Effect of Operating Conditions on Dimethyl Ether Steam Reforming in

Feb 11, 2014 - This paper studies the effect of operating conditions in dimethyl ether (DME) steam reforming on a bifunctional catalyst synthesized wi...
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Effect of Operating Conditions on Dimethyl Ether Steam Reforming in a Fluidized Bed Reactor with a CuO−ZnO−Al2O3 and Desilicated ZSM‑5 Zeolite Bifunctional Catalyst Jorge Vicente, Javier Ereña,* Lide Oar-Arteta, Martin Olazar, Javier Bilbao, and Ana G. Gayubo Departamento de Ingeniería Química, Universidad del País Vasco UPV/EHU, Apartado 644, 48080 Bilbao, Spain ABSTRACT: This paper studies the effect of operating conditions in dimethyl ether (DME) steam reforming on a bifunctional catalyst synthesized with CuO−ZnO−Al2O3 metallic function and a HZSM-5 zeolite treated with NaOH to moderate acidity. The experimentation has been carried out in a fluidized bed reactor in the 225−325 °C range, with space time between 0.1 and 2.2 gcatalyst h/gDME, steam/DME molar ratio between 3 and 6, and DME partial pressure between 0.08 and 0.25 bar. The 275−300 °C range is suitable for obtaining high values of DME conversion and H2 yield with minimum CO formation and deactivation by coke and avoiding hydrocarbon formation. Stable values of DME conversion (0.85), H2 yield (0.81), and H2 production rate (180 mmolH2/(gcatalyst h)) are obtained during 48 h at 300 °C with a steam/DME ratio of 4 and space time of 0.60 gcatalyst h/gDME. The main cause of deactivation is coke deposition on the metallic function.

1. INTRODUCTION The steam reforming of dimethyl ether (SRD) is a promising way of producing hydrogen-rich gas for fuel cell systems, as it can take place at low temperatures only slightly higher than those required for the steam reforming of methanol (SRM).1 The use of dimethyl ether (DME) as a raw material for producing H2 by steam reforming has several advantages compared to methanol, such as high hydrogen content (13 wt % vs 12.5 wt % of methanol), no toxicity or hazard factor, gaslike property, liquid-storage density, and available handling infrastructure (similar to liquefied petroleum gas).2 Furthermore, the single-step synthesis of DME on a bifunctional catalyst is considered a suitable process for the large-scale valorization of CO2, given it is thermodynamically more favorable than the synthesis of methanol, which allows cofeeding CO2 together with the syngas.3−6 The SRD reaction proceeds over bifunctional catalysts via the hydrolysis of DME over the acid function, followed by steam reforming of methanol (MeOH) over the metallic function. The individual reactions are

conversion and high H2 selectivity by minimizing the formation of CO (a poison for the anode catalyst in proton exchange membrane (PEM) fuel cells) and CH4. The more widely studied metallic functions for SRD are those commonly used in the methanol reforming process (no bifunctional catalyst is required), which may be grouped into two types:8,9 (i) those containing Cu and (ii) those containing metals from groups 8− 10. The most widely studied Cu-based metallic function is CZA (CuO−ZnO−Al2O3) with numerous commercial and laboratory-synthesized modifications,1,10−16 although other copperspinel type functions have recently been proposed for SRD, with the aim being a higher resistance to sintering.17−21 Although the latter have their advantages compared to Cubased catalysts, such as their higher thermal and long-term stability, most studies report their lower activity and selectivity for methanol reforming, given that they predominantly catalyze methanol decomposition,8,9 thus producing CO and H2. A water gas shift (WGS) reaction takes place in the presence of water, partially converting CO to CO2, but a significant amount of CO is produced via decomposition. Nevertheless, concerning Cu-based catalysts, there is some agreement in the literature on the existence of a pathway through a methyl formate intermediate, which directly releases CO2 and H2. Consequently, the CO produced during methanol steam reforming over Cu-based catalysts is formed by the reverse WGS reaction. Other catalysts, such as Pd/ZnO alloys used in methanol reforming (not assayed for SRD), provide similar results to those obtained with Cu-based ones.22 Furthermore, Pd metallic functions in SRD lead to higher yields of byproducts (especially CO),23 so subsequent WGS steps are required to purify the H2 stream.

