.V. All rights reserved. doi:10.1016/j.fuproc.2009.06.021
reactors. It is obvious that, hot spots have detrimental consequences on the operation of the reactor, such as temperature runaway, catalyst deactivation, undesired side reactions and thermal decomposition of products. Therefore, heat management and temperature control are the key factors for the selection of the reactor type for OCM process. Temperature homogeneity may be achieved by applying fluidized-bed reactors in which intensive mixing of the particulate solids results in a relatively fast dissipation of the generated heat within the catalytic reactor. Some researchers have even claimed that OCM may be carried out isothermally in the fluidized-bed reactors [6–8]. Apart from selective catalysts, the reactor type and the mode of contact between the reactants are considered to play key roles in selective oxidation processes. It is often found that the reaction takes place using oxygen from the catalyst lattice, and the contribution of gas-phase oxygen is generally considered a determinant to selectivity. In the OCM reaction, the catalyst is more selective than the gas phase towards formation of C2 products [9]. This has led to the reactor designs that attempt to reduce the concentration of gas-phase oxygen in the reaction. Over the past two decades, several groups have studied OCM on catalytically active membrane which changes the OCM reaction mechanism or minimizes the presence of the gas-phase oxygen [10,11]. In these studies, the C2 selectivity was much higher than that obtained in the cofeed reactors. An alternative approach is to avoid the simultaneous presence of hydrocarbons and oxygen over the catalyst. In this approach reducible (redox) catalysts may be used either in a circulating bed reactor, in which the catalyst is transferred between the reactor and the regenerator, or in a two-zone fluidized-bed reactor in which it uses segregated feeds for the oxidant and the hydrocarbon in a single
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Nomenclature dp H Hb Hm Umf mcat. Dreactor QSTP ρc
Particle diameter, µm Reactor height, cm Height of the catalyst bed under fluidization condition, cm The height at which methane is introduced in the TZFB, cm Minimum fluidization velocity, cm/s Amount of the catalyst in the reactor, g Reactor diameter, cm Total feed flow rate at standard condition, ml/min Catalyst density, g/cm3
reactor vessel. TZFBR allows hydrocarbon oxidation make use of lattice oxygen of the catalyst in the reduction zone of the bed (above the hydrocarbon inlet) in absence of the gas-phase oxygen. The oxygen depleted (reduced) catalyst is then regenerated (reoxidized) by internal recirculation to the oxidation zone at the bottom part of the reactor where oxygen is fed. Two-zone fluidized-bed reactors were previously used in the laboratory studies for oxidative dehydrogenation of butane and propane and it was reported that the selectivity of desired products increased [12–14]. Step-change experiments on the Mn/Na2WO4/SiO2 catalyst revealed that the catalyst is capable of performing OCM reaction, with desirable C2 selectivity, in the absence of gas-phase oxygen [15]. In this work, oxidative coupling of methane over Mn/Na2WO4/ SiO2 catalyst is investigated in a two-zone fluidized-bed reactor and the results are compared with those obtained in a conventional fluidized-bed reactor.
quartz probe, located approximately near the reactor axis, was used to take samples at different bed heights which were sent to GC for analyses. The both probes were used for TZFBR experiments. The reactor was placed in an electrically heated furnace. Temperature of the reactor was controlled using a thermocouple mounted in the outer wall of the reactor, centered along the catalyst bed. The temperature of the furnace remains almost constant. Another thermocouple was placed in a quartz tube sheathe and located inside the bed to monitor the catalyst bed temperature. Flow rates of all feed gases, CH4 (99.999%), O2 (99.999%) and Ar (99.995%), were regulated by mass flow controllers (Unit Instrument UFC-1200GI). The effluent of the reactor was first passed through a condenser to remove almost all of the water vapor and then analyzed with an on-line Carle 400A GC equipped with a Propack-Q packed column, a methanizer and an FID detector. CO and CO2 were converted to methane in the methanizer to provide an accurate detection of them by using FID detector. The feed was also analyzed by a bypass stream. In the TZFBR, methane was fed at the mid-point along the fluidized bed with Hm height from the bottom of the reactor, and O2–Ar mixture was fed at the bottom of the reactor, Fig. 1. In the FBR, the feed gas mixture, i.e. methane, oxygen and argon were premixed and fed through the quartz distributor at the bottom. The main operating variables studied are: bed temperature, methane/oxygen ratio (Rmo), and the height at which methane is introduced (Hm), gas flow rate (as gas relative velocity), as well as inert gas ratio (as mole %) in the feed. The experiments performed on 7.5 g of the catalyst and a total flow rate of 300 sccm/min, except for the experiments which are done for the effects of flow rate. This amount of the catalyst corresponds to a height equivalent to 10 cm.
2. Experiments 2.1. Catalyst preparation The 4.0 wt%Mn/5.0 wt%Na2WO4/SiO2 catalyst was prepared by dry impregnation method, using Davisil grade 645 silica gel (Aldrich). An aqueous solution of Mn(NO3)2·4H2O (Merck) with appropriate concentration was added dropwise to silica gel and dried at 130 °C overnight. The catalyst was then impregnated with an aqueous solution of Na2WO4·2H2O (Merck) with appropriate concentration, dried again at 130 °C overnight and calcined at 800 °C for 8 h. The catalyst was sieved to make a desired particle size with diameters of 150–250 µm. The catalyst density was measured to be 1.1 g/cm3. X-ray diffraction (XRD) patterns of the catalyst at the following conditions were obtained by a Philips diffractometer with Cu-Kα target in the range of 4–70o (2θ): a—the fresh catalyst, b—the catalyst under O2 and inert gas at 800 °C was fluidized and then cooled at the same condition, i.e. fully oxidized catalyst, and c—the catalyst under methane and inert gas at 800 °C was fluidized and then cooled at the same condition, i.e. fully reduced catalyst. The BET surface area, pore volume and pore size distribution of the catalyst have been reported elsewhere [16]. 2.2. Setup A schematic of the setup for two-zone fluidized-bed reactor (TZFBR) and fluidized-bed reactor (FBR) is shown in Fig. 1. The reactor was 20 mm I.D. and 25 cm length quartz tube equipped with a sintered quartz distributor plate with nominal pore size between 10 and 16 μm and an axial movable quartz probe to introduce the methane feed at the desired bed height (Hm). This probe terminates with a bulb with 8 fine holes uniformly located on its surface, to ensure a good radial distribution of methane. A second movable
Fig. 1. Schematic diagram of the setup used as a two-zone fluidized-bed reactor (TZFBR) and fluidized-bed reactor (FBR). The H, Hb, and Hm are reactor height, catalyst height under fluidization condition, and the height at which methane is introduced, respectively.
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The following runs were conducted: I. Variation of temperature at Rmo = 5.0. II. Variation of Rmo at 800 °C. III. Variation of the height at which the methane was introduced (Hm) at 800 °C and Rmo = 7.0. IV. Variation of flow rate of the feed gas (as relative gas velocity) at 800 °C and Rmo = 7.0. V. Variation of inert gas ratio (as mole %) in the feed at 800 °C and Rmo = 7.0. VI. Variation of methane normalized concentration at 800 °C and Rmo = 7.0.
