Catalytic performance on iron-based Fischer–Tropsch catalyst in fixed-bed and bubbling fluidized-bed reactor

Applied Catalysis B: Environmental 103 (2011) 169–180

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Applied Catalysis B: Environmental journal homepage: www.elsevier.com/locate/apcatb

Catalytic performance on iron-based Fischer–Tropsch catalyst in fixed-bed and bubbling fluidized-bed reactor Suk-Hwan Kang a , Jong Wook Bae b,c,∗ , Joo-Young Cheon b , Yun-Jo Lee b , Kyoung-Su Ha b,∗∗ , Ki-Won Jun b , Dong-Hyun Lee c , Byung-Woo Kim c a

Plant Engineering Center, Institute for Advanced Engineering (IAE), Suwon, Kyonggi-do 443-749, Republic of Korea Petroleum Displacement Technology Research Center, Korea Research Institute of Chemical Technology (KRICT), PO Box 107, Sinseongno 19, Yuseong, Daejeon 305-600, Republic of Korea c School of Chemical Engineering, Sungkyunkwan University, Suwon, Kyonggi-do 440-746, Republic of Korea b

a r t i c l e

i n f o

Article history: Received 14 March 2010 Received in revised form 17 December 2010 Accepted 14 January 2011 Available online 21 January 2011 Keywords: Fischer–Tropsch synthesis Iron-based catalyst Light olefin Clean fuel Fixed-bed reactor Bubbling fluidized-bed reactor

a b s t r a c t Fischer–Tropsch synthesis (FTS) for the co-production of C2 –C4 olefins and clean fuels such as gasoline and middle distillate from syngas was investigated on four different iron-based catalysts in a fixed-bed and a bubbling fluidized-bed reactor. The catalysts were prepared by wet-impregnation using the Al2 O3 , SiO2 and iron ore (FeOx ) supports with the active components of Fe, K and (or) Cu, and each K/FeCuAlOx catalyst was prepared by a co-precipitation method. Except for the catalyst impregnated on SiO2 , CO conversion of the other catalysts is similar regardless of the type of reactor. However, the yield of light olefins in the range of C2 –C4 hydrocarbons and clean fuels is closely related to the formation of active iron carbide species and to surface acidity. The impregnated K/FeOx catalyst is found to be one of the promising catalysts to be applied in a bubbling fluidized-bed reactor for middle temperature FTS reaction due to its high resistance to catalyst attrition with a high catalytic performance. © 2011 Elsevier B.V. All rights reserved.

1. Introduction With the increase in the cost of petroleum products due to the depletion of oil reserves, an alternative method for synthesizing hydrocarbons through Fischer–Tropsch synthesis (FTS) has again received considerable attention. FTS is an important technology in the production of clean liquid fuels and valuable chemicals from syngas derived from the gasification of coal or biomass and the reforming of natural gas or other carbon-containing components [1–3]. The general production process of light olefins and clean fuels such as gasoline (C4 –C12 hydrocarbons) and middle distillate (C10 –C20 hydrocarbons) is composed of FTS reaction and subsequent cracking process of wax component [4]. These processes are significantly important to produce sulfur-free clean fuels due to the nature of low content of sulfur and nitrogen components in the syngas. In FTS process, the reactor is chosen appropriately

∗ Corresponding author at: Petroleum Displacement Technology Research Center, Korea Research Institute of Chemical Technology (KRICT), PO Box 107, Sinseongno 19, Yuseong, Daejeon 305-600, Republic of Korea. Tel.: +82 42 860 7383; fax: +82 42 860 7388. ∗∗ Corresponding author. Tel.: +82 42 860 7286; fax: +82 42 860 7388. E-mail addresses: fi[email protected] (J.W. Bae), fi[email protected] (K.-S. Ha). 0926-3373/$ – see front matter © 2011 Elsevier B.V. All rights reserved. doi:10.1016/j.apcatb.2011.01.024

according to the targeted products such as light olefin, gasoline, diesel or heavy hydrocarbons including wax, and the process also requires the appropriate selection of catalyst preparation method depending on the type of reactor, such as plug flow (fixed-bed or supercritical), fluidized-bed or slurry bubble column reactor [5,6]. In the development of FTS reactor, obtaining the optimum temperature gradient in catalyst-bed by adopting the proper cooling method, such as co-current flow of coolant in plug flow reactor, is critical to eliminate hot-spot formation. Even though the backmixing of catalyst and reactant gases in fluidized-bed or bubble column reactor can happen and will result in reducing the reactor efficiency, the appropriate design of catalyst particle size and formulation in fluidized-bed (or entrained flow) reactor or liquid circulation in bubble column reactor can efficiently provide a well-controlled isothermal temperature profile in each reaction system. In addition, the multi-stage cooling method with a cooling liquid along the length of the catalyst-bed is one of the alternative options for the removal of generated heat. In particular, it has been learned that the fluidized-bed reactor can be utilized efficiently for various multiphase reactions, because it is characterized by higher heat and mass transfer rates due to the effective contacting and mixing property in comparison with other plug flow systems [7–9]. The fluidized-bed reactor has also been proposed as an excellent process in view of the high conversion of syngas

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and selectivity for valuable chemicals through high temperature FTS reaction with an economically feasible process in multiphase reactions [10,11]. Several types of metal-incorporated catalysts can be used for FTS reaction; however, only iron- and cobalt-based catalysts appear to be economically feasible on industrial application. Although cobalt-based catalysts mainly produce valuable high molecular weight straight-chain hydrocarbons, iron-based catalysts are more useful when the H2 /CO ratio is low (because of its characteristics of water–gas-shift activity on iron-based catalyst), these iron-based catalysts are also useful for producing alkenes, oxygenates, and branched hydrocarbons, depending upon the kind of promoters, reaction conditions and reactor types employed. Especially, FTS reaction in fluidized bed reactor has been mainly using iron-based catalyst which has been prepared by spray drying or wet-impregnation method because of the low attrition property of those catalysts [12,13]. For iron-based FTS catalysts, various promoters are used to enhance the reducibility of iron species, to increase chain growth probability of hydrocarbons, to increase the extent of active species like iron carbides and to enhance the catalytic stability during FTS reaction. Copper or manganese or potassium can be one of the representative promoters, and additional structural promoters such as Al2 O3 or SiO2 are included with an appropriate quantity to impart structural stability as well as to improve mechanical properties [12,13]. The effect of the concentration of K promoter on iron-based catalyst was also investigated from our previous works [14], and the results revealed that catalytic performance was not much altered above 6 wt% K with a trivial variation of C2 –C4 olefin. The general role of K is known to be enhancing the reducibility of iron oxides and olefin selectivity, however, the high concentration of K is responsible for the suppressed catalytic performance. In particular, iron-based catalyst for high temperature Fischer–Tropsch (HTFT) reaction around 330 ◦ C, which were prepared by the fusion of iron oxide together with the chemical promoter, K2 O, and structural promoters such as MgO or Al2 O3 , have been used for many years by Sasol company [15]. In the present investigation, FTS reaction with the various ironbased catalysts for the co-production of light hydrocarbons and clean fuels from syngas at middle temperature Fischer–Tropsch (MTFT) around 300 ◦ C was carried out in a fixed-bed reactor (FBR) and a bubbling fluidized bed reactor (BFBR) to find out efficient production method of clean chemicals without adopting the additional wax cracking process. Even though many researchers have tried to find out proper catalytic system intensively with respect to finding an optimum catalyst composition in a fixed-bed reaction system, the present paper is focused on how to develop the proper catalytic system in both the type of supports and reactors by considering the characteristics of reactor performance. The ironbased catalyst was prepared by a conventional wet-impregnation method using an aqueous solution containing Fe–Cu–K metal precursors over support such as Al2 O3 and SiO2 , by the co-precipitation of Fe–Cu–Al metal precursors with sodium carbonate as precipitate, and by the wet-impregnation of potassium carbonate over iron ore. The objective of this work is to find out the influence of the preparation method of FTS catalysts on the physicochemical properties such as the pore structure and the acidity of catalysts, and on its catalytic performance in FTS reaction from syngas in FBR and BFBR. The properties of those catalysts are further substantiated by using the characterization tools before and after FTS reaction and the different catalytic activity values are systematically compared in both FBR and BFBR. In addition, the characteristics of two different FTS rectors are compared in the view of catalysts prepared by different methods for MTFT reaction, and finally we derive the proper catalytic system to obtain high catalytic performance as well as to increase the production of useful olefin chemicals and clean fuels.

