Ind. Eng. Chem. Res. 2000, 39, 4063-4069
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Impact of Iron Deposition on Pd/δ-Al2O3 in Selective Hydrogenation of Pyrolysis Gasoline Wei-Bin Su, Wan-Rong Chen, and Jen-Ray Chang* Department of Chemical Engineering, National Chung Cheng University, Chia-Yi, Taiwan
In selective hydrogenation of pyrolysis gasoline, Pd/Al2O3 is often deposited with iron species due to corrosion of equipment. The impact of deposited iron on the Pd/Al2O3 is examined. TPR (temperature-programmed reduction) results suggest that the clusters of deposited iron remain segregated from Pd clusters. By use of a model of Fe2O3-Pd/Al2O3 in hydrogenation of pyrolysis gasoline, the deposited iron is shown to increase isoprene conversion, olefin selectivity, and coke formation. TPD (temperature-programmed desorption) of NH3 and TPR results indicate that the deposited iron is in the form Fe2O3, which increases the acidity of the catalyst. The increased acidity promotes coke formation. Excess coke formation leads to high-pressure drops and eventually bed plugging. Introduction A two-stage hydrogenation process unit is used to stabilize pyrolysis gasoline and to reduce its sulfur content.1,2 The first-stage process, which normally uses δ-alumina-supported Pd catalysts for partial hydrogenation, is to saturate the gum precursors such as alkenyl-aromatics and conjugated diolefins to highoctane blending components, such as olefins.3,4 Following the hydrogenation process, depentanizer and deoctanizer were used to separate C5 (hydrocarbon compounds with one molecule containing five carbon atoms) and C9+ (hydrocarbon compounds with one molecule containing nine carbon atoms or more), respectively, from the pyrolysis gasoline stream. After the separation, the second-stage process is used to remove the sulfur-containing compounds by cobalt-molybdate catalysts For the first-stage hydrogenation, the whole process was sometimes shut down, resulting in a great economical loss, because of abnormal pressure drops. It has been reported that when the void fraction of catalyst bed decreases to a critical margin due to the accumulation of coke, crushed particles, or solid contaminants, the pressure drop rises suddenly.5 Since the most common contaminants in a refinery are corrosive iron, resulting from chloride corrosion, sulfide corrosion, or naphthenic acid corrosion,6 iron deposition was speculated to be one of the contaminants to shrink bed void of the first-stage hydrogenation process. The characterization results of the used catalysts unloaded from top-bed confirmed the speculation; the content of deposited iron was 0.54 wt %, which is even greater than that of palladium loading of 0.23 wt %. Recent research work showed that particles with diameters of 100 µm or larger will be strained out at the top of a reactor comprising 1-2 mm diameter catalysts, whereas particles smaller than 20 µm will penetrate the catalyst bed and undergo deposition on the catalysts.7 Since the deposited iron cannot be removed during the catalyst regeneration-reduction step, the effect of iron deposition on the catalysis of the * Telephone: 886-5-2720411 ext 6239. E-mail: chmjrc@ ccunix.ccu.edu.tw.
palladium catalyst should be investigated further, and this was the goal of this study. To implement the research goals, used commercial Pd catalysts were characterized to understand the nature of iron deposition, such as the amount of iron deposition along the catalyst bed and the existence of Pd-Fe bimetallic interactions. A model Fe2O3-Pd catalyst was prepared to facilitate the investigation of iron deposition effects; by using spent catalyst from commercial plant, the effects of metal agglomeration, metal migration, and coke formation are hard to differentiate from iron deposition. Experimental Section Material and Catalyst Preparation. The δ-Al2O3 support was prepared by calcining γ-Al2O3 with a particle size of about 2 mm (A2U, Osaka Yogyo) at 1000 °C for 6 h. The resulting material had a bulk density of 0.68 g/cm3. The BET surface area and pore volume measured with an Omnisop 360 analyzer were 82.4 m2/g and 0.57 cm3/g, respectively. The eggshell Pd catalysts were prepared by an impregnation technique with an excess of solution.3 Fifteen grams of support was brought to contact with a solution of 0.1068 g of Pd(CH3COO)2 in 150 mL of toluene, followed by removing the solvent by filtration, and then calcining at 350 °C for 6 h. The catalyst