ARTICLE pubs.acs.org/EF
Role of Acidity in the Deactivation and Steady Hydroconversion of Light Cycle Oil on Noble Metal Supported Catalysts Alazne Gutierrez,* Jose M. Arandes, Pedro Casta~no, Andres T. Aguayo, and Javier Bilbao Departamento de Ingeniería Química, Universidad del País Vasco, Apartado 644, 48080 Bilbao, Spain ABSTRACT: The deactivation of noble metal catalysts has been studied in the hydrocracking of the Light Cycle Oil (LCO) obtained as a byproduct in FCC units. The catalyst metallic functions are Pd, Pt, and PtPd, which are supported on acid materials of different porous structure and acidity (HY zeolite, Hβ zeolite, amorphous alumina, and an FCC catalyst). The reaction conditions are 350 °C; 50 bar; H2/LCO molar ratio (nH2), 8.9 molH2 (molLCO)1; space velocity (WHSV), 4 h1; time on stream, 300 min. The roles of the metallic function, porous structure of the support, and, particularly, catalyst acidity in the deactivation by coke deposition have been studied. Deactivation leads the catalyst to a pseudostable state, with significant activity remaining when a support with high acidity is used (a HY zeolite with SiO2/Al2O3 = 5) and a better performance of the PtPd metallic function.
1. INTRODUCTION The sustainability of the oil refining industry requires the development of intensification strategies for the upgrading of crude and secondary interest refinery streams to increase the yield of fuels and petrochemical raw materials and comply with ever more restrictive quality and environmental requirements. Furthermore, this strategy should be complemented by cofeeding streams derived from alternative sources to oil (biomass, coal, natural gas) or from the upgrading of postconsumer wastes (plastics, tires) to the refinery unit.110 Light cycle oil (LCO) is a stream of secondary interest in catalytic cracking units (FCC), whose production is increasing because of the rise in the throughput of FCC units around the world to meet the demand for automotive gasoline and light olefins.11 LCO has commonly been used as blend stock for home heating oil, industrial fuel oil, and diesel. Nevertheless, the increasing use of natural gas for home heating and energy production has reduced its demand. Furthermore, the use of LCO as diesel pool blend stock to meet the demand for automotive fuels is restricted by its low cetane index (around 20) and by the increase in sulfur and aromatic content in the final fuel. Based on these circumstances, there is renewed interest in upgrading LCO. Among the initiatives for improving LCO composition, studies have been carried out on the effect of the feed and catalyst acid function in the FCC unit.12,13 Ancheyta and Rodriguez studied the recirculation of LCO together with vacuum gas oil (VGO) in the FCC unit and observed an increase in the octane number of the gasoline lump, but also an undesired increase in the content of aromatics.14 Furthermore, the solutions for conditioning LCO by distilling sulfur components are costly and, moreover, additives need to be incorporated to increase the cetane number when the treated stream is used as a blending component for diesel and heating oil. The present trend in the upgrading of LCO is the implementation of hydrotreatment units, either for the production under mild conditions of a stream that is suitable for feeding in an FCC unit or for the production by severe hydrotreatment (hydrocracking) of suitable streams for gasoline and diesel pools. r 2011 American Chemical Society