2. Experimental 2.1. Catalyst preparation Iron-based FTS catalysts adopted in the present study were prepared by using the conventional coprecipitation and wetimpregnation methods to evaluate the effects of supports and preparation methods simultaneously and finally find our proper catalytic system to obtain the high selectivity to useful chemicals such as light olefins in the range of C2 –C4 hydrocarbons and clean fuels. (1) FeCuK/␥-Al2 O3 catalyst was prepared by the conventional wetimpregnation method using ␥-Al2 O3 support. The ␥-Al2 O3 support was obtained by calcination of boehmite (high purity of Catapal-B, Condea) in air at 600 ◦ C with surface area of around 231 m2 /g and pore volume is around 0.47 cm3 /g. Aqueous iron nitrate, copper nitrate and potassium carbonate solutions with the required composition were mixed thoroughly. Since the pH of metal solutions has approached 4.0 during catalyst preparation, it can affect the structure of the alumina with the structure collapse. Thus, during the preparation of FeCuK/␥-Al2 O3 catalyst, the pH of metal precursor solutions in the slurry of ␥-Al2 O3 support was maintained at 7.0 using (NH3 )4 OH solution. The active metal impregnation on ␥-Al2 O3 was carried out with continuous stirring at room temperature for 12 h. Each sample was then dried in a rotary evaporator before subjecting it for calcination at 500 ◦ C in air for 5 h. The weight ratio of total iron metal components to ␥-Al2 O3 on the finished catalyst was fixed at 20/100 with the weight ratio of Fe/Cu/K as 1/0.1/0.2. The final catalyst was denoted as FeCuK/Al2 O3 with the composition of 15.9 wt% Fe/1.6 wt% Cu/3.2 wt% K/79.3 wt% Al2 O3 . (2) FeCuK/SiO2 catalyst was prepared by the conventional wetimpregnation method on SiO2 support. The SiO2 (Davisil grade 645, Aldrich) with surface area of 336 m2 /g and pore volume of 1.11 ml/g was used as support. Aqueous iron nitrate, copper nitrate and potassium carbonate solutions, with the required composition, were mixed thoroughly. The active metal impregnation on SiO2 was carried out with continuous stirring at room temperature for 12 h. Each sample was then dried in a rotary evaporator before subjecting it for calcination at 500 ◦ C in air for 5 h. The weight ratio of total iron metal components to SiO2 on the finished catalyst was fixed at 20/100 with the weight ratio of Fe/Cu/K as 1/0.1/0.2. The final catalyst was denoted as FeCuK/SiO2 with the composition of 15.9 wt% Fe/1.6 wt% Cu/3.2 wt% K/79.3 wt% SiO2 . (3) K/FeCuAlOx catalyst (100 Fe:6.6 Cu:15.7 Al by weight ratio) was prepared by using a coprecipitation method with ammonium hydroxide solution as precipitant dissolved in the deionized water. The metal nitrate solution was composed of iron nitrate, copper nitrate and aluminum nitrate and the precipitate was obtained at pH 6.8 by dropping the above two solutions separately. The precipitate was dried in air at 110 ◦ C for 24 h and potassium carbonate (K2 CO3 ) was impregnated on the dried sample by the wet-impregnation method with the fixed weight ratio of 100/4 (Fe/K). The catalyst was dried again at 110 ◦ C for 48 h and subsequently calcined in air at 500 ◦ C for 5 h. The final catalyst was denoted as K/FeCuAlOx with the final composition of 3.2 wt% K/79.2 wt% Fe/5.2 wt% Cu/12.4 wt% Al. (4) K/FeOx catalyst without adding Cu promoter was prepared by the conventional wet-impregnation method with K2 CO3 on iron ore (FeOx ) which is composed of 64.5 wt% iron metal. The iron ore contains iron metal around 64.5 wt% in the form of Fe3 O4 and Fe2 O3 and 3.5 wt% SiO2 , 3.5 wt% Al2 O3 and 0.9 wt% TiO2 with other trivial impurities such as P and S below 0.08 wt%. The impregnation of potassium precursor on iron ore

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Table 1 Physical properties of the iron-based FTS catalysts. Catalystsa

FeCuK/Al2 O3 FeCuK/SiO2 K/FeCuAlOx K/FeOx

Preparation method and composition (wt. ratio of catalyst components)

Impregnation (20/2/4/100) Impregnation (20/2/4/100) Coprecipitation (4/100/6.6/15.7) Impregnation (6/94)

Surface area (m2 /g)

Pore volume (cm3 /g)

Average pore diameter (nm)

Apparent density (g/cm3 )

Bulk density (g/cm3 )

SEM/EDAX wt. ratio (before)b

SEM/EDAX wt. ratio (after)b

K/Fe

Cu/Fe

K/Fe

Cu/Fe

192.1

0.264

4.4

1.53

0.84

0.108

0.097

0.179

0.115

157.8

0.678

13.0

0.84

0.49

0.134

0.116

0.180

0.116

89.5

0.171

5.9

2.43

1.39

0.038

0.077

0.043

0.075

0.6

0.010

14.4

3.81

2.23

0.138

0

0.116

0

FeCuK/Al2 O3 and FeCuK/SiO2 catalysts were prepared by conventional impregnation method at a fixed weight ratio of Fe/Cu/K as 1/0.1/0.2 on ␥-Al2 O3 and SiO2 support with a fixed amount of Fe/support at 20/100. K/FeCuAlOx catalyst was prepared by coprecipitation method using K2 CO3 (Fe:K = 100:4 wt%) and K/FeOx catalyst was made by wet-impregnation of K2 CO3 on iron ore (FeOx ) with K/Fe weight ratio of 6/94. b The chemical composition of FTS catalysts before and after reaction was characterized by using SEM/EDAX analysis and the weight ratio of K/Fe and Cu/Fe was displayed. a

was carried out with continuous stirring at room temperature for 12 h. The sample was then dried in a rotary evaporator before subjecting it for calcination at 500 ◦ C in air for 5 h. The weight ratio of potassium metal to iron ore (K/FeOx ) in the finished catalyst was fixed at around 6/94 (wt/wt). The final catalyst was denoted as K/FeOx . The summarized physical properties of all four catalysts are given in Table 1.

