M Ratio Effects in Methanol Steam

Article pubs.acs.org/IECR

Cite This: Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

Pressure, Diffusion, and S/M Ratio Effects in Methanol Steam Reforming Kinetics Rajesh Thattarathody,† Moris Artoul,† Rafael M. Digilov,† and Moshe Sheintuch*,†,‡ †

Department of Chemical Engineering, Technion-Israel Institute of Technology, Haifa, Israel 32000 College of Engineering, Guangdong Technion Israel Institute of Technology (GTIIT), 243 Da Xue Road, Shantou, Guangdong, P.R. China

‡

ABSTRACT: The kinetics and diffusion effects of methanol steam reforming reaction over commercial Cu/ZnO/Al2O3 catalyst was studied for steam to methanol (S/M) ratios of 0 to 1 and pressures below 6 bar. Our objective is the development of a novel high-pressure propulsion technology based on the concepts of thermochemical recuperation (TCR) and onboard hydrogen production. A simple kinetic model assuming methanol decomposition followed by water−gasshift was used to estimate the rate constants (kMD, kWGS). The apparent activation energy of kMD was estimated as 45−55 kJ/ mol for large pellets and S/M = 1.0, 0.5, and 0.0; kMD for S/M = 0 (and the conversions obtained) were smaller than those of S/M = 1, probably due to CO inhibition. At temperatures above ∼500 K, the WGS is at equilibrium. Strong pore-diffusion limitations are evident at 1 bar for the 3 mm catalyst, evident experimentally as well as by analysis; the apparent kMD is almost diffusion free for particles of 0.7 mm in diameter. This resistance increases, of course, with P. The selectivity of CO (dry basis) increases with W/Fmeth (weight of the catalyst divided by flow rate of methanol in units of kgcat s mmol−1) and increases with decreasing particle size. As the pressure increases, the ratio of CO and CO2 increases moderately up to 4 bar; at 6 bar the change is drastic. Similar observations were made with S/M = 0.5 and 1.0. Deactivation rates and coke formation were also studied and were found to be marginal under atmospheric pressures over a period of 10 h and became evident only at S/M = 0 and 275 °C; at 6 bar the decline was evident already at S/M = 0.5. The source of deactivation was attributed to coking, a conclusion based on TPO of spent catalyst. In the case of S/M = 1 and 0.5, and at 6 bar, there was a shift in the composition from CO2 to CO with time over a period of 10 h at 275 °C. However, the carbon deposition in all cases was estimated to be about the same.

1. INTRODUCTION The internal combustion engine (ICE) is the main power plant in most modern transportation systems. It is well-known that about 30% of the energy introduced to the ICE with the fuel is wasted with the hot exhaust gases. Utilizing part of this energy, also known as waste heat recovery (WHR), can lead to a significant increase in the overall ICE efficiency. One promising approach for that purpose is to use the waste heat to sustain endothermic reactions of fuel reforming, a method referred to as thermochemical recuperation (TCR).1−4 TCR has two main advantages, over other strategies like turbocharging: it provides better improvement in efficiency, and the high hydrogen content in the reforming products (reformate) usually leads to increased flame velocity, higher octane number, wider flammability limits, and reduced combustion irreversibility.5 That results, in turn, in reduction of CO, NOx, and volatile hydrocarbons emissions. This work is a part of a research aimed at the development of a novel high-pressure propulsion technology based on the concepts of waste heat recovery and onboard hydrogen production. Methanol as a hydrogen vector enables fuel distribution in the liquid phase using available infrastructure, thus overcoming the © XXXX American Chemical Society

safety and capacity problems associated with pure hydrogen distribution.6,7 It is easily converted to hydrogen at relatively lower temperatures by decomposition or by steam reforming. Methanol can be considered as a renewable energy source as it can be produced in large quantities from syngas, natural gas, or biomass either directly by fermentation or indirectly by gasification. The methanol steam reforming products composing hydrogen-enriched gases can be directly fed to the internal combustion engine (ICE) instead of using the fuel cell, thereby eliminating the already known problems of on-board H2 storage and reducing expenses. Unlike fuel cell, ICE can efficiently burn different mixtures of hydrogen, carbon monoxide, and other gases. Methanol steam reforming for TCR has been vastly studied in the past few years,8 and prototypes of on board reforming based on methanol have been demonstrated by Toyota. While its benefits have been described above, we should note some of the Received: December 5, 2017 Revised: February 4, 2018 Accepted: February 7, 2018

A

DOI: 10.1021/acs.iecr.7b05033 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Industrial & Engineering Chemistry Research main drawbacks of liquid methanol as an automotive fuel: these include relatively low heating value, poor engine startability at low ambient temperatures, and higher aldehyde emissions. The kinetics of catalytic methanol SR was also a subject of intensive research but mainly at atmospheric pressure (i.e., for FC applications). Here we study the kinetics at moderate pressures. The overall steam reforming reaction (eqs 1−3), comprises methanol decomposition (MD) and water gas shift (WGS) described by ΔH = 49.7 kJ mol−1

CH3OH + H 2O ⇌ CO2 + 3H 2

CH3OH ⇌ CO + 2H 2

ΔH = 90.2 kJ mol−1

CO + H 2O ⇌ CO2 + H 2

ΔH = − 41.2 kJ mol−1

rMSR =

(dehydration)

2CH3OH ⇌ CH4 + 2H 2 + CO2 2CH3OH ⇌ HCO2 CH3 + 2H 2

CH3OH ⇌ CH 2O + H 2

(methanation)

(1) (2) (3)

rMSR =

rWGS =

(5) (6)

(formaldehyde synthesis)

(7)

KM K H2

⎞⎛ ⎟⎜ ⎠⎝

⎞⎛ ⎟⎜ ⎠⎝

⎞⎤ ⎟⎥ H2 ⎠⎦

pM p

⎞ ⎞ ⎟ + √K H2√pH ⎟2 p ⎠ 2 H2 ⎠

(Cs T)2

pM

⎛ kMSR K CH3O*(1)⎜ ⎝

⎞⎛ ⎟⎜1 − pH ⎠⎝ 2

pM

(9)

rMD =

pH 3 pCO

⎞

⎟C kMSR pM pW ⎠ s1 2

2

T

Cs1a T

Den1(1 + KH(1a)1/2pH21/2 )

⎛ ⎛ p ⎜ Den1 = ⎜1 + K CH3O*(1)⎜ M ⎜ p ⎝ H2 ⎝ ⎞⎞ ⎛ pW ⎟⎟ ⎜ * + K OH (1) ⎜ p ⎟⎟ ⎝ H2 ⎠⎠