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Table 1 Particle size distribution of catalyst used under 150 µm sieve. Particle size (µm)
Weight (%)
120–150 90–120 90–60 Lower than 60
64.5 33.1 2.1 0.3
3. Results and discussion
With regards to the catalyst degradation i.e. attrition resistance, in a test run performed for 40 h, the catalyst was weighted and no significant loss in the catalyst weight was noticed in the reactor. The catalyst was then sieved by a 150-µm sieve, and the ratio between the weight of the catalyst particles having smaller size than 150 µm to the initial weight of the catalyst was found to be about 7%. It should be mentioned that the particle size of the fresh catalyst was between 150 and 250 µm. The particle size distribution for those smaller than 150 µm is presented in Table 1. As is evident the amount of fine for particles smaller than 90 µm is less than 3%. These results indicate that the mechanical behavior of the catalyst in the reactor, i.e. the attrition resistance of the catalyst, is quite acceptable.
3.1. XRD characterization
3.3. Temperature uniformity of fluidized bed
XRD patterns of the Mn/Na2WO4/SiO2 catalyst at various conditions are presented in Fig. 2. The fresh catalyst and the one treated in an atmosphere of oxygen at 800 °C have similar patterns which indicate the presence of Mn2O3 and Na2WO4 as main phases. The XRD pattern of Mn/Na2WO4/SiO2 fluidized in a reducing atmosphere of methane at 800 °C shows MnWO4 as a dominant phase. It appears that the Mn2O3 and Na2WO4 are transformed into MnWO4 under reducing condition, indicating that the oxygen of the catalyst lattice is consumed to a large extent. In other words, the oxidation state of manganese is changed from 3 to 2 under reducing conditions. Investigations of Hou et al. [17], which were conducted in a fixed-bed reactor, confirm this explanation. Therefore, Mn/Na2WO4/SiO2 is a promising catalyst for the two-zone fluidized-bed reactor, due to its redox property.
The temperature across the catalytic bed for both fluidized-bed reactor and the two-zone one, the temperature was monitored along the bed in order to ensure the uniformity of the reactor temperature. It was found that when temperature of the reactor was set to be 800 °C, the maximum temperature difference along the bed was measured to be about 7 °C. Above the bed of catalyst, the temperature drops sharply.
The carbon balance based on analyses of feed and product was performed and for almost all of the experiments carried out, was better than 97%. In addition, for most of the experiments performed in our study we had several measurements (at least three measurements) and the results presented in this work are the mean of those measurements. The relative errors for each repeated measurement for calculation of C2 selectivity and methane conversion did not exceed 1.2% and 2.2% respectively.
3.2. Fluidizability and attrition resistance The particle size and density of the catalyst used were 150–250 µm and 1.1 g/cm3 respectively. Based on the Geldart's classification [18], the catalyst with the mentioned specifications belongs to the group A. The minimum fluidization velocity (Umf) was calculated to be 1.8 cm/s [19]. For OCM on the Mn/Na2WO4/SiO2 catalyst at 800 °C and the utilized flow-rate range, the fluidizability was observed to be desirable.
Fig. 2. XRD patterns of Mn/Na2WO4/SiO2 catalyst: a) freshly calcined, b) fluidized under oxygen and inert at 800 °C and c) fluidized under CH4 and inert at 800 °C.
3.4. Comparison of the results obtained in TZFBR and FBR Fig. 3 presents the methane and oxygen conversions while Fig. 4 shows the products selectivity obtained in TZFBR and FBR, at different temperatures. A strong influence of temperature on activity and selectivity is observed for both reactors, in the range of 700–780 °C. With the increasing temperature, methane conversion rises, however, following the complete conversion of oxygen, at about 760 °C in the FBR and about 780 °C in the TZFBR, methane conversion remains unchanged. Fig. 4 reveals that, in the FBR, as the temperature increases, the C2 selectivity increases while that of CO2 decreases. However, when O2 conversion is completed, CO2 selectivity remains almost unchanged, whereas the C2 selectivity decreases gently due to a slight increase in
Fig. 3. Variation of CH4 and O2 conversion with temperature in the TZFBR and the FBR, QSTP = 300 ml/min, U/Umf = 3.5, Hb = 10 cm, Hm = 5.5 cm, % mole inert gas = 88, and Rmo = 5.
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Fig. 4. Variation of different products selectivity with temperature in the TZFBR and the FBR. Other conditions are the same as in Fig. 3.
CO selectivity. It is seen that the variation of CO selectivity in both reactors is not considerable. In the TZFBR, C2 selectivity increases with temperature in the whole range of temperature examined in this set of experiments. The increase in methane conversion and C2 selectivity with the increasing reaction temperature in OCM reaction in the fixed bed has also been observed earlier on many types of catalysts [20–23]. Li et al. [24] obtained almost similar trend for the variation of CH4 conversion and product distribution in OCM reaction on Mn/Na2WO4/SiO2 catalyst. The activation energy for formation of C2H6 over Mn/Na2WO4/ SiO2 catalyst was reported to be much higher compared to that for COx, i.e. 305 and 105 kJ/mol respectively [9]. This is why the OCM reaction should be conducted at high temperature; increasing the reaction temperature is favorable for enhancing C2 selectivity. Comparing the selectivity results in TZFBR and FBR, Fig. 4, it is observed that at temperatures below 780 °C, C2 selectivity in the TZFBR is lower than that for the FBR. However, at temperatures above 780 °C it is higher in TZFBR. At 800 °C, C2 selectivity in TZFBR is 66% against 63% in the FBR. It is seen that CO selectivity of both reactors is almost the same at all temperature. At temperatures higher than 780 °C, CO2 selectivity of TZFBR is lower. Consequently at temperatures above 780 °C, the C2 selectivity in TZFBR is higher than that in the FBR. Salehon et al. [15] investigated the dynamics of Mn/Na2WO4/SiO2 catalyst. They exposed the fully oxidized catalyst to a step change of CH4 in Ar and observed that the catalyst had redox properties. Both hydrocarbon products and carbon oxides were formed in the absence of gas-phase oxygen. It was found that in the absence of gas-phase oxygen, diffusion of bulk lattice oxygens to the surface is very slow. They concluded that, only surface and near-surface oxygens participate in methane conversion. Thus using larger amount of catalyst increases methane conversion. Moreover, it has been reported that the surface lattice oxygens is more selective for the formation of C2 products whereas the bulk lattice oxygens lead mainly to carbon oxides [9,15,25].
A Mars-van Krevelen type mechanism for the redox catalysts has been proposed for both cofeed and redox mode of operations [26,27]. Based on this mechanism, the oxidation of methane to methyl radicals takes place by the oxidized catalyst surface i.e. methyl radicals are formed at surface of the transition metal oxide sites. In the redox mode, i.e. cyclic operation or in TZFBR, metal oxides which have been reduced are reoxidized when the catalyst is exposed to oxidizing atmosphere. C2H6, as the primary hydrocarbon product, are formed from methyl radicals in the gas phase and carbon oxides principally from encounters of C2 with the oxidized surface, with a minor contribution to carbon oxides from oxidation of methyl radicals [28,29]. C2H4, as the secondary product, are formed by dehydrogenation of C2H6. Thus, OCM involves both the heterogeneous (surface catalyzed) and homogeneous (gas phase) reactions [9]. It is apparent that in TZFBR at high temperatures, all or a part of the surface oxygens are available for the OCM reactions. In this case the contribution of gas-phase oxygen in formation of COx products are absent to some extent, and thus the C2 selectivity is improved compared to that in the FBR. In this investigation, due to the laboratory restrictions, a small scale TZFBR was used. It is obvious that having methane and oxygen fully apart is practically impossible. That is why the extent of C2 improvement is not so significant. However, it is expected that in larger reactors, improvement of C2 selectivity is more pronounced. Fig. 5 shows that C2H4/C2H6 ratio for both reactors increases as the temperature increases. The observations are in accordance with those reported in the literature over different catalysts [21,25,29–31]. The increase in the C2H4/C2H6 product ratio with temperature suggests that the conversion of ethane to ethylene is favored at higher temperatures. This is expected due to the higher rate of both oxidative- and nonoxidative-dehydrogenation of ethane to ethylene. In TZFBR the ratio is lower for the entire range examined which suggests that oxidative dehydrogenation of ethane contributes to the formation of ethylene which may not be ignored. Effects of CH4/O2 ratios on methane conversion and products selectivity are presented in Fig. 6. This series of experiments were performed at 800 °C using 88% inert gas (Ar) and Hm = 7 cm for TZFBR. In both the TZFBR and the FBR, with decreasing Rmo, methane conversion increases and C2 selectivity decreases. This behavior is almost similar to the observed behavior in fixed-bed reactors [31,32]. At Rmo about 3, a significant difference is observed in methane conversion for two reactors, due to the large amount of oxygen. Similar to the fixed-bed reactors, higher oxygen concentration leads to higher methane conversion in FBR, whereas in TZFBR, due to the limited capacity of the catalyst for adsorption of oxygen, it does not contribute in higher methane conversion. The C2 selectivity in TZFBR is however, 4–5.5% larger than that in FBR. It is also observed that higher values of Rmo results in lower selectivity to COx products. It appears that in TZFBR, all or most part of O2 is adsorbed on the catalyst surface and the contribution of gas-phase formation of COx is reduced.
