Some Considerations on the Design and Operation of High

Ind. Eng. Chem. Res. 2004, 43, 4657-4667

4657

Some Considerations on the Design and Operation of High-Temperature Catalytic Reverse-Flow Reactors Dirk Neumann,†,‡ Vanessa Gepert,† and Go1 tz Veser*,†,‡ Department of Chemical Engineering, University of Pittsburgh, Pittsburgh, Pennsylvania 15261, and Max-Planck-Institut fu¨ r Kohlenforschung, 45470 Mu¨ lheim an der Ruhr, Germany

Reverse-flow operation of catalytic reactors is known to result in very efficient regenerative heat integration. While reverse-flow operation is in particular ideally suited for high-temperature short-contact-time reactions, the application of this reactor concept under such extreme reaction conditions also requires particular care in the design and operation of the reactor. We report results from an experimental study on some of the main design and operating parameters for high-temperature catalytic reverse-flow reactors. Results for reverse-flow operation are compared to conventional stationary reactor operation. Temporal and spatial temperature profiles at periodic steady state are presented, and the influence of the structure of the catalyst and inert zones, the role of homogeneous reactions, and the frequency of flow switching are investigated and discussed. Introduction Catalytic partial oxidation of hydrocarbons at hightemperature, millisecond-contact-time conditions (T ) 800-1200 °C, τ ) 1-50 ms) has received much attention over the past decade as a novel reaction engineering route to base chemicals in the petrochemical industry. Under these rather extreme reaction conditions, hydrocarbons can be converted with good selectivities and yields over noble metal catalysts to more valuable products such as synthesis gas, olefins, and oxygenates.1-4 The reaction route is of significant technological interest, since direct oxidation of hydrocarbons is a moderately exothermal reaction route, opening the possibility of autothermal process operation. Furthermore, due to the high-temperature conditions, the reaction proceeds with extremely high reaction rates, allowing for unusually high space-time yields (in excess of 106 h-1) and thus very high reactor throughputs in very small and compact reactor configurations. A particularly interesting example of catalytic oxidation of hydrocarbons is the direct oxidation of methane with air or oxygen to synthesis gas, i.e., a mixture of CO and hydrogen.1,5-10

CH4 + 1/2O2 f CO + 2H2

∆HR ) -37 kJ/mol

Beyond the prospect of autothermal reactor operation, which makes this reaction route energetically favorable to the strongly endothermal steam re-forming of methane, the direct oxidation route has the additional advantage of yielding the product gases CO and H2 at a stoichiometric ratio of 1:2, which is the desired ratio for downstream processes such as methanol synthesis or Fischer-Tropsch reactions. Consequently, no additional secondary re-formers or water-gas shift stages * To whom correspondence should be addressed. Address: Department of Chemical Engineering, 1249 Benedum Hall, University of Pittsburgh, Pittsburgh, PA 15261. Tel.: (412) 624-1042. Fax: (412) 624-9639. E-mail: [email protected]. † University of Pittsburgh. ‡ Max-Planck-Institut fu¨r Kohlenforschung.

are necessary, further adding to the simple and compact nature of this process route. Most intriguingly, no thermodynamic limit on syngas yields exists at sufficiently high reaction temperatures and stoichiometric feed conditions, i.e., virtually 100% yield of CO and H2 can be attained at equilibrium from methane/air as well as methane/oxygen mixtures (see Figures 1 and 2). Therefore, in conjunction with an autothermal process route (i.e. no additional heating that would result in additional emissions), catalytic partial oxidation of methane opens the possibility of a true zero-emission process for syngas formation. However, thermodynamic calculations also show that this “zero-emission” process can only be achieved at reaction temperatures above 900 °C, which is well in excess of adiabatic limits (the adiabatic temperature rise for stoichiometric methane/air mixtures, for example, is only ∆Tad ≈ 250 K!). Hence, at adiabatic operation, a fraction of the methane feed must be (internally) combusted to achieve the necessary reaction temperatures. This, however, limits the adiabatically attainable syngas yields to well below optimum yields (see Figures 1 and 2): for reactor operation with air, syngas yields at stoichiometric conditions drop by 30-40%, and even at operation with pure oxygen, maximum yields are reduced by 15-20%. Clearly, superadiabatic reactor operation is required to fully exploit the potential of this promising reaction route. We have previously shown that heat-integrated reactor concepts are very efficient ways to achieve such superadiabatic process conditions. This can be done via integrated recuperative countercurrent heat exchange between the hot effluent gases and the cold feed gases11 or, more efficiently, via regenerative heat exchange in a reverse-flow reactor.12 This dynamic mode of reactor operation, in which the periodic switching of the flow direction of the gases in the catalytic reactor leads to a highly efficient integration of the reaction heat, was pioneered in the early 1980s by Boreskov, Matros and co-workers,13,14 and the functional principle was then elegantly explained in the following years by Eigenberger and co-workers, who studied this reactor concept

10.1021/ie034257s CCC: $27.50 © 2004 American Chemical Society Published on Web 05/25/2004

4658 Ind. Eng. Chem. Res., Vol. 43, No. 16, 2004

Figure 1. Results from equilibrium calculations for methane/air mixtures with varying methane/oxygen ratio. Hydrogen yield (top left), CO yield (top right), and equilibrium temperatures (bottom) are shown for adiabatic conditions (diamonds) and a fixed temperature, at which maximum hydrogen yields at the corresponding methane/oxygen ratio are obtained (triangles).