(4)

(dehydrogenation)

⎛ ⎛ ⎜1 + ⎜ ⎝ ⎝

KM K H2

where CsT represents total surface concentration of site (mol m−2). Peppley et al.,15 who tested only the SR case, developed expressions for the reaction rates of all the three possible reactions (eqs 1−3) mainly based on extensive investigations on methanol synthesis. They assume that active sites for the decomposition reaction are distinct from those for the reforming and the WGS reaction, and that the rate-determining step for both the MSR and the MD is the dehydrogenation of adsorbed methoxy groups. The suggested rate expressions are,

Some side reactions (eqs 4−8) that occur, especially with low steam to methanol (S/M) ratio, can lead to the formation of small amounts of dimethyl ether (eq 4), methyl formate (eq 6), and methane (eq 5), etc. At S/M ∼ 1, a mixture of CO2 and H2 is almost exclusively produced. But methanol reforming to CO2 and H2 is rather undesirable for ICE feed purposes compared to methanol decomposition because CO2 is a diluent gas and is not an energy carrier. However, the absence of steam in the case of methanol decomposition results in some organic side products which, in turn, lead to a larger extent of coking compared to reforming. 2CH3OH ⇌ (CH3)2 O + H 2O

⎡ ⎛ ⎢k CH3O⎜ ⎝ ⎣

(10)

⎞ ⎟+K * HCOO (1)pCO2 pH2 ⎟ ⎠

⎛ p p ⎞⎛ k WGS*K OH*(1)⎜ COp W ⎟⎜1 − ⎝ H2 ⎠⎝

pH pCO ⎞ 2 2 2 ⎟ C T k WGSpCO pW ⎠ s1

(Den1)2 ⎛ kMDK CH3O*(2)⎜ ⎝

(11)

⎛ pH 2 p ⎞⎞ ⎞⎛ ⎜⎜1 − ⎜ 2 CO ⎟⎟⎟Cs2 TCs2a T ⎟ pH ⎠ ⎝ kMDpM ⎠⎠ 2 ⎝

pM

Den3 (12)

CH3OH ⇌ C + H 2 + H 2O

(carbonation)

(8)

⎛ ⎛ p ⎜ Den3 = ⎜1 + K CH3O*(2)⎜ M ⎜ p ⎝ H2 ⎝

The highly active copper-based catalysts are preferred for methanol steam reforming over groups VIII−X catalysts, which show better stability, due to economic reasons.9 Nevertheless, a number of disadvantages are reported for Cu-based catalysts such as pyrophoricity, deactivation by thermal sintering, coke deposition, or change in oxidation state.10−12 The Cu crystallites readily undergo sintering above 300 °C.13 The formation of several organic byproducts as listed above leads to the formation of coke on catalyst surface via pyrolysis. The formation of coke by CO disproportionation reaction (Boudouard reaction) is less favorable in the case of reforming because of the higher CO2/CO ratio compared to decomposition. The changes in oxidation state from Cu(0) to Cu(I) is also reported to affect the activity or selectivity. Most of the studies over Cu/ZnO/Al2O3 were conducted at S/ M = 1 atmospheric pressure, and ignored diffusion resistance effects and correlated the experimental data with the rate equation derived for a single overall SR reaction. In, studies that refer to two reactions there was disagreement on the order: whether SR followed by r-WGS or MD followed by WGS. Jiang et al.,14 after analyzing the reaction mechanism of MSR expressed the reaction rate as the following expression, assuming that the r-WGS reaction is negligible,

⎛ ⎞ ⎟ + K *(2)⎜ pW OH ⎜ p ⎟ ⎝ H2 ⎠

⎞⎞ ⎟⎟ ⎟⎟ ⎠⎠

× (1 + KH(2a)1/2pH 1/2 ) 2

where Ki*(1) = Ki(1)/KH(1a)1/2, [Ki(1) stands for the adsorption coefficient for surface species, i adsorbed on type 1 site], Cs1T represents total surface concentration of type 1 active sites, which catalyze the elementary surface processes involved in the mechanisms for the MSR and WGS, and Cs1aT is that of H2 adsorbing type. Similarly, Cs2T are that of type 2 sites which catalyze surface processes for MD and Cs2aT that of H2 adsorbing type for MD. The values of kMSR, kMD, or kWGS rate constants were based on published data on the heats of adsorption and are not included in the nonlinear regression and kW* = kWKCO(1). Geissler et al.16 developed the following rate expression for MSR on commercial Cu/ZnO/Al2O3 catalyst. ⎡E ⎛ 1 1 ⎟⎞⎤ rMSR (T) = k(513K) exp⎢ a ⎜ − ⎥ ⎝ ⎣R T 513K ⎠⎦ ⎞ ⎛ pH 3 pCO 2 2 ⎟ × pM 0.4 ⎜⎜1 − p ⎟ K (T) p p MSR M W⎠ ⎝ B

(13)


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Industrial & Engineering Chemistry Research Purnama et al.17 for the same catalyst used a reaction scheme involving direct SR reaction followed by rWGS reaction and derived and fitted the results with, rMSR = k1pM m pW n ;

