Ind. Eng. Chem. Res. 1997, 36, 4459-4465
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Catalytic Reformer-Combustor: A Novel Reactor Concept for Synthesis Gas Production L. Mleczko,*,† S. Malcus, and T. Wurzel Lehrstuhl fu¨ r Technische Chemie, Ruhr-Universita¨ t Bochum, D-44780 Bochum, Germany
A novel reactor design, so called a reformer-combustor, has been proposed for performing CO2 reforming of methane to synthesis gas with high yield and energy efficiency. In this reactor that consists of two interconnected fluidized beds, i.e. reformer and combustor, heat that is needed for the highly endothermic CO2 reforming is generated by catalytic combustion of methane over a reforming catalyst. The heat generated in the combustor is supplied to the reformer with circulating solid particles. Circulation of solids is used not only for heat transport but also for catalyst regeneration. The experimental studies of CO2 reforming as well as of catalytic combustion of methane over a Ni (1 wt %)/R-Al2O3 catalyst that were performed in a laboratoryscale fluidized-bed reactor (i.d. ) 5 cm) confirmed the feasibility of this reactor concept. 1. Introduction Conversion of methane to synthesis gas (syngas) still represents the major route for utilizing natural gas as chemical or petrochemical feed stocks (Rostrup-Nielsen, 1994). Currently, syngas is mainly produced by the highly endothermic steam reforming of methane. Since in steam reforming syngas with a H2/CO ratio of 3:1 is generated, further gas makeup is necessary to meet the requirements of the down-stream technologies. In recent time carbon dioxide reforming (see eq 1) has
CH4 + CO2 f 2CO + 2H2 ∆rH°298K ) +247.9 kJ/mol (1) attained renewed interest. It is considered an alternative process for producing synthesis gas with a low H2/ CO ratio (1:1) which is needed for direct processing of syngas, e.g. oxo-synthesis (Edwards, 1995). For various catalysts it has been reported that when dry reforming is performed in atmospheric-pressure, laboratory-scale fixed-bed reactors, syngas yields near the thermodynamic equilibrium were achieved (for a review see Wang et al. (1996)). In spite of these promising results the dry reforming process still has not been commercialized, mainly due to the severe catalyst deactivation caused by carbon deposition. Further constraints for performing this reaction in a fixed bed are given by the strong endothermicity as well as heat and mass transport limitations (Rostrup-Nielsen, 1994). In order to overcome these obstacles, a fluidized-bed reactor has been proposed for syngas production by dry reforming. Experimental results obtained in laboratoryscale units indicated that also in this reactor type syngas yields near the thermodynamic equilibrium could be achieved. Due to the intensive solid circulation, almost isothermal reactor operation was attained (Blom et al., 1994; Mleczko et al., 1997). Moreover, lower deactivation rates than in fixed-bed reactors were observed (Wurzel et al., 1997). However, one of the main obstacles with respect to the large-scale application is given by the demand to supply large amounts of heat into the fluidized bed. According to RostrupNielsen (1993), heat fluxes of approximately 80 kW/m2 that result in tube wall temperatures of approximately † Present address: Bayer AG, ZT-TE4, Building E4.1, 51368 Leverkusen, Germany. E-mail: Leslaw.MLeczko.LM@ bayer-ag.de.
