Oxidative Coupling of Methane in a Fluidized-Bed Reactor over a

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Ind. Eng. Chem. Res. 1996, 35, 54-61

Oxidative Coupling of Methane in a Fluidized-Bed Reactor over a Highly Active and Selective Catalyst L. Mleczko* and U. Pannek Lehrstuhl fu¨ r Technische Chemie, Ruhr-Universita¨ t Bochum, D-44780 Bochum, Germany

V. M. Niemi† and J. Hiltunen† NESTE OY, Technology Centre, P.O. Box 310, SF-06101 Porvoo, Finland

Oxidative coupling of methane to C2+ hydrocarbons over a Zr/La/Sr catalyst was investigated in an atmospheric-pressure shallow fluidized-bed reactor (i.d. ) 5 cm; Hmf ) 1.4-3.2 cm) at temperatures between 800 and 880 °C. The catalyst was mechanically and catalytically stable, but its fluidizability was limited; agglomeration and channeling occurred. The highest C2+ yield amounted to 18.0% (XCH4 ) 36.5%, SC2+ ) 49.4%) and 17.2% (XCH4 ) 36.6%, SC2+ ) 46.9%) for the diluted (pO2 ) 17 kPa, pCH4 ) 41.5 kPa, pN2 ) 41.5 kPa) and undiluted feed (pO2 ) 28 kPa, pCH4 ) 72 kPa), respectively. These yields are among the highest ones reported in the open literature for OCM in fluidized beds. In the whole investigated temperature range higher selectivities and yields were obtained upon reducing partial pressures of methane and oxygen but keeping their ratio constant (pCH4/pO2 ) 2.5). An increased gas velocity (from u/umf ) 6 to 10) or bed height (from 1.4 to 3.2 cm) resulted in a drop of C2+ selectivity. 1. Introduction The large deposits of natural gas and the growing importance of environmental protection caused substantial research on conversion of natural gas to chemical feedstocks and liquid fuels. In particular, the heterogeneously catalyzed oxidative coupling of methane (OCM) to higher hydrocarbons (C2+) achieved a large interest in the last years. The immense research effort has led to a significant increase of the activity and selectivity of the applied catalysts. However, further progress in C2+ selectivity and yield by catalyst improvement or by reaction engineering means is necessary to make the process commercially viable. The main engineering challenge is to perform the reaction in a safe manner and to cope with the large amount of heat released in this process. From this point of view for a large-scale operation a fluidized bed appears to be the most promising reactor type. The feasibility of a fluidized bed for performing the OCM has been confirmed in experimental investigations in laboratory-scale bubbling beds (for review, see Mleczko and Baerns, 1995). However, many active and selective OCM catalysts were not fluidizable, e.g., Li/MgO (Edwards et al., 1990a; Geerts et al., 1992), NaOH/CaO, Na2SO4/CaO, and Na2CO3/CaO (Andorf and Baerns, 1991; Baerns et al., 1993), or did not exhibit sufficient mechanical strength Sm2O3 (Geerts, 1990). Moreover, since the OCM exhibits a complex reaction network, C2+ selectivity and yield depended on the reaction conditions, particularly, it is influenced by the hydrodynamics of a fluidized-bed reactor. Contradictory findings were reported with respect to this point; for some catalysts comparable or, under certain conditions, even higher C2+ selectivities and yields than in fixed-bed reactors were reported (Edwards et al., 1990a; Follmer, 1988; Mleczko et al., 1991). However, also lower C2+ selectivities and yields were measured (e.g., Mleczko et al., * Author to whom correspondence should be addressed. Fax: +234/709 4115. e-mail: [email protected]. † Fax: +358/155417226.

