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1 Chemical Engineering Science 65 (21) Contents lists available at ScienceDirect Chemical Engineering Science journal homepage: Catalytic partial oxidation of ethane to ethylene and syngas over and coated monoliths: Spatial profiles of temperature and composition Brian C. Michael, David N. Nare, Lanny D. Schmidt Department of Chemical Engineering and Materials Science, University of Minnesota, 421 Washington Ave. S.E., Minneapolis, MN 55455, USA article info abstract Article history: Received 7 October 29 Received in revised form 1 March 21 Accepted 22 March 21 Available online 25 March 21 Keywords: Catalytic partial oxidation Oxidative dehydrogenation Steam reforming Ethane Ethylene Syngas Spatial profile measurements Thermodynamics The catalytic partial oxidation of C 2 H 6 over and coated monolithic supports (4.7 wt% M/a-Al 2 O 3 45 PPI) was investigated with a capillary sampling technique for a range of C 2 H 6 /air ratios at constant inlet flow (8 ms contact time), with and without addition. Effluent data clearly indicate the differences in product distribution between catalysts and equilibrium. effectively converts the reactant mixtures to syngas with 8% selectivity, whereas produces with 55% C-atom selectivity, while neither produces thermodynamically favored C. Spatially resolved measurements provide direct evidence of the multi-zone nature of the reactors. With, complete conversion of O 2 occurs to produce mostly, and within the first 3 mm of catalyst, followed by a reforming zone to produce additional syngas. consumes O 2 more slowly, which results in a steady increase in temperature along the reactor. Ethylene formation correlates to reactor temperatures 475 1C, regardless of, in line with the onset of homogeneous reactions. Hydrogen addition tests (C 2 H 6 /O 2 / ¼2/1/2) clearly exhibit preferential oxidation of with O 2 over, which shifts the maximum in temperature upstream while preserving a portion of the C 2 H 6 for production. addition modifies the concentration and temperature profiles minimally on. The main differences between catalysts are the high reforming and O 2 consumption activity with compared to, which are likely responsible for differences in yields. & 21 Elsevier Ltd. All rights reserved. 1. Introduction The efficient utilization of light hydrocarbons becomes increasingly important as liquid fossil resources continue to decrease. One strategy for efficient use is to first convert hydrocarbons into syngas (mixtures of and ), which can then be processed into chemicals and fuels by a variety of technologies. Importantly, since light hydrocarbons often occur as mixtures, either within existing processes or from natural gas, methods for syngas production must be demonstrated to be effective for each of the components. Alternatively, separation of the components provides avenues for direct conversion into valuable chemicals. Several routes have been explored for upgrading the CH 4 component, such as oxidative coupling, and established technologies exist for C 2 H 6 conversion to. With current world production estimated to be over 11 million tons, is the most important and basic building block in the petrochemical industry (Albright et al., 198; Ren et al., 27), so its efficient production is paramount. Corresponding author. address: schmi1@umn.edu (L.D. Schmidt). Current technologies for syngas and production share a few characteristics. Ethylene is produced in large steam crackers (Sundaram et al., 2), which carry out endothermic pyrolysis (Eq. (1) in the presence of steam, while syngas is commonly produced in steam reforming (Eq. (2)) reactors packed with Ni-based catalysts: C 2 H 6 - þ, DHrxn ¼ 136 kj=mol ð1þ C 2 H 6 þ2-2þ5, DHrxn ¼ 347 kj=mol ð2þ Both reactors are large, heat-integrated equipment that require external heating, and relatively long residence times (1 s). This characteristic means that they are only economical on very large scales, and cannot be easily scaled for remote (e.g. stranded gas) applications. In the early 199 s, Huff and Schmidt (1993) demonstrated that syngas could be produced in short contact times (5 ms) over based catalysts by adding O 2 to the feed. With an identical reactor setup, instead using catalysts, could be produced with yields 45%. Both of these processes operate without external heat input because the overall reactions are exothermic. Catalytic partial oxidation (CPO) to syngas (Eq. (3)) requires C 2 H 6 /O 2 ratio of 1, while oxidation to (called oxidative 9-259/$ - see front matter & 21 Elsevier Ltd. All rights reserved. doi:1.116/j.ces

2 3894 B.C. Michael et al. / Chemical Engineering Science 65 (21) dehydrogenation, ODH) (Eq. (4)) requires C 2 H 6 /O 2 ¼2: C 2 H 6 þo 2-2þ3, DHrxn ¼ 136 kj=mol ð3þ C 2 H 6 þ:5o 2 - þ, DHrxn ¼ 15 kj=mol ð4þ The use of sacrificial co-fed fuels has also been shown to increase yields, particularly when the catalyst is judiciously modified. When was added to the feed over a Sn catalyst, 85% selectivity to was obtained at 47% conversion (Bodke et al., 1999), making ODH comparable to a steam cracker without the limitations noted with the conventional route. Additionally, the byproduct can be separated and recycled to ODH systems using LaMnO 3 perovskites to increase selectivity (Donsı et al., 25a). With addition, the reaction was hypothesized by Bodke et al. (2) to be a two-zone process with oxidation in the first part of the catalyst, and pyrolysis downstream giving