ΔH ° = 24 kJ/mol

(CH3)2 O + H 2O ↔ 2CH3OH

(1)

CH3OH + H 2O ↔ 3H 2 + CO2

ΔH ° = 49 kJ/mol (2)

The SRD and the reverse water gas shift reaction (r-WGSR) generally take place over a metallic function, and methane and hydrocarbons are also generated via DME decomposition when a strong acidic function or high temperatures are used:7 CO2 + H 2 ↔ H 2O + CO

(CH3)2 O → CH4 + H 2 + CO

ΔH ° = 41 kJ/mol

(3)

ΔH ° = − 1 kJ/mol

Received: Revised: Accepted: Published:

(4)

Consequently, suitable metallic and acid functions are required in the bifunctional catalysts for attaining high DME © 2014 American Chemical Society

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water) and dried at 110 °C, and subsequently calcined at 550 °C for 3 h in order to obtain its acid form. The treatment of the HZSM-5 with NaOH efficiently decreases the acid strength and generates new mesopores by breaking Si−O−Al bonds (desilication) and maintaining Si−OH−Al bonds.40,41 This treatment under moderate conditions (300 °C and 0.4 M NaOH solution) is suitable for increasing the stability of DME reforming catalysts by minimizing the formation of hydrocarbons and coke.34 The catalyst has been prepared by wet physical mixing of the previously calcined metallic and acid functions, with a mass ratio of 1:1, which is suitable for attaining a synergy between the steps of DME hydrolysis (eq 1) and methanol reforming (eq 2).38 The resulting mixture is dried at 110 °C and then pressed, ground, and sieved in the 150−250 μm range. The following properties have been determined for the catalyst: composition, by inductively coupled plasma atomic electron spectroscopy (ICP-AES) in an ARL Model 3410; BET surface area (SBET) and porous structure, by N2 adsorption− desorption in a Micromeritics ASAP 2010; metal surface area and metal crystal size, by N2O chemisorption in a Micromeritics AutoChem 2920 connected online to a Balzers Instruments Omnistar mass spectrometer; and total acidity and acid strength, by differential adsorption of ammonia at 150 °C in a TA Instruments SDT 2960 connected online to a Balzers Instruments Thermostar mass spectrometer. Table 1 sets out

Concerning the acid function, γ-Al2O3 is the more commonly used one, but due to its low acidity, a high temperature (usually in the 300−400 °C range) is required for DME hydrolysis, which is the rate-limiting step in the overall reaction under these conditions.10,17−19,24 Consequently, the use of γ-Al2O3 as the acid function promotes DME decomposition and the reverseWGS reaction, which considerably increase CH4 and CO formation, as well as the sintering of Cu in the CuO−ZnO− Al2O3 metallic function. HZSM-5 zeolite is an attractive acid function because it allows reforming at a lower temperature (approximately 100 °C lower) than using γ-Al2O3 does. This is due to its higher acidity,12,25 enhancing the hydrolysis reaction, which even in this case is slower than methanol reforming. However, it may favor the formation of hydrocarbons, via MTH (methanol to hydrocarbons) reactions,26−30 which drastically diminish H2 production,7,31−33 and also may contribute to increasing deactivation by coke deposition.13 The use of HZSM-5 zeolites with a high Si/Al ratio (>90), i.e., of moderate acidity, attenuates the formation of hydrocarbons and increases catalyst stability.12,25 It has been previously reported that an alkaline treatment of the HZSM-5 zeolite with a 0.4 M NaOH solution for 300 min is suitable for the use of this zeolite as the acid function in DME steam reforming,34 as this alkaline treatment moderates HZSM-5 acidity and contributes to attenuating the formation of coke.35,36 Consequently, the bifunctional catalysts based on CuO− ZnO−Al2O3 as the metallic function and an alkali treated HZSM-5 zeolite as the acid function are considered promising catalysts for SRD. Given the catalyst’s kinetic performance is a consequence of the synergism between the activities of the metallic and acid functions, previous studies have paid special attention to the preparation and composition of the CuO− ZnO−Al2O3 metallic function,37 as well as to the mass ratio between this function prepared under the best conditions (Cu/ Zn/Al atomic ratio = 4.5:4.5:1.0) and the treated HZSM-5 zeolite,34 with the optimum zeolite/metallic function mass ratio being 1:1.38 This paper addresses the effects that operating conditions (temperature, space time, DME partial pressure, steam/DME molar ratio) have on the performance (DME conversion, yield and selectivity of H2, and stability) of the catalyst prepared under the conditions optimized in the previous studies mentioned. The results of this paper allow establishing suitable conditions for the process in order to maximize DME conversion and H2 yield, and minimize both the formation of byproducts and catalyst deactivation. An overall structured study of all these operating variables, paying special attention to catalyst stability, is of great interest, given that it is a key factor for process viability. The use of a fluidized bed reactor is also of interest for ensuring bed isothermicity, and thereby increasing catalyst stability.

