Shape selective methanol to olefins over highly thermostable DDR catalysts

Applied Catalysis A: General 391 (2011) 234–243

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Applied Catalysis A: General journal homepage: www.elsevier.com/locate/apcata

Shape selective methanol to olefins over highly thermostable DDR catalysts Yasukazu Kumita a,b , Jorge Gascon a,∗ , Eli Stavitski a,c , Jacob A. Moulijn a , Freek Kapteijn a a

Catalysis Engineering – Chemical Engineering Department, Delft University of Technology, Julianalaan 136, 2628 BL Delft, The Netherlands Global R&D – Processing Development, Kao Corporation 1334, Minato Wakayama 640-8580, Japan c Inorganic Chemistry and Catalysis Group, Debye Institute for Nanomaterials Science, Utrecht University, Sorbonnelaan 16, 3584 CA Utrecht, The Netherlands b

a r t i c l e

i n f o

Article history: Received 22 February 2010 Received in revised form 12 May 2010 Accepted 13 July 2010 Available online 22 July 2010 Keywords: DDR ZSM-58 MTO ZSM-5 NH3-TPD

a b s t r a c t ZSM-58, having a DDR topology, is shown to be a very attractive catalyst for the direct formation of propylene and ethylene via conversion of methanol. A performance similar to the state of the art SAPO34 catalysts is achieved, while no olefins longer than C4 are formed. In addition, ZSM-58 has a much higher thermostability than SAPO catalysts. Mainly propylene, ethylene and linear butenes (trans-but-2-ene and butadiene) are formed when materials with the DDR topology are used as catalysts during the MTO process. The ratio propylene/ethylene can be tuned by changing the reaction conditions or the degree of catalyst coking. An optimum in performance, in terms of stability and selectivity, is found for catalysts containing one acid site (one Al) per accessible cavity. Deactivation of the catalysts takes place due to formation of coke and homogeneous blocking of the catalysts porosity. Activity is fully recovered after regeneration in air. © 2010 Elsevier B.V. All rights reserved.

1. Introduction Light olefins, namely propylene and ethylene, are among the most demanded raw materials in the chemical industry. Their overall demand, especially that of propylene is forecasted to keep rising during the next decades [1]. Traditionally, they have been produced at very large scale in steam cracking, a non-catalytic process. These processes are robust but the selectivity is modest. Therefore, it is not surprising that catalytic processes to supply these chemicals are of high interest for both industry and academia. FCC (fluid catalytic cracking) processes have been adapted by modifying the catalysts, increasing the production of lower olefins, in particular propylene. Several other approaches appear to be promising. Catalytic dehydrogenation of lower paraffins [2], direct or oxidative, modifications in the catalytic cracking and the direct conversion of methanol to olefins have been proposed as possible processes. The latter, the so-called methanol to olefins (MTO) process has attracted the attention of researchers during the last four decades [3]. Methanol can be made either from syngas (from natural gas or coal) or even via fermentation. Subsequently it can be catalytically processed to gasoline (methanol-to-gasoline, MTG) or to olefins (MTO). The MTG process was initially designed as an alternative to Fischer–Tropsch-synthesis, but with the development of new small pore zeolites, more specifically silico-alumino-phosphates

∗ Corresponding author. Tel.: +31 015 2789820. E-mail address: [email protected] (J. Gascon). 0926-860X/$ – see front matter © 2010 Elsevier B.V. All rights reserved. doi:10.1016/j.apcata.2010.07.023

like SAPO-34, it was discovered that ethylene and propylene can be selectively produced. Most of the open MTO literature deals with the use of either ZSM-5 or SAPO-34 catalysts, while only recently the use of other zeolites like ZSM-11, ZSM-22, SAPO-5, SAPO-11, SAPO-18 and SAPO-35 has been reported [4–6]. It is claimed that the mild acidity of SAPO-34 together with its CHA topology are responsible for the high selectivity of the material towards the formation of ethylene and propylene [3]. This is probably the reason why mainly silico-alumino-phosphates have been explored. However, not much is known about materials with similar topologies containing only Si and Al. From the application viewpoint, it would be desirable to produce a catalyst with a higher thermal stability, since during the MTO process regeneration of the catalysts at high-temperatures in the presence of air is an integral part of the operation and silico-alumino-phosphates have a limited thermal stability compared to silico-aluminate systems [7]. A possible candidate is ZSM-58, a silico-aluminate with low aluminium content, having a DDR topology firstly synthesized by Exxon Mobil [8,9]. The structure of DDR consists of windowconnected cages, in contrast to zeolite frameworks, which makes this material a member of the group of clathrasils. The crystal structure is built by corner-sharing SiO4 tetrahedra that are connected to pseudohexagonal layers of face-sharing pentagonal dodecahedra (512 cages). These layers are stacked in an ABCABC sequence and are interconnected by additional SiO4 tetrahedra that form six-membered rings between the layers. Thus, two new types of