the overall ODH reaction (Eq. (4)). The demonstration of industrially relevant yields by ODH made it a focus of much of the research in C 2 H 6 CPO; however, the mechanism of production is still a subject of some debate. Huff and Schmidt (1996) developed a mechanism consisting purely of catalyzed surface steps, which fit experimental data reasonably well. In contrast, Beretta et al. (21) suggested gas phase free radical chain mechanisms were sufficient to produce in short times once the initial radical pool was formed. This work agrees with observations from low pressure temporal analysis of products (TAP) reactors, which showed no evidence of formation up to 6 1C on, whereas was produced under atmospheric pressure (Silberova et al., 23). Other researchers propose a heterogeneous homogeneous reaction scheme, suggesting that the catalyst was responsible for oxidizing surface C and H species to produce localized heat that drives endothermic homogeneous reactions (Beretta and Forzatti, 21; Vincent et al., 28; Zerkle et al., 2). Part of the reason why uncertainty still exists is the lack of data that shows how products evolve inside the catalyst at relevant operating conditions. These measurements are necessary for validating mechanisms that predict product evolution when incorporated into sound reactor models. Such measurements are elusive because of the complexities associated with short contact time reactors (SCTRs): high reaction temperatures (up to 1 1C), extreme thermal gradients (up to 2 1C/mm) and the possibility of gas phase reactions. Some attempts have been made using thermowells as well as infrared cameras to collect axial temperature profiles (Basile et al., 21; Basini et al., 21; Marengo et al., 23), but these experiments can introduce heat loss by conduction and radiation compared to well-insulated CPO reactors unless special attention is given to the apparatus (Simeone et al., 28). Henning and Schmidt (22) used multiple side ports in a SCTR to find species and temperature profiles downstream of the catalyst, revealing that substantial reaction occurs downstream of the catalyst. Although this work provides useful and direct evidence for gas phase reactions, it was not focused on measuring species within the bed. More recently, Horn et al. (26) developed a method for sampling species within monolithic foam supports by inserting a small, modified capillary into a channel which allows for access to all locations along the axis of a SCTR. A similar system has also been developed by investigators at Oak Ridge National Labs (Choi et al., 25), and is commercially available (SpaciMS). The majority of the investigations using these spatially resolved techniques are related to CH 4 CPO, although one C 2 H 6 CPO profile was reported for a single fuel/oxygen ratio (Horn et al., 26), and a recent work reports spatial profiles of propylene oxidation (Shakir et al., 29). In this paper, we expand on the early work by Horn et al. by investigating the phenomena of C 2 H 6 CPO for a range of fuel/oxygen ratios, over and catalysts, and with the use of co-feed. The goal of this work is primarily two-fold. First, by acquiring spatial species and temperature profiles, we attempt to provide direct evidence of phenomena predicted by theoretical work, and second, we attempt to show how and catalysts ultimately yield different effluent product distributions owing to spatial variation in their oxidation and reforming activity. 2. Experimental To gain insight into the phenomena occurring within an operating reactor, an apparatus was constructed that was based on the one used by Horn et al. (26). The apparatus consisted of a feed system, reactor, and sampling system specifically adapted to sample species expected in the C 2 H 6 CPO system. A schematic, and other details not described here have been reported elsewhere (Michael et al., 29) Catalyst preparation The catalysts were supported on a-al 2 O 3 reticulated foam cylindrical monoliths with 45 pores per linear inch, nominally measuring 17 mm in diameter by 1 mm long. Prior to deposition with catalyst precursors, the monoliths were drilled with an ultrasonic drill to create a hole 75 mm in diameter along the center axis. After drilling, the monoliths were repeatedly rinsed in acetone to remove drilling debris and calcined at 9 1C for 6 h. The catalysts were prepared by wetting the monoliths to the incipient point with an aqueous solution of (NO 3 ) 3 (Alfa Aesar) containing enough atom to produce a final metal loading of 4.7 wt%. The monoliths were then dried in air over night, and heated to 13 1C in air for 2 h, after which they were calcined at 6 1C for 2 h. The catalysts were prepared similarly, following the recipe described by Bodke et al. (2). For these catalysts, the was deposited from Cl 6 solutions (Alfa Aesar), and heat treated as described above. Both catalysts were treated in flowing (5% in Ar at 5 sccm) at 25 1C for 2 h prior to use in the CPO reactor. Although extensive characterization was not conducted in the present work, other studies by our group (Bodke and Schmidt, 1999; Bodke et al., 2; Degenstein et al., 26) have shown that the metal exists on the low surface area support (2m 2 /g) as a broken film prior to reaction, which can transform into spherical particles during reaction in the absence of a stabilizing washcoat. To avoid variability in performance owing to this transformation process, the catalysts were operated at the most extreme conditions for 2 h