HZSM-5: Mono- or Bifunctional

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Methane Dehydroaromatization by Mo/ HZSM-5: Mono- or Bifunctional Catalysis? Nikolay Kosinov, Ferdy J.A.G. Coumans, Evgeny A. Uslamin, Alexandra S. G. Wijpkema, Brahim Mezari, and Emiel J. M. Hensen ACS Catal., Just Accepted Manuscript • DOI: 10.1021/acscatal.6b02497 • Publication Date (Web): 05 Dec 2016 Downloaded from http://pubs.acs.org on December 5, 2016

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Methane Dehydroaromatization by Mo/HZSM-5: Mono- or Bifunctional Catalysis? Nikolay Kosinov*, Ferdy J.A.G. Coumans, Evgeny A. Uslamin, Alexandra S.G. Wijpkema, Brahim Mezari and Emiel J.M. Hensen*

Schuit Institute of Catalysis, Laboratory of Inorganic Materials Chemistry, Eindhoven University of Technology, PO Box 513, 5600 MB Eindhoven, The Netherlands

Corresponding authors: Nikolay Kosinov Eindhoven University of Technology PO Box 513, 5600 MB Eindhoven, The Netherlands Tel: +31-40-2478156 Email: [email protected]

Emiel J.M. Hensen Eindhoven University of Technology PO Box 513, 5600 MB Eindhoven, The Netherlands Tel: +31-40-2475178 Email: [email protected]

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Abstract The active site requirements for methane dehydroaromatization by Mo/HZSM-5 were investigated by employing as catalysts physical mixtures of Mo-bearing supports (HZSM-5, SiO2, γ-Al2O3 and activated carbon) and HZSM-5. Separation of the two catalyst components after activation or reaction was possible by using two different sieve fractions. Our comparison demonstrates that migration of volatile Mo-oxides into the micropores of HZSM-5 is at the origin of the observed catalytic synergy in methane dehydroaromatization for physical mixtures. The propensity of Mo migration depends on the activation method and the Mo-support interaction. Migration is most pronounced for Mo/SiO2. Prolonged exposure of HZSM-5 zeolite to Mo-oxide vapors results in partial destruction of the zeolite framework. Mo-carbide dispersed on non-zeolitic supports afforded predominantly coke with only very small amounts of benzene. The main function of the zeolite is to provide a shape-selective environment for the conversion of methane to benzene. A comparison of Mo/HZSM-5 and Mo/Silicalite-1 demonstrates that aromatization of methane is an intrinsic ability of molybdenum carbides dispersed in the 10membered ring micropores of MFI zeolite. Thus, one important role of the Brønsted acid sites is to promote the dispersion of the Mo-oxide precursor and, accordingly, the active Mo-carbide phase in the micropores of HZSM-5.

Keywords: methane; aromatization; bifunctional mechanism; shape selectivity; Mo/ZSM-5

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Table of Contents Graphic

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Introduction New and improved technologies for natural gas production have led to lower methane pricing and renewed interest in converting the smallest of all hydrocarbons to more valuable products such as building blocks for polymers (olefins) and liquid fuels (mainly paraffins). All of the current commercial processes that convert methane to other chemicals are indirect and convert methane into a mixture of carbon monoxide and hydrogen (syngas) by steam reforming, partial oxidation or combinations thereof, prior to syngas upgrading to fuels or chemicals. Although highly optimized, these processes require economy of scale. In Fischer-Tropsch processes, for example, syngas generation is the most expensive step in the conversion of gas to liquid fuels. As methane is present at many places around the globe in relatively small fields, it would be attractive to have a direct and more flexible process that could better monetize these often remote sources. The shale-gas boom has added further incentive to develop such a process. Many approaches have been explored to convert methane to light olefins (e.g., oxidative coupling [1]), aromatics (mainly benzene [2]) or oxygenates (e.g., methanol [3,4]). Non-oxidative conversion of methane shows greater promise than oxidative approaches, mainly because of the high reactivity of products of the latter approach, such as ethylene and methanol, towards oxygen. Non-oxidative methane conversion includes complete (non-)catalytic pyrolysis to produce elemental carbon and hydrogen [5,6], non-oxidative coupling [7,8] and methane aromatization [9,10,11]. Aromatization of methane to benzene has attracted widespread attention from academia and industry. Thermodynamics dictate that appreciable methane conversion can only be achieved above 650 °C with carbon being the favored product. It is thus imperative to use a catalyst that can activate the C-H bonds in methane and provide good selectivity towards benzene. Mo/HZSM-5 has been the most widely studied catalyst, affording typical benzene selectivities between 60 and 80% at a close to equilibrium methane conversion of 10-12% at 700 °C [12]. Methane dehydroaromatization (MDA) provides a mixture of benzene, toluene, naphthalene and ethane/ethylene as well as hydrogen. A major by-product of the reaction is coke (polyaromatic and graphitic species), which causes rapid deactivation of the Mo/HZSM-5 zeolite catalyst [13,14,15,16,17]. The mechanism of the MDA reaction on Mo-containing zeolites has not been resolved yet. The majority of studies adhere to a bifunctional mechanism involving Mo-carbide or – oxycarbide sites and Brønsted acid sites of the zeolite as the two reaction centers [18,19,20,21,22,23,24]. The Mo-carbide phase activates the C-H bond in methane and converts the resulting CHx intermediates into ethylene. Ethylene is then oligomerized followed by cyclization of the intermediates to benzene and other aromatic products on Brønsted acid sites embedded in the zeolite micropores [25,26,27]. The alternative monofunctional mechanism has been discussed by Mériaudeau et al., emphasizing activation

