Catalyst Equilibration for Transformation of Methanol into

Apr 1, 1996 - (Aguayo et al., 1994a; Guisnet and Magnoux, 1994,. Benito et al., 1996); (2) irreversible loss of activity during reaction steps, due to...
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Ind. Eng. Chem. Res. 1996, 35, 2177-2182

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Catalyst Equilibration for Transformation of Methanol into Hydrocarbons by Reaction-Regeneration Cycles Pedro L. Benito,* Andre´ s T. Aguayo, Ana G. Gayubo, and Javier Bilbao Departamento de Ingenierı´a Quı´mica, Universidad del Paı´s Vasco, Apartado 644, 48080 Bilbao, Spain

A study has been made of the activity recovery of H-ZSM5 zeolite based catalysts, which have been used in reaction-regeneration cycles in the transformation of methanol into gasoline. Catalyst stability is sensitive to catalyst calcination conditions (temperature and time) and to zeolite Si/Al ratio. From experiments carried out on automated reaction-regeneration equipment provided with an isothermal fixed-bed reactor, optimum catalyst equilibration conditions have been determined, in order to minimize the irreversible deactivation in the regeneration step. The effect of this step on the catalyst acid structure has been studied. 1. Introduction The possible causes of catalyst deactivation during the MTG (methanol to gasoline) process carried out in reaction-regeneration cycles are (Allum and Williams, 1988) (1) carbonaceous residue (coke) deposition inside catalyst particles during the reaction step, which is originated by methanol and/or product degradation (Aguayo et al., 1994a; Guisnet and Magnoux, 1994, Benito et al., 1996); (2) irreversible loss of activity during reaction steps, due to the effect of steam (product) upon acid sites and the structure of the zeolite (Nayak and Choudhary, 1984); and (3) zeolite acid structure alteration (transformation of Bronsted sites into Lewis sites) during the regeneration steps by coke combustion (Liederman et al., 1978; Yurchak et al., 1979; Bibby et al., 1992). Irreversible loss of activity (causes 2 and 3) must be eliminated or attenuated as far as possible in the industrial process. Equilibration of the catalyst is the usual procedure, which consists of a thermal treatment (with or without steam) that gives way to an unavoidable decrease in the initial catalyst activity. In this paper, methanol transformation into gasoline in reaction-regeneration cycles has been studied with the aim of establishing thermal equilibration conditions enabling the catalyst to recover its activity. Adopting these measures, coke deposition on the catalyst is expected to be the only cause of catalyst deactivation, which will be reversible; that is, the catalyst will completely recover its activity after regeneration steps. The study has been extended to catalysts prepared with H-ZSM5 zeolites of different Si/Al ratios, with the aim of analyzing the effect of fresh catalyst surface acidity on the equilibration treatment. The study has been carried out in an isothermal fixedbed reactor, because the temperature must be clearly delimited. Extrapolation of results to cyclic operation in the industrial process carried out in an adiabatic fixed bed will require control of the bed temperature profile so as not to exceed the equilibration temperature. 2. Experimental Section 2.1. Catalysts. Five ZSM5 zeolites have been synthesized with different Si/Al ratios, 24, 32, 42, 78, and 154, by following the method detailed in a previous paper (Benito et al., 1994), according to Mobil patents (Argauer and Landolt, 1972; Chen et al., 1973) from sodium silicate, aluminum sulfate, and tetra-n-propylammonium nitrate. S0888-5885(95)00493-3 CCC: $12.00