2.2. Catalytic performance tests Catalytic performance on the four different iron-based FTS catalysts was studied in FBR (I.D. = 10.7 mm) equipped with electric heater with a catalyst amount of 0.3 g possessing a particle size of 53–120 ␮m. In addition, catalytic performance was also carried out in BFBR (0.034 m I.D. × 1.5 m height) equipped with the perforated type of gas distributor and the reaction temperature was controlled by using electric heater. The schematic reaction apparatus for FBR and BFBR is displayed in Fig. 1 ((A) for FBR and (B) for BFBR). The hot trap fixed at 100 ◦ C for separating heavy wax components from unreacted syngas with C1 –C8 hydrocarbons was also equipped in both reactors [6,9]. The catalyst materials possessing the same particle size distribution as that adopted in FBR were loaded with around 100 g in BFBR equipped with a perforated plate containing 11 evenly spaced holes of 1.0 mm diameter, which served as a reactant distributor. Prior to the FTS reaction in both reactors, the catalyst was reduced at 450 ◦ C for 24 h in a flow of 5% H2 balanced with He. After reduction, the synthesis gas (H2 /CO = 2) was fed into the reactors. The FTS reaction was carried out subsequently under the following reaction conditions; T = 300 ◦ C, P = 1.0 MPa and UG (linear velocity of syngas) = 0.2, 0.4, 1.0, 2.0, 4.0, 8.0 and 10.0 cm/s for BFBR and T = 300 ◦ C, P = 1.0 MPa and space velocity (SV) = 2000 ml/gcat /h for FBR. The effluent gas from the reactor was analyzed by using an online gas chromatograph (YoungLin Acme 6000 GC) employing a GS-GASPRO capillary column connected with flame-ionization detector (FID) for analysis of hydrocarbons and Carboxen-1000 packed column connected with thermal conductivity detector (TCD) for the analysis of carbon oxides, hydrogen, and methane and using Ar as an internal standard component with 5 vol.% based on syngas. The CO conversion and the degree of CO2 formation were directly calculated from the results of TCD analysis, and the product distribution in the range of C1 –C8 was analyzed by using FID. The hydrocarbon selectivity above C9 + is calculated based on carbon balance and ASF (Anderson–Schulz–Flory) distribution with the additional analysis of wax components corrected in hot-trap by FID.

2.3. Catalyst characterization The BET surface areas were estimated from nitrogen adsorption isotherm data obtained at −196 ◦ C on a Micromeritics, ASAP-2400 equipment. The calcined samples were degassed at 250 ◦ C in a He flow for 4 h before the measurements. The pore volumes and pore size distribution of the catalysts were determined by BJH (Barett–Joyner–Halenda) model from the data of the desorption branch of the nitrogen isotherms. The pore volumes were determined at a relative pressure (P/P0 ) of 0.99. The particle size distributions of K/FeCuAlOx and K/FeOx before and after FTS reaction were measured using MASTERSIZER 2000 (Malvern Instruments). Before the measurement, the sample was added to anhydrous ethanol and sonicated for 10 min to disperse the catalyst particles in the solvent. The powder X-ray diffraction (XRD) patterns were obtained with a Rigaku diffractometer using Cu-K( radiation to identify the phases of iron-based catalysts and their crystallinity. The temperature programmed reduction (TPR) experiments were performed to determine the reducibility of the surface iron oxides. Prior to TPR experiments, the samples were pretreated in a He flow up to 400 ◦ C and kept for 2 h to remove the adsorbed water and other contaminants, followed by cooling to 50 ◦ C. The reducing gas containing 5% H2 /Ar mixture was passed over the samples at a flow rate of 30 ml/min, with the heating rate of 10 ◦ C/min up to 800 ◦ C, and kept at that temperature for 30 min. The effluent gas was passed over a molecular sieve trap to remove the generated water and analyzed with a GC equipped with TCD. The temperature programmed reduction with CO (CO-TPR) to confirm the degree of carbide formation and reducibility on calcined FTS catalyst was also carried out. Prior to CO-TPR experiments, the samples were pretreated in a He flow up to 400 ◦ C and kept for 2 h, followed by cooling to 100 ◦ C. The reducing gas of CO was passed over the samples at a flow rate of 15 ml/min, with the heating rate of 5 ◦ C/min up to 1000 ◦ C. The effluent gas was analyzed by using a mass-spectrometer equipped with a quadrupole mass detector (GENESYS 42) for monitoring CO2 evolution (m/z = 44). The temperature programmed surface reaction (TPSR) with H2 on the reacted FTS catalysts was tested on FBR for 70 h on stream to confirm the content and kind of deposited hydrocarbons on the catalyst surface with the same conditions of CO-TPR experiment by monitoring CH4 evolution (m/z = 15) instead of another CH4 fragment (m/z = 16) due to the overlapping with H2 O fragment. The temperature programmed desorption of ammonia (NH3 TPD) experiments were performed to determine the surface acidity of the iron-based FTS catalysts. About 0.1 g of the sample was

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Fig. 1. Schematic diagrams of reaction apparatus (A) for FBR and (B) for BFBR.

flushed initially with a He flow at 250 ◦ C for 2 h, cooled to 100 ◦ C, and saturated with NH3 . After adsorption of NH3 , the sample was purged with a He flow until equilibrium, and then NH3 TPD experiments were carried out from 50 to 900 ◦ C at a heating rate of 10 ◦ C/min. The effluent gas was directly connected and analyzed with a GC equipped with TCD without passing any trapping materials to eliminate the concomitant removal of NH3 and H2 O. The surface iron species and their electronic state after FTS reaction were further characterized by using X-ray photoelectron spectroscopy (XPS; ESCALAB MK-II) analysis. During the experiment, the Al-K␣ monochromatized line (1486.6 eV) was adopted and the vacuum level was kept around 10−7 Pa, XPS analysis was

carried out with a step size of 0.02 eV and each step was scanned twice. The low S/N ratio may be due to the contamination of catalyst surface with the deposited hydrocarbon components. The powder samples after FTS reaction were previously passivated with 1% O2 /He gas and pressed in air atmosphere into thin pellets before characterization and the binding energy (BE) was corrected with the reference BE of C 1s (284.4 eV). Furthermore, the carbon species formed during FTS reaction is also compared according to catalyst types by evaluating C 1s peaks. The fresh and used FeCuK/Al2 O3 and FeCuK/SiO2 catalysts were characterized by using the transmission electron microscopy (TEM; TECNAI G2 instrument) to compare the particle size variation during FTS reaction. The chemical compositions of fresh and used

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Fig. 3. TPR profiles of iron-based FTS catalysts. Fig. 2. Pore size distribution of iron-based FTS catalysts.

catalysts used in BFBR reaction are further characterized by using SEM/EDAX analysis. 3. Results The overall FTS reaction can be simplified as the combination of paraffin formation and olefin formation with the water–gas shift (WGS) reaction: Paraffinformation : nCO + (2n + 1)H2 ⇔ Cn H2n+2 + nH2 O

(1)

Olefinformation : nCO + 2nH2 ⇔ Cn H2n + nH2 O

(2)

WGSreaction : CO + H2 O ⇔ CO2 +H2

(3)