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of methane to ethylene or acetylene, followed by aromatization into benzene exclusively over Mo-carbide [28,29,30]. The bifunctional nature of the MDA reaction has been investigated before by physically separating the methane activation function from the acid function. Combinations of HZSM-5 with Mo/SiO2 [31], Mo2C [32] and Mo2C/α-Al2O3 [33] display synergy compared with the individual components. Chen et al. reported that a physical mixture of MoO3 and HZSM-5 was more active than a composite Mo/HZSM-5 catalyst [34]. A recent patent describes a process based on physical separation of the Mo and the acidic function in such way that more effective regeneration of the individual components is possible [35]. On contrary, Wang et al. did not observe any synergy when using a physical mixture of Mo/SiO2 and HZSM-5 [36]. This study and others emphasize the importance of Mo-carbide species embedded in the zeolite micropore space and their association with Brønsted acid sites [37,38,39,40,41]. Many investigations have demonstrated that, in the calcined precursor, part of the MoOx phase is present as highly dispersed cationic complexes replacing protons and that reduction/carburization may partly or completely disturb this interaction [42,43,44,45]. All of these studies hint at the benefit of close proximity of Mo-species and the Brønsted acid sites to obtain good catalytic performance. This proximity requirement may relate to sufficient supply of ethylene to the Brønsted acid sites to attain high rate of oligomerization. It is also conceivable that the Brønsted acid sites mainly serve to anchor the oxidic precursor and possibly the (oxy)carbidic Mo phase during what would then be a monofunctional mechanism by Mo-carbides. In that view, the location of the Mo-carbide species in the micropores is essential to the shape-selective conversion of methane into benzene. We endeavored to resolve the apparent discrepancies arising from previous experiments using physical mixtures by taking a slightly different approach. We prepared physical mixtures of HZSM-5 and various Mo-oxide precursors on different supports (HZSM-5, SiO2, γ-Al2O3 and activated carbon). By choosing different particle sizes for these two components, the mixtures could be separated after the reaction. The upshot of these experiments will be that migration of MoO3 during the heating stage of the precursor from the Mo/support component to HZSM-5 explains the apparent synergistic effect of physical mixtures in the MDA reaction. That is to say, there is no actual synergy in the physical mixtures and the main requirement for high MDA performance is that the Mo-carbide species are dispersed in the micropores of HZSM-5. We discuss the implications of these results and demonstrate, by using Silicalite-1 instead of HZSM-5 as the support, that Brønsted acid sites are not necessary for the methane aromatization, once the Mo-carbide phase is embedded in a suitable shape-selective environment.