Figure 2. Results from equilibrium calculations for methane/ oxygen mixtures with varying methane/oxygen ratio. Hydrogen yield (top left), CO yield (top right), and equilibrium temperatures (bottom) are shown for adiabatic conditions (diamonds) and a fixed temperature, at which maximum hydrogen yields at the corresponding methane/oxygen ratio are obtained (triangles).

for the incineration of waste gases15-17 as well as for the (endothermal) catalytic dehydrogenation of ethylbenzene.18 In the present contribution, we extend our previous report on a high-temperature reverse-flow reactor by studying the influence of main reactor design parameters on the performance of reverse-flow reactors under high-temperature conditions. Experimental Setup The experimental setup is explained in more detail in a previous report and will therefore only briefly be described here.12 The reactor setup is shown schematically in Figure 3. The catalyst used in all experiments reported here is an alumina foam monolith that is coated with a thin layer of platinum via conventional wet impregnation from a salt solution. This catalyst is positioned between two extruded cordierite monoliths that act as inert heat reservoirs. This configuration allows an efficient regenerative heat integration in the following way: cold reactants enter the catalyst bed and the heat of reaction leads to increasing temperatures in the catalytic zone as well as in the inert zone downstream of the catalyst. At reverse-flow operation, the gas flow through the reactor is reversed at one point in time so that the cold reactor feed is now flowing through the preheated inert zone, increasing the temperature of the reactants and at the same time cooling the inert zone. The reactants thus enter the catalyst zone at an already elevated temperature level, which is then further raised by the heat of reaction. Upon exiting the catalyst zone, the hot product gases exchange heat with the (“new”) down-

Figure 3. Schematic view of the experimental setup: the catalyst is positioned in the reactor tube between two inert zones. Two pairs of valves (V1 and V2) are positioned diagonally on either side of the reactor tube and allow the periodic switching of the flow direction. Product gases are detected via mass spectrometer (MS) and gas chromatographs (GC).

stream inert zone, which is thus heated. If this flow reversal is repeated at an appropriate cycling frequency,

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a “periodic steady state” is eventually reached, which is characterized by the fact that the temperature and concentration profiles of two subsequent half-periods are mirror images of each other. In this way, a symmetric temperature profile is achieved that is characterized by very high reaction temperatures in the reaction zone and comparatively cold temperatures at both reactor ends. The necessary flow reversal is accomplished in the laboratory reactor with the aid of four magnetically operated valves, which are synchronized in pairs. Moveable thermocouples allow the measurement of temperature profiles along the reactor axis in the inert zones. The irregular foam structure of the catalyst zone, however, precludes temperature measurements inside the catalyst zone itself. A mass spectrometer (Balzers) is used for qualitative, time-resolved concentration measurements, and a double-oven gas chromatograph (Shimadzu) allows time-averaged, quantitative determination of product concentrations. In a typical experiment, C-, O- and H-atom balances close to better than 3%. The experimental setup (reactor operation and data acquisition) is fully computer-controlled (PC with NI DA/AD-boards and LabView software). All results reported in the following were obtained during reactor operation with air (rather than pure oxygen). For a direct comparison of the performance of a reverse-flow reactor with a conventional stationary reactor, the same reactor setup is used for steady-state experiments without switching of the flow direction. In this way, it is assured that all differences between steady-state results and results at reverse-flow operation are due to differences in reactor operation rather than differences in the experimental setup. Results & Discussion Catalyst Structure. Previously, we had reported results over a Pt-coated alumina foam monolith with 45 ppi (pores per inch) that demonstrated the strong improvement of syngas yields at reverse-flow operation versus stationary reactor operation.12 Monolithic catalysts are ideally suited for high-temperature millisecondcontact-time reactions, as the open porous structure of the monoliths allows for very high through-put with negligible pressure drop while at the same time ensuring reasonably high catalytic surface areas. Varying the number of pores per inch (ppi) of the catalyst, i.e., varying the diameter of the pores, not only is the catalytic surface area changed but, more importantly for regenerative heat integration, the heat storage capacity as well as the heat conductivity in the catalyst zone will be significantly affected. It can therefore be expected that the pore density of the catalyst has a noticeable impact on reactor performance. We studied the influence of pore density on the reactor performance at steady-state (SS) and reverse-flow operation (RFR) for three different Pt-coated alumina foams (30, 45, and 80 ppi). Results for steady-state operation are shown in Figure 4 and results at reverseflow operation are shown in Figure 5, where hydrogen and carbon monoxide yields as well as catalyst exit and entrance temperatures (for the SS reactor) and maximum catalyst exit temperatures (for the RFR) are shown with increasing reactor inlet gas flow from 1 to 5 slm (standard liters per minute). Generally, it can be seen that syngas yields in the RFR are strongly increased in comparison to SS opera-

Figure 4. Syngas yields and temperatures at the catalyst inlet and exit versus total volumetric inlet flow rate in the stationary catalytic reactor with varying catalyst pore diameter: 30 ppi (diamonds and dotted lines), 45 ppi (triangles and dashed lines), and 80 ppi (circles and solid lines) foam monoliths. Experimental conditions: CH4/O2 ) 2.0, 45 ppi foam monolith.

tion at all conditions tested. Clearly, heat integration is beneficial for syngas yields, independent of the catalyst structure. As already noted in our previous report, the improvement of process yields through regenerative heat integration is more and more pronounced as the total gas flow rate increases: while for 1 slm total gas flow the heat integration leads to relatively minor improvements in both yields of about 5-10%, this increases strongly toward 5 slm total flow, where yields are improved by as much as 30-40%! As previously explained, this can be traced back to the increasingly efficient utilization of the inert zone as heat reservoirs for the regenerative heat exchange at reverseflow operation, which leads to increasing catalyst entrance temperatures (see also below, Figure 6) and hence strong improvements in syngas yields. Beyond this difference with increasing flow rate, however, both SS and RFR results show qualitatively identical trends: hydrogen and CO yields increase with increasing pore density, with the increase from 30 to 45 ppi being more pronounced than the increase from 45 to 80 ppi. Maximum reactor temperatures in the RFR are about 200 K above maximum temperatures observed at SS. As expected, heat integration leads to increased reactor temperatures. However, this increase is not as strong as might be expected on the basis of the extremely high catalyst exit temperatures observed at SS operation. This indicates how efficiently the heat integration converts the sensible heat of the product gases into latent heat, i.e., there is an increased selectivity toward partial oxidation products, so only a small fraction of the integrated heat shows up in increased product gas temperatures. It is also noteworthy that even the


Figure 5. Syngas yields and maximum temperatures at the catalyst inlet/exit versus total volumetric inlet flow rate in the catalytic reverse-flow reactor with varying catalyst pore diameter: 30 ppi (diamonds and dotted lines), 45 ppi (triangles and dashed lines), and 80 ppi (circles and solid lines) foam monoliths. Experimental conditions: CH4/O2 ) 2.0, 45 ppi foam monolith, duration of half-period τ/2 ) 15 s.

