1588
Ind. E n g . Chem. R e s . 1990, 29, 1588-1599
Carley, A. F.; Rasaias, S.; Roberts, M. W.; Wang, Tang-Han. Chemisorption of Nitric Oxide by Nickel. Surf. Sci. 1979, 84, L227L230. Egashira, Y.; Komiyama, H. Spontaneous Nitridation of Ultrafine Particles of Ni during NO/CO Reaction. Chem. Lett. 1987, 2413-2416. Holloway, P. H.; Hudson, J. B. Kinetics of the Reaction of Oxygen with Clean Nickel Single Crystal Surface, 1 Ni(100) Surface, 2 Ni(ll1) Surface. Surf. Sci. 1974, 43, 123-149. Kimoto, K.; Kamiya, Y.; Nonoyama, M.; Uyeda, R. An Electron Microscope Study on Fine Metal Particles Prepared by Evaporation in Argon Gas a t Low Pressure. Jpn. J . A p p l . Phys. 1963, 2, 702-713. Labohm, F.; Gijzeman, 0. L. J.; Geus, J. W. The Interaction of Oxygen with Ni(ll1) and the Reduction of the Surface Oxide by Carbon Monoxide and by Hydrogen. Surf. Sci. 1983, 135, 409-427. Morrow, B. A.; Moran, L. G. The Adsorption of NO and NO, on
Silica-Supported Nickel. J. Catal. 1980, 62, 294-303. Morrow, B. A.; Sont, W. N.; Onge, A. St. The Reaction between NO and CO on silica-Supported Nickel. J. Catal. 1980,62, 304-315. Price, G. L.; Baker, E. G. The Chemisorption of Nitric Oxide on (100) Nickel Studied by LEED, AES, UPS and Thermal Desorption. Surf. Sci. 1980, 91, 571-580. Price, G. L.; Sexton, B. A.; Baker, B. G. The Chemisorption of Nitric Oxide on (110) Nickel Studied by LEED, AES, and Thermal Desorption. Surf. Sci. 1976, 60, 506-526. Shelef, M.; Kummer, J. T. The Behavior of Nitric Oxide in Heterogeneous Catalytic Reactions. Chem. Eng. Prog., Symp. Ser. 1971, 67 (115), 74-92. Shelef, M.; Otto, K.; Gandhi, H. The Oxidation of CO and COz and by NO Supported Chromium Oxide and Other Metal Oxide Catalysts. J . Catal. 1968, 12, 361-375.
Receioed for review October 30, 1989 Accepted March 22, 1990
Binder/Support Effects on the Activity and Selectivity of Iron Catalysts in the Fischer-Tropsch Synthesis Dragomir B. Bukur,*pt Xiaosu Lang,+Doble Mukesh,? William H. Zimmerman,? Michael P. Rosynek,' and Chiuping LiI Kinetics, Catalysis and Reaction Engineering Laboratory, Department of Chemical Engineering and Department of Chemistry, Texas A&M University, College Station, Texas 77843
The influence of silica and alumina binders (supports) on the activity and product selectivity of precipitated iron catalysts for Fischer-Tropsch synthesis (FTS) was studied in a fixed bed reactor a t 1.5-3.0 MPa and 220-250 " C , using synthesis gas with H 2 / C 0 = 1 M feed ratio. Both the FTS and the water gas shift activity decreased with increasing support content. The catalyst deactivation rate decreased with catalysts containing 24 and 100 g of Si02 per 100 g of Fe. Secondary reactions (olefin hydrogenation and isomerization) increased with increasing silica content. Hydrocarbon selectivities improved (less gaseous hydrocarbons) with the addition of support, except for the catalyst containing 24Si02/ 100Fe. The results were interpreted in terms of interactions between potassium and/or iron with supports.
Introduction The objective of adding a catalyst support is to provide a large surface area for the formation and stabilization of small metal crystallites in the catalyst. The support may also have significant effects on the catalyst activity and selectivity due to strong metal-support interactions (SMSIs) (e.g., Vannice and Garten, 1980; Bartholomew et al., 1980; Reuel and Bartholomew, 1984). Evidence for SMSIs has been found in studies of Fischer-Tropsch synthesis (FTS) over supported nickel, ruthenium, and cobalt catalysts (e.g., Vannice and Garten, 1979, 1980; Goodwin et al., 1984; Reuel and Bartholomew, 1984). In all of these studies, it was found that the catalyst activity and selectivity were influenced by one or more of the following variables: nature of support, metal dispersion, metal loading, and/or preparation method. Several studies have been made with supported iron catalysts (e.g., Anderson, 1956,1984;Guczi, 1981; Dry, 1981; Jung et d., 1982; Egiebor and Cooper, 1985; Berry et al., 1986), but only a few of them dealt with the effect of support type and content on catalyst activity and detailed product distribution. In general, iron catalysts with high support-tometal ratios have been ineffective FTS catalysts. The
* Author to whom correspondence should be addressed. 'Department of Chemical Engineering. Department of Chemistry. 0888-5885/ 9012629- 1588$02.501 0
major findings from studies of direct relevance to the work presented in this paper are summarized below. Dry (1981) reported some results, from comprehensive studies conducted at SASOL (South Africa), on the effect of supports on the catalytic performance of a series of precipitated iron catalysts (Fe/Cu/K20). Unfortunately, only relative concentrations for some components in the catalysts employed were reported, and the catalyst activity and wax selectivity were also reported as relative quantities. Silica was found to be the best support in terms of both activity and selectivity. Catalysts containing a second support material (CaO, Cr203,A120,, V205,Tho2, MgO, or Ti02)in addition to Si02also showed inferior performance relative to a Fe/Cu/K20/Si02 catalyst. Dry emphasizes that a careful balance between promotional (potassium, in the form of K20) and stabilization (support) additives must be achieved to maximize the desired activity and hydrocarbon product selectivities. Potassium is an essential promoter in iron catalysts for the F'TS, since it enhances the formation of both longer chain and olefinic hydrocarbons (Anderson, 1956; Dry, 1981). When acidic oxides are used as supports, they may react with basic alkali and thereby reduce the promotional effect of potassium. Also, the use of high surface area supports tends to reduce direct contact between iron and potassium since the metal covers only a small fraction of the support surface (McVicker and Vannice, 1980). This would also render the potassium promotion less effective. 0 1990 American Chemical Society
Ind. Eng. Chem. Res., Vol. 29, No. 8, 1990 1589 Recently Egiebor and Cooper (1985) conducted a study on the influence of silica support on selectivity to l-olefins, internal olefins, and branched hydrocarbons in the gasoline range (C5411hydrocarbons)of a precipitated iron catalyst. They found that the l-olefin and branched hydrocarbon selectivities decreased whereas the internal olefin selectivity increased with an increase in the silica content of the catalyst. On the other hand, the aromatic and alcohol contents of the products remained relatively constant for all three catalysts tested (100Fe/4.2Cu/6.7K/xSi02 in parts per weight, where x = 21, 50, or 73). The present study was undertaken to determine effects of silica and alumina as binders (supports) and of process conditions on the activity and product selectivity of precipitated iron catalysts. An unsupported catalyst (100Fe/5Cu/4.2K, in parts per weight) containing same promoter concentrations as the supported catalysts was used as the reference (base line) catalyst to determine the undisguised effect of binders (Si02or Al2O3) and promoters (CuO and KzO). Three silica (8,24, and 100 g of Si02per 100 g of Fe) and two alumina (8 and 24 g of A1203per 100 g of Fe) containing catalysts were synthesized and tested in a fmed bed reactor at reaction conditions similar to those employed in industrial reactors for FTS. The composition of one of the synthesized catalysts was chosen to match that of the Ruhrchemie catalyst (100Fe/5Cu/4.2K/ 24sio2),which was initially employed in commercial fixed bed reactors (ARGE reactors) at SASOL. Particular attention was given to the effects of supports, reaction temperature, pressure, and gas space velocity on FTS and water gas shift (WGS) activity, catalyst deactivation, hydrocarbon product distribution, and olefin and oxygenates selectivities. The results of the catalyst characterization by surface area and pore size distribution measurements, isothermal reduction, temperature-programmed desorption, and X-ray photoelectron spectroscopy are also presented and are used to explain some of the results from catalytic tests.