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Experimental Catalyst preparation As starting materials, HZSM-5 (Si/Al = 13, Süd-Chemie, now Clariant), HMOR (Si/Al = 11, mordenite, AkzoNobel, now Albemarle), HSSZ-13 (Si/Al 20, synthesized according to [46]), silicalite-1 (synthesized according to [47]), silica (Grace 643), γ-alumina (CK-300, Ketjen, now Albermarle) and activated carbon (RX-3-extra, Norit) were used. Incipient wetness impregnation was used to load Mo on these support materials using aqueous solutions of ammonium heptamolybdate tetrahydrate ((NH4)6Mo7O24·4H2O, Merck). The Mo loading as MoO3 was 5 wt% for the zeolites and 10 wt% for the other supports. After impregnation, the samples were dried overnight at 110 °C and then calcined in air at 550 °C for 6 h with the exception of Mo/C which was calcined at 250 °C. The ramp rate during heating was 2 °C/min in all cases. In the following, the resulting catalysts are denoted as Mo/HZSM-5, Mo/HMOR, Mo/HSSZ-13, Mo/Silicalite-1, Mo/SiO2, Mo/Al2O3 and Mo/C. Catalyst characterization The elemental composition of the samples was analyzed by ICP-OES (Spectro CIROS CCD ICP optical emission spectrometer with axial plasma viewing). Prior to measurement, the samples were calcined in air at 550 °C and then dissolved in an 1:1:1 (by weight) mixture of HF (40%):HNO3 (60%):H2O. The amount of coke deposited on spent samples was analyzed by a Mettler Toledo TGA/DSC 1 instrument. For this purpose, 25 mg of sample was placed in an alumina crucible. Uncovered crucibles were heated to 750 °C at a rate of 5 °C/min under 40 mL/min He + 20 mL/min O2 flow. In the case of Mo/C catalysts the amount of formed coke was calculated from the decrease of the remaining inorganic residue. The acidity of the samples was evaluated using infrared spectroscopy of adsorbed pyridine. Spectra were recorded in the 4000 – 400 cm-1 range using a Bruker Vertex 70V spectrometer. Samples were pressed to obtain self-supporting wafers (10-15 mg, diameter 1.3 cm)) and then placed in the environmental cell. After evacuation, the sample was kept at 550 °C for 2 h to remove adsorbed species, followed by cooling to 50 °C. A background spectrum was measured, before exposure to pyridine. The sample was then exposed to pyridine until saturation as followed by IR. Finally, a spectrum was recorded at 50 °C after evacuation of the sample at 350 °C for 1 h. XRD was performed on a Bruker D2 powder diffraction system (Cu Kα radiation, scan speed 0.01 °/sec, 2θ range 5 to 60 °). The relative crystallinity of samples normalized to pristine HZSM-5 was evaluated by summing the areas of the reflections (planes indicated in brackets) at 23.1 ° (051), 23.3 ° (501/501), 23.7 ° (-511/511), 24.0 ° (033) and 24.4 ° (-313/313). 6


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C MAS NMR spectra of coked catalysts were recorded on a 4.7 Tesla Bruker DRX-200 NMR

spectrometer operating at 50.3 MHz. The measurements were performed using a 4-mm MAS probe head with a sample rotation rate of 8-10 kHz. Quantitative spectra were recorded using the high-power proton decoupling direct excitation (HPDEC) pulse sequence with an interscan delay of 10 s. The spectra were referenced to the 13C signal of adamantane at 38.56 ppm. TEM micrographs were obtained with a FEI Tecnai 20 instrument at an electron acceleration voltage of 200 kV. The surface of the materials was probed by X-ray photoelectron spectroscopy (XPS) using a Thermo Scientific K-alpha spectrometer equipped with a monochromatic Al Kα X-ray source and a 180° doublefocusing hemispherical analyzer with a 128-channel detector. Spectra were obtained using an aluminum anode (Al Kα = 1486.6 eV) operating at 72 W with a spot size of 400 µm and pass energy of 200 eV. Catalytic activity measurements Methane aromatization was carried out in a quartz reactor (internal diameter 4 mm; length 400 mm) in downflow mode. In a typical experiment, 0.3 g catalyst was placed in the reactor between quartz wool to support the catalyst. Thermal mass flow controllers were used to supply the gases to the reactor. Reactions were performed in a 15 ml/min flow of CH4/N2 (95/5). N2 served as the internal standard. The products were analyzed using an online gas chromatograph (Interscience Compact GC) equipped with three separate columns and detectors. A thermal conductivity detector (TCD) coupled with a Molsieve 5A column was used to analyze light gases (H2, N2, CH4 and CO). Light hydrocarbons (ethane, ethylene) were analyzed with an RT-Q-BOND column equipped with a TCD, and higher hydrocarbons (benzene, toluene, naphthalene) were analyzed using an FID on an Rtx-1 column. In experiments involving physical mixtures, the fresh Mo-containing samples were mixed with HZSM-5 particles in a 1:1 ratio (by weight). The particle size of the Mo-containing fraction was 0.1-0.2 mm obtained by pelletizing, crushing and sieving. HZSM-5 particles were obtained in the 0.3-0.5 mm size range in the same manner. Catalytic performance of the Mo-containing catalysts (without HZSM-5) was tested by mixing the sample with SiC instead of HZSM-5 to obtain a similar catalyst weight (experiments denoted as SiC). Prior to the reaction, the catalyst bed was either directly pre-heated to 700 °C in methane (denoted as CH4) or it was pre-heated in a He flow at 700 °C for 1 h, 3 h or 6 h (denoted as He-1, He-3 and He-6) followed by exposure to methane. In all cases, the reaction conditions for methane aromatization were atmospheric pressure, a temperature of 700 °C, a methane-based WHSV of 2 h-1 and a total reaction time of 16 h. After cooling to room temperature, the spent catalyst mixture was separated into two fractions using a 0.3 mm sieve. Fig.1 illustrates the approach to separate the spent physical catalyst mixture into the two original fractions. 7


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Figure 1. Illustration of physical mixture separation principle employed in this work.