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1994). This work aimed at determination of the catalytic performance of the Zr/La/Sr catalyst for OCM in a fluidized-bed reactor. This material has been found to be a very active and selective catalyst. When applying this catalyst in a polytropic fixed-bed reactor (i.d. ) 6 mm), a maximum C2+ yield of 21.4% (XCH4 ) 41.8%, SC2+ ) 52.5%) was obtained (pCH4 ) 20 kPa, pO2 ) 9 kPa, pN2 ) 71 kPa, T ) 860 °C, mcat/F ) 0.15 g‚s‚mL-1). However, its performance when fluidized is unknown. Two particular goals were addressed in these investigations: (a) testing the fluidizability and mechanical stability of this catalyst and (b) elucidating the operating conditions maximizing C2+ selectivity and yield. An optimization of reaction conditions is an especially challenging task in a fluidized bed since C2+ selectivity and yields not only depend on temperature and partial pressures of reactants but can also be influenced by gas back-mixing, bed porosity, and mass transport between bubbles and the emulsion phase. 2. Experimental Section 2.1. Catalyst. The Zr/La/Sr catalyst was developed by NESTE OY. Its chemical composition cannot be disclosed in this paper due to proprietary reasons. The catalyst was heat-treated at 900 °C before crashing and sieving it. The particle diameter of the catalyst applied in the experiments ranged from 71 to 160 µm. Catalyst density amounted to 1700 kg/m3. According to Geldart’s classification (Geldart, 1973), it belongs to Group A. The minimum fluidization velocity of 0.018 m/s was experimentally determined at 600 °C for nitrogen as fluidizing gas. 2.2. Apparatus. A laboratory scale (i.d. ) 5 cm) fluidized-bed reactor made of quartz was applied (see Figure 1). The total height of the reactor amounted to 1.35 m. The feed gas was heated in the preheating section (L ) 80 cm, D ) 2 cm) up to 400 °C. A porous quartz plate (dorif ) 40-90 µm) was used as the gas distributor. The bed temperature was controlled by an electric heater placed on the outer side of the reactor wall. In order to reduce particle entrainment, a disengaging section and an internal cyclone were located on © 1996 American Chemical Society

Ind. Eng. Chem. Res., Vol. 35, No. 1, 1996 55

Figure 2. Temperature profiles measured in the fluidized-bed reactor ((9) mcat ) 30 g, mcat/F ) 0.3 g‚s‚mL-1, pCH4 ) 72 kPa, pO2 ) 28 kPa; (2) mcat ) 30 g, mcat/F ) 0.3 g‚s‚mL-1, pCH4 ) 41.5 kPa, pO2 ) 17 kPa, pN2 ) 41.5 kPa; (b) mcat ) 70 g, mcat/F ) 0.7 g‚s‚mL-1, pCH4 ) 72 kPa, pO2 ) 28 kPa).

Figure 1. Laboratory-scale fluidized-bed reactor. Table 1. Experimental Conditions Applied in the Experiments Performed in the Fluidized-Bed Reactor (dR ) 5 cm) pCH4 pO2 pN2 ptot mcat/F T mcat

41.5-72 kPa 17-28 kPa 0-41.5 kPa 100 kPa 0.3-0.7 g‚s‚mL-1 800-880 °C 30-70 g

top of the reaction zone. The postcatalytic section was not heated but isolated in order to avoid condensing of water. Temperature profiles were measured at three axial positions of the fluidized bed by means of thermocouples within a quartz well (o.d. ) 4 mm). The flow rates of the feed gases methane (99.95%), oxygen (99.95%), and nitrogen (99.99%) were stabilized by mass flow controllers. Reactants (CH4, O2, and N2 as diluent) and products (C2H6, C2H4, C3H8, C3H6, H2, CO, and CO2) were analyzed by on-line gas chromatography. H2O was condensed downstream of the reactor. 2.3. Reaction Conditions. The catalyst was tested in the range of experimental conditions listed in Table 1. The height of the settled catalytic bed amounted to 1.4 and 3.2 cm, respectively. The shallow fluidized bed was selected primarily in order to obtain stable fluidization (see below). 2.4. Experimental Procedure. The reactor was heated up in a stream of nitrogen to approximately 650 °C before being switched to the OCM feed. Between experiments the catalyst was kept in a stream of nitrogen at temperatures above 600 °C. The productgas composition was analyzed when the bed temperature was constant within a range of 2 K for approximately 10 min. All data presented in this work are based on an average from three analyses. 3. Results 3.1. Reactor Operation. Chemical Stability. No significant drop in activity or C2+ selectivity was observed after 60 h of operation. However, in order to keep the stability constant, the catalyst had to be retained between the experiments in a stream of nitrogen at temperatures above 600 °C.