Fig. 5. Variation of C2H4/C2H6 ratio with temperature in the TZFBR and the FBR. Other conditions are the same as in Fig. 3.
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Fig. 8. Variation of C2 selectivity versus methane conversion for the TZFBR and the FBR. Other conditions are the same as in Fig. 6.
Fig. 6. Variation of CH4 conversion and different product selectivity with the ratio of CH4/O2 in the TZFBR and the FBR, QSTP = 300 ml/min, U/Umf = 3.5, T = 800 °C, Hb = 10 cm, Hm = 7 cm, % mole inert gas = 88, and mcat. = 7.5 g.
segregation of oxidation and reduction zones for this system. The TZFBR results demonstrate that it is possible to have continuous operation while avoiding the simultaneous presence of oxygen and hydrocarbons in the gas phase. Fig. 9 presents variation of methane conversion and product selectivity versus height of methane inlet (Hm). It is seen that by decreasing the height of methane inlet (Hm) entering TZFBR, the length of reaction zone increases and that of the oxidation zone decreases. As a result, the separation of two zones inside the reactor becomes less distinct and consequently C2 selectivity decreases. On the other hand, due to the increasing of the length of reaction zone, the reactor becomes more similar to a FBR and thus, conversion has been increased. Moreover, by increasing the height of methane inlet (Hm), though separation of two oxidation and reduction zones
Fig. 7 presents variation of C2 yield and C2H4/C2H6 ratio versus Rmo for both reactors. It is seen that the C2 yield of both reactors is almost equal at R mo between 5 and 9. Comparing FBR and TZFBR performance, the TZFBR is preferred because at equal C2 yield it has higher C2 selectivity. Fig. 7 also shows that the C2H4/C2H6 ratio in the TZFBR is less than that in the FBR. As mentioned earlier, ethylene is considered to be a secondary product and is formed by dehydrogenation of ethane. Due to the larger residence time of ethane in FBR, i.e. the whole length of FBR, as opposed to TZBBR, i.e. the upper zone of TZFBR, the extent of ethane dehydrogenation of C2H6 in the TZFBR is less than that in the FBR. Variation of C2 selectivity versus methane conversion in FBR and TZFBR is presented in Fig. 8. It is seen that the TZFBR, at low to medium conversion, gives a higher C2 selectivity at the same methane conversion. This may be considered as one of the advantages of
Fig. 7. Variation of the yield and C2H4/C2H6 ratio with CH4/O2 ratio for the TZFBR and fluidized-bed FBR. Other conditions are the same as in Fig. 6.
Fig. 9. Variation of the CH4 conversion and different product selectivity with the height of methane inlet at the TZFBR, QSTP = 300 ml/min, U/Umf = 3.5, CH4/O2 = 7, T = 800 °C, Hb = 10 cm, % mole inert gas = 88, and mcat. = 7.5 g.
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becomes more distinct, but because of the narrowing of OCM reaction zone and reaching below a specified limit, methane conversion decreases. The more desirable conditions between these two limits are those which results in a relatively higher conversion provided that the higher selectivity is achieved. It can be observed that at Hm of 7 cm, 800 °C and Rmo = 7, C2 selectivity in TZFBR and FBR are 73% and 67.5% at respectively. Fig. 10 shows variation of the CH4 conversion and C2 selectivity and yield with the U/Umf ratio in the TZFBR and the FBR. In the ranges of U/ Umf under study for FBR, with changing flow rate no change was noticed in C2 selectivity, CH4 conversion and yield. However, for TZFBR increase in flow rate in the range of U/Umf = 2–3, C2 selectivity has no appreciable change, noting that the value of C2 selectivity is approximately 5% higher for TZFBR as compared with that of FBR. While for the value of U/Umf N 3, C2 selectivity decreases. It is noticed that the methane conversion for TZFBR remains almost constant and similar to FBR in the range of U/Umf = 2–3, while for the value of U/Umf N 3, methane conversion decreases sharply. For TZFBR, the yield has no appreciable change in the range of U/Umf = 2–3, indicating that the value of the yield is approximately 1.1% higher than that of FBR. While for the value of U/Umf N 3, the yield decreases sharply. It may be argued that higher U/Umf increases the amount of gas in bubble phase, which in turn retards the incorporation of oxygen into the emulsion phase inside the oxidation zone, and may give rise to oxygen bypass into the reaction zone and outside the bed. Thus, there is less separation of oxidation and reduction zones at a given feed composition in the range of U/Umf under investigation. This causes C2 selectivity and CH4 conversion to be decreased. Effects of inert gas ratio in the feed on methane conversion and C2 selectivity and yield for TZFBR as well as FBR are presented in Fig. 11. It is seen that for FBR, feed dilution has no effect on C2 selectivity and methane conversion; although in case of the fixed-bed reactor, it may
Fig. 11. Variation of the CH4 conversion and C2 selectivity and yield with the inert gas mole% in the feed in the TZFBR and FBR, U/Umf = 3, CH4/O2 = 7, T = 800 °C, Hb = 10 cm, Hm = 7 cm, and mcat. = 7.5 g.
be effective for preventing the formation of overheated zones, this will not occur in fluidized-bed reactor due to temperature uniformity. This fact has also been verified by Mleczko et al. [6,33]. For OCM in fluidized-bed reactor they found that the feed dilution had almost no effect on C2 selectivity and yield. In case of TZFBR, it is seen that at an 88% value of the inert gas, C2 selectivity and yield are higher than those of FBR, whereas CH4 conversion is almost identical. By decreasing the inert gas, the three above parameters decrease. C2 selectivity may decrease if separation of the oxidation and reduction
Fig.10. Variation of the CH4 conversion and C2 selectivity and yield with the U/Umf in the TZFBR and the FBR, CH4/O2 = 7, T = 800 °C, Hb = 10 cm, % mole inert gas = 88 and mcat. = 7.5 g.
Fig. 12. Variation of the normalized concentration of methane along the bed at U/Umf = 3, CH4/O2 = 7, T = 800 °C, Hb = 10 cm, and mcat. = 7.5 g.
