Role of Shape Selectivity and Catalyst Acidity in the Transformation of

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Industrial & Engineering Chemistry Research

1

Role of Shape Selectivity and Catalyst Acidity in the Transformation of Chloromethane into Light Olefins

Monica Gamero*, Andres T. Aguayo, Ainara Ateka, Paula Pérez-Uriarte, Ana G. Gayubo, and Javier Bilbao

Department of Chemical Engineering, University of the Basque Country (UPV/EHU). P.O. Box 644, 48080. Bilbao (Spain). *[email protected]

ABSTRACT The transformation of chloromethane into light olefins (C2-C4) has been studied both on HZSM-5 catalysts with different SiO2/Al2O3 ratios (30, 80 and 280) and on SAPO-n catalysts (SAPO-34 and SAPO-18) in order to analyze the role of shape selectivity and acidity in the kinetic behavior. The catalysts have been prepared by agglomerating these acid functions with bentonite and -Al2O3, and the kinetic runs have been performed in a fixed bed reactor under the following operating conditions: 350 and 450 ºC; space time, 2.35, 5.89 and 14.99 gcat h (molCH2)-1; and time on stream, 255 min. A comparison of the reaction indices (conversion of chloromethane, selectivity to light olefins and propylene fraction) at zero time and throughout time on stream using the different catalysts has allowed establishing, on the one hand, the significance of the shape selectivity and acidity of these catalysts (which are more influential in this reaction than in the transformation of methanol), and on the other, the need for a compromise between these properties. A HZSM-5 zeolite catalyst with moderate acidity (SiO2/Al2O3 = 80) has a good kinetic behavior at 350 °C, and recovers its activity by

ACS Paragon Plus Environment


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2 coke combustion with air. However, all the catalysts studied undergo irreversible deactivation at 450 °C by dealumination of the acid function to form AlCl3.

Keywords: methane, chloromethane, olefins, HZSM-5 zeolite, SAPO-n, acidity, shape selectivity.

Highlights -

Shape selectivity and acidity are essential for stability and olefin selectivity

-

The significance of these factors is greater than in the transformation of methanol

-

The acidity and severity of shape selectivity favor deactivation

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A SiO2/Al2O3 ratio of 80 for the HZSM-5 zeolite is suitable

-

Deactivation occurs by coke deposition at 350 ºC and also by dealumination at 450 ºC


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3

Graphical Abstract

45

SAPO-n

HZSM-5 zeolite

M Methane O C2-C4 olefins C5+ HC5+ P C2-C4 paraffins BTX Aromatics (BTX)

40 35 30

Chloromethane

Yield, %

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25 20 15 10 5 0 SiO2/AlCZ-30 2O3 = 30

80 CZ-80 280

CZ-280

HZSM-5 zeolite


CS-18

CS-34

SAPO-34

SAPO-18


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4 1. Introduction The increase in methane reserves (mostly in the form of shale gas) is having a major impact on the energy and commodity sectors, and also improving the prospects for the implementation of methane valorization processes for the production of automotive fuels and raw materials for the petrochemical industry.1 One of the routes for methane valorization involves a two-stage process consisting of chloromethane production by hydrochlorination and the subsequent transformation of chloromethane into light olefins. This route is of interest for meeting the increasing demand for olefins,2,3 given that this integrated process is an attractive alternative to the current route of methane conversion into syngas and the subsequent synthesis of methanol or DME, which are then converted into light olefins by means of the MTO process,4,5 or the DTO process.6 The integrated methane-chloromethanelight olefin process has a potentially greater energy efficiency and the HCl required in the hydrochlorination is recovered in the transformation of chloromethane. The reaction scheme for the transformation of chloromethane into hydrocarbons on acid catalysts is similar to that for the MTO and DTO processes, with light olefins as primary products, which are transformed into paraffins and higher hydrocarbons by secondary oligomerization-cracking reactions and other isomerization, alkylation, hydrogen transfer and coke formation reactions.7-9 Consequently, the studies on the transformation of chloromethane into light olefins have been related to advances in the catalysts preparation and the greater understanding of the reaction mechanism of MTO and DTO processes in order to improve light olefins selectivity and reduce deactivation by coke. The most widely studied catalysts in the transformation of chloromethane are SAPO-3410-16 (used in the MTO process) and HZSM-5 zeolite (used in the DTO process), with different modifications, mainly by means of metal incorporation.17-22 Furthermore, spectroscopic studies show that the reaction mechanism on both catalysts proceeds, as in the MTO and DTO processes, through the formation of


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5 methoxy ions as active species for the formation of polymethylbenzenes as intermediate compounds. These intermediates release light olefins, being also the precursors of the coke that blocks the catalyst micropores.22,23 Although the aforementioned studies reveal many similarities in the kinetic behavior of SAPO-34 and HZSM-5 zeolite in the MTO and DTO processes, the lower activity in the transformation of chloromethane and the faster deactivation should be highlighted. The deactivation of HZSM-5 zeolite catalyst in the transformation of chloromethane has been studied in a previous work, showing similarities with the transformation of methanol according to the mechanisms involving reaction steps and coke formation and growth throughout time on stream.24 This work explores the role that the key properties of the catalysts acidity and shape selectivity play in the transformation of chloromethane into olefins, i.e., the effect these properties have on activity, selectivity to light olefins and stability. Therefore, two different families of catalysts with different severity in shape selectivity have been used. On the one hand, the HZSM-5 zeolites have a three-dimensional porous structure (MFI topology) constituted by elliptical straight channels (0.53 x 0.56 nm) and zigzag channels (0.51 x 0.55 nm), which intertwine without cavities at the intersections. In order to study the effect of acidity, three zeolites with different SiO2/Al2O3 ratios have been used. On the other hand, two SAPO-n (SAPO-34 and SAPO-18) have been studied, which have similar structures with a severe shape selectivity because the micropores in both materials are 0.38 x 0.38 nm channels.25 This different spatial distribution leads to different cages at the pore intersections, of 1.27 x 0.94 nm for SAPO-34 and 1.27 x 1.16 nm for SAPO-18.26 Furthermore, the acid sites of SAPO-18 have a lower acid strength than those of SAPO-34, thereby behaving differently in the transformation of methanol into olefins, with a slower deactivation by coke deposition for SAPO-18.27