The final catalyst, as is used in the reactor, is obtained by physical mixing, under high humidity conditions, of the zeolite (25 wt %) with a binder (bentonite, 30 wt %) and an inert (R-alumina supplied by Martinswerk and calcined at 1100 °C, 45 wt %). The denominations of the catalysts are the following: Z1B (Si/Al ) 24), Z2B (Si/Al ) 32), Z3B (Si/Al ) 42), Z4B (Si/Al ) 78), Z5B (Si/Al ) 154). The physical properties of the H-ZSM5 zeolites and contributions to the pore volume of the catalysts from different diameter pores (measured in an ASAP 2000 N2 adsorption-desorption apparatus and in an Autopore 9220 II Hg intrusion porosimeter, both from Micromeritics) are set out in Table 1. The nature of the acidic sites has been studied by FTIR analysis after pyridine adsorption (Ve´drine et al., 1979). The equipment used is a catalytic chamber, Spectra Tech, connected in series with a FTIR Nicolet 740 spectrophotometer. The surface acidic strength distribution and the total acidity have been studied by measurement of the differential adsorption heat of NH3, by combining differential scanning calorimetry (Setaram DSC 111) and FTIR spectrophotometry in order to measure both the heat of neutralization of the acidic sites and the amount of base chemically absorbed (Aguayo et al., 1994b). Once the sample is saturated at 200 °C, the total acidity measurement is also obtained by temperature-programmed desorption (TPD) of the base, by following a ramp of 5 °C/min between 200 and 600 °C, with He flow of 20 cm3/min and using FTIR spectrophotometry for measurement of desorption products. 2.2. Reaction-Regeneration Equipment and Product Analysis. The automated equipment has been detailed in previous papers (Arandes et al., 1990; Gayubo, 1992; Gayubo et al., 1993). The equipment, which is controlled by computer, allows for uninterrupted reaction-regeneration cycles, where the length of time and regimes of flow rates and temperatures can be programmed in each one of the steps (reaction, stripping, combustion, conditioning between steps). A fixed-bed reactor of 0.007 m i.d., which has a coil to preheat the gases, is used, and both are in an oven of electric resistances, in whose design the inertia to the computer-controlled temperature response has been minimized. Temperature is measured by means of three thermocouples introduced in the bed: one in the middle point of the bed axis, another one at the outlet of the bed axis, and, finally, one on the inside of the bed wall. The products pass through a 10-port valve © 1996 American Chemical Society

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Table 1. Properties of the Synthesized Zeolites and Pore Volume Distribution of the Catalysts (Contribution of Pores of Different Diameter, dp) zeolites denomination

Si/Al

BET area, m2/g

H-ZSM5-1 H-ZSM5-2 H-ZSM5-3 H-ZSM5-4 H-ZSM5-5

24 32 42 78 154

420 405 440 365 395

catalysts Vpore, cm3/g

density, g/cm3

0.65 0.62 0.64 0.62 0.63

0.94 0.93 0.96 0.97 0.97

% micropore denomination (dp < 10-3 µm) Z1B Z2B Z3B Z4B Z5B

8.0 8.6 8.8 7.5 8.1

% mesopore % macropore (10-3 µm < dp < 0.01 µm) (0.01 µm < dp < 2 µm) 14.7 15.3 15.5 14.0 14.6

77.2 76.0 75.7 78.4 77.2

Figure 1. Evolution of gaseous products of coke combustion measured by FTIR analysis.

Figure 2. Product distribution for Z1B catalyst regenerated by coke combustion during different times.

that allows for a sample to be sent to the HewlettPackard 5890 Series II chromatograph. The chromatograph has an HP 3390-A integrator, which is provided with a card enabling the results of the analysis to be sent to a computer by means of RS-232-C interface. The feed-reaction-analysis system is controlled by a computer run. The calculation of the weight fraction of the lumps of oxygenates (methanol and dimethyl ether), of light olefins (ethylene, propylene, and butenes), and of the rest of the products is carried out by means of a program in FORTRAN, on the basis of the composition of each individual product obtained in the chromatographic runs.