Table 1. Properties of the Bifunctional Catalysts, with the Same CZA Metallic Function (Cu/Zn/Al = 4.5:4.5:1.0, Nominal) and with HZ30 (Parent Zeolite with SiO2/Al2O3 = 30) and A0.4-300 (Alkali Treated Zeolite) Acid Functions acid function HZ30

A0.4-300

wt % metallic function (nominal) wt % metallic function (real) SBET, m2/gcatalyst Vmicropore, cm3/g Vmesopore, cm3/g dmesopore, Å Smetallic, m2Cu/gCu dCu, nm total acidity, mmolNH3/gcatalyst

50 51.5 230 0.057 0.181 54 56 12.0 0.42

50 48.6 207 0.041 0.330 82 56 11.9 0.34

mean acid strength, kJ/molNH3

125

103

the properties of the catalysts prepared based both on the HZSM-5 zeolite treated with alkali (A0.4-300), which has been used in this paper, and on the parent (untreated) zeolite (HZ30). The physical properties correspond to a combination of the metallic and acid functions. Microporosity is provided by the zeolite acid function and mesoporosity by the metallic and acid functions. Treating the zeolite with alkali considerably increases the mesopore volume and BET surface area and slightly decreases the micropore volume, although zeolite desilication does not cause significant changes in the specific metal surface area or in the Cu particle average size.34 The capacity of the CZA metallic function for NH 3 adsorption is negligible compared to that of the zeolite, and accordingly, the total acidity of the catalyst is almost half that corresponding to its acid function (accounts for 50 wt %). The effect of attenuating the total acidity and average acid strength of the catalyst by treating the zeolite with NaOH is a direct

2. EXPERIMENTAL SECTION 2.1. Catalyst. The metallic function CuO−ZnO−Al2O3 (CZA), with an atomic ratio of Cu/Zn/Al = 4.5:4.5:1.0, has been prepared by coprecipitating the corresponding nitrates with Na2CO3 at pH 7.0 and 70 °C, and then calcining it at 325 °C for 3 h.37−39 The parent HZSM-5 zeolite (SiO2/Al2O3 = 30) has been supplied by Zeolyst International in ammonium form. This zeolite has been subjected to an alkaline treatment with a 0.4 M NaOH solution for 300 min at 80 °C, followed by fast cooling and ion exchange with ammonium nitrate of the resulting NaZSM-5 zeolite.34 The solid is then washed (distilled 3463

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In order to quantify the activity of the metallic function in the bifunctional catalyst, the effective conversion of methanol, XMeOH, has been determined in the second step of the SRD process (i.e., methanol steam reforming, eq 2) as

consequence of the effect observed in the acid function, i.e., a significant decrease in total acidity and a slight decrease in the acid strength with this treatment.34 The content of the coke deposited on the deactivated catalyst has been determined by temperature programmed oxidation (TPO) with air in a TA Instruments SDT 2960 thermobalance connected online to a Balzers Instruments Thermostar mass spectrometer. Prior to combustion, the deactivated catalyst is subjected to sweeping with N2 for 30 min in the reactor itself at the reaction temperature, in order to desorb the reaction products and homogenize the coke.42 The combustion procedure in the thermobalance is as follows: the sample is stabilized for 30 min with air at 75 °C and the temperature is subsequently increased (5 °C/min) to 550 or 600 °C and maintained at this value for 30 min to ensure complete coke combustion. Given that coke combustion is complete, the measurement of coke content has been carried out based on the CO2 signal in the spectrometer. 2.2. Reaction Equipment and Product Analysis. The kinetic runs have been carried out in automated reaction equipment provided with an isothermal fluidized bed reactor with a 22 mm internal diameter connected online to a MicroGC Agilent 3000 for product analysis, provided with four modules for the analysis of the following: (1) permanent gases (O2, N2, H2, CO, CO2, CH4); (2) oxygenates (MeOH, DME), light olefins (C2−C3), and water; (3) C2−C6 hydrocarbons; (4) C6− C12 hydrocarbons and oxygenate compounds. The hydrodynamic properties of the bed have been improved by mixing the catalyst (particle size between 150 and 250 μm) with an inert solid (CSi, with particle size between 60 and 90 μm) at a catalyst/inert ratio of 1/4. Prior to the catalytic runs, the catalyst was reduced in a stream of 5% H2 in He at 300 °C for 2 h with a total flow rate of 60 mL/min. The operating conditions are as follows: temperature between 225 and 325 °C; total pressure, 1.2 bar; partial pressure of DME between 0.08 and 0.25 bar; steam/DME molar ratio between 3 and 6; space time up to 2.2 gcatalyst h/gDME (with a catalyst mass between 0.6 and 4 g and DME mass flow rate between 2.06 and 6.72 gDME/h); He flow rate between 16 and 200 cm3(NC)/min; water flow rate between 0.047 and 0.135 mL/min (2.5−5 times the minimum fluidization velocity); time on stream, up to 24 h (48 h runs when space time is high and so deactivation is slow). The range of operating conditions has been established for attaining vigorous fluidization, while avoiding excessive bubbling and bed segregation. Furthermore, the conditions established correspond to severe deactivation, with the aim being to clearly determine the effect variables have on deactivation. 2.3. Reaction Indices. DME conversion has been calculated based on the flow rates at the reactor inlet (FDME,0) and outlet (FDME): XDME =