contained 0.31 wt % Pd, measured by inductively plasma optical emission spectroscopy using a JobinYvon JY-38 instrument. The model catalysts, Fe2O3Pd/Al2O3, were prepared from the eggshell Pd catalysts using the same impregnation technique; the prepared Pd catalysts were brought to contacted with a solution of Fe2(CH3COO)3, had the their solvent filtered off, and were calcined at 350 °C for 6 h. The model catalyst contained 0.08 wt % iron. Pd(CH3COO)2 was chosen as the catalyst precursor because the affinity between the precursor and the support is higher than that for the precursors commonly used in industry, and thus the catalyst with an eggshell profile can easily be prepared.3 Temperature-Programmed Reduction. The apparatus used for the temperature-programmed reduction (TPR) was described by Jones and McNicol.8 A gas stream of 10% H2 in argon passed through the catalyst
10.1021/ie000292c CCC: $19.00 © 2000 American Chemical Society Published on Web 10/11/2000
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sample (0.3 g) in a quartz reactor heated at 10 °C/min to 600 °C with a temperature-programmed furnace. The water produced by reduction was trapped into a column of silica gel. The amount of H2 consumption was detected with a thermal conductivity detector (TCD). The reduction temperature was monitored by a K-type thermocouple. Temperature-Programmed Desorption. A quartz tube was packed with a small amount (0.3 g) of catalyst sample. The catalyst sample was then reduced under the same operating conditions as those used in the catalytic performance test except at 1 atm. After the reduction, the sample was cooled to room temperature. When the system achieved steady state (20 mL/min He flow rate at 40 °C), 1 mL of NH3 was injected onto catalyst bed through He carrier gas. The injections were repeating until none of chemisorbed NH3 was detected with the TCD. Desorption experiments were then carried out with He flowing at 20 mL/min He at 1 atm by increasing catalyst bed temperature from 40 to 700 °C at 10 °C/min. Evolved NH3 was again monitored with the TCD. CO Chemisorption. A quartz tube was packed with about 1.0 g of catalyst sample. After the reduction, the tube was connected to a three-way ball valve, and the connection ports were carefully purged with He to inhibit any air contamination. After the system become steady (20 mL/min He flow rate and 35 °C), a 0.1-mL pulse of CO was repeatedly injected into the catalyst bed with He carrier gas until none of the pulse was chemisorbed. The amount of chemisorption was then calculated by summing up the proportions of all pulses consumed. Catalytic Performance Test. The activity test for prepared catalysts was carried out in a continuousdownflow fixed-bed reactor. The reactor is 0.924 m long with an inside diameter of 2.07 cm. The reaction temperature was controlled at 95 °C by a three-zone electric furnace with a PID temperature controller. The feed was prepared by mixing 10 wt % isoprene (Merck) in toluene (99.22% purity, no. 4 reforming plant, Chinese Petroleum Corp., Taiwan). The reactor was packed with 10 mL of catalyst mixed with inert SiO2 (Merck) in a ratio of 1:5 by volume. A gradient packing method was used so that the catalyst bed would have a nearly uniform temperature and the wall and bypassing effects would be minimized. The ratio of bed length to catalyst particle diameter was approximately 50; the axial dispersion effects are inferred to have been negligible.9 The upstream part of the reactor was a preheated zone filled with SiO2. The catalyst samples were reduced with 12 L (NTP)/h hydrogen at 100 °C and 30 atm for 10 h. The hydrogenation reaction was then carried out with a weight hourly space velocity of 18 h-1 (g of feed/(h g of catalyst)), at 95 °C, 30 atm, and a H2/isoprene molar ratio of 2.262. Up to 97% feed was recovered as reaction products in the material balance tests. The loss material is attributed to hydrocarbons deposited on the wall of the reaction system, on the catalyst, and on the surface of SiO2. A condenser trapped liquid products at -5 °C. Both gas and liquid samples were collected periodically and analyzed by a gas chromatography (Hewlett-Packard model 5890 A, FID model) coupled with a data processor (SP4270). A capillary column (Petrocol DH 150, 150 m × 0.25 mm i.d., 1.0 µm phase film) was performed with
Figure 1. Sampling point of spent commercial catalysts.
Figure 2. Metal distribution profile along reactor (b Pd, [ Fe).