A challenge in the hydrotreatment units is to strike a balance among activity, selectivity, and stability in the catalysts. The use of catalysts with noble metals as metallic function is of great interest, but their high activity and conversion to aromatics has as a downside, namely, rapid deactivation in the treatment of feeds with a high content of aromatics and heteroatomic compounds.1518 Literature studies on LCO hydrotreatment focus on catalyst discrimination, but due to the experimental difficulties involved in the transformation of real LCO feeds, specifically the rapid deactivation of the catalysts, they address the hydrotreatment of model compounds,19 LCO fractions,20 partially hydroprocessed LCO,21,22 and mixtures of LCO (or partially hydroprocessed LCO) with straight run gas oil (SRGO).23,24 Furthermore, to better understand the role of each function (metallic and acid) in the catalyst and improve the properties of these functions, certain papers study either the first step of LCO hydrocracking (LCO hydrogenation),2530 or the second one involving the cracking or ring-opening of the partially hydroprocessed LCO.21,22,31 Although deactivation is a key factor for process viability, few papers deal with this aspect. Tailleur et al. have proven that the deactivation of a WNiPd/TiO2Al2O3 catalyst is attributed to the deposition of coke on the strong acid sites of the acid function, which evidence the significance of acidity for palliating deactivation.3234 This paper deals with the hydrocracking of LCO on noble metal catalysts (Pd, Pt, and PtPd) supported on several acid materials to further knowledge of deactivation attenuation by investigating the effect on deactivation of the properties of the catalyst’s metallic and acid functions. The results of LCO conversion (hydrocracking and hydrodesulphurization), deactivation, and selectivity of products have been correlated with the properties of the catalyst by regression analysis to assess their effect (particularly that of acidity) on catalyst deactivation and performance. Received: April 6, 2011 Revised: May 27, 2011 Published: June 02, 2011 3389
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The study’s interest lies in the transformation of a real LCO, without any pretreatment by hydrodesulphurization, hydrodenitrogenation, or hydrodemetalation, which means the catalyst is subjected to very severe deactivating conditions because all the heavy heteroatomic components in the LCO reach the catalytic bed. The results will presumably not be directly applicable to industrial implementation, in which LCO pretreatment steps are required (using one or two beds of transition metal bifunctional catalysts), but they will provide limits for operating conditions when noble metal catalysts are used under these severe conditions. Furthermore, the results are a basis for subsequent studies in which the effect of operating conditions will be addressed.
2. EXPERIMENTAL SECTION 2.1. Catalysts. Fifteen catalysts have been prepared based on five different supports: Y12 (HY ultrastable zeolite, CBV712, Zeolyst International, SiO2/Al2O3 = 12), Y5 (HY ultrastable zeolite, CBV500, Zeolyst International, SiO2/Al2O3 = 5), B (Hβ zeolite, CP814E, SiO2/Al2O3 = 25), F (a commercial FCC catalyst, Albemarle), and A (amorphous γ-alumina, Stem Chemicals). The acid supports supplied in ammonium form have been calcined to obtain their acid form, according to the following steps: (i) 2 h at 400 °C (5 °C min1); (ii) 15 h at 500 °C (5 °C min1); and (iii) 2 h at 550 °C (5 °C min1). The aim of the calcination in the case of support F is its regeneration, given that this catalyst supplied by Repsol YPF (Spain) is a spent one from the FCC unit. The bifunctional catalysts have been prepared by wet impregnation of the supports at 80 °C (in the case of catalysts containing Pt and PtPd) and at 40 °C (in the case of those containing Pd) with an aqueous solution of Pt(NH3)4(NO3)2 (Alfa Aesar) and/or Pd(NH3)4(NO3)2 (Stem Chemicals). The required amount of noble metal solution is added to 100 mL of deionized water by keeping pH = 7. Once adsorption equilibrium has been reached, the excess of water is removed in a Rotavapor. The catalyst obtained is dried in an oven for 24 h at 120 °C and subsequently calcined with air at 450 °C (following a ramp of 5 °C min1) for 2 h. The metal content of the catalysts has been determined by