m = 0.6, n = 0.4

rrWGS = k 2pCO PH2 − k −2pW pCO 2

2. EXPERIMENTAL SETUP AND PROCEDURES All tests were conducted with a CuO/ZnO/Al2O3 commercial catalyst (Clariant G-66A ), either with a standard 3 × 3 mm cylindrical pellets or with crushed and sieved powders to evaluate mass transfer limitations. Two types of experimental systems were employed and are described as follows. (i) A laboratory fixed bed reactor, aimed at obtaining kinetic data as well as coking patterns. (ii) Temperature-programmed experiments aimed at the understanding and characterization of the used and spent catalysts as well as studying the nature and behavior of carbon deposits that occur during methanol steam reforming and decomposition. 2.1. Catalytic Testing. Experiments were conducted either at atmospheric pressure or with a system that allows one to vary the pressure; two different fixed bed reactors were employed in the two tasks, as described below, but the feeding and analysis systems were identical. A liquid reactants solution of desired composition stored in a sealed mixed reservoir was fed to a vaporizer by means of a calibrated peristaltic pump. The vaporizer temperature is monitored by a K-type 1/8″ thermocouple and maintained at reaction temperature by a band-type heating element controlled via a PID temperature controller. In the high pressure version, the feed from the reservoir was sent to the reactor using Grundfos DDA pump through a vaporizer. The vaporizer was heated to 150 °C and the line between vaporizer and reactor was heated at 100 °C with the help of heating tapes. The conventional isothermal fixed-bed reactor used for atmospheric pressure was made of stainless steel of nominal i.d of 1.0 cm and of 6.0 cm total length where two stainless steel 10 mesh sieves were used to support the catalyst bed. It was loaded with 1−3 g catalysts. A K-type 1/8″ thermocouple protruded through a collar in the sieve was fixed at the center of the catalyst bed; the reading of this thermocouple is monitored, logged, and used as the reaction temperature for all analyses. The reaction section of the tube is placed in an aluminum heating block made of four 125 W cartridge heaters controlled by another PID temperature controller unit. The gaseous stream exiting the reactor is then passed through a water-cooled condenser and then through a drain filter trap where unreacted methanol and water are separated from gaseous products. The CO and CO2 concentrations are continuously monitored by IR analysers (Edinburgh instruments), the gaseous products are further analyzed by means of a gas chromatograph equipped with a thermal conductivty detector (Thermo Scintific) packed with two in-line coloumns (QS-bond and Shinecarbon) that was calibrated with different mixtures of all components. All measurements were recorded at least 20 min after reaching a thermal steady state. High pressure tests were carried in a vertical, down flow, 0.8 cm in diameter and 14 cm long fixed bed reactor. The catalyst was placed between plugs of quartz wool (bed length 1.1 cm), and the regions below and above the catalyst bed were packed with glass beads. The reactor was heated externally through a heating block. It was loaded with ∼0.7 g crushed catalysts with a particle size range of 1−2 mm. Pressure control was achieved by incorporating a back pressure regulator after the liquid condenser and gauges before and after the reactor. The conversion was determined by weighing the condensate. At the time of condensate collection, initially the flow was stopped by a valve placed in between the reactor and the condenser and then liquid was collected by slightly opening the valve attached to the liquid

(14) (15)

The apparent activation energies obtained were 76 kJ mol−1 for the SR and 108 kJ mol−1 for the rWGS reactions. A major issue for the feasibility and practicality of an on board catalytic methanol reformer is that of catalyst stability. The catalyst selected must ensure high activity and stability of SR and WGS reactions, at temperatures lower than those of the exhaust gas. Purnama et al.17 studied the stability of commercial CuO/ ZnO/Al2O3 catalyst (Süd-Chemie) over 320 h on stream (T = 250 °C, S/M = 1, W/Fmeth = 0.007 kgcat s mmol−1; W/Fmeth = weight of the catalyst divided by flow rate of methanol in units of kgcat s mmol−1) and observed slight decrease in the conversion for the initial 100 h followed by stable activity. Choi et al.12 found on the same catalyst that with S/M = 0.5 and at 250 °C (GHSV = 1100 h−1) deactivation is less evident within the first 100 h on stream than with methanol alone. In the absence of water, methanol conversion declines rapidly during the initial period and reached steady state after 20 h (T = 300 °C, GHSV = 1100 h−1). Studies indicated that most (copper-based) commercial catalysts are prone to some degree of deactivation mainly due to coke deposition during methanol decomposition, significantly shortening catalyst lifetime. The deactivation of catalysts by coking is a very complex process where mechanism and kinetics are still not completely understood. Much of the prior work on the modeling of catalyst deactivation is based on empirical correlations in which the activity of the catalyst is described only as a function of time. Although this empirical quantification of catalyst coking is very useful, its possibilities to predict catalyst activity with time on stream at different reaction conditions are very limited. In this study, catalytic tests and temperatureprogrammed experiments were carried out to measure the coke formed during the reforming reaction. A model was proposed to predict the coke formation and its effect on catalyst activity. Amphlett et al.18 examined the thermodynamic MSR equilibria for pressures up to 30 bar and S/M of 1.5−0.67 assuming that only reactions 1−3 exist; methanol conversion is ∼98% at 10 bar and at 250 °C. The present work follows our previous study19 using the same catalyst but limited to atmospheric pressure with powdered catalyst while using Micromeritics’ Autochem 2920 for temperature-programmed procedures and a mass spectroscopic analysis. That work showed, that aside from H2, CO, and CO2, only methyl formate and dimethyl ether are formed and that their production declined with increasing S/M and vanished at S/M = 1. Isothermal rate oscillations were observed for the first time for this reaction, with a period of 10 min in order of magnitude. This behavior will not be detected, of course, in the present packed bed study. A simple kinetic model was developed for methanol decomposition (MD) using S/M = 0 data, and the rate and activation energies were determined. Some of these results are compared with the present results. The structure of this paper is the following: after describing experimental procedures, we describe results under atmospheric pressure in sections 3.1 and 3.2 (where diffusion resistance effects are shown to be important), while the effect of pressure and of S/ M is described in section 3.3 and stability issues are discussed in section 3.4. C


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Industrial & Engineering Chemistry Research

Figure 1. Methanol conversion and H2 yield (left) and dry product composition (right) vs temperature, [S/M] = 1, W/Fmeth = 0.025 kgcat s mmol−1, and P = 1 bar. Also shown is expected CO concentration under the assumption of WGS equilibrium (line).

Figure 2. Arrhenius plot (left, kMD in cc g−1s−1) and the distance from WGS equilibrium expressed as ln[(CCO2CH2)/(CCOCW)] (right) vs 1/T plot.

which are less accurate than the GC H2 composition measurement. CO selectivity drastically increases at temperatures higher than 260 °C where CO formation is favorable, as predicted from equilibrium. To determine the kinetic parameters, we view the network as composed of two reactions: methanol decomposition (MD), which is practically irreversible at atmospheric pressure and assumed to be first order in methanol concentration (CM), followed by WGS. To determine the MD kinetic parameters, we write the reactor balance assuming a PFR model, where f is the conversion and δ is the relative change in volumetric flow due to reaction

collector. High pressure studies with S/M = 1 were conducted with a total liquid flow rate of 6 mL/h (W/Fmeth of 0.0246 kg s mmol−1) and those with S/M = 0.5 and a total liquid flow rate of 5 mL/h (W/Fmeth of 0.025). The temperature of the catalyst bed was continuously monitored with the help of a K-type 1/8″ thermocouple. Reduction has been carried out at 300 °C under H2 (50 mL/min) for 1 h prior to each experiment. The temperature was then varied to the desired one. Temperature-programmed ramping (reaction, oxidation, or desorption) were conducted with Autochem II 2920.

