Short-Contact-Time Reactor for Catalytic Partial Oxidation of Methane

contact times in the draft tube (, 0.1 g‚s‚mL-1) were sufficient to obtain nearly ... of an ICFB reactor for the partial oxidation of methane to s...
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Ind. Eng. Chem. Res. 1999, 38, 1813-1821

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Short-Contact-Time Reactor for Catalytic Partial Oxidation of Methane K.-J. Marschall and L. Mleczko* Lehrstuhl fu¨ r Technische Chemie, Ruhr-Universita¨ t Bochum, D-44780 Bochum, Germany

An internally circulating fluidized-bed (ICFB) reactor (i.d.eff ) 1.9 cm, Hriser ) 10 and 20 cm, respectively) was applied to investigate the catalytic partial oxidation of methane to synthesis gas over Ni/R-Al2O3 (Ni loading, 1 and 5 wt %; particle diameter, 71-160 and 250-355 µm). The experiments were performed at 800 °C by applying a methane to oxygen ratio of 2:1. The contact times in the draft tube (, 0.1 g‚s‚mL-1) were sufficient to obtain nearly thermodynamic equilibrium values for the methane conversion (XCH4,eq ) 92%) and selectivities to hydrogen (SH2,eq ) 97%) and carbon dioxide (SCO,eq ) 97%). The achieved results indicate the general suitability of an ICFB reactor for the partial oxidation of methane to synthesis gas. However, stable and controllable reactor operation with sufficient solids circulation through the draft tube was possible only over a narrow range of gas velocities (0.75-1.50 m‚s-1). Furthermore, the conversion and selectivities were strongly influenced by the temperature distribution in the ICFB reactor. Especially, the decrease of temperature in the top of the draft tube and in the fountain region promoted backreactions reducing the conversion of methane. Furthermore, catalyst deactivation due to carbon depositions occurred in the ICFB reactor. This effect caused for the 1 wt % Ni catalyst a significant drop of the activity during 150 h of time on stream. The methane conversion decreased in the range of 12% for a catalyst particle fraction of 71-160 µm. It was shown by TEM that two different types of carbon (encapsulating and whisker carbon) were deposited on the surface. The first species resulted in a reversible deactivation (activation of the catalyst by carbon dioxide in the reverse Boudouard reaction was partially possible), whereas the second species resulted in an irreversible deactivation due to a Ni loss by removing the carbon fibers which had Ni sites on top of the filament. 1. Introduction Synthesis gas (syngas), a mixture of hydrogen and carbon monoxide with varying composition, is an important chemical feedstock. Syngas is used for the production of methanol or higher hydrocarbons (Foulds and Lapszewicz, 1994). The produced hydrogen is further utilized for the ammonia synthesis. Nowadays, syngas is mainly produced by steam reforming of methane or naphtha. However, because of its high endothermicity, this process is characterized by high investment costs and energy consumption. Therefore, many alternative routes for syngas production have been proposed (e.g., Rostrup-Nielsen et al., 1994). Among the different alternatives, the catalytic partial oxidation of methane (eq 1.1) seems to be very promising.

CH4 + 0.5O2 f CO + 2H2 ∆rH1000K ) -21.8 kJ‚mol-1 (1.1) Compared to steam reforming, the main advantage of this reaction is the slight exothermicity. This enables autothermal reactor operation and lowers investment costs and emissions of CO2 and NOx. Finally, the produced syngas is characterized by a hydrogen to carbon monoxide ratio of 2:1 that is desired for methanol production and in Fischer-Tropsch synthesis. The current research in the field of partial oxidation of methane to syngas concentrates on catalyst develop* Corresponding author. Present address: Bayer AG, ZT-TE 4.4, D-51368 Leverkusen, Germany. Email: [email protected].

