Preparation and Catalytic Evaluation of Cobalt-Based Monolithic and

Jul 23, 2008 - examples,7–13 monolithic catalysts were compared with com- mercial or .... coriolis mass flow meter (Endress+Hauser, Promass 83F DN8)...
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Ind. Eng. Chem. Res. 2008, 47, 6589–6597

6589

Preparation and Catalytic Evaluation of Cobalt-Based Monolithic and Powder Catalysts for Fischer-Tropsch Synthesis Robert Guettel, Jens Knochen, Ulrich Kunz, Markus Kassing,† and Thomas Turek* Institute of Chemical Process Engineering, Clausthal UniVersity of Technology, Leibnizstrasse 17, 38678 Clausthal-Zellerfeld, Germany

Cobalt-based monolithic and powder catalysts for Fischer-Tropsch synthesis were prepared. The aluminasupported catalysts contained cobalt (18.6 ( 0.9 wt %) and rhenium (1.2 ( 0.1 wt %) as active phases. To ensure the comparability of both catalysts, monolithic and powder catalysts were prepared from the same CoRe/γ-Al2O3 active powder. While the monolith was prepared by dip coating, the slurry for the coating procedure was also used for preparation of the powder catalyst. It could be shown that both catalysts have comparable composition, pore structure, Brunauer-Emmett-Teller (BET) surface area, and active metal surface area. Catalytic measurements with suspended powder catalyst in a stirred tank reactor and monolithic catalyst in a fixed-bed reactor in the slug-flow regime were performed during Fischer-Tropsch synthesis. Higher reaction rates at comparable methane selectivities were obtained with the monolithic catalyst. Estimations show that the advantageous mass-transfer characteristics of the monolithic catalyst in the slug-flow regime are responsible for this reaction rate enhancement. 1. Introduction In heterogeneously catalyzed fluid-phase reactions, different catalyst geometries can be used. Duducovic et al. have suggested distinguishing between fluidized-bed, randomly fixed-bed, and structured fixed-bed catalysts.1 In catalyzed gas-liquid-solid reactions, these catalyst geometries may be used in different suitable catalytic reactors: (a) slurry stirred tank reactor, slurry bubble column (fluidizedbed catalyst) (b) trickle-bed reactor, flooded bed reactor (randomly fixedbed catalyst) (c) monolith reactor (structured fixed-bed catalyst). During the recent attempts of process intensification, structured catalysts have become quite popular for heterogeneously catalyzed multiphase reactions.2,3 Honeycomb monoliths especially are supposed to be attractive alternatives to conventional fixed-bed and fluidized-bed structures in two- and three-phase applications. The main advantages of structured honeycomb catalysts are low pressure drop, high geometric surface area, high mass-transfer coefficients, and short diffusion lengths. In contrast to fluidized-bed catalysts, there is no need for catalyst separation and little danger of catalyst attrition. The low radial heat transfer can be seen as a drawback in some applications. Nevertheless, theoretical investigations based on correlations for mass transfer, hydrodynamics, pressure drop, and intrinsic kinetics showed a high potential for structured catalysts, compared to randomly fixed beds.4–6 In addition to modeling studies, several experimental investigations were conducted to evaluate the potential of monolithic catalysts (see Table 1). During these studies, it was generally assumed that the different catalyst structures had comparable properties. In most cases, the comparability was based on using the same mass of the catalytically active metal in the different * To whom correspondence should be addressed. Tel.:+49 5323 72 2184. Fax: +49 5323 72 2182. E-mail address:turek@ icvt.tu-clausthal.de. † Present address: Institute for Separation Technology, Clausthal University of Technology, Leibnizstrasse 15, 38678 Clausthal-Zellerfeld, Germany.

geometries. However, it can be suspected that the different catalyst geometries in some of the studies summarized in Table 1 did not have exactly the same properties. In the literature examples,7–13 monolithic catalysts were compared with commercial or homemade pellets that were used as fixed-bed or, after crushing, as suspended catalysts. In these cases, it is not guaranteed that the pore structures and the distribution of the active metals in different geometries are the same. In the paper by Boger et al.,14 the details of the powder catalyst used were not given. In the paper by Moulijn and co-workers,15 the same silica source was used for washcoating the monolith and preparing the powder catalyst. However, the use of a binder in case of the monolith may have caused differences to the powder catalyst that was manufactured without binder. In the study of Liu et al.,16 the monolith was extruded from active material and then crushed into particles. This procedure ensures an even distribution of active sites and the same pore structure in both catalyst geometries. On the other hand, the diffusion lengths in powder and monolith were different. Finally, Schanke et al.17 prepared monoliths by washcoating of a cordierite carrier with a catalytically active CoRe/γ-Al2O3 powder. However, the powder catalyst used for comparison was prepared by tabletting the original catalyst powder and crushing it to the desired sieve fractions. During the pressing of tablets, the pore structure of the catalyst may have changed. Nijhuis et al.18 summarized the methods for the preparation of monolithic honeycomb that are used in research and industrial applications. While the most common procedure in research involves the washcoating of the monolithic carrier, followed by impregnation with active metals, the methodology also described how honeycombs can be coated with layers of readymade catalysts. It is well-known that the latter method is quite often used by manufacturers of automotive car exhaust catalysts.19 Our approach intends to focus on the preparation of different catalyst geometries with comparable dispersion of active sites in the catalyst layer and pore structure. For this purpose, the method starts with the preparation of a powder that contains active metal on a carrier. This is followed by washcoating of the monolith carrier with a suspension of this base catalyst. For