Experimental Section Reactor System and Operating Procedure. Experiments were conducted in a l-cm (inside diameter) stainless steel vertical fixed bed reactor with downward flow of the synthesis gas. The experimental apparatus and the product collection and analysis systems have been described in detail elsewhere (Zimmerman et al., 1989; Bukur et al., 1989). Liquid products collected in a high-pressure trap (wax) and a low-pressure trap (aqueous- and organic-phase products) were analyzed by gas chromatography (two Perkin-Elmer GCs and a Varian 3400 GC). The flow rate of the exit gas (noncondensible products and unconverted reactants) was measured periodically with a soap film flowmeter. Following any change in process conditions, at least 14-16 h was allowed to elapse before a mass balance was made. After the conclusion of this unsteady-state period, liquid samples were collected in steady-state traps over a 6-8-h period, and several tail gas samples were collected and analyzed in a Carle AGC 400 gas chromatography. On the basis of measured gas flow rates and the analysis of products collected, the total and individual atomic closures of carbon, hydrogen, and oxygen were calculated in order to determine the accuracy of the results. Atomic closures of 100% f 3% were normally obtained. Reactor Start-up and Process Conditions. Following reduction, the flow was switched to helium, and the bed was cooled to 190 "C. The system was then pressurized to 1.5 m a , the helium flow was stopped, and the synthesis gas (H2:C0 = 1:l) was introduced at a gas space velocity
(SV) of 2 nL/ (g of cata1yst.h). (Gas space velocity is defined as normal liters per gram of catalyst per hour, and 2 nL/ (g of catalyst-h) is equivalent to a gas hourly space velocity of 250 h-l, based on the bed (catalyst + inert) volume.) The bed temperature was gradually increased to 235 "C over a period of 6 h. The first mass balance was conducted after an additional 40 h on stream. Catalytic tests typically lasted about 170 h, during which time six mass balances were made. One of these balances was a repetition of the base set of conditions, viz., 1.5 MPa, 2 nL/(g of catalyst-h), and 235 "C, whereas others were conducted at different process conditions (see Tables I11 and IV for a list of conditions, times on stream, conversions, and selectivities obtained at different process conditions). The effect of reaction temperature (220,235, and 250 "C) was studied at SV = 2 nL/(g of catalyst-h), and the effect of gas flow rate (SV = 2 and 4 nL/(g of cata1yst.h) was studied at 235 "C. One balance was carried out at higher pressure and space velocity (235 "C, 4 nL/(g of catalystnh), 3 MPa), by keeping the ratio of pressure and gas space velocity constant (i.e., constant contact time in the reactor). In the absence of catalyst deactivation, syngas conversion values obtained in balances with P/SV = constant should be nearly the same, since the space time yield in FTS is proportional to pressure (e.g., Anderson, 1956; Dry, 1981; Bukur and Brown, 1987). Catalyst Preparation/Reduction Procedure. Catalyst preparation involved three steps: preparation of the iron-copper precursor, incorporation of binder/support (silica or alumina), and finally potassium impregnation. The constant-pH precipitation technique used to prepare the Fe/Cu catalyst precursor has been described in detail previously (Bukur et al., 1989). In brief, the catalyst precursor was continuously precipitated at 82 "C from a flowing aqueous solution containing iron and copper nitrates at the desired Fe/Cu ratio, using aqueous ammonia. The precipitate was then thoroughly washed with distilled water by vacuum filtration. Impregnation with SiOz binder/support was accomplished by addition of an appropriate amount of dilute (26 wt %) K2Si03solution to undried, reslurried Fe/Cu coprecipitate, followed by adjustment of the pH to 1 6 to ensure complete deposition of the silicate. Incorporation of A1203binder/support was performed by adding the required amount of dried Fe/Cu precipitate to a solution of aluminum sec-butoxide in 2butanol and decomposing the impregnated aluminum sec-butoxide to form A1203 by exposure to air saturated with water vapor. After a vacuum drying step, the potassium promoter was added as aqueous KHC03 solution via an incipient wetness pore-filling technique. The final step was to dry the catalyst at 120 "C for 16 h in a vacuum oven. The dried catalyst was calcined in air at 300 "C for 5 h and then crushed and sieved to a diameter between 0.25 and 0.55 mm (30/60 mesh). Typically, 3.5 g of the catalyst (3 cm3) was diluted 1:8 by volume with glass beads of the same size range and was charged into the reactor. All catalysts were reduced with carbon monoxide at 280 "C and atmospheric pressure for 12 h. Six catalysts were employed in the present study: an unsupported iron/copper/potassium catalyst (100Fe/ 5Cu/4.2K in parts per weight), three silica-supported catalysts (1O0Fe/5Cu/4.2K/xSiO2, where x = 8, 24, and loo), and two alumina-supported catalysts (100Fe/5Cu/ 4.2K/xA1203,where x = 8 and 24). The impregnation technique used to prepare the alumina-containingcatalysts could attain a maximum concentration of only about 30A1203/100Fe, so a catalyst with high alumina concen-
1590 Ind. Eng. Chem. Res., Vol. 29, No. 8, 1990
tration ( 100A120,/ 100Fe) was not available for testing. Catalyst Characterization. The metal concentrations in the catalyst were determined by atomic absorption spectroscopy using a Varian Spectra AA 30 absorption spectrophotometer. The following techniques were used to characterize selected catalysts after calcination and CO reduction: (i) surface area measurements utilizing the BET method on a Quantachrome-Quantasorb apparatus; (ii) pore size distribution by mercury porosimetry on a Micromeritic Pore Sizer 9310; (iii) temperature-programmed/isothermal reduction; (iv) X-ray photoelectron spectroscopy on a Hewlett-Packard (Model 5950A) ESCA spectrometer. A brief description of the last two techniques, including sample-handling methods, is given below. The bulk reduction behavior of each catalyst was studied by using a temperature-programmed/isothermalreduction technique. A 5% CO/He reductant stream, maintained at a flow rate of 12 cm3/min by a mass flow controller, was directed through the reference compartment of a thermal conductivity (TC) cell, then through a 10-15-mg catalyst sample contained in a U-shaped quartz reactor, and finally through the other compartment of the TC cell. Consumption of CO reductant during temperature ramping or isothermal treatment of the sample caused a change in thermal conductivity of the effluent gas stream, resulting in a continuous graphical display of reduction rate as a function of time on an X-Y plotter. An in-line trap, located upstream of the detector and maintained at -196 "C, ensured continuous removal of the C02byproduct formed during iron/copper reduction by CO. While under CO/He flow, the sample temperature was ramped in each experiment at 20 "C/min to 300 "C and then maintained at the latter temperature for 4-12 h, while continuously recording the rate of CO consumption. Detector response was calibrated by measuring the peak area corresponding to total CO consumption during reduction of a dry CuO sample of known weight. X-ray photoelectron spectroscopy (XPS) was used to study surface compositions and oxidation states of the iron catalysts following reduction treatment in CO. Samples were prepared in the form of thin pressed wafers and were retained on a gold-plated copper sample holder by a gold-plated aluminum spring clip. After external reduction treatment in CO at 300 "C, samples were evacuated and then transferred in sealed reactors, without exposure to the atmosphere, to the antechamber of a glovebox, into which the XPS sample probe penetrated through an airtight seal. The antechamber was back-filled with dry nitrogen before transferring the sealed reactor into the main glovebox chamber for sample mounting on the instrument probe. Typical pressure in the XPS analysis chamber during measurement was Torr. The monochromatic X-ray beam power was maintained at 600 W, and an electron flood gun was used to eliminate sample charging effects. The C 1s peak at 284.6 eV, due to adventitious surface carbon, and the 0 1s peak at 531.6 eV, due to oxygen in the Si02or A1203 support, were used as internal references. The Au 4f7,2peak at 83.8 eV from the gold-plated sample holder was used as an external reference.