ment and reaction engineering aspects (e.g., Vernon et al., 1990; Dissanayake et al., 1991; Choudary et al., 1993; Goula et al., 1996; Nichio et al., 1996; Baerns et al., 1997). With respect to reaction engineering, the main problems are connected to temperature control and catalyst deactivation. In fixed-bed reactors steep temperature gradients were observed (Vermeiren et al., 1992; Heitnes et al., 1995). High-temperature spikes can be explained by the two-stage mechanism of the catalytic partial oxidation of methane. The very exothermal deep oxidation in the primary step (eq 1.2) is followed by the endothermal steam (eq 1.3) and carbon dioxide reforming (eq 1.4). The hydrogen to carbon dioxide ratio

CH4 + 2O2 f CO2 + 2H2O ∆rH1000K ) -801.2 kJ‚mol-1 (1.2) CH4 + CO2 h 2CO + 2H2 ∆rH1000K ) 259.7 kJ‚mol-1 (1.3) CH4 + H2O h CO + 3H2 ∆rH1000K ) 224.9 kJ‚mol-1 (1.4) is finally influenced by the water-gas shift reaction. Furthermore, in the second stage the deposition of coke due to the methane pyrolysis can occur. This coke cannot be removed in a fixed-bed reactor and may cause catalyst deactivation. To overcome these obstacles, the application of bubbling fluidized-bed reactors has been proposed (e.g., Goetsch et al., 1989; Olsbye et al., 1993; Mleczko and

10.1021/ie9806646 CCC: $18.00 © 1999 American Chemical Society Published on Web 04/03/1999

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Wurzel, 1997). Because of intensive solid mixing, isothermal reactor operation is achieved by coupling the exothermal oxidation with the endothermal reforming zone. Furthermore, catalyst deactivation due to carbon deposition is avoided. Catalyst particles from the reforming zone are transported to the oxygen-rich distribution zone where carbon depositions are combusted. This reactor concept has successfully been tested in a laboratory-scale by several authors (Slagtern et al., 1994; Bharadwaj and Schmidt, 1994; Santos et al., 1994, 1996; Mleczko and Wurzel, 1997). In all studies isothermal operation and syngas yields near the thermodynamic equilibrium were obtained. However, because of the fast reaction, the thermodynamic equilibrium was achieved in the first few centimeters above the gas distributor. This is also in line with results reported for microcatalytic fixed-bed and monolithic reactors (Bharadwaj and Schmidt, 1994) using residence times as short as 10-5 s. However, for such fast reactions it is expected that the rate of chemical reactions will be influenced by mass- and heat-transfer limitations, especially for an industrial scale. Indeed, simulations of an industrialscale bubbling fluidized-bed reactor showed a strong decrease of the space-time yield in the industrial reactor compared to the laboratory unit. For example, using a Ni/R-Al2O3 catalyst in a laboratory-scale bubbling fluidized-bed reactor (i.d. ) 5 cm, mcat ) 0.1 kg), a maximum space-time yield of 22.0 molsyngas/gNi‚h was reached. In an industrial bubbling fluidized-bed reactor (i.d. ) 4 m, mcat ) 20 tons), the space-time yield dropped to about 0.45 molsyngas/gNi‚h (Wurzel and Mleczko, 1997). The effect of mass transport limitations between the bubble and emulsion phase in a fluidized bed can be minimized by operating the reactor in the high-velocity fluidization regime. Bharadwaj and Schmidt (1994) used a fluidized bed operating close to the turbulent regime. They achieved equilibrium conversion in a bed with a height of 2 cm fluidized with gas velocities of about 0.16 m‚s-1. The possibility of performing the catalytic partial oxidation of methane in an industrial-scale riser reactor was analyzed by means of simulations by Pugsley and Malcus (1997). They showed that equilibrium conversions of methane can be obtained in a riser with 20 m height operated with gas velocities of 6-12 m‚s-1. However, the validity of their predictions is limited because Pugsley and Malcus assumed that the catalyst is in the whole riser reactor in the same oxidation state. Furthermore, the effect of catalyst deactivation was not taken into account. Against this background the catalytic partial oxidation of methane to syngas should be performed at short contact times such as in fixed-bed reactors, ensuring isothermal operation like in bubbling fluidized-bed reactors. Therefore, the use of a short-contact-time reactor like the ICFB is proposed in this work. The ICFB reactor combines the advantages of the fixed bed and the bubbling fluidized bed for the title reaction. In this reactor high gas throughputs resulting in short contact times such as those in monoliths are obtained. The circulation of solids is utilized to handle the reaction heat (Berruti et al., 1993; Mleczko and Marschall, 1997). By this means nearly isothermal operation could be reached by coupling the exothermic total combustion with the endothermic reforming reactions. Furthermore, in contrast to bubbling fluidized-bed and circulating fluidized-bed reactors, the results obtained in a laboratory-scale ICFB reactor can be reproduced in industrial-