He flowing at 20 cm/s, starting at 30 °C for 20 min and then increasing at a 10 °C/min temperature-programming rate until it reached 150 °C for 0.5 h (0.9 µL split 100:1). After the test reactions, the catalysts were unloaded and characterized by the temperature-programmed oxidation (TPO) and thermogravimetric analyzer (Setaram, TG/DTA92/DSC 121). The apparatus used for TPO is the same as that for TPR, whereas the flowing gas is O2. Results and Discussion Characterization of Contaminants along the Catalyst Bed. To understand the penetration depth of contaminants along the catalyst bed, samples were taken from eleven different layers, namely, T1-T4 and B1-B6 (Figure 1), of catalyst bed for analysis. Figure 2 shows the iron deposition profile along catalyst bed, measured by inductive plasma optical emission spectroscopy. The amount of iron deposition decreased sharply from the top of catalyst bed and approached 700 ppm at the bottom, while no migration of Pd had been observed. The result indicated that the catalysts act like a filter allowing the deposition of corrosive iron on Pd catalyst. However, consistent with the work of Narayan et al., some iron particles did penetrate into the catalyst bed.7 The deposition of corrosive iron on the Pd catalysts may change the catalytic properties and needs to be examined further.
Ind. Eng. Chem. Res., Vol. 39, No. 11, 2000 4065 Table 1. Characterizing Results of Spent Commercial Catalysts by XRFa element Al Fe Zn Cr Mn Ni Cu Pb Ca K Cl S Si P TOP layer ∨ T1 layer ∨ B6 layer ∨ a
∨ ∨ ∨
∨ ∨ ∨
∨ ∨ ∨
∨ × ×
∨ × ×
∨ × ∨
× ∨ ×
∨ ∨ ∨
∨ × ∨ ∨ ∨ ∨ ∨ ∨ ∨ ∨ ∨ ∨ ∨ ∨ ∨
Key: (detected), × (not detected).
Figure 3. XRF spectra of spent commercial catalysts: (A) TOP layer and (B) T1 layer, (C) B6 layer.
In addition to iron, which affects the catalytic properties the most, other minor elements such as sulfur, phosphorus, and calcium were also identified by X-ray fluorescence (XRF) analyzer (SIEMENS SRS-300) and were summarized in Table 1. In addition, Figure 3 shows XRF spectra of selected spent commercial catalysts. During the regeneration, a mixture of air and steam was used to burn out the deposited coke via oxidation. Notwithstanding that the deposited iron may come from the chloride corrosion, sulfide corrosion, and naphthenic acid corrosion, the catalyst regeneration step oxidizes all iron species to iron oxide. Characterization of Used Catalysts by Temperature-Programmed Reduction. As shown in Figure 4a, the maximum reduction rates are at 470 and 380 °C for Fe2O3 and Fe3O4, respectively. The TPR of the used catalysts show two peaks at 90 and 470 °C, respectively (Figure 4b). The peak at 90 °C is characteristic of PdO (Figure 4b), and that at 470 °C is characteristic of Fe2O3. The TPR for the PtRe/γ-Al2O3 catalysts indicates that platinum catalytically reduces rhenium oxide resulting in a decrease of reduction temperature for rhenium oxide on the catalyst.10 Inferred from these results, the TPR of the used catalysts shows no significant change in the reduction temperature of Fe2O3, suggesting that segregated Fe2O3 and Pd particles are formed on the alumina surface. Moreover, the iron species deposited on the Pd catalysts is not expected to be reduced, because the maximum commercial plant operation temperature is 150 °C and which is far lower than the reduction temperature required for Fe2O3.11 Characterization of Model Catalysts. It has been reported that the catalyst deactivation of the eggshell Pd catalysts for pyrolysis gasoline selective hydrogenation is mainly caused by coke or gum formation, metal agglomeration, and metal migration.12-14 Except for coke deposition, which can be burnt up in the regenera-
Figure 4. (a) TPR profiles for (A) Fe3O4, (B) Fe2O3. (b) TPR profiles for (A) Fe2O3-PdO/Al2O3, (B) spent commercial catalyst at TOP layer, and (C) PdO/Al2O3, (D) Fe2O3.
