Effect of Pressure on the Hydrocracking of Light Cycle Oil with a Pt–Pd

Article pubs.acs.org/EF

Effect of Pressure on the Hydrocracking of Light Cycle Oil with a Pt− Pd/HY Catalyst Alazne Gutiérrez,* José M. Arandes, Pedro Castaño, Martin Olazar, and Javier Bilbao Departamento de Ingeniería Química, Universidad del País Vasco, Apartado 644, 48080 Bilbao, Spain ABSTRACT: The effect of pressure (35−75 bar) has been studied on the hydrocracking of light cycle oil (LCO) on a Pt−Pd/ HY catalyst. The remaining operating conditions are as follows: 400 °C; H2/LCO molar ratio (nH2), 10 molH2 (molLCO)−1; space velocity (WHSV), 2 h−1; time on stream, 0−24 h. The reaction indices studied are the conversions of hydrocracking and hydrodesulphurisation, the selectivity of naphtha and medium distillates, and the concentrations of these fractions. It has been proven that once an initial deactivation period has elapsed, the Pt−Pd/HY catalyst is very stable and has a high capacity for producing naphtha and medium distillates at 400 °C. Furthermore, the catalyst is fully regenerated by coke combustion with air at 550 °C. An increase in pressure allows reaching a pseudostable state with higher activity for hydrodesulphurisation (conversion 0.96 under 75 bar) and hydrocracking (conversion 0.87). Naphtha selectivity increases as pressure is increased and is 68% under 75 bar for a conversion of 98%, with overcracking being insignificant. The concentration of aromatics in the naphtha is 30 wt %, and it is therefore suitable for the gasoline pool. Lower values of space velocity are to be used in the hydrocracking to reduce aromatic concentration further.

1. INTRODUCTION The oil industry faces the challenge of an increasing demand for automotive fuel and raw materials in a global market in which oil depletion and severe environmental regulations on fuel composition have a major impact.1 Within this scenario, sustainable refineries should adopt measures such as optimization and unit integration, the upgrading of heavy oil and secondary streams, and cofeeding nonstandard streams, such as biomass or materials derived from postconsumer wastes (plastics and tires). The highest output refinery units involve catalytic cracking (FCC), hydrocracking, and the coker unit, and therefore, efforts are focused on these units by planning and refurbishing existing units to adapt them to new feeds and research on the technological development of new units.2−6 In order to further the aforementioned objectives, considerable attention has been paid in the literature to the study of cracking under FCC conditions by feeding atmospheric and vacuum distillation residues,7,8 naphtha and coker and visbreaker gas oil,9,10 as well as nonstandard feeds, such as bio-oil (liquid product from the flash pyrolysis of lignocellulosic biomass),11−13 waxes produced by polyolefin pyrolysis (which are fed pure or dissolved in vacuum gasoil, VGO),14,15 and the plastics themselves dissolved in VGO.16−18 FCC units have undergone technological changes consisting of improving catalysts and the design of reaction−regeneration units in order to intensify the production of light olefins and cofeed feeds that are heavier than standard ones. Nevertheless, high yields of heavy aromatic fractions are obtained due to the low cracking capacity of aromatic rings.19,20 Likewise, apart from coke, the coker unit generates naphtha and gasoil with a high aromatic content whose upgrading by cracking in the FCC unit is hindered by the rapid deactivation of the catalyst caused by coke deposition.21,22 Light cycle oil (LCO) is a secondary stream in the catalytic cracking unit (FCC), with a boiling point similar to diesel. Its © 2012 American Chemical Society