3. EXPERIMENTAL RESULTS 3.1. Atmospheric Pressure. Experiments were conducted with a 3 g of catalyst, with a stream of 1:1 steam to methanol [S/ M] molar ratio within the temperature range 200−300 °C, showed that methanol conversion is almost complete at 280 °C (Figure 1), hydrogen partial pressure (dry basis) is in the range 0.72−0.75, CO concentration is strongly affected by temperature and can be well-predicted by WGS equilibrium. Two values are used to quantify the catalyst performance (Figure 1): hydrogen yield defined as moles of hydrogen obtained per 1/3 mol methanol fed and carbon monoxide selectivity defined as the carbon monoxide percentage of total carbon species produced. SCO =

[CO] [CO] + [CO2 ]

hydrogen yield =

H2 3MeOH

F(CM, in)

k C df = MD M = kMDC M , in (1 − f)/(1 + δ f) dW D (17)

yielding kMDW = F

∫0

f

(1 + f )df = −2 ln(1 − f ) − f 1−f

(18)

where F is feed flow rate and CM,in is methanol feed concentration. The dilution by products of the MD reaction (M → CO + 2H2), for this case where the methanol is diluted with steam in equal parts, is described by δ = 1. D is an inhibition term due to adsorption of other components, but for now we assume D = 1. Arrhenius plot (Figure 2, left) showed the apparent activation energy to be 45 kJ/mol, compared with values of 75−100 kJ/ mol15,16,20,21 in the literature suggesting strong diffusional limitations in the 3 mm pellets, as will be verified in the next subsection. Under strong diffusion limitations, the apparent activation energy is half the true one. Figure 2, right, presents the measured distance from WGS equilibrium, by plotting calculated

(16)

Hydrogen yield approaches its higher stoichiometric limit (3 mol of hydrogen per mole methanol) as the temperature increases and conversion reaches unity. The hydrogen yield appears to be higher than methanol conversion, but the small difference is within the experimental error. Recall that conversion is based on measurements of condensed methanol in exit stream, D


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The kMD value is determined from the data casted according to linear relation of a first-order reaction (eq 18) while using only the low conversion points; the kMD values are highly sensitive to the high conversion data. An important measure of the reforming process is the change in energy, expressed by the lower heating value (LHV) of the products relative to methanol molar LHV. For each composition there is a critical value of conversion after which the gaseous stream (i.e., not containing unreacted methanol) has a higher LHV in comparison with the combustion of liquid methanol. The increase in LHV (∼8% above methanol) at S/M = 1 is due to the higher hydrogen content in the product, which is more energetic than CO. 3.2. Diffusion Limitations. The purpose of this section is to find the optimal catalyst size that will yield a good effectiveness factor with a low pressure drop. The commercial cylindrical pellets of the 3 × 3 mm used in previous experiments were crushed and sieved into two different powders: powder no. 1: 0.9−1.1 mm; catalyst loading in reactor: 5 g. Powder no. 2: 0.6− 0.8 mm; catalyst loading in reactor 1.65 g. Catalytic testing and activity measurements were performed with the same experimental procedure conducted with the pelletized catalyst. The conversion data versus W/Fmeth, for the three catalyst sizes (Figure 5), already indicates strong porediffusion limitations. Quantifying this trend by estimating kMD, while ignoring the high conversion points, shows indeed a factor of almost 3 between the kMD on pellet and powder 2. That still does not imply that the latter is free of diffusion resistance, and we pursue this point below. Figure 5 (right) shows that the CO selectivity (dry basis) increases with W/Fmeth and increases with decreasing particle size. The difference between the two powders is relatively small. The experimental results are compared with predictions based on the assumption of instantaneous equilibrium at the measured conversion, showing good agreement. The instantaneous equilibrium is the solution of eq 19,

KWGS,eq using measured concentrations, showing that at temperatures above ∼500 K the WGS is at equilibrium while at lower temperatures the distance becomes appreciable. Note that the CCO is above its equilibrium value supporting the reaction mechanism composed of methanol decomposition followed by WGS. Comparison of the rate constants (kMD, Figure 3) obtained here with the 3 × 3 cylindrical catalyst and S/M = 1 with those

Figure 3. Comparison of kMD (in cc g−1 s−1) calculated from conversion in Figure 1 (S/M = 1; 3 × 3 mm catalyst) with values obtained with S/M = 0 in an Autochem in a recent work (crushed catalyst pellets, 1−2 mm19). W/Fmeth in the legend is given in kgcat s mmol−1.

obtained in the Autochem with crushed catalytic pellets of 1−2 mm in size and in the absence of steam, but diluted in Ar,19 show the former to be more active by a factor of ∼3, but the apparent activation energy is similar. The difference in activity is reasonable for the difference in operating conditions of the two systems. A kinetic study at constant reaction temperature (270 °C) was conducted by varying volumetric feed rate, showing (Figure 4, S/ M = 1) that as the ratio W/Fmeth increases, conversion approaches 100%, in agreement with the irreversibility of MD at atmospheric pressure, and hydrogen content in the product approaches the stoichiometric 3/4 (Figure 4). For W/Fmeth values exceeding 0.024 (kgcat s mmol−1), corresponding to a residence time of 0.65 s, conversion exceeds 90%. This residence time will be even shorter for smaller pellets or catalyst wash coat. Moreover as conversion increases, hydrogen yields approach their upper limit of 3 mol/mol M, and CO increases steeply with residence time as shown in Figure 4 (left). Using this value to predict f(W/F) shows a reasonable agreement (Figure 4).

R = [x(SM(1 + δf ) − f + x) − (3f − x)(f − x)/KWGS] = 0

(19)

where x = CCO/CM,in, SM = S/M (=1 in this case). The figure also shows the ultimate (large W/F or f = 1) equilibrium. The Thiele modulus is defined as ϕ=

k1ρ R Deff q

(20)

Figure 4. Methanol conversion, hydrogen yield, CO selectivity and predicted conversion by a 1st-order kinetics (left) and dry product composition (right) vs W/Fmeth ([S/M] = 1, T = 270 °C). Product compositions of CO and CO2 are predicted well by WGS equilibrium (not shown). E


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Figure 5. Catalyst size effects: (left) methanol conversion (points) and fitted conversion by 1st-order kinetics (kMD, line), and (right) experimental (points) CO selectivity vs W/Fmeth for powdered and pelleted catalysts; T = 280 °C, S/M = 1; solid lines of CO selectivity represent instantaneous equilibrium.

Figure 6. (left) Plot of kMD (determined as η(ϕ)kMD,true in units of cc g−1 s−1) for the pellet and powdered catalysts vs particle radius at 280 °C. Also shown are the kMD values of the three catalysts in Figure 5. (Right) calculated pressure drop in a fixed bed 40 cm long and 1.5 cm in diameter at design flow rate.