Figure 6. Experimentally measured temperature profiles in the catalytic fixed-bed reactor at stationary operation (top) and at dynamic reverse-flow operation (bottom, period length ) 30 s) for three different total inlet gas flow rates. Due to the irregular foam structure of the catalyst (Pt-coated 45 ppi foam monolith), temperatures in the 10-mm-wide catalyst zone at the center cannot be measured. Experimental conditions: total gas flow 3 slm, CH4/O2 ) 2.0, 45 ppi foam monolith.

highest reactor temperatures observed in these experiments (T < 1300 °C) did not pose a problem for reactor design and materials in the sense that no degradation of reactor materials was observed at any conditions over the operating time scales presented in this report, an often stated concern for the operation of high-temperature, heat-integrated reactors. Finally, one can see in the temperature curves that the maximum temperatures in the RFR and the catalyst exit temperatures at SS operation decrease with increasing pore density, as expected due to the increasing selectivity of the reaction and the correspondingly decreasing heat of reaction. For the reverse-flow reactor, temperatures decrease by as much as 230 K, emphasizing again that maximum temperatures do not need to be a concern in a high-temperature RFR as long as the reaction proceeds with sufficient selectivity. Interestingly, however, the catalyst entrance temperatures at SS increase with increasing pore density, giving rise to a scissor-like behavior of the catalyst entrance and exit temperature at SS, and at lowest flow rates and highest pore density, the catalyst entrance temperature even exceeds the catalyst exit temperature. This seems to indicate that with decreasing flow rate and increasing pore density the reaction zone shifts toward the catalyst entrance. This can be traced back to the counteracting effects of convective heat transport by the gas flow and the conductive heat transport in the solid catalyst/support material: The convective heat transport by the gas flow leads to a cooling of the catalyst front edge, so that with increasing gas flow the reaction front is pushed further into the catalyst bed. Heat conduction in the catalyst, on the

other hand, tends to counteract any steep temperature gradient in the catalyst zone and thus predominantly leads to heat transport toward the catalyst entrance, where extremely steep temperature gradients develop due to the high reaction rate. Since denser foams allow for more efficient heat conduction, this “conductive flattening” of the steep temperature gradient is most pronounced for the 80 ppi monolith. (This can in fact be regarded as a simple “heat integration” at work, even at stationary reactor operation!) Therefore, the combination of the increased heat conduction with the decreased convective cooling with decreasing gas flow and increasing monolith density eventually leads to a catalyst entrance temperature that is higher than the catalyst exit temperature. Overall, these results indicate that not only the selection of an appropriate high-temperature-stable, highly selective catalyst is a necessary precondition for hightemperature RFR operation, but also the physical structure of the catalyst is an important design parameter with significant impact on reactor operation and performance. Temperature Profiles. The effect of heat integration on reactor temperature is apparent in more detail in the measured axial temperature profiles shown in Figure 6, where temporally averaged temperature profiles as measured via moveable thermocouples inside the (extruded) inert zones are shown for three different total gas flow rates (1, 3, and 5 slm) at SS operation (upper graph) and at periodic steady-state in the RFR (bottom graph). (Due to the foam structure of the catalyst, no measurements are possible in the catalyst zone, which is located at z ) 11-12 cm.)


Figure 7. Experimentally measured temperature profiles in the catalytic reverse-flow reactor during two half-periods at periodic steady state (τ/s ) 15 s). The position of the catalyst is indicated by the shaded area between z ) 11 and 12 cm. Experimental conditions: total gas flow 3 slm, CH4/O2 ) 2.0, 45 ppi foam monolith. (Temperatures inside the catalyst between z ) 11 and 12 cm cannot be measured and are therefore linearly interpolated between the catalyst exit and entrance temperatures purely for reasons of graphic representation.).

It is obvious from a comparison of the temperature profiles at SS and RFR that heat integration leads to strongly increased entrance temperatures at the catalyst zone. This increase is about 400 K at all three flow rates. At the same time, the catalyst exit temperatures are only mildly increased by about 100 K at all three gas flows. This again indicates that the increased feed temperatures lead to a very efficient improvement of syngas yields, thus converting much of the increased thermal energy of the feed gases into chemical energy. Furthermore, it can be seen from these graphs that while catalyst entrance and exit temperatures are increased in the RFR, the reactor exit temperatures are decreased by as much as 400 K at the highest flow rates. It is also apparent that this effect is in fact strongly dependent on the gas flow rate, with the strongest temperature decrease occurring at the highest flow rate. As we already observed in our previous experiments,12 reverse-flow operation is most efficient at high flow rates and is thus a perfect match for these high-temperature ultrashort-contact-time reactions! Overall, the temperature profiles indicate that reverseflow operation leads to a true integration of the reaction heat into the process, with increased temperatures in the catalytic reaction zone (as indicated by the increased entrance and exit temperatures) and decreased temperature in the inert zones surrounding the catalyst (the so-called “cold in-cold out” principle of heat integration). Figure 7 offers a detailed look at the temporal development of the temperature profiles over one full period at periodic steady-state in the reverse-flow reactor for a 45 ppi catalyst at 3 slm total gas flow with a switching frequency of 1/15 Hz. Beyond the highly symmetric shape of the temperature profiles on either side of the catalyst (as already apparent in the temporally averaged profiles in Figure 6), one can see that the temperature profile at the beginning and the end of the period (i.e. at t ) 0 and 30 s) are identical, as required by the definition of the “periodic steady-state”. The results further demonstrate that the temperature oscillations within one period at