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zones becomes less distinct. On the other hand, due to a narrower reaction zone in TZFBR as compared to FBR, with decreasing inert gas and increasing the ratio of reactants, conversion will also decrease. Variations in the normalized concentration of methane along the reactor length are presented in Fig. 12. As can be seen, the inlet methane concentration attains the maximum value, and inside OCM reaction zone its concentration decreases by being converted to products. On the other hand, due to diffusion of methane around the inlet point, it is seen at a distance of about 1 cm lower than the inlet point, where the concentration is decreased rapidly and diminished. 4. Conclusion In this research, the effects of feed segregation, the so called TZFBR, on the performance of a fluidized-bed reactor for OCM reaction were investigated and compared to the performance of a conventional fluidized-bed reactor. The Mn/Na2WO4/SiO2 catalyst employed exhibited a redox property which is the main requirement for the operation of TZFBR. It was seen that the C2 yield of both reactors was almost equal at CH4/O2 in the range of 5–9. However, TZFBR resulted in a higher C2 selectivity at equal C2 yield. The C2 selectivity in TZFBR was 4–5.5% larger than that in FBR. The reaction of methane with selective lattice/ surface oxygen species may have led to the higher selectivity. The TZFBR results demonstrate that it is possible to have a continuous operation for OCM reaction while avoiding the simultaneous presence of oxygen and hydrocarbons in the gas phase. Acknowledgment The authors appreciate a partial financial support from Iranian National Petrochemical Company under contract No. 84131. References [1] G.E. Keller, M.M. Bhasin, Synthesis of ethylene via oxidative coupling of methane: I. Determination of active catalysts, J. Catal. 73 (1982) 9–19. [2] X. Fang, S. Li, J. Lin, Y. Chu, Oxidative coupling of methane on W–Mn catalysts, J. Mol. Catal. (China) 6 (1992) 427–433. [3] D. Wang, M.P. Rosynek, J.H. Lunsford, Oxidative coupling of methane over oxidesupported sodium–manganese catalysts, J. Catal. 155 (1995) 390–402. [4] S. Pak, J.H. Lunsford, Thermal effects during the oxidative coupling of methane over Mn/Na2WO4/SiO2 and Mn/Na2WO4/MgO catalysts, Appl. Catal., A 168 (1998) 131–137. [5] G. Follmer, L. Lehmann, M. Baerns, The application of laboratory scale catalytic fixed and fluidized bed reactors in the oxidative coupling of methane, ACS Div. Petrol. Chem. 33 (1988) 453–459. [6] L. Mleczko, U. Pannek, V.M. Niemi, J. Hiltunen, Oxidative coupling of methane in a fluidized-bed reactor over a highly active and selective catalyst, Ind. Eng. Chem. Res. 35 (1996) 54–61. [7] A. Santos, J. Santamaria, M. Menendez, Oxidative coupling of methane in a vibrofluidized bed at low fluidizing velocities, Ind. Eng. Chem. Res. 34 (1995) 1581–1587. [8] R. Andorf, L. Mleczco, D. Schweer, M. Baerns, Oxidative coupling of methane in a bubbling fluidized bed reactor, Can. J. Chem. Eng. 69 (1991) 891–897. [9] S. Li, Reaction chemistry of W–Mn/SiO2 catalyst for the oxidative coupling of methane, J. Nat. Gas Chem. 12 (2003) 1–9.
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[10] F.T. Akin, Y.S. Lin, Oxidative coupling of methane in dense ceramic membrane reactor with high yields, AIChE J. 48 (2002) 2298–2305. [11] Y.P. Lu, A.G. Dixon, W.R. Moser, Y.H. Ma, Oxidative coupling of methane using oxygen-permeable dense membrane reactors, Catal. Today 56 (2000) 297–305. [12] M.L. Pacheco, J. Soler, A. Dejoz, J.M. Lopez Nieto, J. Herguido, M. Menendez, J. Santamaria, MoO3/MgO as a catalyst in the oxidative dehydrogenation of n-butane in a two-zone fluidized bed reactor, Catal. Today 61 (2000) 101–107. [13] O. Rubio, J. Herguido, M. Menendez, Two-zone fluidized bed reactor for simultaneous reaction and catalyst reoxidation: influence of reactor size, Appl. Catal., A Gen. 272 (2004) 321–327. [14] J. Gascon, C. Tellez, J. Herguido, M. Menendez, A two-zone fluidized bed reactor for catalytic propane dehydrogenation, Chem. Eng. J. 106 (2005) 91–96. [15] V. Salehoun, A.A. Khodadadi, Y. Mortazavi, A. Talebizadeh, Dynamics of Mn/ Na2WO4/SiO2 catalyst in oxidative coupling of methane, Chem. Eng. Sci. 63 (2008) 4910–4916. [16] A. Malekzadeh, M. Abedini, A.A. Khodadadi, M. Aamini, H.K. Mishra, A.K. Dalai, Critical influence of Mn on low-temperature catalytic activity of Mn/Na2WO4/SiO2 catalyst for oxidative coupling of methane, Catal. Lett. 84 (2002) 45–51. [17] S. Hou, Y. Cao, W. Xiong, H. Liu, Y. Kou, Site requirements for the oxidative coupling of methane on SiO2-supported Mn catalysts, Ind. Eng. Chem. Res. 45 (2006) 7077–7083. [18] D. Geldart, Types of gas fluidization, Powder Technol. 7 (1973) 285–292. [19] D. Kunii, O. Levenspiel, Fluidization Engineering, 2nd Ed, Butterworth-Heinemann, Boston, 1991, pp. 69–70. [20] S.J. Korf, J.A. Roos, L.J. Veltman, J.G. Van Ommen, J.R.H. Ross, Effect of additives on lithium doped magnesium oxide catalysts used in the oxidative coupling of methane, Appl. Catal. 56 (1989) 119–135. [21] V.R. Choudhary, V.H. Rane, Y.P. Meenakshi, Comparison of alkali metal promoted MgO catalysts for their surface acidity/basicity and catalytic activity/selectivity in the oxidative coupling of methane, J. Chem. Technol. Biotechnol. 68 (1997) 177–186. [22] V.H. Rane, S.T. Chaudhari, V.R. Choudhary, Comparison of the surface and catalytic properties of rare earth-promoted CaO catalysts in the oxidative coupling of methane, J. Chem. Technol. Biotechnol. 81 (2006) 208–215. [23] V.R. Choudhary, V.H. Rane, Surface properties of CaO (or BaO) La2O3·MgO catalysts and their performance in oxidative coupling of methane, J. Chem. Technol. Biotechnol. 69 (1997) 63–69. [24] J. Wang, L. Chou, B. Zhang, H. Song, J. Zhao, J. Yang, S. Li, Comparative study on oxidation of methane to ethane and ethylene over Na2WO4-Mn/SiO2 catalysts prepared by different methods, J. Mol. Catal., A Chem. 245 (2005) 272–277. [25] S. Pak, P. Qiu, J.H. Lunsford, Elementary reactions in the oxidative coupling of methane over Mn/Na2WO4/SiO2 and Mn/Na2WO4/MgO catalysts, J. Catal. 179 (1998) 222–230. [26] C.A. Jones, J.J. Leonard, J.A. Sofranko, The oxidative conversion of methane to higher hydrocarbons over alkali-promoted Mn/SiO2, J. Catal. 103 (1987) 311–319. [27] J.A. Sofranko, J.J. Leonard, C.A. Jones, A.M. Gaffney, H.P. Withers, Catalytic oxidative coupling of methane over sodium-promoted Mn/SiO2 and Mn/MgO, Catal. Today 3 (1988) 127–135. [28] C.H. Lin, K.D. Campbell, J.X. Wang, J.H. Lunsford, Oxidative dimerization of methane over lanthanum oxide, J. Phys. Chem. 90 (1986) 534–537. [29] T. Ito, J.X. Wang, C.H. Lin, J.H. Lunsford, Oxidative dimerization of methane over a lithium-promoted magnesium oxide catalyst, J. Am. Chem. Soc. 107 (1985) 5062–5068. [30] V.R. Choudhary, V.H. Rane, Oxidative coupling of methane over Sm-promoted MgO: influence of composition and preparation conditions, J. Chem. Technol. Biotechnol. 70 (1997) 325–330. [31] V.R. Choudhary, V.H. Rane, Acidity/basicity of rare earth oxides and their catalytic activity/selectivity in oxidative coupling of methane to C2-hydrocarbons, J. Catal. 130 (1991) 411–412. [32] V.R. Choudhary, S.A.R. Mulla, B.S. Uphade, Oxidative coupling of methane over alkaline earth oxides deposited on commercial support precoated with rare earth oxides, Fuel 78 (1999) 427–437. [33] L. Mleczko, U. Pannek, M. Rothaemel, M. Baerns, Oxidative coupling of methane over a La2O3/CaO catalyst. Optimization of reaction conditions in a bubbling fluidized-bed reactor, Can. J. Chem. Eng. 74 (1996) 279–287.