R is the particle radius, ρ is catalyst density, q is a correction for the catalyst particle geometry (q = 3 for spheres, q = 2 for infinitely long cylinders, while the hydraulic radius for finite cylinder is Rh = πR2L/(2πRL+2πR2) and for this pellet with L = 2R, q = 3, and Deff is the effective diffusivity of methanol in CO2.

Deff = Dm

ε τ

where L is the length of the bed, ρ is the density of fluid, μ is the dynamic viscosity, Vs is the superficial velocity, εb is bed void fraction, and DP is the particle diameter. Figure 6 compares the reduced apparent rate with increase pellet size versus the decline in pressure drop, indicating that particle sizes should be limited to 1−2 mm. 3.3. High Pressures and Methanol-Rich Feed. Recall that results in this section were obtained with crushed catalysts (1−2 mm) with a different reactor than that of section 3.1. Methanol conversion under atmospheric pressure and high pressures was determined by weighing the condensate product. 3.3.1. Pressure Effects (S/M = 1). The pressure-dependence of the dry product composition at a temperature of 250 °C (Figure 7) shows almost a steady 75% H2 with little CO; the CO concentration increases substantially upon change from 4 to 6 bar, accompanied by a reduction in both CO2 and H2. The conversion shows a nonmonotonic behavior, which may be due to experimental fluctuations. The limiting equilibrium conversion at 6 bar was calculated to be 97%. The effect of increasing temperature on product composition was studied at 6 bar pressure (Figure 8), showing increase in conversion and CO concentration and decline in CO2. As temperature and conversion are increased, several counteracting effects occur: water concentration declines pushing the WGS equilibrium toward CO while the higher temperature shifts it in the opposite direction. 3.3.2. S/M Effects. Conversions with S/M = 0.5 and atmospheric pressure (Figure 9, left) are similar to those with stoichiometric feed (Figure 1). Kinetic analysis to determine its temperature dependence shows a line almost overlapping that of S/M = 1 (Figure 10, right). Product distribution is similar to what

(21)

Dm is the diffusion coefficient of methanol (taken here to be in CO2, Dm = 0.117 cm2/s), ε and τ are the catalyst porosity (estimated as ε ∼ 0.5) and tortuosity (estimated by the approximation τ = 1/ε). The Thiele modulus is then used to estimate effectiveness factor η, the ratio of apparent reaction rate to that in the absence of diffusion resistance (i.e., with sufficiently small particles), η = (tanh ϕ)/ϕ

(22)

To find the true diffusion-free rate constant, kMD,true, let us use the powder 2 value, kMD = 8.47 (cc g−1 s−1 units), catalyst density ρ = 2.12 g/cm3, catalyst effective radius R = 0.035 cm, and q = 3 to obtain ϕ = 0.29. The true value is not far from the apparent one, kMD,true = 8.5 cc g−1 s−1. The effective reaction constant dependence on catalyst size is shown in Figure 6. To find the optimal catalyst size we should compare the apparent rate decline with increasing size to the decrease of pressure drop (ΔP). To that end, we apply the Ergun equation ΔP =

150μ(1 − εb)VsL εb2Dp2

+

1.75(1 − ε)ρVs2L εb2Dp

(23) F


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depleted. The inhibition is strong at S/M = 0, some CO2 is produced, probably by Boudouard reaction. As suggested, we model this behavior by rMD = kMDCM/(1 + KCOPCO). To estimate KCO, we use the S/M = 0 data, for which PCO = Pf/3, and using kMD values at 250 °C (Figure 10, line of S/ M = 1 or 1/2) we can integrate f

kMDW = F

∫

(

(1 + δf ) 1 +

K COPf 3

)df

1−f

0

⎡ K P⎤ = [−(1 + δ)ln(1 − f ) − δf ]⎢1 + CO ⎥ ⎣ 3 ⎦ K P⎛ δf 2 ⎞ ⎟ − CO ⎜f + 3 ⎝ 2 ⎠

Figure 7. Methanol conversion and dry product composition vs pressure. Temperature = 250 °C, S/M = 1, and W/Fmeth = 0.0246 kgcat s mmol−1.

(24)

−1

We estimated KCO ∼ 25 bar by proper plotting the equation above, which is linear in P and indeed obtained good agreement of the f(P) line (Figure 11, S/M = 0). Similar to the atmospheric pressure cases, small percentages of CH4 (less than 0.5%) were observed at 6 bar when water was absent in the feed. 3.4. Deactivation and Coke Formation. This section presents observations of deactivation and coke deposition on the Cu/ZnO/Al2O3 catalyst under operating conditions that are relevant for on-board hydrogen production for internal combustion engine. 3.4.1. Preliminary Characterization. To determine the amount of reducible species present in the catalyst and reveal the temperature at which reduction occurs, temperatureprogrammed reduction was performed on 1 g fresh [i.e., before being exposed to methanol (Figure 12)]. Reduction occurs at a peak temperature of 270 °C, evident by opposite peaks of TCD and H2 (measured by MS); since hydrogen is the only species present, the area of the TCD signal is used to calculate the catalyst hydrogen uptake. This value is then used to assess the rereducibility of the catalyst after reaction and reactivation. To assess the effect of sintering on deactivation, a second TPR was performed after oxidizing the sample: the hydrogen uptake profile was identical to Figure 12, indicating that the reducibility of the catalyst did not change, suggesting that the deactivation was due to carbon deposits rather than catalyst sintering. In order to find the source of deactivation, then after an isothermal methanol decomposition run of several hours, a temperature-programmed desorption was performed in a He atmosphere to both degas the catalyst sample and to monitor the species desorbed from the sample. This procedure was

was expected: at full conversion, it is expected to be 5/7 H2 and 2/7 CO or CO2, while at f = 0.6 (as at 225 °C), it is 2/3 and 1/3, respectively. In the absence of water in feed (S/M) = 0, the conversion is significantly smaller (Figure 10, left) than that at S/M = 1 (Figure 1), while CO concentration is surprisingly smaller than expected. The expected product distribution, in the absence of other reactions, is 2/3 H2 and 1/3 CO (no CO2). The low conversion can be attributed to strong inhibition by CO (e.g., rate eq 18 but with D = 1 + KCOPCO). It may also result from not accounting for other products like methane and DME, as well as from the coking reaction 2CO → CO2 + C. We rule out the former explanation; indeed small amounts of methane (less than 0.5%) were also detected by the GC, but they cannot account for the difference. The activation energies were determined as 45, 46, and 55 kJ/ mol for S/M = 1, 0.5, and 0, respectively, using the Arrhenius plot (Figure 10, right). We single out the inhibition effect of CO on the rate since CO concentration becomes significant in the absence of steam. This deviates from other studies in the literature (see Introductions, eqs 9−15), which accounted for inhibition by hydrogen or by methanol; however, H2 and methanol concentrations do not differ significantly from S/M = 0 to that in S/M = 1 and will not account for this effect. This is supported further by the study below. Pressure studies (Figure 11) shows that the conversion under pressure is inhibited significantly when steam is deficient, when compared with results at S/M = 1. The inhibition is moderate at S/M = 0.5, the conversion decline with increasing P due to CO inhibition, and there is a shift from CO2 to CO as water is

