Enhanced Methane Dehydroaromatization via Coupling with

May 4, 2017 - The highest combined yield for all aromatic compounds (benzene, toluene, and naphthalene) produced is achieved at a C6H6/CH4 ratio of âˆ...
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Enhanced Methane Dehydroaromatization via Coupling with Chemical Looping Casper Brady, Brian M Murphy, and Bingjun Xu ACS Catal., Just Accepted Manuscript • Publication Date (Web): 04 May 2017 Downloaded from http://pubs.acs.org on May 4, 2017

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Enhanced Methane Dehydroaromatization via Coupling with Chemical Looping Casper Brady, Brian Murphy, and Bingjun Xu* Center for Catalytic Science and Technology, Department of Chemical and Biomolecular Engineering, University of Delaware, 150 Academy Street, Newark, Delaware ABSTRACT: Methane dehydroaromatization (DHA) is a desirable process for the conversion of methane in stranded natural gas to valuable and easy-to-transport liquid aromatic compounds. However, the reaction is severely limited by thermodynamics. We present a method to circumvent the thermodynamic limitation by coupling DHA with chemical looping to achieve reactive separation of H2 from the products of DHA. The proposed strategy involves a four-step cycle, i.e., DHA on the Mo/H-ZSM-5 catalyst, hydrogen removal and regeneration on the Fe3O4/FeO redox pair, and water removal on Zeolite 5A. The feasibility of this process is validated by investigating each step individually. The results indicate that an aromatics yield up to >40% is possible with the proposed process.

KEYWORDS: dehydroaromatization, chemical looping, reactive separation, methane activation, Mo/H-ZSM-5

INTRODUCTION U.S. production of natural gas has increased dramatically in the past two decades,1 owing largely to the development and maturing of hydraulic fracturing for oil and shale gas extraction. One major barrier to using this new-found feedstock is the high transportation cost of natural gas from remote extraction sites. Thus, modular processes that convert methane to liquid products locally at extraction sites without the need for extensive infrastructure to handle high pressure gases or co-fed reactants are highly desirable. Methane is typically the main component in natural gas, and the most challenging to upgrade. Non-oxidative dehydroaromatization (DHA) of methane is a promising approach as no other reactant besides methane is needed and liquid aromatic compounds are produced in a single step. The DHA reaction was first reported in 1993 by passing methane over molybdenum oxide supported H-ZSM-5 (Mo/H-ZSM-5),2 which remains one of the most effective DHA catalysts.3-5

While much progress has been made, the implementation of DHA has to overcome two main obstacles, i.e., thermodynamic limitations, rapid catalyst deactivation. The thermodynamic barrier restricts the yield to benzene to roughly 12-13% at 700 °C.3 Accumulation of coke on the surface of catalysts generally accounts for 10-20% of methane conversion4, 6 and deactivates the catalyst in a few hours, which necessitates frequent catalyst regenerations. Much effort has been devoted to overcome these barriers, e.g., varying the type of zeolite framework 3-4, 6-9 and the silica to alumina ratio10, using membrane reactors to remove produced H2 during reaction to shift the thermodynamic equilibrium11-14, periodic regeneration by H2 or oxidants15-17 and removing aluminum from the frame work to decrease the amount of coke formed.18 However, none of the reported strategies yielded sufficient improvements to make DHA a commercially viable process in the foreseeable future. Membrane reactors are of particular interest because they can in principle enhance the equilibrium yield to aromatics11-14, however, their potential is limited

Figure 1: (a) Proposed combined DHA and chemical looping cycle for methane upgrading. (b) Thermodynamic equilibrium of methane aromatization with stages of hydrogen removal in between aromatization stages with 100% hydrogen removal at different temperatures. (c) Thermodynamic equilibrium of methane aromatization at 700°C with stages of hydrogen removal in between aromatization stages with varying degrees of hydrogen removal. All analyses assume 100% selectivity to benzene.

