The influence of hydrogen-permeable membranes and pressure on methane dehydroaromatization (MDA) in packed-bed catalytic reactors

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1 The influence of hydrogen-permeable membranes and pressure on methane dehydroaromatization (MDA) in packed-bed catalytic reactors Benjamin Kee a, Canan Karakaya a, Huayang Zhu a, Steven C. DeCaluwe a, Robert J. Kee a, a Mechanical Engineering, Colorado School of Mines, Golden, CO 1, USA Abstract Computational simulations are developed and applied to study the coupling of packed-bed methane dehydroaromatization reactors with hydrogen-selective membranes, for the production of value-added fuels such as benzene. Detailed chemical kinetics for reforming over bi-functional Mo/H-ZSM- catalysts are validated against published literature, and simulations explore the effect of hydrogen removal, operating pressure, reactor temperature, and gas-hourly space velocity (GHSV). Although results reveal that membrane integration significantly increases methane conversion, the desired benzene selectivity decreases, due to the increased yield of undesired byproducts such as naphthalene. The benzene-to-naphthalene ratio depends strongly and nonlinearly on the membrane hydrogen removal, and simulations demonstrate that hydrogen membranes are most beneficial at relatively high GHSV and relatively low catalyst temperature. Increasing pressure decreases conversion and benzene selectivity, but increases benzene production rates and doesn t affect naphthalene selectivity. While the single-pass benzene yield remains low, results predict that advanced multi-pass reactor designs with hydrogen membranes and increased pressure can operate continuously to increase benzene production rates greatly. Keywords: Methane dehydroaromatization, MDA, Hydrogen membrane, Benzene production 1. Introduction The primary objective of this paper is to explore the effects of hydrogen removal through a membrane for the conversion of methane to benzene via methane dehydroaromatization (MDA), based on a detailed reaction mechanism. The results provide new quantitative insight about the interactions between catalyst performance, operating conditions, and membrane performance. The technical approach is based upon computational modeling, with particular attention to shell-and-tube packed-bed membrane reactors. An elementary-step chemical reaction mechanism is used to represent heterogeneous chemistry on Mo/zeolite catalysts. Transport fluxes within the packed bed involve convective and diffusive contributions. Effective diffusion coefficients consider both micro- and macro-porosity, incorporating the transport within and around the catalyst pellets. The operating conditions varied are gas-hourly space veloc- Corresponding author address: rjkee@mines.edu (Robert J. Kee) 1 ity (GHSV), pressure, and temperature. Membrane flux is varied by adjusting the permance. Several important results emerge from the study. Membranes that enable large hydrogen fluxes increase methane conversion, but at the expense of benzene selectivity. Benzene selectivity competes with undesired naphthalene formation, with the tradeoff depending on membrane performance. High pressures reduce conversion, but increase benzene molar production rates. Short catalyst-bed residence times and moderate hydrogen removal rates increase benzene yields. Increasing the pressure allows for operation in non-coking regimes, while increasing benzene production rates Background The recent abundance and inexpensive supply of natural gas presents new opportunities for conversion into useful chemical products. However, transportation of natural gas is impractical, which has sparked interest in gas-to-liquids technology. Natural gas from oil production and refining is usually flared to meet environmental restrictions. As a prominent component in natural gas, methane is the focus of catalytic conversion processes

2 Preprint of article published in Ind. Eng. Chem. Res., CH C H + 9 H2. Conversion (%) CH CH 3 CH CH atm 9 1 Temperature (K) Figure 1: Constant temperature and pressure equilibrium conversion and selectivity calculations. pressure is increased from 1 to atm. For the typical conditions, sustained production of benzene is difficult. Operating conditions and hydrogen removal can assist in improved benzene production. Consistent with Le Cha telier s principle, implementing a hydrogen membrane to remove H2 from an MDA reactor can increase methane conversion. Reaction 1 shows that for every mole of benzene produced, 9 moles of hydrogen are also produced. By removing H2 via a hydrogen-selective membrane, the thermodynamic equilibrium shifts toward the products. Figure 2 shows a potential implementation for an MDA membrane reactor. As illustrated, the catalytically active packed-bed is housed in the reactor s annular region. Hydrogen produced within the catalyst bed is