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Results and discussion Physical mixtures Temperature-programmed surface reaction (TPSR) in methane was used to study the reduction/carburization behavior of the Mo-containing samples. Fig. 2 shows CH4 conversion and C6H6 formation during CH4 TPSR. CH4 consumption started around 600 °C for all materials. The amount of consumed CH4 was somewhat lower for Mo/C than for the other catalysts, which may point to involvement of the carbon support in Mo carburization [48]. Only Mo/HZSM-5 produced a significant amount of benzene. Very small amounts of benzene were observed for the other samples. Benzene formation on Mo/HZSM-5 started after a delay with respect to the onset of carburization (~650 °C). This indicates that MoCx species are the active sites for methane activation in line with literature [49,50].

Figure 2. CH4 TPR profiles of corresponding catalysts (solid lines) and benzene yields obtained in these experiments (dashed). Conditions: 5 °C/min, 15 ml/min of CH4:N2 (95:5), 0.15 g of catalyst + 0.15 g of SiC.

Figure 3 shows the benzene yield as a function of time on stream for the Mo/SiO2 and Mo/HZSM-5 containing physical mixtures. Corresponding plots for physical mixtures of Mo/C and Mo/Al2O3 with HZSM-5 are given in the Supporting Information (Figs. S1,S2). HZSM-5 and SiC cannot activate methane under the chosen reaction conditions (Table S1). The catalytic performance of the Mo-containing samples was determined in a bed using SiC as diluent in order to maintain a similar total bed weight. The time on stream behavior observed in these experiments is typical for Mo/HZSM-5. After an initial increase of the benzene yield, the activity of the catalyst strongly decreased. The various aspects of the strong deactivation observed for such catalysts have already been extensively discussed in the literature [51,52,53,54,55]. The time on stream plots show that the catalysts behave differently in their physical mixture with HZSM-5 than as a single catalyst. To better appreciate these performance differences, we 9


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integrated the amount of methane converted and the products – benzene, toluene, naphthalene, aliphatics (mainly ethylene) and coke – formed during 16 h on stream (Fig. 4). As expected, Mo/HZSM-5 was the best catalyst for the production of aromatics from methane. In total, 22.7 mmol CH4 gcat-1 was converted and the overall benzene selectivity was 52%. The Mo catalysts prepared using the non-zeolitic supports produced minor amounts of aromatics and aliphatics. Coke was the dominant product of methane conversion. Despite the significant differences in product selectivity, all catalysts can activate methane and the activity decreased in the order Mo/HZSM-5 >> Mo/SiO2 > Mo/Al2O3 > Mo/C.

Figure 3. Benzene yield as a function of time on stream for (a) Mo/HZSM-5+HZSM-5 and (b) Mo/SiO2+HZSM-5 series of physical mixture catalysts. Conditions: 700°C, atmospheric pressure, 0.3 g of catalyst mixture (1:1), WHSV 2.0 h-1.

The catalytic activity and selectivity considerably changed upon mixing the Mo-containing samples with HZSM-5 zeolite. Direct activation of the physical mixture by pre-heating in methane increased the total amount of methane converted as compared to the experiments without HZSM-5. The influence of mixing with HZSM-5 was much larger for Mo/SiO2 and Mo/HZSM-5 than for Mo/C and Mo/Al2O3. In all cases, the aromatics selectivity of the physical mixtures was higher than in the corresponding reference measurement. As we surmised that sublimation of Mo-oxides from the Mo-containing sample to HZSM-5 caused the performance increase, we varied the pretreatment by holding the physical mixtures in He flow at 700°С for varying times, before replacing the He flow by the CH4/N2 reactant flow. This procedure further enhanced the catalytic activity and the selectivity for all physical mixtures.

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Figure 4. Distribution of hydrocarbon products and coke obtained after 16h MDA runs. Conditions: 700°C, atmospheric pressure, 0.3 g of catalyst mixture (1:1), WHSV 2.0 h-1, 16h.