XMeOH =

rp,H2 =

Fiυi−1 FDME,0

(7)

FH2

(8) W The selectivity of each j component in the C1 lump (CO, CO2, and CH4) has been defined as the molar fraction of each component in this lump:

Sj =

Fj FCO + FCO2 + FCH4

(9)

3. RESULTS 3.1. Kinetic Behavior at Zero Time on Stream. 3.1.1. Effect of Temperature. Figure 1 shows the results at

Figure 1. Effect of temperature on the values at zero time on stream of DME conversion, MeOH effective conversion, and product (H2, CO2, and CO) mole fractions. Conditions: W/FDME,0 = 0.60 gcatalyst h/gDME, steam/DME = 3, and PDME = 0.25 bar.

zero time on stream of DME conversion, MeOH effective conversion, and product (H2, CO2, and CO) molar fractions at different temperatures. These results have been obtained by extrapolating the values of composition versus time obtained in long runs to zero time on stream. The DME hydrolysis reaction rate at 225 °C is very low, and DME conversion (0.22) is only slightly higher than that corresponding to thermodynamic equilibrium. As temperature is increased, the DME hydrolysis equilibrium shifts to the right due to methanol reforming on the metallic function. Consequently, DME conversion increases sharply in the 225− 275 °C range, and moderately above 275 °C, to a value of 0.87 at 325 °C. The effect increasing temperature has on MeOH effective conversion is largely insignificant due to the high MeOH reforming reaction rate at 225 °C. The H2 molar

(5)

The yields of H2 and CO have been calculated from the flow rates of these compounds (Fi) at the reactor outlet: Yi =

FMeOH,0

where FMeOH,0 is the methanol molar flow rate corresponding to DME conversion, which is calculated as twice the number of DME moles converted, according to the stoichiometry of DME hydrolysis (eq 1). The hydrogen production rate, rp,H2, has been defined as the moles of hydrogen formed by time and catalyst mass unit:

FDME,0 − FDME FDME,0

FMeOH,0 − FMeOH

(6)

where υi is the stoichiometric coefficient of i component produced in SRD (υ = 6 for H2 and υ = 2 for CO and CO2). 3464

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fraction is low below 250 °C, but it increases sharply as the temperature is increased, recording a value of 0.66 at 325 °C. Furthermore, as the temperature is increased, the CO concentration in the outlet stream increases. The molar fraction of CO at 225 °C is insignificant (0.001), and at 325 °C it is 0.033 (0.044 on a water-free basis). CO is produced through the reverse WGS reaction (eq 3),8,9 which takes place in series with methanol reforming (eq 2). On the basis of the aforementioned results, this reaction should be contemplated in the overall SRD reaction scheme, especially above 275 °C, and when H2 and CO2 concentrations in the reaction environment are high. 3.1.2. Effect of DME Partial Pressure. Figure 2 shows the effect of DME partial pressure on the results at zero time on Figure 3. Effect of steam/DME molar ratio on the values at zero time on stream of DME conversion, MeOH effective conversion, H2 yield, and CO2 selectivity based on C1 products. Conditions: 300 °C, W/ FDME,0 = 0.20 gcatalyst h/gDME, and PDME = 0.16 bar.

considerably from 0.0089 to 0.0039 on a water-free basis as the steam/DME ratio is increased from 3 to 6. Nevertheless, the MeOH effective conversion hardly changes because the methanol reforming reaction (eq 2) is almost full under these conditions. Although an increase in the steam/DME ratio increases DME conversion and H2 yield, these advantages should be considered together with the disadvantage of a higher cost incurred by steam generation and product separation, given that product molar fractions decrease due to dilution (Figure 4).

Figure 2. Effect of DME partial pressure on the values at zero time on stream of DME conversion, MeOH effective conversion, H2 yield, and CO2 selectivity based on C1 products. Conditions: 300 °C, W/FDME,0 = 0.30 gcatalyst h/gDME, and steam/DME = 3.

stream of DME conversion, MeOH effective conversion, H2 yield, and CO2 selectivity based on C1 compounds. The study has been carried out with different flow rates of inert gas (He), corresponding to a range of gas relative velocities (u/umf) in the fluidized bed between 2.5 and 5 and the same value for the remaining operating conditions. As observed, an increase in DME partial pressure slightly increases DME conversion and H2 yield. MeOH conversion does not change and, therefore, an increase in DME partial pressure increases the hydrolysis reaction rate, but it does not affect the methanol reforming reaction rate. Nevertheless, CO2 selectivity based on C1 compounds decreases slightly as the DME partial pressure is increased, given that higher concentrations of H2 and CO2 shift the WGS reaction to the left, i.e., it favors CO formation. Consequently, the molar fraction of CO increases from 0.0036 to 0.0055 on a water-free basis as the DME partial pressure is increased from 0.09 to 0.21 atm, respectively. 3.1.3. Effect of Steam/DME Molar Ratio in the Feed. The effect of the steam/DME molar ratio has been studied in the range between 3 (minimum according to stoichiometry) and 6 by carrying out runs using different values of water and inert gas (He) flow rates, with the remaining operating conditions being fixed. Figure 3 shows the results at zero time on stream of DME conversion, MeOH effective conversion, H2 yield and CO2 selectivity based on C1 compounds. An increase in the steam/DME ratio slightly increases DME conversion and H2 yield by shifting the DME hydrolysis equilibrium (eq 1). This increase is less significant for a steam/ DME ratio higher than 4. Furthermore, higher water content hinders the reverse WGS reaction by shifting its equilibrium to the left, which explains why the CO concentration decreases