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cages arise, a small decahedron, 43 56 61 cage, and a large 19-hedron, 43 512 61 83 cage. By connecting the 19-hedra cavities through a ˚ a two-dimensional single 8-ring with an aperture of 4.4 A˚ × 3.6 A, porous system is formed, which is accessible to small molecules after thermal treatment at 773 K [10,11]. The selective adsorption of propylene over propane and trans2-butene and butadiene over other butane and butene isomers together with the high uptakes of ethylene on the all-silica DD3R were firstly reported by our group by studying single component adsorption using a Tapered Element Oscillating Microbalance (TEOM) [12–14]. Olson et al. [15] studied the sorption properties of some 8-ring zeolites, among which ZSM-58, in which the high selectivity for propylene was also suggested from single component adsorption. Later we reported the first propane–propylene mixture separation using DDR in a packed bed adsorption operation [16]. Up to now, little is known about the catalytic properties of ZSM58 in the open literature, to the best of our knowledge only its use in the catalytic cracking of n-octane in its acidic form [17] and the application of the Mn exchanged form in the catalytic oxidation of n-hexane [18] have been reported. In this work we present the application of ZSM-58 for the direct synthesis of olefins via methanol conversion. The performance of this DDR-structure type catalyst is comparable in terms of activity to the best SAPO-34 catalysts, it shows high selectivities to propylene and ethylene, while no product longer than linear butenes is formed. Moreover, the ratio propylene/ethylene can be tuned either by changing the reaction temperature or by controlling the coke level in the catalyst.

2. Experimental 2.1. General All chemicals were obtained from Sigma–Aldrich and were used without further purification. Scanning Electron Microscopy (SEM) was measured in a JEOL JSM 6500F setup coupled to an Energy Dispersive Spectrometer (EDS) for micro-analysis. Nitrogen adsorption at 77 K in a Quantachrome Autosorb-6B unit gas adsorption analyzer was used to determine the textural properties as BET surface area between 0.05 and 0.15 relative pressures and pore volume at 0.95 relative pressure. The t-plot was used to calculate the external surface area of the catalyst particles (calculated as the surface area of pores larger than micropores). The crystalline structures were analyzed by X-ray diffraction (XRD) using a Bruker-AXS D5005 with CuK␣ radiation. Thermogravimetric analysis (TGA) was performed by means of a Mettler Toledo TGA/SDTA851e, on samples of 10 mg under flowing 60 ml/min of air at a heating rate of 10 K/min up to 873 K. Elemental analysis was carried out by means of inductively coupled plasma optical emission spectroscopy (ICP-OES). The samples were digested in duplo in an aqueous mixture of 1% HF and 1.25% H2 SO4 and analyzed with an ICP-OES Perkin Elmer Optima 3000dv in order to determine the amount of Si and Al present in the structure. DRIFT spectra were recorded in a Bruker model IFS66 spectrometer, equipped with a high-temperature cell with CaF2 windows. The spectra were recorded after accumulation of 128 scans and a resolution of 4 cm−1 . A flow of helium at 10 ml/min was maintained during the measurements. Before collecting the spectra the different samples were pretreated in a helium flow at 393 K for 30 min. KBr was used as background. The acid site concentration was determined using temperatureprogrammed desorption of ammonia (NH3 -TPD). NH3 -TPD was carried out on a Micromeritics TPR/TPD 2900 apparatus equipped with a thermal conductivity detector (TCD). Approximately 25 mg