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Fuel Processing Technology j o u r n a l h o m e p a g e : w w w. e l s ev i e r. c o m / l o c a t e / f u p r o c
Comparative study of the two-zone fluidized-bed reactor and the fluidized-bed reactor for oxidative coupling of methane over Mn/Na2WO4/SiO2 catalyst A. Talebizadeh, Y. Mortazavi ⁎, A.A. Khodadadi Oil and Gas Center of Excellence, School of Chemical Engineering, University College of Engineering, University of Tehran, P.O. BOX 11155-4563, Tehran, Iran
a r t i c l e
i n f o
Article history: Received 25 January 2009 Received in revised form 3 June 2009 Accepted 23 June 2009 Keywords: Oxidative coupling Methane Two-zone Fluidized Mn/Na2WO4/SiO2 Redox
a b s t r a c t In this work, oxidative coupling of methane over Mn/Na2WO4/SiO2 catalyst is studied in a two-zone fluidized-bed reactor (TZFBR) and its performance is compared with a fluidized-bed reactor (FBR). Diluted oxygen in argon was introduced into the bottom of the TZFBR through a quartz ferrite and methane was entered at higher positions along the fluidized bed. The catalyst circulated between the oxygen-rich and methane-rich zones in the TZFBR reactor. The effects of the main operating variables including bed temperature, the methane/oxygen ratio (Rmo), and the height at which methane was introduced into the reactor (Hm) were investigated. It is found that under some operating conditions the TZFBR gives a higher C2 selectivity than that obtained in the FBR reactor. Reaction of methane with lattice oxygen of the Mn/ Na2WO4/SiO2 redox catalyst in the methane-rich zone may have led to the higher selectivity. © 2009 Elsevier B.V. All rights reserved.
1. Introduction A great deal of attention has focused on the production of C2+ hydrocarbons, mainly C2H6 and C2H4, from oxidative coupling of methane (OCM) since the early work of Keller and Bhasin [1]. It is noticeable that in this process CO2 and CO are undesirable byproducts which cause decrease of the C2 selectivity. A lot of investigation has been done for increasing C2 selectivity and yield by means of different catalyst formulations and/or by reaction engineering in order to make the process commercially feasible. One of the most effective catalysts for the OCM is believed to be Mn/Na2WO4/SiO2, which was first introduced by Fang et al. [2]. It was reported that with this catalyst in the fixed-bed reactor, it is possible to reach a C2+ selectivity up to 80% at a CH4 conversion of about 20% for periods longer than 90 h [3,4]. The OCM reactions are highly exothermic and predominantly carried out in the presence of a heterogeneous catalyst. In the fixedbed reactors, especially in the industrial scale ones, due to the large amount of heat released in the course of reaction, the operation of the reactor is accident prone. Pak and Lunsford [4] described the thermal effects and magnitude of the hot spots for Mn/Na2WO4/SiO2 and Mn/ Na2WO4/MgO catalysts. They found that the exothermic reactions that occur during the OCM over these catalysts resulted in a small region in the catalyst bed which had a temperature that was 150 °C higher than the reactor wall. Also, Follmer et al. [5] observed steep axial and radial temperature gradients in the catalytic fixed-bed
⁎ Corresponding author. Tel.: +98 21 61112193; fax: +98 21 66967793. E-mail address: [email protected] (Y. Mortazavi). 0378-3820/$ – see front matter © 2009 Elsevier B.V. All rights reserved. doi:10.1016/j.fuproc.2009.06.021
reactors. It is obvious that, hot spots have detrimental consequences on the operation of the reactor, such as temperature runaway, catalyst deactivation, undesired side reactions and thermal decomposition of products. Therefore, heat management and temperature control are the key factors for the selection of the reactor type for OCM process. Temperature homogeneity may be achieved by applying fluidized-bed reactors in which intensive mixing of the particulate solids results in a relatively fast dissipation of the generated heat within the catalytic reactor. Some researchers have even claimed that OCM may be carried out isothermally in the fluidized-bed reactors [6–8]. Apart from selective catalysts, the reactor type and the mode of contact between the reactants are considered to play key roles in selective oxidation processes. It is often found that the reaction takes place using oxygen from the catalyst lattice, and the contribution of gas-phase oxygen is generally considered a determinant to selectivity. In the OCM reaction, the catalyst is more selective than the gas phase towards formation of C2 products [9]. This has led to the reactor designs that attempt to reduce the concentration of gas-phase oxygen in the reaction. Over the past two decades, several groups have studied OCM on catalytically active membrane which changes the OCM reaction mechanism or minimizes the presence of the gas-phase oxygen [10,11]. In these studies, the C2 selectivity was much higher than that obtained in the cofeed reactors. An alternative approach is to avoid the simultaneous presence of hydrocarbons and oxygen over the catalyst. In this approach reducible (redox) catalysts may be used either in a circulating bed reactor, in which the catalyst is transferred between the reactor and the regenerator, or in a two-zone fluidized-bed reactor in which it uses segregated feeds for the oxidant and the hydrocarbon in a single
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Nomenclature dp H Hb Hm Umf mcat. Dreactor QSTP ρc
Particle diameter, µm Reactor height, cm Height of the catalyst bed under fluidization condition, cm The height at which methane is introduced in the TZFB, cm Minimum fluidization velocity, cm/s Amount of the catalyst in the reactor, g Reactor diameter, cm Total feed flow rate at standard condition, ml/min Catalyst density, g/cm3
reactor vessel. TZFBR allows hydrocarbon oxidation make use of lattice oxygen of the catalyst in the reduction zone of the bed (above the hydrocarbon inlet) in absence of the gas-phase oxygen. The oxygen depleted (reduced) catalyst is then regenerated (reoxidized) by internal recirculation to the oxidation zone at the bottom part of the reactor where oxygen is fed. Two-zone fluidized-bed reactors were previously used in the laboratory studies for oxidative dehydrogenation of butane and propane and it was reported that the selectivity of desired products increased [12–14]. Step-change experiments on the Mn/Na2WO4/SiO2 catalyst revealed that the catalyst is capable of performing OCM reaction, with desirable C2 selectivity, in the absence of gas-phase oxygen [15]. In this work, oxidative coupling of methane over Mn/Na2WO4/ SiO2 catalyst is investigated in a two-zone fluidized-bed reactor and the results are compared with those obtained in a conventional fluidized-bed reactor.