The mechanism of FTS reaction on iron-based catalyst is still not clear due to the complex phase changes of iron species starting from iron metal precursor or iron oxides and subsequent phase transformation such as the formation of metallic iron during reduction step and the formation of various active iron carbides during FTS reduction [1–3,16–20]. In addition, several investigators suggest that the formation of iron carbide species is responsible for obtaining high FTS activity, and that the iron oxides such as magnetite (Fe3 O4 ) and hematite (␣-Fe2 O3 ) are the most active phases for WGS reaction [17]. The formation of CO2 by WGS is proportional to the partial pressure of H2 O produced during FTS reaction [18]. Thus, at high CO conversion, the formation rate of CO2 generally increases up to the equilibrium level. In order to produce clean fuels from FTS wax products, cracking process with or without feeding of hydrogen is generally employed using solid acid catalysts [4]. 3.1. Textural properties of calcined iron-based FTS catalyst The textural properties such as BET surface area, pore volume and average pore diameter of four different catalysts prepared using different preparation methods are summarized in Table 1 and their pore distribution patterns are also shown in Fig. 2. The catalytic performances during FTS reactions have been generally influenced by their characteristics, such as textural properties, extent of active species formation and acidic sites, which are generally monitored to elucidate the effect of parameters with respect to the activities and selectivities [18]. The larger surface areas were observed on the impregnated catalysts: such as FeCuK/Al2 O3 and FeCuK/SiO2 around 192.1 and 157.8 m2 /g separately with a large pore volume. The lower average pore diameter on the impregnated FeCuK/Al2 O3 and co-precipitated K/FeCuAlOx catalysts was observed around 4.4 and 5.9 nm respectively. The lower average

pore diameter on FeCuAl/Al2 O3 catalyst with high surface area is mainly attributed to the presence of bimodal pore structure with the co-existence of micro and mesopores in the regions of 3–5 nm and 10–20 nm. However, K/FeOx catalyst, which does not include copper promoter, prepared by using iron ore with impregnation method showed the lowest surface area of 0.6 m2 /g, but the highest average pore diameter of around 14.4 nm. The apparent and bulk densities of four different catalysts are in the following order; K/FeOx > K/FeCuAlOx > FeCuK/Al2 O3 > FeCuK/SiO2 . These textural properties are significantly influencing the behavior of fluidization and mass-transfer efficiency and eventually will alter the catalytic performances, especially in BFBR. 3.2. Temperature-programmed analysis (H2 -TPR and NH3 -TPD) The reduction behavior on the four different FTS catalysts was investigated by H2 -TPR and is shown in Fig. 3. Except for K/FeOx catalyst, the reduction patterns of the other catalysts display two sharp peaks below 550 ◦ C and one broad peak above 550 ◦ C. In the case of copper promoted iron-based catalysts, the reduction process under H2 atmosphere occurs in two distinctive stages. The first stage, occurring between 250 and 370 ◦ C, corresponds to the combined reduction of hematite (␣-Fe2 O3 ) to magnetite (Fe3 O4 ) and CuO to Cu [19]. Cu is used as a promoter for the iron-based FTS catalysts to facilitate easy reduction of iron oxides on the catalyst employed for the production of gasoline-range hydrocarbons. Cu not only decreases the reduction temperature of iron oxides but also decreases the operating temperatures, compared to catalyst prepared without adding Cu promoter. In addition, co-promoters of Cu and K are generally known to stabilize FTS activity and selectivity [20]. The first reduction stage, at the somewhat high temperature region of 370–550 ◦ C, represents the reduction of magnetite (Fe3 O4 ) to FeO. The second reduction stage at the temperature region above 550 ◦ C represents reduction of FeO to metallic Fe. In the first reduction stage, the reduction temperature on the impregnated FeCuK/Al2 O3 and FeCuK/SiO2 catalysts showed lower values than that of the coprecipitated K/FeCuAlOx catalyst. The maximum peak temperatures are around 367 and 471 ◦ C on K/FeCuAlOx catalyst. However, K/FeOx catalyst, which does not contain Cu promoter and is generally adopted in the commercial process with the composition of fused iron oxides and the promoters such as K, Mg and Al, showed different reduction behavior. As shown in Fig. 3, K/FeOx catalyst is reduced at much higher temperature (above 600 ◦ C) compared to the other catalysts and this fact reveals the bulk reduction of FeOx to metallic Fe. The reduction of the supported iron-based FTS catalysts on Al2 O3 or SiO2 is known to be

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S.-H. Kang et al. / Applied Catalysis B: Environmental 103 (2011) 169–180 Table 2 The amount of active acidic sites on four different FTS catalysts measured by NH3 TPD.a Catalysts

FeCuK/Al2 O3 FeCuK/SiO2 K/FeCuAlOx K/FeOx

Active acidic site (mmol NH3 /gcat ) Weak (first peak; I)

Strong (second peak; II)

Total

0.180 0.099 0.022 0.018

1.260 0.661 0.588 0.532

1.440 0.760 0.610 0.550

a The desorption peak of NH3 in Fig. 4 is separated in three parts in the regions of 140–300 ◦ C (I), 300–550 ◦ C (II) and above 550 ◦ C (III). The desorption peaks appearing at 140–550 ◦ C (peaks I and II) are assigned to the active acid sites (total acidic sites) and peak (III) is not considered as an active site due to the possible desorption character of H2 O.

3.3. Catalytic performance in FBR and BFBR Fig. 4. NH3 -TPD profiles over iron-based FTS catalysts.

much difficult due to the strong interaction with supports and the inhibited reduction behavior of hematite to magnetite. These characters are also responsible for the less formation of active iron carbides [5,6,19], however, the coprecipitated K/FeCuAlOx and K/FeOx catalysts also showed a much hard reducibility on our system. It suggests that the high surface concentration of iron oxides on the impregnated FeCuK/Al2 O3 and FeCuK/SiO2 catalysts is responsible for showing lower reduction temperature than that of coprecipitated K/FeCuAlOx . The H2 uptakes of FeCuK/Al2 O3 , FeCuK/SiO2 , K/FeCuAlOx and K/FeOx catalysts are 3.39, 2.04, 5.80 and 5.59 mmol H2 /gcat respectively. But it is not possible to compare the degree of reduction because the metallic Fe content of each catalyst is different. Even though the facile reducibility was observed on FeCuK/Al2 O3 , FeCuK/SiO2 , and K/FeCuAlOx catalysts compared to that of K/FeOx catalyst, the different phase transformation to the active iron carbide species on the surface during FTS reaction should be chiefly responsible for showing the different catalytic performance, as explained in the following section (CO-TPR, XPS and TPSR experiments). The NH3 -TPD patterns of all catalysts are shown in Fig. 4. Three stage desorption of NH3 is observed in the regions of 140–300 ◦ C (I), 300–550 ◦ C (II) and above 550 ◦ C (III) respectively. The three distinctive desorption peaks of the NH3 which was analyzed by TCD were assigned to different acidic sites on the catalyst surface [21–23]. The desorption peak below 300 ◦ C corresponds to the acid sites contributed by the matrix of Al2 O3 or SiO2 sites, while the second peak between 300 and 550 ◦ C is attributed to the catalytic species present on the surface. Moreover, the higher desorption peak temperature above 550 ◦ C can probably be attributed to the amount of H2 O evolution from metal oxide component and framework of Al2 O3 or SiO2 ; this dependence was previous confirmed by our investigation using a mass spectrometer [23]. Therefore, the desorption peaks appearing in the range of 140–550 ◦ C (peaks I and II) are assigned to the active acid sites induced from acidic components and catalytic species on Al2 O3 or SiO2 support and its content are probably correlated with the active acidic sites. Table 2 gives the number of acid sites expressed as mmol NH3 /gcat with respect to two categories of weak and strong acidic sites. It can be observed that the number of weak acid sites and the number of strong acid sites are found to be lowest on K/FeOx catalyst. These trends can be related to the hydrocarbon selectivity, especially in the range of C1 –C4 , on the four different catalysts. The content of total acidic sites on four different catalysts is in the following order: FeCuK/Al2 O3 > FeCuK/SiO2 > K/FeCuAlOx > K/FeOx .