Figure 4. Effect of steam/DME molar ratio on the values at zero time on stream of product molar fractions. Conditions: 300 °C, W/FDME,0 = 0.20 gcatalyst h/gDME, and PDME = 0.16 bar.

3.1.4. Effect of Space Time. The effect of space time has been determined by runs using different catalyst mass, with the remaining conditions being fixed. Figure 5 shows the results at 300 °C and zero time of DME conversion, MeOH effective conversion, H 2 yield, and CO 2 selectivity based on C1 compounds. DME and MeOH conversions and H2 yield increase sharply as space time is increased. Above W/FDME,0 = 2.2 gcatalyst h/gDME, DME and MeOH conversions and H2 yield asymptotically approach the values of 0.95, 1.0, and 0.94, respectively. Furthermore, an increase in space time also enhances the reverse WGS reaction (eq 3), thus increasing the CO yield. Thus, the CO concentration increases from 0.003 to 0.033% on a water-free basis for an increase in space time from 0.1 to 2.2 gcatalyst h/gDME. Therefore, CO2 selectivity decreases. A factor contributing to this increase in CO concentration is the increase 3465

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Table 2. Operating Conditions and Coke Content (Cc) in the Catalyst (For Time on Stream = 24 h) in the Runs for Determining Catalyst Stability

Figure 5. Effect of space time on the values at zero time on stream of DME conversion, MeOH effective conversion, H2 yield and CO2 selectivity based on C1 products. Conditions: 300 °C, steam/DME = 4, and PDME = 0.16 bar.

no.

T, °C

steam/DME molar ratio

PDME, bar

W/FDME,0, gcatalyst h/gDME

Cc, mg/g

1 2 3 4 5 6 7 8 9 10

275 300 325 325 300 300 300 300 300 300

3 3 3 6 4 4 4 4 4 4

0.16 0.16 0.16 0.13 0.08 0.18 0.25 0.25 0.25 0.25

0.13 0.13 0.13 0.13 0.08 0.08 0.08 0.20 0.60 0.60

1.0 1.8 4.8 1.8 1.8 2.1 2.2 1.3 0.2 0.3a

a

catalyst

For 48 h time on stream.

thermodynamic equilibrium) in order to clearly observe the decrease in catalyst activity with time on stream. Figure 6a shows the decrease in DME and methanol conversions with time on stream for the three temperatures studied, with this decrease being more pronounced at 325 °C. The decrease in DME and methanol conversions with time on stream evidences the deactivation of both the acid and the metallic functions, although the deactivation of the metallic function is more significant as the temperature is increased, which gives way to a considerable reduction in the conversion of methanol with time on stream at 325 °C. The decrease in the production of H2 (Figure 6b) is also more severe at 325 °C, given that it is reduced by half after approximately 24 h. Therefore, although hydrogen production at zero time on stream is highest at 325 °C, the average cumulative production for 24 h at this temperature is 395 mmolH2/(gcatalyst h), which is very similar to the average production corresponding to 300 °C and higher than that corresponding to 275 °C (256 mmolH2/(gcatalyst h)). The selectivity of CO2 based on C1 compounds (Figure 6c) at zero time on stream decreases as the temperature is increased, due to the enhancement of the reverse WGS reaction, as mentioned above for the results in Figure 1. This selectivity of CO2 is almost the same for 24 h at 275 °C, which evidences that catalyst deactivation at this low temperature has a similar effect on the reforming and reverse WGS reactions. Nevertheless, the selectivity of CO2 at 300 and 325 °C increases with time on stream (in a more pronounced way at 325 °C) because deactivation affects the reverse WGS reaction to a greater extent than the reforming reaction. The deactivation results in Figure 6 are consistent with the significant increase in coke deposition as the temperature is increased, with coke content values being 1.0, 1.8, and 4.8 mg/ gcatalyst, for 275, 300, and 325 °C, respectively (Table 2). The cause of deactivation has been studied from the TPO curves of the coke for determining the nature and location of the coke and by N2O chemisorption by pulses for quantifying possible Cu sintering. The TPO curves for the catalyst deactivated at different temperatures (Figure 7) show two mean peaks: (i) one at low temperature (∼260 °C), corresponding to the coke deposited on the metallic function, whose combustion is activated by the metal, and (ii) the other one at high temperature (410−450 °C range), corresponding to the coke deposited on the acid function.23,49,50 A third peak is also observed at an intermediate temperature (around 350 °C),