transported across the membrane into an inert sweep channel, and the desired MDA products exit downstream. (1) This is an endothermic reaction that requires high temperatures (7- C) and relatively low gas-hourly 1 space velocities (7 ml g 1 cat hr ). MDA was first reported by Wang et al. [7] in 1993 using molybdenum supported on H-ZSM- zeolite, claiming 7-% methane conversion and % benzene selectivity. Despite various lab-scale advances since its discovery, low benzene yield and poor catalyst stability currently limit scale-up of MDA. This is primarily due to competition between reaction 1 and naphthalene (C H ) production: CH C H + 1 H2, Selectivity (%) to higher-value hydrocarbons [1]. Fischer-Tropsch (FT) is the commercial standard for forming larger hydrocarbons. The issue with F-T processes is that methane must first be converted to synthesis gas (CO+H2 ) before conversion to larger hydrocarbon products. The multiple unit processes involve in F-T generate inherent inefficiencies. Methane dehydroaromatization is a potentially viable route for converting methane to benzene [1 ]. Benzene has many attractive uses, including as a precursor for ethylbenzene and cumene, both of which are produced on commercially large scales, and as a common additive to gasoline and jet fuels. MDA is a direct, nonoxidative path to form higher hydrocarbons, which can enable more efficient conversion, relative to F-T. MDA can be expressed globally as: (2) where naphthalene acts as a precursor for larger polyaromatic hydrocarbons, the production of which leads to catalyst coking and deactivation. Identifying operating conditions that promote high CH conversion while maintaining high benzene selectivity is a major challenge, and one of the primary objectives of this work MDA thermodynamics Figure 1 shows the equilibrium conversion and selectivities at constant temperature and pressure, for pressures of and atm. At standard operating conditions (7 C and ), MDA is equilibrium-limited to 12% conversion []. Increasing temperature increases conversion, but decreases benzene selectivity, while increasing pressure reduces conversion and increases benzene selectivity. Increasing flow rate decreases conversion but can increase benzene yield [9]. For both pressures, the selectivity toward benzene is fairly similar across the temperature range. However, the selectivity to naphthalene decreases around percentile as Figure 2: Membrane coupled packed-bed reactor in an annulus design MDA is typically employed at atmospheric pressure. Once again considering Le Cha telier s principle, increasing pressure will reduce conversion, due to the large number of moles of hydrogen produced. Because the number of moles produced from benzene (Eq. 1) is less than that from naphthalene (Eq. 2), increasing pressure favors benzene selectivity over naphthalene. Furthermore, the driving force for membrane flux is partial 2

3 pressure; increasing the operating pressure increases the hydrogen removal rate, which should ameliorate conversion losses with increasing pressure. 2. Computational Model To explore the design of membrane-coupled MDA, a computational model is developed and applied to simulate the detailed MDA kinetics in a packed-bed annular reactor. As illustrated in Fig. 2, the integrated packedbed MDA reactor contains an annular porous packed bed, an H 2 -permeable selective membrane, which is normally supported on the porous structure, and the internal sweeping gas-flow channel. Fig. 3 shows the radially symmetric computation domain. The model assumes one-dimensional, isothermal, axial transport within the packed bed, represented as species mass fluxes j k. Species transport through the membrane is expressed as a radial mass flux j k,m. The inner and outer radii of the packed-bed annulus are denoted as R i and R o, respectively. The cross-sectional flow area of the packed bed and the hydrodynamic perimeter of the membrane exterior can be expressed as A B = π ( ) R 2 o R 2 i and PB = 2πR i, respectively. Preprint of article published in Ind. Eng. Chem. Res., ( ) ε g ρ g K g K g + j k = A s ṡ k W k + P B t A B k=1 k=1 K g k=1 j k,m. () The independent variables are time t and the axial coordinate z. The dependent variables are the mass fractions Y k and the gas-phase density ρ g. The porosity of the packed bed is ε g. Chemical production rates via heterogeneous reactions are represented by ṡ k. Surface reaction rates are evaluated at the isothermal surface temperature. The specific surface area of the active catalysts (i.e., active surface area per unit packed-bed total volume) is A s and W k are the species molecular weights. The gas-phase pressure is determined from the idealgas equation of state. For operating temperatures from C to C, the homogeneous gas-phase chemistry is assumed to be negligible in comparison to the heterogeneous surface chemistry []. The surface species coverages θ k,m are