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of each sample was pretreated at 923 K and rapidly cooled to 473 K followed by loading with ammonia applying a flow of 30 mL (STP)/min for about 10 min. A helium flow of 50 mL/min was applied for 10 min to remove weakly adsorbed NH3 . The adsorption process was repeated 3 times. After the loading procedure, a linear temperature program was started (473–973 K at 10 K/min) in a flow of He (50 ml/min), and the ammonia desorption was followed by the TCD. 2.2. Synthesis of the catalysts For the synthesis of ZSM-58 [19], first the structure directing agent (SDA), methyltropinium iodide (MTI), was prepared by adding 25.1 g methyl iodide (99 wt.%) dropwise to a solution of 25.0 g tropine (98 wt.%) in 100 g ethanol at 273 K under stirring and keeping the suspension under reflux for 72 h. After cooling and filtration, the resulting crystalline MTI was washed with 100 g ethanol and dried at 353 K. The molar ratio of the synthesis mixture was MTI:SiO2 :Al2 O3 :Na2 O:H2 O = 17.5:70:x:11.5:2800. It was prepared as follows: 10 g of Ludox HS-40 was added to a solution of 4.7 g MTI and 8.7 g demineralised H2 O and stand overnight while stirring (solution A). 0.78 g sodium hydroxide and the corresponding amount of sodium aluminate (Al2 O3 : 50–56, Na2 O: 40–45 wt%) were dissolved in 33.3 g H2 O (solution B). The resulting solution (A + B) is stirred for 30 min, placed in a teflon lined autoclave and heated under hydrothermal conditions (autogenous pressure) to 433 K for 3–7 days depending on the Al content (the higher the Al content the longer the synthesis time, i.e. 4 days synthesis for Si/Al = 392 and 7 days for Si/Al = 80). The as synthesized materials were calcined (2 K/min) at 973 K for 10 h in a static oven. A commercial ZSM-5 from Zeolyst International (CBV28014, SiO2 /Al2 O3 = 560) was used as benchmark catalyst. 2.3. Experimental setup Experiments were carried out with a 1/4 in. stainless steel reactor with a length of 6 cm. 100 mg of catalyst particles (pelleted zeolite, crushed and sieved to a size of 0.75–1 mm) was placed in the reactor between quartz wool plugs for packing purpose and gas distribution. An ISCO liquid pump was used to feed the liquid to the reactor system. In a typical experiment, the liquid feed stream (methanol or methanol + water) was pumped to the reaction section located in a heated chamber (373 K) to evaporate the liquid while diluting it with N2 . In this chamber all valves and reactor oven were located to avoid condensation. The temperature of the reactor was varied between 573 and 723 K in the different experiments. The product mixture was analyzed online with an Interscience Compact GC over a 60 m RTX® -1 (1% diphenyl-, 99% dimethylpolysiloxane) column. Conversion (%) was calculated as follows: X = 100 ×

NMeOHin − NMeOHout NMeOHin

The overall olefin selectivity (%) was based on the carbon atoms in the products and calculated as follows: S = 100 ×

2NC2 H4 + 3NC3 H6 + 4NC4 H8 NMeOHin − NMeOHout

With Ni being the number of moles. Similarly that for the individual olefins is defined. The yield (%) of olefin i is then simply the product of conversion and its selectivity:

Yi =

Si X 100

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Fig. 1. SEM micrographs (a–d) and XRD patterns (e) of the different ZSM-58 catalysts: SiO2 /Al2 O3 = (a) 64, (b) 110, (c) 215 and (d) 392.

3. Experimental results 3.1. Characterization SEM micrographs of the ZSM-58 catalysts under study are shown in Fig. 1 together with their XRD patterns. In Table 1 the

textural properties before and after 6 reaction hours are summarized. Despite minor differences, every catalyst shows a well developed crystalline DDR pattern, while the particle size increases with decreasing the Al content, ranging from 0.5 to 7 ␮m. From the ICP/MS analyses it is calculated that 2, 1, 0.6 and 0.3 atoms of Al are present per accessible cavity in the catalyst for the samples with a SiO2 /Al2 O3 ratio of 64, 110, 215 and 393, respectively.

Table 1 Textural properties of the different ZSM-58 catalysts. SiO2 /Al2 O3

Fresh catalyst 2

64 110 215 392

Spent catalyst 2

2

SBET (m /g)

Sexternal (m /g)

Smicro (m /g)

307 332 357 396

35 28 19 18

271 305 337 378

SBET (m2 /g) 23 12 6 6

Sexternal (m2 /g)

Smicro (m2 /g)

Coke (wt.%)

21 14 8 9

2 – – –

12.1 9.8 6.5 6.6

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Fig. 2. Ammonia TPD analyses of the different ZSM-58 and ZSM-5 catalysts used, indicated by their SiO2 /Al2 O3 ratio.