Results and Discussion Catalyst Characterization. Each catalyst preparation was analyzed for actual iron, copper, potassium, and silica or alumina contents by atomic absorption spectroscopy. Table I summarizes the intended nominal and actual analyzed compositions of each catalyst used in this study. For the purpose of simplicity, each catalyst will be referred to by its nominal composition. The total surface areas,
Table I. Compositions and Properties of Synthesized Iron Catalysts surface composition, parts by wt area: metal nominal analvzed m2/e exDosureO 100Fe/5Cu/4.2K 100Fe/5.1Cu/4.1K 30 0.015 100Fe/5Cu/4.2K/ 100Fe/5.1Cu/4.OK/ 95 0.022 SO2 7.8Si02 100Fe/5Cu/4.2K/ 100Fe/5.4Cu/4.6K/ 150 0.021 24sio2 ' 28@ ' 100Fe/5Cu/4.2K/ 100Fe/5.3Cu/4.1K/ 250 100Si0, 96.5Si0, 100Fe/5&/4.2K/ 100Fe/5.0&/3.9K/ 85 0.040 8A1;03 7.7A1203 100Fe/5Cu/4.2K/ 100Fe/5.1Cu/4.OK/ 60 0.051 25A12O3 20.5A120, I -
'
'
Following calcination in air a t 300 OC for 5 h and treatment in flowing CO for 4 h at 300 "C.
determined by application of the BET method to N2 adsorption isotherms a t -196 "C, were obtained for each catalyst, following reduction treatment in CO a t 300 "C. These data, also presented in Table I, demonstrate that increasing the content of Si02support results in a substantial increase in catalyst surface area. The surface area of the unsupported 100Fe/5Cu catalyst precursor, prior to calcination and reduction, is about 300-350 m2/g. In the absence of an added binder/support, the surface area of the precursor decreases to 30 m2/g after calcination and reduction, due primarily to removal of water during dehydration that causes collapse of the highly porous, high surface area Fe00H/Fez03structure. Introduction of Si02 into the pores of the precipitate by impregnation with K.$i03 provides a rigid matrix that helps to prevent complete collapse of the original pore structure during subsequent thermal treatments, thus maintaining most or all of the original surface area of the precipitate, depending on the amount of Si02added. Thus, the high surface areas of the Si02-supported catalysts are primarily due to partial or complete maintenance of the original high surface area of the Fe/Cu precipitate and are not an inherent property of the SiOz itself. In addition to surface area measurements, mercury porosimetry was employed to obtain the pore size distributions of calcined and reduced catalysts. Following calcination and subsequent reduction in CO for 4 h a t 300 "C, the distribution of the pore volume for each of the three silica-supported catalysts differed, depending upon the Si02content. The 8SiOzmaterial exhibited a well-defined maximum at a pore diameter of 150-200 A, while the 24sio2 and 100Si02catalysts displayed bimodal pore size distributions. In addition to a maximum at 100 A, each of the latter two materials had a substantial fraction of its total pore volume accounted for by pores 160 8, in diameter (the lower limit of detection of the porosimeter employed for these measurements). The 8A1203catalyst displayed a broad size maximum centered a t -125 A, while the catalyst showed a maximum a t 8Alz03 (8SiOz) > 24AlZo3> 24sio2 > 100Si02. Possible explanations for the decrease in catalyst activity with the increasing support content are (1)a lower degree of reduction and (2) a reduction in the effective potassium content of the catalyst. Catalysts containing 24 or more parts support per 100 parts iron had a lower degree of reduction than the unsupported catalyst and the two catalysts containing small amounts of binder (8Si02or 8Al2O3) as shown in Table 11. According to recent studies with a variety of iron FT catalysts (Berry et al., 1986; Dictor and Bell, 1986; Snel, 1989), the activity increases with the degree of reduction. The observed decrease in activity with increasing support content in our study is in qualitative agreement with these studies. Potassium promotion, up to a certain amount, is known to increase the FTS activity (e.g., Anderson et al., 1952; Dry, 1981; Bukur et al., 1990). The FTS activity of the unsupported iron catalyst (100Fe/5Cu/4.2K) employed in the present study is markedly higher than that obtained in our previous studies (Bukur et al., 1987, 1990) with catalysts containing smaller amounts of potassium (i.e., 100Fe/3Cu/xK, x = 0.2 or 0.5). Therefore, the observed decrease in FTS activity with increasing amount of Si02 might be attributed in part to the reduction in the effective potassium content of the catalyst. The latter is due to the reaction between potassium and silica (Dry and Oosthuizen, 1968) or to decreased Fe/K contact on high surface area SiOz-supportedcatalysts (McVicker and Vannice, 1980). In contrast to our results, Egiebor and Cooper (1985) found that the reactant conversions changed only slightly as the support concentration increased from 21 to 73 parts SiOz per 100 parts Fe. The H2 CO conversions in tests conducted at 0.7 MPa, 300 "C, 240 h-l, and Hz/CO = 1 were between 51.4% and 54.5%. These conversions are similar to that obtained in our study with the 24sioz/ lOOFe catalyst at 235 "C, indicating that our catalyst was more active. The differences in trends (activity versus support concentration) in Egiebor and Cooper's study and in the present one may be due to the differences in methods of catalyst preparation, activation procedures, and process conditions employed in the two studies. The base-line conditions in our study were repeated between 143 and 169 h on stream, and the activity of all catalysts had declined (Table 111). The unsupported catalyst, the 8SiO2/lOOFe catalyst, and the two aluminacontaining catalysts had lost between 56% and 65% of their activity at 40-48 h on stream, whereas the loss in activity, over the same period of time, was much smaller (- 12-17 % ) with catalysts containing 24 and 100 parts SiOz per 100 parts Fe. The 24SiOZ/100Fe catalyst exhibited excellent stability during tests in fixed bed and slurry bed reactors at 250 "C, 1.5 MPa, 2 nL/(g of catalystah), and H2/C0 = 0.67 over a period of 480 h, as reported elsewhere (Bukur et al., 1988). The loss in activity may be attributed in part to the high potassium content of these catalysts, since potassium fosters carbon depos-
+
40
lo
1-1 c
0 ' 210
220
230
240
250
260
Temperature (OC)
Figure 3. Effect of catalyst support and reaction temperature on water gas shift activity.