Main Products and Accuracy of Analysis. C2H6, C2H4, H2, H2O, CO, and CO2 were detected as the main reaction products. The selectivity to C3+ hydrocarbons was lower than 4%. The inaccuracy of the balance of carbon and oxygen was lower than 2 and 5%, respectively. Fluidizability and Mechanical Stability. Agglomeration of the catalyst was observed at reaction conditions as well as when fluidized with air at room temperature. After defluidization channeling of gas occurred. Fluidization could be kept when operating with a shallow bed (Hmf < 3.5 cm) and high gas velocities (u/umf > 6). No special tests were performed in order to determine the mechanical stability. The loss of catalyst during 60 h of operation due to entrainment and attrition amounted to 2.6 g for an initial bed mass of 30 g. Temperature Profiles. The temperature profiles at reaction conditions are shown in Figure 2. When fluidizing the catalyst in a stream of nitrogen, negligible temperature gradients (∆T < 2 K) were measured. When a catalyst mass of 30 g (Hmf ) 1.4 cm) was applied, a temperature difference of about 10 K was observed within the bed. The temperature was constant within the first centimeter but then decreased. At the first two centimeters above the bed, the temperature dropped about 200 K. When a diluted feed (pCH4 ) 41.5 kPa, pO2 ) 17 kPa, pN2 ) 41.5 kPa) and the bed containing 30 g of the catalyst were applied, the profile within the bed was similar. The temperature drop above the bed, however, was not so fast. Increasing the catalyst mass to 70 g (Hmf ) 3.2 cm) and using undiluted feed, the temperature measured directly above the gas distributor was lower than the setpoint temperature. The bed was almost isothermal within the first two centimeters (0.5 cm < h < 2 cm). In the upper part of the bed (h > 2 cm), the temperature decreased. 3.2. Effect of Temperature and Feed-Gas Composition. This series of experiments aimed at determination of the temperature and feed-gas composition yielding the best reactor performance. The experiments were performed by applying a bed height of 1.4 cm fluidized with a linear gas velocity of u800°C ) 0.18 m/s (u/umf ) 10). Temperatures were varied between 800 and 880 °C. Two different feed-gas compositions were applied: (a) undiluted feed as is expected to be used in the commercial operation and (b) feed gas containing 41.5% N2 as diluent. In both cases, the methane-tooxygen ratio amounted to 2.5: (a) pCH4/pO2 ) 72/28 and (b) pCH4/pO2 ) 41.5/17 (kPa/kPa).

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Ind. Eng. Chem. Res., Vol. 35, No. 1, 1996

Figure 3. Influence of temperature on oxygen conversion at different feed-gas compositions (mcat ) 30 g, dp ) 71-160 µm, mcat/F ) 0.3 g‚s‚mL-1, u/umf )10, (4) pCH4 ) 41.5 kPa, pO2 ) 17 kPa, pN2 ) 41.5 kPa, (O) pCH4 ) 72 kPa, pO2 ) 28 kPa, pN2 ) 0 kPa).

Figure 5. Influence of temperature on C2+ selectivity at different feed-gas compositions (mcat ) 30 g, dp ) 71-160 µm, mcat/F ) 0.3 g‚s‚mL-1, u/umf ) 10, (4) pCH4 ) 41.5 kPa, pO2 ) 17 kPa, pN2 ) 41.5 kPa, (O) pCH4 ) 72 kPa, pO2 ) 28 kPa, pN2 ) 0 kPa).

Figure 4. Influence of temperature on methane conversion at different feed-gas compositions (mcat ) 30 g, dp ) 71-160 µm, mcat/F ) 0.3 g‚s‚mL-1, u/umf ) 10, (4) pCH4 ) 41.5 kPa, pO2 ) 17 kPa, pN2 ) 41.5 kPa, (O) pCH4 ) 72 kPa, pO2 ) 28 kPa, pN2 ) 0 kPa).

Figure 6. Influence of temperature on C2+ yield at different feedgas compositions (mcat ) 30 g, dp ) 71-160 µm, mcat/F ) 0.3 g‚s‚mL-1, u/umf ) 10, (4) pCH4 ) 41.5 kPa, pO2 ) 17 kPa, pN2 ) 41.5 kPa, (O) pCH4 ) 72 kPa, pO2 ) 28 kPa, pN2 ) 0 kPa).

Oxygen and Methane Conversion. Conversion of oxygen increased with temperature for both feed-gas compositions (see Figure 3). The largest gradients were observed between 800 and 840 °C; conversion increased from 83 to 94% and from 89 to 99% for the diluted and undiluted feed, respectively. For all investigated temperatures, oxygen conversions were approximately 5% higher for the undiluted feed. The conversion of methane also increased with temperature for both investigated cases (see Figure 4). A strong increase was only observed for temperatures up to 840 °C. In this temperature range, slightly higher methane conversions were found when applying undiluted feed. Above 840 °C, the conversions do not differ significantly. Using feed-gas containing nitrogen methane, conversion rose from 30% at 800 °C to 37% at 880 °C. When applying undiluted feed, values between 32% (800 °C) and 37% (880 °C) were obtained.