The activity of the Mo/SiO2+HZSM-5 mixture held in He for 1 h was higher than that of the Mo/HZSM-5 reference zeolite. At the same time, the amount of aromatics was higher and the amount of coke slightly less. Prolonging the hold time in He before starting the reaction had a negative effect on activity and aromatics selectivity. A similar trend was observed for the physical mixture of Mo/HZSM-5 and HZSM5. The best catalytic performance was achieved after a He hold time of 1 h. Lower performance was observed for longer hold times. A difference with the Mo/SiO2+HZSM-5 case is that the increased activity is nearly completely due to increased coke formation. It is important to note that the catalytic performance of the Mo/SiO2+HZSM-5 and Mo/HZSM-5+HZSM-5 mixtures after a hold time in He of 1 h were very similar. The spent physical mixtures were then characterized to understand the synergetic effect observed in methane dehydroaromatization. For this purpose, the larger HZSM-5 fraction was separated from the fraction of Mo-containing catalyst (Fig. 1). Elemental analysis of the former fraction showed that the Si/Al ratio varied only slightly in the various experiments (Fig. 5a), demonstrating that there was no measurable contamination between the two fractions. Further evidence for the proper separation of the two fractions is provided in the Supporting Information (Figs. S3-S4 and discussion therein).

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Figure 5. Si/Al ratio (a) and Mo content (b) of HZSM-5 fractions obtained after separation of corresponding MDAspent physical mixtures. The nominal Si/Al ratio of the parent HZSM-5 sample is 13. Error bars for Mo/SiO2+HZSM-5 samples were derived from 3 independent measurements.

Fig. 5b shows the Mo content of the HZSM-5 fraction following catalytic testing in a mixture with Mocontaining samples. Very small amounts of Mo were observed on HZSM-5 after catalytic testing in mixtures with Mo/C or Mo/Al2O3. The concentration of migrated Mo did not appreciably increase after 1 h for these samples. In contrast, the amount of migrated Mo increased with the He hold time when Mo/HZSM-5 and Mo/SiO2 were mixed with HZSM-5. The total amount of Mo migrated from Mo/SiO2 was much larger than from Mo/HZSM-5. Mo migration occurs due to sublimation of MoO3, whose vapor pressure is about 13 Pa at 700 °C [56]. On contrary, Mo-carbides are not volatile at the used temperature, explaining why Mo migration is nearly absent for directly reduced/carburized samples. The difference in the amount of migrated Mo between the various samples is related to the interaction of Mo-oxide with the support. The volatility of the Mo-oxide phase on the different supports was confirmed by TGA analysis of the Mo-containing samples in He. Fig. 6 shows that for Mo/Al2O3 and Mo/HZSM-5 samples Mo-oxides sublime above 800 °C (the weight loss at lower temperature is due to loss of associated water). These results attest to the relatively strong interaction of Mo-oxide with γ-Al2O3 and HZSM-5 zeolite [57,58].

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Figure 6. TG (a) and DTG (b) profiles of supported Mo catalysts. Conditions: 10 °C/min, 40 ml/min He flow.

The much weaker interaction between Mo-oxide and silica is evident from more pronounced weight loss at lower temperatures. In fact, the weight loss starts at about 730 °C, which corresponds well with the sublimation temperature of bulk MoO3 (Figs. S5-S6). For Mo/C, there is a gradual weight loss already between 350 °C and 750 °C and then a sharp decrease at 760 °C. The low temperature weight loss is due to decomposition of ammonium molybdates [59]. Unlike the other samples, the Mo/C sample was calcined at lower temperature (250 °C) during preparation and the molybdate precursor had not been fully decomposed. The other feature at higher temperature is due to carbothermal reduction of MoO3 into Mo2C in an inert flow, which is characterized by a weight loss of about 45% [60]: 2 MoO3 + 7 C → Mo2C + 6 CO By taking the characterization and activity data together, we can conclude that the migration of Mo-oxide from Mo/SiO2 to HZSM-5 strongly improves the catalytic performance of the Mo/SiO2+HZSM-5 mixture. Strong interaction of MoO3 with γ-Al2O3, carbon and HZSM-5 limits MoO3 sublimation at 700 °C and, accordingly, its migration to HZSM-5. In these cases, it is likely that the Mo-oxide precursor is already at least partially carburized before sublimation can occur. The observation that optimum catalytic performance of the Mo/SiO2+HZSM-5 and Mo/HZSM5+HZSM-5 mixtures was reached after a hold time of 1 h, whereas the Mo content of HZSM-5 was highest after a hold time of 6 h demands an explanation. For this purpose, we recorded XRD patterns of the spent zeolite fractions obtained from Mo/SiO2+HZSM-5 and Mo/Al2O3+HZSM-5 mixtures and both components of the Mo/HZSM-5+HZSM-5 mixture. Fig. 7 compares the crystallinity as determined by comparison of the intensity of the five dominant diffraction peaks of MFI relative to the corresponding