ition, which, in turn, blocks the active sites on the surface (Anderson et al., 1952; Dry, 1981; Arakawa and Bell, 1983; Bukur et al., 1990). The increase in stability with addition of silica support may be the result of stabilization of iron crystallites during synthesis and/or interaction between the potassium and silica. The latter decreases the surface basicity of the catalyst and thus reduces the rate of carbon deposition. Catalyst deactivation did not have a marked effect on product selectivities (hydrocarbon product distribution, olefin selectivities, CO conversion to COz and oxygenates) as shown in Table I11 by comparing the results at different TOS (40-48 h versus 143-169 h). Therefore, it may be assumed that product selectivities obtained during other mass balances, conducted between 50 and 140 h on stream, were not affected by partial catalyst deactivation. Water Gas Shift Activity. The effects of reaction temperature and support type and content on the water gas shift (WGS) activity are shown in Figure 3, where the CO conversion to COz (COzselectivity) is used as a measure of the WGS activity. The WGS activity increased with temperature for all catalysts and decreased with increasing silica content. The unsupported catalyst, the two alumina-supportedcatalysts, and the 8SiO2/l00Fecatalyst had high WGS activity, even at 220 "C. Since potassium is known to promote the WGS activity of iron FischerTropsch catalysts (Anderson et al., 1952; Arakawa and Bell, 1983; Bukur et al., 1990), high WGS activity of these catalysts may be attributed to their high potassium content and lack of interaction between potassium and support. On the other hand, the observed decrease in the WGS activity with increasing amount of silica is an indication of partial neutralization of alkali promoter by acidic sites on silica or decreased Fe/K contact due to the high catalyst surface area. The observed differences in WGS activities may be disguised by differences in CO and/or H2 CO conversions (see Tables I11 and IV). However, from our data at constant temperature (235 "C), it appears that the effect of conversion on COz selectivity is not so marked (Table 111), and comparisons of C 0 2 selectivities of different catalysts at constant CO conversions reveal that the trends depicted in Figure 3 reflect the intrinsic differences among the catalysts that are not disguised by differences in conversions. For example, the COz selectivity of the 24sio2 catalyst at 40 h on stream is 39%, whereas COP selectivities of the unsupported catalyst and the 8Si02and 8A1203catalysts at 144 h are all higher (42.3-44.1 % ). Note that in all four experiments the values of CO and H2 + CO conversions were nearly the same. Also, the C02selectivity of the 100SiOzcatalyst at 250 "C is much smaller than that of the 24Si02 catalyst at 220 "C, 24.8% versus 32.3%,
+
1594 Ind. Eng. Chem. Res., Vol. 29, No. 8, 1990 Table IV. Effect of SuDDort and Process Conditions on Catalyst Activity and Selectivity catalystb unsupported 8Si02 24sio2 100SiOz time on stream, h 95 temp, "C 250 pressure, MPa 1.5 S V , nL/(g cat..h) 4.0 CO conversion, % 78.8 Hz + CO conversion, 70 , 65.4 2.30 exit molar H2/C0 ratio extent of WGS (pcO2/pco2+ p ~ ~ 0.78 o ) hydrocarbon selectivities, wt 7' 5.6 CH, 21.6 c2-c4 c5-c:: 23.7 ClZC 49.1 CO conversion to products, % hydrocarbons 48.2 oxygenates 2.4 45.3 CO2 4.1 unaccounted olefin selectivity, wt 70 78.9 cs-c,, 80.2
8Alz03
24A1203
120 220 1.5 2.0 37.3 33.0
169 235 3.0 4.0 55.6 54.3
96 250 1.5 4.0 80.5 65.1 1.17 1.22 2.77 0.70 0.68 0.83
220 1.5 2.0 73.6 62.1 2.01 0.79
145 235 3.0 4.0 46.4 42.7 1.25 0.73
96 250 1.5 2.0 78.2 67.6 1.97 0.79
120 220 1.5 2.0 29.7 31.1 0.96 0.47
96 250 1.5 2.0 30.8 35.2 0.87 0.37
119 220 1.5 2.0 11.3 15.8 0.90 0.12
169 235 3.0 4.0 22.4 29.4 0.82 0.15
96 250 1.5 4.0 69.8 57.8 1.79 0.85
119 220 1.5 2.0 36.9 31.7 1.17 0.71
168 235 3.0 4.0 42.2 40.0 1.08 0.64
96 250 1.5 2.0 88.6 72.8 3.75 0.85
120 220 1.5 2.0 24.7 22.1 1.07 0.71
168 235 3.0 4.0 33.0 30.5 1.10 0.67
4.6 19.0 25.7 50.7
4.6 23.1 19.5 52.8
3.6 15.4 14.7 66.3
2.5 14.4 14.3 68.8
2.9 17.7 16.8 62.6
7.3 27.0 21.5 44.2
4.9 23.2 20.7 51.2
5.7 18.0 25.3 51.1
4.6 15.8 18.3 61.3
4.3 14.8 14.8 66.1
4.6 18.1 20.5 56.8
3.4 14.6 16.5 65.5
4.4 20.6 21.3 53.7
4.7 20.2 24.3 50.8
4.0 19.5 21.8 54.7
4.8 21.1 18.5 55.6
52.7 3.0 40.9 3.4
53.3 1.9 40.3 4.5
52.2 2.0 43.9 1.9
54.4 2.1 42.4 1.1
58.3 2.2 42.2 -2.7
51.8 1.8 43.1 3.3
68.9 3.6 32.3 -4.8
65.5 3.0 24.8 6.7
80.3 4.5 11.0 4.2
77.9 4.9 10.9 6.3
48.8 2.3 44.6 4.3
53.4 2.4 39.0 5.2
54.4 2.5 37.9 5.2
49.2 1.7 46.5 2.6
52.8 2.0 44.5 0.66
54.6 2.1 37.6 5.7
121
74.6 74.3 79.7 77.7 75.4 72.7 79.1 76.1 77.8 76.7 80.5 76.4 69.0 80.7 75.8 75.2 76.2 76.1 78.6 76.9 75.9 75.1 73.0 76.6 76.0 73.5 79.7 16.7 75.2 80.1 75.8 74.7
OSpace velocity based on unreduced catalyst. bActive metal and promoter amounts for all catalysts: 100Fe/5Cu/4.2K (in parts by weight).
whereas the conversions are about the same. It should be noted that the 100Si02catalyst was tested at much higher temperature (which would tend to increase the WGS activity) than that used for the 24Si02 catalyst, yet it has significantly lower WGS activity. Egiebor and Cooper (1985) did not discuss the effect of silica content on the WGS activity; however, they reported values of H, and CO conversions. From these data, the H2/C0 usage ratios can be calculated, and for the three catalysts (21-73 parts Si02per 100 parts Fe) employed in their study the values of the usage ratio, U , were between 1.5 and 1.7. The value of the usage ratio is indicative of the extent of the WGS reaction and lies in the range (y/4) < U < (1 + y/2) (Bukur and Brown, 1987). y is the average hydrogen to carbon atomic ratio in hydrocarbon products, and its value is typically between 2.2 and 2.4. The lower limiting value of the usage ratio corresponds to high WGS activity (all water produced by FTS reaction is converted to COz by the WGS reaction), whereas the upper limit corresponds to the case where there is no WGS reaction a t all (no CO, production). The usage ratios obtained in Egiebor and Cooper's study were rather high, implying that the WGS activity of their catalysts was low. H,/CO usage ratios in our study were much lower; e.g., at 250 "C, the usage ratios were between 0.65 and 0.73 for the unsupported catalysts and catalysts containing 8-24 parts support (Si02or A1203),whereas the 100Si02/100Fe catalyst had a usage ratio of 1.3. The latter value is still significantly lower than those obtained in Egiebor and Cooper's study, indicating that the interaction between potassium and silica was much weaker with catalysts employed in our study than in theirs. Carbon Number Product Distribution. The products of FTS follow the Anderson-Schulz-Flory (ASF) distribution, which is characterized either by a single value or by two values of the chain growth probability factor (Anderson, 1956; Huff and Satterfield, 1984). We found that whenever all products collected were analyzed, including those in the high-pressure trap, two chain growth probabilities were needed to characterize the product distribution. An example of the ASF distribution is shown in Figure 4, for the data obtained in tests with two silica-containing catalysts (8Si02/lOOFe and 24Si02/100Fe)9
1
T-236OC. SV-2 nL/g-cat.h. P-1.6 MPa. H2XO-l
100 Fe/6 Cu/42 K/24 Si02
c
.-c
0.1
0
t
i: 0
0.01
"."".
J
100 F d 6 Cd4.2 K/B SO2
.-c0
0.1
E
-
U
0
I
001
0
4
E
12
16
20
24
28
32
Carbon Number
Figure 4. Anderson-Schdz-Flory distribution for selected catalysts.