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6 2. Experimental 2.1. Catalysts Five catalysts corresponding to different acid functions have been studied. These active functions are: three commercial HZSM-5 zeolites (supplied by Zeolyst International in the ammonium form with SiO2/Al2O3 ratios of 30, 80 and 280), and two SAPO-n (SAPO-34 and SAPO-18). The commercial zeolites have been subjected to a calcination step (550 ºC, 3 h) in order to remove the ammonia and obtain the acid form. The SAPO-n have been synthesized using H3PO4 (85 %) as the phosphorous source and Silicasol (40 wt%) as the silicon source. The SAPO-18 has been synthesized following the method by Chen et al., 28 from a source of aluminum (Al2O3 3H2O) and N,N-diisopropylethylamine (iPr2EN) as organic template. The preparation method of SAPO-34 has been the proposed by Wendelbo et al.,29 using aluminum isopropoxide (Al(OC3H7)3) as aluminum source, tetraethylammonium hydroxide (TEAOH, 20 wt%) as organic template and HCl (35 wt%) for pH control. The chemical composition of the SAPO-n synthesized, before the removal of the organic template, is as follows: (SiO2)0.60 (Al2O3) (P2O5)0.90 (C8H19N)1.60 50H2O

for

SAPO-18;

and

(SiO2)0.08 (Al2O3)2.8 (P2O5)0.98 (HCl)0.02 (C8H20NOH) 41H2O for SAPO-34. The acid functions have been agglomerated with bentonite as binder (Exaloid) (30 wt%) and alumina (Prolabo, calcined at 1000 ºC to become inert) as inert charge (45 wt%). The catalyst particles have been obtained by wet extrusion, using a high-pressure hydraulic piston, through 0.8 mm diameter holes. The extrudates are firstly dried at room temperature for 24 h, and later in an oven at 110 ºC for 24 h. Then the catalyst has been sieved to a particle size between 0.15 and 0.3 mm, suitable for been used in a fixed bed reactor. Finally the catalyst has been calcined for 3 h at 575 ºC. This temperature is reached following a ramp of 5 ºC min1

. The agglomeration of the acid functions with bentonite and alumina provides mechanical

resistance to the catalyst particle and creates a matrix of mesopores and macropores that


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7 improves the accessibility of the reactants and facilitates the location of coke, thus attenuating deactivation of the acid function by coke blockage at the micropores entrance.30,31 The zeolites have been denominated Z-30, Z-80 and Z-280, corresponding to the SiO2/Al2O3 ratios of 30, 80 and 280, respectively. The catalysts prepared by agglomeration have been denoted as CZ-30, CZ-80 and CZ-280. Similarly, SAPO-18 and SAPO-34 have been denoted S-18 and S-34 in a simplified form and, CS-18 and CS-34, respectively, the catalysts prepared by agglomerating the acid functions. The surface area and porous structure of the acid functions and the catalysts (Table 1) have been determined by N2 adsorption-desorption (Micromeritics ASAP 2000) and Hg porosimetry (Micromeritics Autopore 9220). The micropore volume of the catalysts corresponds to the acid function, whereas the volume of meso- and macropores corresponds to the matrix of the catalyst (bentonite and alumina). As observed in Table 1, the BET surface area and the porous structure of the two groups of acid functions are different. Thus, BET surface areas of HZSM-5 zeolites are lower than those corresponding to SAPO-n. In addition, certain differences are observed in the properties of the zeolites, with a smaller micropore volume for Z-280. Comparing the properties of the two SAPO-n, micropore volume and BET surface area of SAPO-18 are significantly higher. As a consequence of the agglomeration, the catalysts have different pore volume distribution than the acid functions, with presence of mesopores and macropores in the bentonite and -Al2O3 and without appreciating a significant blockage of the acid function micropores. Table 1 The total acidity (or acid site amount) and acid strength distribution of the acid functions and the catalysts (Table 2) have been determined by combining the techniques of thermogravimetric analysis and differential scanning calorimetry applied to the adsorption of NH3. In addition, studies of temperature programmed desorption (TPD) of NH3 were also conducted.



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8 Figures 1 and 2 show the results of acid strength distributions (graphs a) and curves of TPD of NH3 (graphs b), for the HZSM-5 zeolite and SAPO-n catalysts, respectively. These analyses have been carried out under the conditions described in previous works, by using a Setaram TG-DSC 111 calorimeter connected on-line (through a hot insulated line) with a Thermostar mass spectrometer from Balzers Instruments.27,32,33 The nature of the acid sites (Brönsted or Lewis) has been determined from the FTIR spectrum of the adsorbed pyridine by analyzing the 1400-1700 cm-1 region, which has been obtained using a Specac catalytic chamber connected on-line with a Nicolet 6700 FTIR spectrometer. The Brönsted/Lewis (B/L) site ratio value at 150 ºC (also shown in Table 2) has been determined for the corresponding zeolites and catalysts from the ratio between the intensity of pyridine adsorption bands at 1545 cm-1 and 1450 cm-1,34 and taking into account the molar extinction coefficients of both adsorption bands proposed by Emeis et al.35: B= 1.67 cm µmol-1 and L= 2.22 cm µmol-1. For the SAPO-n this analysis is not possible due to diffusion limitations of pyridine in the pores. Table 2 shows that total acidity, average acid strength and the Brönsted/Lewis ratio of the different zeolites notably decrease as the SiO2/Al2O3 ratio of the HZSM-5 zeolite is increased, which is a trend reported in the literature.36,37 Table 2 Figure 1 Figure 2 A comparison of the results for the acid functions and for the catalysts reveals that the agglomeration with bentonite and alumina does not provide significant acidity (0.013 mmolNH3 g-1) to the final catalyst, so the total acidity (or acid site amount) remains constant (defined by mass unit of the acid function). Consequently, the total acidity per catalyst mass unit is approximately a quarter of the corresponding zeolite. However, the agglomeration has a remarkable effect of redistribution of the acid strength, thereby removing


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9 the strongest acid sites, which are identified in the acid functions by an adsorption heat of NH3 above 140 kJ (molNH3)-1 for Z-30 zeolite and SAPO-n (not shown). In addition, in the TPD curves of the acid functions (not shown), these sites are identified by a peak at high temperature (T3 in Table 2). Once agglomerated, these strong acid sites are not identified and the acid strength distribution is more homogeneous in the catalysts than in the acid functions. In line with the mentioned elimination of the strongest acid sites, Brönsted/Lewis ratio (Table 2) also decreases when the zeolites have been agglomerated. Furthermore, the SAPO-n catalysts have a total acidity and average acid strength similar to those of the zeolite catalyst with an intermediate value of SiO2/Al2O3 ratio (Z-80 catalyst). It should be noted that when both SAPO-n catalysts are compared, the total acidity and average acid strength are significantly higher for CS-34, corresponding to SAPO-34. 2.2. Reaction equipment and product analysis The runs have been carried out in an automated reaction equipment (Figure 3) composed of the following sections: i) a reactants and inert feed system; ii) a hot box containing: a system to preheated the feed, a 6-port valve to drive the reactants and products flow, the catalytic reactor and beds for HCl adsorption, and; iii) a products analysis system, consisting of a micro-GC for the online analysis of the reaction products. Figure 3 The inputs of gases and liquids and a safety exit are shown on the left in Figure 3. The feed streams are: i) N2/He: inert gas; ii) Air: oxidizing gas for coke combustion; iii) chloromethane: reactant gas; iv) N2/He: inert gas that drags a sample from the reactor output to a chromatograph for online analysis, and; v) He: inert gas, carrier for the micro-GC. The flows of gas inputs are measured by mass flow meters (Bronkhordt High-Tech BV Series), which are complemented with a valve and a control system in the range from 3 to