S0888-5885(97)00245-5 CCC: $14.00
1050 °C have to be achieved in conventional fixed-bed tubular steam reformers. Due to the higher reaction enthalpy of dry reforming even higher heat fluxes are necessary for the heat supply. Therefore, heating with gas-fired burners usually applied in conventional largescale fixed-bed reformers seems to be not practicable. Also heat supply by means of heat pipes that were proposed for heating fluidized-bed steam reformers (McCallister, 1984) has not found industrial acceptance. Against this background a new reactor concept for dry reforming of methane, so called a catalytic reformercombustor, has been proposed. This design consists of two interconnected fluidized-bed reactors; i.e. in the first reactor endothermic CO2 reforming is carried out, whereas in the second one combustion of natural gas takes place. The reactors should be connected by circulation of the catalyst between both reactors. The circulation of solids will be used for transferring the heat from the combustor to the reformer and for catalyst regeneration by burning-off carbon deposited during the reforming reaction. When influence of carbon deposition on the energy balance is neglected, 8.15 MW has to be supplied to the reformer in order to produce syngas by dry reforming in a reactor with the capacity of 124.2 molsyngas/s (H2/CO ) 0.94). Approximately the same amount of heat (802 MW) can be achieved by combustion of 10.2 mol/s. The heat transport between the reactors can be controlled by the difference of the temperature in the reactors and by the circulation rate of the catalyst; e.g. for a temperature difference between the reformer and the combustor of 50 K, the circulation rate of the catalyst with a heat capacity of 79 J‚K-1‚mol-1 has to amount to 210 kg/s. This rate can be achieved easily in a fluidized bed with solids circulation (Berruti et al., 1995). The concept of the interconnected fluidized beds for coupling exo- and endothermic reactions and for catalyst regeneration is well-known and has been already successfully applied, e.g. for the FCC process (e.g. see King (1992)), for dehydrogenation of paraffins (Sanfilippo et al., 1992), and for hydrogen and syngas production (Matsukata et al., 1995, 1996). A solids circulation system has also been proposed for steam reforming of methane (Guerrieri, 1970). However, in the patent of Guerrieri the catalyst is heated by flue gases generated in a burner placed below the gas distributor. In the arrangement proposed in this paper the combustion of methane will be performed catalytically directly in the © 1997 American Chemical Society
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fluidized bed over the reforming catalyst. This design utilizes all advantages of a flameless combustion; i.e. thermal NOx emissions and soot formation are avoided, and thermal stress is minimized (Pfefferle and Pfefferle, 1987). The work reported in this paper aimed at experimental validation of the feasibility of the concept of a reformer-combustor reactor for CO2 reforming. The experimental results describing the influence of reaction and hydrodynamic conditions on the catalytic performance of a fluidized-bed reactor for the CO2-reforming reaction are summarized elsewhere (Mleczko et al., 1997; Wurzel, 1997). Since the concept of the catalytic reformer-combustor requires a catalyst which is active and selective for both reactions, experimental investigations of the catalytic combustion of methane in a fluidized-bed reactor over a Ni (1 wt %)/R-Al2O3 catalyst that was found to be active for CO2 reforming (Mleczko et al., 1997) were performed. Furthermore, the effect of changing gas atmosphere, i.e. from a reducing one in the reformer to an oxidizing one in the combustor and vice versa on the activity and stability of the circulating catalyst was studied. The experiments were performed in laboratory-scale bubbling beds. In this study circulation of solids was simulated by altering the composition of the feed gases. 2. Experimental Section 2.1. Catalyst. The Ni (1 wt %)/R-Al2O3 catalyst was prepared by incipient wetness technique from R-Al2O3 (Janssen) with an aqueous solution of Ni(NO3)2‚6 H2O (Merck). After drying for 12 h at 100 °C, the catalyst precursor was calcined in air at 470 °C for 10 h. XRD analysis performed for the fresh catalyst after calcination indicated that the nitrate was completely decomposed to form nickel(II) oxide which was the dominant nickel compound on the fresh catalyst (Mleczko et al., 1997). 