3.3.1. Catalytic performance in FBR The catalytic performance of the four different iron-based catalysts in FBR was investigated at H2 /CO = 2, SV = 2000 ml/gcat h, P = 1.0 MPa and T = 300 ◦ C using 0.3 g of catalyst. The activity and selectivity to hydrocarbons were tested for over 70 h. The results are summarized in Table 3, where the yields to olefins in the range of C2 –C4 hydrocarbons are also included. All catalysts, except for FeCuK/SiO2 catalyst which showed a rapid deactivation from 85.5% to 37.2%, have shown remarkably higher initial activity with stability, as shown in Fig. 5. The catalytic activity was stabilized after 15 h on stream, except for K/FeOx catalyst. Due to its low reducibility this catalyst had a longer induction period of around 30 h starting from lower CO conversion around 20%. With respect to the selectivity of hydrocarbons, K/FeCuAlOx and K/FeOx catalysts show higher selectivity to C2 –C4 olefins of 79.5 and 85.7% respectively, whereas FeCuK/Al2 O3 catalyst shows lowest C2 –C4 olefin selectivity of 52.6% and FeCuK/SiO2 catalyst shows an intermediate value. The ␥-Al2 O3 in FeCuK/Al2 O3 catalyst plays significant roles for acidic component as well as structural promoter. The existence of abundant strong acidic sites on FeCuK/Al2 O3 catalyst as shown in Table 2, can be the reason for the suppressed C2 –C4 olefin formation with a high CH4 selectivity by the enhanced production of isomers and aromatics in the C12 + hydrocarbon range, at the expense of olefins formed during FTS reaction [1–3], or by the possible hydrogenation activity of C2 –C4 olefins to form paraffinic components. Moreover, the decrease in acidic sites on K/FeCuAlOx catalyst increases olefin

Fig. 5. CO conversion with time on stream on FBR.

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Table 3 FTS reaction over iron-based catalysts in fixed-bed reactor (FBR).a Catalysts

Conversion of CO

Conversion to CO2

FeCuK/Al2 O3 FeCuK/SiO2 K/FeCuAlOx K/FeOx

96.1 37.2 95.4 96.5

37.8 16.1 36.1 33.8

a b c

Selectivity in hydrocarbons C1

C2 –C4

C5 +

14.1 14.2 9.9 7.5

37.4 27.1 30.9 25.9

48.5 58.7 59.2 66.6

O/(O + P) in C2 –C4 (mol%)b

Yield to olefins (mol%)c

52.6 77.1 79.5 85.7

11.5 4.4 14.6 13.9

FTS reaction was carried out under the following reaction conditions; H2 /CO = 2, SV = 2000 ml/gcat h, P = 1.0 MPa and T = 300 ◦ C. O/(O + P) represents the olefin selectivity (C-mol%) defined as the mole percentage of olefin/(olefin + paraffin) in the range of C2 –C4 hydrocarbons. The yield to olefin (mol%) is the reacted CO to olefins in the range of C2 –C4 hydrocarbons after subtracting the extent of CO conversion to CO2 .

selectivity in C2 –C4 hydrocarbons range by the above mentioned reaction. The lower selectivity to C2 –C4 olefins on FeCuK/SiO2 could also be attributed to the higher concentration of strong acidic sites, its fast deactivation may be due to the phase transformation of metallic iron to inactive iron silicate species [24] which significantly interacted with unstable amorphous silica species during FTS reaction (shown in XPS analysis in the following section). Interestingly, high CO conversion and C2 –C4 olefin selectivity were observed on K/FeOx catalyst, even though it does not contain any copper promoter and it possesses a low surface area, due to the facile transformation of surface metallic iron to active iron carbide species (shown in XRD, CO-TPR and TPSR analyses in the following section). Therefore, one can state that a lower concentration of strong acidic sites with a facile formation of iron carbide species should be more preferable factors for increasing C2 –C4 olefin selectivity compared to the effects of physicochemical properties such as surface area and degree of reduction in FBR.

3.3.2. Catalytic performance in BFBR In the present investigation, the catalytic performance on iron-based FTS catalysts in BFBR was also studied at T = 300 ◦ C, P = 1.0 MPa and H2 /CO = 2 to compare with that in FBR at a different superficial gas velocity values (linear velocities of syngas). It is well known that iron-based FTS catalysts exhibit high WGS activity, and hence they are of increasing interest for FTS reaction, even more than cobalt-based catalysts for low H2 /CO feed stocks produced from gasification of coal or biomass. The particle size and density of the four different iron-based FTS catalysts according to the viscosity of syngas have influenced the minimum fluidization velocity (Umf ). The Umf values of FeCuK/Al2 O3 , FeCuK/SiO2 , K/FeCuAlOx and K/FeOx catalysts at T = 300 ◦ C and P = 1.0 MPa are found to be 0.33, 0.19, 0.54 and 0.85 cm/s respectively. Therefore, the catalytic activity measurement in BFBR was carried out over Umf to maximize the fluidization behavior in a catalyst-bed. The catalytic performance on those catalysts was tested for over 20 h at each superficial gas velocity, and the results are summarized in Table 4. All the catalysts except for FeCuK/SiO2 have no direct relation with the superficial gas velocity and show high CO conversion with high catalytic stability. As shown in Table 4, CO conversion on FeCuK/SiO2 decreases remarkably with increasing the superficial gas velocity with the high formation rate of CH4 . This is similar to the result obtained in FBR; the lower activity on FeCuK/SiO2 with high CO2 and CH4 formation is mainly attributed to the difficult transformation of iron oxides to the active iron carbide species [4]. The observed catalytic performance on FeCuK/Al2 O3 in BFBR is similar to the result of FBR with low C2 –C4 olefin selectivity. The K/FeCuAlOx and K/FeOx catalysts that showed a high catalytic performance with respect to selectivity to C2 –C4 olefins and CO conversion in FBR are also found to be showing a high catalytic performance in BFBR. With the increase of superficial gas velocity above Umf , the yield to C2 –C4 olefins increased marginally without a significant loss of catalytic activity as shown in Table 4. Although the above two catalysts are adequate for applying to BFBR due to their high cat-

alytic performance compared to that of supported catalysts such as FeCuK/Al2 O3 and FeCuK/SiO2 , further consideration should be given to the resistance to catalyst attrition during FTS reaction. 3.4. Characteristics of iron-based catalysts In order to further understand the characteristics of iron-based FTS catalysts, we carried out XRD studies on the four different catalysts collected before and after FTS reaction and the results are shown in Fig. 6. The freshly calcined iron-based FTS catalysts showed hematite (␣-Fe2 O3 ) as the major detectable iron phase, with the characteristic peaks appearing at 2 values of 24.2◦ , 33.1◦ , 35.6◦ , 40.8◦ , 49.5◦ , 54.0◦ , 57.6◦ , 62.5◦ and 64.0◦ ; these values agree

Fig. 6. XRD patterns of iron-based FTS catalysts (A) before and (B) after FTS reaction.