tion, catalyst deactivation caused by other factors cannot be recovered. If used and fresh catalysts were compared in the reaction tests, we are unable to differentiate the effects of metal agglomeration and metal migration from iron deposition. Hence, model catalysts, eggshell Pd/ Al2O3 and Fe2O3-Pd/Al2O3, were prepared for test reactions in this study. Table 2 shows that the impregnation of iron precursor and the preceding calcination does not affect the surface area and average pore size of the support. Moreover, as shown in Figure 5, the XRD results indicated that both catalysts remained in the original δ-Al2O3 form after the treatment steps of catalyst preparation. Pd(CH3COO)2 and Fe2(CH3COO)3 were purposely used as catalyst precursors. The high affinity between these organometallic compounds and δ-Al2O3 retards the diffusion of these species into the interior of the δ-Al2O3 pellet, leading to the formation of catalysts with eggshell metal profile.3 No catalytic reduction of Fe2O3 was observed from TPR characterizing the model Fe2O3Pd/Al2O3 catalysts. This result suggests that the model catalysts do not have bimetallic interactions and is consistent with the properties of the used catalyst unloaded from the commercial plant. Thus, the Fe2O3-
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Table 2. Physicochemical Properties of Model Catalysts and Commercial Catalysts support BET surface area, m2 pore size, Å pore volume, mL/g bulk density, g/cm3 Pd, wt % Fe, wt % support type metal distribution CO uptake
model catalysts
commercial catalysts
Al2O3
PdO/Al2O3
Fe2O3-PdO/Al2O3
fresh
spent TOP layer
82.4 94.0 0.57 0.78 ***** ***** δ-form ***** *****
79.6 95.8 0.44 0.84 0.31 ***** δ-form egg-shell 0.47
79.8 93.0 0.41 0.87 0.308 0.08 δ-form egg-shell 0.49
80.1 ***** 0.53 0.82 0.306 ***** δ-form egg-shell *****
68.9 88.6 0.43 0.854 0.305 0.54 δ-form egg-shell *****
Figure 7. Reaction scheme for the selective hydrogenation of isoprene.
Figure 5. XRD spectra of model catalysts: (A) PdO/Al2O3 and (B) Fe2O3-PdO/Al2O3.
main TPD peak for the Fe2O3-Pd/Al2O3 catalysts was consistent with that of Pd/Al2O3, whereas an additional peak at about 380 °C representing a new acid site was observed. Consistent with the results of Fe-modified AlPO4, the introduction of Fe2O3 leads to catalysts of higher acidity.16 Together with all these characterization results, we conclude that the changes in catalytic properties resulting from iron deposition are mostly caused by the new acid sites. Effect of Iron Deposition on Catalysis. Isoprene was chosen as the model compound for the selective hydrogenation of pyrolysis gasoline. The reasons for choosing isoprene have been explained in our previous paper.13 The reactions involve hydrogenation of isoprene and subsequent isomerization of isopentenes. Concomitantly with the hydrogenation, isoprene polymerizes to coke or gum, leading to catalyst deactivation. The simplified reaction scheme is depicted in Figure 7. On the basis of the reaction scheme, the isoprene conversion, selectivity to pentenes, and selectivity to coke formation are formulated by the following equations.
isoprene conversion ) (IP reacted)/(IP fed) (1) selectivity to pentenes ) (pentenes produced)/ (IP reacted) (2) Figure 6. NH3 TPD profiles for δ-Al2O3 (-), Pd/Al2O3 (- - -), Fe2O3-Pd/Al2O3 (‚ ‚ ‚).
selectivity to coke ) 1 - (pentenes + pentane produced)/(IP reacted) (3)
Pd/Al2O3 prepared is a good model catalyst for the commercially used catalysts. As shown in Table 2, there are no significant differences in palladium loading and metal dispersion for both catalysts, suggesting that both catalyst samples have the same active Pd sites. The acidities of δ-Al2O3 support, Fe2O3-Pd/Al2O3, and Pd/Al2O3 catalysts were probed by using NH3 TPD, since the NH3 desorption temperature and the gas evolved are considered, respectively, as indexes of acids strength and of acid amount.15 Comparison of the NH3 TPD chromatograms of the three samples allows us to reach conclusion on the effects of Fe2O3 deposition on the catalyst acidity. As shown in Figure 6, the NH3 desorption peak for the δ-Al2O3 support and the Pd/Al2O3 catalysts is at about 110 and 300 °C, respectively. The
The total conversion of isoprene catalyzed by Fe2O3Pd/Al2O3 and Pd/Al2O3 is shown as a function of time on stream in the flow reactor (Figure 8a). For the reaction catalyzed by Pd/Al2O3, the isoprene conversion increased with time on stream to about 35 h on stream and then declined slightly. Within the induction period, the catalysts were wetted gradually by isoprene, and thus the isoprene conversion increases with time on stream. At the same time, initially high coking decreases reaction rate significantly. Then, a maximum of conversion was observed. In contrast, the Fe2O3-Pd/ Al2O3 catalysts also presented a maximum conversion, whereas the induction period is much shorter and the variation of conversion during the test period is not as pronounced as that for Pd/Al2O3. The detail kinetic study for the induction period has not been attempted. However, we might speculate that the increase of the
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Figure 9. TPO spectra for the used Fe2O3-Pd/Al2O3 (-) and used Pd/Al2O3 catalysts (- - -).