prior to collecting profile data. In general, the integral reactor results are insensitive to metal loadings above 1%, as well as slight variations in the preparation procedure (Bodke et al., 2) Reactor configuration The reactor consisted of a quartz tube (2 mm I.D. by 25 mm long) containing a catalyst-coated monolith between two blank monoliths, which were wrapped in aluminosilicate cloth to hold the monoliths in place and to prevent gas bypass. A quartz capillary was introduced 12 mm downstream of the monolith assembly through a septum-sealed side port. A second quartz capillary was fed through an inlet reactor cap and the aligned holes of the monoliths. This capillary was sealed at the downstream tip and received a 15 mm orifice 4 mm upstream from the tip for gas sampling. The entire reactor was wrapped in insulation to prevent heat loss, and typically achieved 95%

3 B.C. Michael et al. / Chemical Engineering Science 65 (21) adiabatic efficiency defined as a ¼ T ad T in T obs T in where T ad is the adiabatic temperature calculated for the observed products, T in is the inlet temperature, and T obs is the observed outlet temperature. A more rigorous assessment of the adiabatic efficiency would require knowledge of the effect of heat loss on the observed products; hence the metric used herein is likely an upper bound on the efficiency. Gases were fed to the reactor with mass flow controllers (Brooks 585) through a pre-heater to maintain a constant inlet temperature of 1 1C at a total inlet flow of 5 SLPM for all tests. All of the gases were CP grade or better and used as-supplied. For most tests, air was used as an oxidant to moderate the adiabatic temperature, and was simulated by feeding O 2 and Ar at a ratio of 21: Profile acquisition and sampling considerations Three measurements were taken along the axis of the steadystate CPO reactor: gaseous species flow using the orifice capillary described above, temperature via fiber optic pyrometry, and temperature via thermocouple. For the first measurement, gases were pulled through the sampling capillary, which was connected to a tee and a microvolume needle valve, to a quadrupole mass spectrometer (QMS, SRS RGA2). The line connecting the needle valve to the QMS was a heated 1/16 316L stainless steel line with an internal bore large enough to allow flow control with the needle valve. Gas analysis was performed by classical least squares multiple regression of the mass spectrum using a range of m/z from 2 to 46. Calibration of the QMS was performed prior to analysis using Ar as an internal inert reference for m/z¼4 and was validated with an independent calibration gas mixture. The species considered for regression were,, CH 4, 2,C 2,,C 2 H 6,O 2, and 2. Water was determined by closing the O atom balance. Typically, C atom balances closed to within 7% along the axis of the reactor. The catalyst surface temperature was estimated by fiber optic pyrometry. A quartz fiber with a 451 polished tip was inserted into the axial capillary, which transmitted radially directed radiation to a single wavelength IR thermometer (Mikron MI-GA 5-LO). This device was calibrated in a furnace. Gas temperatures were estimated by inserting a 254 mm K-type thermocouple into the capillary. Each of the three measurements was taken independently. To obtain a spatial profile, the capillary probe (either the orifice, thermocouple tip, or optical fiber tip) was positioned at the front face of the monolith assembly, then a measurement was taken after a dwell of 1 s and incremented.2 mm to the next position. This process was repeated until the entire monolith assembly was measured. For each feed mixture, a profile was taken, along with an analysis of the effluent at the side port capillary via GC (HP 589, 3 HayeSep D, TCD). Profiles were repeated for at least three monoliths, and representative profiles are presented here. Particular attention was placed on the correct representation of the species measurements via the axial capillary sampling system since the possibility of homogeneous reactions is high in the C 2 H 6 CPO system. Sampled gases enter the capillary and flow counter currently in the absence of a catalyst, but at temperatures that are sufficient to support homogeneous chemistry. For this reason, the residence time of the gases in the sample flow must be as small as possible, and must be quenched as rapidly as possible for the mixture to be representative of only the reactor. ð5þ To optimize the sample flow through the capillary (and therefore the residence time), the orifice was placed at the same position as the side sample capillary during a steady state reactor run. The flow through the axial capillary was then increased with the needle valve while the conversion was monitored by the QMS until it matched the side sample conversion to within 5%. The capillary flow was then estimated ex situ with a small bubble column, and was found to be 1 sccm, or approximately 2% of the total reactor flow. Although this procedure does reduce the effect of capillary reactions, it is likely that reactions in the capillary cannot be completely eliminated. For this reason, the conversions depicted in the spatial profiles represent upper bounds for the true values. 