Regarding specific surface area, a clear trend is observed: the lower the amount of aluminium the higher the BET area; the observed external surface area correlates well with the mean particle size. In Fig. 2, the acidic properties of the synthesized catalysts are compared with that of the studied commercial ZSM-5 catalyst. Three desorption steps are observed. A major one with a maximum shifting from 650 to 700 K and increasing with increasing Al content, a desorption between 800 and 900 K similar for all samples, and a desorption around 550 K increasing with increasing Al content. The acidity, i.e. the amount of ammonia desorbing, correlates with the amount of aluminium in the framework for the ZSM-58 samples, while a lower acidity is obtained for the ZSM-5 sample: apart from a much higher uptake, the temperature of maximum ammonia desorption is also higher for the ZSM-58 catalysts. Both facts point out at a stronger intrinsic acidity of the DDR framework. 3.2. MTO on ZSM-58 Fig. 3 shows the reaction results for the different catalysts at 673 K feeding methanol using N2 as diluting gas. Experiments were

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performed at least 2 times with an intermediate regeneration step in air at 723 K. Activity was completely recovered in every case. For the sake of clarity and comparison between different operational conditions, results are presented as a function of the amount of methanol fed to the reactor per gram of catalyst (‘throughput’). Main products observed were propylene, ethylene and linear butenes/butadienes (mostly trans-but-2-ene and buta-1,3-diene). Formation of coke was observed and quantified by TGA at the end of every experiment. Selectivity is given as overall selectivity to C2, C3 and C4 olefins, the balance is assumed to be ‘coke’. In every case during the first reaction minutes, up to approximately 0.75 g MeOH/g cat no product breakthrough can be observed. The length of this period does neither follow any trend with the aluminium content, nor with the specific surface area of the samples. While the catalysts with the lowest acidity (SiO2 /Al2 O3 = 215 and 393) deactivated very fast, the other two (SiO2 /Al2 O3 = 64 and 110) are in terms of conversion and selectivity to olefins stable for longer periods: in both cases after 5 g MeOH fed per gram of catalyst, strong deactivation occurs and conversion drops from 100% to approximately 40%. With respect to the product distribution, at the beginning of the reaction more propylene than ethylene is formed and during time-on-stream this ratio changes and ethylene becomes the more important product. In every case the yield of butenes was lower than 15% on a molar carbon basis and decreased with time-onstream. On the other hand, when conversion dropped below 100%, dimethyl ether (DME) started to be observed at the exit of the reactor, becoming the most important product at the end of the experiments. In view of these screening results, the sample with a SiO2 /Al2 O3 = 110 was selected as the optimal one in terms of stability (resistance to deactivation) and selectivity to olefins, therefore further experiments were performed over this catalyst. Fig. 4 shows the product distribution with time for the optimal ZSM-58 catalyst at four different temperatures. Temperature plays a major role, not only on the activity and stability of the catalyst, but also on the distribution of products. While at low temperatures propylene is the preferred product (reaching yields above 40%), at

Fig. 3. MTO reaction results for the different ZSM-58 catalysts: SiO2 /Al2 O3 = 64, 110, 215 and 392 (from left to right and from top to bottom). Mcat = 100 mg, FMeOH = 0.25 ml/h (liquid) (WSHV=2 h−1 ), MeOH:N2 = 1:4, T = 673 K.

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Fig. 4. C2 –C4 olefin distribution during the MTO reaction over a ZSM-58 catalyst: SiO2 /Al2 O3 = 110, Mcat = 100 mg, FMeOH = 0.25 ml/h (liquid) (WSHV=2 h−1 ), MeOH:N2 = 1:4. (a) T = 573 K, (b) 623 K, (c) 673 K and (d) 723 K.

Fig. 5. Effect of water addition on the MTO reaction results of a ZSM-58 catalyst: SiO2 /Al2 O3 = 110, T = 673 K. Mcat = 100 mg, FMeOH = 0.15 g/h (WSHV=2 h−1 ), (a) MeOH:H2 O:N2 = 1:5:6 and (b) MeOH:N2 = 1:11.

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the highest temperature ethylene is preferentially formed. In general, the higher the amount of propylene formed, the higher the amount of butenes generated and the lower the amount of ethylene. With respect to stability, at the lowest and highest temperatures (573 and 723 K) the catalyst deactivates faster, while at intermediate temperatures, and especially at 673 K, the catalyst is much more stable. It is generally observed that the addition of water enhances the catalyst stability during the MTO reaction and that the product distribution may change upon varying the partial pressure of MeOH [3]. The effects of these two variables are presented in Fig. 5, showing the results of experiments at lower MeOH partial pressure and in the presence or absence of H2 O. Nitrogen was used as diluent, keeping the partial pressure of MeOH constant. The addition of water, in the case of ZSM-58, increases the selectivity to lower olefins, approaching 100% during the stage of high methanol conversions. The percentual formation of the longer butenes is suppressed, while hardly any DME is formed. Further, the stability