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10 300 cm3 min-1. The equipment also allows the supply of a stream of liquid (such as methanol), not used in this work. The gases are preheated and mixed, reaching a T-shaped connection where, once homogenized, are sent to a 6-port valve. Depending on the position of the above mentioned valve the reactants are sent to the reactor or bypassed to the chromatograph. The isothermal fixed bed reactor (316 stainless steel with an internal diameter of 9 mm and 10 cm of effective length) is located inside a ceramic covered stainless steel cylindrical chamber, which is heated by an electric resistance and can operate up to 100 atm and 700 ºC with a catalyst mass of up to 5 g. The catalytic bed consists of a mixture of catalyst and solid inert, carborundum (CSi, with an average particle diameter of 0.105 mm) in order to ensure an isothermal bed and provide sufficient height to behave as an ideal flow under low space time conditions. The temperature is controlled by a digital TTM-125 Series controller and measured by a thermocouple (K-type) situated in the catalyst bed. There are two other temperature controllers: one for the furnace chamber and the other one for the transfer line between the reactor and the micro-GC. The pressure meter (Bronkhordt High-Tech BV Series) measures overpressures from 100 mbar to 4 bar. The pressure controller (P-600 Series) acts on a needle valve depending on the gas flow through the reactor. The operating variables are controlled by software specifically designed for this process called Adkir. At the outlet of the catalytic reactor, the reaction products pass through an adsorption equipment (two fixed beds), which retains the HCl byproduct of the reaction. This step is necessary in order to avoid damaging the product analysis equipment. The adsorption beds are composed of a mixture of CaCO3 and Laponite (5 wt%), previously calcined at 800 ºC to increase surface area. The gases leaving the HCl adsorption reactors pass through a 5 μm particle filter and subsequently are cooled to 170 °C. Reaction product samples (diluted in a


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11 He stream of 20 cm3 min-1) are continuously analyzed in a gas chromatograph (microchromatograph Varian CP-4900). The remaining stream of reaction products passes through a Peltier cell at 0 ºC. The amount of liquid condensate is controlled by a level sensor and the non-condensable gas flow is vented. The micro-chromatograph is controlled by Star Toolbar software and provided with four analytical modules with the following columns: a molecular sieve (MS-5) where H2, O2, N2, CO and methane are separated; Porapak Q (PPQ), where the lighter products are separated (CO2, methane, ethane, ethylene, propane, propylene, water, HCl, butanes and butenes); 5CB (CPSIL), where C5-C10 fraction and BTX aromatics are separated; and Alumina (Al2O3) where C2-C5 light paraffins and olefins are separated. Product identification is performed using a GC-MS (Shimadzu GC-MS QP2010), with a DB1-MS column, using the mass spectra library NIST 02. The quantification of the compounds is carried out based on calibration standards of known concentration. 3. Results Experiments have been carried out by feeding pure chloromethane at atmospheric pressure under the following operating conditions: temperature, 350 and 450 ºC; space time (defined as the ratio between the catalyst mass and the molar flowrate in the feed expressed in CH2 equivalent units), 2.35, 5.89 and 14.99 gcat h (molCH2)-1; time on stream, 255 min. The composition of the product stream has been quantified by considering the concentration of chloromethane (MCl) and the following product lumps: light olefins (ethylene, propylene and butenes) (O), light paraffins (ethane, propane, isobutane and butane) (P), C5-C10 aliphatics, which include all the olefins and paraffins with more than 5 carbon atoms (C5+), and aromatics (benzene, toluene and xylenes) (BTX). Moreover, the methane (M) formed by decomposition of chloromethane has also been considered.



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12 The kinetic behavior of the catalysts has been quantified with the following reaction indices: The conversion of chloromethane: X 

FMCl inlet  FMCl outlet 100 FMCl inlet

(1)

where chloromethane molar flow rate in the feed, (FMCl)inlet, and in the reactor outlet stream, (FMCl)outlet, have been expressed in CH2 equivalent units. The yield of each i lump: Yi 

Fi

FMCl inlet

100

(2)

where Fi is the molar flow rate of the i lump at the reactor outlet stream, expressed in CH2 equivalent units. The selectivity of each i lump: Si 

Fi 100  Fi

Fraction of propylene in the olefin lump: Fpropylene 

(3) S propylene Sethylene  S butenes

(4)

3.1. Kinetic results for the fresh catalysts Figure 4 shows the yields of the product fractions (graph a) and the selectivity to light olefins and the propylene fraction (graph b), for the different catalysts. These results correspond to 350 ºC and a space time of 5.89 gcat h (molCH2)-1. It has been taken into account that the reaction has an induction period of about 10 min under these conditions, so that the results correspond to this value of time on stream, for which the deactivation is considered still incipient and the catalyst keeps the properties of the fresh one (Tables 1 and 2). In addition, in order to facilitate a quantitative comparison of the results, the values of the most significant reaction indices, defined by eqs 1-4, have been gathered in Table 3. Figure 4a and Table 3 show that an increase in the SiO2/Al2O3 ratio of the HZSM-5 zeolite from 30 to 80 involves a slight decrease in the conversion of chloromethane and yields of C2-C4 olefins (from 47.69 % to 44.80 % and from 16.03 % to 15.23 %, respectively). The


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13 yields of C2-C4 paraffins and aromatics also decrease, whereas the yield of the non-aromatic C5-C10 fraction (around 19 %) remains practically constant. These results are explained by the decrease in the acid site amount of the catalysts (and also in its average acid strength) (Table 2), whose consequence is a lower reaction extent. Figure 5 shows the scheme of the transformation of chloromethane into hydrocarbons, which is similar to that well-established for the transformation of methanol into hydrocarbons.5 In this scheme, light olefins are the primary products, which act as reactants for the formation of higher molecular weight byproducts, by oligomerization-cracking, isomerization and alkylation reactions, in which hydrogen transfer reactions have a great impact. These hydrogen transfer reactions require acid sites with significant acid strength, which explains the notable decrease in the yield of byproducts (particularly aromatics and paraffins) with the increase in the SiO2/Al2O3 ratio of the zeolite catalyst due to the decrease in the average acid strength of the acid sites (Table 2). Figure 4 Table 3 Figure 5 As a result of the low acid site amount and average acid strength, and although the yield of olefins is low for CZ-280 catalyst, a high selectivity to olefins (54.71 %) is obtained. In addition, it should be noted that the selectivity to the non-aromatic C5-C10 fraction is high (41 %) and those corresponding to C2-C4 paraffins (2.3 %) and aromatics (1.2 %) are low. This result is very interesting for industrial operation because the non-aromatic C5-C10 fraction can be fed back, thus increasing the yield of light olefins, as has been determined in the transformation of methanol.38 Additionally, propylene is the prevailing olefin in the product stream, with a fraction composition (eq 4) of 0.63. This result is of special relevance to meet the increasing demand for propylene.