3. Results and discussion 3.1. Integral reactor performance To make an overall comparison between the and catalysts in C 2 H 6 CPO, the integral reactor performance was measured by GC at a position 42 mm downstream of the front face of the monolith assembly. Reactant conversion, H atom selectivity, and C atom selectivity as a function of inlet C 2 H 6 :O 2 ratio (C/ O) with O 2 :Ar of 21:79 is depicted in Fig. 1. Selectivity is defined here as n i,j F out j S i,j ¼ P j a C 2 H 6 n i,j Fj out ð6þ where S i,j is the i atom selectivity to species j, n i,j is the number of i atoms in species j, Fj out is the outlet flow of species j. The product distributions depicted in Fig. 1 are consistent with the results published in other works on C 2 H 6 CPO with diluted feeds (Beretta et al., 21; Beretta and Forzatti, 24; Huff and Schmidt, 1993). Ethane conversion decreases monotonically from 98% to 41% over the catalyst and similarly decreased from 86% to 41% on with increasing from 1 to 2. Oxygen conversion is 499% over for every ; however, it decreases from 97% to 85% with for. In contrast to the similar trends in conversion, a clear difference in the product spectrum for the two catalysts exists. The catalyst converts C 2 H 6 primarily to syngas, with minor products being 2,, and. Platinum yields considerable quantities of and comparable selectivity to and (3% on average). At low, is favored, while at high, temperature drops, and selectivity to 2 increases, thus a maximum is exhibited in the selectivity. The conversions and selectivities shown in Fig. 1 are lower than those expected for CPO of C 2 H 6 with un-diluted feeds (Bodke et al., 1999, 2; Huff and Schmidt, 1993; Yokoyama et al., 1996). Detailed dilution tests have been conducted in other studies and have shown that inert content is a sensitive parameter in the reactivity of C 2 H 6 /O 2 mixtures in CPO (Bodke et al., 2; Donsı et al., 25a). While dilution serves to moderate the reactor temperature, it also lowers the average reaction rate. Since the air oxidizer employed here results in an inert content of 65 55% for of 1 and 2, respectively, these effects are likely responsible for the lower conversion. In addition, the location of the sampling point used here may be farther upstream from that of other works, which is another important parameter to consider when comparing performance, since reactions can continue downstream of the catalyst (Henning and Schmidt, 22) Equilibrium considerations In addition to the experimental values, results from equilibrium calculations for the same feeds are plotted in Fig. 2. These

4 3896 B.C. Michael et al. / Chemical Engineering Science 65 (21) Conversion 1% 8% 6% 4% C 2 H 6 O 2 S C 1% 8% 6% 4% 2% 2% % 2. % 2. S H 1% 8% 6% 4% 2% S C 1% 8% 6% 4% 2% % 2. % 2. S H 1% 8% 6% 4% 2% % 16% 14% 12% 1% 8% 6% 4% 2% % 2. S C C 2 CH Fig. 1. Reactor performance measured at z¼42 mm for inlet flow of 5 SLPM, T in ¼1 1C over 4.7 wt% /45 PPI a-al 2 O 3 (filled markers) and 4.7 wt% /45 PPI a-al 2 O 3 (open markers). calculations were performed in the HSC Chemistry software package (Oy, 27) with constant enthalpy and constant pressure constraints, rather than constant temperature and pressure, in accord with the low heat loss of the experimental reactor. The species included in the calculations were Ar,,,, 2, C (solid),c 2,,C 2 H 6,CH 4, and O 2. Neither of these catalysts yield effluent compositions that agree entirely with the equilibrium mixtures. Chemical equilibrium predicts considerable selectivity to solid C at 4; however, no solid C accumulation was observed in these tests. In deficient O 2, decomposition of C 2 H 6 results in solid C and in addition to, 2, and. Close examination of the effluent data for shows that the quantities of, 2 and are similar to those for equilibrium, while the difference in and solid C correspond to the incomplete conversion of C 2 H 6. Additionally, can be an intermediate in the decomposition of C 2 H 6 toward equilibrium, which is seen in the experimental results, but not the calculations. The effluent data, therefore, represent products that are dictated by competing kinetic processes, such as C C and C H bond cleavage, although gives products similar to equilibrium. To understand how the reactants progress to the products observed at the effluent, and what processes may dictate the differences between catalyst type and equilibrium, spatial reactor profiles of temperature and species flow were acquired for the same feeds as in Fig Spatial species profiles at low For either catalysts, no reaction is observed within the noncatalytic front heat shield (ozo9 mm), as indicated by the constant flow of C 2 H 6 and O 2 in that region of Figs. 3a and b. In the catalytic region, the reaction begins abruptly, but yields differences in product distribution between the two catalysts. For the catalyst, the major products in the first few millimeters are and, whereas the major products for are and. In addition to the product spectrum, the rate of reaction is different for the two catalysts; more rapidly consumes O 2 (to completion in the first 3 mm), whereas C 2 H 6 and O 2 react slower over with incomplete O 2 conversion. The O 2 flow curve in Fig. 3a also displays a transition between a high and low consumption rate, which is not present in Fig. 3b. Nonetheless, the relatively rapid reaction in the first parts of the two catalysts are exothermic, and appear to be largely responsible for the differences in,, and species observed at the effluent (i.e. in Fig. 1). Further downstream, a second distinct region exists with either catalysts; however, the overall reaction is, again, catalyst dependent. The rapid consumption of O 2 on the catalyst produces intermediate, which in the second region reacts with the remaining C 2 H 6 to form additional and. This reaction proceeds with a stoichiometry approximately equal to steam reforming (Eq. (2)). Conversely, after approximately 5% of