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by membrane stability at DHA temperature(s) and high costs. Chemical looping has been researched extensively as a method to separate combustion products via the use of solid oxygen carriers, which are generally reducible metal oxides.19-21 This is achieved by first flowing a hydrocarbon fuel over a metal oxide at high temperatures, which reduces the metal oxide while the hydrocarbon is oxidized to CO, CO2 and water. The reduced oxide is then re-oxidized using air in a separate reactor. Similarly, thermochemical cycles utilize the redox cycle of a metal and very high temperature to achieve water or CO2 splitting.22-26 This is accomplished by first heating an oxide to very high temperature (1500-2500 °C) to release oxygen.22 The reduced solid is then exposed to steam or CO2 to recover the metal oxide and produce H2 or CO. In this work, we show that chemical looping is an effective reactive separation strategy to remove H2 produced in DHA (Figure 1a), and increase the equilibrium methane conversion and aromatics yield. First methane flows over a Mo/H-ZSM-5 catalyst for the DHA reaction at 700 °C, producing an equilibrated mixture of methane, aromatics and H2 (step 1). H2 is then selectively oxidized in the DHA effluent by passing through a metal oxide, e.g., Fe3O4, to produce a reduced oxide, e.g., FeO, and water (step 2). The reduced oxide is then re-oxidized using water to produce H2 (step 3). The net result of steps 2 and 3 is separating H2 from the effluent from step 1. Finally, water produced in step 2 is removed from the gas stream using a high temperature water adsorbent (step 4), which can be regenerated by desorbing the water using temperature swing. The effluent from step 4, which ideally only contains methane and aromatics, is used as the feed to the DHA unit in the subsequent cycle, as the thermodynamic driving force for DHA is restored owing

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to the lack of H2. Repeating these four steps could convert methane with a theoretical 100% atom efficiency to valuable products, i.e., aromatics and H2. For ease of discussion, we refer to each iteration of these four steps as a stage.

RESULTS AND DISCUSSION A simple thermodynamic analysis shows that equilibrium methane conversion increases with the number of stages of the proposed cycle (Figure 1b). The equilibrium conversion of the DHA reaction is detailed in the Supporting Information (Figure S1) and is calculated assuming equilibrium conversion of methane and 100% selectivity to benzene. The equilibrium conversion of methane is iteratively calculated for each stage, assuming 100% H2 removal between stages, at three commonly used DHA temperatures (700, 750 and 800 °C, Figure 1b). Methane conversion higher than 60% could be achieved within 10 stages at 700 °C. The number of stages required to reach a given methane conversion decreases as temperature increases, because the equilibrium methane conversion in DHA increases with temperature.3 Unsurprisingly, the number of stages needed to reach a given methane conversion increases as the degree hydrogen removal decreases, because the presence of hydrogen reduces the thermodynamic driving force for DHA (Figure 1c). To demonstrate the feasibility of the proposed DHA/chemical looping cycle, we investigated each step individually in a fixed bed configuration (Figure S2) with real or simulated feeds. Product distribution of DHA with a pure methane feed on a 2 wt% Mo/H-ZSM-5 catalyst at 700 °C is consistent with literature (Figure 2a), with the methane conversion ranging from 6% to 9%, and the benzene yield ranging from 4% to 6%.27-29 Methane conversion decreases with time on stream due to coking of

Figure 2: (a) Time dependent reactivity data for a feed stream of pure methane. (b) Total conversion (including benzene and methane on a per carbon basis), and yields for all aromatic species and (c) Total percentage of carbon in the feed and outlet in aromatic compounds at C6H6:CH4 molar ratios to simulate the feed after various cycles of H2 removal on Mo/HZSM-5 at 700 °C. Conversions and yields are calculated from data at 100 min time on stream (time dependent data can be found in Figure S3).

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the catalyst. Benzene yield peaks at ~100 minutes time on stream, as the molybdenum carbide sites necessary for methane activation are formed3, followed by a decrease due to catalyst coking. Mixtures of methane and benzene with varying molar ratios were used as feeds to simulate the feed in subsequent stages (stages ~1.5, 2, 3, and 6, Figure 2b). Assuming benzene is the only aromatic product and no carbon loss to other products, including coke, the composition of simulated feed can be determined, e.g., a molar C6H6/CH4 ratio of 0.021 is employed to simulate the feed for 2nd stage based on the equilibrium methane conversion in the 1st stage. The lack of H2 in the simulated feed reflects the composition of the effluent after step 4 of the proposed cycle, assuming complete H2 removal. Product distributions in DHA with simulated feeds show that H2 removal dramatically increases the yield to aromatics. Similar methane conversions at each stage were observed, because the lack of H2 in simulated feeds restores thermodynamic driving force for DHA (Figure 2b). High concentrations of benzene in the feed, which simulates later stages, favor the formation of larger aromatics (toluene and naphthalene) and coke. Benzene is likely an intermediate in the production of alkylated aromatics (e.g., toluene), polyaromatic compounds (e.g., naphthalene) and coke.3 The net yield of benzene, i.e., the yield of benzene at the effluent minus that in the simulated feed, becomes negative at C6H6/CH4 ratios higher than 0.045 (stage 3), which indicates the rate of benzene consumption in methylation, oligomerization and coking pathways outweighs that of benzene production in DHA. The highest combined yield for all aromatic compounds (benzene, toluene and naphthalene) produced is achieved at a C6H6/CH4 ratio of ~0.13 (stage 6), beyond which the amount of aromatic carbon lost to coking becomes higher than that produced by DHA. This indicates that the maximum yield of aromatics (benzene, toluene and naphtha-