ICP-AES (inductively coupled plasmaatomic emission spectroscopy) in a HORIBA Jobin Yvon Activa. Prior to the analysis, the sample had been subjected to an acid digestion with HF (Merck) at 90 °C. N2 adsorptiondesorption isotherms have been obtained at 196 °C in a Micromeritics ASAP 2010 once the samples have been degassed under a vacuum of 150 °C for 8 h. The surface area has been calculated according to BET equation, whereas pore volume and its distribution have been determined according to BJH methodology. The macropores of the FCC catalyst matrix used as support have been determined by Hg intrusion porosimetry in a Micromeritics Autopore II 9220. The metal surface has been determined by H2 chemisorption in a Micromeritics AutoChem II, according to the double isotherm method.35 The analyses for Pt catalysts have been carried out at 100 °C,36 and for Pd catalysts have been carried out at 70 °C.37 Crystallinity and metal particle size have been determined by XRD diffractometry in a Philips X’Pert MPD, using BraggBrentano Theta 2Theta geometry, with a ka1 radiation of 1.540598 and a ka2 of 1.544426, a PW3123/00 secondary graphite monochromator, and a PW3011 scintillation counter. The TEM images of the catalysts have been obtained by transmission electron microscopy provided with a super twin-lens, LaB6 filament, a point-to-point resolution of 0.235 nm, and line-to-line resolution of 0.144 nm. Slope (40°. Total acidity and acid strength distribution have been measured by temperature programmed desorption (TPD) of NH3 adsorbed at 150 °C in a TG-DSC Setaram 111 calorimeter provided with a Harvard
Figure 1. Distribution of LCO components grouped according to families and carbon atom numbers. Apparatus syringe, which was online with a Balzers Quadstar 422 mass spectrometer. Prior to analysis, each sample was subjected to sweeping with He at 550 °C for 0.5 h for impurity removal. It was then cooled to 150 °C and an NH3 flow of 50 μL min1 was introduced. Once the sample had been saturated, NH3 TPD was carried out and desorption was quantified by means of a Balzers Quadstar 422 spectrometer. Accordingly, a He flow of 20 mL min1 was passed at the same time as the temperature was raised to 550 °C following a heating ramp of 5 °C min1.38 Acid strength has been determined by simultaneously monitoring NH3 differential adsorption by calorimetry and thermogravimetry at 150 °C in a Setaram TG-DSC111.39 The type of acid sites has been determined by Fourier transform infrared (FTIR) spectroscopic analysis of adsorbed pyridine in a Nicolet 740 SX FTIR provided with a Specac transmission catalytic cell equipped with a temperature controller and a Vacuubrand rotary vacuum pump. The Br€onsted-Lewis (B/L) site ratio has been calculated from the vibrational bands of pyridine adsorbed at 1547 cm1 (pyridine associated with Br€onsted sites) and at 1455 cm1 (pyridine associated with Lewis sites) using the molar extinction coefficients by Emeis et al.40 Prior to the analysis, a pellet made up of approximately 1 mg sample and KBr (transparent to IR radiation) was subjected to vacuum at 400 °C for 0.5 h. The sample was then saturated with pyridine at 150 °C and the spectrum was recorded. 2.2. LCO Properties. The LCO has been supplied by Repsol YPF. The analysis has been carried out by gas chromatography/mass spectrometry (Shimazdu-QP2010-S) and the results allowed calculating the concentrations of component fractions in wt %, Figure 1. The total aromatic content (calculated as the sum of mono-, di-, tri-, and polyaromatics shown in Figure 1) is 66.83 wt %, with two-ring aromatics (methyl naphthalene as main component) being the major fraction (28.32 wt %), although the amount of three-ring aromatics is also considerable (16.44 wt %). The total sulfur content has been determined by X-ray fluorescence (XRF) (Philips MiniPal PW-4025). The elemental analysis of the LCO has been carried out in a EuroVector Euro EA Elemental Analyzer (CHNS) and has confirmed the results of sulfur content obtained by XRF. Density has been determined according to the ASTM D 4052 standard, and the cetane index according to ASTM D 4737 and D 976 standards. Simulated distillation has been carried out as per the ASTM D 3390