Figure 8. Effects of iron deposition on (a) total conversion of isoprene hydrogenation and selectivity to (b) coke and (c) pentenes. Conditions: temperature ) 95 °C, P ) 30 atm, LHSV ) 18 h-1 (2 Pd/Al2O3, 1 Fe2O3-Pd/Al2O3).
affinity between isoprene and the catalysts caused by the deposition of Fe2O3 promotes the dispersion rate of reactants, resulting in a shorter induction period. The selectivity to coke formation for the Fe2O3-Pd/ Al2O3 catalysts decreased with time on stream, whereas that for the Pd/Al2O3 catalysts increased slightly to a maximum at about 15 h on stream and then declined smoothly (Figure 8b). The difference in catalytic proper-
ties observed for these two catalysts was clearly caused by the deposition of Fe2O3; Fe2O3 on the catalysts enhance coke formation resulting in a higher isoprene conversion. Since the TPR characterizing the Fe2O3Pd/Al2O3 catalysts indicates that Pd and Fe2O3 are segregated, Fe2O3 may act as active sites for coke formation or promote the formation of coke on support. However, the coke formation rate declines with the coverage of Fe2O3 by coke leading to reduction in the differences between two catalyst samples in total conversion and selectivity to coke. Similar to the effects of sulfur, water, and CO poisoning,13,14,17-19 the deposition of coke lessens the adsorption of hydrogen on Pd sites and thus retards further hydrogenation of isopentenes. As shown in Figure 8c, for both catalysts, the selectivity to pentenes increases with coke deposition. At the start of the run, the Fe2O3-Pd/Al2O3 catalysts have lower selectivity to isopentenes because of higher conversion of isoprene to coke. However, after more Pd sites were covered by coke, a higher selectivity to isopentenes was observed for Fe2O3-Pd/Al2O3. In this study, coke deposited on the catalysts can roughly be categorized as soft and hard coke. During the temperature-programmed oxidation (TPO) experiments, coke deposited on the catalysts was burnt out with the consumption of O2 and the evolution of H2O, CO, and CO2. Except H2O, which was adsorbed on silica gel, all the other effluent gas contributed to TPO signals; consumption of O2 contributed to a positive signal while evolution of CO and CO2 contributed to a negative signal. For soft coke, the ratio of hydrogen to carbon is relatively higher and easier to be burnt out at lower temperatures. Since the TPO signal from O2 consumption is more pronounced than those of CO and CO2, and thus a positive peak at 250 °C was observed. In contrast, the negative peak at about 550 °C is the characteristic peak for burning hard coke. An extra negative peak at about 300 °C was also observed for the Fe2O3-Pd sample. During the selective hydrogenation of isoprene on Pd catalysts, coke may be formed by free radical polymerization of olefins, catalytic polymerization of olefins, and successive Diels-Alder cycloaddition (the reaction of a diene with an alkene to form a cyclohexane) associated with hydrogen transfer reactions.20,21 Be-
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void space of catalyst bed, leading to an excessive pressure drop. Acknowledgment The support of the National Science Council of the Republic of China (Contract No. NSC 88-2214-E-194005) and the Refining & Manufacturing Research Center of the Chinese Petroleum Corp. is acknowledged. Nomenclature
Figure 10. TGA spectra for the used Fe2O3-Pd/Al2O3 (-) and used Pd/Al2O3 catalysts (- - -).
cause of the complexity of the reactions and difficulty in coke characterization, we are unable to discriminate the structure of the coke that contributes to the peak at 300 °C from the others, and relate coke structure with different active sites. The TPO profiles characterizing Pd/Al2O3 and Fe2O3Pd/Al2O3 catalysts indicated that the area of the peak at 250 °C for Fe2O3-Pd/Al2O3 is about 2.2 times that for Pd/Al2O3 whereas the peak at 550 °C is only about 1.2 times (Figure 9). These results suggested that the deposition of Fe2O3 on the Pd catalysts promotes mostly soft coke formation. As shown in Figure 10, the TGA experiments were also consistent with the conclusion. Conclusion The characterization results of spent catalysts taken from pyrolysis gasoline selective hydrogenation units indicate that smaller iron fine particles (