in CO2 concentration in the reaction medium due to a higher conversion.16 The aforementioned results and those in the literature evidence the good performance of catalysts made up of CuO− ZnO−Al2O3 and HZSM-5 zeolite for SRD when they are prepared using a suitable method and composition.14,34,37,38 The reaction indices are better than those reported by Matsumoto et al. using Cu/CeO2−H-mordenite catalysts, given that these authors obtained 86% DME conversion and a hydrogen production rate of 71 mmolH2/(gcatalyst h) at a temperature of 250 °C and using a very high value of space time (0.35 gcatalyst and a flow rate of DME/H2O/N2 = 2.3:7.7:16.0 mL/min (at 25 °C)).23 Studies in the literature deal with the improvement of the catalyst by modifying the CuO−ZnO−Al2O3 metallic function. Thus, Feng et al. incorporate ZrO2 and obtain full DME conversion and a H2 yield of more than 90% with a CO2 selectivity of around 95% in a fixed bed at 270 °C and using a space velocity lower than 2461 mL/(gcatalyst h).15,16 These results are open to improvement based on future innovations and show the potential interest of DME reforming for producing hydrogen to be used in fuel cells and as a raw material. 3.2. Catalyst Stability. The evolution of reaction indices with time on stream (up to 24 h) has been studied within operating condition ranges established for significant deactivation in order to clearly analyze the effect of these conditions. 3.2.1. Effect of Temperature. The formation of hydrocarbons from oxygenates (DME and MeOH) gives way to a significant decrease in H2 yield.32−34 Furthermore, the formation of hydrocarbons takes place presumably through the formation of methoxy ions, which are also intermediates in the formation of coke in DME synthesis,43 and in the transformation of oxygenates on the HZSM-5 zeolite through well-known mechanisms activated by the zeolite acid sites and, particularly, by the sites of higher acid strength.44−46 In the literature, the presence of methoxy ions as intermediates in DME reforming has been determined by spectroscopy and they have been attributed a coke precursor role.47,48 The effect of temperature on catalyst deactivation has been studied based on runs 1−3 in Table 2. It should be noted that the moderate acidity of the catalyst allows operating at up to 325 °C by avoiding hydrocarbon formation.34 The runs have been carried out with a low value of space time (far from 3466

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Al2O3 support of the metallic function, as has been established in the study of the coke deposited on CuO−ZnO−Al2O3/γAl2O3 catalysts used in the synthesis of DME.50 An increase in the reaction temperature enhances total coke deposition (Table 2), but mainly the coke deposited on the metallic sites (peak at 260 °C in Figure 7). This result explains the effect of temperature by decreasing methanol conversion with time on stream (Figure 6a). Furthermore, the mechanism of coke formation occurs presumably through methoxy ion intermediates, as proposed by Agarwal et al. for methanol steam reforming.51 In addition to directly enhancing the evolution of methoxy ions to form coke, the temperature increase also contributes to increasing the concentration of these ions by favoring methanol formation by DME hydrolysis. Therefore, methanol reforming is almost in equilibrium at 325 °C and methoxy ions are in excess for this reaction, thereby favoring coke formation. As a consequence of deactivation, the metal surface area of the catalyst deactivated at 325 °C (24 h) is 30 m2/gCu. This lower metal surface area compared to that of the fresh catalyst, 56 m2/gCu, is due to the fact coke deposition hinders N2O adsorption on the metal, which has been proven because the metal surface area corresponding to the fresh catalyst is recovered through catalyst regeneration by coke combustion at 300 °C with diluted O2. This result dismisses Cu sintering as a cause of deactivation under the operating conditions studied (with 325 °C as the ceiling reaction temperature, as it is the temperature used in catalyst calcination). The aforementioned results allow concluding that under the conditions studied coke deposition on the metal surface area is the main cause of catalyst deactivation, which attenuates its activity for reforming and WGS reactions, with Cu sintering being insignificant. The coke is formed presumably by degrading methoxy ions, which are intermediates in the transformation of oxygenates into hydrocarbons, and catalyst stability lasts longer below 300 °C, as methoxy ion formation is attenuated. 3.2.2. Effect of Steam/DME Molar Ratio. In addition to increasing the reforming reaction rate, the presence of water in the reaction medium contributes to attenuating coke deposition. Sierra et al. have studied this effect in the synthesis of DME on a CuO−ZnO−Al2O3/γ-Al2O3 catalyst and attributed it to the inhibition in the formation of methoxy ions, which are intermediates in the formation of hydrocarbons and coke.43 Steam is also important for attenuating coke deposition in the transformation of methanol and ethanol into hydrocarbons on HZSM-5 zeolite and SAPO catalysts.26,27,52−55 Nevertheless, excess steam in the SRD reaction medium may partially oxidize the metal active sites.8 Consequently, the effect of the steam/DME ratio on catalyst stability should be studied in order to delimit the conditions that minimize coke deposition without deteriorating the metallic function. Figure 8 shows the evolution with time on stream of DME and methanol conversions (Figure 8a) and H2 production (Figure 8b) for steam/DME ratio values of 3 and 6. The decrease in the conversions with time on stream is less pronounced for the steam/DME ratio of 6, which is consistent with the hypothesis whereby coke is formed through methoxy ions as intermediates and an increase in the amount of steam in the reaction medium contributes to inhibiting the formation of these ions. Furthermore, an increase in the formation of H2 contributes to attenuating coke deposition, given that coke