10.1021/ie800377n CCC: $40.75  2008 American Chemical Society Published on Web 07/23/2008

6590 Ind. Eng. Chem. Res., Vol. 47, No. 17, 2008 Table 1. Summary of Experimental Studies Comparing Different Catalyst Geometries in Gas-Liquid-Solid Three-Phase Reactorsa ref(s) 7

reaction

compared reactor types

Broekhuis et al.

hydrogenation of glucose to sorbitol

MR, SSR

Enache et al.8

hydrogenation of pyrolysis gasoline

MR, RFBR

Liu9

hydrogenation of styrene and 1-octene MR, RFBR

Marwan and Winterbottom10 hydrogenation of butyne-1,4-diol 11

DBC (monolithic, packed, slurry) MR, RFBR

Mazzarino et al.

hydrogenation of AMS

Nijhuis et al.12,13

hydrogenation of AMS, consecutive hydrogenation of benzaldehyde hydrogenation of edible oil

MR, RFBR

hydrogenation of 3-methyl-1pentyn-3-ol dehydrogenation of ethylbenzene to styrene Fischer-Tropsch synthesis

MSR, SSR

14

Boger et al.

Hoek et al.15 Liu et al.

16

Schanke et al.17

MSR, SSR

MR, RFPB MR, RFBR

catalyst MR: monolith + washcoat + Ru Powder: commercial MR: monolith + washcoat + Pd/Ni Pellets: commercial MR: monolith + 4 wt % NiO/Al2O3 washcoat, 70 µm Pellets: 3.2 mm and 2 sieve fractions, no composition given MR: monolith + washcoat + 1 wt % Pd Pellets: 16 mm Raschig rings + 1 wt % Pd MR: commercial 1 wt% Pd on monolith Pellets: commercial egg-shell Pd/alumina MR: monolith + washcoat + 1 wt % Ni Pellets: commercial extrudates + 7 wt % Ni MR: monolith + washcoat + Pd Powder: not specified MR: monolith + washcoat + binder + Pd Powder: Si-source + Pd, no binder MR: extrusion from active material, wall thickness ) 640 µm Pellets: crushed monolith, 420-250 µm base catalyst: commercial alumina + Co/Re MR: monolith + base catalyst washcoat Pellets: crushing and sieving of base catalyst

a Abbreviations used in this table: AMS, R-methylstyrene; MR, monolithic reactor; SSR, stirred slurry reactor; MSR, monolithic stirred reactor; DBC, down flow bubble column; RFPB, radial flow packed bed; and RFBR, randomly fixed bed reactor.

Figure 1. Preparation method for powder and monolithic catalysts.

improved stability of the catalyst layer, a binder was used. To achieve the highest degree of comparability, the powder catalyst was prepared from the same binder-containing suspension. As an example of industrial relevance the preparation of cobalt-based and rhenium-promoted catalysts for FischerTropsch synthesis (FTS) was chosen.17 The Fischer-Tropsch reaction can be described as the formation of paraffinic or olefinic chains (see eq 1): CO + 2H2 f -CH2- + H2O

(1)

In the present contribution, the textural and catalytic properties of powder catalysts and washcoats applied on cordierite carriers are compared. The prepared catalysts have been tested in an experimental setup that allows for measurements with suspended catalysts in a stirred tank reactor and with fixed-bed monoliths, according to the loop-reactor concept proposed by Kapteijn et al.20 In the latter case, the liquid phase is recirculated at high liquid flow rates to intensify the mass-transfer rates in the

desirable slug-flow regime inside the monolith channels, which was realized, for the first time, during Fischer-Tropsch synthesis. 2. Experimental Section 2.1. Catalyst Preparation. Powder and monolithic catalysts with approximate contents of 20 wt% cobalt and 1 wt% rhenium on γ-Al2O3 were prepared as described by Holmen et al.17 (See Figure 1.) This allows comparison to previous investigations of monoliths in Fischer-Tropsch synthesis by the groups of Kapteijn and Moulijn 21,22 and Holmen.23–25 First, a base powder catalyst was prepared (step 1). The base catalyst and additional binder material was suspended (step 2) to coat a cordierite monolith carrier by dip coating (step 3). The suspension was also used to produce the powder catalyst by drying and calcining (step 4). For preparation of the base powder catalyst, 139.5 g of alumina powder (γ-Al2O3, 5 µm, Puralox UF 5/230, Sasol)