periodic steady-state are about 200 K at the catalyst entrance and exit and less than 100 K at the reactor ends. Thus, excessive temperature maxima are clearly avoided at all times during the period. (Although temperatures inside the catalyst bed could not be measured, detailed reactor simulations indicate that maximum temperatures within the catalyst bed never exceed temperatures in the inert zones by more than 200 K19). Homogeneous Reactions. Another concern when dealing with high-temperature catalytic oxidation of hydrocarbons is the occurrence of homogeneous gasphase reactions. These reactions generally need to be avoided, since homogeneous reactions typically result in selectivity losses and, more importantly for the current reaction system, in a loss of reactor controllability or even serious safety concerns due to the occurrence of open flames or explosions if reactant concentrations cannot be guaranteed to be outside the flammability range everywhere in the reactor at all times. While typical catalytic reactions are conducted at temperatures well below the ignition temperatures for homogeneous reactions (often this is the very reason for the use of a catalyst), high-temperature catalysis operates under conditions where homogeneous reactions are possible and hence the reactor has to be designed and operated in a way that ensures prevention of these reactions. We have previously shown in detailed reactor simulations that for steady-state reactor operation in catalytic partial oxidation, homogeneous reactions are avoided due to the fact that homogeneous oxidation of methane shows a significant ignition delay time. For millisecondcontact-time reactors, this ignition delay is on the order of the whole catalyst contact time in the reaction system.20 Therefore, while homogeneous reactions are in principle allowed, competition between catalytic and homogeneous reaction pathways leads to complete catalytic conversion of methane due to the absence of an ignition delay for the catalytic reaction as well as the significantly higher catalytic reaction rates. At reverse-flow operation, however, the situation is significantly altered by the fact that the catalyst is now surrounded by long, catalytically inert zones (the extruded cordierites in our setup). Additionally, the regenerative heat-exchange leads to strongly increased temperatures in the inert zones (see Figure 6), thus increasing the probability of precatalytic homogeneous reactions. Great care must therefore be taken in the reactor design and operation not to exceed the ignition delay time at the corresponding temperature inside the upstream inert zone in the RFR. Figure 8 summarizes results from a simple reactor design calculation for this purpose: shown is the ignition delay time against CH4/O2 ratio and reaction temperature for methane/air mixtures in a simple isothermal plug-flow reactor simulation with detailed homogeneous chemistry (GRI-mechanism 3.0). As apparent from the 3D surface plots in the upper graph, ignition delay times decrease strongly with decreasing CH4/O2 ratio and increasing temperature, from almost 400 ms for very fuel rich mixtures with CH4/O2 ) 3 at a temperature slightly below 900 °C (1160 K) to about 30 ms for the leanest mixture (CH4/O2 ) 1) and a temperature of 1300 K. In addition to the ignition delay times, the graphs show the residence times of the reactor feed gases for


Figure 8. Ignition delay for methane/air mixtures as a function the CH4/O2 ratio and isothermal temperature (shaded surface in top graph and shaded contours in bottom graph) along with residence time of the gases in the inert zones of the laboratory reactor (white surface in upper graph and vertical contour lines in lower graph). Results from numerical PFR simulations with detailed homogeneous gas-phase chemistry (GRI-mechanism 3.0).

our laboratory reactor geometry (at 3 slm total inlet flow) for the given gas temperature (flat, white surface in the upper graph and vertical contour lines in the lower graph). The intersection of the two planes in the 3D plot signifies the maximum temperature that must not be exceeded by a given CH4/O2 ratio (or a minimum value that the CH4/O2 ratio must not fall below, respectively) in the inert zones to avoid the occurrence of homogeneous reactions, since to the lower right of this intersection the ignition delay becomes shorter than the residence time in the inert zones and hence homogeneous reactions can occur. One can see that for the given reactor geometry, homogeneous reactions can be avoided for all methane/air mixtures as long as inert zone temperatures remain below 1190 K. The contour plot at the bottom allows finding the critical residence time for any methane/air mixture and corresponding temperature and thus offers a detailed guideline for safe reactor design. As apparent from the experimental results in Figures 6 and 7, maximum temperatures in the inert zones in our laboratory reactor do in fact exceed the allowable temperatures as calculated above. However, the inert zones are also far from isothermal and show these

temperatures only in a very short zone right in front of the catalyst. Obviously, a detailed reactor design would have to take the experimentally observed temperatures profiles inside the inert zones into account. However, capping the allowable temperatures inside the inert zones below the isothermal temperature limits obtained in these simple simulations offers a conservative estimate and simple guideline for safe reactor design and operation. Finally, it should be pointed out that while the increased temperatures in the upstream inert zone of an RFR might pose concerns with regard to the occurrence of undesired homogeneous reactions and therefore require careful design, temperatures in the downstream inert zone are strongly decreased in the RFR in comparison with stationary reactor operation (see also Figure 6). While, particularly at steady-state operation with high flow rates, homogeneous reactions in the hot product gas downstream of the catalyst can be of concern (both for safety reasons and for reasons of selectivity loss), the rapidly decreasing temperatures in the downstream inert zone at reverse-flow operation leads to a quick quenching of any potential homogeneous reactions and thus ensures a better control of the reaction conditions. Catalyst Length. An issue that is often raised in discussions of high-temperature millisecond-contacttime catalysis is the question whether the thermodynamically allowed syngas yields would be obtained or at least approximated to a better degree if the catalyst contact time of the gases would be increased. The easiest way to achieve this is obviously a longer catalyst zone. (Reducing the total reactor gas flow is not an appropriate way to do this, since it will at the same time significantly alter the temperature profiles in the reactor, as already seen in our results above). We therefore investigated the catalytic partial oxidation of methane with varying length of the catalyst zone at steady-state and reverse-flow operation. The results are shown in Figure 9, where syngas yields are shown for steady-state reactor operation (“SS”, top row) and reverse-flow operation (“RFR”, bottom row) for a 1- and 3-cm-long catalyst zone, corresponding to a tripling of the catalyst contact time. One can see that for both catalyst lengths, RFR operation leads to similarly strong improvements in both syngas yields. Furthermore, for both catalyst length the improvements in syngas yields are particularly pronounced at high flow rates, as already explained above. Surprisingly, however, lengthening the catalyst zone, i.e., increasing the catalyst contact time leads to decreasing syngas yields, i.e., it has the opposite effect of what might be expected. This decrease is similar for steady-state operation and reverse-flow operation with roughly 10% loss in both CO and H2 yields. A hint to the reason for this unexpected result can be gained from the curves at stationary reactor operation: while for a 1-cm-long catalyst the maximum in syngas yields is achieved at a volumetric inlet flow rate of about 3 slm, this maximum seems to shift toward higher reactor through-puts with increasing catalyst length, i.e., for the 3-cm-long catalyst the curves just start to flatten out at 5 slm (our experimental setup was limited to 5 slm total gas flow). This shift in the curves toward higher flow rates can also be recognized in the results for the RFR, where the curves for the 1- and 3-cm-long catalyst appear to be parallel-shifted toward


Figure 11. Hydrogen and CO yields in the stationary reactor (“SS”, left graph) and in the reverse-flow reactor (“RFR”, right graph) for different length of the catalyst zone. The x-axis is a “rescaled” volumetric inlet flow, calculated by dividing the volumetric inlet flow by the square root of the catalyst length.