C2+ Selectivity and Yield. C2+ selectivity increased with temperature for both investigated cases (see Figure 5). When applying diluted feed, SC2+ rose from 41% at 800 °C to 49% at 880 °C. For undiluted feed it increased from 39% to 47%. The difference between the two curves amounted to approximately 2% and remained almost constant in the whole temperature range. In both cases only a slight increase of C2+ selectivity was measured above 860 °C; the C2+ selectivity runs into a plateau. For both feed-gas compositions, an increase of C2+ yield with temperature was observed (see Figure 6). In the whole temperature range higher yields were obtained for the diluted feed. C2+ yield increased from 12.4% at 800 °C to 18% at 880 °C. Using undiluted feed, yields between 12.2% (800 °C) and 17.2% (880 °C) were achieved. 3.3. Influence of the Gas Velocity. The effect of the gas velocity on the catalytic performance was investigated for velocities of u800°C ) 0.12 m/s and u800°C ) 0.18 m/s. The respective fluidization numbers and


Figure 7. Influence of temperature on oxygen conversion determined for different gas velocities (u/umf ) 6/10, mcat ) 30 g, dp ) 71-160 µm, mcat/F ) 0.3/0.45 g‚s‚mL-1, pCH4 ) 72 kPa, pO2 ) 28 kPa).

Figure 8. Influence of temperature on methane conversion determined for different gas velocities (u/umf ) 6/10, mcat ) 30 g, dp ) 71-160 µm, mcat/F ) 0.3/0.45 g‚s‚mL-1, pCH4 ) 72 kPa, pO2 ) 28 kPa).

contact times varied between u/umf ) 6-10 and mcat/F ) 0.3-0.45 g‚s‚mL-1. The undiluted feed (pCH4 ) 72 kPa, pO2 ) 28 kPa, pCH4/pO2 ) 2.5) and a catalyst mass of mcat ) 30 g (Hmf ) 1.4 cm) were applied. Oxygen and Methane Conversion. Upon decreasing the gas velocity, higher oxygen conversions were obtained (see Figure 7). Applying u ) 0.18 m/s, values between 89% (800 °C) and 99% (880 °C) were obtained. Using the lower gas velocity (u ) 0.12 m/s), oxygen conversion increased from 95% at 800 °C to 99.5% at 880 °C. However, at temperatures above 840 °C the differences are negligible. The influence of temperature on methane conversion for both investigated gas velocities is presented in Figure 8. For the low gas velocity the conversion increased from 32% at 800 °C to 35% at 820 °C. At higher temperatures, XCH4 remained almost constant. Upon applying the high gas velocity, a continuous rise of methane conversion from 32% at 800 °C to 37% at 880 °C was observed. At temperatures below 840 °C higher conversions were achieved at low

Figure 9. Influence of temperature on C2+ selectivity determined for different gas velocities (u/umf ) 6/10, mcat ) 30 g, dp ) 71-160 µm, mcat/F ) 0.3/0.45 g‚s‚mL-1, pCH4 ) 72 kPa, pO2 ) 28 kPa).

Figure 10. Influence of temperature on C2+ yield determined for different gas velocities (u/umf ) 6/10, mcat ) 30 g, dp ) 71-160 µm, mcat/F ) 0.3/0.45 g‚s‚mL-1, pCH4 ) 72 kPa, pO2 ) 28 kPa).

gas velocities, whereby at elevated temperatures the higher gas velocity promoted methane conversion. C2+ Selectivity and Yield. In the whole temperature range higher C2+ selectivities were obtained when applying the lower gas velocity (see Figure 9). However, the qualitative dependence of C2+ selectivity is different for the two investigated cases; for the low gas velocity the selectivity strongly increased up to 820 °C. At higher temperatures C2+ selectivity passed through a wide maximum of approximately 47% around the temperature of 860 °C. Applying the higher gas velocity, a continuous increase of the selectivity with temperature was observed; at 880 °C it approached the value measured for the lower gas velocity. The influence of temperature and gas velocity on C2+ yield is shown in Figure 10. Upon applying the lower gas velocity, a wide maximum which amounted to 16.8% was detected at 860 °C. For temperatures below 840 °C higher yields were obtained using the lower gas velocity. Above this temperature, the superior yield was obtained when

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Figure 11. Influence of temperature on oxygen conversion determined for different bed heights (mcat ) 30/70 g, dp ) 71-160 µm, mcat/F ) 0.3/0.7 g‚s‚mL-1, u/umf ) 10, pCH4 ) 72 kPa, pO2 ) 28 kPa).