2.2. Apparatus. The laboratory-scale fluidized-bed reactor (i.d. ) 5 cm, H ) 122 cm) was made of quartz. The reactants were distributed through a porous quartz plate. The catalytic bed (Hmf ) 3 cm) was heated electrically through the reactor wall. At the end of the reactor a disengaging section was located which was additionally equipped with an internal cyclone. Thermocouples within quartz tubes were located in three axial positions of the fluidized bed for measuring the temperature profiles. Reactants (CH4, CO2, N2, O2) and products (CO, H2, CO2) were analyzed on-line by gas chromatography. The amount of water which was condensed downstream was calculated from the oxygen balance. The accuracies of the carbon and hydrogen mass balances were better than (5% and (5%, respectively, for CO2-reforming and better than (8% and (8% for methane combustion, respectively. For every data point three gas samples were analyzed. Details of the experimental equipment are given elsewhere (Mleczko et al., 1990). 2.3. Experimental Conditions. The reactor was operated under atmospheric pressure. Particle diameters in the range from 71 to 160 µm were used. Combustion of methane was studied for a methane-toair ratio of 1:10, i.e. a methane-to-oxygen ratio of 1:2. Reaction temperature was varied between 650 and 850 °C, whereas the mass of the catalyst amounted to 106.2 g (Hmf ) 3 cm, mcat./V˙ ) 2.9-4.8 g‚s‚mL-1). CO2 reforming during the alternated reactor operation was conducted at 800 °C at a contact time of 4.4 g‚s‚mL-1
(mcat. ) 106.2 g, Hmf ) 3 cm). The feed gases (pCH4:pCO2 )1:1) were diluted with 30 kPa nitrogen. 2.4. Method of Investigation. Both reactions were examined separately as well as in the alternated operation mode by switching feed gases between CO2 reforming and methane combustion. For determination of the influence of reaction conditions on catalyst properties (coke depositions, oxidation state, catalyst phases), the catalyst was analyzed by XRD, TEM, and combustion analysis. In order to elucidate the effect of reactions taking place in freeboard in all combustion experiments, methane and oxygen conversions were measured at the reactor outlet and on a height of 3 cm above the gas distributor. Furthermore, for measurement of concentration profiles, gas samples were sucked from the fluidized bed by means of an axial movable sample tube (o.d. ) 0.8 cm). 3. Results and Discussion 3.1. Reactor Operation. Mechanical Stability. The Ni/R-Al2O3 catalyst was well-fluidizable. The minimum fluidization velocity determined experimentally at 800 °C in a nitrogen stream amounted to 0.6 cm/s. Visual observations indicated that in the range of gas velocities applied in the investigations of both reactions (u/umf ) 6.5-11.8) the reactor was operated in the freely bubbling regime. The catalyst exhibited also quite good mechanical and thermal stability. No effect of the gas atmosphere on the mechanical stability of the catalyst was detected. The average attrition rate of the catalyst determined during 129 h of operation amounted to 0.041 g/h. Energy Balance and Temperature Control. Because heat losses in laboratory reactors are much higher compared to industrial units, they are not representative with respect to the conditions yielding autothermal operation. Nevertheless, in order to check if conditions applied in these investigations are representative for a future large-scale operation, energy balances for both investigated reactors were analyzed. During the alternated operation the reformer consumed 58 W (XCH4 ) 60%). Simultaneously, during the combustion 82.3 W were released. At thermodynamic equilibrium (XCH4 ) 90.4%) energy supply has to be increased to 87.4 W. In both reactions the good mixing of solids resulted in almost isothermal operation; maximum temperature gradients measured in the bed amounted to 2 K (see Figure 1). In the freeboard a monotonous decline of temperature occurred. However, during the combustion stable reactor operation was only possible for temperatures below 780 and above 820 °C. Above 780 °C a steep increase to 820 °C occurred; this effect indicates that at this temperature ignition occurred. 3.2. Catalytic Combustion of Methane. Catalytic Performance. Conversions of methane and oxygen depended strongly on the bed temperature (see Figure 2). Only low conversions (XCH4, XO2 < 23%) were measured at temperatures below 780 °C. However, when the temperature was raised up to 820 °C, a steep increase of conversions took place. At this temperature almost complete conversions of methane and oxygen were determined. Moreover, the conversions did not depend on contact time (4.8 g‚s‚mL-1 > mcat./V˙ > 2.9 g‚s‚mL-1) or hydrodynamic conditions. Catalytic Stability. Conversions of methane and oxygen measured on a height of 3 cm above the gas distributor dropped during 4.5 h of operation from 31
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Figure 1. Axial temperature profile measured during methane combustion (u/umf ) 8.6, TR ) 850 °C, pair:pCH4 ) 10:1, pCH4:pO2 ) 1:2).
Figure 2. Effect of reaction temperature on methane (9) and oxygen (b) conversion (Hmf ) 3 cm, TR ) 850 °C, u/umf ) 6.8, pair:pCH4 ) 10:1, pCH4:pO2 ) 1:2).