solved from the heterogeneous reaction rates θ k,m t = ṡk Γ m, () where Γ m is the available site density for surface site m (m = Mo 2 C or H-ZSM-) Gas-phase transport The gas-phase species mass fluxes due to convection and diffusion are represented as j k = ρ g Y k v D e k W k [X k ] () where [X k ] are the species molar concentrations and D e k are the effective diffusion coefficients. The apparent velocity v due to the pressure gradient p through the porous packed-bed structure can be represented by Darcy flow approximation Figure 3: Illustration of the computational domain, including labeling of fluxes and reactor geometry Conservation equations The model is formulated as a series of continuum partial differential equations, which are integrated over a sufficiently long time span to simulate quasi-steadystate operation of the reactor. The transient species and overall mass balances are represented as ( ε g ρ g Y k ) t + j k = A s ṡ k W k + P B A B j k,m, (3) 3 v = B g p. (7) µ where µ is the viscosity. The bed permeability B g can be determined by the Kozeny-Carman relationship B g = ε 3 gd 2 p 72τ g ( 1 εg ), () where d p is the mean particle diameter and τ g is the tortuosity. Diffusion within the MDA packed bed occurs on both the macro- and microscale. The space between catalyst particles is represented as the macroporosity ε g. Within the catalyst particles, molecules navigate through the

4 catalyst pellets toward active catalyst sites, which is represented as the microporosity ε m. Effective diffusion coefficients are calculated from the random pore model [11, 12] to incorporate both types of porosity. ( ) 1 + 3εg D e k = ε2 gd k,m + ε2 m D k,kn, (9) 1 ε g The Knudsen diffusivities D k,kn are defined as D k,kn = 2 3 r p RT πw k, () where r p is the average pore radius. The mixtureaveraged diffusion coefficients D k,m are determined from kinetic theory, retrieved from Cantera Hydrogen-permeable membrane The hydrogen removal method incorporated into this model is a Sieverts law correlation j H2,M = k M p n H 2,React pn H 2,Perm L M W H2, (11) where partial pressures of hydrogen within packed bed ( reactant ) and the sweep channel ( permeate ) are p H2,React and p H2,Perm, respectively. The membrane thickness is L M. The exponent n (. n 1.) depends on the hydrogen partial pressure. For hydrogen partial pressures less than 1 bar, which are common for most MDA applications, the diffusion of atomic hydrogen is the rate limiting step, then n =.. If hydrogenhydrogen interactions within the bulk become significant, n may increase. The hydrogen permeability of the membrane can be expressed in the Arrhenius form, ( k M = k M exp E ) a, (12) RT where E a is the apparent activation energy. In this study, the H 2 permeate flux is assumed to be much lower than the inert sweep gas flow, and hence the sweep side hydrogen concentration is assumed negligible ( p H2,Perm = ). Additionally, the permeability and thickness were abstracted into a single variable A M = k M /L M which represents the permeance, resulting in an updated Sieverts law correlation 1 j H2,M = A M p. H 2,React W H 2. (13) 2.. Performance metrics The results are generally reported in terms of outlet gas-phase compositions, methane conversion, and selectivity to major products. The present study uses methane conversion, in addition to selectivities and yields for both benzene and naphthalene as basic comparison metrics. Methane conversion is evaluated as X CH = J CH,in J CH,out, (1) J CH,in where J CH,in and J CH,out are the inlet and outlet methane molar fluxes. The carbon selectivity of product species k is evaluated as N C,k J k,out S k =, (1) J CH,in J CH,out where N C,k is the carbon number of species k (e.g., N C = for ) and J k is the species outlet molar flux. The product yields are evaluated as Y k = X CH S k = N C,kJ k,out. (1) J CH,in The membrane effectiveness is reported as the ratio of hydrogen removed through the membrane to the hydrogen exiting the end of the packed bed ṁ H2,M % H 2 removed =, (17) ṁ H2,M + ṁ H2,out where ṁ H2,M is the total mass flow rate of H 2 through the membrane and ṁ H2,M is the packed bed outlet mass flow rate of H 2. Another important metric is the ratio of the selectivity toward benzene to the selectivity toward naphthalene Benzene/Naphthalene ratio = S S C H. (1) Finally, the local transport resistances in the reactor bed can be characterized via the Weisz-Prater (C WP ) and Mears (C M ) criteria. The Weisz-Prater criterion quantifies the internal diffusion transport limitations C WP = ṡ CH ε g d 2 p D e CH [ XCH ], (19) where ṡ CH is the apparent destruction rate of methane, ε g is the packed bed porosity, d p is the particle diameter, D e CH is the diffusion coefficient of methane, and [ ] X CH is the methane concentration. For C WP 1, internal diffusion limits the reaction. As flow rates increase, the transport becomes a limiting factor. For the conditions