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Figure 2. Schematic of the experimental setup.

was suspended in a solution of 174.6 g of cobalt nitrate (Co(NO3)2 · 6 H2O) and 4.56 g of perrhenic acid (54 wt%, H.C. Starck) in 222 mL of deionized water. Subsequently, this suspension was dried at 90 °C and calcined for 4 h at 400 °C in air. The slurry for dip coating was prepared by suspending 60 g of the base powder catalyst in 54 mL of deionized water and 34 mL of colloidal alumina binder (pseudo-boehmite, 20 wt%, 50 nm, pH 7, Alfa Aesar). A pH of 7 in an aqueous environment was optimal, as will be shown later. The suspension was dispersed with an ultrasound sonotrode for 1 min at 100 W. The resulting slurry contained 45 wt% of solid, while 10 wt% of the solid fraction consisted of colloidal binder material. As carriers, monoliths (cordierite, 400 cpsi, Corning) with a diameter of 16 mm and a length of 50 mm were used. After dip coating for 1 min, the monolith channels were blown out with pressurized air to remove the excess slurry. The outer surface of the monolith was cleaned to remove the coating. Finally, the coated monoliths were dried for 1 h at 90 °C and calcined for 4 h at 400 °C in air. The coating procedure was repeated to achieve higher washcoat mass fractions; each step was followed by drying and calcination. For comparison, monolith samples were also coated with pure alumina washcoats. The mechanical stability of the washcoats was examined by flow-through tests with water. After 3 h at a velocity of 12 cm/ s, followed by 1 h at a velocity of 1 m/s, a loss of only 0.3 wt% was measured. Hence, the washcoat on the monolithic carrier seems to be sufficiently stable. The final powder catalyst was prepared using the same slurry as that used for the dip coating. The slurry was dried for 6 h at 90 °C, calcined for 4 h at 400 °C, and then crushed and sieved to the desired fraction. 2.2. Catalyst Characterization. The zeta potential of slurries was measured with a Nano-ZS zeta sizer (Malvern Instruments). The resulting catalysts and the main intermediates were analyzed by different methods. The chemical composition of the catalysts was determined by inductively coupled plasma-optical emission spectroscopy (ICP-OES). The pore structure and BrunauerEmmett-Teller (BET) surface area were measured using nitrogen adsorption. X-ray diffraction (XRD) measurements were performed to determine the phase structures. The distribution and the layer thickness of the washcoat were examined by scanning electron microscopy (SEM).

Temperature-programmed reduction (TPR) was performed to obtain information about the phases present in the active material. The samples were heated with a heating rate of 10 K/min, from 50 °C to 550 °C in a flow (30 mL/min, NTP) of 10% H2 in argon at atmospheric pressure. Pulsed chemisorption measurements with CO were used to determine the metal surface area, dispersions, and mean cluster sizes. The pulsed chemisorption was performed at 50 °C with 10% CO in helium after reducing the catalyst for 1.5 h at 350 °C under pure hydrogen. For both techniques, a BELCAT-M (BEL Japan, Inc.) catalyst analyzer was used. 2.3. Catalytic Measurements. Catalytic measurements were performed to compare the reaction rates and methane selectivities obtained with monolithic and powder catalysts during Fischer-Tropsch synthesis. The powder catalyst was crushed and sieved to a fraction of 50-140 µm. The monolithic catalyst was investigated both as fixed-bed and suspended-bed catalysts. In the latter case, the monolithic catalyst was crushed to allow its use in a stirred tank reactor. To avoid separation of the washcoat layer and the inert cordierite carrier, the crushed material was not sieved but completely used for the catalytic measurements. The powder catalysts were reduced in an external oven in a flow of 10% H2 in N2 at atmospheric pressure and a temperature of 623 K for 36 h after heating in N2 at a rate of 1 K/min. After reduction, the catalyst was cooled in N2 to room temperature, covered with Squalane (C30H62, CAS No. 111-013) to prevent oxidation and immediately transferred to the reactor. The monolithic catalyst for fixed-bed investigations was reduced in the experimental setup using the same procedure. The experiments were performed in a continuously operated setup that combines a fixed-bed reactor with a stirred-tank reactor, according to the loop reactor concept (see Figure 2). The experiments with a suspended catalyst were performed in a stirred-tank reactor with a volume of 1 L, which was equipped with a gas injection stirrer. For monolithic catalysts, a tubular reactor (18 mm ID) with a catalyst bed length of 50 cm was equipped with a static mixer (Fluitec Georg AG, CSE-X/8G 12.4 mm, five elements) at the top and a sieve at the bottom. The reactor was heated with heating tape from Horst GmbH. During operation in fixed-bed mode, the stirred tank was used as a phase separator for the gas and liquid phases. Gas and liquid can be recycled independently, using a compressor and a pump,