using the higher gas velocity. The maximum yield of 17.2% was obtained at 880 °C for the higher gas velocity. 3.4. Influence of the Bed Height. The effect of the bed height on the catalytic performance was investigated by applying catalyst masses of 30 and 70 g. The resulting bed heights amounted to 1.4 and 3.2 cm, respectively. An undiluted feed (pCH4 ) 72 kPa, pO2 ) 28 kPa, pCH4/pO2 ) 2.5) was applied. Other conditions were the same as in the previous series of experiments. Oxygen and Methane Conversion. The influence of the bed height on oxygen conversion is presented in Figure 11. Higher values were achieved upon applying higher beds. However, the difference decreased with temperature; at 860 and 880 °C approximately the same conversions were measured. For a catalyst mass of 30 g (Hmf ) 1.4 cm) oxygen conversion increased from 89% at 800 °C to 99% at 880 °C. Applying a catalyst mass of 70 g (Hmf ) 3.2 cm) between 800 and 880 °C, an increase of conversion from 94% to 99% was observed. Methane conversion increased with temperature for both investigated cases in the range 800-860 °C (see Figure 12). Above 860 °C no increase was detected for the lower bed and only a slight growth of methane conversion was measured for the higher bed. The conversions varied between 32% and 37% for 30 g of catalyst and 32% and 36% for 70 g. At 800 °C, a higher methane conversion was achieved for a catalyst mass of 70 g. Above 800 °C, slightly higher conversions were achieved when applying the lower bed. C2+ Selectivity and Yield. For both bed heights, C2+ selectivity increased with temperature in the whole investigated range (see Figure 13). The lower bed resulted always in higher C2+ selectivities; SC2+ increased from 39% at 800 °C to 47% at 880 °C. Upon applying the higher bed (Hmf ) 3.2 cm), a rise from 37% to 46% was measured. The difference between the selectivities obtained for two bed heights was almost independent of the temperature and amounted to less than 3%. An increase of the C2+ yield with temperature was observed for both bed heights (see Figure 14). Applying the lower bed (Hmf ) 1.4 cm), higher yields were obtained in the whole temperature range. The highest yield amounted to 17.2% at 880 °C and 16.4% at 880 °C for Hmf ) 1.4 and 3.2 cm, respectively. The

Figure 12. Influence of temperature on C2+ selectivity determined for different bed heights (mcat ) 30/70 g, dp ) 71-160 µm, mcat/F ) 0.3/0.7 g‚s‚mL-1, u/umf ) 10, pCH4 ) 72 kPa, pO2 ) 28 kPa).

Figure 13. Influence of temperature on methane conversion determined for different bed heights (mcat ) 30/70 g, dp ) 71-160 µm, mcat/F ) 0.3/0.7 g‚s‚mL-1, u/umf ) 10, pCH4 ) 72 kPa, pO2 ) 28 kPa).

difference between the two curves grew from 0.2% at 800 °C to 0.8% at 880 °C. 4. Discussion 4.1. Reactor Operation. Similar to other OCM catalysts (Andorf et al., 1991; Edwards et al., 1992; Mleczko et al., 1992; Santos et al., 1994) in the whole range of methane conversions (XCH4 < 35%) and upon applying an undiluted feed, a good temperature control was realized. Almost isothermal operation was achieved even in shallow beds (Hmf ) 1.4 cm). For comparison, when performing the OCM reaction over this catalyst in a polytropic fixed bed (i.d. ) 6 mm), temperature gradients of more than 100 K were observed (mcat/F ) 0.3 g‚s‚mL-1, pCH4 ) 41.5 kPa, pO2 ) 17 kPa, pN2 ) 41.5 kPa). Furthermore, temperature spikes made impossible operation with the undiluted feed (CH4/O2 ) 2.5). The catalyst was mechanically stable, however, a further increase of mechanical strength would be necessary


Figure 14. Influence of temperature on C2+ yield determined for different bed heights (mcat ) 30/70 g, dp ) 71-160 µm, mcat/F ) 0.3/0.7 g‚s‚mL-1, u/umf ) 10, pCH4 ) 72 kPa, pO2 ) 28 kPa).