The experimental data are well represented by the three-parameter model of Huff and Satterfield (1984). (cy1 and a2are chain growth probabilities associated with type 1 and type 2 sites, respectively, and p is the fraction of type 1 sites on the catalyst.) Similar results were obtained with the other catalysts employed in the present study. Hydrocarbon Product Distribution. Figure 5 illustrates the effects of support type and content and reaction temperature on hydrocarbon product distribution. For a given catalyst, the reaction temperature did not have a marked effect on hydrocarbon selectivity as shown in Figure 5 and Tables I11 and IV. In most cases, the product distribution shifted toward lower molecular weight hydrocarbons (more methane and Cz-C4 hydrocarbons and less Clz+products) with increase in temperature. The
Ind. Eng. Chem. Res., Vol. 29, No. 8, 1990 1595
T("C):
220 235 250
220 235 250
220 235 250
220 235 250
8SiO2
24Si02
100Si02
Catalyst: Unsupported
a
lOOFe/5Cu/4 2K
0
100Fs15Cu14 2KIlOOS102
L
100Fs15Cu/4 ZKl8A1203 100Fa15Cu/4 2K124A1203
2
3
4
5
6
T=235OC, SV=2 nL/g-cat h P = l 5 MPa. H2/CO=1
7
8
9 1 0 1 1 1 2 1 3 1 4 1 5
Carbon Number
Figure 6. Effect of support type and content on olefin selectivity.
1
T("C):
220 235 250
220 235 250
Catalyst:
Unsupported
8AIzO3
220 235 250 24A1203
P = 1.5 MPa, SV = 2 or 4 nL/g-cat.h, H2/CO = 1
Figure 5. Effect of support type and content and reaction temperature on hydrocarbon selectivity: (a) silica-supported catalysts and (b) alumina-supported catalysts.
amount of liquid and solid (wax) products produced was high with all catalysts tested. This may be attributed to their high potassium content, since the latter is known to favor production of high molecular weight products (Anderson et al., 1952; Dry, 1981; Arakawa and Bell, 1983; Bukur et al., 1990). All supported catalysts, except the 24Si02/lOOFe catalyst, produced less methane and gaseous (C2-C4) hydrocarbons than the unsupported catalyst at all three reaction temperatures. This trend is unexpected, since the promotional effect of potassium is expected to decrease with the addition of support. The expected trend was observed in a test with the 24Si02/100Fe catalyst, where the selectivity toward lower molecular weight products was greater than that obtained with the unsupported catalyst. With the 1O0SiO2/lOOFe catalyst, one would expect to see even more methane and C2-C4 hydrocarbons; however, this catalyst produced somewhat less C1-C4 hydrocarbons than the unsupported one under all conditions tested. This suggests a possibility of some Fe-Si02 (metal/support) interactions, the result of which is an enhancement in selectivity of long-chain hydrocarbons. This enhancement appears to be large enough to compensate for reduction in surface basicity of the catalyst. The workers at SASOL showed that a higher surface basicity correlates with a lower methane selectivity in the Fischer-Tropsch synthesis (Dry and Oosthuizen, 1968; Dry, 1981). They found that the addition of silica depresses surface basicity and, as a result, methane selectivity increased in tests with fused iron catalysts. However, Dry (1981), in discussing the influence of different supports on FT synthesis of precipitated iron catalysts, reported that the Si02-supported catalyst had the highest wax selectivity (i.e., more higher molecular weight products) amongst the various catalysts tested, including that with no support. The unsupported and the silica-supported catalyst had different amounts of potassium (0.67- and 10K20, re-
spectively), so in their case, it is not clear whether the increased production of high molecular weight products was due to iron-silica interactions or to differences in the amount of potassium. Our study indicates that the addition of silica at constant potassium loading relative to iron does not necessarily result in increased methane selectivity. Additional studies are required to explain these observations. Changes in gas space velocity (2 versus 4 nL/(g of catalyst-h) at 235 "C and 1.5 MPa) or reaction pressure (1.5 versus 3.0 MPa at 235 "C and 4 nL/(g of catalystoh)) had only a minor effect on hydrocarbon product distributions in tests with all catalysts as shown in Tables I11 and IV. The presence of the silica support also influenced the selectivity of branched hydrocarbon products in the gasoline fraction (C5-Cll hydrocarbons). These products were calculated as the remainder after linear paraffins, olefins, and alcohols have been identified. The branched hydrocarbon selectivity decreased with the increase in the silica content of the catalyst, whereas such an effect was not observed with the two alumina supported catalysts. Our observations are in agreement with the results of Egiebor and Cooper (1985), who reported a linear decrease in branched hydrocarbon selectivity (in C5-Ci1 range) with an increase in silica content. The results from both studies are consistent with those of Dictor and Bell (1986), who reported that potassium promotion enhances branched hydrocarbon selectivity. Olefin Selectivity. 1-Alkenesare the primary products of the FT synthesis reaction over iron-based catalysts, and they may undergo two main secondary reactions: hydrogenation to n-paraffins and isomerization to 2-alkenes. Figure 6 illustrates the effect of support (content and type) on the total olefin selectivity, expressed as the mass fraction of linear olefins of the same carbon number. Potassium promotion is known to supress olefin hydrogenation (Anderson et al., 1952; Dry, 1981; Arakawa and Bell, 1983), which explains the high olefin selectivity ( 80% for c4-c15) obtained in the test of the base-line (unsupported) catalyst. In our recent study with unsupported precipitated iron catalysts (Bukur et al., 1990),the maximum value of the olefin content in the C4-CI5 range was approximately 70%, and was obtained with a 100Fe/ 1K catalyst. The increased olefin selectivity in the present study with the 100Fe/5Cu/4.2K catalyst is due to a higher potassium content of this catalyst. The two alumina-supported catalysts and the 8Si02/100Fe catalyst have essentiallythe same olefin content as the unsupported catalyst, whereas the olefin contents of the catalysts containing 24 and 100 parts Si02per 100 parts Fe are lower, particularly at higher carbon numbers. The bell-shaped curve obtained with the 100Si02/100Fe catalyst is typical N
1596 Ind. Eng. Chem. Res., Vol. 29. No. 8, 1990
t
/
n l
I
9
10
11
12
13
14
15
Carbon Number
Figure 7. Effect of support type and content on 2-olefin selectivity.