Figure 9. Syngas yields with varying volumetric inlet gas flow for a 1- and a 3-cm-long catalyst. Experimental results for stationary reactor operation (“SS”, top graphs) are shown in comparison to results at reverse-flow reactor operation (“RFR”, bottom graphs).

Figure 12. PFR simulation results with a detailed homogeneous reaction mechanism (GRI-mechanism 3.0) for a stoichiometric methane/air mixture (CH4/O2 ) 2) at T ) 1500 K. Even though the length of the reactor has been increased by more than 2 orders of magnitude in comparison to the experimental setup, equilibrium yields (which are >99% at these conditions) are not achieved.

Figure 10. Catalyst inlet (full symbols) and exit temperatures (open symbols) at stationary reactor operation (left graph, “SS”) and maximum catalyst exit temperatures at reverse-flow operation (right graph, “RFR”) with varying volumetric inlet gas flow for a 1-cm (diamonds and full lines) and a 3-cm-long catalyst (circles and dashed lines).

higher flow rates. Similarly, the temperature curves in both cases appear shifted toward higher flow rates. We tested this assumption by correlating the experimentally measured reactor yields with differently scaled functions of the inlet volumetric flow and the catalyst length and found a very good correlation with the volumetric flow divided by the square root of the catalyst length. This correlation is shown in Figure 11, where syngas yields at steady-state reactor operation (left graph) and reverse-flow operation (right graph) are plotted against this “rescaled volumetric flow rate”. The graph shows in addition to the data for 1- and 3-cmlong catalyst from Figure 10 also data for a 2-cm-long catalyst (omitted from the previous figure for reasons of better readability of the graph). One can see that this

rescaled data falls both for SS and RFR operation onto single curves, with some divergence for data at higher flow rates at SS operation. At this point, we do not have a consistent explanation for this observation. However, it seems likely that this rescaling stems from some form of complex interplay between convective cooling with increasing gas flow and conductive heat transport in the catalyst. Whatever the exact explanation for the observed trend is, the data demonstrate clearly that longer catalyst residence times do not help to achieve better syngas yields in this reaction system. To further support this experimental observation, we conducted a simple plug-flow reactor simulation with a detailed homogeneous reaction mechanism to confirm that the reaction system is indeed kinetically hindered from reaching thermodynamic equilibrium: Figure 12 shows hydrogen and CO yields along the reactor axis for a stoichiometric mixture of methane and air at a typical reaction temperature of 1500 K with detailed homogeneous chemistry (GRI-mechanism 3.0). The simulation results demonstrate that even if the reactor length is increased by as much as 2 orders of magnitude beyond the length of the reaction zone in these experiments, H2 and CO yields do not exceed ∼50% and ∼60%, respectively, while thermodynamic equilibrium yields are >99.8% in both cases. This also agrees well with a


previous simulation study with a detailed coupled heterogeneous-homogeneous reaction mechanism, which had also indicated that even in the presence of a Pt catalyst the reaction is virtually complete after only a few millimeter length into the catalyst bed.20 Thus, both experimental and simulation results demonstrate clearly that increasing the length of the reaction zones does not help kinetically, but rather appears to pose significant problems for the thermal operation of the reactor, resulting in decreasing syngas yields. Inert Zones. As beautifully explained by Eigenberger and Nieken,17 reverse-flow reactors can be rendered particularly efficient by introducing inert zones on either side of the catalyst bed. This elegant reactor design not only allows a significant reduction in the amount of (typically expensive) catalyst but furthermore allows complete decoupling of the kinetic and thermal aspects of the reactor design. As seen above, lengthening the catalyst bed to improve heat integration is not a viable option for high-temperature catalytic partial oxidation, and the introduction of inert zones as heat reservoirs for the regenerative heat exchange therefore becomes a necessity. The thermal capacity of these heat reservoirs is obviously another reactor design parameter with strong consequences for the efficiency of the reactor concept. High heat capacity material, longer inert zones, or denser structures in the inert zone will result in improved heat storage in these zones and thus generally result in improved reactor performance. The choice of suitable materials for inert zones at the extreme temperatures observed in high-temperature catalysis is, however, limited to certain high-temperature ceramics, as hightemperature steel or other metals show catalytic activity under these conditions and are therefore not suitable. We decided to test the influence of the heat conductivity of the inert zones in a particularly simple way: rather than change the material of the inert zones, we exchanged the 11-cm-long extruded cordierite monoliths against 11 1-cm-long pieces of the same material and shape. In this way, the inert zone is completely unaltered in every regard except the heat conductivity, which is now hampered through the suboptimal contact between the individual pieces of cordierite. (While segmentation of a monolith is also known to change the hydrodynamics of the gas flow, this should be of minor importance in the present case of the completely unreactive inert zones.21,22) Therefore, this configuration allows us to “simulate” the behavior of an inert zone with reduced heat conductivity (a disadvantage of this procedure is obviously that the exact value of the heat conductivity of this setup remains unknown). Figure 13 shows the experimental results for the single-piece inert zone (circles) versus the cut inert zone (triangles) at reverse-flow operation. As expected, the SS reactor shows identical results for the two cases. Therefore, only one curve is shown for SS (diamonds). One can see that the separation of the inert zone monoliths into several pieces leads to a small but significant improvement in both syngas yields by 5-8% over the whole range of methane/oxygen ratios. Apparently, the steeper temperature gradient that is to be expected for an inert zone with lowered heat capacity is beneficial to the efficiency of the heat integration. As the unchanged catalyst temperatures at RFR operation indicate, this improvement in heat integration is furthermore virtually completely converted into increased