Figure 1. Schematic representation of the laboratory-scale ICFB reactor.

scale reactors by applying the direct scale-up concept of multitubular reactors (Mleczko and Marschall, 1997). In this paper the general suitability of an ICFB reactor for the catalytic partial oxidation of methane to syngas will be examined. Hydrodynamic conditions which allowed a stable and controllable reactor operation will be determined. As catalyst, 1 and 5 wt % Ni supported by R-Al2O3 were selected, which were used in former studies in bubbling fluidized-bed reactors (e.g., Mleczko and Wurzel, 1997). For this catalyst the influence of the hydrodynamics on the activity and stability was examined. 2. Experimental Setup 2.1. Internally Circulating Fluidized-Bed Reactor. A detailed description of the ICFB reactor (see Figure 1) has been given elsewhere (Mleczko and Marschall, 1997). Therefore, only the main characteristics of this reactor will be described here. The ICFB reactor consisted of a conically shaped vessel made of quartz with a maximum diameter of 5 cm. The feed gas entered the reactor through a porous quartz plate (pore size 40-90 µm). In the reactor axis a tube (i.d.eff ) 1.9 cm, Hriser ) 10 or 20 cm) was placed. Two different kinds of draft tubes were applied in the ICFB reactor. For preliminary studies of the long-time stability of the catalyst, a draft tube with a baffle above the top was used. However, for further experiments draft tubes without baffles were constructed to reduce the pressure drop (Mleczko and Marschall, 1997). As a consequence less gas bypassing through the annulus was obtained, accepting the higher entrainment of fine particles. A constant length of the catalyst entrance window (entrainment region) was used during the investigations (h ) 8 mm). An internal cyclone was placed in the disengaging section to reduce the entrainment of solids. The temperature in the annulus was controlled by means of outer electric heaters. The height of the heated section corresponded to the length of the draft tube. This arrangement was used in order to minimize the temperature gradient in the tube. The temperature distribution in the draft tube was measured by two thermocouples placed in a quartz well in the axis of the riser. Furthermore, the temperature was measured in the annulus and in the disengaging section.

Ind. Eng. Chem. Res., Vol. 38, No. 5, 1999 1815 Table 1. Contact and Residence Times in the Draft Tube of the ICFB Reactor 10 cm draft tube

20 cm draft tube

contact times contact times mcat/V˙ (kg‚s‚m-3) mcat/V˙ (kg‚s‚m-3) ugas800°C residence residence (m‚s-1)  ) 0.85  ) 0.90 time (s)  ) 0.85  ) 0.90 time (s) 0.83 0.95 1.07 1.35