5 Table 1: MDA Reaction Mechanism on the Mo/H-ZSM- by Karakaya, et al. [13] Molybdenum Site Chemistry 1. CH + (m) CH (m) 2. CH (m) CH + (m) 3. CH (m) CH 2 (m) + H 2. CH 2 + H 2 CH (m). CH 2 (m) + CH C 2 (m). C 2 (m) CH 2 (m) + CH 7. C 2 (m) C 2 H + H 2 + (m). C 2 H + H 2 + (m) C 2 (m) Zeolite Site Chemistry 9. C 2 H + H(z) C 2 H (z). C 2 H (z) C 2 H + H(z) 11. C 2 H (z) + C 2 H C H 9 (z) 12. C H 9 (z) C 2 H (z) + C 2 H 13. C H 9 (z) + C 2 H H 13 (z) 1. H 13 (z) C H 9 (z) + C 2 H 1. H 13 (z) H 12 + H(z) 1. H 12 + H(z) H 13 (z) 17. H 12 + H(z) H 11 (z) + H 2 1. H 11 (z) + H 2 H 12 + H(z) 19. H 12 + C 2 H (z) H 11 (z) + C 2. H 11 (z) + C 2 H 12 + C 2 H (z) 21. H 11 (z) + cy- H 11 (z) 22. cy- H 11 (z) H cy- H 11 (z) H + H(z) 2. H + H(z) cy- H 11 (z) 2. C H 9 (z) C H + H(z) 2. C H + H(z) C H 9 (z) 27. C H + C 2 H (z) H 13 (z) 2. H 13 (z) C H + C 2 H (z) 29. C H + H(z) C H 7 + H 2 3. C H 7 + H 2 C H + H(z) 31. H + H(z) C H 9 (z) + H C H 9 (z) + H 2 H + H(z) 33. H 9 (z) H 7 (z) + H 2 3. H 7 (z) + H 2 H 9 (z) 3. H 7 (z) + H(z) 3. + H(z) H 7 (z) C H 7 (z) C H 13 (z) 3. C H 13 (z) + C H 7 (z) 39. C H 13 C H 12 + H(z). C H 12 + H(z) C H C H 12 C H 11 + H(z) 2. C H 11 + H(z) C H C H 11 C H + H(z). C H + H(z) C H 11. C H 11 (z) C H 9 (z) + H 2. C H 9 (z) + H 2 C H 11 (z) 7. C H 9 (z) C H + H(z). C H + H(z) C H 9 (z) 9. CH + H(z) CH 3 + H 2. CH 3 + H 2 CH + H(z) 1. + CH 3 (z) C 7 H 9 (z) 2. C 7 H 9 (z) + CH 3 (z) 3. C 7 H 9 (z) C 7 H + H(z). C 7 H + H(z) C 7 H 9 (z) explored in this study, the Weisz-Prater criterion is typically less or near to 1. The Mears criterion quantifies the external diffusion transport limitations, C M = ṡ CH ε g d p 2k c [ XCH ], () where k c is the mass transfer coefficient. For C M <.1, external mass transfer effects can be neglected. 2.. MDA kinetics A thermodynamically consistent mechanism is used for calculating the heterogeneous chemical reactions rates, as shown in Table 1 [13]. The mechanism consists of reactions, which encompasses both the molybdenum and zeolite sites. There are 12 gas phase species and 1 surface species. The reactions are presented as irreversible pairs, but achieve thermodynamic consistency. 3. Results The results of this study are first presented without the influence of hydrogen removal. Then membrane effects are explored. Finally, the effects of pressure are explored with and without a membrane. Table 2 summarizes the major variables and parameters for the results (picked to match experimental operating conditions [1]). The temperatures, pressures, and GHSV may vary between parameter studies. Although the current mechanism does not incorporate coking reactions, the two ring aromatic naphthalene is generally accepted as the monomer of carbon-deposit-forming PAH species. Hence, increased naphthalene formation is used in this study as an indication that the reactor is operating within a catalyst fouling regime.

6 Table 2: Major variables and parameters for dry MDA simulations Dimensions Length. cm Annulus Inner Diameter 1. cm Annulus Outer Diameter 1. cm Inlet Temperature 7-7 C Pressure 1-7 atm Mole Fractions.9 CH. N 2 GHSV 7-7 ml g 1 Catalyst Macroporosity.3 Microporosity. Molybdenum Loading % Specific Area 9. + m 1 Pore Radius 2.2 m Particle Diameter 7 m 3.1. No hydrogen removal, atmospheric pressure We first explored the effect of operating conditions (GHSV and temperature) on reactor performance without hydrogen-permeable membrane coupling. For these cases, the GHSV was typically varied between 7 GHSV 2 ml g 1. Figures,, and show the predicted effects of varying GHSV and reactor temperature on methane conversion (Fig. ), benzene selectivity (Fig. ), benzene yield, and the /C H yield ratio (Fig. ). The individual symbols in Figure correspond to the specific GHSV values from Figure. Taken together, the figures convey the great difficulty in optimizing reactor operating conditions. High GHSV leads to low conversion and high benzene selectivity, but low GHSV leads to the production of catalyst-deactivating species. Typical operating conditions for MDA include high temperatures near 7 C and low gas hourly space velocities between 7-1 ml g 1. These conditions assure high methane conversion, but at the expense of low benzene yield. Figures and show that increasing the temperature increases the methane conversion. However, there is a simultaneous decrease in the selectivity toward benzene, as seen in Fig.. Additionally, increasing reactor temperature increases the likelihood of catalyst fouling via coking (shaded region in Fig. ). For long and stable operation times, the naphthalene formation rate and the concentrations should be low (i.e. the to C H yield ratio should be high). Figure shows that, while benzene yield increases with temperature, the formation CH conversion (%) C 72 C 7 C 7 C GHSV (ml g cat hr ) Figure : Effect of GHSV and reactor temperature on methane conversion rate for an MDA reactor