quartz probe, located approximately near the reactor axis, was used to take samples at different bed heights which were sent to GC for analyses. The both probes were used for TZFBR experiments. The reactor was placed in an electrically heated furnace. Temperature of the reactor was controlled using a thermocouple mounted in the outer wall of the reactor, centered along the catalyst bed. The temperature of the furnace remains almost constant. Another thermocouple was placed in a quartz tube sheathe and located inside the bed to monitor the catalyst bed temperature. Flow rates of all feed gases, CH4 (99.999%), O2 (99.999%) and Ar (99.995%), were regulated by mass flow controllers (Unit Instrument UFC-1200GI). The effluent of the reactor was first passed through a condenser to remove almost all of the water vapor and then analyzed with an on-line Carle 400A GC equipped with a Propack-Q packed column, a methanizer and an FID detector. CO and CO2 were converted to methane in the methanizer to provide an accurate detection of them by using FID detector. The feed was also analyzed by a bypass stream. In the TZFBR, methane was fed at the mid-point along the fluidized bed with Hm height from the bottom of the reactor, and O2–Ar mixture was fed at the bottom of the reactor, Fig. 1. In the FBR, the feed gas mixture, i.e. methane, oxygen and argon were premixed and fed through the quartz distributor at the bottom. The main operating variables studied are: bed temperature, methane/oxygen ratio (Rmo), and the height at which methane is introduced (Hm), gas flow rate (as gas relative velocity), as well as inert gas ratio (as mole %) in the feed. The experiments performed on 7.5 g of the catalyst and a total flow rate of 300 sccm/min, except for the experiments which are done for the effects of flow rate. This amount of the catalyst corresponds to a height equivalent to 10 cm.
2. Experiments 2.1. Catalyst preparation The 4.0 wt%Mn/5.0 wt%Na2WO4/SiO2 catalyst was prepared by dry impregnation method, using Davisil grade 645 silica gel (Aldrich). An aqueous solution of Mn(NO3)2·4H2O (Merck) with appropriate concentration was added dropwise to silica gel and dried at 130 °C overnight. The catalyst was then impregnated with an aqueous solution of Na2WO4·2H2O (Merck) with appropriate concentration, dried again at 130 °C overnight and calcined at 800 °C for 8 h. The catalyst was sieved to make a desired particle size with diameters of 150–250 µm. The catalyst density was measured to be 1.1 g/cm3. X-ray diffraction (XRD) patterns of the catalyst at the following conditions were obtained by a Philips diffractometer with Cu-Kα target in the range of 4–70o (2θ): a—the fresh catalyst, b—the catalyst under O2 and inert gas at 800 °C was fluidized and then cooled at the same condition, i.e. fully oxidized catalyst, and c—the catalyst under methane and inert gas at 800 °C was fluidized and then cooled at the same condition, i.e. fully reduced catalyst. The BET surface area, pore volume and pore size distribution of the catalyst have been reported elsewhere [16]. 2.2. Setup A schematic of the setup for two-zone fluidized-bed reactor (TZFBR) and fluidized-bed reactor (FBR) is shown in Fig. 1. The reactor was 20 mm I.D. and 25 cm length quartz tube equipped with a sintered quartz distributor plate with nominal pore size between 10 and 16 μm and an axial movable quartz probe to introduce the methane feed at the desired bed height (Hm). This probe terminates with a bulb with 8 fine holes uniformly located on its surface, to ensure a good radial distribution of methane. A second movable
Fig. 1. Schematic diagram of the setup used as a two-zone fluidized-bed reactor (TZFBR) and fluidized-bed reactor (FBR). The H, Hb, and Hm are reactor height, catalyst height under fluidization condition, and the height at which methane is introduced, respectively.
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The following runs were conducted: I. Variation of temperature at Rmo = 5.0. II. Variation of Rmo at 800 °C. III. Variation of the height at which the methane was introduced (Hm) at 800 °C and Rmo = 7.0. IV. Variation of flow rate of the feed gas (as relative gas velocity) at 800 °C and Rmo = 7.0. V. Variation of inert gas ratio (as mole %) in the feed at 800 °C and Rmo = 7.0. VI. Variation of methane normalized concentration at 800 °C and Rmo = 7.0.
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Table 1 Particle size distribution of catalyst used under 150 µm sieve. Particle size (µm)
Weight (%)
120–150 90–120 90–60 Lower than 60
64.5 33.1 2.1 0.3
3. Results and discussion
With regards to the catalyst degradation i.e. attrition resistance, in a test run performed for 40 h, the catalyst was weighted and no significant loss in the catalyst weight was noticed in the reactor. The catalyst was then sieved by a 150-µm sieve, and the ratio between the weight of the catalyst particles having smaller size than 150 µm to the initial weight of the catalyst was found to be about 7%. It should be mentioned that the particle size of the fresh catalyst was between 150 and 250 µm. The particle size distribution for those smaller than 150 µm is presented in Table 1. As is evident the amount of fine for particles smaller than 90 µm is less than 3%. These results indicate that the mechanical behavior of the catalyst in the reactor, i.e. the attrition resistance of the catalyst, is quite acceptable.
3.1. XRD characterization
3.3. Temperature uniformity of fluidized bed
XRD patterns of the Mn/Na2WO4/SiO2 catalyst at various conditions are presented in Fig. 2. The fresh catalyst and the one treated in an atmosphere of oxygen at 800 °C have similar patterns which indicate the presence of Mn2O3 and Na2WO4 as main phases. The XRD pattern of Mn/Na2WO4/SiO2 fluidized in a reducing atmosphere of methane at 800 °C shows MnWO4 as a dominant phase. It appears that the Mn2O3 and Na2WO4 are transformed into MnWO4 under reducing condition, indicating that the oxygen of the catalyst lattice is consumed to a large extent. In other words, the oxidation state of manganese is changed from 3 to 2 under reducing conditions. Investigations of Hou et al. [17], which were conducted in a fixed-bed reactor, confirm this explanation. Therefore, Mn/Na2WO4/SiO2 is a promising catalyst for the two-zone fluidized-bed reactor, due to its redox property.
The temperature across the catalytic bed for both fluidized-bed reactor and the two-zone one, the temperature was monitored along the bed in order to ensure the uniformity of the reactor temperature. It was found that when temperature of the reactor was set to be 800 °C, the maximum temperature difference along the bed was measured to be about 7 °C. Above the bed of catalyst, the temperature drops sharply.
The carbon balance based on analyses of feed and product was performed and for almost all of the experiments carried out, was better than 97%. In addition, for most of the experiments performed in our study we had several measurements (at least three measurements) and the results presented in this work are the mean of those measurements. The relative errors for each repeated measurement for calculation of C2 selectivity and methane conversion did not exceed 1.2% and 2.2% respectively.
3.2. Fluidizability and attrition resistance The particle size and density of the catalyst used were 150–250 µm and 1.1 g/cm3 respectively. Based on the Geldart's classification [18], the catalyst with the mentioned specifications belongs to the group A. The minimum fluidization velocity (Umf) was calculated to be 1.8 cm/s [19]. For OCM on the Mn/Na2WO4/SiO2 catalyst at 800 °C and the utilized flow-rate range, the fluidizability was observed to be desirable.
Fig. 2. XRD patterns of Mn/Na2WO4/SiO2 catalyst: a) freshly calcined, b) fluidized under oxygen and inert at 800 °C and c) fluidized under CH4 and inert at 800 °C.