production is increasing in step with the predicted role of the FCC unit for meeting the increasing demand for light olefins and the increasing trend of feeding heavy feeds into the FCC unit.23 Nevertheless, due to the high content of aromatics (with a significant content of polycyclic compounds) and sulfur, and a low cetane index,24 the use of LCO in the formulation of automotive diesel fuel is severely limited. Several strategies have been followed to correct the imbalance between production and demand: (i) production of a better quality LCO in FCC;25 (ii) recirculation in FCC;26 (iii) hydrocracking.27−32 Hydrocracking is a process with good perspectives for upgrading oil-derived heavy and aromatic (hydrodearomatization) streams as well as other feeds derived from biomass or plastic waxes.1 The hydrocracking of complex feeds with a low S content, such as the pyrolysis gasoline (the aromatic stream produced in steam cracking), has proven to be versatile for producing fuels and raw materials.33−35 Nevertheless, the implementation of hydrocracking units in refineries faces severe limitations associated with their unit cost and difficulties related to the activity and stability of catalysts, given that they undergo severe deactivation in the valorization of heavy component streams with significant S and N content.36,37 Consequently, the hydrocracking of real feeds needs to be studied in order to improve catalysts and establish process conditions that minimize these problems. Although bifunctional catalysts with transition metals (Co, Mo, Ni, V, W, or their combinations) are efficient for hydrodesulphurisation, they have moderate activity for the transformation by hydrocracking of LCO into naphtha and medium distillate streams of suitable composition for the pools of gasoline and refinery diesel.38,39 The noble metal catalysts Received: June 4, 2012 Revised: July 30, 2012 Published: July 31, 2012 5897

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In order to synthesize the catalyst, wet impregnation of the support has been carried out at 80 °C with an aqueous solution of Pt(NH3)4(NO3)2 (Alfa Aesar) and Pd(NH3)4(NO3)2 (Stem Chemicals). The amount of noble metal required is added to 100 mL of deionized water by maintaining pH = 7. Once the adsorption equilibrium has been reached, the excess water is removed in a Rotavapor. The catalyst obtained is dried in an oven for 24 h at 120 °C and subsequently calcined in air for 2 h at 450 °C reaching this temperature following a ramp of 5 °C min−1. The results obtained in the characterization of the catalyst are summarized in Table 1. Interestingly, the volume of micropores

(Pt and Pd) are more active for aromatic hydrocracking, which is enhanced using HZSM-5, HY, and Hβ zeolites as support.35 Nevertheless, these noble metal catalysts are of low stability and liable to fast deactivation in the hydrocracking of feeds with heteroatomic aromatics containing S and N. Accordingly, these feeds require steps prior to hydrocracking (hydrodesulphurisation, hydrodenitrogenation, and hydrodemetallisation) using transition metal catalysts, hindering the industrial implementation of processes for upgrading streams such as LCO.40,41 To further investigate on the hydrocracking of LCO in a single-step (without prior steps for removing S, N, and metals), studies have been carried out on the suitable composition of the catalyst and effect of operating conditions, such as space velocity and temperature.27−31 The role played on catalyst stability by the acid properties of HY and Hβ zeolites used as support has been determined in order to reach a pseudostable state of the catalyst with significant residual activity proportional to the acidity of the fresh catalyst.27 These results complement those obtained in the literature for the hydrocracking of model compounds, which reveal an increase in the thioresistance of Pt and Pd catalysts when acid zeolites are used as support.35 Furthermore, a good performance has been reported for high acidity HY zeolites in attenuating coke formation, which is attributed to the capacity of these zeolites for cracking coke precursors, diminishing their concentration and attenuating their condensation toward polyaromatics structures.32 Furthermore, the composition of the coke depends heavily upon the features of the support: acidity and micropore topology. It has also been proven that the configuration of the metallic function obtained by combining both metals, Pt−Pd, improves both the hydrocracking of LCO and the thioresistance of catalysts,42−45 which is also consistent with the results in the literature for the hydrocracking of pure aromatics.46−49 Studying the effect of space velocity (WHSV) when Pt−Pd/ HY catalyst is used at 350 °C and 40 bar30 has revealed its thioresistance and the need for working with a space velocity lower than 1 h−1 for obtaining high yields of naphtha and medium distillates. However, moderate dearomatization is achieved at 350 °C. As temperature is increased in the 350− 400 °C range, S retention on the metallic function decreases, which is effective for increasing the activity remaining in Pt− Pd/HY catalysts at pseudostable state in order to give way to high yields of naphtha and medium distillates (45 wt % each at 400 °C), with low capacity for overcracking to LPG and dry gas. A temperature increase in the 350−400 °C range improves the composition of naphtha and especially of medium distillates, by decreasing the concentration of aromatics and increasing that of paraffins and naphthenes.31 In order to increase the selectivity of the naphtha stream in the hydrocracking of LCO and decrease its concentration of aromatics (hydrodearomatization), which is required for its incorporation into the pools of gasoline and diesel oil, this paper studies the effect of pressure and H2/LCO molar ratio (nH2) on the composition of these streams.