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14 The results in Figure 4 and Table 3, corresponding to SAPO-n catalysts, are very different to those reported for HZSM-5 zeolite catalysts, and evidence that conversion is lower but selectivity to light olefins (Figure 4b) higher, with C5-C10 aliphatics and light paraffins being byproducts and total absence of aromatics (BTX). The severity in the shape selectivity of the SAPO-n microporous structure plays an important role to explain the low conversion obtained compared to the corresponding HZSM-5 zeolite catalysts. This is deduced by comparing the conversion of SAPO-n catalysts with the CZ-280 catalyst. Thus, conversion is higher for this HZSM-5 catalyst, although its total acidity and average acid strength are lower. This lower conversion for SAPO-n has already been observed in the literature with experiments at 400 °C. 22 Furthermore, the more severe shape selectivity for the SAPO-n hinders bimolecular hydrogen transfer reactions, favoring the increase in the selectivity to light olefins and especially to propylene because it is the most reactive olefin. The difference in the results for the two SAPO-n must be attributed to their different acidity, given that the difference in their microporous structure is insignificant.25,26 SAPO-34 has higher acid site amount and average acid strength than SAPO-18 (Table 2). Therefore, conversion is higher (8.28 % versus 4.59 %), as well as the yield of C5-C10 hydrocarbons and light paraffins, which are the products of light olefin oligomerization and of their transformation by hydrogen transfer, respectively. In addition to a higher selectivity to light olefins, as shown in Figure 4b, the propylene fraction obtained is significantly higher for SAPO-n catalysts than for HZSM-5 zeolite catalysts. The maximum value of propylene fraction is obtained with CS-18 (1.35, corresponding to a propylene selectivity of 48.47 %), which is higher than the one for CS-34 (0.97). This advantage is explained because the lower acidity of SAPO-18 limits the development of propylene oligomerization-cracking reactions, whose final olefin is ethylene.8,39 For the same reason, a comparison of HZSM-5 zeolite catalysts shows that the


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15 highest propylene fraction (0.63) corresponds to CZ-280, the one with the lowest acid site amount and acid strength of the studied catalysts. The increase in space time also has a remarkable effect on the increase in conversion and product distribution. Table 3 shows that the conversion for 14.99 gcat h (molCH2)-1 is also the highest for CZ-30 catalyst (70.33 %) among the HZSM-5 zeolite catalysts and for CS-34 (17.23 %) among the SAPO-n catalysts. It is also observed that the yield of olefins increases by increasing the space time for all the catalysts, except for the most active one (CZ-30). This is explained because, under these reaction conditions, an increase in space time selectively enhances different steps of the reaction scheme in Figure 5. For the least acid catalyst, CZ280, the intermediate steps of the hydrocarbon pool mechanism formation and light olefin formation are favored. However, for the most active CZ-30 catalyst, these steps are not limiting and the subsequent ones are favored, specifically those of transformation of light olefins into byproducts, so that the yield of olefins decreases as space time is increased. As a consequence of the effect that CZ-30 and CZ-80 catalysts have on product distribution, olefin selectivity decreases by increasing space time. Furthermore, an increase in space time has not a significant effect on the value of propylene fraction. Figure 6, corresponding to the reaction at 450 ºC, extends the comparison of the catalysts, providing results for yields of product fractions (graph a) and selectivities to light olefins and propylene fraction (graph b), for a higher conversion of chloromethane than in Figure 4. As shown in Table 3, corresponding also to these conditions, chloromethane conversion for HZSM-5 zeolite catalysts is much higher than that for SAPO-n catalysts. Thus, the conversion is almost full (95.5 %) for CZ-30 catalyst and one fifth of this value for SAPOn catalysts. Comparing the results of HZSM-5 zeolite catalysts at this temperature, the maximum yield of C2-C4 olefins (Figure 6a) (41.62 %) corresponds to CZ-80 catalyst. This catalyst has an intermediate acidity and acid strength among the catalysts studied, which



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16 confers sufficient activity upon the primary step of olefin formation in the scheme in Figure 5, but not enough for byproduct formation step. This catalyst allows attaining a high conversion of chloromethane (88.23 %), with the yield of light olefins (41.62 %) being much higher than that observed at 350 ºC and for the same space time (15.23 %). A greater fraction of propylene (0.50) is also obtained, which reveals that it is a versatile catalyst for achieving a high olefin yield and selectivity under high conversion conditions. CZ-280 catalyst, of low acidity, is hardly active at 450 ºC, even for olefin formation. However, the high yield of aromatics is remarkable under these conditions of high conversion, particularly with CZ-30 catalyst, of higher acidity and acid strength, whereas the major byproducts with CZ-80 catalyst are C5-C10 aliphatics, which may potentially be recycled, as has been mentioned before, thereby increasing the interest of this catalyst. Furthermore, the conversion of chloromethane on SAPO-n catalysts is notably higher at 450 ºC than at 350 ºC, with values of 19.12 % and 21.21 % for CS-18 and CS-34, respectively. The yield of olefins also increase (13.17 % and 15.44 %, for CS-18 and CS-34, respectively), although the selectivity to olefins decreases (68.87 % and 72.98 %, respectively), and propylene fraction is also lower than that corresponding to 350 ºC. Moreover, the yield of methane is considerable at 450 ºC with all the studied catalysts (Figure 6a), with a maximum yield of 7 % for CZ-80 catalyst, although in an industrial process this methane would be fed back to the hydrochlorination step. It should be noted that the yield of methane in the transformation of chloromethane is much higher than in the transformation of methanol at the same temperature and on similar catalysts.8 The small yield of methane obtained with these catalysts is attributed to chloromethane cracking, which is a competitive reaction with the steps of hydrocarbon pool mechanism in the scheme in Figure 5. Figure 6