Figure 8. Methanol conversion and dry product composition (left) vs temperature and the distance from WGS equilibrium expressed as ln[(CCO2CH2)/ (CCOCW)] (right) vs 1/T plot; pressure = 6 bar, S/M = 1, W/Fmeth = 0.0246 kgcat s mmol−1. G


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Figure 9. Methanol conversion and dry product composition (left) vs temperature and the distance from WGS equilibrium expressed as ln[(CCO2CH2)/ (CCOCW)] (right) vs 1/T plot; S/M = 0.5, pressure = 1 bar, W/Fmeth = 0.025 kgcat s mmol−1.

Figure 10. Methanol conversion and dry product composition (left) vs temperature, S/M = 0, pressure = 1 bar, W/Fmeth = 0.025 kgcat s mmol−1, and Arrhenius plot (right, kMD in cc g−1s−1) for various S/M ratio.

Figure 11. Methanol conversion and dry product composition vs pressure, S/M = 0.5 (left), S/M = 0 (right, conversion predicted using eq 24 is also shown; temperature = 250 °C, W/Fmeth = 0.025 kgcat s mmol−1.

energies and sites. TPD profile of the sample that was used for methanol decomposition at 300 °C still showed similar carbon monoxide desorption behavior, but an additional peak of hydrogen accompanied by a CO2 peak that occurred at 220 °C; that can be related with a hydrocarbon formation that might be associated with the observed catalyst deactivation. Temperature-programmed oxidation was performed to verify whether any residual carbon depositions were formed during the isothermal reaction runs. The samples showed similar TPO profiles (Figure 14) with a maximum in CO2 signal at 570 °C accompanied by a minimum in the oxygen signal. After calibrating the CO2 signal, the amount of measured CO2 was then converted to carbon percent weight of spent catalyst. 3.4.2. Characterization of Catalysts Used in Catalytic Tests. This procedure was performed on 4 catalyst samples collected in previous catalytic tests at (300 °C and atmospheric pressure) in a flow reactor; the results are shown in Table 1 (set 2) showing that the presence of steam inhibited carbon deposition by an order of

Figure 12. Hydrogen consumption profile, fresh catalyst TPR.

performed at 3 different temperatures (200, 250, and 300 °C) with three different fresh catalyst samples (Figure 13). Carbon monoxide and hydrogen desorbed in a similar fashion at a peak temperature of 270 °C, indicating similar desorption H


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Figure 13. TPD profiles of used catalyst after 3 h reactions at 200 °C (left, solid lines) or 250 (left, dashed lines) and 300 °C (right).

6 bar and at lower S/M. Full conversion stable behavior was observed with a stoichiometric feed ratio, over a period of 10 h at P = 1 bar (Figure 15A). Product composition is, as expected, 75% H2 complemented by CO2 and some CO. Lower initial conversion and a strong decline in methanol conversion with time was observed at 6 bar; higher CO concentrations are evident in this case with similar product composition (Figure 15B). Similar results were observed under the same conditions with S/M = 1/2: stable behavior at atmospheric pressure with expected product composition 5/7 H2 complemented by CO2, CO (Figure 15C). At 6 bar, we observed an oscillatory behavior around the expected (5/7) value of hydrogen composition with some shift from CO2 to CO (Figure 15D). At the end of the run most of the carbon is in the form of CO. The higher CO concentration at high pressure suggests inhibition of WGS reaction under these conditions. The time on stream study with S/M = 0 (Figure 15E) at 275 °C and at atmospheric pressure indicates small catalyst deactivation over a period of 10 h. While the H2 mole-fraction is around its expected value (2/3), the CO fraction is below its expectation and CO2 is not expected if Boudouard reaction is ignored. At 6 bar, the deactivation is significant, with significant CO2 production. After the time on stream experiments at atmospheric pressure and at 6 bar, a TPD experiment was carried out in helium (50 mL/min) by ramping up to 400 °C and keeping the temperature for 30 min to remove all the adsorbed materials from the catalyst surface. After cooling in helium, TPO was carried out with 10% O2/He (100 mL/min) up to 750 °C and kept for 20 min to quantify the amount of coke produced on the catalyst surface during reaction. The TPO thermogram (Figure 16, left) showed only one peak in the MS CO2 signal, at ∼680 °C which can be attributed to graphite formation. The amount of coke produced is 0.020 g of carbon per gram of catalyst for the 6 bar case and 0.021 g of carbon per gram of catalyst for the atmospheric pressure test. Analysis of the TPO profiles (Figure 16, right) obtained for the S/M = 0.5 case revealed that the amount of coke produced was 0.020 and 0.018 g of carbon per gram of catalyst for the atmospheric and 6 bar pressures, respectively (Table 1). However, the TPO conducted with the spent catalyst with S/ M of 0 after TPD at 400 °C revealed the formation of similar amount of coke (0.020 g C/gcat) on the catalyst surface as that for other S/M ratios. The deactivation that was observed at low S/M, and the coke formation that was verified visually as well as reported in previous studies showing heavy catalyst deactivation when operating under dry conditions, led us to conduct a catalytic activity test for over 50 h at 290 °C. With a stoichiometric water to methanol feed ratio, the conversion was quantified periodically and the

Figure 14. TPO profile of the used catalyst after 3 h methanol decomposition reaction at 300 °C followed by TPD.