in order to reduce the high elutration rate. Similar to many other OCM catalysts (see Edwards et al., 1990a; Andorf and Baerns, 1991; Mleczko et al., 1994), also the Zr/La/Sr catalyst exhibited only limited fluidizability. In order to avoid this obstacle, a shallow bed had to be used. The hydrodynamic conditions realized in the investigated reactor are characterized by short contact times (0.1-0.3 s) and the low back-mixing of gas. In an industrial scale these conditions can be approached in a turbulent bed. 4.2. Effect of Temperature and Feed-Gas Composition. The experiments confirmed the high activity and selectivity of the Zr/La/Sr catalyst. Temperature profiles indicated that oxygen was almost completely converted in the first few millimeters after entering the bed. In contrast to the fixed-bed reactor, no dilution was necessary in order to manage the heat released by the OCM reaction. A C2+ yield of 17.2% obtained for the undiluted feed is among the highest one reported for fluidized-bed reactors in the open literature. However, there is a group of materials exhibiting similar catalytic performance, e.g., for the “CSIRO-catalyst” and for the La2O3/CaO catalysts the highest yields amounted to YC2+ ) 19% (Do et al., 1995) and YC2+ ) 16% (Mleczko et al., 1996), respectively. Improvement of C2+ selectivity and yield by diluting the feed in a fluidized-bed reactor was also reported for the PbO/γ-Al2O3 catalyst (Andorf, 1992). This effect can be explained by the nonlinear dependence of the reaction rates of the selective and nonselective primary steps on the partial pressures of oxygen and methane. However, the increase of C2+ selectivity and yield when diluting the feed is not generally valid; e.g., for La2O3/CaO no effect of the feed dilution on C2+ selectivity was observed when carrying out the reaction in a fluidized bed (Mleczko et al., 1996). The increase of C2+ selectivity with temperature can be explained by a higher activation energy of the selective reaction step compared to the nonselective steps (for the proposed reaction scheme, see Mleczko and Baerns, 1995). This kinetic effect was observed for the majority of OCM catalysts (Mleczko and Baerns, 1995). However, the increase of C2+ selectivity at temperatures above 860 °C has not been observed in fluidized-bed reactors until now. For other OCM catalysts investi-

gated in this reactor type the selectivity passed through a maximum at lower temperatures. The temperature of the maximum selectivity depended on the catalyst and on the reaction conditions. Typically this temperature varied between 740 and 840 °C for PbO/γ-Al2O3 (Andorf et al., 1991), from 800 to 840 °C for La2O3/CaO (Mleczko et al., 1992), and from 830 to 850 °C for the CSIRO catalyst (Edwards et al., 1992). For different catalysts containing Na, e.g., NaOH/CaO and Na2SO4/ CaO, the maximum which was located between 710 and 820 °C or a plateau in this temperature range was achieved (Andorf and Baerns, 1991; Andorf et al., 1991). The difference between the above-discussed characteristics for the Zr/La/Sr and other catalysts can be explained by both different reaction pathways and different hydrodynamic conditions. Concerning the reaction pathway, it can be postulated that this catalyst is less active for the consecutive reactions, e.g., compared to La2O3/CaO which exhibits similar activity for the primary steps (Mleczko et al., 1992; Schweer et al., 1994). This is an important feature of the catalyst when considering application of the distributed feed of oxygen as a mean for improving C2+ selectivity (Baerns and Hinsen, 1984). The different extents of the consecutive reactions can be, however, also caused by a lower backmixing in the shallow bed compared to the bubbling beds applied in previously cited studies. The effect of temperature and feed-gas composition was similar to the one determined in the fixed-bed reactor; i.e., for both reactor types in the temperature range up to 880 °C, conversion of methane, C2+ selectivity, and yield rose with temperature and when diluting the feed. However, when applying the same partial pressures of reactants and comparing the performance of both reactor types at the same characteristic temperature, i.e., the hot-spot in the fixed bed and the bed temperature in the fluidized-bed reactor, lower C2+ selectivities and yields were measured in the fluidized bed. For example, for pCH4 ) 41.5 kPa, pO2 ) 17 kPa, pN2 ) 41.5 kPa, and T ) 880 °C the maximum yield in the fixed bed amounted to 19.7% (XCH4 ) 37.9%, SC2+ ) 52.0%) compared to YC2+ ) 18.0% (XCH4 ) 36.5%, SC2+ ) 49.4%) in the fluidized bed. The drop of C2+ selectivity and yield compared to the fixed-bed reactor is due to the mass transport limitation between the dilute (bubbles) and dense (emulsion) phases, which is not negligible even when applying a bed height comparable with the height of the bubble formation zone (Hmf ) 1.4 cm). Furthermore, C2+ selectivity and yield are influenced to a large extent by gas back-mixing. Both effects, i.e., mass transport limitation and gas back-mixing, are detrimental for a high selectivity to an intermediate product. This is confirmed by the higher hydrogen selectivities and CO-to-COx ratios in the fluidized bed; e.g., at the above conditions a hydrogen selectivity of 15% and a CO-to-COx ratio of 0.23 was achieved compared to 5% and 0.16 in the fixed bed. However, also the effect of the nonselective gas phase reactions promoted by a higher average porosity compared to a fixed bed cannot be excluded. 4.3. Influence of the Gas Velocity. The effect of the gas velocity on C2+ selectivity and yield is especially important when considering scale-up. In order to increase the space-time yield, industrial fluidized-bed reactors are mostly operated at much higher gas velocities compared to laboratory reactors. Moreover, in case of the La2O3/CaO (Mleczko et al., 1996) and PbO/γ-Al2O3 (Andorf, 1992) catalysts, an increased gas velocity