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17 3.2. Catalyst stability along time on stream Figure 7 compares the evolution with time on stream of chloromethane conversion for HZSM-5 zeolite (graph a) and SAPO-n (graph b) catalysts, at 350 ºC and for two values of the space time (5.89 and 14.99 gcat h (molCH2)-1). Figure 8 compares the results for all the catalysts at 450 ºC for the same value of space time (5.89 gcat h (molCH2)-1). Firstly, it is observed that conversion peaks goes through a maximum value, which is characteristic of the well-known hydrocarbon pool mechanism for the transformation of methanol.40 After an initial step of methoxy ion formation, intermediate polymethylbenzenes are generated, which leads to the release of olefins as primary products in the gas phase. Deactivation takes place firstly by blockage of acid sites, and subsequently by blockage of the micropores in the catalyst due to the presence of polymethylbenzenes. These compounds evolve into inert structures (hexamethylbenzenes) for the reaction and gradually condense into polyaromatic structures. This evolution takes place by means of mechanisms activated by acid sites and favored by the presence of cavities in the porous structure, which allow the location of the condensed compounds.23 Consequently, three periods are observed in Figures 7 and 8 for all the catalysts, in which the reaction is controlled successively by the steps of: i) induction (polymethylbenzene formation); ii) high conversion (with maximum release of light olefins), and; iii) decrease in conversion (when the deterioration of acidity and the porous structure of the catalysts is faster than the formation of active intermediates). Gayubo et al.41 have established a kinetic model to quantify the evolution of these steps in the transformation of methanol into hydrocarbons over a SAPO-18 catalyst. However, although the same periods as in transformation of methanol are qualitatively observed, for the transformation of chloromethane the deactivation rate (Figures 7 and 8) is remarkably faster, which has already been reported in literature.15,22,42,43 This difference can be related to the absence of steam in the reaction medium, which has a favorable effect by attenuating coke deposition in the



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18 transformation of methanol. This is explained by steam competence with coke precursors in their adsorption on the acid sites.

44,45

Consequently, the MTO process is performed co-

feeding water with methanol.46 Furthermore, a previous work dealing with the causes of deactivation (reversible and irreversible) for a HZSM-5 zeolite catalyst reports the presence of polymethylbenzenes and its derivatives in the coke.24 Figure 7 Figure 8 Furthermore, a significant effect of the catalyst on the conversion profiles is observed in Figures 7 and 8. Firstly, the induction period is longer for the least active catalysts, for which a slower formation of intermediate polymethylbenzenes is observed. Comparing the results for the three HZSM-5 zeolite catalysts (Figures 7a and 8) and those for SAPO-n catalysts (Figures 7b and 8), the difference may be associated with the different acidity of the catalysts, which favors the formation of methoxy ions and the subsequent generation of polymethylbenzenes. Consequently, the duration of the induction period for each family of catalysts is lower when acidity is increased. However, the severity of shape selectivity has a small impact. The different decay rate of chloromethane conversion with time on stream for the different catalysts is a relevant result, with the decay rate for each family of catalysts being related to acidity significance. Thus, the deactivation rate order for HZSM-5 zeolite catalysts is: CZ-30 > CZ-80 > CZ-280. For the SAPO-n catalysts: CS-34 > CS-18. This result may be attributed to the role of acidity in the formation of polymethylbenzenes and activation of their condensation towards structures that are inert for the formation of olefins, but block the micropores of the catalyst. This blockage is faster for SAPO-n catalysts, in whose microporous structure there are cavities that allow the location of polyaromatics,47 whereas in


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19 the HZSM-5 zeolites this blockage is hindered by the well-known capacity for heavy components displacement in the reaction medium towards the outside of the crystals.48,49 Therefore, from the perspective of minimizing the deactivation, both acidity and porous structure play an important role. Taking this into account, a HZSM-5 zeolite with a moderate value of acid site amount and strength is an interesting catalyst. Consequently, the CZ-80 catalyst has low deactivation (insignificant for 255 min time on stream at 350 ºC) and, as mentioned above, performs well concerning activity and selectivity. Concerning minimization of deactivation, space time has a minor effect, given that an increase in space time (Figure 7) decreases the duration of the induction period and the conversion of chloromethane increases. The decrease in conversion with time on stream has the same trend for lower values of space time. Furthermore, an increase in temperature from 350 to 450 ºC (Figure 8) has the negative effect of accelerating deactivation, which is very fast for all the catalysts, by increasing the rate of polymethylbenzene condensation. Table 4 sets out the content of carbonaceous material (Cc) deposited on the catalysts used under the different conditions. The values have been determined by combustion with air at 575 ºC, in a TGA Q 5000 TA thermobalance. These coke deposits are studied in catalytic processes because of their role in catalyst deactivation. However, in the transformation of methanol and chloromethane, the deposits corresponding to the induction period will correspond to active polymethylbenzenes in the hydrocarbon pool mechanism. These deposits will gradually lose its capacity as reaction intermediates and play the role of deactivating coke. Comparing the results in Table 4 with the deactivation profiles for the different catalysts in Figures 7 and 8, it can be concluded that the carbonaceous material content on the catalyst for 255 min of time on stream is consistent with the deactivation profiles and, as mentioned above, coke deposition is also related to acidity and porous structure. It may also be concluded that coke contents below 3.0 wt% do



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20 not have a significant effect on the conversion decrease with time on stream (they correspond to active polymethylbenzenes) and the conversion decrease is noticeable for higher coke contents. The effect of reaction temperature on coke deposition should also be noted, i.e., the highest contents are obtained at 450 ºC, especially for catalysts with high acidity, so that the coke content for CZ-30 catalyst is 20.41 wt%. Moreover, great differences are observed in the coke contents for the two families of catalysts, which may explain the different deactivation rate. This fact explains the essential role of shape selectivity by enhancing the blockage of SAPO-n crystals and, in the case of HZSM-5 zeolite, the flow of coke precursors towards the exterior of the crystals, thereby minimizing micropores blockage. 3.3. Stability in reaction-regeneration cycles Surely, the industrial implementation of the transformation of chloromethane into olefins will require the catalyst regeneration by coke combustion and the subsequent use of the regenerated catalyst in successive reaction-regeneration cycles, similarly to the MTO process (with SAPO-34 catalyst in a fluidized bed reactor and regeneration in another fluidized bed),50 and to the DTO process (with HZSM-5 zeolite catalyst, in a fixed bed reactor with in situ regeneration).51,52 The stability of the catalysts has been studied by operating in successive reaction-regeneration cycles, under the following conditions: Reaction step: 350 and 450 ºC; space time, 2.35 and 5.89 gcat h (molCH2)-1; time on stream, 75 min. The regeneration step involves two stages: i) coke aging by sweeping the catalyst with He (20 min at the reaction temperature with a flow rate of 30 cm3 min-1) in order to remove the volatile components, and; ii) coke combustion in situ in the reactor with air (1 h at 550 ºC with a flow rate of 30 cm3 min-1). It was found that after reaction at 350 °C, all the catalysts fully recover the initial activity, and the results of the evolution of conversion with time on stream are almost