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is slightly improved, though in both cases conversion drops after more than 4 g of methanol have been fed per gram of catalyst. Lower MeOH partial pressures have a negative effect on the catalyst performance when comparing the results presented in Fig. 5b with those at the same temperature in Figs. 3 and 4: the stability is slightly lower in terms of methanol throughput, while the yield to propylene decreases and the overall selectivity to olefins is also lower, so more coke is formed. 3.3. ZSM-5 as benchmark catalyst In order to benchmark the catalytic properties of the studied ZSM-58 catalysts, experiments with a commercial high silica ZSM5 catalyst were performed. Differences in terms of topology are expected to play an important role in the product distribution and in the stability of the catalyst. In Fig. 6 the experimental results for the ZSM-5 catalyst at different MeOH partial pressures and in the presence and absence of water are presented, under simi-

Fig. 6. MTO reaction results on a high Si ZSM-5 catalysts: SiO2 /Al2 O3 = 140, Mcat = 100 mg, T = 673 K. (a) FMeOH = 0.15 g/h (WSHV=2 h−1 ), MeOH:N2 = 1:4; (b) FMeOH = 0.15 g/h (WSHV=2 h−1 ), MeOH:H2 O:N2 = 1:5:6 and (c) FMeOH = 0.15 g/h (WSHV=2 h−1 ), MeOH:N2 = 1:11.

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Fig. 7. Product composition during conversion of ethylene (left) and propylene (right) with time-on-stream on a fresh ZSM-58 catalyst. SiO2 /Al2 O3 = 110, T = 673 K. Mcat = 100 mg, Ftotal = 15 ml (STP)/min; 8% olefin in feed. Yield is defined per carbon atom of reactant.

lar conditions as used for the ZSM-58 catalysts. The results when using ZSM-5 differ quite from the ones obtained for ZSM-58. Main products in order of decreasing amount are propylene, butenes and ethylene. The activity of ZSM-5 is fairly stable, while the total selectivity to olefins is constant: after feeding more than 10 g of MeOH per gram of catalyst, the activity is still high and the distribution of products remains nearly invariant. The addition of water seems to improve the stability but has little effect on the product distribution; the selectivity to olefins is similar, in contrast to the 20 points higher selectivity observed when adding water to the ZSM-58 system.

3.4. Reactivity of reaction products From the previous experiments, large differences are observed upon comparing the performance of the ZSM-58 and ZSM-5 catalysts. These differences might be due to the different topologies of the samples, but also to the strength of the active sites present in each catalyst. Ammonia TPD characterization indicates that the acidity of the ZSM-5 catalyst is the lowest one, as it could be expected from its high Si/Al ratio. However, while in the case of the ZSM-58 catalysts such low acidities resulted in a very poor reaction performance and a fast deactivation, in the case of ZSM-5 this mild acidity is enough for the conversion of methanol.

It is well known that for consecutive reactions, valuable information can be extracted by feeding intermediate reaction products as starting reactants [20]. In this case, we performed several experiments with both catalysts feeding ethylene and propylene. Surprisingly, while the ZSM-58 catalysts presented a high activity, the ZSM-5 sample was hardly active. Results for the SiO2 /Al2 O3 = 110 ZSM-58 catalyst are presented in Fig. 7. For the sake of clarity yields are given on a carbon atom basis (i.e.: for the experiments feeding ethylene, the yield to propylene is calculated as: YC3 H6 = X3NC3 H6 /(2(NC2 H4 − NC2 H4 )), where X is the conversion of ethylene). The evolution of the reactants is plotted in the graph as (100 − X). At the initial stages of the reaction with ethylene, conversions higher than 60% are obtained and large amounts of propylene and butenes are observed at the reactor exit. Also a large selectivity to coke is observed in the first reaction minutes. Although propylene turned out to be less reactive, initial conversions near 40% are achieved with formation of coke, butenes and ethylene. 3.5. Formation of coke Fig. 8 shows the evolution of the coke formation and the decrease in specific surface area for the SiO2 /Al2 O3 = 110 ZSM-58 catalyst sample upon reaction with MeOH, ethylene or propylene at different reaction times. The amount of coke formed (wt.%) does not correspond with the decrease in specific surface area and fol-

Fig. 8. Evolution of the specific surface area (left) and the coke amount on the catalysts (estimated by TGA) (right) with time-on-stream for a ZSM-58 catalyst (SiO2 /Al2 O3 = 110) for different reaction feeds. T = 673 K, Mcat = 100 mg. Ftotal = 15 ml (STP)/min.