with no membrane. of undesired products naphthalene and PAH are favored at high temperatures. Selectivity (%) 7 7 Increased GHSV 7 C 7 C 72 C 7 C 12 1 CH conversion (%) Figure : Effect of GHSV on benzene selectivity and methane conversion an MDA reactor with no membrane. Increase or decrease in methane conversion corresponds to specific GHSV values between 7-2 ml g 1 as it is shown in Figure. Although decreasing the GHSV increases the methane conversion (Figure ), the impacts on benzene selectivity, yield, and propensity for catalyst coking (Figure and ) are more complex. Figure shows that for each temperature there appears an optimal GHSV, corresponding to the maximum benzene selectivity. This is due to competition between naphthalene and benzene formation. As GHSV increases from low to moderate values, the naphthalene formation from benzene becomes transport limited. The benzene diffusion surpasses the naphthalene formation rate, and ultimately the reaction favors benzene formation, which appears as a local maximum in the benzene selectivity. Further increasing the GHSV decreases the methane conversion rate as well as the benzene selectivity, as the benzene formation rate also becomes internally transport limited, as indicated by the Weisz-Prater criterion dependence on GHSV. 1

7 Benzene yield (%) 9 Increased GHSV 7 C 7 72 C 7 C 7 C /C H yield ratio Figure : Effect of GHSV and reactor temperature on benzene yield for an MDA reactor with no membrane. Each data point on benzene to naphthalene ratio corresponds to the specific GHSV between 7-7 ml g 1 as it is shown in Figure. The transitions between reaction rate controlled regimes and transport controlled regimes differ for each temperature range. At low temperatures, the transition is sharper. A slight increase in the GHSV can easily cause the transition between kinetically-controlled and transport-controlled regimes. Increasing temperature often widens the range of GHSV values in the nontransport controlled regime. Using the catalyst properties and reaction conditions for the present study, the optimal sensitivity at 7 C corresponds to GHSV in the range of 3 - ml g 1. At 7 C, the range expands to 3 - ml g 1, at 72 C, the range is 3-9 ml g 1 cat hr 1, and finally at 7 C, the optimum selectivity range is at - 1 ml g 1. Figure highlights the challenges involved in optimizing MDA reactor operating conditions. The conditions with the highest benzene yields coincide with high carbon deposition rates (indicated by /C H ratios below 2.). The ratio was determined from experimental measurement of the benzene-to-naphthalene ratio immediately after carburization. At high GHSV, the /C H ratio is high and the reactor operates in a reduced coking regime, but the process is transport limited. Therefore, methane conversion and benzene yields are low. At high temperatures, methane conversion, and hence benzene yield, increase, but the reaction heavily favors naphthalene production, resulting in significant coking and catalyst deactivation. Even at very high temperatures (7 C) and low GHSV (7 ml g 1 ), the maximum benzene yield remains below 9%. Based on Figures and, MDA reactors should be operated at low temperatures (i.e., below 72 C for long-term time on stream). However, at temperatures below 72 C, the 7 methane conversion and consequently the benzene yield is too low to meet economic expectations. The compromise between high benzene yield and catalyst stability is therefore a major issue to achieve economic viability Hydrogen membrane, atmospheric pressure Figures 7a and b show the effects of hydrogen removal on MDA reactor performance. Methane conversion and major species yields are plotted as a function of hydrogen removed, for a reactor operating at 1 ml g 1 and the temperature varies between 7-7 C. Hydrogen removal increases the methane conversion as well as naphthalene and benzene yields. The increases in both naphthalene and benzene yields follow a nearly linear trend below % hydrogen removal, where the methane conversion is below %. Further removal of hydrogen significantly increases methane conversion, which causes the product distribution to significantly favor naphthalene. Even though high benzene yields ( 1%) are achieved at 9% hydrogen removal, the corresponding naphthalene yield reaches approximately %. At high hydrogen removal rates, the catalyst deactivation via PAH formation is highly probable. It is well accepted that the coke formation rates are related to the H 2 concentration [1]. Graphitic coke as well as PAH, often form at low hydrogen concentration ranges. a) CH conversion (%) b) Benzene yield (%) 7 C 72 C 7 C 7 C H 2 removal (%) 7 C 1 72 C 7 C 7 C 7 C 72 C 7 C 7 C H 2 removal (%) Napthalene yield (%) Figure 7: Effect of hydrogen removal on (a) Methane conversion, (b) Benzene and naphthalene yields at temperatures 7-7 C and 1 atm. GHSV values for all temperature points are 1 ml g 1.