The minimum regeneration time at 550 °C has been determined by operating in reaction-regeneration cycles with regeneration steps of different lengths of time. Reaction step conditions are as follows: Z1B catalyst, calcined at 550 °C; temperature, 350 °C; contact time, 0.03 (g of catalyst) h (g of methanol)-1; methanol partial pressure in the feed, 88 kPa; time on stream, 6 h. The different regeneration steps have been carried out with increasing lengths of time: 0, 5, 10, 15, 20, 30 and 40 min, with an air flow of 50 cm3/min. After each regeneration, during the conditioning period until the reaction temperature is reached, 10 min, the catalyst is subjected to a sweeping-stabilizing treatment with a He stream of 50 cm3/min. Similarly, prior to every regeneration step, a sweeping treatment with He (50 cm3/min) at 550 °C has been carried out. In Figure 2, the distribution of products expressed as weight fraction on a water-free basis is shown for the successive reaction cycles as a function of the length of time of the regeneration step. It can be observed that, for 30 min coke combustion, an activity level (or product composition) is reached which does not improve with longer combustion times. Comparing these results with those of Figure 1, it is evident that the catalyst tolerates the presence of a small amount of nonoxidized residual coke. This result has already been observed (Copperthwaite et al., 1986). A possible explanation lies in the location of the residual coke within the intersections of the zeolite channels (Ortega et al., 1996), which is due to combustion difficulty of the internal coke. There is little probability of the existence of external residual coke due to oxygen accessibility to the external surface of zeolite crystals. The small amount of residual coke is inert and does not inhibit the flow of the reaction components and the development of the reaction mechanism, which takes place on the acid sites preferably located on the intersections of the zeolite channels. 3.2. Effect of Si/Al Ratio on the Catalyst Performance for Reaction-Regeneration Cycles. The catalyst performance has been studied in six successive

3. Results 3.1. Regeneration Conditions. The selection of the conditions (temperature and time) for regeneration by coke combustion in air is based on two opposing criteria that directly influence cycle economy: (1) the higher the temperature used, the shorter the time needed for regeneration; (2) regeneration temperature affects the acid structure and, consequently, the catalyst activity. The study of the conditions needed for coke combustion has been carried out in a Setaram DSC 111 differential scanning calorimeter on line with a Nicolet 740 FTIR spectrophotometer for analysis of the gaseous products (CO, CO2, H2O). The results at 550 °C are shown in Figure 1 for coke deposited in a Z1B catalyst deactivated under the following conditions: temperature, 350 °C; contact time, 0.052 (g of catalyst) h (g of methanol)-1; time on stream, 2 h. Prior to combustion, the deactivated catalyst was subjected to sweeping in the calorimeter, with He (50 cm3/min) at 550 °C during 1 h, in order to stabilize the coke structure by aging (Ortega et al., 1996). In Figure 1, it can be observed that at 550 °C total combustion of the coke can be reached, although a period longer than 1 h is needed. Nevertheless, more than 90 wt % of the initial coke has been oxidized in the first 30 min.

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Figure 3. Product distribution in the operation under reaction-regeneration cycles with the five catalysts studied.

reaction-regeneration cycles using the five catalysts prepared with zeolites synthesized with different Si/Al ratios. The reaction has been carried out under the following conditions: temperature, 350 °C; contact time, 0.052 (g of catalyst) h (g of methanol)-1; methanol partial pressure in the feed, 88 kPa; time on stream, 2 h. The sweeping treatments with He have been carried out as detailed in the previous section. Results of product distribution at zero time on stream, which have been expressed as weight fraction on a water-free basis, are shown in Figure 3 against the number of cycles. An obvious deterioration of the activity of the catalysts, when subjected to the first regeneration, can be observed in Figure 3. This activity deterioration is less important as the Si/Al ratio increases, which means a higher stability in reaction-regeneration cycles of the catalysts with a smaller Al content. The activity loss during the first cycle is a result of the effect that the regeneration treatment has on the acid sites of the catalyst, due to the formation of hot spots in the individual particles during the coke combustion (al-