47 41 36 29

31 27 24 19

0.12 0.11 0.09 0.07

94 82 73 58

62 54 48 38

0.24 0.22 0.18 0.14

The ICFB reactor was integrated in a fully automatized apparatus. Details of the equipment have already been given by Mleczko et al. (1990). Gas flow rates were adjusted by means of mass flow controllers. Reactants (CH4 and O2) and products (H2, CO, and CO2) were analyzed by on-line gas chromatography (GC). All presented data points in the paper represent an average value of at least three GC analyses. The yield of water was not determined experimentally but calculated from the hydrogen balance. In all measurements, the carbon and oxygen balances were kept within an error of (2% and (5%, respectively. A PC-based system was used for data acquisition and control functions. For reference measurements and catalyst treatment, a laboratoryscale bubbling fluidized-bed reactor (i.d. ) 5 cm) was used. For details see, e.g., Mleczko and Wurzel (1997). 2.2. Catalyst. The experiments were carried out by applying a Ni/R-Al2O3 catalyst with Ni loadings of 1 and 5 wt %, respectively. The catalyst was produced by in situ reduction with a hydrogen/nitrogen mixture (1:1) of the NiCO3/R-Al2O3 precursor. The precursor was treated for 7 h with this gas mixture, increasing the temperature in several steps up to 400 °C. By this means the carbonate species was destroyed. The obtained nickel oxide was further reduced to elementary nickel by hydrogen. Two particle fractions were prepared (71 < dp < 160 µm and 250 < dp < 355 µm). 2.3. Experimental Conditions. All experiments were performed at 800 °C by applying a methane to oxygen ratio of 2:1. A total of 150 g of catalyst of the desired particle size distribution was used in the ICFB reactor, resulting in a static bed height of 8 cm. A stable and controllable circulation of solids with minimal gas bypassing was obtained by applying draft tubes of 10 and 20 cm length, respectively. Gas velocities from 0.75 to 1.47 m‚s-1 (800 °C) were applied in the experiments, resulting in residence and contact times given in Table 1. For the estimation of the contact times, porosities in the draft tube were assumed (Mleczko and Marschall, 1997; Marschall and Mleczko, 1999). 3. Results and Discussion 3.1. Fluidizability and Hydrodynamic Properties. These experiments were performed at room temperature and aimed at the determination of hydrodynamic conditions for a stable and controllable reactor operation. The hydrodynamic conditions were strongly influenced by the gas velocity, and stable operation was obtained only over a narrow range of this parameter (0.75-1.50 m‚s-1). Different hydrodynamic regimes were observed visually. No circulation of solids through the draft tube but only slugging was observed for low gas velocities (1.5 m‚s-1) resulted in a fluidization of the annular bed. Similar hydrodynamic regimes were observed also in a former study (Mleczko and Marschall,

1997). Experiments performed at higher temperatures (max 650 °C) confirmed that the same hydrodynamic regimes occurred under these conditions. Furthermore, the strong decrease of the temperature gradient in the draft tube could be used as an indicator for the onset of the solids circulation. However, in contrast to the former work, in the present study stable circulation of solids was possible for all particle fractions (71 < dp < 160 µm and 250 < dp < 355 µm). This was explained by the increased height of the annulus and, in turn, increased static pressure of the catalytic bed surrounding the draft tube. A better solid entrainment into the centered gas jet was enabled, and gas bypassing around the draft tube was reduced. Furthermore, even for fine particles (71 < dp < 160 µm) an agglomeration in the ICFB reactor was avoided. After the catalytic experiments, the mass of the remaining catalyst in the reactor was determined. With the help of the time on stream, the average entrainment rate was estimated to 0.4 be g‚h-1, which is in the range of the earlier reported values (Mleczko and Marschall, 1997). However, this entrainment rate is much higher than the ones in bubbling fluidized-bed reactors (5%) was obtained, but not in all cases was the initial activity or the thermodynamic equilibrium reached. Also a second species of carbon was identified by the TEM photographs (see Figure 6b). From the literature it is known that this carbon species grows on nickel surfaces and represents a kind of high-temperature graphitic carbon (e.g., Snoeck et al., 1997). In contrast to the encapsulating carbon, the whisker carbon does not result in an instantaneous deactivation of the catalyst because free access to the active nickel sites on the top of the filaments is further possible (e.g., Chesnokow et al., 1995). However, these filaments are brittle and may break in the draft tube of an ICFB reactor under the mechanical stress, resulting in the loss of the active component. This interpretation is supported by TEM pictures of broken carbon filaments; e.g., Figure 6c shows the bottom of such a filament on the surface of the deactivated catalyst. The mechanical destruction of filaments is probably responsible for the loss of the active component. An ICP analysis of the fresh and deactivated catalyst showed after 40 h a decrease of the Ni loading from 0.93% Ni for the fresh catalyst to 0.91 wt % for the deactivated one. Furthermore, after the carbon dioxide treatment, the Ni loading decreased to 0.82 wt %. According to these facts, the lower activity after the regeneration was explained by the loss of the Ni crystallites because of the “burning” of the carboneous connection between the support and the nickel sites. Such a mechanism was already suggested as an explanation of the nickel loss during the catalytic partial oxidation of methane to syngas in a bubbling bed reactor (Santos et al., 1996; Wurzel, 1998). Consequences for the Operation of an ICFB Reactor. For an industrial application of an ICFB reactor, the catalytic stability of the applied catalyst is