of catalysts having high olefin hydrogenation activity. Ethylene is more reactive than other low molecular weight olefins (C3-C6),whereas the decrease in olefin content with the increase in molecular weight may be attributed to either the increased reactivity of higher molecular weight olefins or to their greater absorptivity (Schulz et al., 1982; Schulz and Gokcebay, 1984; Dictor and Bell, 1986). The data indicate that the interaction between potassium and support (alumina or silica) is minimal at low concentrations of the latter (up to 24 parts support per 100 parts iron), but it increases with an increase in support content. The olefin selectivity decreases with an increase in silica content due to the reduction in surface basicity of the catalyst and/or decreased Fe/K contact due to high catalyst surface area. Egiebor and Cooper (1985) reported that the l-olefin content decreases with the increase in silica content and with increasing molecular weight for hydrocarbons in the C&, range. This is consistent with the results obtained in our study and with the statement made earlier in this paper that their catalysts had much stronger interactions between potassium and silica. Changes in reaction conditions (temperature, 220-250 "C; gas space velocity, 2-4 nL/(g of catalyst-h); and pressure, 1.5-3.0 MPa) had only a small effect on the olefin content in tests with all catalysts, as shown in Tables I11 and IV. This indicates that the differences in conversion did not have a marked effect on the results and that the trends depicted in Figure 6 reflect intrinsic differences among the catalysts. For example, low conversions obtained with the lO0SiO2catalyst, relative to those obtained in tests with the other five catalysts (see Table 111), are expected to favor primary products of synthesis. In spite of this, the 100Si02catalyst had the lowest olefin content amongst all catalysts tested. Also, differences in the exit H2/C0 molar ratio did not have an effect on product selectivities. In tests of the unsupported catalyst and the two catalysts containing small amounts of binder (8Si02 and 8Al2O3),the exit H2/C0 ratios during the first mass balance period (40-48 h) were much higher than those obtained a t 144 h on stream (Table 111). However, this did not have a marked effect on the olefin content nor on any other product selectivities (hydrocarbon, oxygenates, C02, 2-olefin), probably due to the fact that only a small portion of the catalyst bed was exposed to the syngas with high hydrogen content. Olefin Isomerization. Figure 7 illustrates the effect of support on 2-olefin selectivity as a function of carbon number. The results in the carbon number range 4-8 were omitted, as they showed erratic trends due to poor chromatographic separation between 1-and 2-alkenes during analysis. The addition of silica or alumina support in small
quantities (8-24 g of support per 100 g of Fe) did not have an effect on the selectivity of the unsupported catalyst. All these catalysts had a low isomerization activity ( 2 4 % internal olefins) that did not vary with carbon number. This behavior is the result of the high level of potassium promotion and lack of interaction between potassium and support. The catalyst with the highest support content ( 100Si02/100Fe) had a higher isomerization activity, and the 2-olefin selectivity increased markedly with carbon number, whereas the selectivity trends observed with the 24Si02/ lOOFe catalyst were between these two extremes. Egiebor and Cooper (1985) also reported the increase in the internal olefin selectivity of hydrocarbons in the C5-Cll range with an increase in the silica content. The results from both studies show that secondary reactions (olefin hydrogenation and isomerization) increase with increasing silica content of the catalyst, which may be attributed to the reduction in the effective potassium content of the catalyst. These trends are consistent with the results obtained in studies on the effect of potassium promotion (Dictor and Bell, 1986; Bukur et al., 1990), where it was observed that both olefin hydrogenation and isomerization decrease with an increase in potassium content of the catalyst. The potassium promotion enhances the absorption of CO and decreases the absorption of H2 and thus supresses olefin hydrogenation and isomerization reactions. The influence of the reaction conditions (temperature, gas space velocity and pressure) was also investigated, and an illustration of the results is presented in Figure 8. Changes in process conditions had only a very small effect on the 2-olefin selectivity of the unsupported catalyst (parts a and b of Figure 81, which is typical of catalysts with high potassium content. These catalysts show the same primary product selectivity over a wide range of process conditions. On the other hand, in tests of the 100Si02/lOOFe catalyst, which showed some propensity toward olefin hydrogenation and isomerization, the 2-olefin selectivity varied with process conditions. The internal olefin selectivity (i.e., isomerization activity) increased with an increase in temperature (Figure 8 4 , and decreased with increase in either the space velocity or the pressure (Figure ad). These trends are in agreement with those reported in the earlier studies with a variety of iron-based catalysts (e.g., Schulz et al., 1982; Schulz and Gokcebay, 1984; Dictor and Bell, 1986; Bukur et al., 1990) and can be rationalized by considering the effects of process conditions on secondary reactions. The increase in gas space velocity favors primary reactions, and therefore, the 2-olefin selectivity is expected to decrease. The internal olefin selectivity decreases with increase in pressure, probably due to increase in CO partial pressure (i.e., its higher surface concentration). With increasing temperature, the rates of both primary l-olefin formation and secondary reactions increase, and the internal olefin selectivity may either increase, decrease, or remain unchanged, depending on the values of the activation energies for these two reactions. With catalysts that have low isomerization activity, the internal olefin selectivity decreases with an increase in temperature (Figure 8a), whereas with catalysts that have higher isomerization activity, the internal olefin formation increases with temperature (Figure 8c). The results obtained with the two alumina-cintaining and 8Si02/lOOFe catalysts followed the trends observed with the unsupported catalyst (Le., changes in process conditions had a small effect on the 2-olefin selectivity), whereas the behavior of the 24sio2/lOOFe catalyst was qualitatively similar to that of the 100Si02/100Fecatalyst.
Ind. Eng. Chem. Res., Vol. 29, No. 8, 1990 1597 60
.
0 c z
0 0
-1.6 MPa, SV P-1.5 MPa. SV P-3.0 MPa. SV
--
2 nUg-cat.h 4 nUg-cat.h 4 nUa-cat.h
Catalyst 100 Fe/6 Cu/4 2 K/lOO Si02
c
c
c
N
-2 '
-P
1
15-
.-e1 .-e
-8
0
1
0
220°C. SV-2 nUg-ca1.h
0
0 26OoC ,SV-4 nL/g-cat.h
.
1.5 MPa. H2/CO-l. 2 nL/g-cclt.h 22ooc 25OoC
1.6 MPa. H2/CO-1
10
9
10
11
12
,
I
,
13
14
15
9
Carbon Number
10
11
12
13
14
15
Carbon Number
Figure 8. Effect of process conditions on 2-olefin selectivity. Influence of temperature on (a) 100Fe/5Cu/4.2K catalyst and (c) 100Fe/ 5Cu/4.2K/100Si02 catalyst. Influence of gas space velocity on (b) 100Fe/5Cu/4.2K catalyst and (d) 100Fe/5Cu/4.2K/100SiOz catalyst. 2
1
Carbon Number
Figure 9. Effect of support on alcohol selectivity.
However, the magnitude of changes in the 2-olefin selectivity was significantly smaller with the 24Si02/lOOFe catalyst in comparison to those observed with the 1O0SiO2/lOOFe catalyst. Oxygenates Selectivity. Oxygenates comprise only a small fraction of the products formed (only 2-5% of the CO converted, as shown in Tables I11 and IV) and consist primarily of n-alcohols and small amounts of aldehydes. The effects of support and process conditions on oxygenates selectivity were not clearly discernible, except for the fact that the 100Si02/lOOFe catalyst produced more oxygenates than other catalysts at all process conditions. Egiebor and Cooper (1985) reported that the alcohol contents of products in the C5-Cll range remained relatively constant with increasing amount of silica (2173Si02/FeFe). The alcohol selectivity, expressed as the percent of CO conversion to alcohols, as a function of carbon number is shown in Figure 9 for selected catalysts. The silica-supported catalyst produced significantly more methanol and ethanol than the unsupported and the 24A120,/ lOOFe catalysts, whereas there were no significant differences in selectivity of C4+products. The above results, interpreted in terms of potassiumsupport interactions, imply that selectivity of the oxygenates increases with a decrease in surface basicity (or with decreasing potassium content of the catalyst). This type of behavior would be expected from the CO insertion mechanism for formation of oxygenates (Anderson, 1984).
However, there is no general concensus in the literature on the effect of potassium on the selectivity of oxygenates, as can be seen from the following discussion. In our recent study on promoter effects on precipitated unsupported iron catalysts (Bukur et al., 1990), we could not correlate the selectivity of oxygenates with the potassium content of the catalyst. However, we observed that potassium promotion suppresses the formation of low molecular weight alcohols (methanol and ethanol) by shifting the product distribution toward higher molecular weight products. Yeh et al. (1985) reported that potassium promotion results in lower alcohol yields (they operated their reactor in a differential mode, and methanol and ethanol were the only alcohols detected). Likewise, Dictor and Bell (1986) and Donnelly and Satterfield (1989) reported that potassium promotion reduces methanol production. On the other hand, there have been several studies in which it was reported that the selectivity of oxygenates increases with potassium promotion. For example, Anderson et al. (1952) reported that the fraction of oxygenated molecules dissolved in the hydrocarbon fractions increased with increasing alkali content, but they did not provide any data on the selectivity of low molecular weight products. Dry (1981, p 182) presented results that show that the acid production (in the aqueous phase) increases with alkali content, particularly for the unsupported catalysts. For the silica-supported catalyst, the alcohol content first increased from 2.5 wt % a t a K 2 0 level of 12 (relative quantity) to 3.4 wt YO a t a K 2 0 level of 21 and then decreased, reaching 2.0% at a K 2 0 level of 32. Arakawa and Bell (1983) reported that the formation of methanol is essentially unaffected by potassium promotion, whereas the proportion of ethanol increases with potassium content of the alumina-supported catalyst (20 w t % of Fe on alumina, 1.5-14 g of K per 100 g of Fe). It appears that seemingly contradictory conclusions concerning the effect of potassium on the selectivity of oxygenates are caused in part by differences in the way selectivities were reported.