Figure 13. Syngas yields (top) and reactor temperatures (bottom graph) versus methane/oxygen ratio in the reverse flow reactor with one 11-cm-long extruded cordierite monolith in each inert zone (circles) and 11 1-cm-long extruded monolith pieces in each inert zone (triangles) in comparison to results at steady-state operation of the reactor (squares). (Vin ) 3 slm, 45 ppi monolith, τ/2 ) 15 s)

selectivity. (Unfortunately, due to alignment problems with the cut monolith pieces, no measurement of temperature profiles in the piecewise inert zones was possible to demonstrate the increasing steepness of the temperature gradient). The fact that even such a small perturbation of the inert zones leads already to detectable changes in reactor performance indicates that the “engineering” of the inert zones is a crucial part of the design of a reverse-flow reactor. Furthermore, Eigenberger et al. have previously demonstrated some of the problems that can occur when operating with ceramic monoliths at very high temperature and with steep temperature gradients:23 while most appropriate ceramics withstand high temperatures fairly well, ceramics tend to be very brittle and do not withstand spatial or temporal gradients well. The use of short individual monolith pieces in the inert zones allows one to counter any potential problems of this kind, since shorter pieces of material tend to accommodate the same local thermal stress significantly better than longer pieces. Frequency of Flow Switching: General Considerations. Obviously, one of the most important reactor operating parameterssand a truly unique parameter to reverse-flow operationsis the frequency or periodicity of the flow switching. One problem associated with the flow switching in an RFR is the fact that at each flow reversal, unconverted reactants that were located in the upstream inert zone are flushed out of the reactor and mixed with the product gas. This unconverted gas thus caps the maximum attainable conversions in a reverse-flow reactor. To avoid this problem, RFRs are typically operated with


an additional “flushing phase” in which the unconverted reactants are pushed through the catalytic zone and thus converted by flowing an inert gas through the reactor. While this avoids the mentioned problem of unconverted reactants in the product stream (and also ensures well-defined reactant concentrations inside the reactor at each point in time inside the reactor itself), it also results in a dilution/contamination of the product stream by the inert gas, and it furthermore complicates an already rather complex mode of reactor operation. We therefore operated our reactor without a flushing phase to keep the process as simple as possible. However, this results in an intrinsic upper limit of achievable reactant conversions, which is dependent on the relative length of one half-period and the residence time of the gases inside one inert zone. This limit is independent of the individual reaction system and can be calculated as follows (for simplicity shown here for the example of a reaction with instantaneous, full conversion of reactants)

Xmax)

V˙ (τP/2-τI) V˙ (τP/2-τI)+V˙ τI

)

2τI τP/2-τI )1 τP/2 τP

where V˙ denotes the volumetric flow rate of the respective reactant; τP/2 is the duration of one half-period, i.e., the length of time between two subsequent flow switchings; and τI is the residence time of the reactant inside one inert zone. The numerator of the first term thus denotes the amount of reactants converted, which for a “true” (kinetic) conversion of 100% is equal to the amount of reactant flowing through the reactor flowing through the catalyst minus the amount that is present in the upstream inert zone at flow reversal. The denominator sums converted plus unconverted reactants, i.e., the amount in the numerator plus the amount of (unconverted) gas in the inert zone. The resulting conversion thus indicates the maximum obtainable conversion (i.e. overall or “nominal” process conversion) for a reaction with a kinetic conversion of 100%. (An extension onto a reaction system with a conversion of less than 100% is straightforward and hence not given here). From this formula, one can already immediately deduce one limiting case for RFR operation: for τP/2 ) τI, the maximum attainable conversion drops to zero, since the reactants never actually reach the catalyst zone, but their flow direction is reversed while they are still in the upstream inert zone (for values τP/2 < τI the formula above yields negative values, which should be replaced by zero, i.e., XRFR ) max(Xmax,0)). This is an obvious limiting case. For sufficiently fast reactions, this limit can in principle be shifted toward very short switching periods by operating the RFR with a completely catalytically active fixed bed, i.e., by omitting the inert zones. However, this case is not only typically uneconomic, due to the high price of catalyst, butsas seen abovesnot efficient for the current (and many similar) reaction systems. Figure 14 shows the above deduced maximum attainable conversion against volumetric flow rate and (full) periodicity for the dimensions of our laboratory RFR. For better visibility, volumetric flow rates start at 0.1 slm and reactor periodicity starts at 1 s (as apparent from the above equation, Xmax would drop toward zero for either of the two quantities going toward zero). One can see that overall process conversions drop rapidly toward very short switching frequencies and

Figure 14. Maximum achievable methane conversions versus periodicity of flow reversal (duration of a full period) and volumetric inlet flow in a reverse-flow with dimensions of our laboratory reactor. One can see that for short switching periods and low volumetric flow rates the flow reversal (without a flushing phase) leads to significant losses in theoretically achievable conversion.

very low flow rates. For the typical reactor switching periodicity of τ/2 ) 15 s used in our experiments, maximum attainable conversions vary from 95% for V ) 1 slm to 99% at 5 slm. Thus, while the effect of flushing out unconverted reactants is almost negligible at high flow rates, it has a significant impact on conversions at lower flow rates. One can also see that the efficiency of the process increases quickly with increasing duration of the periodicity. For example, doubling the duration of a half-period from 15 to 30 s increases the maximum attainable yields to 97.5% for V ) 1 slm and to 99.5% at 5 slm. Clearly, high flow rates and a long switching period must be the aim of reverseflow operation! It should finally be mentioned that process selectivities are not affected by this effect as long as no (desired or undesired) product gases are present in the reactor feed. However, absolute concentrations of products in the product stream will of course be reduced due to the presence of unconverted reactants. Frequency of Flow Switching: High-Temperature Partial Oxidation. Generally, two limiting cases exist for the switching periodicity in an RFR: the limit of infinitely slow switching, which ultimately results in simple steady-state operation of the reactor, and the limit of infinitely fast switching, which results in a stagnating gas flow and hence extinction of the reaction (as seen above). For efficient reactor operation one therefore needs to find the optimal switching frequency between these two limits at which maximum process yields can be obtained. We tested the dependency of process yields on the RFR periodicity for a Pt-coated 45 ppi monolith and a Rh-coated 45 ppi monolith. Results are shown in Figure 15 (note the logarithmic x-axis). One can see that process yields are indeed strongly depending on the periodicity of the reactor operation, with a pronounced maximum in yields for the Ptcatalyzed process at a duration of one half-period (i.e. the length of time during which the gas flows in one direction) of τ/2 ≈ 10 s. Furthermore, it is interesting to note that process yields drop quite precipitously toward shorter periodicities (by about 10% from τ/2 )