of primary importance. Therefore, the formation of whisker carbon should be avoided. This can be achieved by the selection of suitable catalysts. Reduction of the filament growth or even a total suppression of the formation of this carbon species should be possible by application of other precursors. Nichio et al. (1996) showed that the formation of carbon fibers was significantly reduced using a nickel acetylacetonate precursor instead of nickel nitrate whereas the activity was not effected. Furthermore, when partial oxidation of methane was performed over a perovskite-type nickel catalyst, the formation of carbon filaments was totally suppressed (Hayakawa et al., 1997). However, for these kinds of catalysts, a lower activity is expected than the one measured in this study. On the other hand, also reaction engineering means should be used to avoid the deactivation of the catalyst. Therefore, the concept of an ICFB reactor with additional fluidization of the annulus is proposed in this paper to continuously regenerate the catalyst. Carbon dioxide can be introduced in the surrounding annular bed to remove carbon deposition from the catalyst surface. Besides carbon dioxide, oxygen can also be considered as an effective regenerating agent for introduction into the annular bed (Mleczko et al., 1997). However, the partial pressure of the introduced oxygen has to be limited to avoid a total oxidation of the metallic nickel to nickel oxide, which is not active for the reforming reactions. Introduction of an auxiliary gas into the annulus is also advantageous for the solids flow. Gas bypassing through the annulus can be reduced, and solids circulation can be intensified (Yang and Keairns, 1983). Furthermore, it is known from the literature that for the production of carbon fibers nucleation on the catalyst surface and diffusion of carbon through the subsurface are necessary. This takes a certain time. Snoeck et al. (1997) reported that the carbon formation will significantly increase within a time scale of several minutes. Therefore, it is expected that because of low residence times of the catalyst in the draft tube (99.5%). The methane conversion and selectivity to syngas were independent from the gas velocity in the reactor equipped with the 10 cm long draft tube. However, when the 20 cm long draft tube was used, the methane conversion slightly increased with decreasing gas velocity. Hydrogen and carbon monoxide selectivities also increased for this reactor arrangement. It is interesting that, although the contact times in a 10 cm long draft tube were sufficient to approach the thermodynamic equilibrium, the application of a longer draft tube that doubled the contact time resulted in lower conversions and selectivities. The weak influence of the contact time (superficial gas velocity for a given length of the draft tube) on the

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Figure 6. TEM pictures of deactivated Ni/R-Al2O3 catalysts: (a) carbon shell covering the active sites; (b) whisker carbon on the surface; (c) bottom of a “broken” carbon needle.

Figure 7. Methane conversion and hydrogen and carbon monoxide selectivities in dependence on the gas velocity for 10 cm (open symbols) and 20 cm long (filled symbols) draft tubes (150 g of Ni/ R-Al2O3 (1 wt % Ni), CH4/O2 ) 2/1, T ) 800 °C, 250 < dp < 355 µm).

conversions and selectivities indicates that residence times of gas below 0.3 s corresponding to contact times of about 0.020-0.070 g‚s‚ml-1 in the draft tube are sufficient to perform the catalytic partial oxidation of methane to syngas. This result is in line with findings of fixed-bed studies. For example, Dissanayake et al. (1993) investigated this reaction at 800 °C and at contact times up to 0.060 g‚s‚mL-1. The thermodynamic equilibrium was reached under these conditions. Also Bharadwaj and Schmidt (1994) obtained in a monolithic fixed-bed reactor the equilibrium composition at residence times in the range of 10-5 s. Temperature Profiles and Temperature Distribution. Similar temperature profiles were observed in all experiments. Figure 8 shows the temperature profiles in the 10 (a) and 20 cm (b) long draft tubes of the ICFB reactor. The gradients decreased with increasing gas velocity. Because the 10 cm long draft tube was insulated over nearly its whole length by the surrounding annulus, only a small temperature gradient was measured (∆T ) 7-15 K). Above the tube a sharp decrease of the temperature was observed, and temperatures of 495 °C (ugas ) 0.83 m‚s-1) and 556 °C (ugas ) 1.36 m‚s-1), respectively, were measured 15 cm above the top of the draft tube. Because the longer draft tube