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21 identical in the six cycles studied (not shown, because they are similar to those in Figure 7). However, after reaction at 450 °C, none of the catalysts recover completely their activity after coke combustion, as observed in Figure 9. This figure shows the chloromethane conversion evolution with time on stream for the fresh catalyst and after being regenerated once and twice. Since the results are qualitatively similar for all the catalysts, Figure 9 shows as an example the results corresponding to CZ-30 catalyst (Figure 9a) and CS-34 catalyst (Figure 9b), for two different values of space time. These results demonstrate the existence of an irreversible deactivation in the reaction step at 450 ºC. The cause of this deactivation is the catalyst dealumination by AlCl3 formation, as determined by different analytical techniques (SEM-EDX, XPS and NMR).24 At this high temperature, AlCl3 is formed by reaction of HCl byproduct with the Al present in the acid function and in the other components of the catalyst particle (bentonite and -Al2O3). Figure 9 4. Conclusions The kinetic scheme of the transformation of chloromethane into light olefins has great analogies with the transformation of methanol, with a characteristic initial step of polymethylbenzene formation as active intermediate in the hydrocarbon pool mechanism and with the subsequent formation of coke (derived from polymethylbenzene condensation) as the main cause of the deactivation. Consequently, the effect of shape selectivity and acidity of the catalysts on the kinetic behavior (activity, selectivity, deactivation) is qualitatively similar for both reactions. However, certain significant differences are appreciated, such as the lower chloromethane reactivity compared to methanol and the faster deactivation in the transformation of chloromethane than in methanol transformation, due to the lower capacity for the formation of methoxy active species and the absence of steam in the reaction medium, respectively. These differences explain the more pronounced effect of shape selectivity and



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22 catalyst acidity on the kinetic behavior. Therefore, HZSM-5 zeolite catalysts, with medium size micropores, have higher activity and slower deactivation than SAPO-n catalysts, with higher severity in the shape selectivity of their microporous structure and with boxes in the micropore intersections, which facilitate the deposition of coke. Similarly, the difference in olefin selectivity is also remarkable, being higher for SAPO-n catalysts but with a much lower chloromethane conversion than for HZSM-5 zeolite catalysts. These differences, observed for catalysts of similar acidity are attributable to the effect of shape selectivity in the mechanism for the formation of olefins from chloromethane. For each family of catalysts, acidity is a key factor for the control of the kinetic behavior, especially for increasing olefin and propylene selectivity, and for attenuating the deactivation by coke by decreasing the amount and strength of the acid sites. Accordingly, a HZSM-5 zeolite catalyst (CZ-80) with medium SiO2/Al2O3 ratio (80) strikes a suitable balance between reaction indices. Similarly, the SAPO-18 catalyst has a better kinetic behavior than the SAPO-34 due to its moderate acidity. At 350 ºC the unique cause of deactivation is the deposition of coke and the catalyst can operate in successive reaction-regeneration cycles. Therefore, when the reaction is performed at this temperature and the regeneration by coke combustion at 550 °C, the catalyst recovers its kinetic behavior, with high values of both conversion and selectivity to light olefins and propylene. However, the reaction temperature is limited by the dealumination in the reaction step, which is considerable at 450 ºC for all the catalysts (with Al in the acid function) due to the formation of AlCl3 at high temperature by reaction with HCl byproduct. Avoidance of this problem is a great challenge for the design of new catalysts.


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23 Acknowledgements All authors received funding from the Basque Government (Project IT748-13) and the University of the Basque Country (UFI 11/39). Authors M. Gamero, A.T. Aguayo, A. Ateka, P. Pérez-Uriarte and J. Bilbao received funding from the Ministry of Economy and Competitiveness of the Spanish Government and the ERDF funds (CTQ2013-46173-R). M. Gamero is grateful for the postdoctoral grant from the University of the Basque Country (No. UPV/EHU2013).



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24 Nomenclature Cc = content of carbonaceous material deposited on the catalyst, wt % dp = pore diameter, Å Fpropylene = fraction of propylene in the olefin lump (eq 4) SBET = BET surface area, m2 g-1 Si = selectivity to each i lump, in CH2 equivalent units (eq 3) T = temperature, ºC Vm, Vp = micropore and pore volume, cm3 g-1 X = conversion of chloromethane, in CH2 equivalent units (eq 1) Yi = yield of each i lump, in CH2 equivalent units (eq 2)

Abbreviations MCl, M, O, P, C5+, BTX = chloromethane, methane, C2-C4 olefins, C2-C4 paraffins, C5-C10 aliphatics, aromatics (benzene, toluene and xylenes), respectively


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25 References (1) Melikoglu, M. Shale Gas: Analysis of its Role in the Global Energy Market. Renewable Sustainable Energy Rev. 2014, 37, 460. (2) Chang, C. D. Mechanism of Hydrocarbon Formation from Methanol. Stud. Surf. Sci. Catal. 1988, 36, 127. (3) Taylor, C. E.; Noceti, R. P.; Schehl, R. R. Direct Conversion of Methane to Liquid Hydrocarbons through Chlorocarbon Intermediates. Stud. Surf. Sci. Catal. 1988, 36, 483. (4) Reddy Keshav, T.; Basu, S. Gas-to-Liquid Technologies: India's Perspective. Fuel Process. Technol. 2007, 88, 493. (5) Stöcker, M. Methanol-to-Hydrocarbons: Catalytic Materials and their Behavior. Microporous Mesoporous Mater. 1999, 29, 3. (6) Khadzhiev, S. N.; Kolesnichenko, N. V.; Ezhova, N. N. Manufacturing of Lower Olefins from Natural Gas through Methanol and its Derivatives (review). Pet. Chem. 2008, 48, 325. (7) Quann, R. J.; Green, L. A.; Tabak, S. A.; Krambeck, F. J. Chemistry of Olefin Oligomerization over ZSM-5 Catalyst. Ind. Eng. Chem. Res. 1988, 27, 565. (8) Aguayo, A. T.; Mier, D.; Gayubo, A. G.; Gamero, M.; Bilbao, J. Kinetics of Methanol Transformation into Hydrocarbons on a HZSM-5 Zeolite Catalyst at High Temperature (400550 ºC). Ind. Eng. Chem. Res. 2010, 49, 12371. (9) Epelde, E.; Aguayo, A. T.; Olazar, M.; Bilbao, J.; Gayubo, A. G. Kinetic Model for the Transformation of 1-Butene on a K-modified HZSM-5 Catalyst. Ind. Eng. Chem. Res. 2014, 53, 10599. (10) Wei, Y.; Zhang, D.; Xu, L.; Liu, Z.; Su, B. L. New Route for Light Olefins Production from Chloromethane over HSAPO-34 Molecular Sieve. Catal. Today 2005, 106, 84.