Figure 3: Fe 3O4 Temperature Programmed Pulse experiment results. Pulses of simulated DHA effluent in stage 1 are introduced to Fe3O 4 at a range of temperatures.

lene) possible with staged hydrogen removal with the Mo/H-ZSM-5 catalyst employed in this study is ~43% (Figure 2c), which is more than a factor of 5 higher than yields reported in one-pass DHA literature.27-29 It should be noted that, as experiments with simulated feeds include only benzene as a cofed aromatic, in practice the yield of aromatics of the process is likely lower as heavier aromatics (such as toluene and naphthalene) are oligomerized to form carbonaceous species more easily than benzene. The Fe3O4/FeO pair is effective in selectively removing H2 from DHA effluents without consuming hydrocarbons. Pulses of simulated DHA effluent from stage 1 were introduced to a bed of Fe3O4 at a range of temperatures from 300-800 °C (Figure 3). Unless otherwise noted the molar composition of the simulated stage 1 DHA effluent is 82% CH4, 16% H2, and 2% C6H6. Calculations of the thermodynamic equilibrium conversion are detailed in the Supporting Information (Figure S4). Experimentally observed H2 conversion roughly follows the values predicted by thermodynamic equilibrium (solid and dashed red lines, respectively, Figure 3) within the temperature range, indicating the reaction between H2 and Fe3O4 is kinetically facile. No CH4 conversion is detected below 800 °C, despite the fact that the conversion of CH4 to CO and H2 is thermodynamically favorable at a range of temperatures from 400-800 °C (black lines, Figure 3). This is likely due to the high activation barrier of methane activation and the lack of an active site on Fe3O4. The thermodynamically predicted onset temperature for benzene oxidation is ~400 °C, however, benzene conversion is not observed until ~650 °C (blue lines, Figure 3). This is likely a result of high activation barriers to activate C-C and CH bonds in benzene in the absence of Brønsted acid sites.30 Similar results were observed using pulses with a higher benzene concentration to simulate the DHA effluent from later stages (stage 5), i.e., 79% CH4, 8% C6H6,

Figure 4: Fe 3O4 Isothermal pulse experiment results. Pulses of DHA effluent are introduced to a bed of Fe3O4 at 650 ⁰C until roughly 60% conversion of the Fe 3O4. The bed is then re-oxidized by steam. Reduction and oxidation cycles are repeated twice.

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13% H2 (Figure S5). Thus, selective H2 oxidation in DHA effluent streams without consuming any hydrocarbons can be achieved from 450-650 °C. Isothermal redox pulse experiments at 650 °C show that the redox cycle of Fe3O4 and FeO is highly reversible. Pulses of simulated DHA stage 1 effluent were introduced to a bed of Fe3O4 at 650 °C. The initial hydrogen conversion is close to the equilibrium value. Hydrogen conversion (Figure 4, pink symbols) decreases with the increasing number of pulses introduced owing to the decreasing Fe3O4/FeO ratio in the solid (Figure 4, green symbols). A lower Fe3O4/FeO ratio leads to lower equilibrium H2 conversion. No detectable level of benzene and methane consumption is observed, except for the first pulse, which is likely due to the presence of a trace amount of Fe2O3 in the commercial Fe3O4 sample (control experiment confirms that Fe2O3 reacts with benzene at 650 °C). The fraction of Fe3O4 consumed after each pulse can be estimated based on the cumulative hydrogen consumption, assuming FeO being the only product (Figure 4, green symbols). ~65% of Fe3O4 is consumed after the introduction of 20 pulses, at which point the per-pulse hydrogen conversion is ~20%. An excess of steam is then introduced to the solid bed at 650 °C to convert all solid back to Fe3O4 and release hydrogen. The amount of hydrogen produced during the steam treatment matches closely with that consumed in the 20 pulses (within 5%, red bars in Figure 4), indicating that the redox cycle is complete reversible. Total recovery of the initial Fe3O4 phase is supported by the identical XRD patterns of the initial sample and after the steam treatment (Figure S6). Two more cycles of pulse experiments and steam treatment are conducted, and very similar results are observed, demonstrating the recyclability of the redox cycle. A small decrease in hydrogen conversion is observed in 2nd and 3rd cycle as compared to 1st cycle, which could be caused by sintering or slow formation of a carbon or car-