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respectively. A gear pump (Gather Industrie) transports the liquid phase from the stirred tank reactor and pumps it toward the tubular reactor. The liquid mass flow is measured by a coriolis mass flow meter (Endress+Hauser, Promass 83F DN8). The gas is recycled by a compressor (Haskel, HUAGD-4) and the mass flow is measured by a coriolis mass flow meter (Bronkhorst, Cori-Flow). The reactants CO and H2, as well as an internal standard (CO2), were fed to the setup via mass flow controllers (Bronkhorst El-Flow F-230M). The liquid level in the setup was maintained by removing the produced liquid from the stirred tank reactor by a continuous gas outlet stream. The device used allows the separation of catalyst and liquid by a 10 µm frit. It is described in detail elsewhere.26 Downstream from the stirred-tank reactor, the setup was equipped with a wax and a water separator operated at 150 and 0 °C to take wax and water samples, respectively. The tubing between the stirredtank reactor and water separator was heated to 150 °C. The stirred-tank reactor was initially filled with ∼450 mL of squalane as the liquid phase. Squalane was chosen as the starting solvent because it is a good solvent for Fischer-Tropsch products and inert in Fischer-Tropsch synthesis.26 After addition of the catalyst, the reactor was purged with nitrogen and pressurized afterward to 21 bar, using a backpressure regulator (Tescom 44-1700) with synthesis gas with a H2/CO ratio of 2. Subsequently, the catalyst was activated for each 24 h at 160 and 170 °C under synthesis gas, respectively, before the temperature was increased further. During the catalytic experiment, a stirring speed of 1000 min-1 was used to minimize external masstransfer limitations. In fixed-bed operation mode, the gas and liquid superficial velocity was set to 0.05 m/s to achieve the advantageous slug-flow regime inside the monolith channels. Because the gas recycle ratio at the chosen conditions is ∼100 and the conversion per pass is low, both the monolith-loop reactor and the stirred-tank reactor are well-mixed. The experiments were conducted at temperatures of 190 and 200 °C at a modified residence time of 12000 kgcat s/mSG3. For monolithic catalysts in fixed-bed operation, the mass of active catalyst material was limited to ∼10 g, because of the dimensions of the reactor. In the stirred-tank reactor, ∼14.5 g of active catalyst material (powder or washcoat) were used. In the case of the crushed monolith, the total solid fraction consisted of inert carrier material and active catalyst. This lead to a mass fraction of ∼18.6 wt% solids in the suspension, compared to ∼4.1 wt% in the case of a powder catalyst where no inert solid was present. A gas chromatograph (HP 5890 GC) was used to take online gas samples every 2 h after passing the water separator. The gas chromatograph was equipped with a thermal conductivity detector (TCD) and a flame ionization detector (FID), a 250 µL sample loop, and a HP-PLOT/Q column (Agilent, 30 m × 0.533 mm × 40 µm) to analyze for the presence of CO, H2, CO2, and CH4. Under the assumption that no CO2 is formed during the reaction, the CO conversion (X) and methane selectivity (S) can be calculated using eqs 2 and 3: X)

S)f

R0,CO - RCO R0,CO

(

R0,CH4 R0,CO - RCO

(2)

)

(3)

Here, R is the peak area ratio of species CO or CH4 and the internal standard, and f is a specific calibration factor.

Figure 3. Zeta potentials of the base materials and catalysts.