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26 (11) Wei, Y.; He, Y.; Zhang, D.; Xu, L.; Meng, S.; Liu, Z.; Su, B. L. Study of Mn Incorporation into SAPO Framework: Synthesis, Characterization and Catalysis in Chloromethane Conversion to Light Olefins. Microporous Mesoporous Mater. 2006, 90, 188. (12) Wei, Y.; Zhang, D.; Liu, Z.; Su, B. L. Highly Efficient Catalytic Conversion of Chloromethane to Light Olefins over HSAPO-34 as Studied by Catalytic Testing and in situ FTIR. J. Catal. 2006, 238, 46. (13) Wei, Y.; Zhang, D.; Liu, Z.; Su, B. L. Mechanistic Elucidation of Chloromethane Transformation over SAPO-34 Using Deuterated Probe Molecule: A FTIR Study on the Surface Evolution of Catalyst. Chem. Phys. Lett. 2007, 444, 197. (14) Wei, Y.; Zhang, D.; Xu, L.; Chang, F.; He, Y.; Meng, S.; Su, B. L.; Liu, Z. Synthesis, Characterization

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27 (20) Jaumain, D.; Su, B. L. Direct Catalytic Conversion of Chloromethane to Higher Hydrocarbons over a Series of ZSM-5 Zeolites Exchanged with Alkali Cations. J. Mol. Catal. A: Chem. 2003, 197, 263. (21) Noronha, L. A.; Souza-Aguiar, E. F.; Mota, C. J. A. Conversion of Chloromethane to Light Olefins Catalyzed by ZSM-5 Zeolites. Catal. Today 2005, 101, 9. (22) Xu, T.; Zhang, Q.; Song, H.; Wang, Y. Fluoride-treated H-ZSM-5 as a Highly Selective and Stable Catalyst for the Production of Propylene from Methyl Halides. J. Catal. 2012, 295, 232. (23) Jiang, Y.; Huang, J.; Reddy Marthala, V. R.; Ooi, Y. S.; Weitkamp, J.; Hunger, M. In situ MAS NMR-UV/Vis Investigation of H-SAPO-34 Catalysts Partially Coked in the Methanol-to-Olefin Conversion under Continuous-flow Conditions and of their Regeneration. Microporous Mesoporous Mater. 2007, 105, 132. (24) Ibañez, M.; Gamero, M.; Ruiz-Martínez, J.; Weckhuysen, B. M.; Aguayo, A. T.; Bilbao, J.; Castaño, P. Simultaneous Coking and Dealuminating of the HZSM-5 Catalyst Used in the Transformation of Chloromethane to Olefins. Catal. Sci. Technol. 2015 Accepted. (25) Wragg, D. S.; Akporiaye, D.; Fjellvag, H. Direct Observation of Catalyst Behaviour under Real Working Conditions with X-ray Diffraction: Comparing SAPO-18 and SAPO-34 Methanol to Olefin Catalysts. J. Catal. 2011, 279, 397. (26) Chen, J.; Li, J.; Wei, Y.; Yuan, C.; Li, B.; Xu, S.; Zhou, Y.; Wang, J.; Zhang, M.; Liu, Z. Spatial Confinement Effects of Cage-type SAPO Molecular Sieves on Product Distribution and Coke Formation in Methanol-to-Olefin Reaction. Catal. Commun. 2014, 46, 36. (27) Aguayo, A. T.; Gayubo, A. G.; Vivanco, R.; Olazar, M.; Bilbao, J. Role of Acidity and Microporous Structure in Alternative Catalysts for the Transformation of Methanol into Olefins. Appl. Catal. A: Gen. 2005, 283, 197.



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28 (28) Chen, J.; Thomas, J. M.; Wright, P. A.; Townsend, R. P. Silicoaluminophosphate Number Eighteen (SAPO-18): a New Microporous Solid Acid Catalyst. Catal. Lett. 1994, 28, 241. (29) Wendelbo, R.; Akporiaye, D.; Andersen, A.; Dahl, I. M.; Mostad, H. B. Synthesis, Characterization and Catalytic Testing of SAPO-18, MgAPO-18, and ZnAPO-18 in the MTO Reaction. Appl. Catal. A: Gen. 1996, 142, L197. (30) Guisnet, M.; Costa, L.; Ribeiro, F. R. Prevention of Zeolite Deactivation by Coking. J. Mol. Catal. A: Chem. 2009, 305, 69. (31) Kim, J.; Choi, M.; Ryoo, R. Effect of Mesoporosity Against the Deactivation of MFI Zeolite Catalyst During the Methanol-to-Hydrocarbon Conversion Process. J. Catal. 2010, 269, 219. (32) Gayubo, A. G.; Aguayo, A. T.; Atutxa, A.; Prieto, R.; Bilbao, J. Role of Reactionmedium Water on the Acidity Deterioration of a HZSM-5 Zeolite. Ind. Eng. Chem. Res. 2004, 43, 5042. (33) Mier, D.; Aguayo, A. T.; Gayubo, A. G.; Olazar, M.; Bilbao, J. Catalyst Discrimination for Olefin Production by Coupled Methanol/n-Butane Cracking. Appl. Catal. A: Gen. 2010, 383, 202. (34) Busch, O. M.; Brijoux, W.; Thomson, S.; Schuth, F. Spatially Resolving Infrared Spectroscopy for Parallelized Characterization of Acid Sites of Catalysts via Pyridine Sorption: Possibilities and Limitations. J. Catal. 2004, 222, 174. (35) Emeis, C. A. Determination of Integrated Molar Extinction Coefficients for Infrared Absorption Bands of Pyridine Adsorbed on Solid Acid Catalysts. J. Catal. 1993, 141, 347. (36) Zhu, X.; Liu, S.; Song, Y.; Xu, L. Catalytic Cracking of C4 Alkenes to Propene and Ethene: Influences of Zeolites Pore Structures and Si/Al2 Ratios. Appl. Catal. A: Gen. 2005, 288, 134.