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Table 1. LCO Properties density, kg L1 (ASTM D 4052)
0.936
simulated distillation (ASTM D86-05) IBP-FBP, °C
101.5466.0
T50T95, °C
277.3392.7
cetane number ASTM D 4737
27.64
ASTM D 976
27.31
chemical analysis (wt%) C
88.7
H
10.2
S
0.5
N
0.2
2887 standard. Table 1 and Figure 2 show the results of these analyses. As observed, LCO is a stream with a significant sulfur content (0.5 wt %) and low cetane index due to its high aromatic content. 2.3. Reaction Equipment and Conditions. The automated reaction equipment used has been described elsewhere41,42 and is equipped with a down-flow fixed bed reactor of 15 cm3 (8 mm internal diameter and 303 mm length), with the reaction conditions being 350 °C; 50 bar; H2/LCO molar ratio (nH2), 8.9 molH2 (molLCO)1; space velocity (WHSV), 4 h1; and time on stream (TOS), 300 min. These operating conditions have been chosen for catalyst discrimination, given that they correspond to significant deactivation. The catalyst (0.150.30 mm pellets) is mixed (1/1 mass ratio) with CSi (0.5 mm). 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 cm3 min1 of H2, 50 cm3 min1 of N2), raising the temperature following a ramp of 5 °C min1 from ambient temperature to 350 °C and keeping 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 microGCGC 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 C5C10 hydrocarbons. The liquids are analyzed in a Hewlett-Packard 6890 gas chromatograph provided with a FID detector and a PONA capillary column (50 m 2 mm 0.5 mm). 2.4. Coke Analysis. The content and characteristics of the coke deposited on the catalysts have been determined by temperature programmed oxidation (TPO) in a Setaram SDT 2960 thermobalance online with a Balzers Instruments Thermostar mass spectrometer for the analysis of the gaseous stream produced in the combustion.43 The sample (30 mg) was heated to the initial combustion temperature 150 °C in an inert atmosphere and combustion followed a ramp of 3 °C min1 to 650 °C.
3. RESULTS 3.1. Catalyst Properties. As shown in Table 2, Pt content (M) in the catalysts is between 0.45 and 1.52 wt %, whereas that of Pd is between 0.37 and 1.25 wt %. Pore volume (Vp) is higher for the catalysts with support B (ca. 0.74 cm3 gcat1), compared with the catalysts supported by Y and A (ca. 0.43 cm3 gcat1), and F (ca. 0.13 cm3 gcat1), respectively. In the case of the catalysts with support Y, 60% of the volume corresponds to the micropore volume (Vmp), whereas this volume accounts for 21% in the catalysts with supports B and F, and only 14% in the catalysts with support A (amorphous structure). The results evidence a
Figure 2. TBP curve of LCO simulated distillation.
better Pt deposition on the supports with higher micropore volume (supports Y and B), whereas Pd is better deposited as the structure becomes less microporous (supports F and A). The catalysts with supports Y and B have the higher values of BET surface area (Sg), > 500 m2 gcat1; those with support A have a surface area of approximately 300 m2 gcat1, and the lower values of BET surface area correspond to the catalysts with support F, ca. 100 m2 gcat1. The values of metal surface area (Sm) are very different depending on the total amount of metal deposited, and it should be noted that in the case of the catalysts with support Y metal dispersion is better in the bimetallic catalysts than in the monometallic ones. The values of average metal particle size (dm in Table 2) are very different and depend on the support. The catalysts with supports Y and B have very low values of metal dispersion (Table 2), which is consistent with the high values of the metal