5 B.C. Michael et al. / Chemical Engineering Science 65 (21) S H X or S C 1% 8% 6% 4% 2% % 1% 8% 6% 4% X ethane X oxygen C (s) CH conversion is roughly at the same position, suggesting that O 2 conversion before and after this point is dominated by different processes. In addition to surface catalyzed routes, gas phase reactions to and may exist downstream. Although most of the O 2 is converted in the center monolith, nearly 5% of the overall C 2 H 6 conversion takes place in the back heat shield, along with a substantial fraction of the production. In contrast, essentially no reaction can be observed downstream of the coated monolith (z419 mm in Fig. 3d). Henning and Schmidt (22) observed similar activity downstream of a catalyst under ODH conditions, namely that 2% and 5% of the overall O 2 and C 2 H 6 conversions (respectively) occurs up to 8 cm downstream, along with production of. In that work, a back heat shield was not employed, so the observed activity was assigned to gas phase reactions; however, in the current work, the back heat shield may be subject to unintentional deposition from the upstream monolith. The close agreement with Henning et al. suggests that, if did exist in the back heat shield herein, the contribution to the overall reaction is small. The second region of the two catalysts are similar in the regard that formation of occurs only in that region, but the overwhelming difference in apparent reaction also means that the heat integration in the adiabatic reactor is specific to the catalyst. It is hypothesized that these heat integration behaviors manifest themselves in unique temperature profiles. Accordingly, temperature measurements for a range of, along with selected species profiles, were acquired with the two catalysts and are presented next. 2% % the O 2 is converted over, its rate lowers distinctly, while, and additional and are formed downstream of this point. The C balance closes with minor production of CH 4, 2, and trace production of C 2 (not shown). As observed by Horn et al. (26), production is not produced 1:1 with. For this reason, those authors argued that production did not occur as direct oxidative dehydrogenation (Eq. (4)), but from pyrolysis of C 2 H 6. As a result, the formation of from the remaining O 2 accounts for a major route of O 2 consumption, which demands that some CH 4 and 2 is formed as well to satisfy mass action Spatial species profiles at high 2. Fig. 2. Equilibrium mixtures based on adiabatic calculations for variable with C 2 H 6 /air feeds. Equilibrium mixtures predict 1% C 2 H 6 and O 2 conversions with no or C 2. The main features associated with C 2 H 6 CPO at ¼ are present at high. Figs. 3c and d depict the consumption of reactants and formation of products for a ratio of 2 for the and catalysts, respectively. This feed ratio corresponds to the stoichiometry of ODH. Reaction rates are highest at the inlet of the catalytic region with nearly complete O 2 consumption over, and with much lower O 2 conversion over than at ¼. In general, the reactivity of the system over the catalyst is lower in region 2 (downstream of the O 2 slope change) than for the lower, while the observed reaction rates over are similar at both. With, formation at ¼2 occurs at a position farther downstream than at ¼, while the locus of 5% O Temperature profiles Temperature profiles, taken with an optical pyrometer and thermocouple, are presented for and in Figs. 4a and 5a, respectively. For either catalyst, both the pyrometer and thermocouple measurement increase in the front heat shield, with a higher slope exhibited in the thermocouple. Near the catalyst entrance, the thermocouple slope increases. Inside the catalytic domain, the temperature profiles differ in shape and value for the two catalysts, although the general trend of decreasing reactor temperature with increasing is respected in either case. The pyrometer measurement of the surface reaches a maximum at z¼1 mm of 14 1C at ¼ and decreases to 95 1C at ¼2 while shifting downstream to z¼11 mm. The thermocouple measurement increases inside the catalyst, and equilibrates with the pyrometer at 5 mm into the catalyst. With, both measurements increase over a larger portion of the catalyst such that the maximum in either measurement is at z¼18 mm for ¼ while shifting gradually downstream to z¼2 mm for ¼2. The maximum temperature is 1 1C for the former case, and 875 1C for the latter. For either catalyst, the front heat shield serves to preheat the gases entering the catalytic region. Heat is conducted from the hot spot to the front of the first monolith while simultaneously transferring heat to the gas. Since the front heat shields are identical, and the preheat is the same for each test, the magnitude of the heat transfer, indicated by the slopes of the pyrometer as well as the temperature difference between the pyrometer and the thermocouple, are nearly the same for either catalyst (6 1C regardless of or catalyst). The sharp peak in surface temperature, and the absence of one for, is related to the rate of reaction and the corresponding product distribution in the first part of the catalyst. Fig. 3 shows that the catalyst quickly consumes O 2 to form,, and, while produces, and a similar amount of with little.