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peaks in the parent HZSM-5 zeolite (in case of Mo/HZSM-5 samples the crystallinity was related to the as-calcined Mo/HZSM-5 catalyst). Clearly, prolonged exposure of HZSM-5 to MoO3 vapor deteriorates the zeolite structure. As demonstrated by Ma et al., progressive dealumination of the framework due to reaction with Mo-oxides results in the formation of Al2(MoO4)3 [61]. In comparison, the decrease in the crystallinity of the HZSM-5 fraction that was in contact with Mo/Al2O3 was much less pronounced. Thus, the strong decrease of the performance for hold times longer than 1 h is due to the interaction of Brønsted acid sites with volatile Mo-oxides, leading to a gradual degradation of the zeolite framework with increasing Mo content (Fig. S7 and discussion therein). In turn, as Fig. S8 shows there was no optimum in the catalytic performance, when pure Mo/HZSM-5 catalyst was kept under He at 700 °C for increasing durations. Thus, the observed maximum for Mo/HZSM-5+HZSM-5 (He-1) sample can be exclusively attributed to the migration of some Mo to the pristine zeolite phase, while the crystal structure of Mo/HZSM-5 was still relatively intact.

Figure 7. Relative crystallinity of HZSM-5 fractions derived from XRD: a) HZSM-5 fractions separated from physical mixtures with Mo/Silica and Mo/Alumina and b) HZSM-5 and Mo/HZSM-5 fractions separated from each other.

To put these conclusions on surer footing, we performed an additional experiment. The Mo/SiO2+HZSM5 physical mixture was kept under He flow at 700 °C for 1 h, followed by cooling the reactor to room temperature and separating the fractions. The zeolite fraction was mixed with SiC and then evaluated in the MDA reaction. Fig. 8 shows that the catalytic results are very similar to the experiment using the physical mixture, confirming that the reaction is mainly catalyzed by the HZSM-5 fraction to which part of the Mo-oxide has migrated.

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Figure 8. Distribution of hydrocarbon products and coke obtained after MDA tests for Mo/SiO2+HZSM-5 and the same mixture after removal of the Mo/SiO2 fraction. Conditions: 700°C, atmospheric pressure, 0.3 g of catalyst mixture (1:1), WHSV 2.0 h-1, 16 h.

The goal of this work was to better understand the synergy observed in physical mixtures of HZSM-5 with a supported Mo-oxide precursor in methane dehydroaromatization. The results clearly show that the earlier reported synergy in such physical mixtures [28-31] can be mainly attributed to Mo migration from the Mo-bearing support to HZSM-5 via sublimation of Mo-oxides. Thus, we conclude that close proximity of Mo-carbides or Mo-oxycarbides and Brønsted acid sites is a requirement for good MDA performance. The ability of Mo-oxide to migrate depends on the Mo-support interactions, the heating procedure and the gas in which the precursor is heated. Heating in CH4 leads to carburization of Mo-oxide and limits Mo migration. The supported Mo-oxides by themselves can be carburized, but the resulting carbides mainly convert methane into coke, presumably because of the absence of a shape-selective environment. All of these Mo-carbides can convert methane at comparable rates, while only Mo/HZSM-5 yields appreciable amounts of benzene. Our data also show that the migration of Mo-oxide leads to dispersion of the volatile species inside the micropores through their interaction with Brønsted acid sites. Moreover, in line with earlier works [62,63] the interaction of the framework Al sites in the zeolite with MoO3 vapor results in dealumination, formation of extraframework Al-molybdates and eventually degradation of the zeolite crystal structure. This explains why prolonged exposure of HZSM-5 to MoO3/SiO2 leads to a decreased catalytic performance, suggesting that a lower Mo loading is preferred in maintaining the zeolite structure. Altogether, these data provide strong evidence for the requirement of close proximity between the Mo-carbide particles and Brønsted acid sites in a shape-selective zeolite

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micropore environment. At this stage, we wished to gain more insight into the requirements of the zeolite component. Influence of zeolite topology For this purpose, we prepared three additional catalysts from different zeolites: (i) pure-silica MFI (Silicalite-1, 10MR) and HZSM-5 5.1 x 5.5 Å pores; (ii) HSSZ-13 (CHA, 8MR) with pore entrances of 3.8 Å x 3.8 Å separating larger cages of 12 Å; and (iii) HMOR with a one-dimensional system of pores of 7.0 Å x 6.5 Å. These zeolites were loaded with ammonium heptamolybdate to obtain a Mo loading of 5 wt%, calcined at 550 °C and tested in the MDA reaction. Fig. 9 shows the methane conversion, benzene yield and overall distribution of products obtained during the test. Clearly, only the shape selective environment of the 10MR zeolites (HZSM-5 and Silicalite-1) is suitable for the conversion of methane to benzene. Mo/HMOR displayed low hydrocarbon selectivity, which may be due to extensive coke formation in the large pores that do not provide the shape-selective environment for benzene formation. The low hydrocarbon yield of Mo/HSSZ-13 is most likely related to the inability of benzene to leave the small CHA pore openings. The coke formation rate of all the catalysts is comparable in line with the results presented above. It should be noted that all the studied aluminosilicate catalysts have comparable Brønsted acidity (Fig. S9), [46]. The most interesting result is the considerable benzene selectivity displayed by Mo/Silicalite-1 (Fig. 9b). The Silicalite-1 sample was prepared with tetraethylorthosilicate as the Si source, which implies that the Al content is negligible. Accordingly, the sample does not contain Brønsted acid sites (Fig. S9).