Figure 3. Dependence of methane (9) and oxygen (b) conversion on time-on-stream (Hmf ) 3 cm, u/umf ) 6.8, TR ) 850 °C, pair: pCH4 ) 10:1, pCH4:pO2 ) 1:2).
to 23% and from 33 to 27%, respectively (see Figure 3). Simultaneously, the catalyst particles changed their color from gray for the fresh catalyst to a light green color at the end of the experiment. According to Tourky et al. (1970), this change of color indicates that the chemical composition of the nickel oxide changes during reaction. At the beginning of the reaction NiO (graycolored) represents the calcination product of Ni(NO3)2 at 470 °C, which contains more oxygen than the sto-
ichiometric one (NiOx, x > 1 (Tourky et al., 1970; Deren and Stoch, 1970)). It is assumed that the high bed temperature during methane combustion promotes the transition of the nonstoichiometric compound to a green one which is assumed to represent stoichiometric nickel oxide (NiO (Tourky et al., 1970)). Therefore the observed decline of catalytic activity with time on stream can be explained by the loss of excess oxygen from nickel oxide as already postulated by Krauss (1949) who investigated the oxidation of ammonia over nickel oxide. Similar conclusions were also presented by Gavalas et al. (1984) who reported that nonstoichiometric NiO containing excess oxygen and, in turn, Ni3+ sites is the active component during oxidative conversion of methane. In order to verify this hypothesis, the deactivated catalyst was treated at 600 °C for 1 h with an oxygen/nitrogen (1:1) mixture. After this treatment an increase of methane conversion (measured at a height of 3 cm above the distributor) from 23.0% for the deactivated catalyst to 27.0% was achieved. Although this value is lower than the initial activity that amounted to 31.0%, it can be concluded that the stoichiometry of the nickel oxide has significant influence on the catalytic activity. Nevertheless, also other factors, e.g. adsorption of water (Firth and Holland, 1969) may contribute to the drop of activity. Concentration Profiles. Concentration profiles were measured for two gas velocities (u/umf ) 6.5, 8.6) at 850 °C (Figure 4a-c). For both experiments steep concentration gradients occurred in the distributor zone. These gradients were more pronounced for the higher gas velocity. For the lower gas velocity a continuous increase of conversion of reactants was measured in the whole bed. For the higher gas velocity only small changes of gas composition occurred in the dense bed above the distributor zone, which indicates that the extent of reaction is influenced by mass transfer between the bubble and emulsion phases. The oxygen concentration profiles indicate that for both gas velocities almost complete conversions were achieved at a height of 6.5 cm (freeboard), whereas conversions at the end of the catalytic bed (H ) approximately 4 cm) amounted to only 50% (see Figure 4a). The increase of the concentrations of water and carbon dioxide was complementary to the decrease of the concentrations of methane and oxygen (see Figure 4b). Significant amounts of carbon monoxide, i.e. up to 1.2 vol % were formed, especially nearby the gas distributor (see Figure 4c). Even at the end of the catalytic bed the concentration of carbon monoxide amounted to 50% of the one of carbon dioxide. In the freeboard zone carbon monoxide was almost completely oxidized to carbon dioxide. In contrast, only small amounts of hydrogen (see Figure 4c) that did not exceed 0.4 vol % were detected in the dense bed and were almost completely consumed in the freeboard zone. Catalytic or Thermal Combustion. In order to elucidate whether or not NiO/Al2O3 is catalytically active for the methane combustion, comparative measurements applying a bed of quartz particles were performed. In contrast to the experiments with the NiO/Al2O3 catalyst, the flameless combustion was not possible. Above the ignition temperature (840 °C) that was significantly higher than in the catalytic bed an explosive combustion occurred. The explosive combustion in freeboard of gases that leave the fluidized bed has been already reported for the combustion of natural
4462 Ind. Eng. Chem. Res., Vol. 36, No. 11, 1997
Figure 4. (a) Concentration profiles of methane (9, 2) and oxygen (b, 1) (Hmf ) 3 cm, TR ) 850 °C, pair:pCH4 ) 10:1, pCH4:pO2 ) 1:2). (b) Concentration profiles of water (b. 1) and carbon dioxide (2, 9) (Hmf ) 3 cm, TR ) 850 °C, pair:pCH4 ) 10:1, pCH4:pO2 ) 1:2). (c) Concentration profiles of hydrogen (2, [) and carbon monoxide (9, b) (Hmf ) 3 cm, TR ) 850 °C, pair:pCH4 ) 10:1, pCH4:pO2 ) 1:2).