6 3898 B.C. Michael et al. / Chemical Engineering Science 65 (21) C 2 H 6 O C 2 H 6 O 2 C 2 H 6 C 2 H 6.6 O 2.6 O 2 H Fig. 3. Spatial profile of major species flows for ¼ (a, b) and ¼2. (c, d) over 4.7 wt% /45 PPI a-al 2 O 3 (a, c) and 4.7 wt% / 45 PPI a-al 2 O 3 (b, d) for inlet flow of 5 SLPM, T in ¼1 1C. The catalytic monolith is located within 9ozo19. The overall exotherm is therefore higher, and confined to a smaller region with the catalyst. The apparent reforming reaction observed with occurs in the absence of exothermic reactions (see Figs. 3b and d), hence it is dependent on convection of heat from the exothermic zone of the reactor. This leads to a decrease in temperature, as observed with the combustion/ reforming mechanism for CH 4 CPO (Beretta and Forzatti, 24; Horn et al., 26; Michael et al., 29). By comparing the and C 2 H 6 curves in Figs. 5b and c, it can be seen that the rate of reaction is nearly invariant to in the second zone of the reactor, which agrees with the similar slopes of pyrometer temperature in the same zone. Qualitatively different behavior occurs with C 2 H 6 CPO on. Oxygen consumption occurs throughout the catalyst in exothermic reactions, resulting in an increase in temperature along the length of the reactor as depicted in Fig. 4a. The gas phase production of via pyrolysis is endothermic, thus the C 2 H 6 CPO system on becomes directly heat integrated between gas and surface in this scenario, rather than between two zones of surface reactions as in the system. Further evidence for gas-surface heat integration can be found by comparing the curves in Figs. 4a and b. As described before, the formation of begins at distances farther downstream with increasing (see Fig. 4b). The gas temperature at these onset locations is approximately constant with at a value of 75 1C. Beretta et al. (21) report ignition of C 2 H 6 /O 2 /N 2 mixtures to form can occur at temperatures as low as 65 1C without a catalyst at similar contact times; similarly, Morales and Lunsford (1989) report that gas phase reactions dominate a lithium-promoted magnesium oxide catalytic system above 675 1C. Although these temperatures are lower than in the current work, Lødeng et al. (1999) report gas phase ignition over blank monoliths at 8 1C. These observations suggest that, regardless of the exact feed composition, the exothermic catalytic reactions that increase the temperature are only indirectly related to formation via heat production. The spatial profiles herein reveal that the feed composition changes the rate of heat liberation over, thereby changing the locations where gas phase production can begin addition Many of the features described in the above discussion are indicative of an exo-endothermic reaction scheme throughout the reactor. Over, the zones are well defined with respect to a first region where O 2 is present, and a second zone with no O 2 and prevailing reforming activity. Profiles over indicate that O 2 is consumed more slowly, but with two distinct zones, and formation occurs downstream in the absence of reforming when temperatures are 75 1C. Earlier work has shown that cofeeding (Bodke et al., 1999, 2) or(donsı et al., 25a) in various ODH systems can boost the selectivity to, by a mechanism proposed by Bodke et al. (2) consisting of preferential oxidation of the co-fed fuel to provide heat for pyrolysis. To test this proposal, and to further probe the differences in catalytic activity between and, direct observation of the hallmarks of this mechanism were sought by obtaining spatial species and temperature profiles with addition.

7 B.C. Michael et al. / Chemical Engineering Science 65 (21) Increasing 11 1 Increasing Temperature ( C) Pyrometer Thermocouple Temperature ( C) Pyrometer Thermocouple C 2 H Increasing Increasing C 2 H Increasing Increasing Fig. 4. Comparison of reactor temperature profiles (a), measured by pyrometer (solid lines) and thermocouple (dashed lines), with formation (b) and C 2 H 6 consumption (c) for multiple (arrows indicate ranging from to 2. by.2 increments) over the 4.7 wt% /45 PPI a-al 2 O 3 catalyst. Dotted lines indicate isotherm/position relationship for formation. Spatial profiles for 2:1:2 C 2 H 6 :O 2 : mixtures over and are presented in Figs. 6a and b, respectively. In these tests, the C 2 H 6 and O 2 flows were the same as the previous tests at ¼2, while a portion of the Ar was replaced with to form the feed Fig. 5. Comparison of reactor temperature profiles (a), measured by pyrometer (solid lines) and thermocouple (dashed lines), with production and consumption (b) and C 2 H 6 consumption (c) for multiple (arrows indicate C/ O ranging from to 2. by.2 increments) over the 4.7 wt% /45 PPI a-al 2 O 3 catalyst. above. For either catalyst, the addition of improves the selectivity to at the effluent, as indicated in Table 1, however, the spatial profiles show a much more pronounced effect for the catalyst in comparison to the profile without addition.