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bide species on the surface of the oxide,31 which could block surface oxides sites for hydrogen oxidation. Water produced by the hydrogen oxidation step needs to be removed because the presence of water 1) limits the equilibrium conversion of hydrogen with Fe3O4, and 2) reduces the yield to aromatics in the subsequent DHA step. Zeolite 5A has been shown to be able to completely adsorb pulses of water introduced in the temperature range from 200-300 °C (Figure S7a). This is consistent with the mass loss of water saturated Zeolite 5A during thermogravimetric analysis (TGA, Figure S7b); at temperatures >200 °C the molecular sieve desorbs roughly 20% of its measured ~13 mmol/g water adsorption capacity, in good agreement with literature.32 Thus, introducing a bed of zeolite 5A at 200 °C allows for complete water removal from the feed to the DHA step, and this bed could be regenerated by temperature swing to 500 °C. No detectable conversion of methane or benzene is observed when pulses of simulated DHA effluent in stage 1 are introduced to Zeolite 5A at 250 °C (Figure S7c). The low hydrogen conversion when reacting the DHA effluent with the Fe3O4 bed (Figure 4) in a single pass could be improved with a multi-pass configuration with multiple beds of Fe3O4 and Zeolite 5A in series between stages of DHA to increase the degree of H2 removal (Figure 5a). A zeolite 5A bed is placed after each Fe3O4 bed to remove water produced in the effluent of the hydrogen oxidation step. Repeating the hydrogen removal and water adsorption cycle is expected to increase the degree of hydrogen removal. The feasibility of this concept is demonstrated by a 3-bed, 2-pass configuration (Figure 5b), in which pulses of simulated DHA effluent in stage 1 are introduced to a first bed of Fe3O4 at 650 °C, followed by a second bed of Zeolite 5A at 200 °C, and then a third bed of Fe3O4 at 650 °C. The initial H2 conversion increases from 51% in the 1-pass experiment to 82% in the 2-pass experiment, which is roughly consistent with com-

Figure 5: (a) Graphical depiction of a Fe 3O4/Zeolite 5A multi-pass setup for higher H2 conversion. (b) Measured H2 conversion over Fe3O4 at 650 °C and Zeolite 5A at 200 °C for the 1 pass and 2 pass cases and equilibrium H2 conversion for the 1-3 pass cases.

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plete water adsorption between the two Fe3O4 beds. Assuming an equilibrated mixture of H2/H2O is produced after passing each Fe3O4 bed, 3 passes are needed to reach >80% H2 removal.

CONCLUSIONS In conclusion, we proposed a 4-step cycle that combines DHA of methane and a chemical loop to enable the reactive separation of hydrogen from the effluent of DHA. We demonstrated the feasibility of each step in the proposed scheme individually, and estimated the upper bound aromatics yield for the current configuration to be ~43%, which is more than a factor of 5 higher than current optimal single-pass results.

EXPERIMENTAL All reactivity experiments were conducted in a fixed bed flow microreactor, either in the continuous flow mode or the pulse mode. The Mo/H-ZSM-5 catalyst was pre-carburized prior to reaction in a stream of 1:1 CH4:H2 during a temperature ramp from room temperature to 700 °C at 10 °C/min and 15 minutes at 700 °C.10 A loading of 0.4 g of Mo/H-ZSM-5 was employed as the catalyst in the DHA studies. Pulse volume of 0.30 and 3.25 mL were used for the temperature programmed pulses, and the isothermal pulses respectively. 0.15 g and 0.05 g of Fe3O4 were loaded for the temperature programmed pulses and the isothermal pulses respectively. All liquids were introduced using Cole Parmer syringe pumps. The 2-pass experiment was performed by mounting 2 reactor tubes in the same furnace with the temperature of the 2 Fe3O4/FeO beds and the Zeolite 5A bed controlled independently at 650 °C and 200 °C, respectively.

AUTHOR INFORMATION Corresponding Author * [email protected]

Funding Sources We gratefully acknowledge the support of the ACS PRF grant 55467-DNI9.

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ASSOCIATED CONTENT Supporting Information. Additional experimental details are available in the Supporting Information. This material is available free of charge via the Internet at http://pubs.acs.org.

ACKNOWLEDGMENT We appreciate Prof. Raul Lobo and Prof. Douglas Buttrey for useful discussions.

REFERENCES (1) Bruijnincx, P. C.; Weckhuysen, B. M., Angew. Chem. Int. Edit. 2013, 52, 11980-11987. (2) Wang, L.; Tao, L.; Xie, M.; Xu, G.; Huang, J.; Xu, Y., Catal. Lett. 1993, 21, 35-41. (3) Spivey, J. J.; Hutchings, G., Chem. Soc. Rev. 2014, 43, 792-803. (4) Ma, S.; Guo, X.; Zhao, L.; Scott, S.; Bao, X., J. Energy Chem. 2013, 22, 1-20.

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