Table 1. Properties of the Catalyst Used property

Y

SiO2/Al2O3 Pt (gmetal 100 gcat−1) Pd (gmetal 100 gcat−1) Vp (cm3 gcat−1) Vmp (cm3 gcat−1) Sg (m2 gcat−1) Sm (m2 gcat−1) dm (nm) acidity μmolNH3 gcat−1 Sa (kJ (molNH3)−1) B/L

5.2

0.41 0.30 773.5

928 116 2.49

Pt−Pd/Y 5.2 1.12 0.93 0.40 0.25 667.8 0.85 5 686 109 0.68

decreases more noticeably than that of pores, indicating that the deposition of Pt−Pd is inside or in the mouth of the micropores. On the other hand, the acidity decreases more severely than the area, indicating that Pt−Pd is mainly depositing on the strongest acid sites (see also the decrease of the acid strength) and blocking the accessibility to Brönsted and weaker acid sites. Metal content has been determined by ICP-AES (inductively coupled plasma-atomic emission spectroscopy) in a Horiba Jobin Yvon Activa. Prior to the analysis, the solid sample has been subjected to acid digestion with HF (Merck) at 90 °C. The properties of the porous structure (Figure 1) have been determined from the N2 adsorption−desorption isotherms recorded at −196 °C in a Micromeritics ASAP 2010. The pore size distribution in terms of volume has been measured using BJH method, so that the micropores are excluded from the analysis. The metal surface and metal particle size have been determined by H2 chemisorption in a Micromeritics ASAP 2010C, according to the double isotherm method.50 The analysis has been carried out at 100 °C,51 or 70 °C,52 depending on the metallic phase. Crystallinity has been determined by XRD diffractometry in a Philips X’Pert MPD, using Bragg−Brentano Theta-2Theta geometry, a PW3123/00 secondary graphite monochromator, and a PW3011 scintillation counter. Figure 2 shows that the incorporation of Pt and Pd causes a severe loss of HY zeolite crystallinity. Furthermore, the diffraction peaks characteristic of Pt are masked and Pd peaks overlap those of HY zeolite. The metal particle size, calculated by H2 chemisorption, was 10 ± 3 nm. These values were verified by microscopy (vide infra). Figure 2 shows the peaks corresponding to Pt and Pd. However the higher intensity of the peaks corresponding to the support (HY zeolite), masks the peaks of the metals and makes it inappropriate to estimate metal particle sizes from these data. Total acidity has been measured by the temperature programmed desorption (TPD) of NH3 adsorbed at 150 °C in a TG-DSC Setaram 111 calorimeter provided with a Harvard Apparatus syringe pump and connected inline with a Balzers Quadstar 422 mass spectrometer.53 Acid strength has been determined by using calorimetry and thermogravimetry to simultaneously monitor the differential adsorption of NH3 at 150 °C in a Setaram TG-DSC111.54 The BrönstedLewis (B/L) acid site ratio has been calculated from the vibrational bands of pyridine adsorbed at 1547 cm−1 (pyridine associated with

2. EXPERIMENTAL SECTION 2.1. Catalyst. The acid support HY (ultra stable FAU zeolite, CBV500, Zeolyst International, SiO2/Al2O3 = 5) supplied in ammonium form has been calcined to obtain its acid form according to the following steps: (i) 2 h at 400 °C (5 °C min−1), (ii) 15 h at 500 °C (5 °C min−1); (iii) 2 h at 550 °C (5 °C min−1). 5898

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Brönsted) and that at 1455 cm−1 (pyridine associated with Lewis) using the molar extinction coefficients propounded by Emeis et al.55 Use has been made of a Nicolet 740 SX FTIR provided with a Specac transmission catalytic cell equipped with a temperature controller and a Vacuubrand rotary vacuum pump. 2.2. LCO Properties. The LCO has been supplied by Repsol YPF (Tarragona, Spain), and its composition (Figure 3) has been