Conclusions The addition of silica or alumina to the precipitated unsupported iron catalyst (100Fe/5Cu/4.2K) influences
1598 Ind. Eng. Chem. Res., Vol. 29, No. 8, 1990
the catalyst activity, stability, and selectivity during Fischer-Tropsch synthesis. Changes in the catalyst performance become pronounced only a t sufficiently high binder (support) concentrations and are primarily due to interactions between support and potassium and/or between metal and support. The specific FTS activity decreased with increasing support content, which may be ascribed to a lower degree of iron reduction following the catalyst pretreatment with CO and decrease in the effective potassium content of the catalyst. However, the catalyst stability improved with the addition of silica either because of the stabilization of iron crystallites during synthesis or due to a decrease in the carbon deposition rate resulting from the reduction in surface basicity. The WGS activity decreased with increasing silica content of the catalyst (catalysts with 24 and 100 g of SiOz per 100 g of Fe), due to reduction in the effective potassium content of the catalyst. Product selectivities also changed markedly with increasing silica content of the catalyst, whereas no significant changes were observed with the two alumina-containing catalysts. The total olefin content and the fraction of branched hydrocarbons both decreased, whereas the fraction of internal olefins increased with an increase in the silica content of the catalyst. These observations reflect the expected trends arising from the reduction of surface basicity of the catalyst with increasing silica content and are in agreement with the results obtained in previous studies with similar catalysts (Dry, 1981; Egiebor and Cooper, 1985). Hydrocarbon selectivities showed some unusual trends with increasing silica content of the catalyst. Catalysts containing 8 and 100 parts SiOzper 100 parts Fe produced less low molecular weight hydrocarbons (C,-C,) than the unsupported catalyst, whereas the catalyst containing 24 parts SiOz produced more gaseous hydrocarbons. These results are indicative of complex metal-support interactions that require further investigation.
Acknowledgment We are grateful to Air Products and Chemicals, Inc., and the U.S. Department of Energy for financial support of this work under contract DE-AC22-85PC80011 and to Dr. E. B. Yeh for synthesizing the catalysts employed in this study. Registry No. CO, 630-08-0; Fe, 7439-89-6; Cu, 7440-50-8; K , 7440-09- 7”
Literature Cited Anderson, R. B. Catalysts for the Fischer-Tropsch Synthesis. In Catalysis; Emmett, P. H., Ed.; Van Nostrand-Reinhold: New York, 1956; Vol IV, pp 29-255. Anderson, R. B. The Fischer-Tropsch Synthesis; Academic Press: New York, 1984. Anderson, R. B.; Seligman, B.; Schultz, J. F.; Elliot, M. A. FischerTropsch Synthesis. Some Important Variables of the Synthesis on Iron Catalyst. Znd. Eng. Chem. 1952, 44, 391-397. Arakawa, H.; Bell, A. T. Effect of Potassium Promotion on the Activity and Selectivity of Iron Fischer-Tropsch Catalyst. Ind. Eng. Chem. Process Des. Dev. 1983,22, 97-103. Bartholomew, C. H.; Pannel, R. B.; Butler, J. L. Support and Crystallite Size Effects in CO Hydrogenation on Nickel. J . Catal. 1980,65, 335-349. Berry, F. J.; Liwu, L.; Dongbai, L.; Chengyu, W.; Renyuan, T.; Su, Z. An Investigation of Metal-Support Interactions in Some Fe, Ru and Fe-Ru Catalysts by in situ Fe5’ Mijssbauer Spectroscopy. Appl. Catal. 1986,27, 195-205. Brown, R.; Cooper, M. E.; Whan, D. A. Temperature Programmed Reduction of Alumina-Supported Iron, Cobalt and Nickel Bimetallic Catalysts. Appl. Catal. 1982, 3, 177-186. Bukur, D. B.; Brown, R. F. Fischer-Tropsch Synthesis in a Stirred
Tank Reactor-Reaction Rates. Can. J. Chem. Eng. 1987, 65, 604-612. Bukur, D. B.; Rosynek, M. P.; Lang, X.; Rossin, J. A,; Yeh, E. B.; Addiego, W. P.; Li, C.; Zimmerman, W. H. Activation and Promotion Studies in Fixed Bed Reactors with Precipitated Iron Fischer-Tropsch Catalysts. Seventh DOE Indirect Liquefaction Contractors Meeting Proceedings, Pittsburgh, Dec 7-9,1987; pp 41-68. Bukur, D. B.; Lang, X.; Patel, S. A.; Zimmerman, W. H.; Rosynek, M. P.; Withers, H. P. Development and Process Evaluation of Improved Fischer-Tropsch Slurry Catalysts. Eighth DOE Zndirect Liquefaction Contractors Meeting Proceeding, Pittsburgh, NOV15-17, 1988; pp 453-482. Bukur, D. B.; Lang, X.; Rossin, J. A.; Zimmerman, W. H.; Rosynek, M. P.; Yeh, E. B.; Li, C. Activation Studies with a Promoted Precipitated Iron Fischer-Tropsch Catalyst. Znd. Eng. Chem. Res. 1989, 28, 113C-1140. Bukur, D. B.; Mukesh, D.; Patel, S. A. Promoter Effects on Precipitated Iron Catalysts for Fischer-Tropsch Synthesis. Znd. Eng. Chem. Res. 1990,29, 194-204. Dictor, R. A,; Bell, A. Fischer-Tropsch Synthesis Over Reduced and Unreduced Iron Oxide Catalysts. J . Catal. 1986, 97, 121-136. Donnelly, T. J.; Satterfield, C. N. Product Distributions of the Fischer-Tropsch Synthesis on Precipitated Iron Catalysts. Appl. Catal. 1989,52, 93-114. Dry, M. E. The Fischer-Tropsch Synthesis. in Catalysis Science and Technology I; Anderson, J. R., Boudart, M., Eds.; SpringerVelag: New York, 1981; pp 159-255. Dry, M. E.; Oosthuizen, G. J. The Correlation Between Catalyst Surface Basicity and Hydrocarbon Selectivity in the FischerTropsch Synthesis. J. Catal. 1968, 11, 18-24. Egiebor, N. 0.;Cooper, W. C. Fischer-Tropsch Synthesis on a Precipitated Iron Catalyst: Influence of Silica Support on Product Selectivities. Can. J . Chem. Eng. 1985, 63, 81-85. Goodwin, J. G., Jr.; Chen, Y. W.; Chuang, S. C. Catalyst Support Effects of Selectivity in the Fischer-Tropsch Synthesis. In Catalytic Conoersions of Synthesis Gas and Alcohols t o Chemicals; Herman, R. G., Ed.; Plenum: New York, 1984; pp 179-189. Guczi, L. Structure and Catalytic Properties of Iron-Containing Bimetallic Catalysts. Catal. Reo.-Sci. Eng. 1981, 23, 329-376. Huff, G. A.; Satterfield, C. N. Evidence for Two Chain Growth Probabilities on Iron Catalysts in the Fischer-Tropsch Synthesis. J. Catal. 1984, 85, 370-379. Hughes, I. S. C.; Newman, J. 0. H.; Bond, G. C. The Characterization of Unsupported Iron and Manganese-Promoted Iron Catalysts by X-Ray Photoelectron Spectroscopy and Temperature-Programmed Reduction. App. Catal. 1987, 30, 303-311. Jung, H. J.; Walker, P. L.; Vannice, M. A. CO Hydrogenation Over Well-Dispersed Carbon-Supported Iron Catalysts. J. Catal. 1982, 75, 416-422. Kock, A. J.; Fortuin, H. M., Gecis, J. W. The Reduction Behavior of Supportrred Iron Catalysts in Hydrogen or Carbon Monoxide Atmospheres. J . Catal. 