3. Results and Discussion 3.1. Dip Coating. To optimize the pH conditions for the dipcoating procedure, zeta potential measurements were performed for the base materials by varying the pH between 3.5 and 7, using hydrochloric acid. The zeta potential is a reference value for the adhesive forces between particles. Particles with different or low absolute potentials have a tendency to agglomerate, whereas a suspension with particles with a potential of more than (25-50 mV can be assumed to be stable.27 The results of the measurements are shown in Figure 3. The monolithic carrier was crushed to particles to measure its potential. Between pH 3.5 and 7, the potential of the carrier decreases from -20 mV to -35 mV, while the potential of the base powder decreases from 40 mV to 20 mV. The colloidal binder material has a potential of 55 mV at pH 7. After the base powder is impregnated with active material, the potential of the base powder catalyst is lower than that of the base powder and decreases between pH 3.5 and pH 7 from 40 mV to 0 mV. Adding the binder material to the base powder catalyst leads to an increase in the potential of the final powder catalyst that remains almost constant at 40 mV over the considered pH range. These results reveal that a suspension of the final powder catalyst is stable. Because of the large difference in the potentials of monolith carrier and final powder catalyst at pH 7, a good adhesion between the washcoat and the cordierite carrier should result. Christiani et al.28 analyzed the effect of aging time in γ-Al2O3 slurries for dip coating and suggested using a pH of 3.5. However, in contrast to the experiments described in this paper, Christiani et al.28 used base alumina powder without an active metal and colloidal binder. In conclusion, for coating monoliths with the CoRe/Al2O3 powder catalyst with added AlOOH binder, a pH of 7 should be chosen. The mass of the washcoat on the monolithic carrier was measured gravimetrically. The different curves shown in Figure 4 were obtained by repeating the experiments under identical conditions. A comparison of coatings with base alumina powder and powder catalyst shows that higher washcoat mass fractions are achieved using the pure alumina slurry. This can be explained by the higher zeta potential difference between alumina powder and monolithic carrier, compared to the powder catalyst (see Figure 3). The high washcoat mass fraction increase during the sixth coating step with alumina powder can be explained by blocking of some channels. In all other cases, the increase in the washcoat mass fraction gradually decreases with the number of coating steps. Two explanations may be given

Ind. Eng. Chem. Res., Vol. 47, No. 17, 2008 6593

Figure 4. Mass fraction of washcoat on monolithic catalysts.

for this phenomenon. First, if a surface is coated with a porous layer, the zeta potential of the entire system will change until, eventually, the values of the coating are attained. Consequently, the driving force for agglomeration will decrease. Second, the surface area of the channel also diminishes as the washcoat mass fraction increases. This leads to a decrease in coating capacity of the capillary and, thus, in a lower achievable washcoat mass fraction per coating step. 3.2. Catalyst Composition. The catalyst composition was determined by ICP-OES measurements. According to the elemental analysis, the prepared catalyst had a cobalt content of 18.6 ( 0.9 wt % and a rhenium content of 1.2 ( 0.1 wt % on the γ-Al2O3 support. These values are in good agreement with the desired composition of 20 wt % cobalt and 1 wt % rhenium. The XRD measurements show that the crystal structure is similar to the results obtained by Storsaeter et al.25 In addition to the Co3O4 and γ-Al2O3 that was found by Storsaeter et al.,25 the colloidal alumina binder phase could be identified. 3.3. Pore Structure and Surface Areas. Data describing the pore structure and BET surface area of catalysts and intermediates are summarized in Table 2. The washcoat could not be characterized separately, because it is attached to the monolith carrier. Therefore, the corresponding data were calculated from values for the monolith carrier and catalyst using eqs 4–6: awc ) Vpore,wc )

amcat - (1 - xwc)amc xwc

Vpore,mcat - (1 - xwc)Vpore,mc xwc

εwc )

εmcat - (1 - ywc)εmc ywc

(4) (5) (6)

In eq 6, the volume fraction of the washcoat in the catalyst ywc is used. This value can be calculated from the mass fraction xwc, as shown in eq 7, where F is the apparent density. From the measured densities and porosities given in Table 2, a value of ywc ) 0.27 is obtained: ywc )

1 Fwc 1 - xwc 1+ Fmc xwc

( )(

)

(7)

The results in Table 2 show that the pore volume of the washcoat is smaller than that of the final powder catalyst, although the calculated porosities are very similar. This