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Table 4 FTS reaction over iron-based catalysts in bubbling fluidized-bed reactor (BFBR).a Catalyst

FeCuK/Al2 O3

FeCuK/SiO2

Superficial gas velocity, Ug (cm/s) CO conversion (%) Selectivity (C-mol%) CO2 –CH–b Hydrocarbons distribution (C-mol%) C1 C2 C2 C3 C3 C4 C4 C5 –C8 C9 –C11 C12 + Olefin in C2 –C4 (C-mol%) Selectivity: O/(O + P)c % yieldd

0.4 98.4

2.0 95.8

0.2 98.1

1.0 60.7

4.0 53.0

1.0 97.2

4.0 95.9

8.0 96.1

1.0 97.5

4.0 95.2

10.0 96.1

12.2 87.8

18.7 78.7

28.7 71.3

39.8 60.2

37.7 62.3

32.4 67.6

34.1 65.9

34.2 65.8

30.0 70.0

31.5 68.5

34.2 65.8

20.3 10.8 0.0 16.6 0.1 11.8 1.7 15.6 10.0 13.1

24.0 11.7 0.1 14.1 3.7 7.9 6.8 12.3 8.2 11.2

34.0 11.7 0.9 9.7 5.4 6.2 4.6 9.8 8.1 9.6

32.6 7.7 3.0 3.9 8.1 2.7 5.3 12.8 9.6 14.3

33.6 4.8 4.4 2.6 6.2 1.7 4.0 17.8 9.8 15.1

17.7 5.3 5.0 3.2 13.2 2.5 14.4 14.6 9.2 14.9

16.3 4.3 5.9 2.2 13.1 1.9 12.8 17.5 10.2 15.8

13.5 3.5 6.5 2.0 14.7 1.9 13.5 17.9 10.4 16.1

21.3 7.5 6.4 5.6 13.1 2.6 11.9 12.1 8.4 11.1

13.2 2.6 6.4 2.3 13.2 2.3 14.7 21.3 8.4 11.1

13.5 3.5 6.5 2.0 14.7 1.9 13.5 18.4 10.4 15.6

4.3 1.6

23.9 8.0

28.2 7.6

53.3 6.0

61.7 4.8

74.8 21.4

79.3 20.1

82.4 21.9

66.7 21.4

82.0 22.4

82.4 21.9

a b c d

K/FeCuAlOx

K/FeOx

FTS reaction was carried out under the reaction conditions of H2 /CO = 2, P = 1.0 MPa and T = 300 ◦ C at a different superficial gas velocity. –CH– represents the total hydrocarbons selectivity. O/(O + P) represents olefin selectivity (C-mol%) defined as the mole percentage of olefin/(olefin + paraffin) in the range of C2 –C4 hydrocarbons. The yield to olefin is the reacted CO moles to olefins in the range of C2 –C4 hydrocarbons after subtracting the extent of CO conversion to CO2 .

with data reported in the JCPDS files [17]. The peak patterns on FeCuK/Al2 O3 and FeCuK/SiO2 catalysts are much broader than that of K/FeCuAlOx and K/FeOx catalysts due to the high dispersion of iron species with strong interaction with porous supporting materials and the relatively lower concentrations of iron metal. In Fig. 6, the XRD patterns of four different catalysts after FTS reaction reveal the presence of iron carbide species and metallic iron together. Particularly, the broad peak at 2 = 44◦ is a strong indication of the iron carbide formation during FTS reaction. In particular, the peak around 2 = 26◦ is related to the metallic iron species. Catalytic activity is strongly dependent on the extent of iron carbide formation [25] and its intensity is significantly higher on K/FeCuAlOx and K/FeOx catalysts which showed high CO conversion and C2 –C4 olefin selectivity. The TPR experiment with CO (CO-TPR) in order to further investigate the formation of iron carbide species was conducted in the temperature range of 100–1000 ◦ C with a heating rate of 5 ◦ C/min. The CO under CO-TPR condition also reduces the iron oxides to metallic iron and further carburizes the metallic iron to iron carbide species depending on the temperature according to the Boudards reaction (2CO = C + CO2 ). The evolved CO2 is a measure of carbon formation on catalyst surface and is also a measure of the tendency to transform metallic iron to iron carbide species, which is well established as an active species during FTS reaction [26]. During CO-TPR experiments, CO2 evolution was monitored with a fragment of m/z = 44 as a function of temperature; the CO-TPR profiles are shown in Fig. 7. It can be seen that CO2 evolution is found to be broad: three different peaks can be assigned to peaks of I, II and III. The peak of CO2 evolution below 400 ◦ C (peak I) is mainly attributed to the reduction of hematite to magnetite, peak II around 500 ◦ C is assigned to the formation of iron carbide species by carburization or polymeric carbon formation and peak III above 600 ◦ C could be assigned to the carbon deposition on the catalyst surface of iron-based FTS catalyst [27,28]. The lowest temperature of CO2 evolution peak (I) was observed on the coprecipitated K/FeCuAlOx catalysts due to the weak interaction between iron and supporting materials; however, a higher temperature on SiO2 support was observed due to the strong interaction of iron and SiO2 [24]. The different reduction behavior on CO-TPR compared to that of H2 -TPR is mainly due to the different iron concentrations on catalyst surface on each FTS catalyst. Even though FeCuK/Al2 O3 and FeCuK/SiO2 catalysts showed a facile reducibility on H2 -TPR due to the high surface

concentration of iron oxides compared to those of K/FeCuAlOx and K/FeOx catalysts, their low reduction behavior on CO-TPR may be attributed to the strong interaction of iron oxides with supporting materials, especially for SiO2 support by forming inactive iron silicate [24]. The CO2 evolution peak (II) related with the formation of active iron carbide species is correlated with the result of H2 consumption from H2 -TPR and its facile formation is observed on all FTS catalysts. Interestingly, the carbon deposition assigned to peak (III) is much more severe on K/FeOx catalyst even though it showed high FTS activity; this could be attributed to the low surface area on K/FeOx . Since the peak (III) is known to be related with the formation of less active polymeric carbons, the further characterization was carried out to elucidate the kinds of carbons by using TPSR on the reacted catalysts, as expected in the following section. From the results of H2 -TPR and CO-TPR, the facile reducibility of iron oxides to metallic iron and the easy formation of iron carbide species are mainly responsible for the high initial activity of iron-based FTS catalyst, except for K/FeOx catalyst. The weight ratio of K/Fe and Cu/Fe on the fresh and used FTS catalysts is further characterized by using surface-sensitive SEM/EDAX analysis and the summarized values are presented in Table 1. The

Fig. 7. CO-TPR profiles of iron-based FTS catalysts using a mass spectrometer.

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177

Fig. 9. TPSR profiles of iron-based FTS catalysts after FTS reaction on BFBR using a mass spectrometer.

Fig. 8. XPS spectra of Fe 2p and C 1s on iron-based FTS catalysts after FTS reaction. (A) Fe 2p on the used catalysts. (B) C 1s on the used catalysts.