8 39 B.C. Michael et al. / Chemical Engineering Science 65 (21) C 2 H 6 C 2 H 6 H 2 H 2 O.6 O 2.6 O 2.4 C.2 2 H Fig. 6. Spatial profile of major species flows for C 2 H 6 /O 2 / ¼2/1/2 over 4.7 wt% /45 PPI a-al 2 O 3 (a) and 4.7 wt% /45 PPI a-al 2 O 3 (b) for inlet flow of 5 SLPM, T in ¼1 1C. The catalytic monolith is located within 9ozo19. Table 1 Comparison of reactor performance measured at z¼42 mm and temperatures for and catalysts for C 2 H 6 /O 2 / ¼2/1/ and 2/1/2. C 2 H 6 /O 2 / 2/1/ 2/1/2 2/1/ 2/1/2 X i (%) C 2 H O S C,I (%) CH T max (1C) Pyrometer Thermocouple z(t max )(mm) Pyrometer Thermocouple In the first mm over the catalyst, rapid consumption of and O 2 occurs with minimal consumption of C 2 H 6. Water is the primary product in this region, which further indicates a preferential oxidation of. Further downstream, C 2 H 6 consumption begins to form,, and along with minor amounts of 2, CH 4 and. The consumption of C 2 H 6 begins prior to complete conversion of ; however, a distinct minimum can be observed in the profile, clearly partitioning the reactor. Production of and in the second zone occur at approximately 1:1, and slow to their final values by the end of the catalyst. Reactions on display similar behavior to the case without co-feed with the exception of the curve, which is translated upward by approximately the amount fed. A subtle minimum can be observed in the flow near the entrance of the catalyst; however, it is too small to discern whether or not it occurs in the absence of C 2 H 6 consumption with certainty. A maximum is present in the curve, indicating the presence of downstream steam reforming. The spatial temperature profiles (Fig. 7) corroborate the shift in activity for (Panel a), and the lack thereof for (Panel b) compared to the profiles without co-feed. For, the temperature maximum increases to 97 1C and shifts upstream (see Table 1) in accord with the existence of the large exotherm associated with rapid production. In contrast to this behavior, the temperature profiles for are similar to those with no co-feed. The opposing trends of and production early in the and catalysts in Fig. 3 has also been observed in CH 4 CPO profiles (Horn et al., 27). Hickman et al. (1993) proposed a mechanism that attributed the difference to the higher activation energy for H (s) oxidation on than on. This attribute is likely responsible for the difference in consumption between and. From the profiles, it can be inferred that the rate of desorption (originating either from C 2 H x surface species or from adsorption of fed ) is competitive with oxidation of H (s), which leads to a relatively small minimum in the curve over. With, on the other hand, desorption does not compete, so a larger initial consumption rate is present. Hydrogen co-feed tests also support the overall mechanism proposed in the literature of a preferential oxidation of sacrificial fuel to provide heat for endothermic production via pyrolysis. As described earlier, the processes involved in liberating heat, either from C 2 H 6 or the co-feed, determine where formation occurs. Fig. 7a reveals that the gas phase temperature reaches the 75 1C point much closer to the front of the catalyst, allowing for gas phase reactions to proceed throughout the reactor. Importantly, since C 2 H 6 is not consumed for heat, the yield of increases dramatically. Donsı et al. (25a) described an analogous process involving preferential oxidation over perovskites gives similar yields as co-feed over, suggesting that enhancement is not unique to /, but rather heat generation in general. In addition, a recent experimental and modeling study by Vincent et al. (28) revealed that the majority of catalytic processes in the ODH of C 2 H 6 over is dominated by oxidation, preserving the C 2 H 6 for gas phase formation via pyrolysis at a distance about halfway through the catalyst. Similarly, Zerkle et al. (2) assert that the dominant gas phase production of occurs via pyrolysis; however, those authors report up to 5% of the overall production occurs on the surface The roles reforming and O 2 in production The basic observation that formation occurs to a large extent over some catalysts (, Sn (Yokoyama et al., 1996), Perovskites (Cimino et al., 28; Donsı et al., 25a), and to a much lesser extent on others (), while also occurring in the absence of a catalyst (Beretta et al., 21; Lødeng et al., 1999), prompts the interrogation of the homogeneous and catalytic processes important for production. In this work, and were compared, and a number of differences and similarities can be observed in the product and temperature evolutions. The profiles are similar in the respect that (1) exothermic reactions occur rapidly very near the entrance of the catalyst leading to,

9 B.C. Michael et al. / Chemical Engineering Science 65 (21) Temperature ( C) H2 /O 2 = 2 1 /O 2 = /O 2 = 7 Pyrometer /O 2 = Pyrometer Thermocouple 1 Thermocouple Fig. 7. Comparison of reactor temperature profiles, measured by pyrometer (solid lines) and thermocouple (dashed lines) for ¼2, with and without -addition over the 4.7 wt% /45 PPI a-al 2 O 3 catalyst (a) and 4.7 wt% / 45 PPI a-al 2 O 3 (b). Temperature ( C) Scheme 1 2,, and, (2) the average catalyst temperatures are high enough for C 2 H 6 pyrolysis, (3) low C 2 selectivities are observed, and (4) coke accumulation was not observed. The profiles are different in the respect that (1) O 2 is completely consumed in the first few mm of the catalyst while it remains throughout the catalyst, (2) steam reforming dominates the downstream section of the catalyst while it is apparently absent on, (3) temperature increases sharply over while more slowly over, except with mixtures with addition. A large number of investigations report that production of occurs primarily in the gas phase (Donsı etal., 25b; Lødeng et al., 1999; Vincent etal., 28; Zerkle et al., 2), and some report that it isnotproducedonatall(silberova