Table 1. Carbon Deposition at Various Reaction Conditionsa set 1: carbon deposition (g of C/g catalyst) after 10 h reaction methanol steam reforming methanol steam reforming methanol decomposition

S/M

set 2: carbon deposition (g of C/ g catalyst) after 0.5 h

1 bar

6 bar

1 bar

1

0.021

0.02

0.005 (3 mm) or 0.01 (0.7 mm)

0.5

0.020

0.018

0

0.020

−

0.06 (3 mm)

a

Conditions: set 1 (described in section 3.3):1−2 mm particles, 275 °C, 10 h TOS, W/Fmeth = 0.025 kg s mmol−1; set 2 (described in section 3.2): 3 or 0.7 mm particles, 300 °C, 3 h TOS).

magnitude. The observation that C deposition rates is higher on a powder catalyst (0.7 mm) when compared with that on pelleted catalyst requires verification. The decrease of carbon formation for both the presence of steam and larger catalyst particles can be associated with the lower carbon monoxide selectivity, as the carbon monoxide presence directly increases the selectivity for the Boudouard reaction, believed to be the main route to carbon deposition in methanol reforming. Also presented in Table 1 are carbon deposition obtained over 10 h during the experiments described below with smaller particles, at 275 °C (vs 300 °C in set 2). The deposition rates were similar in order of magnitude to the data in the other set that was obtained in 30 min, suggesting that the process reaches saturation quite quickly. Also, no coherent trend of coke deposition on P or S/M was observed. 3.4.3. Catalyst Stability. The stability of the catalyst was tested at 275 °C at 1 or 6 bar (left, right columns, Figure 15) and various S/M ratios: the decline in activity is more pronounced at I


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Figure 15. Methanol conversion and dry product composition vs time on stream (S/M = 1 (A and B), S/M = 0.5 (C and D) and S/M = 0 (E and F), pressure = 1 bar (A, C, and E) or 6 bar (B, D, and F), T = 275 °C, W/Fmeth = 0.025 kgcat s mmol−1).

Figure 16. MS spectra corresponding to CO2 (mass 44) during TPO under 10% O2/He (flow rate 100 mL/min) after methanol steam reforming at 275 °C for 10 h at atmospheric pressure with S/M = 1 (left) or 0.5 (right), each at the two indicated pressures.

concentrations of the dry product monitored continuously (Figure 17). Methanol conversion drops rapidly in the initial 5−6 h. Then there is a slower decline for the following 10 h followed by a stable activity. The monitoring of the carbon species in the reactor effluent is also shown. In order to further investigate the relationship between carbon deposition and catalytic activity, four separate isothermal runs, each 15 h long, of methanol decomposition reactions were performed in the Autochem in the range of 250−330 °C. The

masses (0−44 a.m.u) were monitored simultaneously, showing the presence of 3 species in the effluent gas (methanol, carbon dioxide, and hydrogen). All experiments show a behavior similar to that described in Figure 17, right, a slower decrease in methanol conversion throughout the time span. Calibration of the mass spectrometer signal was performed based on prepared mixtures of precisely known composition. Gas-phase composition was calculated from the mass spectrometer signal, the hydrogen signal was in linear correlation to J


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Figure 17. Catalyst stability: (left) methanol conversion, CO, and CO2 content vs time on stream (S/M = 1, T = 290 °C, 1 bar), (right) methanol conversion vs time on stream during methanol decomposition at T = 220 °C.

its molar fraction, fragmentation of the different carbon species (CO and methanol) was calibrated, and contributions from other than the indicated ones were subtracted. The molar composition was then used to calculate methanol transient conversion as shown in Figure 17, right. The hourly average deactivation rate (i.e., the decrease in conversion per hour) is shown in Figure 18 together with the

Figure 19. CO2 profile−TPO after reaction at 330 °C.

that the source of these carbon deposits might be hydrocarburic rather than filamentous carbon deposits that could cause the third peak which was not accompanied by H2O formation. A previous study showed that carbon deposits with a burn off temperature of around 700 °C were associated with graphite formation on catalyst surface. Again to assess the effect of sintering on deactivation, a second TPR was performed, and the same hydrogen uptake profile was obtained indicating that the reducibility of the catalyst did not change, suggesting that the deactivation was due to carbon deposits rather than catalyst sintering.

Figure 18. Average hourly deactivation and carbon deposited at various reaction temperatures (3 mm pellets, 1 bar, S/M = 1, 15 h experiment).

total amount of carbon deposited expressed in terms of percent carbon/catalyst weight; the amount of carbon deposited was quantified by TPO after 15 h time on stream. Apparently, there is no simple (linear) relation between the two. A possible explanation is that the rate of coke formation increases with temperature because of the higher production rate of CO which can finally be converted into coke by the Boudouard reaction. The Arrhenius equation parameters calculated from the average carbon production rates are shown in Table 2.

4. CONCLUSIONS The kinetics and diffusion effects of methanol steam reforming reaction over commercial Cu/ZnO/Al2O3 catalyst was studied for steam to methanol (S/M) ratios of 0 to 1 and pressures below 6 bar, with the objective of development of a novel high-pressure propulsion technology based on the concepts of thermochemical recuperation (TCR) and onboard hydrogen production. Better ICE efficiencies are expected with a feed rich in syngas (H2 + CO), which in this case implies that methanol decomposition products are better than those of SR, that, however, will accelerate deactivation (see below). A simple kinetic model assuming methanol decomposition followed by water−gas-shift was used to estimate the rate constants (kMD, kWGS); the apparent activation energy of kMD was estimated as 45−55 kJ/mol for S/M = 1, 1/2, 0; kMD for S/M = 0 (and the conversions obtained) was smaller due to CO inhibition. At temperatures above ∼500 K, the WGS is at instantaneous equilibrium. As the pressure increases, the ratio of CO and CO2 increases moderately up to 4 bar; at 6 bar, the change is drastic. Similar observation was made with S/C = 0.5. Strong pore-diffusion limitations are evident at 1 bar for the 3 mm catalyst, evident experimentally as well as by analysis; the

Table 2. Arrhenius Rate Constant for the Coke formation parameter activation energy Ea pre exponential factor k0

30 kJ/mol 8.01 × 103 mg coke/gcat h

The nature of the carbon deposits on the catalyst surface is not yet clear in the literature; the carbon dioxide profiles measured in the TPO after the reactions at 330 °C showed a different behavior in comparison to the ones at lower temperatures (Figure 19). Three CO2 peaks were observed (at 260, 400, and 700 °C) corresponding to three minima in O2 signal. The first two CO2 peaks were accompanied by an increased H2O signal, indicating K