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resulted in an improved C2+ selectivity. This effect was not reproduced when applying the Zr/La/Sr catalyst; the higher gas velocity caused a drop of the C2+ selectivity. However, due to the increased methane conversion, higher yields were achieved. The improvement of methane conversion in spite of the lower selectivity was possible due to the shift in production of nonselective products from CO2 and H2O to CO and H2. In order to explain the above effects, the hydrodynamics of the investigated reactor has to be taken into account. According to Kwauk (1992), four zones can be recognized in a shallow fluidized bed: gas cushion, dense bed, cloud, and fog. The gas cushion is formed from jets which at high gas velocities merge directly above the distributor. This phase has a very sparse population of particles. It is a matter of discussion if this phase is formed in a fluidized bed with a porous plate distributor as in the applied reactor. The dense phase which is located above the gas cushion contains highly expanded bed. Although Kwauk reports that in shallow beds bubbles in their usual forms were not identifiable, the visual observation indicates that besides the dense phase also the gas flow in a dilute phase (gas snakes, plumes) occurred. The dense-phase region corresponds to the bubbles formation zone as proposed by Yates (1983) in a bubbling-bed reactor. The dense zone is topped by a cloud layer with diminishing solids concentration until it lapses into a foggy formation. Upon applying the model of the shallow-bed hydrodynamics, the most probable explanation for the observed effects is as follows: the elevated gas velocity caused an increase of the volume of the gas cushion and in porosity in the dense part of the bed. This, in turn, resulted in a higher concentration of oxygen promoting nonselective reaction steps and nonselective gas-phase reactions. Since the main products of these reactions are CO and H2, the measured increase of CO-to-COx ratios and selectivity to hydrogen with rising gas velocity confirm the above explanation. The same opinion that high bed porosity might be detrimental for C2+ selectivity was presented for the PbO/γ-Al2O3 catalyst (Tjatjopoulos and Vasalos, 1992). 4.4. Influence of the Bed Height. The application of higher beds resulted in lower C2+ selectivities and yields; it should be noted that for temperatures above 820 °C no significant difference in oxygen conversion was observed. This effect can be explained either by different hydrodynamic conditions in the oxygen-rich zone or by consecutive reactions in the oxygen-free zone. However, the decrease of methane conversion observed at temperatures above 820 °C indicates that the drop of C2+ selectivity was caused mainly by nonselective oxidation reactions. The different hydrodynamic conditions and their effect on the catalytic performance can be explained by applying the previously introduced model of a shallow fluidized bed and the model of a bubbling bed. According to the model of a shallow-bed reactor, plug-flow of gas can be expected in experiments performed for Hmf ) 1.4 cm. The increased static bed height results mainly in a larger dense-bed zone. According to Yates (1983), the formation of bubbles is completed at the distance from the gas distributor which amounts to two initial bubble diameters; for the Zr/La/ Sr catalyst the initial bubble diameter predicted by the correlation of Mori and Wen (1975) varies at a temperature of 800 °C between 0.5 and 1.0 cm for the lowest and highest gas velocities, respectively. Therefore, for