Figure 9. Methane conversion (a), benzene yields (b) and overall product distributions (c) obtained during 16h of MDA tests with different 5%Mo/zeolite catalysts. Conditions: 700°C, atmospheric pressure, 0.3 g of catalyst, CH4 WHSV 2.0 h-1, no dilution with SiC.

These results demonstrate that Mo-carbides embedded in properly sized micropores can catalyze the aromatization of methane to benzene, that is to say, Brønsted acid sites are not a prerequisite. We 16


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investigated the chemical nature of the carbonaceous deposits on spent Mo/Silica, Mo/Silicalite-1 and Mo/HZSM-5 catalysts by

13

C MAS NMR. The resulting spectra shown in Fig. 10 are dominated by a

signal in the 124-128 ppm range due to aromatic carbon. The weak feature around 275 ppm is due to carbon from molybdenum carbide [64]. The polyaromatic nature of coke formed on Mo/Silicalite-1 and Mo/SiO2 shows that aromatization of methane is an intrinsic property of molybdenum carbide.

Figure 10. 13C MAS NMR spectra of spent Mo/SiO2, Mo/Silicalite-1 and Mo/HZSM-5 samples. Number of scans for Mo/Silicalite-1 and Mo/SiO2 was 8192, for Mo/HZSM-5 – 6144. The spectra are normalized by number of scans and sample weight. Asterisks denote spinning sidebands. Pretreatment conditions: 700°C, atmospheric pressure, 0.3 g of catalyst, CH4 WHSV 2.0 h-1, no dilution with SiC.

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Figure 11. Mo/(Si+Al) XPS ratio as a function of the Mo loading for Mo/Silicalite-1 and Mo/HZSM-5. Bulk ratios are also given. The samples were prepared by impregnation followed by calcination in air at 550°C for 6 h.

To understand the superior catalytic performance of Mo/HZSM-5, we compared Mo/HZSM-5 and Mo/Silicalite-1 zeolites with varying Mo loading (0.25% - 5%) pretreated by calcination in air at 550 °C. XPS was used to analyze the location of the Mo species. As XPS is a surface-sensitive (several nm) technique, it can be used to probe the diffusion of Mo-oxides into the micropores of the much larger zeolite crystals. Fig. 11 shows how the amount of Mo at the external zeolite surface changes with Mo content for the two sets of calcined zeolites (see Table S2 for numerical data). Clearly, the surface of the Silicalite-1 crystals contains much more Mo than HZSM-5 irrespective of the Mo loading. This shows that Brønsted acid sites in HZSM-5 stabilize Mo-oxide species in the micropores, shifting the distribution of the Mo-oxide phase between the external zeolite surface and the micropore space. Fig. 12 shows the catalytic results in the MDA reaction for these catalysts. The time on stream behavior is comparable for Mo/HZSM-5 and Mo/Silicalite-1. The activities of Mo/HZSM-5 are higher than those of Mo/Silicalite-1. In Fig. 12c, we compare the maximum reaction rate of benzene normalized on the Mo loading. The equilibrium line in Fig. 12c corresponds to the reaction rate for which the equilibrium between methane and benzene is achieved. For Mo/HZSM-5, the reaction rate first increased with Mo loading up to a loading of about 1 wt% and then decreased due to the progress towards the equilibrium. This implies that active Mo-carbide sites for methane activation were obtained with increasing Mo loading. The trend was very different for the Mo/Silicalite-1 zeolites. The amount of active Mo-carbide sites was too low to reach the equilibrium reaction rate irrespective of the Mo loading. With decreasing Mo loading, the reaction rate increased, which shows that the Mo-carbide phase should be very different from the one in Mo/HZSM-5. This is consistent with the very different product distributions between the

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ACS Catalysis

two sets of samples. The low dispersion of the Mo-carbide phase in Mo/Silicalite-1 is due to its predominant location on the external zeolite surface. This explains the high coke yield. Only a small part of Mo is embedded in the zeolite, which is responsible for benzene formation. As a much greater part of Mo is in the zeolite micropores, the Mo/HZSM-5 samples yield much more benzene. For both sets, it is observed that the benzene selectivity increased with Mo loading (Fig. S10).