Figure 3. LCO composition. determined by gas chromatography/mass spectrometry (ShimadzuQP2010-S). The total aromatic content is 66.83 wt %, with two-ring aromatics (methyl naphthalene as main component) as the major fraction (28.32 wt %). Furthermore, there is a significant amount of three-ring aromatics (16.44 wt %). Table 2 shows the results of the LCO analysis. As observed, it is a stream with a high sulfur content (0.5 wt %) and a low cetane number

Table 2. Properties of the LCO density, kg L−1 (ASTM D 4052) simulated distillation (ASTM D86-05) IBP−FBP, °C T50−T95, °C cetane number ASTM D 4737 ASTM D 976 chemical analysis (wt %) C H S N

Figure 1. Isotherms of N2 adsorption−desorption (a) and pore size distribution (b) for Y zeolite support and Pt−Pd/Y catalyst.

0.936 101.5−466.0 277.3−392.7 27.64 27.31 88.7 10.2 0.5 0.2

due to its high aromatic content. Total sulfur content has been determined by X-ray fluorescence (XRF) (Philips MiniPal PW-4025). The elemental analysis has been carried out in a EuroVector Euro EA Elemental Analyzer (CHNS), confirming the results of sulfur content obtained by XRF. Density has been determined according to the ASTM D 4052 standard, and the cetane number according to ASTM D 4737 and D 976 standards. Simulated distillation has been carried out as per the ASTM D 2887 standard. 2.3. Equipment and Conditions for Reaction and Analysis. The automated reaction equipment used has been described elsewhere31 and is provided with a 15 cm3 downflow fixed bed reactor (8 mm internal diameter and 303 mm length). The reaction conditions are as follows: 400 °C; 30−75 bar; H2/LCO molar ratio (nH2), 10 molH2 (molLCO)−1; space velocity (WHSV), 2 h−1; time on stream (TOS), 0−24 h. In order to attain bed isothermicity, the catalyst (0.15−0.30 mm pellets) is mixed (1/1 mass ratio) with CSi

Figure 2. X-ray diffractograms of Pt−Pd/Y catalyst and Y zeolite.

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(0.5 mm). In order to avoid gas bypassing and heat losses, CSi layers of around 20 mm thickness are placed above and below the catalyst bed. Prior to the reaction, the catalyst is activated in situ under atmospheric pressure with a stream of H2/N2 mixture (30 mL min−1 of H2, 50 mL min−1 of N2), raising the temperature following a ramp of 5 °C min−1 from ambient temperature to 350 °C and maintaining this temperature for 4 h. The reaction products are sent to a gas/liquid separator, and the gases are analyzed online in a VARIAN CP-4900 MicroGC provided with four channels: (i) a molecular sieve to separate the permanent gases H2, O2, N2, methane, and CO2; (ii) a Porapak Q to separate C2 hydrocarbons, CO2 and H2S; (iii) Al2O3 to separate C3 and C4 hydrocarbons; and (iv) CPSiL to separate C5−C10 hydrocarbons. The liquids are analyzed in a Hewlett-Packard 6890 gas chromatograph provided with an FID and a PONA capillary column (50 m × 2 mm × 0.5 mm).

3. RESULTS 3.1. Product Fractions and Reaction Indices. In order to facilitate the comparison of the results corresponding to the different conditions studied, the components of the LCO and reaction products have been grouped into the following lumps according to their boiling point (approximately corresponding to carbon atom number): dry gas (DG) (C1−C2); liquefied petroleum gas (LPG) (C3−C4); naphtha (N) (C5−C12) (36− 216 °C); medium distillates (MD) (C13−C20) (216−343 °C); heavy cycle oil (HCO) (C21+) (343 °C+).31 The analysis of the product stream by online GC/MS allows identifying of the components and their grouping into the aforementioned lumps. The main objective of this work is to convert the heaviest fraction of LCO, which is HCO, so that the conversion has been defined as the disappearance of HCO. Although many other definitions are possible, we observed that this conversion also provides an estimation of the hydrocracking of low-addedvalue compounds (boiling point lower than 260 °C). X=