though acceptable bed isothermicity is reached). This subject deserves attention, although it is beyond the scope of this paper. In a previous paper (Gayubo et al., 1996) it was observed that the Si/Al ratio increase within the 24-154 range has, as a main consequence, a notable reduction in the amount of active sites. The sites are of similar strength, and only a very slight increase in the proportion of very strong acid sites is observed. An assumption is made that the sites not recuperated in the first regeneration are in positions within zeolite crystals, in which a permissable site density is exceeded. Under these conditions, the proximity between sites and between coke combustion sites facilitates dehydroxylation. The increase in catalyst thermal stability as the Si/Al ratio increases can be attributed to its lower total acidity and lower density of Bronsted and Lewis sites. As the Bronsted sites are separated further from each other, their capacity for dehydroxylation and for conversion into Lewis sites (less active for hydrocarbon generation) decreases. Consequently, the Z5B catalyst does

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Figure 4. Effect of calcination temperature upon product distribution for Z1B catalyst by operating under reaction-regeneration cycles.

Figure 5. Effect of calcination time at 550 °C upon product distribution for Z1B catalyst by operating under reactionregeneration cycles.

not suffer any measurable loss of activity in the first regeneration. It must be pointed out that the results of Figure 3 correspond to catalysts previously calcined at 550 °C during 2 h. These conditions do not correspond to the maximum initial activity, and they already imply a severe equilibration of the catalyst. These conditions have been established because regeneration by combustion in order to remove the coke completely must be carried out at 550 °C. The fact that all the catalysts studied recover their activity during the second and ensuing regenerations indicates that they do not suffer any irreversible deactivation because of the presence of steam (product) in the reaction steps (at 350 °C). The deactivation in the reaction step is exclusively due to coke deposition, and, consequently, it is reversible. Nevertheless, strong acid sites that remained after the previous calcination (carried out at the same temperature, 550 °C) are removed during the first regeneration. A great effect of the Si/Al ratio on selectivity is observed in Figure 3. As the Si/Al ratio increases, selectivity to light olefins increases and the production of hydrocarbons decreases. Z1B is the most interesting catalyst with respect to production of hydrocarbons in the gasoline boiling point range (Benito, 1995). 3.3. Catalyst Equilibration. From the results shown in previous sections the best performance in reaction-regeneration cycles corresponds to Z1B catalyst (calcined at 550 °C during 2 h). With the aim of avoiding the deterioration of the catalyst activity during the first regeneration, the catalyst equilibration has been studied by changing the temperature and time of calcination. The equilibration treatment will be of interest if it does not provoke a severe deterioration in initial activity. The distribution of products in successive cycles carried out with Z1B catalyst calcined during 2 h at 550, 560, 570, and 600 °C has been compared in Figure 4. The experiments have been carried out under the conditions detailed in the previous section. It can be observed that, as the calcination temperature rises from 550 to 560 °C, the catalyst initial activity is smaller, but, despite this, the catalyst also suffers activity deterioration after the first cycle. It is observed in Figure 4 that, when calcining at 570 and at 600 °C, the catalyst equilibration is achieved (it does not suffer activity deterioration in the first reaction-regeneration cycle). Nevertheless, an unwanted consequence is that the catalyst calcined at these temperatures has a lower initial activity than the catalyst calcined at 550 °C.