8 Overall reactor performance is typically based on desired product benzene yield. Yield is an important comparison parameter, but is not sufficient to quantify the reactor performance. Another important parameter is the product selectivity. Figure 7b shows that although the benzene yield increases, the selectivity decreases below %, and the methanes value is wasted to form the unwanted product naphthalene. For cases in which the benzene yield is low but selectivity is high, the methane preserves its value as a fuel. Although the single-pass yield is low, multi-pass stage reactors can be designed to further convert methane to benzene. Another important parameter is the benzene formation rate. Typically, due to low inlet GHSV, the benzene formation rates are low ( 1 µmol s 1 ). Figure shows predicted benzene formation rates at GHSV = 1 ml g 1 cat hr 1 for temperatures between 7 and 7 C, while varying hydrogen removal ratio. At each hydrogen removal rate, the benzene formation rates is plotted against the /C H yield ratio. Figure shows that the benzene formation rate can increase up to times with membrane integration. However, in the regimes with the highest benzene formation rates, /C H yield ratios are below., indicating that naphthalene is the primary product. Coke formation is inevitable in this range. The trend is similar for all temperatures studied, indicating that membrane-coupled MDA can not run efficiently at low GHSV values. Figure 9 displays the effect of increasing GHSV at (a) ml g 1 and (b) ml g 1. The effect of membrane integration on benzene yield and formation rate is more pronounced at lower GHSV. Increasing gas-hourly space velocity decreases the benzene yield, but increases the maximum /C H yield ratios as well as increasing the benzene formation rates. Figure, 9a, and 9b show that, independent of temperature and GHSV, membrane integration decreases the /C H yield ratios, due to the decrease in benzene selectivity as hydrogen removal increases. Figures 9a and 9b also show that reactor efficiency is higher at temperatures below 7 C. At lower temperatures, the benzene yield decreases, but the reactor can run for much longer time periods due to the reduced coking propensity, while maintaining reasonable benzene production rates. a) ml g cat hr Increased H 2 removal ratio formation (μmol s ) /C H yield ratio 7 b) ml g cat hr formation (μmol s ) Dashed lines: Yield Solid lines: Rate Non-coking regime 7 C 7 C 72 C 7 C /C H yield ratio yield (%) C yield (%) formation (μmol s ) Increased H 2 removal ratio PAH formation region 7 C 7 C 72 C 7 C /C H yield ratio Figure : Effect of hydrogen removal on benzene formation and product distribution at a GHSV of 1 ml g 1 Figure 9: Effect of GHSV on yield and formation rates in a membrane integrated MDA reactor. (a) ml g 1 cat hr 1 and (b) ml g 1 Considering benzene yield, formation rate, and selectivity as decision criteria, this study suggests that, for a membrane-coupled MDA reactor with the given reactor and catalyst properties, suitable operating conditions are 7 C and ml g 1. Figure provides a comparison between common GHSV and those suggested by this study. Typical gas-hourly space velocities for lab-scale experiments are near 7 ml g 1. Such low GHSVs lead to high benzene yields but suffer from poor selectivity to benzene, causing faster catalyst deactivation. Increasing the GHSV to ml g 1 results in higher benzene formation rates within noncoking regimes. At the safer operating range, where the /C H yield ratio is greater than 2., the benzene yield is around %. At 7 C and ml g 1 cat hr 1, where /C H = 3, the methane conversion is around 9%, benzene selectivity is %, and naphthalene selectivity is %. Half of the hydrogen is removed from the reactor, at an average rate of 1. µmol s 1.