(mHCO)i − (mHCO)o (mHCO)i

(1)

where (mHCO)i and (mHCO)o are the mass flow rates of the HCO fraction in the feed and reactor outlet streams, respectively. The hydrodesulphurization conversion has been defined as XHDS =

Si − S o Si

Figure 4. Effect of pressure on the evolution of conversion with time on stream for hydrodesulphurisation (a) and hydrocracking (b).

conversion (Figure 4a) and around 10 h for hydrocracking conversion (Figure 4b). Qualitatively similar results have been obtained under other reaction conditions,28−31 and they are explained by deactivation occurring for two reasons:32 (i) coke deposition and (ii) sulfurization of the metallic function. Both coke deposition and sulfurization reach a pseudoequilibrium state depending on the reaction conditions. Coke deposition reaches an equilibrium level due to the hydrocracking of coke precursors, which is enhanced by the acidity of the support,28,32 whereas the equilibrium of Pt−Pd metallic function sulfurization depends mainly on temperature, whose increase gives way to a decrease in adsorption capacity. It should be noted that there is a dependency between the thioresistance of the catalyst and its remaining capacity for hydrocracking coke precursors. Accordingly, the catalyst used with a thioresistant (Pt−Pd) metallic function and a strongly acid (HY zeolite) function allows reaching pseudoequilibrium states with considerable residual activity, which remains constant with time on stream for hydrodesulphurisation and hydrocracking. The hydrocracking of model compounds revealed that the incorporation of Pd into Pt catalysts improves thioresistance,42−44 which is

(2)

where Si and So are the mass flow rates of sulfur in the feed and product streams, respectively. The selectivity of each i fraction is defined by considering the yields of all the product fractions: Si =

Yi YDG + YLPG + YN + YMD

(3)

where the yield of each i product fraction is calculated from the mass flow rate of i at the reactor outlet, (mi)o: Yi =

(mi)o (mLCO)i

(4)

3.2. Effect of Pressure on Deactivation. Figure 4 shows the effect of pressure on the evolution of hydrodesulphurisation and hydrocracking conversions with time on stream. Both conversions decrease with time, until they reach a pseudostable state that remains constant with time on stream. The time required to reach this state is 6−8 h for hydrodesulphurization 5900

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due to a synergetic effect of the support acidity on the thioresistance of the metallic function by the formation of electron-deficient metallic particles, Ptδ+ or Pdδ+.45,46 Figure 4 shows the favorable effect of increasing pressure on the residual activity of the catalyst at the pseudostable state for hydrodesulphurization (Figure 4a) and hydrocracking (Figure 4b). Under 75 bar, the conversion of hydrodesulphurization at the pseudostable state is 96%, and 87% for hydrocracking. The literature reports that an increase in pressure enhances the hydroprocessing reactions of residual feeds and heavy streams and hinders the catalyst deactivation rate;56 furthermore, the results also show that increasing hydrogen pressure does not have a negative effect on conversion due to site blockage.57 Furthermore, it should be noted that the runs in Figure 4 correspond to a space velocity, WHSV, of 2 h−1. In a previous paper,30 the significant effect of space velocity has been reported, and therefore, the results in Figure 4 would be considerably improved with a lower value of space velocity. 3.3. Effect of Pressure on the Conversion and Selectivity of Naphtha and Medium Distillates. From an industrial point of view, the more interesting results correspond to the pseudostable state of the catalyst (Figure 5). In this state, the conversion of hydrodesulphurization increases almost linearly with pressure in the 35−75 bar range (Figure 5a). The effect of pressure on product stream composition is highly significant above 45 bar (Figure 5b), i.e., the concentration of naphtha increases and that of medium distillates decreases, with the concentration of HCO fraction decreasing significantly as pressure is increased from 60 to 75 bar. Figure 5c shows an exponential increase in hydrocracking conversion and naphtha selectivity as pressure is increased, reaching values of 0.87 and 0.74, respectively, under 75 bar. Figure 6 shows the effect of pressure on the relationship between naphtha selectivity and hydrocracking conversion. As mentioned above, an increase in pressure gives way to an increase in naphtha selectivity. Selectivity peaks under 75 bar, with its value being of around 0.68 for a conversion of 98%. Above this conversion under the conditions studied (400 °C, WHSV = 2 h−1), overcracking of naphtha and medium distillate fractions takes place by forming dry gases and LPG. In addition to increasing naphtha selectivity, pressure increase shifts this overcracking state to higher conversion values. 3.4. Effect of Pressure on Product Composition. Figure 7 shows the effect of pressure on the composition of naphtha (Figure 7a) and medium distillates (Figure 7b). The results correspond to the pseudostable state of the catalyst. Under 30 bar, aromatics account for 50 wt % of the naphtha, whereas the concentrations of the families corresponding to n-paraffins, indanes, and biphenyls are of around 10 wt %. As pressure is increased above 45 bar, there is a significant decrease in the concentration of indanes and an increase in those of i-paraffins and cycloalkanes, with n-paraffins increasing to a lesser extent. Pressures above 60 bar are required for decreasing the concentration of aromatics, and therefore, their concentration is 30 wt % under 75 bar. Benzene, whose presence is being increasingly restricted in gasoline, is decreased from 1 wt % at 30 bar, down to 0.7 wt % at 75 bar. The RON index is above 90 and increases linearly as pressure is increased. Considering the results obtained in previous papers, studying the effect of other operating conditions on naphtha composition,27−31 space velocity is a significant variable, and therefore, values of space velocity lower than those used here (WHSV = 2 h−1) should be applied to decrease the content of