Figure 6. Effect of temperature on the evolution with time on stream of (a) DME conversion and MeOH effective conversion, (b) H2 production rate, and (c) CO2 selectivity based on C1 compounds. Conditions: W/FDME,0 = 0.13 gcatalyst h/gDME, steam/DME = 3, and PDME = 0.16 bar.

Figure 7. TPO profiles of coke combustion for catalyst deactivated at different temperatures. Conditions: W/FDME,0 = 0.13 gcatalyst h/gDME, steam/DME = 3, PDME = 0.16 bar, and time on stream, 24 h.

which should be attributed to the coke deposited at the interface between the metallic and acid functions, or on the 3467

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steam/DME ratio has on attenuating deactivation is more severe for methanol reforming than for DME hydrolysis. Furthermore, in order to prove there is no partial oxidation of Cu during reforming with a steam/DME ratio of 6, the deactivated catalyst has been treated with H2 and no water has been observed in the chromatographic analysis of the product stream. As Cu crystals are very stable and Cu does not undergo sintering, as proven above, as well as the fact the catalyst fully recovers its activity subsequent to combustion, evidence that the sole cause of catalyst deactivation is coke deposition. Consistent with the results in Figure 8, the coke content in the catalyst (determined by TPO) (Table 2) is higher for a steam/DME ratio of 3 (4.8 mg/gcatalyst) than for a ratio of 6 (1.8 mg/gcatalyst). A comparison of the TPO profiles corresponding to the catalysts deactivated in runs with different steam/DME ratios in the feed (Figure 9) reveals that an increase in the

Figure 8. Effect of steam/DME ratio on the evolution with time on stream of (a) DME conversion and MeOH effective conversion and (b) H2 production rate. Conditions: 325 °C, W/FDME,0 = 0.13 gcatalyst h/gDME, and PDME = 0.16 bar. Figure 9. TPO profiles for the catalyst deactivated under different values of steam/DME molar ratio. Reaction conditions: 325 °C, W/ FDME,0 = 0.13 gcatalyst h/gDME, and PDME = 0.16 bar.

formation takes place following mechanisms based on aromatic dehydrogenation and condensation.45,56 The higher DME and MeOH conversions obtained as the steam/DME ratio is increased (Figure 8a) evidence that catalyst deactivation is attenuated, although this attenuation of deactivation has presumably different effects on each step (DME hydrolysis and methanol reforming) of the global SRD process. The assessment of the effect that this attenuation of deactivation has on each step is a complex task, given that there is a synergy between the two steps; i.e., the decrease in DME conversion observed with time on stream is due to deactivation affecting not only the DME hydrolysis step (first), but also the methanol reforming step, given that the latter shifts the DME hydrolysis equilibrium. In spite of this synergy, the results in Figure 8a allow a quantitative comparison to be made of the effect that the attenuation of deactivation caused by the increase in the steam/DME ratio has on each step. Thus, for a steam/ DME ratio of 3, DME conversion decreases by approximately 40% over 24 h (from 0.58 to 0.35), whereas for a steam/DME ratio of 6 it decreases by 27% (from 0.63 to 0.46). Consequently, as the steam/DME ratio is doubled from 3 to 6, the decrease in the DME conversion of the overall SRD process attenuates by 13%. Furthermore, for a steam/DME ratio of 3, methanol conversion decreases by approximately 21% over 24 h (from 0.97 to 0.77), whereas for a steam/DME ratio of 6 it decreases by 11% (from 0.97 to 0.86), which means that there is a difference of 10% in the attenuation of deactivation when the steam/DME ratio is doubled from 3 to 6. The 3% difference between these two figures for attenuation (13 and 10%) may be attributed to the smaller influence of deactivation on DME hydrolysis. It is therefore concluded that the effect the