et al., 23). Fewer studies focus on the catalytic processes for consumption; however, it can be argued that these routes are at least as important as gas phase production in determining selectivity in SCTRs since both contribute to the net production. It follows, then, that consumption activity serves to distinguish catalysts in C 2 H 6 CPO. The catalyst produces temperature profiles that resemble those for with addition; therefore, production should show similar rates on as by the pyrolysis ignition correlation described earlier (Section 3.5). However, much lower selectivity is observed with. It is likely that the high reforming activity of compared to (Beretta and Forzatti, 24; Jones et al., 28) competes with the pyrolysis reaction that occurs in the gas phase. The competition can occur in two ways according to Scheme 1: the C 2 H 6 consumption via reforming (r 2 ) may be higher than pyrolysis (r 1 ), or the consumption by reforming (r 3 ) may compete with its production by pyrolysis. For any net production to occur, the average reforming rates r 2 and r 3 must be lower than the pyrolysis rate r 1. From the profiles, it can be observed that the production occurs early in the catalyst where temperatures are highest, and remains relatively constant throughout the rest of the reactor. The total reforming rate (indicatedbythe curve) drops severely through the remaining portion, and is independent of C 2 H 6 concentration. The rate of production fits well with the rate of C 2 H 6 consumption, indicating that other consumption routes (e.g. pyrolysis) are equal to the corresponding reforming rate. Relatively few works focus on steam reforming of C 2 H x on and catalysts compared to steam reforming of CH 4 ;however,graf et al. (27) have shown that the relative activity in reforming is C 2 H 6 4 4CH 4 over a /YSZ catalyst. This ultimately implies that C 2 H 6 reforming is faster than C 2 H 6 pyrolysis except where temperatures are sufficiently high in the gas phase and mass transfer to the surface controls the reforming rate. Importantly, the same authors note that the rate of reforming on is far lower than on based on conversion in isothermal tests, supporting the hypothesis that reforming activity plays an important role in the selectivity to observed with different catalysts. The routes for gas phase production of (e.g. pyrolysis in Scheme 1) also influence the overall yield in SCTRs. In particular, oxidative routes have been proposed to ignite at lower temperatures and proceed with higher rates than pyrolysis (Beretta et al., 21; Donsı et al., 22) on the basis of the activation energy of ethyl radical formation from OH (3 kcal/mol) versus H (9.5 kcal/ mol). Furthermore, O 2 assisted dehydrogenation of the C 2 H 5 radical to form results in the products of ODH (i.e. and rather than and ). Although the profiles in the current work indicate that the overall stoichiometry is not entirely assignable to ODH, the profiles over, wherein selectivity is highest, show O 2 present throughout the reactor with some concurrent production. Similarly, is formed over prior to the complete O 2 consumption. In addition, the reforming activity cannot alone account for the lower selectivity with since formation is not observed in the downstream heat shield where temperatures are still high enough to expect pyrolysis of unreacted C 2 H 6. The relative contributions of gas phase and catalyzed reactions are difficult to precisely determine from the sampling method used here, and remain a challenge in the study of C 2 H 6 CPO. Nonetheless, it is clear from the comparison of spatial profiles that reforming activity is an important distinguishing characteristic for and in determining the yield of in SCTRs. Similarly, gas phase radical reactions are sensitive to O 2 concentration for production, which correlates well with the O 2 consumption characteristics for and : consumes O 2 quickly to completion and yields very little, whereas does not completely consume O 2, but yields considerably higher quantities of. Further investigation of these relationships is warranted and is currently in progress.

10 392 B.C. Michael et al. / Chemical Engineering Science 65 (21) Conclusions Spatial concentration and temperature profiles were taken for CPO of C 2 H 6 /air mixtures for a variety of fuel/o 2 ratios over and -based monolithic catalysts. These profiles provide direct evidence that CPO of C 2 H 6 occurs in a multi-zone fashion in SCTRs. Over, rapid oxidation of C 2 H 6 to,, and occurs in the first few mm of the catalyst to completely consume O 2, followed by a reforming zone to produce additional syngas. The oxidation is slower over, which results in a slower release of heat with primary oxidation products being and. These two different oxidation behaviors give rise to a sharp peak in surface temperature for, and a more gradual increase in temperature for. A definite correlation can be made between production and reactor temperature, namely that, regardless of, the temperature required for production over is 475 1C, which agrees with other reports of the onset of homogeneous reactions. The locations at which reactor temperature is 475 1C increase with and therefore changes the catalyst volume available for production. Hydrogen addition experiments confirm the hypothesis originally proposed by Bodke and modified by Donsí, wherein preferential oxidation shifts the temperature profile maximum upstream allowing for production throughout the catalyst, boosting yield, while not completely consuming O 2 in the process. The apparent reforming activity of, and lack thereof with explains the major differences in production observed with either catalyst, while the presence of O 2 also correlates with production. Acknowledgements The authors thank Dr. Dave Rennard and Joshua Colby for insightful discussions related to the sampling technique. 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