Figure 5. Effect of pressure on the reaction indices at the pseudostable state of the catalyst: (a) hydrodesulphurization conversion; (b) concentration of product fractions; (c) hydrocracking conversion and naphtha and medium distillate selectivity. 5901

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The medium distillates produced under 30 bar (Figure 7b) have an aromatic concentration of 61 wt % and, as mentioned for the naphtha fraction, pressures above 60 bar are required for a significant decrease in this concentration (30 wt % under 75 bar) and, thereby, for their incorporation into the diesel pool together with other less aromatic refinery streams. Pressures above 45 bar exponentially enhance the concentration of iparaffins and n-paraffins. The concentration of the remaining components is lower, and the effect of pressure is insignificant. 3.5. Catalyst Regeneration. A combination of techniques (temperature-programmed oxidation coupled with mass spectrometry and Fourier transformed infrared spectroscopy, ultraviolet−visible spectroscopy, two-dimensional gas chromatography/mass spectrometry, and analysis of the soluble coke) has been used previously to study the relationship between coke deposition (content in the catalyst, composition, and location on the metallic and acid sites) and catalyst properties.32 The aforementioned study revealed a suitable behavior of the catalyst used here, given that it has a high activity for hydrocracking coke precursors. Consequently, once an initial time has elapsed in which a given amount of coke is deposited (conditioned by the properties of the catalyst and the reaction conditions), the deposition reaches a pseudoequilibrium and, subsequently, the catalyst maintains a considerable residual activity with time on stream for LCO hydrocracking. Nevertheless, catalyst regeneration will presumably be required when very long reactions are carried out under industrial operating conditions. It has been proven that when performing this regeneration by coke combustion at 500 °C for 3 h, the catalyst fully recovers its kinetic behavior, at least in the first 10 reaction−regeneration cycles carried out. The runs have been carried out under 75 bar with 24 h time on stream, and the results of the evolution of hydrodesulphurization and hydrocracking conversions with time on stream are similar to those shown in Figure 4 under the same pressure. The results of product distribution at the pseudostable state of the catalyst subjected to successive regenerations are also similar to those corresponding to the fresh catalyst (Figures 5−7). This proven absence of irreversible deactivation evidence the catalyst stability and is interesting for use at industrial scale. Figure 8 shows the transmission electron microscopy (TEM) images corresponding to the fresh (Figure 8a), used (Figure 8b), and regenerated catalysts (Figure 8c). The crystals of the metallic function are observed in Figure 8b and c (red arrows), and the visible coke deposited on the outside of the catalyst particles is indicated in Figure 8b (black arrows). This external coke could coexist with one deposited inside the micropores of the catalysts. However, the latter could not be visualized by TEM. The combustion of the coke in the used catalyst allows recovering crystallinity, as observed in Figure 9, where X-ray diffractograms are compared for the used and regenerated catalysts (the latter similar to the fresh catalyst, Figure 2). Likewise the catalyst recovers its physical properties and acidity by coke combustion.