1985, 96, 261-275. Leith, I. R.; Howden, M. G. Temperature-Programmed Reduction of Mixed Iron-Manganese Oxide Catalysts in Hydrogen and Carbon Monoxide. Appl. Catal, 1988,37, 75-92. Li, C. Effect of Potassium and Copper Promoters on Reduction Behavior of Precipitated Iron Catalysts. Ph.D. Dissertation, Texas A&M University, College Station, 1988. Madon, R. J.; Taylor, W. F. Fischer-Tropsch Synthesis on a Precipitated Iron Catalyst. J . Catal. 1981,69, 32-43. McVicker, G. B.; Vannice, M. A. The Preparation, Characterization, and Use of Supported Potassium-Group VI11 Metal Complexes as Catalysts for CO Hydrogenation. J. Catal. 1980, 63, 25-34. Reuel, R. C.; Bartholomew, C. H. Effect of Support and Dispersion on the CO-Hydrogenation Activity/Selectivity Properties of Cobalt. J . Catal. 1984, 85, 78-88. Schulz, H.; Gokcebay, H. Fischer-Tropsch CO-Hydrogenation as a Means for Linear Olefins Production. in Catalysis of Organic Reactions, Kosak, J. R., Ed.; Dekker: New York, 1984; pp 153-169. Schulz, H.; Rosch, S.; Gokcebay, H. Selectivity of the FischerTropsch COHydrogenation. In Coal: Phoenix of ‘805, Proc. 64th C. I. C. Coal Symp.; AI Taweel, Ed.; Ottawa, 1982; pp 486-493. Snel, R. The Nature of Hydrocarbon Synthesis by Means of Hydrogenation of Carbon Monoxide on Iron Based Catalysts-Part I. J. Mol. Catal. 1989, 53, 129-141. Vannice, M. A.; Garten, R. L. Metal-Support Effects on the Activity and Selectivity of Ni Catalysts in CO/H2 Synthesis Reactions. J. Catal. 1979, .56, 236-248.
I n d . Eng. Chem. Res. 1990, 29, 1599-1606 Vannice, M. A,; Garten, R. L. Influence of the Support on the Catalytic Behaviour of Ruthenium in CO/H2 Synthesis Reactions. J . Catal. 1980,63, 255-260. Yeh, E. B.; Schwartz, L. H.; Butt, J. B. Silica Supported Iron Nitride in Fischer-Tropsch Reactions. 11. Comparison of the Promotion Effects of K and N on Activity and Selectivity. J. Catal. 1985, 91, 241-253.
1599
Zimmer", W. H.; R w i n , J. A.; Bukur, D. B. Effect of Particle Size on the Activity of a Fused Iron Fischer-Tropsch Catalyst. Ind. Eng. Chem. Res. 1989,28,406-413. Received for review September 26, 1989 Revised manuscript received January 29, 1990 Accepted April 9, 1990
Thermal Dehydration of Calcium Hydroxide. 1. Kinetic Model and Parameters Angel Irabien,* Javier R. Viguri, and Inmaculada Ortiz Departamento d e I n g e n i e r h Quimica, Facultad de Ciencias, Universidad del Pab Vasco, Apdo. 644, Bilbao 48080, S p a i n
In this work, the kinetic model describing the behavior of the dehydration reaction of calcium hydroxide in the range of temperatures 330-450 "C is reported. Two different types of solids have been utilized in dehydration tests: commercial calcium hydroxide, S = 8.3 f 1 m2-g-l and calcium hydroxide reagent obtained in the laboratory under controlled conditions, S = 18.7 f 1.4 m2*g-l. A discrimination of the reaction model using the structural parameter )I of the random pore model led to the best results for a value of = 0, which corresponds to a pseudohomogeneous kinetic model of the form
+
dx
- = kso exp(-E,/RT) dt
PO*exp(-AH/RT) RT So(1- x )
with kinetic parameters k, = 1.81 X lozocm-s-', E, = 280.4 kJ.mol-', po* = 1.834 X lo8 atm, and -AH = 138.5 kJ*mol-'. The suitability of the kinetic model to describe the behavior of calcium hydroxide during the dehydration process was confirmed by the results obtained in the correlation of the experimental data of calcium hydroxide reagent, a different solid with a higher surface area.
Introduction The most commonly used reagents in flue gas desulfurization dry processes are limestone and lime (Ortiz et al., 1987), although recent interest has focused on the hydration of calcium oxide to give calcium hydroxide and the subsequent calcination to produce a highly porous form of CaO (Marsh and Ulrichson, 1985). Lime sulfation is accompanied by product layer expansion, and thus, the final conversion level is favored by a highly porous structure. The reaction mechanism is not yet fully understood, but several suggestions have been reported (Bhatia and Perlmutter, 1981; Borgwardt and Bruce, 1986; Simons et al., 1987). The physicochemical properties of the reagent such as specific surface, pore volume, pore size distribution, and particle size distribution are the determining factors in the degree of gas desulfurization as well as in the utilization of the solid sorbent. Several authors have reported that CaO derived from hydrated lime, Ca(OH)2,is more reactive than that derived from the respective limestone, CaCO,, and even from commercial CaO, yielding higher ultimate conversions of the solid sorbent and thus resulting in greater sulfur capture by Ca(OH)2(Viguri et al., 1988; Bruce et al., 1989). Beruto et al. (1980) reported that solid products of decomposing Ca(OH)2powder at 320 "C under vacuum are particles that have approximately the same exterior dimensions as the parent Ca(OH)2particles, showing slitshaped geometry and having higher internal surface areas than does calcium oxide obtained from CaCO, calcination under vacuum at 510 OC.
* Author t o whom
correspondence should be addressed. 0888-5885/90/2629-1599$02.50/0
Under the conditions usually found in desulfurization processes, the reagent is calcium oxide produced during thermal treatment of calcium hydroxide or calcium carbonate:
A CaO + H 2 0
Ca(OH)2 CaC03
A
CaO
(1)
+ COP
(2) The kinetic behavior of the desulfurization reaction depends on the initial stage of calcium hydroxide dehydration or calcium carbonate calcination. Specially in the cme of calcium hydroxide reagent, dehydration takes place at medium temperatures (350-600 "C), depending on the concentration of water vapor in the gas phase, and sintering can interact due to the higher temperature of the desulfurization process (Borgwardt and Bruce, 1986; Mai and Edgar, 1989). In previous papers, Criado and Morales (1976) studied the mechanism df the thermal decomposition of calcium hydroxide and Mu and Perlmutter (1981) reported the thermal dehydration kinetics of a high-purity calcium hydroxide (98% by weight) under nonisothermal conditions, giving the temperature range of every compound resulting from the decomposition of the initial solid and the reaction kinetics. The results were extremely sensitive to heating rates and decomposition temperatures (325-415 "C), and dehydration kinetics were obtained at the very slow heating rate of 1 "C/min. The dehydration results were fitted to the linear form of the kinetic equation (Mu and Perlmutter, 1981) dx/dt = Ito exp(-E/RT)(l - x ) 2 / 3 (3) The estimated parameters were the activation energy, E 0 1990 American Chemical Society