contrasting result may be caused by measurement errors, especially for the monolithic catalyst and, therefore, also for the washcoat. As can be seen in Figure 5, there are only minor differences in the pore size distributions of final powder and monolithic catalyst. The BET surface area of the washcoat and the final powder catalyst are also very similar (see Table 2). Thus, it can be concluded that the monolithic catalyst and the final powder catalyst have comparable pore structure properties. Becuase of the fact that the same active catalyst powder was used for the washcoating and preparation of the powder catalyst, the distribution of active sites must also be very similar. The results of TPR measurements of final powder and monolithic catalysts are shown in Figure 6. The profile of monolithic catalyst shows smaller peaks for Co(NO3)2, Co3O4, and rhenium, which may be caused by its longer overall calcination time. The results of the chemisorption experiments obtained with the final powder and monolithic catalysts are summarized in Table 3. In all cases, the small rhenium content was neglected and it was assumed that the reduced catalyst consisted of cobalt metal on alumina. It can be seen that the measured active metal surface area of the final powder catalyst is somewhat higher than that for the monolithic catalyst. Consequently, a higher metal dispersion and a smaller cobalt cluster size are obtained. The results of chemisorption measurements are predominantly comparable to the literature, although a higher cobalt dispersion of 10.2% has been reported.29 This may be caused by lower cobalt and rhenium contents (12 wt % Co, 0.5 wt % Re) and a different chemisorption method using hydrogen adsorption at 40 °C. 3.4. Characteristic Diffusion Length. The characteristic diffusion length in porous catalysts is usually expressed by the ratio of catalyst volume and its external surface area. If one assumes a spherical geometry of the powder catalyst, the resulting diffusion length amounts to one-sixth of the mean particle diameter. In our case, a coated square monolith channel typically shows a catalyst thickness that is ∼5 times thicker in the corners (∼225 µm), compared to the planar walls (∼45 µm) (see Figure 7). The characteristic diffusion length of the monolith washcoat was calculated using eq 8. Here, the numerator represents the washcoat volume in one cell of the monolith. In the denominator, the external surface area is calculated using the void space of a coated capillary. The open frontal area (OFA) of an uncoated monolith capillary is typically 0.75 and the diameter of a monolith cell (dcell) is 1.27 mm for the 400 cpsi monolith used in our study. This calculation accounts for the distribution of the washcoat thickness, and it should give a reliable value for the characteristic diffusion length. δwc,char )

ywcdcell 4√OFAmc - ywc

(8)

For the crushed monolithic catalysts, it can be expected that the diffusion length of the major fraction is smaller than the washcoat layer thickness on the planar walls (i.e., 45 µm). However, it cannot be excluded that part of the crushed catalyst layer, especially from the corners of the square monolith channels, has a higher diffusion length. Overall, we assume a characteristic diffusion length of 50 µm for the crushed monolith. 3.5. Catalytic Investigations. Table 4 summarizes conditions and results of the catalyst testing experiments. Steady-state conditions during the measurements were achieved after ∼125-150 h. The methane selectivity in all experiments ranges

6594 Ind. Eng. Chem. Res., Vol. 47, No. 17, 2008 Table 2. Properties of Final Catalyst and Intermediates BET surface area, a (m2/g) monolith carriera monolith catalysta base powdera base powder catalysta final powder catalysta washcoatb a

0.28 33.6 197 (208)c 144 167 167

pore volume, Vpore (cm3/g)

average pore diameter, dpore (nm)

0.16 0.22 1.00 0.80 0.67 0.46

2528 730 288 541 325

porosity, ε

apparent density, F (g/cm3)

0.26 0.36 0.64 0.65 0.62 0.63

1.60 1.65 0.645 0.806 0.920 0.920d

Measured. b Calculated. c Data taken from data sheets. d As final powder catalyst.

Figure 5. Distribution of specific pore volume in the final powder catalyst and the washcoat.

°C. One reason to explain these differences could be the different active metal surface areas for monolithic and powder catalysts. Taking this into account, using a modified, area-specific reaction rate, a better agreement between powder and crushed monolithic catalysts is obtained at 190 °C. Investigations reported in the literature30 show a higher reaction rate (by a factor of 3-4) for monoliths. However, these results were obtained in the absence of a liquid phase and, therefore, are not directly comparable. The higher reaction rate of the monolithic catalyst in fixedbed operation mode, compared to the crushed monolithic catalyst, could also be caused by stronger external mass-transfer resistances in the stirred-tank reactor. To test this hypothesis, the mass-transfer coefficients in the stirred-tank reactor and monolith reactor have been estimated. Based on literature correlations,31,32 the overall mass-transfer coefficient kov for the stirred-tank reactor (STR) (eq 9) and the monolith reactor (MR) (eq 10) under reaction conditions was estimated. The value for kaGL,STR was obtained from measured data reported in the literature.33 The density, viscosity, and surface tension of the liquid phase, as well as the Henry coefficient for CO, were calculated with the literature correlations for Sasol wax.33 1 1 1 ) + kaov,STR kaGL,STR kaLS,STR

(9)

where kaLS,STR )

6Dycat dp2

(2 + 0.3Reε0.75Sc0.33) Reε )

( ) εdp4FL3

1⁄3

ηL3

kaGL,STR ) 0.1 s-1

and kaov,MR ) kaGLS,MR + kaGS,MR )