gas in the bed of inert particles (Sadilov and Baskakov, 1973). Makhorin and Glukhomanyuk (1978) found that smooth combustion of methane in the inert fluidized bed was possible only at temperatures above 1000 °C. The difference between ignition temperatures in the inert and catalytic beds confirms that Ni/R-Al2O3 is catalytically active for the oxidation of methane. Also the lower conversion in the inert bed supports the conclusion that the combustion is ignited catalytically and not thermally. Finally, this result is in line with outcomes of the investigations of partial oxidation of methane (Dissanayake et al., 1991). For this reaction which proceeds in two consecutive steps, i.e. the total oxidation and reforming of methane, Ni/R-Al2O3 was
found to be active for both reaction steps. Its activity depended, however, on the oxidation state of nickel. The high concentrations of reactants and intermediate products in the dense part of the catalytic bed indicate that catalytically initiated reactions proceed further in the gas phase. This explanation is supported by the high concentration of CO that is known to be a major product of methane oxidation in the gas phase at low temperatures in a oxygen poor atmosphere (Zanthoff and Baerns, 1990). The low conversion level at the end of the dense bed can be explained by the high mobility of the solid particles that promote the termination of radicals. The importance of the termination of radicals on the solid surface for the extent of gas phase reactions has been reported already by Couwenberg et al. (1996). Furthermore, Hesketh and Davidson (1991) reported that the particulate phase of an incipiently fluidized bed of sand inhibits combustion of propane and methane. In turn, combustion took place mainly in the bubble phase and above the bed surface. This hypothesis is furthermore supported by the high efficiency of a fluidized bed for flame quenching (Donsaı` et al., 1975). Also Feugier et al. (1987) found for the combustion of methane in a circulating fluidized bed that with decreasing porosity higher temperatures had to be applied to achieve almost complete conversions. It cannot be excluded that the high concentration of hydrogen and carbon monoxide in the catalytic bed is caused by partial oxidation of methane or gasification of carbon deposited on the catalyst particles. However, this explanation is less probably since the H2/CO ratio is lower than 1. Such a low hydrogen-to-carbon monoxide ratio is rather untypical for partial oxidation but corresponds well to the results of the gas phase reaction of methane and oxygen. In the experiments of methane partial oxidation performed at 800 °C in a fluidized bed consisting of inert material (R-Al2O3), the H2/CO ratio amounted to 0.5 (Mleczko and Wurzel, 1997). Furthermore, the catalyst in the combustor is in a highly oxidized form; i.e. it consists mainly of nickel oxide, which is rather inactive for the reforming reactions (Dissanayake et al., 1991). Since on the catalyst particles no carbon deposits were detected, neither by combustion nor by TEM analysis, the contribution of this reaction to the overall conversion should be negligible. Furthermore, the influence of the methanation reaction on the drop of the concentration of hydrogen and carbon monoxide in the freeboard is much less probably, since no increase of methane concentration was observed. The temperature for achieving complete conversions of methane by applying Ni/Al2O3 was significantly higher than the one reported for methane combustion in a turbulent bed over a Pd/Al2O3 catalyst (Foka et al., 1994). They reported that stable reactor operation was obtained already at 500 °C. This difference can be caused not only by the higher activity of palladium compared to nickel but also by the different fluidization regimes. 3.3. Simulation of Operation of ReformerCombustor Reactor. This series of experiments aimed at elucidation of the effect of changing gas atmosphere on the activity and stability of the Ni/R-Al2O3 catalyst. Reaction temperature during combustion (cycles 2, 4, 6) amounted to 850 °C, whereas CO2 reforming (cycles 1, 3, 5, and 7) was studied at 800 °C.