Even though the result of Figure 4 is conclusive about the equilibration of the catalyst at 570 °C under the conditions in which the cycles have been carried out, a reasonable doubt arises about the validity of this result under all reaction conditions (temperature, contact time, time on stream). Figure 5 shows the distribution of products in successive reaction-regeneration cycles carried out under the same conditions as those in the previous section with Z1B catalyst calcined at 550 °C for different times. It is seen that the effect of the calcination time is less important than the effect of temperature (Figure 4). In any case, a slight equilibration improvement of the catalyst is appreciated when the calcination time increases. Unfortunately, this effect is accompanied by a slight deterioration of the activity of the fresh catalyst. It must be pointed out that the results of calcination in a steam environment are of no interest, because the dealuminating effect of steam causes structural alterations in the H-ZSM5 zeolite and the total conversion to hydrocarbons drastically decreases. 3.4. Characterization of the Equilibrated Catalyst. Analysis of the porous structure and surface area of the five catalysts after six reaction-regeneration cycles indicates that, within experimental error, fresh catalyst values are maintained. This result shows that the small amount of residual coke remaining after the regenerations does not inhibit the internal diffusion of the reaction components, which explains the previously obtained result that the catalyst performance is unaltered. Therefore, the activity loss of the catalysts during the first regeneration must be attributed to surface acidity deterioration. Differential heat curves for NH3 adsorption are shown in Figure 6 for fresh and regenerated Z1B catalyst (after being used for 2 h at 350 °C with a contact time of 0.052 (g of catalyst) h (g of methanol)-1 and a methanol partial pressure in the feed of 88 kPa). A decrease in the total acidity from 0.52 to 0.48 mmol of NH3 (g of zeolite)-1 is observed, which is a consequence of the disappearance of a fraction of strong acid sites of the fresh catalyst (those with an adsorption heat higher than 160 kJ (mol of NH3)-1). The decrease in the total acidity after the first reaction-regeneration cycle has also been measured by temperature-programmed desorption (TPD) of NH3. This measurement decreases with the catalyst regeneration from 0.50 to 0.45 mmol of NH3 (g of zeolite)-1. FTIR spectra of pyridine adsorption on fresh and regenerated Z1B catalyst are shown in Figure 7. An

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Si/Al ratio of 24 at 550 °C during 2 h, an optimum equilibration state of the catalyst is reached. This catalyst has a high initial activity which slightly decreases after the first regeneration; the remaining activity is recovered in the following regeneration steps. This result is a consequence of the loss of strong Bronsted acid sites in the first regeneration step. The catalyst instability in the first regeneration can be avoided by means of a severe thermal treatment of equilibration, but at the temperature needed for this, 570 °C, the decrease of the catalyst activity is noteworthy. Acknowledgment Figure 6. Differential adsorption heat of NH3 for fresh and regenerated Z1B catalyst.

This work was carried out with financial support from the University of the Basque Country/Euskal Herriko Unibertsitatea (Project UPV 069.310-EB004/92) and from DGICYT (Project PB90-0655). Literature Cited

Figure 7. FTIR spectra of pyridine adsorption for fresh and regenerated Z1B catalyst.

attenuation of the pyridine adsorption bands is observed for the regenerated catalyst. These bands correspond to pyridine bound to Bronsted sites (bands at 3256, 3180, and 3078 cm-1 and peaks at 1628 (not well resolved), 1545, and 1489 cm-1); to pyridine coordinated to Lewis sites (peaks at 1610, 1489 and 1453 cm-1) (Ve´drine et al., 1979; Datka and Tuznik, 1986; Datka and Piwowarska, 1988). A greater decrease of the band corresponding to Bronsted sites, 1545 cm-1, than the one to Lewis sites, 1453 cm-1, is observed. Consequently, the Bronsted/Lewis site ratio decreases with regeneration from 3.1 to 2.9. 4. Conclusions For catalysts prepared from H-ZSM5 zeolites with different Si/Al ratios, it has been verified that the minimum conditions for the reactivation of the catalyst by coke combustion with air are as follows: temperature, 550 °C; time, 30 min. After regeneration under these conditions, a small amount of residual coke (presumably in the channel intersections of the zeolite crystals) still remains, which is tolerated by the catalyst. Catalyst behavior in reaction-regeneration cycles is sensitive to temperature and calcination time and to the Si/Al ratio of the zeolite. Increasing the Si/Al ratio within the range from 24 to 154 and/or increasing the calcination temperature increases catalyst stability at the expense of the initial activity deterioration. By calcining the catalyst prepared from the zeolite with a

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Received for review August 8, 1995 Revised manuscript received January 18, 1996 Accepted February 1, 1996X IE950493U

X Abstract published in Advance ACS Abstracts, April 1, 1996.