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Fig. 9. DRIFT spectra of a spent ZSM-58 catalyst after 6 reaction hours under different reactant conditions.

lows a clear trend with the partial pressure of the hydrocarbon used as reactant. In the case of methanol, already after 180 min (corresponding with a throughput of 5 g MeOH/g catalyst) the total surface area of the catalyst is lost, though the amount of coke on a weight basis still increases further with time-on-stream. In the case of the olefins, still some surface area of the catalyst is available when no further reaction is observed. In order to characterize the deposited coke, DRIFT spectroscopic analysis of the spent catalysts after different reaction conditions was carried out. Fig. 9 summarizes the most important results. Absorbance bands at 2961, 2930 and 2869 cm−1 common for the C–H stretching on the ␣ position of alkylbenzenes [21] can be assigned to the aromatic compounds formed in the zeolite cavities. The low intensity of the absorption above 3000 cm−1 due to stretching vibration of aromatic C–H bond, suggest that the aromatic species are heavily methylated [22]. A band at around 2960 cm−1 was previously observed in MOR, FAU and BEA zeolites, and was assigned to the propylated benzenes [21]; large pores of 12MR zeolites allow for the formation of larger alkyl substituents. In the case of ZSM-58 the accessible cavities of this material would also be able to host large alkyl substituted benzenes. A strong absorption around 1600 cm−1 (so-called “coke” band) arises from a combination of ı(C–H) modes of a mixture of carbonaceous hydrogen-deficient deposits [23–25] and aromatic (C–C) modes [26]. The intensity of the latter band has been used as a measure of the amount of coke deposited in zeolite channels [23]. Finally, the adsorption at 1567 cm−1 can be ascribed to alkylnaphtalenes or polyphenylene structures [23,27,28]. No IR bands around 1500 cm−1 are observed which are characteristic of highly condensed aromatic deposits [26]. The IR band intensities of coke in the case of methanol as reactant are noticeably higher than in the other cases. 4. Discussion As inferred from the similar topology and the 8MR window cage size of ZSM-58 as SAPO-34 this catalyst has a good performance in the MTO reaction. Furthermore, mainly propylene, ethylene and linear butenes (trans-but-2-ene and butadiene) are formed, in agreement with our adsorption results on DD3R [12–14,16,29,30], where only these components entered this structure. In every

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MTO experiment products breakthrough did not take place until a time corresponding with a throughput of approximately 0.5–1 g methanol/g catalyst. This is attributed to the time needed to vaporize the liquid fed together with adsorption of methanol and the first reaction products. Similar phenomena have been reported for SAPO-34 catalysts [20]. The fact that in the experiments performed with ethylene and propylene this delay is not observed is related to the negligible effect of delay time in the case of gaseous reactants and to the much lower adsorption coverage of these olefins, more than 5 times smaller at the studied temperatures in the case of SAPO-34 [13,20]. From Fig. 3 it is clear that the Si/Al ratio plays a key role in the stability of the catalyst: an optimum is found for one aluminium atom per accessible cage: lower amounts of aluminium result in a prompt deactivation of the catalysts, while higher amounts do not improve further the performance: even considering the smaller particle size of the SiO2 /Al2 O3 = 64, the stability of the catalyst is not better and a larger amount of ethylene is produced compared to propane. This fact can be related with a higher cracking activity when more than one acid site is present per cavity of the material, leading to a faster production of coke (see Table 1) and therefore to a faster deactivation. When comparing the effect of the Si/Al ratio with other catalysts with similar topologies [6], this effect is much stronger in the case of ZSM-58. NH3 -TPD confirms the strong acid character of the sites in ZSM-58, as compared to those of ZSM-5. When comparing the NH3 -TPD of the different ZSM-58 samples, it can be observed that the third desorption peak, centered at 850 K, hardly changes with increasing the amount of Al in the framework, while the desorption peak centered at 675 K increases linearly with the amount of aluminium. The NH3 -TPD results point at the presence of, at least two different acid sites: the strongest ones (peak at 850 K) would suffer a strong and quick deactivation, while the softer ones (peak at 675 K) would be responsible for the MTO activity. This fact would also account for the swift deactivation of the low Al content catalysts. IR spectroscopy indicates the elevated branching degree of the aromatics deposited in the catalysts, pointing out the high acidity of this material. During the period of full conversion a change in the product distribution is observed. In general production of propylene and butenes decreases with time, while the production of ethylene increases. Considering that DDR is a shape selective catalysts and sorbent, this trend is attributed to the decreasing available space in the cavities as the reaction proceeds, allowing only the production of smaller olefins by hampering the longer ones to escape from the catalyst pores [31]. This fact can be used to tune the final product distribution simply by controlling the total amount of coke in the catalyst during operation. The integral behaviour of the plug flow reactor used in this study is not easy to decouple from the reaction results: while in the initial part of the bed the methanol is converted into olefins, these may be converted further downstream in the reactor: the longer the time-on-stream the smaller the region of olefin conversion. After deactivation of the catalysts DME becomes the most important reaction product, but at this point most of the catalyst porosity is occupied by coke (see Fig. 8). This suggests that this DME is mainly formed at the external surface of the particles on Lewis sites that have not been deactivated by the deposition of coke. The total amount of coke observed correlates better with the particle size than with the Si/Al ratio, similarly as reported for SAPO-34 [32]: the larger the particle size the higher the probability of closing the available porosity. This is actually a very important factor that may lead to incorrect conclusions when studying large zeolite crystals: recently Mores et al. introduced two different coke formation mechanisms for large ZSM-5 and SAPO-34 crystals (larger than 20 ␮m) and suggested that SAPO-34 is deactivated due to formation of coke deposits at the outer regions of the crystals. Though this might be true for huge crystals, considering the large