9 formation (μmol s ) 1... Increased H 2 removal ratio 7 ml g cat hr ml g cat hr. ml g cat hr 7 ml g.2 cat hr Non-coking regime /C H yield ratio 1 C yield (%) C 2 species involved in MDA kinetics are ethane (C 2 ) and ethylene (C 2 H ). Figure 12 shows that increasing pressure decreases benzene selectivity while increasing selectivity to the C 2 hydrocarbons. However, the naphthalene selectivity stays nearly constant. From a coking standpoint, the steady naphthalene selectivity is beneficial since independent of pressure, the coking propensity doesn t increase significantly. Considering multipass reactors, C 2 hydrocarbons can be recycled for further conversion to benzene. Figure : Comparison of effect of GHSV on yield and formation rate in a membrane integrated MDA reactor. Darker region indicates the desired operating range. Designing a practical reactor for MDA would require high flow rates and moderate hydrogen removal rates. At theses conditions, conversion would be low, but formation of coking precursors would be relatively reduced, increasing operation time. Staged multi-pass reactors could overcome some of the low-conversion issues. However, the benzene output from each individual reactor would be small Pressure effects Increasing the packed-bed operating pressure reduces the MDA rates of progress toward the products. Typical MDA reactors operate at atmospheric pressure. Higher pressure would decrease naphthalene production, but also decease conversion. Figure 11 shows that as pressure is increases up to 3 atmospheres, the conversion decreases as expected. The benzene yield drops quicker than the naphthalene yield. Conversion or Yield (%) Benzene Yield Conversion Naphthalene Yield Pressure (atm) Figure 11: Pressure effects on conversion and yield at 7 C and 1 ml g 1 At high pressure, unwanted MDA side products, such as the C 2 hydrocarbons, become more prominent. The 9 Low/High Hydrocarbon Ratio Ratio Increasing pressure Benzene Naphthalene C 2 Hydrocarbons 3 7 Conversion (%) 7 3 Selectivity (%) Figure 12: Selectivity dependence on pressure from 1 to 7 atm at 7 C and 1 ml g 1 Figure 12 also shows the ratio of low to high hydrocarbons, where the small hydrocarbons consist of ethylene, ethane, and benzene. As pressure increases, the kinetics begin to favor the higher hydrocarbons (naphthalene) and are less selective to the smaller hydrocarbons. Approaching the highest pressures, the kinetics begin to favor the lower hydrocarbons again, as shown by the local minima. The pressure doesn t have a large effect on the naphthalene selectivity, which indicates that instead of preventing the conversion of benzene to naphthalene, the conversion of C 2 hydrocarbons to benzene is reduced. Benzene formation (μmol s ) Increasing pressure GHSV 7 7 GHSV 1 GHSV 2 1 GHSV 3 7 Conversion (%) Benzene Selectivity (%) Figure 13: Selectivity dependence on pressure from 1 to 7 atm at 7 C and GHSV of 1 and 7 ml g 1

10 Figure 13 shows the effect of gas-hourly space velocity with varying pressure on the benzene formation rates and the benzene selectivity. At low GHSV, the benzene formation rates and selectivity are low, compared to the high GHSV. As the conversion decreases with higher pressure, benzene formation rates increase and selectivity decreases. The higher GHSV consistently shows higher benzene production. Because there are not major losses in benzene selectivity, moderate pressures and flow rates should be considered for high benzene production. 3.. Pressure effects with hydrogen removal Incorporating the high operating pressure with a hydrogen removal membrane can increase conversion and the benzene formation rates. Figure 1 shows the relationship between hydrogen removal, conversion, and low/high hydrocarbon ratio at ml g 1 cat hr 1 and 7 C. The grey area denotes the operating regime where coking is likely. As hydrogen removal increases, the conversion increases, but the low-to-high hydrocarbon ratio favors the coking species. The conversion and hydrocarbon ratio are slightly lower for higher pressures. Fig. 1 hows the relationship between hydrogen removal and conversion. Even though increasing pressure does not increase conversion or benzene selectivity, if indirectly increases the driving force for membrane transport. The Sievert s law correlation for membrane flux depends on the partial pressure of hydrogen in the packed bed. Increasing operating pressure increases hydrogen partial pressure assuming constant hydrogen composition. Conversion (%) 7 7 GHSV = ml g cat hr 7 atm 3 7 atm H 2 removal (%) Low/High Hydrocarbon Ratio a weak function of hydrogen removal rate. As the hydrogen removal is increases with increasing membrane permeance, the benzene production rates increase significantly for the high pressure case. The benzene production rates do not change significantly for the atmospheric pressure case. The benzene selectivity gener- Benzene formation (μmol s ) Low/High Hydrocarbon Ratio Increasing H 2 Removal 7 atm 7 atm GHSV = ml g cat hr 2 7 atm 1 3 