Figure 6. TPR profiles of the final powder and the monolithic catalyst. Table 3. Active Metal Dispersion, Active Surface Area, and Mean Cluster Size of Final Powder Catalyst and Monolithic Catalyst

final powder catalyst monolith catalyst

active metal dispersion (%)

active surface area (mCo2/gcat)

mean cluster size (nm)

3.77 ( 0.38 3.01 ( 0.08

5.09 ( 0.51 4.08 ( 0.11

26.7 ( 2.69 33.2 ( 0.92

from 12.4% to 14.7%. These values are somewhat higher than those reported for comparable catalysts in the literature.30 The powder catalyst shows the lowest methane selectivity, followed by the crushed monolithic catalyst and the monolithic catalyst used in fixed bed operation mode. However, these differences are relatively small. From the measured conversion and inlet molar flow rate of CO, the observed reaction rate can be obtained. Comparing powder and crushed monolithic catalysts, differences in synthesis gas conversion and reaction rate are observed. These differences are more pronounced at the higher reaction temperature of 200

where

(

δf ) dch

0.66Ca2⁄3 1 + 3.33Ca2⁄3

4D δfdch

(10)

) Ca )

ηLuTP σL

The influence of external mass transfer on the observed reaction rate can be analyzed using the time constants for mass transfer (1/kov), the observed reaction rate (1/kobs), and the intrinsic reaction rate (1/kint) (eq 11). The simplified first-order reaction rate constant (kobs) can be calculated from the observed reaction rate, the catalyst density, and the CO concentration in the liquid phase. The gas phase was assumed to be ideally mixed and consist of CO, H2, and H2O only. The intrinsic reaction rate constant may still contain internal mass-transfer effects, especially for the monolithic catalyst. However, at the small

Ind. Eng. Chem. Res., Vol. 47, No. 17, 2008 6595

Figure 7. Scanning electron microscopy (SEM) photomicrograph of the washcoat thickness in the corners of the channels. Table 4. Conditions and Results for Catalyst Testing Powder parameter

at 190 °C

mass catalyst/solid (g) mass fraction solid characteristic diffusion length (µm) catalyst particle size (µm) metal surface area (m2/gcat) conversion selectivity to methane mass specific reaction rate (molCO kgcat-1 s-1) area specific reaction rate (molCO Pa-2 mCo-2 s-1) activation energy (kJ/mol)

14.3/14.3 0.041 16 50 - 140 5.09 0.319 0.127 3.96 × 10-4 0.778 × 10-4 110

diffusion lengths for the powder catalyst, one can assume that pore diffusion does not influence the observed reaction rate.22 Table 5 summarizes the rate constants for the powder and crushed monolithic catalysts in stirred-tank reactors and the monolithic catalyst in fixed-bed reactors at 190 and 200 °C. 1 1 1 ) + kobs kaov kint

Crushed Monolith at 200 °C

(11)

These simple calculations show that mass-transfer resistances obviously influence the observed reaction rate. Comparison between a crushed monolithic catalyst in a stirred-tank reactor and a monolithic catalyst in a fixed-bed reactor with the same active metal surface area show a higher influence of mass transfer on the observed reaction rate in the stirred-tank reactor. At comparable intrinsic reaction rates, the observed reaction rate in monolith reactor is significantly higher. These differences are more pronounced at the higher reaction temperature of 200 °C. The time constants for powder catalyst in the stirred-tank reactor also are summarized in Table 5. These results also show

at 190 °C

0.509 0.124 6.31 × 10-4 1.24 × 10-4

Monolith

at 200 °C

14.6/65.0 0.186 50 350 4.08 0.268 0.133 3.32 × 10-4 0.814 × 10-4 81.8

at 190 °C

at 200 °C

10.1 120 0.393 0.138 4.87 × 10-4 1.19 × 10-4

4.08 0.363 0.145 4.51 × 10-4 1.11 × 10-4 138

0.650 0.147 8.07 × 10-4 1.98 × 10-4

Table 5. Constants for Mass Transfer, Observed and Intrinsic Reaction Rate Value (s-1) reaction rate constant

at 190 °C

at 200 °C

Monolithic Catalyst kobs kaov kint

0.0160 0.1410 0.0181

0.0342 0.1420 0.0451

Powder Catalyst kobs kaov kint

0.0156 0.0737 0.0197

0.0285 0.0740 0.0462

Crushed Monolithic Catalyst kobs kaov kint

0.0127 0.0378 0.0191

0.0198 0.0381 0.0414

a higher influence of mass-transfer resistances on the observed reaction rate, compared to the monolithic catalyst in a fixed-