Figure 15. Syngas yields versus periodicity of flow reversal (duration of a half period) in the reverse-flow reactor for a Ptcoated (left) and a Rh-coated (right) 45 ppi catalyst at CH4/O2 ) 2 and total inlet gas flow of 3 slm.

10 to 2 s), while the decrease toward longer periodicities is comparatively flat (by less than 10% from τ/2 ) 10 to 100 s). The decreasing trends toward short switching periods can be explained by the rather slow thermal dynamics of the reverse-flow reactor in comparison to the extremely fast high-temperature catalytic kinetics, which results in an inefficient utilization of the heat capacity of the inert zones at very fast switching periods and hence results in a strong reduction of the thermal efficiency of the reactor. The decrease is further accelerated by the above-mentioned decreasing maximum attainable process conversions due to the back-flushing of unconverted reactants, which explains the steepness of this drop. Toward longer switching periods, the decreasing trend is due to the fact that the heat capacity of the inert zones now has been fully exhausted; further lengthening of the periods therefore does not lead to any further improvement in heat integration, and the process yields therefore start to decline toward the limit of the stationary process. Clearly, increasing the heat capacity of the inert zones through the use of higher heat capacity material or simply an increase in the mass of the inert material (i.e. denser monolith structures) will allow one to extend the switching frequency in the reverse-flow process toward longer periods. This can be of significant practical importance, since it significantly reduces the mechanical stress and thus extends the lifetime of the valves that enable the flow reversal. Comparing the results for the Pt-coated monolith with results from a study using a Rh-coated monolith (Figure 15, right graph), one observes that the optimum in syngas yields is shifted toward longer switching periods, from τ/2 ≈ 10 s for Pt to τ/ 2 ≈ 20 s for Rh. Furthermore, syngas yields are much higher over Rh than over Pt (about 70% vs about 50%, respectively). The fact that Rh is a better catalyst for syngas production than Pt has been observed previously by Hickman and Schmidt.24 The authors explained the difference between these catalysts by the differences in the potential energy surfaces of the reactions over these catalysts, in particular the higher activation energy for OH formation on Rh surfaces.1 Clearly, heat integration does not change these characteristics. In general, both the fact that the relative order of the catalysts is unchanged and the fact that for both catalysts the maxima in CO and H2 yields occur in parallel indicate that the process improvements via reverse-flow opera-

tion occur only due to the heat integration (i.e. thermal aspects of the reactor operation) rather than any influence of the unsteady-state operation on the reaction kinetics or mechanism. The shift toward longer periodicity with Rh can be explained by the reduced amount of heat that evolves in the more selective reaction over that catalyst. As explained above, the decrease in syngas yields toward longer switching frequencies is due to an “exhaustion” of the heat reservoirs offered by the inert zones. A more selective reaction, however, results in lower temperatures of the product gases, and therefore, it takes longer to “fill” the heat reservoirs. This allows operating the reactor at longer switching periods, which, as explained above, is not only beneficial for the lifetime of the switching valves but also reduces losses in reactant conversion due to flushing out of unconverted reactants. It seems interesting to note that an improved reaction leads therefore to improved or facilitated process operation. As a result, it seems that development of improved reactor concepts should go hand-in-hand with the development of improved and more selective catalysts for a reaction. Summary We have presented results from an investigation into some of the main reactor design and operating parameters for a reverse-flow reactor for high-temperature catalytic reactions. Length and structure of the catalyst and the structure of the inert zones were varied, the importance of homogeneous reactions was discussed, and the periodicity of flow reversal as a unique parameter in RFR operation was studied. We demonstrated that maximum temperatures in a high-temperature catalytic RFR are not necessarily a problem, with maximum temperatures at the catalyst ends in any experiment reported here being below 1300 °C. Maximum temperature could furthermore be reduced by more than 200 °C through the use of denser catalyst structures. Generally, denser catalyst structures turned out to be beneficial not only at reverseflow operation but also for steady-state operation. The combination of relatively high solid density with lowpressure drop and well-mixed fluid flow therefore makes foam monoliths ideal support structures for high-temperature reactions for both types of reactor operation.25 Furthermore, our results showed that local temperature variations during one (semi-)cycle at periodic steadystate are also well-behaved, with maximum temperature differences remaining below 200 °C at any point in the reactor. To our surprise, we found that extending the length of the catalyst zone not only does not improve reaction yields but in fact leads to reduced yields. This result is in qualitative agreement with simulation results that show that even at these high temperatures the partial oxidation of methane is kinetically hindered from reaching equilibrium syngas yields. While we found that our results scale with the ratio of volumetric flow rate over the square root of the catalyst length, we currently have no consistent explanation for this scaling. It does, however, seem to indicate the importance of convective heat transport in the gas phase versus conductive heat transport in the solid phase and thus the importance of the “thermal design” of a RFR. Finally, we showed that the optimal periodicity for reverse-flow operation is dependent on the catalyst used,