crystallite size observed in the TEM images (Figure 3). Furthermore, the catalysts with supports F and A (with lower micropore volume and lower crystallinity) have high values of metal dispersion. Two factors contribute to this disparity in the results:44 (i) Pt has a higher atomic size than Pd and a lower capacity for entering the porous framework, where it is deposited as 314 nm size crystallites when metal contents are of around 1 wt %; (ii) Pd crystallites are liable to migrate and undergo sintering during the catalyst calcination step, causing large agglomerates. The TEM analysis (Figure 3) shows that the results for metal crystal size (dm) are consistent with those above obtained by H2 chemisorption. The crystal planes are better distinguished for Pt (Figure 3a), whereas Pd and PtPd form particles of irregular size (Figure 3b, c, d). In the case of Y12 zeolite catalysts, the crystallite size ranges are Pt, 520 nm (Figure 3a); Pd, 520 nm (Figure 3b); PtPd, 220 nm (Figure 3c, d). A good PtPd dispersion is noteworthy and, furthermore, Pt is better dispersed on Y12 zeolite than on Y5 zeolite. Crystal size is lower in the catalysts with support B, especially for Pt. The size ranges are Pt, 15 nm; Pd, 210 nm; PtPd, 210 nm. It should be noted that this last size is similar to that corresponding to supports Y. 3391
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Table 2. Textural and Metallic Properties of Fresh Catalystsa M (gmetal/100 gcat) catalyst
support
SiO2/Al2O3
Pt
Pt/Y12
HY zeolite
12
Pd/Y12
HY zeolite
12
PtPd/Y12
HY zeolite
12
Pt/Y5
HY zeolite
5.2
Pd/Y5
HY zeolite
5.2
PtPd/Y5
HY zeolite
5.2
Pt/B
Hβ zeolite
25
Pd/B PtPd/B
Hβ zeolite Hβ zeolite
25 25
0.67
Pt/F
FCC catalyst
4
1.02
Pd/F
FCC catalyst
4
PtPd/F
FCC catalyst
4
Pt/A
γ-Al2O3
Pd/A
γ-Al2O3
PtPd/A
γ-Al2O3
Pd
Vp (cm3 gcat1)
Vmp (cm3 gcat1)
Sg (m2 gcat1)
Sm (m2 gcat1)
dm (nm)
1.21 0.63
0.50
0.28
743
0.09
13
1.25
0.48
0.29
749
0.02
11
0.58
0.34
0.18
511
0.26
5
0.40
0.26
667
0.20
18
1.52 1.12
0.93
0.42
0.28
719
0.40
10
0.46
0.40
0.25
668
0.85
5
0.71
0.16
547
0.56
4
0.71 0.80
0.18 0.15
607 562
0.50 0.09
7 6
0.13
0.03
99
0.51
5
0.13
0.03
106
0.17
17
1.02 0.87 0.57 1.22 0.45
0.42
0.14
0.02
90
0.33
8
0.46
0.07
311
0.21
4
1.05
0.39
0.06
294
0.54
8
0.37
0.48
0.06
285
0.37
6
0.99 0.46
M = metal content in the catalyst, wt%; Vp = pore volume, cm3 g1; Vmp = micropore volume, cm3 g1; Sg = BET surface area, m2 g1; Sm = metal surface area, m2 g1; dm = mean diameter of metal crystals, nm. a
The TEM images for the catalysts with support F show (particularly in the sample corresponding to PtPd/F) the HY zeolite particles embedded in the matrix that makes up the catalyst. It is also observed that the metal crystals are preferably deposited on the zeolite crystals and, to a lesser extent, on the inert matrix of the support. The size ranges of the metal crystals are between 5 and 20 nm for Pt, Pd, and PtPd. In the case of catalysts with support A, the size of the crystals observed is small for Pt, between 2 and 5 nm, and slightly higher for Pd, between 2 and 10 nm, and between 4 and 10 nm for PtPd. The results of NH3 TPD (Table 3) allow establishing the following order for total catalyst acidity, which corresponds to the order of support acidity: Y5 > Y12 ≈ B > A > F. Table 3 also shows the values of weak (Rw), medium (Rm), and strong (Rs) acidity, which have been determined as the amounts of NH3 desorbed in three temperature ranges: 150280, 280420, and 420550 °C, respectively. It is observed that the catalysts with supports Y12 and B have identical values of total acidity (R), although the values of average acid strength (Sa) and Br€onsted/ Lewis site ratio (B/L) are higher for the catalysts with support Y12 than those with support B. The catalysts based on Y5 zeolite have a high total acidity and a significant acid strength with a higher average than the other catalysts due to the low Si/Al ratio of the support. On the other hand, the catalysts with support A have a lower total acidity, acid strength, and Br€onsted/Lewis ratio than the catalysts based on Y and B zeolites. The catalysts with support F are those with the lower acidity indices, which is due to their lower zeolite content, given that the support is an FCC catalyst with a USY zeolite content of approximately 15 wt %. 