Figure 6. Effect of pressure on the relationship between naphtha selectivity and hydrocracking conversion.

Figure 7. Effect of pressure on the composition of naphtha (a) and medium distillates (b).

4. CONCLUSIONS Pressures in the 35−75 bar range have a significant effect on LCO hydrocracking over a Pt−Pd catalyst supported on a HY zeolite. An increase in pressure effectively attenuates deactivation by reaching a pseudoequilibrium state of the catalyst with a high residual activity for hydrodesulphurisation and hydrocracking, whose conversions under 75 bar are 0.96

aromatics.30 The temperature (400 °C) and catalyst used in this study are suitable for operating under high hydrocracking rates and avoiding the overcracking enhanced by higher temperatures and a more acid catalyst. 5902

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Figure 8. TEM micrographs of Pt−Pd/Y catalyst: (a) fresh, (b) used, and (c) regenerated. Black arrows, coke. Red arrows, Pt−Pd metal particles.

Furthermore, the combustion of the coke deposited on the catalyst for 24 h fully restores catalyst activity. Although this result is encouraging for the use of this catalyst in industry, operation for long runs is required to confirm the absence of irreversible deactivation.

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AUTHOR INFORMATION

Corresponding Author

*Telephone: +34-946015341. Fax: +34-946013500. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

■

ACKNOWLEDGMENTS This work was carried out with the financial support of the Ministry of Science and Education of the Spanish Government (Project CTQ2006-03008/PPQ), the University of the Basque Country (UFI 11/39), and the Basque Government (Project GIC07/24-IT-220-07).

■ Figure 9. X-ray diffractograms of deactivated and regenerated Pt−Pd/ Y catalysts.

and 0.87, respectively. In this state, both conversions remain constant with time on stream, which provide evidence for the thioresistance and low deactivation by coke of the Pt−Pd/HY catalyst, which is consequently very stable under high pressure. As pressure is increased, the selectivity of the naphtha fraction increases considerably, which under 75 bar and 400 °C is 68% for a LCO conversion of 98%, with overcracking being mild under these conditions. This naphtha has an aromatic concentration of 30 wt % and is therefore suitable for incorporating into the pool of gasoline together with other less aromatic refinery streams. The results are encouraging for upgrading the LCO fraction by indirect hydrocracking, although the content of aromatics under the conditions studied is excessive for considering the naphtha obtained as a commercial gasoline, given that it exceeds the maximum concentration of aromatics allowed. In order to meet this requirement, values of space velocity lower than those in this paper are advised, based on the results of previous studies on the effect of other operating conditions on naphtha composition.

NOTATION B/L = Brönsted/Lewis site ratio (molB/molL). dm = average diameter of metal crystals (nm) DG, HCO, LCO, LPG, MD, N = dry gas, heavy cycle oil, light cycle oil, liquefied petroleum gas, medium distillates, and naphtha, respectively HDM, HDS = hydrodemetallization and hydrodesulphurization, respectively M = metal content in the catalyst (wt %) mi = mass flow rate of i lump (gi h−1) nH2 = H2/LCO molar ratio (molH2 (molHC)−1) Rp = pore radius (Å) Sa = average acid strength of the catalyst (kJ (molNH3)−1) Si = selectivity of i lump, eq 3 Si, S = mass flow rate of i lump at the reactor inlet and outlet, respectively (g h−1) Sg = BET surface area (m2 g−1) Sm = metal surface area (m2 g−1) Vp, Vmp = pore and micropore volume, respectively (cm3 g−1) X, XHDS = hydrocracking and hydrodesulphurization conversion, eqs 1 and 2, respectively WHSV = weight hourly space velocity (gLCO gcat−1 h−1) Yi = yield of i lump, eq 4

Subscripts and Superscripts

i, o = inlet and outlet 5903

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dx.doi.org/10.1021/ef3009597 | Energy Fuels 2012, 26, 5897−5904