with more selective catalysts leading to longer optimal switching periods. This is beneficial not only for the lifetime of the flow-switching valves but also allows operating the RFR without a flushing-phase between half-periods, thus greatly simplifying reactor operation. Overall, we see reverse-flow reactors and high-temperature catalytic reactions as a perfect match: reverseflow operation not only allows one to overcome the intrinsic autothermal limitations of high-temperature partial oxidation reactions12 but it also benefits from the high flow rates that can are possible in high-temperature catalysis due to the extremely high reaction rates. Furthermore, the steep temperature gradients in the inert zones on either side of the catalyst help to quench potential homogeneous reactions in the hot product gas downstream of the catalyst and thus help to localize the reaction in the catalyst zone. Finally, the compactness of this integrated reactor concept perfectly complements the compactness of high-temperature short-contact-time reactors. Acknowledgment Financial support by the Deutsche Forschungsgemeinschaft (DFG), the “Verband der Chemischen Industrie” (VCI), and the Global Engineering Education Exchange program (Global E3) are gratefully acknowledged. This paper is dedicated to G. Eigenberger on the occasion of his 65th birthday. Literature Cited (1) Hickman, D. A.; Schmidt, L. D. Production of syngas by direct catalytic oxidation of methane. Science 1993, 259(5093 Jan 15), 343-346. (2) Huff, M.; Torniainen, P. M.; Schmidt, L. D. Partial oxidation of alkanes over noble metal coated monoliths. Catal. Today 1994, 21(1 Aug 30), 113-128. (3) Schmidt, L. D.; Huff, M.; Bharadwaj, S. Partial Oxidation Reactions and Reactors. Chem. Eng. Sci. 1995, 49, 3981-3994. (4) Goetsch, D.; Schmidt, L. D. Microsecond Catalytic Partial Oxidation of Alkanes. Science 1996, 271, 1560-1562. (5) Aasberg-Petersen, K.; Hansen, J.-H. B.; Christensen, T. S.; Dybkjaer, I.; Christensen, P. S.; Nielsen, C. S.; Madsen, S. E. L. W.; Rostrup-Nielsen, J. R. Technologies for large-scale syngas conversion. Appl. Catal. 2001, A221, 379-387. (6) Ashcroft, A. T.; Cheetham, A. K.; Foord, J. S.; Green, M. L. H.; Grey, C. P.; Murrell, A. J.; Vernon, P. D. F. Selective oxidation of methane to syngas using transition metal catalysts. Nature 1990, 344, 319-321. (7) Basini, L.; Damore, M.; Fornasari, G.; Matteuzzi, D.; Sanfilippo, D.; Trifiro, F.; Vaccari, A. Syngas production by partial oxidation of methane: dependence of reactivity on catalyst properties and contact time. Stud. Surf. Sci. Catal. 1997, 107, 429-434.

(8) Bharadwaj, S. S.; Schmidt, L. D. Catalytic Partial Oxidation of Natural Gas to Syngas. Fuel Proc. Technol. 1995, 42, 109-128. (9) Hickman, D. A.; Schmidt, L. D. Syngas Production by Direct Oxidation of Methane on Monoliths. J. Catal. 1992, 138, 267282. (10) Veser, G.; Frauhammer, J.; Friedle, U. Syngas formation by direct oxidation of methane. Reaction mechanisms and new reactor concepts. Catal. Today 2000, 61(1 Aug), 55-64. (11) Friedle, U.; Veser, G. Counter-current heat-exchange reactor for high-temperature partial oxidation reactions. I. Experiments. Chem. Eng. Sci. 1999, 54(10), 1325-1332. (12) Neumann, D.; Veser, G. Dynamic Reactor Operation and Short Contact-Time Catalysis: Methane Partial Oxidation in a Reverse-Flow Reactor. AIChE J. Submitted. (13) Boreskov, G. K.; Matros, Y. S. Flow reversal of reaction mixture in a fixed catalyst bedsA way to increase the efficiency of chemical processes. Appl. Catal. 1983, 5(3), 337-343. (14) Boreskov, G.; Matros, Y. Unsteady-State Performance of Heterogeneous Catalytic Reactions. Catal. Rev.-Sci. Eng. 1983, 25, 551. (15) Eigenberger, G.; Nieken, U. Catalytic combustion with periodic flow reversal. Chem. Eng. Sci. 1988, 43(8), 2109-2115. (16) Nieken, U.; Kolios, G.; Eigenberger, G. Control of the ignited steady state in autothermal fixed-bed reactors for catalytic combustion. Chem. Eng. Sci. 1994, 49(24B), 5507-5518. (17) Eigenberger, G.; Nieken, U. Catalytic cleaning of polluted air: reaction engineering problems and new solutions. Int. Chem. Eng. 1994, 34(1 Jan), 4-16. (18) Kolios, G.; Frauhammer, J.; Eigenberger, G. Autothermal Fixed-Bed Reactor Concepts. Chem. Eng. Sci. 2000, 55(24), 59455967. (19) Neumann, D.; Veser, G. Numerical simulation of catalytic oxidation of methane in a reverse-flow reactor. Manuscript in preparation. (20) Veser, G.; Frauhammer, J. Modelling steady state and ignition during catalytic methane oxidation in a monolith reactor. Chem. Eng. Sci. 2000, 55(12), 2271-2286. (21) Andersson, S. L.; Scho¨o¨n, N. H. Methods to increase the efficiency of a metallic monolithic catalyst. Ind. Eng. Chem. Res. 1993, 32, 1081. (22) Holmgren, A.; Andersson, B. Mass transfer in monolith catalystssCO oxidation experiments and simulations. Chem. Eng. Sci. 1998, 53, 2285. (23) Frauhammer, J.; Eigenberger, G.; von Hippel, L.; Arntz, D. A new reactor concept for endothermic high-temperature reactions. Chem. Eng. Sci. 1999, 54(15-16), 3661-3670. (24) Hickman, D. A.; Haupfear, E. A.; Schmidt, L. D., Synthesis Gas-Formation by Direct Oxidation of Methane over Rh Monoliths. Catal. Lett. 1993, 17(3-4), 223-237. (25) Schmidt, L. D.; Dietz, A., III In Monoliths for partial oxidation catalysis, Synthesis and Properties of Advanced Catalytic Materials, 1995; Materials Research Society Symposium Proceedings: 1995; p 299-307.

Received for review November 18, 2003 Revised manuscript received April 8, 2004 Accepted April 12, 2004 IE034257S

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