Conversion (%) 7 7 Benzene Selectivity (%) Figure 1: Benzene formation rate, selectivity, and hydrocarbon ratio dependence on hydrogen removal and pressure at 7 C and ml g 1 ally decreases with increasing conversion. The C 2 hydrocarbons also decrease as more hydrogen is removed. However, naphthalene selectivity increases as shown by the low/high hydrocarbon ratio, which leads to coking regimes for large hydrogen removal. The major benefit of the hydrogen removal at increased pressure is the increase in benzene molar production rate. Moderate hydrogen removal rates, along with increased pressure and moderate flow rates are best for large benzene production rates and reduced selectivity toward coking species.. Conclusions Figure 1: Conversion and low/high hydrocarbon ratio as a function of percent hydrogen removed at ml g 1 and 7 C. Figure 1 shows the benzene formation rates, selectivity, and hydrocarbon ratio at ml g 1 and pressures of 1 and 7 atm. Benzene production rate is The objective of this paper was to investigate alternative operating conditions for MDA reactors, with an emphasis in hydrogen removal and high pressure. Based on the mechanism developed by Karakaya et al. [13], a detailed chemical kinetics model is coupled with a membrane packed-bed model. Increasing the gas-hourly

11 space velocity decreases conversion but increases selectivity to benzene. Although hydrogen can be removed to increase conversion, coke-inducing species formation increases, reducing benzene selectivity. However, operating at moderate gas-hourly space velocities ( ml g 1 cat hr 1 ) and moderate hydrogen removal rates, benzene is produced within a non-coking regime at improved rates. Increasing pressure significantly increases benzene production rates in non-cooking regimes. Best conditions are near GHSV of ml g 1, around % hydrogen removal, and 7 atm.. Acknowledgements This work was supported by the Air Force Office of Scientific Research (FA92-9) and CoorsTek, Inc. We also acknowledge the insightful discussion with Dr. Grover Coors (CoorsTek Membrane Sciences) and Selene Hernández Morejudo (University of Oslo). [11] N. Wakao and J.M. Smith. Diffusion in catalyst pellets. Chem. Eng. Sci, 17:2 3, 192. [12] A. Srinivasan and C. Depcik. One-dimensional pseudohomogeneous packed-bed reactor modeling: I. chemical species equation and effective diffusivity. Chem. Eng. Technol., 3:22 32, 13. [13] C. Karakaya, S. Hernández-Morejudo, H. Zhu, and R.J. Kee. Catalytic chemistry for methane dehydroaromatization (MDA) on a bifunctional Mo/HZSM- catalyst in a packed bed. Ind. Eng. Chem. Res., :99 99, 1. [1] S.H. Morejudo, R. Zanón, S. Escolástico,, I. Yuste-Tirados, H. Malerød-Fjeld, P.K. Vestre, W.G. Coors, A. Martínez, T. Norby, J.M. Serra, and C. Kjølseth. Direct conversion of methane to aromatics in a catalytic co-ionic membrane reactor. Science, 33:3, 1. [1] S. Natesakhawat, N.C. Means, B.H. Howard, M. Smith, V. Abdelsayed, J.P. Baltrus, Y. Cheng, J.W. Lekse, D. Link, and B.D. Morreale. Improved benzene production from methane dehydroaromatization over Mo/HZSM- catalysts via hydrogenpermselective palladium membrane reactors. Catal. Sci. Technol, :23 3, 1.. References [1] S. Ma, X. Guo, L. Zhao, S. Scott, and X. Bao. Recent progress in methane dehydroaromatization: From laboratory curiosities to promising technology. J. Energy Chem., 22:1, 13. [2] C. Karakaya, H. Zhu, and R.J. Kee. Kinetic modeling of methane dehydroaromatization chemistry on Mo/Zeolite catalysts in packed-bed reactors. Chem. Engr. Sci., 123:7, 1. [3] X. Guo, G. Fang, G. Li, H. Ma, H. Fan, L. Yu, C. Ma, X. Wu, D. Deng, M. Wei, D. Tan, R. Si, S. Zhang, J. Li, L. Sun, Z. Tang, X. Pan, and X. Bao. Direct, nonoxidative conversion of methane to ethylene, aromatics, and hydrogen. Science, 3:1 19, 1. [] S. Majhi, P. Mohanty, H. Wang, and K.K. Pant. Direct conversion of natural gas to higher hydrocarbons: A review. J. Energy Chem., 22:3, 13. [] K.S. Wong, J.W. Thybaut, E.Tangstad, M.W. Stocker, and G.B. Marin. Methane aromatization based upon elementary steps: Kinetic and catalyst descriptors. Micropor. Mesopor. Mat., 1:32 312, 12. [] J.-P. Tessonnier, B. Louis, S. Rigolet, M.J. Ledoux, and C. Pham-Huu. Methane dehydro-aromatization on Mo/ZSM- : About the hidden role of Brønsted acid sites. Appl. Catal. A, 33:79,. [7] L. Wang, L. Tao, M. Xie, G. Xu, J. Huang, and Y. Xu. Dehydrogenation and aromatization of methane under non-oxidizing conditions. Catal. Lett., 21:3 1, [] P.M. Bijani, M. Sohrabi, and S. Sahebdelfar. Thermodynamic analysis of nonoxidative dehydroaromatization of methane. Chem. Eng. Technol., 3:12 132, 12. [9] A.A. Stepanov, L.L. Korobitsyna, and A.V. Vosmerikov. Features of non-oxidative conversion of methane into aromatic hydrocarbons over Mo-containing zeolite catalysts. Energ. Environ. Sci., 3:1, 1. [] L. Li, R. Borry, and E. Iglesia. Reaction-transport simulations of non-oxidative methane conversion with continuous hydrogen removal homogeneous-heterogeneous reaction pathways. Chem. Engr. Sci., :19 11, 1. 11