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amounts of coke found after reaction in smaller crystals, (in general more than 10 wt.%, see Table 1 and Refs. [32,33]) it is clear that when dealing with micrometer sized crystallites most of the catalyst porosity is occupied by coke deposits, especially in the case of the high Al content catalysts (smallest particles). In the case of the ZSM-58 catalysts, the accessible [43 512 61 83 ] cavity has a free volume of about 0.35 nm3 and a free cross sectional diameter of 0.875 nm by assuming a sphere-like cage. There are six of these cavities present per unit cell, and considering that every large cavity can be filled with two stacked benzene molecules at maximum (or by one heavily branched benzene molecule), this value would correspond to a 13 wt.% of coke in the spent catalysts, almost the value calculated by TGA analysis of the spent catalysts containing one acid site per large cavity. This clearly demonstrates that full blocking of the pores takes place. The addition of water, even at low concentrations, improves to a large extent the performance of the ZSM-58 catalyst. Though the period of stable operation does not change, almost 100% selectivity is found to propylene and ethylene. This fact could be related with a faster desorption of the olefins in the presence of water, that might compete for the same (acid) adsorption sites, avoiding further consecutive reaction of these products. Methanol conversion already produces water, so this influence is already present during reaction. At lower methanol pressure this intrinsic water influence may be less effective, explaining the slightly faster deactivation of the catalyst. The presence of added water suppresses the DME formation over the deactivated catalysts due to thermodynamic limitations. Upon comparing the results obtained for ZSM-58 with the ones obtained for ZSM-5, it is clear that the behaviour of both catalysts is very different and that physical constraints in terms of topology dominate product distribution and catalyst stability. First important difference is that the needed acidity as estimated from ammonia TPD in the case of ZSM-5 is far less than for ZSM-58. The lower reactivity of propylene and ethylene on the MFI together with the lower acidity demonstrate that the oligomerization and coke formation activity of this catalyst are also lower, accounting for the higher stability of the ZSM-5 catalyst. With respect to the distribution of products similar amounts of propylene are formed in both catalysts, but far less ethylene is produced over MFI, while larger amounts of longer chain olefins are formed, illustrating the shape selective character of ZSM-58. The addition of water in the case of MFI clearly has a different effect than in the case of ZSM-58. The affinity of DDR materials for short chain olefins is stronger than that of ZSM-5 [34,35], so the coverage of propylene and especially of ethylene [13] on the ZSM-58 catalysts is much more lowered by the competition with water. Hence, the consecutive reaction of these olefins over the strong acid sites is suppressed. The comparison in terms of performance of the best ZSM58 catalysts with results published in literature for other 8-membered ring (8MR) alumino-silicates and silico-aluminophosphates [3–6,20,31,33,36] yields an interesting observation: when using ZSM-58 the amount of butenes formed is lower than reported for SAPO-34, state of the art 8MR MTO-catalyst, while no hydrocarbons longer than C4 are detected in the case of DDR in contrast to catalysts with the CHA topology [31], and their stabilities (in terms of deactivation) are similar under comparable reaction conditions. Considering that the main target of the MTO process is the selective formation of propylene, followed by ethylene, the lower selectivity to butenes is an advantage in the case of ZSM-58. Moreover, by varying the temperatures the yield of propylene may be double that to ethylene (see Fig. 4b). When comparing both zeotypes several differences are found in terms of topology that we believe are responsible for the different selectivities: DDR is built up of cavities communicating through ˚ small ellipsoidal windows with a pore opening of 3.6 A˚ × 4.4 A,