deviation of K/Fe weight ratio on fresh FTS catalysts suggests the strong effects of support contribution to metal dispersion. The high concentration of K on K/FeOx catalyst of K/Fe ratio around 0.138 was observed due to the small surface area of iron ore and vice versa on K/FeCuAlOx catalyst. The similar value is also observed on FeCuK/SiO2 catalyst and low value on K/FeCuAlOx catalyst is attributed to the high concentration of iron. The similar trend of Cu/Fe weight ratio was also observed on fresh FTS catalysts. However, the weight ratio of K/Fe increased on all FTS catalysts after reaction, except for K/FeOx catalyst. The decrease of K/Fe ratio on K/FeOx after reaction suggests that K component was rearranged during FTS reaction with better dispersion compared to that of fresh catalyst. The increase of K/Fe weight ratio on the other catalysts suggests the possible segregation of K component. But, the variation of Cu/Fe weight ratio is not significant on all catalysts after FTS reaction. The redistribution of K component during FTS reaction is also contributed to the different catalytic performance by showing a high catalyst activity on K/FeOx catalyst. In addition, the concentration of K on iron-based catalyst alters the catalytic performance with the increase of olefin selectivity up to 6 wt% K on iron-based catalyst confirmed by our previous investigation [14], however, the higher concentration of K is responsible for showing a low catalytic performance. In order to understand the iron phase transformation that might have occurred during FTS reaction, we found Fe 2p3/2 and Fe 2p1/2 values from XPS analysis are taken for the reacted FTS catalysts and the spectra are shown in Fig. 8(A). The binding energy (BE) of metallic iron has been reported to be around 707 eV for Fe 2p3/2 and at

around 711 eV for the corresponding iron oxides [29]. Especially, the BE of Fe 2p3/2 on K/FeOx catalyst showed somewhat higher value at 709.7 eV and the formation of iron carbides is also not significant as compared to other catalysts. This behavior of K/FeOx catalyst is also supported by the results of XRD, CO-TPR and TPSR analyses which is due to the difficult reduction of bulk iron ore and the small amount of iron carbide formation on the outmost surface of K/FeOx catalyst. Furthermore, the spin orbit splitting values for Fe 2p3/2 and Fe 2p1/2 peaks are known to be around 11.3 for iron metal and around 11.8 for the iron oxides, and these values are not much altered on our reacted iron-based catalysts. The BE shown in Fig. 8(A) is slightly higher than that of metallic iron species but significantly lower than that of iron oxides. Similarly the spin orbit splitting value is somewhat lower than those of reported iron oxides and metallic iron species. These results suggest that the reacted iron-based catalysts, even after passivation in the air environment, contain predominantly metallic iron species, but the surface iron oxides formed during the passivation step or unreduced iron oxides are also contributing to the shift in BE. It is difficult from BE values and spin orbit splitting values alone to say anything about formation of surface carbides on the passivated catalysts, but the lower spin orbit splitting values and BE suggest the presence of other phases of iron species which may be iron carbides. The high intensity of the Fe 2p3/2 peaks on FeCuK/Al2 O3 , K/FeCuAlOx and K/FeOx catalysts suggests the high concentration of iron species on the catalyst surface compared to that of FeCuK/SiO2 catalyst; this difference is related with the results of H2 -TPR and CO-TPR. Further to confirm the correlation between the extents of iron carbide formation and catalytic activity, XPS analyses of carbons species on the used catalysts were carried out and the results are shown in Fig. 8(B). Even though the detailed assignment of carbon species on the used catalyst surface is difficult due to the significant transformation during O2 passivation, the characteristic two C 1s peaks are observed on the used FTS catalysts around 284.4 and 288.5 eV (assigned to peak (I) and peak (II) respectively). The characteristic XPS peaks of C 1s, as reported by Zhang et al. [30], appear at the biding energy of around 284 eV for carbidic carbon and 288 eV for carbonate species. However, on our passivated catalysts after FTS reaction, the peaks could be assigned to the adsorbed carbon species for the peak shown at around 284.4 eV (peak (I)) and deposited carbon species as an active iron carbide at 288.5 eV (peak (II)). The active iron carbides could be easily transformed to carbonates during O2 passivation step and the peak at 288.5 eV could be possibly assigned to active iron carbide species. The intensities of peak (I) at 284.4 eV and that of peak (II) at 288.2 eV

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Fig. 10. TEM images of fresh and used iron-based catalysts of FeCuK/Al2 O3 and FeCuK/SiO2 .

are compared to elucidate the extent of active carbide formation. As shown in Fig. 8(B), the intensity ratio of peak (II)/peak (I) is larger on K/FeCuAlOx and K/FeOx catalysts than on FeCuK/Al2 O3 and FeCuK/SiO2 catalysts. The values are found to be 0.188, 0.402, 0.137 and 0.143 for K/FeOx , K/FeCuAlOx , FeCuK/SiO2 and FeCuK/Al2 O3 catalysts respectively. The higher value of peak (II)/peak (I) on K/FeOx , K/FeCuAlOx catalysts suggests the facile formation of active iron carbide species and resulted in a high catalytic performance. The amount of active iron carbide species on iron-based FTS catalyst plays an important role in catalytic activity and product distribution. Xu and Bartholomew [31] reported that the carbon species revealed during temperature-programmed hydrogenation include (i) adsorbed atomic carbon at the peak temperature around 250–400 ◦ C, (ii) polymeric carbons at 420–460 ◦ C, (iii) bulk iron carbides at 480–650 ◦ C and (iv) graphitic surface carbons above 650 ◦ C. They reported that the reactivity of carbon species decreased with the increase of peak temperature. To elucidate the kinds of carbon species on our iron-based FTS catalysts, we carried out TPSR studies on the reacted catalysts used in FBR, as shown in Fig. 9. The most active species on catalyst surface assigned to peak (I) of adsorbed atomic carbon below the peak temperature of 400 ◦ C is dominant on the FeCuK/Al2 O3 catalyst with the peak temperature below 200 ◦ C and on FeCuK/SiO2 catalysts at around 350 ◦ C. These catalysts also showed large CH4 formation compared to the other catalysts such as K/FeCuAlOx and K/FeOx catalysts. The abundant presence of a second active carbon species assigned to peak (II) on K/FeCuAlOx at around 450 ◦ C and on K/FeOx at around 480 ◦ C is related to the facile formation of higher hydrocarbons. Especially, peak (III) values assigned to the iron carbide species are larger on K/FeOx catalyst. The active carbon species (peak (II) and (III) assigned to polymeric carbon and iron carbides) are dominant on K/FeOx catalyst even though it contains a high quantity of least active graphitic carbons. The high quantity of active carbon species assigned to peaks (II) and

(III) is beneficial to produce the higher hydrocarbons in the range of C2 –C4 products with a high CO conversion. Although the catalytic performance on iron-based FTS catalysts is significantly related with the extent of carbide formation during reaction which is confirmed by CO-TPR, XPS and TPSR experiments, the variation of iron particle size before and after FTS reaction on FeCuK/Al2 O3 and FeCuK/SiO2 is further characterized by TEM analysis and the images are shown in Fig. 10. The carbon formation with the significant aggregation of iron particles is observed on FeCuK/Al2 O3 after reaction, however, the variation of iron particle size on FeCuK/SiO2 catalyst is not significant after FTS reaction. The transformation of metallic iron species to the active iron carbides could be simultaneously occurring with the aggregation of iron particles, and the lower CH4 formation with concomitantly higher C5 + selectivity was observed on the larger iron particles as reported by our previous work [32]. Therefore, the low aggregation of iron particles on FeCuK/SiO2 after FTS reaction is possibly responsible for the inactive carbide formation and resulted in showing a low catalytic performance with the formation of large amount of CH4 product. 4. Discussion FTS reaction for the production of C2 –C4 olefins and clean fuels from syngas was investigated on the four different iron-based FTS catalysts in FBR and BFBR to elucidate the roles of iron carbide formation and to select an appropriate type of reactors and catalysts. The co-production of clean fuels and light olefins could be achieved by FTS reaction with the combination of cracking process of wax component [4]. Although physical properties such as surface area, average pore size and density of FTS catalyst particle are significantly influencing the behavior of fluidization and mass-transfer efficiency, the catalytic performances in FBR and

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179

Fig. 11. The yield to C2 –C4 olefin with respect to the superficial gas velocity in FBR and BFBR.