Figure 12. Catalytic performance of Mo/HZSM-5 and Mo/Silicalite-1 series of catalysts: (a) and (b) benzene yield vs time on stream, (c) maximum benzene formation rate and (d) total distribution of hydrocarbon products and coke (note different y-axes for Mo/HZSM-5 and Mo/Silicalite-1). Conditions: 700°C, atmospheric pressure, 0.3 g of catalyst, CH4 WHSV 2.0 h-1, no dilution with SiC.

The catalytic results also show that the Mo/Silicalite-1 catalysts deactivate much faster than the Mo/HZSM-5 ones. As overall coke formation on Mo/Silicalite-1 is comparable to or lower than on Mo/HZSM-5, we speculate that the faster deactivation in Mo/Silicalite-1 is due to the weaker interaction of the active Mo-carbide phase in the micropores and the zeolite surface. This is confirmed by TEM analysis of spent samples (Fig. 13), showing formation of large Mo-carbide agglomerates on the surface of 1%Mo/Silicalite-1. On contrary, on spent 1%Mo/HZSM-5 such agglomerates were not observed.

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These differences emphasize the importance of the interaction between the Mo-(oxy)carbide phase and Brønsted acid sites in HZSM-5. This also suggests that the conversion of dispersed Mo-oxo complexes bound to the zeolite exchange sites to Mo2C might not be complete.

Figure 13. TEM micrographs of (a,b) spent 1%Mo/Silicalite-1 and (c) spent 1%Mo/HZSM-5 catalysts. Pretreatment conditions: 700°C, atmospheric pressure, 0.3 g of catalyst, CH4 WHSV 2.0 h-1, no dilution with SiC.

In summary, Mo-(oxy)carbides can catalyze the conversion of methane to ethylene, benzene, toluene, naphthalene and heavier polyaromatic compounds. Placing the Mo-(oxy)carbides is important for two reasons. First, conversion of methane on the active phase in the confinement of the 10MR zeolite shifts the selectivity to benzene. Clearly, Mo-(oxy)carbides at the external surface of the zeolite mainly produce heavy hydrocarbons and coke deposits. Second, the dispersion of the active phase is much higher, resulting in much higher reaction rate. Brønsted acid sites are essential, because they stabilize dispersed Mo-oxide species and, in this way, shift the distribution of the Mo-oxide precursor phase towards the micropores. Moreover, we found indications that the interaction remains after conversion of the Mo-oxo phase into an (oxy)carbidic one, preventing extensive sintering. Based on these observations, we argue that it might be that the mechanism of the MDA reaction is monofunctional. However, considering the role that Brønsted acid sites can play in converting olefins into aromatics [65,66], at this stage we cannot completely exclude that acid sites are also involved in olefin conversion reactions.

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ACS Catalysis

Conclusions In this work we have demonstrated that the apparent synergy obtained by physically mixing a Mocontaining non-zeolitic support with HZSM-5 zeolites originates from the migration of Mo-oxide into the micropores of HZSM-5 zeolite. Mo migration strongly depends on the metal-support interaction and is greatest for Mo/SiO2. Prolonged exposure of HZSM-5 to MoO3 vapor leads to partial destruction of the zeolite framework. Accordingly, we conclude that an essential requirement for methane dehydroaromatization is the embedding of the Mo-carbide function in the zeolite micropores. At 700 °C Mo-carbide species can convert methane into a range of aromatic products. Good benzene selectivity can only be obtained in the shape-selective environment provided by 10MR zeolites, without shape selectivity the main product is coke. The key function of the Brønsted acid sites is to facilitate the dispersion of Mospecies into the micropores before and during the reaction. From the observation that Mo-carbide dispersed in (non-acidic) Silicalite-1 can also convert methane into benzene and aromatic coke, we infer that Brønsted acid sites are not required for methane dehydroaromatization. Thus, it might be that the conversion of methane to benzene takes place in a monofunctional manner on highly dispersed Mocarbide species embedded in the 10MR zeolite micropores.

Supporting Information Available Detailed information regarding catalytic performance, TGA, XRD, FTIR and XPS. This material is available free of charge via the Internet at http://pubs.acs.org.

Acknowledgements Financial support from the Sabic-NWO CATC1CHEM CHIPP project is gratefully acknowledged. We thank Dr. Christoph Dittrich (SABIC), Dr. Frank Mostert (SABIC), Dr. Xander Nijhuis (SABIC), Prof. Dr. Jorge Gascon (TU Delft) and Ina Vollmer (TU Delft) for fruitful discussion.

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HZSM-5: Mono- or Bifunctional

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