the higher bed (Hmf ) 3.2 cm) fluidization in the bubbling regime can be expected. Since in a bubbling bed back-mixing of gas is much more intensive compared to the shallow bed (Kwauk, 1992), ethylene and ethane can be transported from the upper part of the bed into the oxygen-rich distributor zone where they are combusted. Lower C2+ selectivity by the similar distribution of nonselective products caused lower conversions of methane due to stoichiometric reasons. The increased CO-to-COx ratios measured when applying the higher bed indicate that the oxidation of back-mixed C2+ hydrocarbons can also take place in the gas phase, since CO is the preferred product of gas-phase oxidation reactions. Moreover, longer contact times, especially in the emulsion phase of the bubbling bed, promote the dehydrogenation of ethane to ethylene. This explains the higher ethylene-to-ethane ratios and higher selectivities to hydrogen obtained for the higher bed. Also steam reforming of C2+ hydrocarbons which was identified to take place at high temperatures and long contact times, e.g., when applying La2O3/CaO (Stansch, 1995, Mleczko et al., 1992), cannot be excluded as a reason for the lower C2+ selectivities obtained upon applying the higher bed. The results obtained in the laboratoryscale reactor indicate that a bubbling bed would not be a reactor of choice for scale-up application. A large diameter of an industrial reactor promotes circulation of solids and, in turn, back-mixing of gas. Furthermore, significantly higher beds compared to the investigated reactor will be necessary for two reasons: the application of an industrial gas distributor will result in lower mass-transfer coefficients and a high bed will be necessary in order to achieve total conversion of oxygen and for temperature control by an immersed heat exchanger. Calculations performed for the La2O3/CaO catalyst which exhibits activity similar to that of the catalyst investigated in this work indicated that the application of the industrial distributor resulted in the increase of the bed height necessary in order to achieve total conversion of oxygen from about 3 cm to 1 m (Pannek and Mleczko, 1995). Such a high bed can be detrimental for C2+ selectivity due to the long contact times in the emulsion phase which, in turn, promote steam reforming of C2+ hydrocarbons. The influence of the contact time in fluidized-bed reactors was also studied for other catalyst systems, e.g., the CSIRO-catalyst (Edwards et al., 1990b) and the La2O3/CaO catalyst (Mleczko et al., 1994). Dependencies similar to those for the Zr/La/Sr catalyst were observed: a decrease in contact time resulted in an increase of the C2+ selectivity and yield. Based on above results, a common conclusion can be drawn that for the active OCM catalysts like Zr/La/Sr, “CSIRO”, and La2O3/CaO a short-contact-time reactor like turbulent bed, riser, or spouted bed will be a preferred choice for the scale-up application. However, as previously indicated for the Zr/La/Sr catalyst (see above) and also for other OCM catalysts (Tjatjopoulos and Vasalos, 1992), the catalytic performance of these reactors might be influenced by gas-phase reactions which in these reactor types are promoted by a higher porosity of the bed. 5. Conclusions The oxidative coupling of methane over the Zr/La/Sr catalyst was studied in a laboratory-scale fluidized-bed reactor. The catalyst was catalytically stable; however, it was fluidizable only in a limited range of reaction conditions. Stable fluidization was achieved upon ap-


plying a shallow bed (Hmf < 3.5 cm) and high gas velocities (u/umf > 6). When performing the OCM in the shallow fluidized bed, isothermal conditions were achieved even at high methane conversions (XCH4 < 35%, pCH4 ) 72 kPa). The highest C2+ yield obtained in the shallow fluidized-bed reactor amounted to 18.0% (pCH4 ) 41.5 kPa, pO2 ) 17 kPa, pN2 ) 41.5 kPa, T ) 880 °C, mcat/F ) 0.3 g‚s‚mL-1, XCH4 ) 36.5%, SC2+ ) 49%). When applying undiluted feed and keeping the methane-to-oxygen ratio constant (pCH4/pO2 ) 72 kPa/ 28 kPa ) 2.5, T ) 880 °C, mcat/F ) 0.3 g‚s‚mL-1), lower selectivities and yields were obtained (YC2+ ) 17.2%, XCH4 ) 36.6%, SC2+ ) 46.9%) compared to the undiluted feed. The highest C2+ selectivity and yield were measured at the highest temperature investigated; no maximum in dependence of C2+ selectivity and yield on temperature was observed. C2+ selectivity depended on the hydrodynamic conditions. With increasing bed height (from 1.4 to 3.2 cm) C2+ selectivity and yield dropped due to the growing back-mixing of gas. An increased gas velocity (from u/umf ) 6 to u/umf ) 10) led to lower C2+ selectivities, indicating that a high bed porosity might be detrimental for C2+ selectivity. C2+ yields obtained in this study are among the highest ones reported in the open literature for OCM in a fluidized-bed reactor. Although the application of a shallow bed reactor minimized the back-mixing of gas, lower C2+ selectivities were obtained compared to the fixed-bed reactor. Acknowledgment The authors thank Prof. Dr. M. Baerns for encouraging the work and critical discussions during performance of this project. Thanks are also due to Mr. Guido Gayko for performing the cited experiments in the fixedbed reactor. Nomenclature dp ) particle diameter, µm FSTP ) volumetric flow, mL/s Hmf ) bed height at minimum fluidization conditions, cm pi ) partial pressure of component i, kPa Si ) selectivity to component i,% T ) reaction temperature, °C u ) gas velocity, m/s umf ) minimum fluidization velocity, m/s mcat ) catalyst mass, g Xi ) conversion of component i, % YC2+ ) yield of higher hydrocarbons, %

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Received for review February 27, 1995 Revised manuscript received August 2, 1995 Accepted October 4, 1995X IE950145S

X Abstract published in Advance ACS Abstracts, December 1, 1995.

Oxidative Coupling of Methane in a Fluidized-Bed Reactor over a

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