3.4. Comparison of the results obtained in TZFBR and FBR Fig. 3 presents the methane and oxygen conversions while Fig. 4 shows the products selectivity obtained in TZFBR and FBR, at different temperatures. A strong influence of temperature on activity and selectivity is observed for both reactors, in the range of 700–780 °C. With the increasing temperature, methane conversion rises, however, following the complete conversion of oxygen, at about 760 °C in the FBR and about 780 °C in the TZFBR, methane conversion remains unchanged. Fig. 4 reveals that, in the FBR, as the temperature increases, the C2 selectivity increases while that of CO2 decreases. However, when O2 conversion is completed, CO2 selectivity remains almost unchanged, whereas the C2 selectivity decreases gently due to a slight increase in
Fig. 3. Variation of CH4 and O2 conversion with temperature in the TZFBR and the FBR, QSTP = 300 ml/min, U/Umf = 3.5, Hb = 10 cm, Hm = 5.5 cm, % mole inert gas = 88, and Rmo = 5.
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Fig. 4. Variation of different products selectivity with temperature in the TZFBR and the FBR. Other conditions are the same as in Fig. 3.
CO selectivity. It is seen that the variation of CO selectivity in both reactors is not considerable. In the TZFBR, C2 selectivity increases with temperature in the whole range of temperature examined in this set of experiments. The increase in methane conversion and C2 selectivity with the increasing reaction temperature in OCM reaction in the fixed bed has also been observed earlier on many types of catalysts [20–23]. Li et al. [24] obtained almost similar trend for the variation of CH4 conversion and product distribution in OCM reaction on Mn/Na2WO4/SiO2 catalyst. The activation energy for formation of C2H6 over Mn/Na2WO4/ SiO2 catalyst was reported to be much higher compared to that for COx, i.e. 305 and 105 kJ/mol respectively [9]. This is why the OCM reaction should be conducted at high temperature; increasing the reaction temperature is favorable for enhancing C2 selectivity. Comparing the selectivity results in TZFBR and FBR, Fig. 4, it is observed that at temperatures below 780 °C, C2 selectivity in the TZFBR is lower than that for the FBR. However, at temperatures above 780 °C it is higher in TZFBR. At 800 °C, C2 selectivity in TZFBR is 66% against 63% in the FBR. It is seen that CO selectivity of both reactors is almost the same at all temperature. At temperatures higher than 780 °C, CO2 selectivity of TZFBR is lower. Consequently at temperatures above 780 °C, the C2 selectivity in TZFBR is higher than that in the FBR. Salehon et al. [15] investigated the dynamics of Mn/Na2WO4/SiO2 catalyst. They exposed the fully oxidized catalyst to a step change of CH4 in Ar and observed that the catalyst had redox properties. Both hydrocarbon products and carbon oxides were formed in the absence of gas-phase oxygen. It was found that in the absence of gas-phase oxygen, diffusion of bulk lattice oxygens to the surface is very slow. They concluded that, only surface and near-surface oxygens participate in methane conversion. Thus using larger amount of catalyst increases methane conversion. Moreover, it has been reported that the surface lattice oxygens is more selective for the formation of C2 products whereas the bulk lattice oxygens lead mainly to carbon oxides [9,15,25].
A Mars-van Krevelen type mechanism for the redox catalysts has been proposed for both cofeed and redox mode of operations [26,27]. Based on this mechanism, the oxidation of methane to methyl radicals takes place by the oxidized catalyst surface i.e. methyl radicals are formed at surface of the transition metal oxide sites. In the redox mode, i.e. cyclic operation or in TZFBR, metal oxides which have been reduced are reoxidized when the catalyst is exposed to oxidizing atmosphere. C2H6, as the primary hydrocarbon product, are formed from methyl radicals in the gas phase and carbon oxides principally from encounters of C2 with the oxidized surface, with a minor contribution to carbon oxides from oxidation of methyl radicals [28,29]. C2H4, as the secondary product, are formed by dehydrogenation of C2H6. Thus, OCM involves both the heterogeneous (surface catalyzed) and homogeneous (gas phase) reactions [9]. It is apparent that in TZFBR at high temperatures, all or a part of the surface oxygens are available for the OCM reactions. In this case the contribution of gas-phase oxygen in formation of COx products are absent to some extent, and thus the C2 selectivity is improved compared to that in the FBR. In this investigation, due to the laboratory restrictions, a small scale TZFBR was used. It is obvious that having methane and oxygen fully apart is practically impossible. That is why the extent of C2 improvement is not so significant. However, it is expected that in larger reactors, improvement of C2 selectivity is more pronounced. Fig. 5 shows that C2H4/C2H6 ratio for both reactors increases as the temperature increases. The observations are in accordance with those reported in the literature over different catalysts [21,25,29–31]. The increase in the C2H4/C2H6 product ratio with temperature suggests that the conversion of ethane to ethylene is favored at higher temperatures. This is expected due to the higher rate of both oxidative- and nonoxidative-dehydrogenation of ethane to ethylene. In TZFBR the ratio is lower for the entire range examined which suggests that oxidative dehydrogenation of ethane contributes to the formation of ethylene which may not be ignored. Effects of CH4/O2 ratios on methane conversion and products selectivity are presented in Fig. 6. This series of experiments were performed at 800 °C using 88% inert gas (Ar) and Hm = 7 cm for TZFBR. In both the TZFBR and the FBR, with decreasing Rmo, methane conversion increases and C2 selectivity decreases. This behavior is almost similar to the observed behavior in fixed-bed reactors [31,32]. At Rmo about 3, a significant difference is observed in methane conversion for two reactors, due to the large amount of oxygen. Similar to the fixed-bed reactors, higher oxygen concentration leads to higher methane conversion in FBR, whereas in TZFBR, due to the limited capacity of the catalyst for adsorption of oxygen, it does not contribute in higher methane conversion. The C2 selectivity in TZFBR is however, 4–5.5% larger than that in FBR. It is also observed that higher values of Rmo results in lower selectivity to COx products. It appears that in TZFBR, all or most part of O2 is adsorbed on the catalyst surface and the contribution of gas-phase formation of COx is reduced.
Fig. 5. Variation of C2H4/C2H6 ratio with temperature in the TZFBR and the FBR. Other conditions are the same as in Fig. 3.
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Fig. 8. Variation of C2 selectivity versus methane conversion for the TZFBR and the FBR. Other conditions are the same as in Fig. 6.
Fig. 6. Variation of CH4 conversion and different product selectivity with the ratio of CH4/O2 in the TZFBR and the FBR, QSTP = 300 ml/min, U/Umf = 3.5, T = 800 °C, Hb = 10 cm, Hm = 7 cm, % mole inert gas = 88, and mcat. = 7.5 g.
segregation of oxidation and reduction zones for this system. The TZFBR results demonstrate that it is possible to have continuous operation while avoiding the simultaneous presence of oxygen and hydrocarbons in the gas phase. Fig. 9 presents variation of methane conversion and product selectivity versus height of methane inlet (Hm). It is seen that by decreasing the height of methane inlet (Hm) entering TZFBR, the length of reaction zone increases and that of the oxidation zone decreases. As a result, the separation of two zones inside the reactor becomes less distinct and consequently C2 selectivity decreases. On the other hand, due to the increasing of the length of reaction zone, the reactor becomes more similar to a FBR and thus, conversion has been increased. Moreover, by increasing the height of methane inlet (Hm), though separation of two oxidation and reduction zones
Fig. 7 presents variation of C2 yield and C2H4/C2H6 ratio versus Rmo for both reactors. It is seen that the C2 yield of both reactors is almost equal at R mo between 5 and 9. Comparing FBR and TZFBR performance, the TZFBR is preferred because at equal C2 yield it has higher C2 selectivity. Fig. 7 also shows that the C2H4/C2H6 ratio in the TZFBR is less than that in the FBR. As mentioned earlier, ethylene is considered to be a secondary product and is formed by dehydrogenation of ethane. Due to the larger residence time of ethane in FBR, i.e. the whole length of FBR, as opposed to TZBBR, i.e. the upper zone of TZFBR, the extent of ethane dehydrogenation of C2H6 in the TZFBR is less than that in the FBR. Variation of C2 selectivity versus methane conversion in FBR and TZFBR is presented in Fig. 8. It is seen that the TZFBR, at low to medium conversion, gives a higher C2 selectivity at the same methane conversion. This may be considered as one of the advantages of
Fig. 7. Variation of the yield and C2H4/C2H6 ratio with CH4/O2 ratio for the TZFBR and fluidized-bed FBR. Other conditions are the same as in Fig. 6.