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bed reactor. Furthermore, very similar intrinsic reaction rates as calculated for the monolithic catalysts are obtained. The influence of external mass-transfer resistances on the observed reaction rate can also be seen from the apparent activation energies, which were calculated from the first-order reaction rate constants (kobs). The results range from 81.8 kJ/ mol for the crushed monolithic catalyst, over 110 kJ/mol for the powder catalyst to 138 kJ/mol for the monolithic catalyst. The influence of mass-transfer resistances in the stirred-tank reactor was also investigated experimentally by changing the stirrer speed between 500 min-1 and 1500 min-1. No significant increase of conversion could be observed during these additional experiments. However, the enhancement in mass transfer by increasing the specific power input is very limited. With eq 9, it can be calculated that kaLS,STR increases by 20% if the power input is enhanced by a factor of 3. This higher mass transfer rate leads to an improvement of the overall rate constant kaov,STR of ∼5% only. Therefore, it can be concluded that the intrinsic kinetics cannot be directly measured with the stirred-tank reactor used in the present study. 4. Conclusions Monolithic and powder catalysts for Fischer-Tropsch synthesis based on cobalt (18.6 ( 0.9 wt %) and rhenium (1.2 ( 0.1 wt %) on γ-Al2O3 have been prepared. The monolithic catalysts were obtained by dip coating in an aqueous suspension of base powder catalyst and colloidal binder followed by a calcination step. The washcoat thickness was adjusted by repeating this procedure. It was found that the washcoat mass fraction achieved in each coating step decreases with the number of steps. This can be explained by decreasing surface area during coating and lower differences in zeta potential between the slurry particles and the monolith carrier. After six coating steps, a washcoat mass fraction of ∼0.20 gwc/gcat was achieved. The final powder catalyst was prepared by drying and calcination of the same catalyst slurry. The composition of the monolithic washcoat and powder catalyst is equal, because the same slurry was used for both preparations. It could also be shown that the pore structure and the internal surface area of the final powder catalyst and the monolith washcoat are very similar. Some differences were found in the reduction behavior and the active metal dispersion by temperature-programmed reduction (TPR) and CO chemisorption measurements, which might be caused by different overall calcination times. The comparability of powder catalyst and monolith washcoat may be further increased by preparing the final powder catalyst in different layers on a flat support in the same manner as the washcoat. These layers can afterward be removed from the support, crushed, and sieved to the desired fraction. During catalytic measurements with powder and crushed monolithic catalysts in the stirred-tank reactor, comparable methane selectivities and reaction rate constants could be observed. Larger deviations at higher reaction temperatures are probably caused by pore diffusion effects in a fraction of the crushed monolithic catalyst with higher diffusion length. Measurements with monolithic catalysts in fixed-bed operation mode show a higher reaction rate and activation energy at similar methane selectivities, compared to the powder catalyst. This reaction rate enhancement is most probably caused by the advantageous mass-transfer characteristics of the monolithic catalyst in the slug-flow regime that was, for the first time, experimentally realized. In conclusion, the present study shows that the preparation of powder and monolithic catalysts with identical physical and

catalytic properties is challenging. However, the achieved properties of the catalysts are very similar and provide a good basis for a more detailed and systematic comparison of powder and monolithic catalysts used for Fischer-Tropsch synthesis. Moreover, the results obtained with the monolithic catalyst reveal that structured catalysts are a very promising alternative to suspended powder catalysts. Nomenclature Parameters a ) specific surface area (m-1) Ca ) capillary number d ) diameter (m) D ) diffusion coefficient (m2/s) f ) specific calibration factor k ) reaction rate constant (s-1) ka ) mass transfer coefficient (s-1) OFA ) open frontal area p ) pressure (Pa) R ) peak area ratio in TCD Re ) Reynolds number S ) selectivity Sc ) Schmidt number u ) superficial velocity (m/s) V ) volume (m3) x ) weight fraction X ) conversion y ) volume fraction. Greek Symbols ε ) porosity  ) specific power input (W/kg) δ ) thickness (m) η ) dynamic viscosity (Pa s) F ) density (kg/m3) σ ) surface tension (N/m) Indexes 0 ) inlet conditions c ) capillary cat ) catalyst cell ) cell of the monolithic carrier ch ) channel char ) characteristic GL ) gas-liquid GLS ) gas-liquid-solid GS ) gas-solid f ) film int ) intrinsic L ) liquid LS ) liquid-solid mc ) monolithic carrier mcat ) monolithic catalyst MR ) monolithic reactor obs ) observed ov ) overall p ) particle pore ) pore SG ) synthesis gas STR ) stirred tank reactor TP ) two-phase wc ) washcoat

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ReceiVed for reView March 7, 2008 ReVised manuscript receiVed June 6, 2008 Accepted June 6, 2008 IE800377N