content of steam in the reaction medium attenuates the deposition of coke on the metallic sites to a greater extent (peaks at 260 °C), which is consistent with the more severe attenuation of the methanol reforming reaction than DME hydrolysis, as observed in Figure 8a. On the basis of these results, it is concluded that an increase in the steam/DME ratio favors catalyst stability in the SRD reaction. In light of the results obtained here on the effect of the steam/DME molar ratio on the values at zero time on stream of H2 yield (Figure 3), product concentration (Figure 4), and attenuation of deactivation by coke deposition (Figure 8), a steam/DME ratio in the 4−5 range is established as suitable, given that above this value the energy cost is likely to be excessive. It is noteworthy that a fluidized bed reactor contributes to attenuating coke formation due to uniform catalyst activity, so the effect of attenuating coke formation by steam is therefore more efficient than in a fixed bed, as has been proven in the transformation of methanol into hydrocarbons.52 3.2.3. Effect of DME Partial Pressure. Figure 10 shows the evolution of DME conversion, MeOH effective conversion (Figure 10a), and H2 production (Figure 10b) with time on stream for three values of DME partial pressure. An increase in the DME partial pressure increases DME conversion and H2 production at zero time on stream, whereas MeOH effective conversion is maintained unaltered. These results are consistent with those in Figure 2, for a higher value of space time (0.30 gcatalyst h/gDME). 3468

dx.doi.org/10.1021/ie402509c | Ind. Eng. Chem. Res. 2014, 53, 3462−3471

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Figure 10. Effect of DME partial pressure on the evolution with time on stream of (a) DME conversion and MeOH effective conversion and (b) H2 production rate. Conditions: 300 °C, W/FDME,0 = 0.08 gcatalyst h/gDME, and steam/DME = 4.

Furthermore, Figure 10 shows there is no significant effect of DME partial pressure on catalyst stability and the decrease in the reaction indices with time on stream is similar for the three values of DME partial pressure. Moreover, the coke contents for the catalysts deactivated under the different DME partial pressures are similar, around 2 mg/gcatalyst (Table 2), although there is a very slight increase in the content as the DME partial pressure is increased. The fact that the DME partial pressure does not have a significant effect on coke deposition is due to two opposing effects of this variable. On the one hand, as the DME partial pressure is increased the hydrolysis reaction rate increases significantly, without affecting the reforming reaction. Consequently, the methanol concentration in the medium increases, and therefore the concentration of coke precursor methoxy ions also increases. On the other hand, H2 production increases, which inhibits the dehydrogenation reactions involved in coke formation. It should be noted that the difference in the values of DME partial pressure studied is small, given that this variable is restricted in order to maintain steam/DME constant and due to the need of operating with a gas flow rate that ensures a suitable hydrodynamic performance in the catalytic bed (u/umf relative velocity between 2.5 and 5). Consequently, the difference in methoxy ion concentration in the catalyst is limited under the conditions studied. 3.2.4. Effect of Space Time. An increase in space time significantly attenuates the decrease with time on stream of DME and MeOH conversions, with both conversions (Figure 11a) and H2 yield (Figure 11b) being almost constant for 24 h for a value of 0.60 gcatalyst h/gDME. Furthermore, CO2 selectivity (Figure 11c) remains constant for low values of space time (0.075 gcatalyst h/gDME) and increases with time on stream for

Figure 11. Effect of space time on the evolution with time on stream of (a) DME conversion and MeOH effective conversion, (b) H2 yield, and (c) CO2 selectivity based on C1 compounds. Conditions: 300 °C, steam/DME = 4, and PDME = 0.25 bar.

higher values of this variable. This result evidences that the deactivation by coke deposition has a greater impact on the reverse WGS reaction than the reforming reaction, which is consistent with the results in section 3.2.1 corresponding to the study of the effect of temperature on catalyst stability. Although the results are not shown, the catalyst is stable for long duration experiments (48 h) with a space time of 0.60 gcatalyst h/gDME. Clearly, as space time is increased deactivation becomes patent for higher values of time on stream. The aforementioned results concerning the effect of space time on deactivation are consistent with the coke content in the catalyst, which sharply decreases as space time is increased (Table 2). This result is in agreement with the hypothesis that the coke precursors are the methoxy ions formed from the oxygenates (DME and MeOH). Thus, the higher concentrations of DME and MeOH in the reaction medium for a low value of space time favors the formation of coke precursor methoxy ions. 3469

dx.doi.org/10.1021/ie402509c | Ind. Eng. Chem. Res. 2014, 53, 3462−3471

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4. CONCLUSIONS An increase in temperature (between 275 and 325 °C), space time (between 0.075 and 2.2 gcatalyst h/gDME), steam/DME ratio in the feed (between 3 and 6), and, to a lesser extent, DME partial pressure in the feed (between 0.08 and 0.25 bar) gives way to a significant increase in the conversion and in the yield of H2 in the reforming of DME on the catalysts studied, although CO yield also increases by enhancing the reverse WGS reaction. The problems related to the stability of the catalyst are due to coke deposition, given there is no Cu sintering under the conditions studied (