resulting in a 2D pore system, in the case of CHA the cavities communicate through 3.8 A˚ × 3.8 A˚ windows in a 3D pore structure. This difference in topology has an impact on diffusion. Olson et al. [15] reported much faster diffusion of both propane and propylene in CHA than in DDR, demonstrating that even a small difference of 0.2 A˚ together with a different porous system (2D vs. 3D) makes a large difference in the behaviour of adsorbates with kinetic diameters similar to the size of the pores. This fact obviously accounts for the different product distributions in ZSM-58 and CHA. Summarizing, ZSM-58 is shown to be a very attractive catalyst for the selective formation of olefins from methanol, with a selective formation of propylene and ethylene and the possibility of tuning this ratio. The stability to deactivation, similar in terms of throughput to the one reported for SAPO-34 catalysts [31], should be overcome by a proper reactor design. In this sense FCC- or twozone fluidized bed reactors [37,38] would be the most appropriate ones to maintain the catalyst at a certain level of deactivation (coke content), enabling even to tune the ratio of olefins on demand. In comparison with other possible candidates for this kind of reactor technology, ZSM-58 presents as important additional advantage the higher thermal stability of alumino-silicates (higher than 973 K for ZSM-58) in comparison with silico-alumino-phosphates (not higher than 773 K). 5. Conclusions ZSM-58 is a very attractive catalyst for the direct formation of propylene and ethylene via conversion of methanol. Mainly propylene, ethylene and linear butenes (trans-but-2-ene and butadiene) are formed when materials with the DDR topology are used as catalysts for the MTO process. Moreover, the ratio propylene/ethylene can be tuned by changing the reaction conditions or controlling the coke level if an appropriate reactor is used. An optimum in performance, in terms of stability and selectivity is found for catalysts containing one acid site per accessible cavity. Water increases the olefin selectivity due to competitive adsorption for the strongly acid sites and suppressing the consecutive reaction of the olefins. Deactivation of the catalysts takes place due to formation of coke and homogeneous blocking of the catalysts porosity. Activity is fully recovered after coke combustion. The similar performance of ZSM-58 with respect to the state of the art SAPO-34 catalysts together with the higher thermal stability of pure alumino-silicates may form the basis for the development of more robust catalysts for the MTO process. Acknowledgments KAO company is gratefully acknowledged for financial support. J.G. gratefully acknowledges the Netherlands National Science Foundation (NWO) for his personal VENI grant. The X-ray facilities of the Department of Materials Science and Engineering of the Delft University of Technology is acknowledged for the XRD analyses References [1] J.S. Plotkin, Catal. Today 106 (2005) 10–14. [2] K.J. Caspary, H. Gehrke, M. Heinritz-Adrian, M. Schwefer, in: G. Erlt, H. Knözinger, F. Schüth, J. Weitkamp (Eds.), Handbook of Heterogeneous Catalysis, Wiley–VCH Verlag GmbH & Co., Weinheim, 2008, pp. 3207–3227. [3] M. Stöcker, Micropor. Mesopor. Mater. 29 (1999) 3–48. [4] Z.M. Cui, Q. Liu, W.G. Song, L.J. Wan, Angew. Chem., Int. Ed. 45 (2006) 6512–6515. [5] A.T. Aguayo, A.G. Gayubo, R. Vivanco, M. Olazar, J. Bilbao, Appl. Catal. A: Gen. 283 (2005) 197–207. [6] Z. Zhu, M. Hartmann, L. Kevan, Chem. Mater. 12 (2000) 2781–2787. [7] B.V. Vora, T.L. Marker, P.T. Barger, H.R. Nilsen, S. Kvisle, T. Fuglerud, Nat. Gas Convers. IV 107 (1997) 87–98. [8] P.G. Rodweald, E.W. Valyocisk, Conversion of Oxygenate(s) to Hydrocarbon(s) using ZSM-58 Zeolite catalyst, Mobil Oil Corp., 1987.

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