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(5) Rakopoulos, D. C.; Rakopoulos, C. D.; Kakaras, E. C.; Giakoumis, E. G. Effects of Ethanol−diesel Fuel Blends on the Performance and Exhaust Emissions of Heavy Duty DI Diesel Engine. Energy Convers. Manage. 2008, 49, 3155−3162. (6) Cheng, W. Development of Methanol Decomposition Catalysts for Production of H2 and CO. Acc. Chem. Res. 1999, 32, 685−691. (7) Olah, G. A. Beyond Oil and Gas: The Methanol Economy. Angew. Chem., Int. Ed. 2005, 44, 2636−2639. (8) Joensen, F.; Rostrup-Nielsen, J. R. Conversion of Hydrocarbons and Alcohols for Fuel Cells. J. Power Sources 2002, 105, 195−201. (9) Sa, S.; Silva, H.; Brandao, L.; Sousa, J. M.; Mendes, A. Catalysts for Methanol Steam Reforminga Review. Appl. Catal., B 2010, 99, 43− 57. (10) Valdes-Solis, T.; Marban, G.; Fuertes, A. B. Nanosized Catalysts for the Production of Hydrogen by Methanol Steam Reforming. Catal. Today 2006, 116, 354−360. (11) Cao, W.; Chen, G.; Li, S.; Yuan, Q. Methanol-steam Reforming over a ZnO−Cr2O3/CeO2−ZrO2/Al2O3 Catalyst. Chem. Eng. J. 2006, 119, 93−98. (12) Choi, Y.; Stenger, H. G. Fuel Cell Grade Hydrogen from Methanol on a Commercial Cu/ZnO/Al2O3 Catalyst. Appl. Catal., B 2002, 38, 259−269. (13) Twigg, M. V.; Spencer, M. S. Deactivation of Copper Metal Catalysts for Methanol Decomposition, Methanol Steam Reforming and Methanol Synthesis. Top. Catal. 2003, 22, 191−203. (14) Jiang, C. J.; Trimm, D. L.; Wainwright, M. S.; Cant, N. W. Kinetic Mechanism for the Reaction between Methanol and Water over a CuZnO-Al2O3 Catalyst. Appl. Catal., A 1993, 97, 145−158. (15) Peppley, B. A.; Amphlett, J. C.; Kearns, L. M.; Mann, R. F. Methanol-steam Reforming on Cu/ZnO/Al2O3 Catalysts. Part 2. A Comprehensive Kinetic Model. Appl. Catal., A 1999, 179, 31−49. (16) Geissler, K.; Newson, E.; Vogel, F.; Truong, T. B.; Hottinger, P.; Wokaun, A. Autothermal Methanol Reforming for Hydrogen Production in Fuel Cell Applications. Phys. Chem. Chem. Phys. 2001, 3, 289−293. (17) Purnama, H.; Ressler, T.; Jentoft, R. E.; Soerijanto, H.; Schlogl, R.; Schomacker, R. CO Formation/Selectivity for Steam Reforming of Methanol with a Commercial CuO/ZnO/Al2O3 Catalyst. Appl. Catal., A 2004, 259, 83−94. (18) Amphlett, J. C.; Evans, M. J.; Jones, R. A.; Mann, R. F.; Weir, R. D. Hydrogen Production by the Catalytic Steam Reforming of Methanol Part 1: The Thermodynamics. Can. J. Chem. Eng. 1981, 59, 720−727. (19) Thattarathody, R.; Sheintuch, M. Kinetics and Dynamics of Methanol Steam Reforming on CuO/ZnO/alumina Catalyst. Appl. Catal., A 2017, 540, 47−56. (20) Takeguchi, T.; Kani, Y.; Inoue, M.; Eguchi, K. Steam Reforming of Methanol on Copper Catalysts Supported on Large-surface-area ZnAl2O3. Catal. Lett. 2002, 83, 49−53. (21) Frank, B.; Jentoft, F. C.; Soerijanto, H.; Krohnert, J.; Schlogl, R.; Schomacker, R. Steam Reforming of Methanol over Copper-containing Catalysts: Influence of Support Material on Microkinetics. J. Catal. 2007, 246, 177−192.

apparent kMD is almost diffusion free for particles smaller of 0.7 mm in diameter. Thus, the following model and rate expressions are suggested: the main, methanol decomposition, reaction rate is described by rMD = kMD,trueρη(ϕ)CM /(1 + K COPCO) kMD,true = 8.3 exp[−Ea /R(1/T − 1/553)]ccg −1s−1

where the activation energy is Ea = 50 kJ/mol, the effectiveness factor η(ϕ), and Thiele modulus are outlined in eqs 22 and 20. The CO inhibition coefficient KCO was estimated to be ∼25 bar−1 at 250 °C (for the case of strong CO inhibition, the η(ϕ) solution of first-order kinetics is not valid anymore and should be modified). The rate of the following reaction, water gas shift, rWGS = k WGS(CCOC W − CCO2C H2 /Keq)

is fast at the range of temperatures and W/F studied and results are well-described by the assumption of instantaneous equilibrium (for the measured or predicted conversion). Deactivation rates and coke formation were also studied and were found to be marginal under atmospheric pressures over a period of 10 h and became evident only at S/M = 0 and 275 °C; at 6 bar, the decline was evident already at S/M = 0.5. The source of deactivation was attributed to coking, a conclusion based on TPO of spent catalyst. In the case of S/M = 1 and 0.5, and at 6 bar, there was a shift in the composition from CO2 to CO with time over a period of 10 h at 275 °C. However, the carbon deposition in all cases were estimated to be about the same.

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. ORCID

Rajesh Thattarathody: 0000-0002-6922-5473 Notes

The authors declare no competing financial interest.

■

ACKNOWLEDGMENTS This research was supported by the Israel Science Foundation through Grant No. 1728/12 and by the Ministry of Energy and Water Resources. The Micromeritics instrument was supported by the Wolfson Foundation and by the I-Core project of the ISF through The Nancy and Stephen Grand Technion Energy Program (GTEP).

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REFERENCES

(1) Ivanic, Z.; Ayala, F.; Goldwitz, J.; Heywood, J. Effects of Hydrogen Enhancement on Efficiency and NOx Emissions of Lean and EGRdiluted Mixtures in a SI Engine. SAE Tech. Pap. Ser. 2005, 10.4271/ 2005-01-0253. (2) Poran, A.; Artoul, M.; Sheintuch, M.; Tartakovsky, L. Modeling Internal Combustion Engine with Thermo-chemical Recuperation of the Waste Heat by Methanol Steam Reforming. SAE Int. J. Engines 2014, 7, 234−242. (3) Chakravarthy, V. K.; Daw, C. S.; Pihl, J. A.; Conklin, J. C. Study of the Theoretical Potential of Thermochemical Exhaust Heat Recuperation for Internal Combustion Engines. Energy Fuels 2010, 24, 1529− 1537. (4) Poran, A.; Tartakovsky, L. M. Energy Efficiency of a Directinjection Internal Combustion Engine with High-pressure Methanol Steam Reforming. Energy 2015, 88, 506−514. L


M Ratio Effects in Methanol Steam

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