BFBR are largely correlated with the extent of iron carbide formation and with the acidity of FTS catalyst. The high formation rate of active carbon species assigned to polymeric carbons and iron carbides on K/FeCuAlOx catalysts (conformed by XRD, CO-TPR, XPS and TPSR studies) is responsible for higher catalytic performance with respect to CO conversion and selectivity to higher hydrocarbons. Even though the small amount of iron carbide formation and difficult reducibility was observed on K/FeOx catalyst with a low BET surface area, the active iron carbide sites existing at the outermost catalyst surface are beneficial to suppress hot-spot formation and carbon deposition in inner-pores of K/FeOx catalyst, especially in FBR. The summarized results of C2 –C4 yield with respect to superficial gas velocity values in FBR and BFBR are shown in Fig. 11. The hydrocarbon selectivity in BFBR (Table 4) such as the selectivity and yield to C2 –C4 olefin (value of O/(O + P) defined as the mole ratio of olefin/(olefin + paraffin)) and C5 + on FeCuK/Al2 O3 catalyst shows a lower value than that of other catalysts. As mentioned earlier, the strong acidic sites on that catalyst (Table 2) accelerate the formation of isomer and aromatics in the gasoline range with a maximum chain length up to C12 + at the expense of olefins [19] or enhance the possible hydrogenation activity of C2 –C4 olefins to form paraffinic hydrocarbons, thus reducing the olefin content in C2 –C4 hydrocarbons. The yields to C2 –C4 olefin on FeCuK/SiO2 catalyst are not much altered in FBR and BFBR, due to their intrinsically lower activity with a low formation rate of iron carbide species and a strong inhibition effect of iron oxide reduction to metallic iron species confirmed by CO-TPR. However, the yields to C2 –C4 olefin on K/FeCuAlOx and K/FeOx catalysts in BFBR are found to be much higher than that of FBR, as shown in Fig. 11. These beneficial effects are the main advantages of BFBR for the production of C2 –C4 olefins. The higher formation rate of active iron carbide species with facile reducibility of iron oxides and appropriate acidity plays an important role to obtain high catalytic performance in FBR and BFBR. The hydrocarbon selectivity, especially for C2 –C4 olefins, on K/FeCuAlOx and K/FeOx catalysts can be easily seen from the property of BFBR. For fuel and chemical products by FTS reaction, Davis [33] has reported that the available control points to vary product distribution are the operating conditions such as temperature, pressure, space velocity and an appropriate type of reactor with the help of highly selective catalyst development. It is possible to combine the optimum process parameter for the selective synthesis toward light C2 –C4 olefins or heavy paraffin obtained with different type of reactors. Influence of process parameters on the product distribution with different type of reactors is shown, and in particular, they are recommended operations in BFBR. In our investigation, CH4 selectivity on K/FeCuAlOx and K/FeOx catalysts decreases with

Fig. 12. Particle size distributions of iron-based catalysts before and after FTS reaction: (A) K/FeCuAlOx ; (B) K/FeOx .

the increase of superficial gas velocity, while the O/(O + P) value increases as shown in Table 4. At the bubbling regime operating in the present investigation, the increase of superficial gas velocity causes the increased bed voidage and this is finally reflected in the expansion of the catalyst bed. The reduction of feed gas volume is difficult due to the mixing between feed gas and catalyst particles without a sufficient supply of feed gas. This mixing is closely connected to the residence time of gas in the catalyst bed. A further increase in gas velocity shows a further transition of fluidization regime from bubbling regime to turbulent regime and the residence time is becoming short. From the viewpoint of heat transfer in bubbling regime of general fluidized bed reactor, the heat transfer coefficient has the maximum valve on that regime, and also the mixing effect between feed gas and catalyst reached the maximum valve [34]. Therefore, the optimum gas velocity to obtain the best catalytic performance in BFBR could be altered according to type of catalysts such as density of catalyst particle as well as the high intrinsic activity of iron-based catalyst such as K/FeCuAlOx and K/FeOx catalysts. On those catalysts, the catalytic performance is marginally different according to the superficial gas velocity above Umf . However, further considerations for the high resistance to catalyst attrition should be taken into account to select a proper iron-based catalyst for applying to BFBR. It is already reported that the precipitated iron-based catalyst as K/FeCuAlOx is suitable to low temperature Fischer–Tropsch (LTFT) at around 250–280 ◦ C [35–37] in a FBR, while the fused iron-based catalyst such as K/FeOx is suitable to HTFT at around 330 ◦ C in a BFBR. These two catalysts have a great difference in the catalyst attrition with the operating conditions and duration. The catalyst of K/FeCuAlOx and K/FeOx , which is

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suitable for applying in FBR for LTFT, should be properly designed to prevent a hot-spot formation in catalyst-bed and a coke formation. However, K/FeOx catalyst is only beneficial for obtaining a high C2 –C4 olefin selectivity in BFBR due to a high resistance for catalyst attrition even at a high velocity. Particle size distributions before and after FTS reaction on K/FeCuAlOx and K/FeOx catalysts are shown in Fig. 12. The average particle size of fresh K/FeCuAlOx catalyst was 98.1 ␮m, while that after FTS reaction for 60 h in FTS was decreased to 92.5 ␮m by physical and chemical attrition of FTS catalysts. However, the average particle size of calcined K/FeOx catalyst was 97.2 ␮m and its size is not much reduced after FTS reaction for 60 h and is found to be around 96.4 ␮m. Therefore, the high catalytic performance with respect to CO conversion and product distribution, and the resistance to catalyst attrition observed on K/FeOx catalyst are beneficial for applying to BFBR on iron-based FTS reaction. 5. Conclusions FTS reactions for co-production of C2 –C4 light olefins and clean fuels from syngas were carried out in FBR and BFBR for MTFT reaction at around 300 ◦ C with four different iron-based catalysts. Under similar operating conditions such as temperature and pressure, the C2 –C4 selectivity and olefin selectivity in BFBR are higher than those in FBR on the K/FeCuAlOx and K/FeOx catalysts. Although the textural properties and degree of reduction are not significantly related with FTS activity and selectivity to C2 –C4 olefin, K/FeCuAlOx and K/FeOx catalysts reveal the higher CO conversion due to the facile formation of iron carbide species on the outer surface of catalyst, especially on K/FeOx . The density of active acid sites is lower on K/FeCuAlOx and K/FeOx catalysts, hence C2 –C4 olefin selectivity increases. However, the high concentration of acidic sites on the supported catalysts (FeCuK/Al2 O3 and FeCuK/SiO2 ) is not beneficial for obtaining high olefin selectivity due to the high formation rate of paraffinic hydrocarbons in both FBR and BFBR. These results are verified by XRD, NH3 -TPD, CO-TPR, TPSR and XPS analyses on the calcined and reacted FTS catalysts to confirm the extent of iron carbide formation and the contribution of acidic sites. The supported iron-based catalysts such as FeCuK/Al2 O3 and FeCuK/SiO2 are not appropriate for MTFT due to the high CH4 and CO2 selectivity induced from the inhibition of iron carbide formation that originates from strong interactions between iron and supporting materials. In the view of C2 –C4 olefin selectivity and CO conversion, K/FeCuAlOx and K/FeOx catalysts can be candidates for applying in BFBR for MTFT reaction. However, K/FeOx catalyst that is simply impregnated with potassium promoter shows a high resistance for catalyst attrition with high catalytic performance and this catalyst can be one of the best catalysts for applying MTFT reaction using BFBR.

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Acknowledgement The authors would like to acknowledge that this work was supported by the Technology Innovation Program (Project No. 10028400) funded by the Ministry of Knowledge Economy (MKE, Korea).

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