3.2. Catalyst Deactivation. The LCO components and reaction products have been grouped into the following fractions according to the boiling point range: dry gas (C1C2); liquefied petroleum gas (LPG) (C3C4); naphtha (C5C12) (36216 °C); medium distillates (MD) (C13C20) (216343 °C); heavy cycle oil (HCO) (C21+) (343 °C+)
Given that the main objective is the conversion of the heavy fraction (HCO) in the LCO, the hydrocracking conversion has been defined as X HC ¼
HCOi HCOo HCOi
ð1Þ
where HCOi and HCOo are the molar flow rates of the HCO fraction in the reactor inlet and outlet streams. The hydrodesulphurization conversion is defined as X HDS ¼
Si So Si
ð2Þ
where Si and So are the mass flow rates of the sulfur contained in the feed and product streams. Figure 4 shows the evolution of LCO hydrocracking conversion with time on stream for the different catalysts (Figure 4a, catalysts with supports A and F; Figure 4b, catalysts with supports Y and B). The hydrocracking conversion corresponding to the fresh catalyst (zero time on stream) is almost complete for all the catalysts studied. Nevertheless, due to catalyst deactivation, conversion decreases with time on stream for all the catalysts to a value that remains constant subsequent to 5 h time on stream. In this situation, the catalysts reach a pseudostable state with constant activity. Deactivation also affects hydrodesulphurization conversion, as observed in Figure 5 (corresponding to the runs shown in Figure 4). Catalyst deactivation is caused by several factors:45 (i) adsorption of sulfur and metals on the metallic function, with the former presumably being more significant; (ii) adsorption of nitrogen compounds on both the metallic function and the acid function; and (iii) deposition of coke, presumably on both functions. Figures 4 and 5 show that the types of metal and support influence deactivation, especially the support. Deactivation is faster for the catalysts with supports F and A (Figure 4a) and, 3392
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Figure 3. TEM micrographs of the catalysts (a) Pt/Y5, (b) Pd/Y5, (c) PtPd/Y5, (d) PtPd/B, (e) PtPd/F.
between these, it is faster for support F, which is also the one of the lowest hydrocarbon conversion when the catalyst is at pseudostable state. Furthermore, the hydrocarbon conversion order for each support when the catalysts are at pseudostable state is as follows: PtPd > Pt > Pd. For the catalysts with zeolites as support (Figure 4b), deactivation follows an order according to the supports: B > Y12 > Y5, and consequently, hydrocarbon conversion at the catalyst pseudostable state follows the reverse order. Concerning the metallic function, hydrocarbon conversion order at this pseudostable state is PtPd > Pt > Pd. Consequently, the trends for the effect of acidity and metallic function are qualitatively similar for both the amorphous (A and F) and the crystalline zeolite supports (Y and B). Figure 5 shows that hydrodesulphurization conversion at pseudostable state for the Pd catalyst is much lower than that corresponding to Pt and PtPd catalysts with the different supports. The catalysts with support A have a higher activity than their counterparts with support F, both at zero time on stream and at the pseudostable state (Figure 5a). Concerning hydrodesulphurization on the catalysts with zeolite supports (Figure 5b), a lower activity is observed at zero time on stream for the catalysts