Figure 8. Temperature profiles in the 10 cm (a) and 20 cm (b) long draft tubes of the ICFB reactor for different gas velocities (150 g of Ni/R-Al2O3 (1 wt % Ni), CH4/O2 ) 2/1, T ) 800 °C, 250 < dp < 355 µm).

was insulated only in the lower part, greater temperature gradients (∆T ≈ 25-40 K) occurred. The gradients also decreased with increasing gas velocity. The decrease of the temperature gradient is due to the more intensive circulation of solids. A similar effect was observed when the oxidative coupling of methane in the ICFB reactor (Mleczko and Marschall, 1997) was investigated. The temperature distribution in dependence on the superficial gas velocity in the ICFB reactor is shown in Figure 9 for the 10 (a) and 20 cm (b) long draft tubes. In principle, the temperature in the surrounding annulus was equal or slightly lower than the temperature at the bottom of the draft tube for all gas velocities. The

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Figure 10. Methane conversion and hydrogen and carbon monoxide selectivities in dependence on the height in a 10 cm long draft tube (ugas ) 0.91 m‚s-1, 150 g of Ni/R-Al2O3 (1 wt % Ni), CH4/ O2 ) 2/1, T ) 800 °C, 250 < dp < 355 µm).

Figure 9. Temperature distribution in the ICFB reactor in dependence on the superficial gas velocity: (a) 10 cm long draft tube and (b) 20 cm long draft tube (150 g of Ni/R-Al2O3 (1 wt % Ni), CH4/O2 ) 2/1, T ) 800 °C, 250 < dp < 355 µm).

higher temperature at the bottom of the draft tube compared to the one in the annulus indicates that the reaction mainly took place in the riser. At the bottom of the tube, heat was generated by the exothermic total oxidation, which is the first step of the two-stage reaction scheme (eq 1.2). This heat was transported to the upper part of the draft tube by the circulating solids and was consumed by the endothermic reforming reactions. By this means a thermal coupling of the two reaction zones was established and the formation of hot spots such as in fixed-bed reactors (e.g., Dissanayake et al., 1991; Heitnes et al., 1995) can be avoided. This explanation is confirmed by the flat temperature profiles measured in the draft tube. Nevertheless, the difference between the beginning and the top of the riser was higher for the 20 cm long draft tube than for the 10 cm long one. In some experiments an even higher temperature was measured in the annulus than at the bottom of the draft tube. This indicates that the reaction partly took place in the surrounding catalyst bed and can be explained by instabilities in the entrainment region, resulting in bypassing of gas through the annulus. This phenomenon has been reported in the literature (Ye et al., 1992; Berruti et al., 1993; Mleczko and Marschall, 1997). Influence of the Temperature on the Conversion and Selectivities. The conversion of methane measured in the ICFB reactor equipped with a 10 cm long draft tube corresponds to the thermodynamic methane conversion of 89% at 770 °C. When experimental uncertainties are taken into account, this temperature correlates with the temperatures measured 1-3 cm above the top of the draft tube. At this height a catalyst fountain well-known from spouted beds with a draft tube was formed (Mathur and Epstein, 1974). The height and the concentration of solids in the fountain region increased with the gas velocity (He et al., 1994). This promoted backreactions, lowering the methane