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reforming reaction. The absolute level of methane and CO2 reforming is relatively low since these experiments were performed in a very shallow bed. The static experiments indicated that contact times longer than 7 g‚s‚mL-1 are necessary in order to achieve conversions approaching thermodynamic equilibrium. During the following cycles a quasi stationary state was attained. In this state the highest conversions were measured at the beginning of each cycle, i.e. after the feed gas was changed. These results confirm that a high activity of the Ni/R-Al2O3 catalyst for the reforming reaction could be restored in the combustion atmosphere. On turn, during the reforming reaction the catalyst was regenerated with respect to its activity for the combustion of methane. However, during each combustion and reforming cycle conversions decreased. With respect to the deactivation rate, the catalyst deactivated faster during the combustion than during the reforming reaction. The drop of catalytic activity for the reforming reaction amounted to 20% (relative) during a period of 400 min. Over a period of 40 min deactivation during combustion resulted in a relative decrease of conversions of methane and oxygen of 16 and 25%, respectively. Effect of Alternating Reactor Operation on Catalyst Properties. In order to explain the effect of the varying gas atmospheres on the catalytic activity, the surface and bulk properties of the catalyst were analyzed (see Table 1). The XRD pattern indicated that the catalyst applied for CO2 reforming consisted of reduced and oxidized nickel crystallites (Figure 6a). On the reduced nickel crystallites carbon was deposited as confirmed by TEM analysis. Since this carbon encapsulated the Ni crystallites, it can be expected that this is the origin of the drop of catalytic activity with time on stream. The amount of carbon deposited on the catalyst was quantified by means of combustion analysis to 0.1%. Because no carbon was found on the support but only on nickel crystallites, the carbon loading related to the nickel content amounted to 10 wt %. Furthermore, the crystallite size determined by TEM amounted to approximately 30-50 nm. In contrast, no carbon was found on the catalyst samples that were taken from the reactor during catalytic combustion. Furthermore, the average crystallite size was significantly lower; it amounted to less than 5 nm. The XRD pattern of this catalyst did not exhibit any peak of metallic nickel (see Figure 6b). However, the size determination of Ni crystallite surrounded by carbon by means of an optical device is impeded by the difficulty to distinguish between nickel and carbon. Therefore, the increase of the crystallite size after the reforming cycles can also be attributed to the deposition of carbon, whereas the decrease of the crystallite size results from the combustion of carbon deposits. In order to obtain further information on the influence of the gas atmosphere on the state of the catalyst, one probe sampled during the combustion period was further treated with methane (800 °C). Furthermore, a sample taken from the fresh (unreduced) catalyst was
Figure 5. (a) Effect of methane combustion on methane (9) and carbon dioxide (b) conversion during CO2 reforming (Hmf ) 3 cm, TR ) 800 °C, u/umf ) 7.5, pCH4:pCO2:pN2 ) 1:1:0.86). (b) Effect of carbon dioxide reforming on methane (9) and oxygen (b) conversion during catalytic combustion (Hmf ) 3 cm, TR ) 850 °C, u/umf ) 6.8, methane:air ) 1:10, pCH4:pO2 ) 1:2).
Conversions in the Alternated Operation Mode. When applying a fresh (unreduced) catalyst for CO2reforming, conversions of methane and carbon dioxide exhibited a slow increase during the first cycle of reforming (see Figure 5a). It has been shown that the increases of conversions are caused by the in-situ reduction of the catalyst (Mleczko et al., 1997). In the following combustion period fast deactivation of the catalyst occurred (see Figure 5b). The low level of the conversions during combustion is due to the fact that the gas composition was measured on a height of 3 cm above the gas distributor. Although always complete conversions of methane and oxygen were measured at the reactor outlet, a very fast drop of conversions measured in the bed occurred. The high conversions measured during the second reforming cycle indicates that the oxidizing atmosphere in the preceding combustion promoted high activity of the catalyst for the
Table 1. Properties of Different Catalyst Samples during Alternating Reactor Operation sample reforming combustion fresh, reduced in H2 after combustion, treated with CH4
carbon deposits
nickel species
10 wt %, encapsulated
Ni0,
Ni2+
10 wt %
Ni2+ Ni0, Ni2+ Ni2+
crystallite size, nm 30-50