Fig. 9. Variation of the CH4 conversion and different product selectivity with the height of methane inlet at the TZFBR, QSTP = 300 ml/min, U/Umf = 3.5, CH4/O2 = 7, T = 800 °C, Hb = 10 cm, % mole inert gas = 88, and mcat. = 7.5 g.
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becomes more distinct, but because of the narrowing of OCM reaction zone and reaching below a specified limit, methane conversion decreases. The more desirable conditions between these two limits are those which results in a relatively higher conversion provided that the higher selectivity is achieved. It can be observed that at Hm of 7 cm, 800 °C and Rmo = 7, C2 selectivity in TZFBR and FBR are 73% and 67.5% at respectively. Fig. 10 shows variation of the CH4 conversion and C2 selectivity and yield with the U/Umf ratio in the TZFBR and the FBR. In the ranges of U/ Umf under study for FBR, with changing flow rate no change was noticed in C2 selectivity, CH4 conversion and yield. However, for TZFBR increase in flow rate in the range of U/Umf = 2–3, C2 selectivity has no appreciable change, noting that the value of C2 selectivity is approximately 5% higher for TZFBR as compared with that of FBR. While for the value of U/Umf N 3, C2 selectivity decreases. It is noticed that the methane conversion for TZFBR remains almost constant and similar to FBR in the range of U/Umf = 2–3, while for the value of U/Umf N 3, methane conversion decreases sharply. For TZFBR, the yield has no appreciable change in the range of U/Umf = 2–3, indicating that the value of the yield is approximately 1.1% higher than that of FBR. While for the value of U/Umf N 3, the yield decreases sharply. It may be argued that higher U/Umf increases the amount of gas in bubble phase, which in turn retards the incorporation of oxygen into the emulsion phase inside the oxidation zone, and may give rise to oxygen bypass into the reaction zone and outside the bed. Thus, there is less separation of oxidation and reduction zones at a given feed composition in the range of U/Umf under investigation. This causes C2 selectivity and CH4 conversion to be decreased. Effects of inert gas ratio in the feed on methane conversion and C2 selectivity and yield for TZFBR as well as FBR are presented in Fig. 11. It is seen that for FBR, feed dilution has no effect on C2 selectivity and methane conversion; although in case of the fixed-bed reactor, it may
Fig. 11. Variation of the CH4 conversion and C2 selectivity and yield with the inert gas mole% in the feed in the TZFBR and FBR, U/Umf = 3, CH4/O2 = 7, T = 800 °C, Hb = 10 cm, Hm = 7 cm, and mcat. = 7.5 g.
be effective for preventing the formation of overheated zones, this will not occur in fluidized-bed reactor due to temperature uniformity. This fact has also been verified by Mleczko et al. [6,33]. For OCM in fluidized-bed reactor they found that the feed dilution had almost no effect on C2 selectivity and yield. In case of TZFBR, it is seen that at an 88% value of the inert gas, C2 selectivity and yield are higher than those of FBR, whereas CH4 conversion is almost identical. By decreasing the inert gas, the three above parameters decrease. C2 selectivity may decrease if separation of the oxidation and reduction
Fig.10. Variation of the CH4 conversion and C2 selectivity and yield with the U/Umf in the TZFBR and the FBR, CH4/O2 = 7, T = 800 °C, Hb = 10 cm, % mole inert gas = 88 and mcat. = 7.5 g.
Fig. 12. Variation of the normalized concentration of methane along the bed at U/Umf = 3, CH4/O2 = 7, T = 800 °C, Hb = 10 cm, and mcat. = 7.5 g.
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zones becomes less distinct. On the other hand, due to a narrower reaction zone in TZFBR as compared to FBR, with decreasing inert gas and increasing the ratio of reactants, conversion will also decrease. Variations in the normalized concentration of methane along the reactor length are presented in Fig. 12. As can be seen, the inlet methane concentration attains the maximum value, and inside OCM reaction zone its concentration decreases by being converted to products. On the other hand, due to diffusion of methane around the inlet point, it is seen at a distance of about 1 cm lower than the inlet point, where the concentration is decreased rapidly and diminished. 4. Conclusion In this research, the effects of feed segregation, the so called TZFBR, on the performance of a fluidized-bed reactor for OCM reaction were investigated and compared to the performance of a conventional fluidized-bed reactor. The Mn/Na2WO4/SiO2 catalyst employed exhibited a redox property which is the main requirement for the operation of TZFBR. It was seen that the C2 yield of both reactors was almost equal at CH4/O2 in the range of 5–9. However, TZFBR resulted in a higher C2 selectivity at equal C2 yield. The C2 selectivity in TZFBR was 4–5.5% larger than that in FBR. The reaction of methane with selective lattice/ surface oxygen species may have led to the higher selectivity. The TZFBR results demonstrate that it is possible to have a continuous operation for OCM reaction while avoiding the simultaneous presence of oxygen and hydrocarbons in the gas phase. Acknowledgment The authors appreciate a partial financial support from Iranian National Petrochemical Company under contract No. 84131. References [1] G.E. Keller, M.M. Bhasin, Synthesis of ethylene via oxidative coupling of methane: I. Determination of active catalysts, J. Catal. 73 (1982) 9–19. [2] X. Fang, S. Li, J. Lin, Y. Chu, Oxidative coupling of methane on W–Mn catalysts, J. Mol. Catal. (China) 6 (1992) 427–433. [3] D. Wang, M.P. Rosynek, J.H. Lunsford, Oxidative coupling of methane over oxidesupported sodium–manganese catalysts, J. Catal. 155 (1995) 390–402. [4] S. Pak, J.H. Lunsford, Thermal effects during the oxidative coupling of methane over Mn/Na2WO4/SiO2 and Mn/Na2WO4/MgO catalysts, Appl. Catal., A 168 (1998) 131–137. [5] G. Follmer, L. Lehmann, M. Baerns, The application of laboratory scale catalytic fixed and fluidized bed reactors in the oxidative coupling of methane, ACS Div. Petrol. Chem. 33 (1988) 453–459. [6] L. Mleczko, U. Pannek, V.M. Niemi, J. Hiltunen, Oxidative coupling of methane in a fluidized-bed reactor over a highly active and selective catalyst, Ind. Eng. Chem. Res. 35 (1996) 54–61. [7] A. Santos, J. Santamaria, M. Menendez, Oxidative coupling of methane in a vibrofluidized bed at low fluidizing velocities, Ind. Eng. Chem. Res. 34 (1995) 1581–1587. [8] R. Andorf, L. Mleczco, D. Schweer, M. Baerns, Oxidative coupling of methane in a bubbling fluidized bed reactor, Can. J. Chem. Eng. 69 (1991) 891–897. [9] S. Li, Reaction chemistry of W–Mn/SiO2 catalyst for the oxidative coupling of methane, J. Nat. Gas Chem. 12 (2003) 1–9.
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