conversion. In the case of the 10 cm long draft tube, this region was heated by the outer electrical heater. This resulted in nearly the same temperature in the fountain and, in consequence, the same conversion for all investigated gas velocities. To validate this hypothesis, concentration profiles were measured with the help of a capillary (see Figure 10). In the draft tube the methane conversion decreased from about 91% at the bottom (corresponding to a contact time of about 0.009 g‚s‚mL-1) to about 89% at the end of the tube, whereas the temperature decreased from 800 to about 775 °C. The observed decrease was close to the experimental error. However, for all analyses the conversions at the top of the draft tube were lower than those at the bottom so that the trend is described in the right way. A similar trend was observed for the selectivity to syngas. The methane conversion of about 89% was also determined at the reactor outlet, indicating that no further reaction took place in the particle-free reactor head. This experiments further confirmed that the partial oxidation of methane is a very fast reaction and the equilibrium was already achieved at the beginning of the draft tube. Furthermore, these results agree reasonably well with the findings of bubbling fluidized-bed studies, which showed that the reaction was already finished straight above the gas distributor (e.g., Olsbye et al., 1994; Santos et al., 1994; Mleczko and Wurzel, 1997). The slight decrease of the methane conversion with increasing height in the draft tube indicates that backreactions at the colder draft tube end or in the fountain region might take place, which shift the thermodynamic equilibrium to lower values. 4. Conclusions The ICFB reactor was proven to be generally suited to perform the catalytic partial oxidation of methane. Methane conversions and selectivities to syngas in the range of the thermodynamic equilibrium values were achieved. A hydrogen to carbon monoxide ratio of 2:1 desired for downstream processes was obtained. Furthermore, like in bubbling fluidized beds a nearly isothermal reactor operation was possible using the catalyst as a heat carrier. The circulation of solids through the draft tube resulted in thermal coupling of the exothermic deep oxidation in the lower part of the draft tube with endothermic reforming reactions in the upper part. By this means hot spots were avoided. Nevertheless, a strong influence of the temperature distribution, especially of the temperature in the fountain region, was observed. Because of the lower temperature in the fountain in comparison to the draft tube

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bottom (especially for the 20 cm long draft tube), backreactions occurred. As a consequence, the conversion of methane and selectivity to syngas were lowered. The use of draft tubes with different lengths showed that the contact times (,0.100 g‚s‚mL-1) were always sufficient for this very fast, heterogeneously catalyzed gas-phase reaction. No significant influence of the contact times on the conversion and selectivities was observed. Therefore, an ICFB reactor equipped with a 10 cm long draft tube is sufficient to perform the partial oxidation of methane to syn gas. However, the catalyst stability was strongly influenced by a deactivation due to carbon depositions on the catalyst surface. Two species were identified: encapsulating and whisker carbons. The formation of the last species resulting in Ni loss due to mechanical stress should be avoided by using, e.g., other catalyst precursors, whereas the first species could be removed by carbon dioxide treatment, regenerating the catalyst. When the modular scale-up concept for the ICFB reactor is taken into account, which enables a direct transfer of the results obtained in the laboratory scale to an industrial scale, the ICFB is the most suited reactor design for large-scale operation. An “on-line” regeneration of the Ni catalyst is proposed by carbon dioxide introduction into the annulus surrounding the draft tube of an ICFB reactor, which further improves the solids circulation and diminishes gas bypassing. However, experiments have to be performed to check the suitability of this reactor concept for the POM. Furthermore, studies of the catalyst long-time stability and of the influence of the hydrodynamic conditions in an ICFB reactor on the catalyst activity and a further improvement of the reactor design reducing the solids entrainment are necessary. Nomenclature dp ) particle diameter, µm dpore ) pore size of the quartz distributor, µm mcat ) mass of catalyst, g h ) height of the catalyst entrance window, mm H ) height, cm Hriser ) height of the riser, cm ICP ) inductive coupled plasma K ) equilibrium constant S ) selectivity, % T ) temperature, °C, K TEM ) transmission electron microscopy TPH ) temperature-programmed hydrogenation u ) superficial gas velocity, m/s V˙ ) volume flow, m3/s X ) conversion, %  ) porosity

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Received for review October 19, 1998 Revised manuscript received January 22, 1999 Accepted February 3, 1999 IE9806646