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29 (37) Al-Dughaither, A. S.; De Lasa, H. HZSM-5 Zeolites with Different SiO2/Al2O3 Ratios. Characterization and NH3 Desorption Kinetics. Ind. Eng. Chem. Res. 2014, 53, 15303. (38) Mier, D.; Aguayo, A. T.; Gayubo, A. G.; Olazar, M.; Bilbao, J. Synergies in the Production of Olefins by Combined Cracking of n-Butane and Methanol on a HZSM-5 Zeolite Catalyst. Chem. Eng. J. 2010, 160, 760. (39) Tabak, S. A.; Krambeck, F. J.; Garwood, W. E. Conversion of Propylene and Butylene over ZSM-5 Catalyst. AIChE J. 1986, 32, 1526. (40) Aguayo, A. T.; Gayubo, A. G.; Vivanco, R.; Alonso, A.; Bilbao, J. Initiation Step and Reactive Intermediates in the Transformation of Methanol into Olefins over SAPO-18 Catalyst. Ind. Eng. Chem. Res. 2005, 44, 7279. (41) Gayubo, A. G.; Aguayo, A. T.; Alonso, A.; Bilbao, J. Kinetic Modeling of the Methanol-to-Olefins Process on a Silicoaluminophosphate (SAPO-18) Catalyst by Considering Deactivation and the Formation of Individual Olefins. Ind. Eng. Chem. Res. 2007, 46, 1981. (42) Svelle, S.; Aravinthan, S.; Bjørgen, M.; Lillerud, K.-P.; Kolboe, S.; Dahl, I. M.; Olsbye, U. The Methyl Halide to Hydrocarbon Reaction over H-SAPO-34. J. Catal. 2006, 241, 243. (43) Zhang, D.; Wei, Y.; Xu, L.; Chang, F.; Liu, Z.; Meng, S.; Su, B. L. MgAPSO-34 Molecular Sieves with Various Mg Stoichiometries: Synthesis, Characterization and Catalytic Behavior in the Direct Transformation of Chloromethane into Light Olefins. Microporous Mesoporous Mater. 2008, 116, 684. (44) Gayubo, A. G.; Aguayo, A. T.; Sánchez Del Campo, A. E.; Tarrío, A. M.; Bilbao, J. Kinetic Modeling of Methanol Transformation into Olefins on a SAPO-34 Catalyst. Ind. Eng. Chem. Res. 2000, 39, 292.



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30 (45) Gayubo, A. G.; Aguayo, A. T.; Morán, A. L.; Olazar, M.; Bilbao, J. Role of Water in the Kinetic Modeling of Catalyst Deactivation in the MTG Process. AIChE J. 2002, 48, 1561. (46) Tian, P.; Wei, Y.; Ye, M.; Liu, Z. Methanol to Olefins (MTO): From Fundamentals to Commercialization. ACS Catal. 2015, 5, 1922. (47) Mores, D.; Stavitski, E.; Kox, M. H. F.; Kornatowski, J.; Olsbye, U.; Weckhuysen, B. M. Space- and Time-resolved in-situ Spectroscopy on the Coke Formation in Molecular Sieves: Methanol-to-Olefin Conversion over H-ZSM-5 and H-SAPO-34. Chem. A Eur. J. 2008, 14, 11320. (48) Castaño, P.; Elordi, G.; Olazar, M.; Aguayo, A. T.; Pawelec, B.; Bilbao, J. Insights into the Coke Deposited on HZSM-5, Hβ and HY Zeolites During the Cracking of Polyethylene. Appl. Catal. B: Environ. 2011, 104, 91. (49) Elordi, G.; Olazar, M.; Lopez, G.; Castaño, P.; Bilbao, J. Role of Pore Structure in the Deactivation of Zeolites (HZSM-5, Hβ and HY) by Coke in the Pyrolysis of Polyethylene in a Conical Spouted Bed Reactor. Appl. Catal. B: Environ. 2011, 102, 224. (50) Chen, J. Q.; Bozzano, A.; Glover, B.; Fuglerud, T.; Kvisle, S. Recent Advancements in Ethylene and Propylene Production Using the UOP/Hydro MTO Process. Catal. Today 2005, 106, 103. (51) Biryukova, E. N.; Goryainova, T. I.; Kulumbegov, R. V.; Kolesnichenko, N. V.; Khadzhiev, S. N. Conversion of Dimethyl Ether into Lower Olefins on a La-Zr-HZSM5/Al2O3 Zeolite Catalyst. Pet. Chem. 2011, 51, 49. (52) Kolesnichenko, N. V.; Goryainova, T. I.; Biryukova, E. N.; Yashina, O. V.; Khadzhiev, S. N. Synthesis of Lower Olefins from Dimethyl Ether in the Presence of Zeolite Catalysts Modified with Rhodium Compounds. Pet. Chem. 2011, 51, 55.


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31

FIGURE CAPTIONS Figure 1. Acid strength distribution of HZSM-5 zeolite catalysts, measured by calorimetrythermogravimetry of differential adsorption (a) and by TPD (b) of NH3. Figure 2. Acid strength distribution of SAPO-n catalysts, measured by calorimetrythermogravimetry of differential adsorption (a) and by TPD (b) of NH3. Figure 3. Reaction equipment. Figure 4. Comparison of product yields (a) and selectivity to light olefins and propylene fraction (b) for different catalysts. Reaction conditions: 350 ºC; space time, 5.89 gcat h (molCH2)-1; time on stream, 10 min.

Figure 5. Evolution of product formation with the reaction extent. Figure 6. Comparison of product yields (a) and selectivity to light olefins and propylene fraction (b) for different catalysts. Reaction conditions: 450 ºC; space time, 5.89 gcat h (molCH2)-1; time on stream, 10 min.

Figure 7. Evolution with time on stream of chloromethane conversion, for two space time values, for HZSM-5 zeolite (a) and SAPO-n (b) catalysts, at 350 ºC. Figure 8. Comparison of the evolution of chloromethane conversion with time on stream for different catalysts. Reaction conditions: 450 ºC; space time, 5.89 gcat h (molCH2)-1. Figure 9. Evolution of chloromethane conversion with time on stream in three successive reaction steps, with coke combustion between steps, for CZ-30 (a) and CS-34 (b) catalysts. Reaction conditions: 450 ºC and 2.35 gcat h (molCH2)-1 (a) and 5.89 gcat h (molCH2)-1(b).



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32

Table 1. Physical properties of the acid functions and catalysts.

SBET, m2g-1 Z-30 Z-80 Z-280 S-18 S-34

497 540 485 798 611

CZ-30 CZ-80 CZ-280 CS-18 CS-34

208 203 219 216 207

Acid functions Vm, Vp, cm3 g-1, 3 -1 cm g 1.7

Role of Shape Selectivity and Catalyst Acidity in the Transformation of

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