supported on Hβ zeolite, which also undergo faster deactivation than those supported on HY zeolite. The latter gives way to complete conversion at zero time on stream and activity is lower for those of Pd when pseudostable state is reached. Based on the aforementioned causes of deactivation, the pseudostable state of the catalysts is reached when the following reversible and parallel deactivation reactions are in equilibrium: (i) Metallic function (m1) sulphuration: m1 ðsÞ + H2 SðgÞ T m1 SðsÞ + H2 ðgÞ
ð3Þ
(ii) Coke deposition: LCO f products T coke
ð4Þ
The equilibrium constant for poisoning the metal m1 is as follows: ðm1 SÞPH2 ð5Þ K p ðTÞ ¼ ðm1 ÞPH2 S The poisoning level corresponding to the pseudostable state decreases as temperature (S adsorption is exothermic) and 3393
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Table 3. Acidity of the Fresh Catalystsa (μmolNH3 gcatalyst1) catalyst
R
Rw
Rm
Rs
Sa kJ (molNH3)1
B/L
Pt/Y12
498
154
326
18
99
0.90
Pd/Y12
600
182
305
113
96
0.95
PtPd/Y12
550
90
385
75
98
0.70
Pt/Y5
678
255
259
164
110
0.56
Pd/Y5
803
251
334
218
108
0.48
PtPd/Y5
686
311
252
123
109
0.68
Pt/B
583
240
332
11
87
0.64
Pd/B PtPd/B
612 598
182 84
348 448
82 66
85 87
0.82 0.49
Pt/F
61
57
4
65
0.65
Pd/F
68
32
32
4
65
0.40
PtPd/F
65
35
27
3
65
0.40
Pt/A
295
78
203
14
80
0.80
Pd/A
357
108
212
37
81
0.85
PtPd/A
347
100
195
52
80
0.48
Sa = acid strength of the catalyst, kJ (molNH3)1; B/L = Br€onsted/ Lewis site ratio, molB/molL; R, Rw, Rm, Rs = total, weak, medium, and strong acidity of the catalysts, respectively, μmol NH3 gcatalyst1. a
PH2/PH2S partial pressure ratio are increased (the latter due to the increase in the H2/LCO molar flow ratio in the feed). It is well-reported that Pt has a higher affinity to sulfur (higher Kp) than Pd, and that PtPd alloy has a lower Kp than Pt, which decreases as the Pd/Pt ratio is higher, with a minimum value that depends on the support and reaction conditions.31,46 Yasuda and Yoshimura have determined that the optimum Pd/Pt ratio is 4 for attenuating the poisoning by sulfur in the hydrocracking of tetraline.47 Coke deposition (eq 4) affects mainly the acid function, in which the coke is formed by condensation reactions activated by the hydrogen transfer capacity of this function.48,49 The formation of coke in a catalyst with hydrocracking capacity is a reversible reaction, given that the coke precursors and intermediates are hydrocracked and the products are displaced toward the outside of the catalyst particles. The fact that coke formationhydrocracking is reversible is essential for explaining the pseudostable state reached by the catalysts. Furthermore, the higher coke content of Pd catalysts compared to Pt is a well-reported finding in the literature on hydrotreatment reactions,44 which is explained by the lower capacity of Pd for the hydrocracking of coke precursors. Furthermore, the synergism of PtPd alloy for hydrogenation and hydrocracking (well-proven for LCO) also occurs for coke precursors and, consequently, the bimetallic catalyst will be more efficient for attenuating coke condensation. According to these circumstances, the same factors (catalyst composition and operating conditions) affect poisoning (eq 3) and coke deposition equilibria (eq 4), and the catalyst reaches a pseudostable state with higher residual activity with (i) the bimetallic function PtPd, (ii) a higher temperature, and (iii) a higher H2 partial pressure. 3.3. Product Selectivity. Figure 6 shows the results of selectivity evolution with HCO conversion for the medium distillate lumps (Figure 6a) and naphtha (Figure 6b) for the PtPd catalysts on different supports. Each curve gathers the experimental points corresponding to different deactivation
Figure 4. Evolution of hydrocarbon conversion with time on stream for the different catalysts: (a) with supports A and F and (b) with zeolitic supports Y and B (350 °C; 50 bar; nH2, 8.9 molH2/molLCO; WHSV, 4 h1).
states of the catalyst, given that conversion decreases with time on stream. The selectivity of each i lump is calculated from the yields of product lumps, as follows: Si ¼
Yi Y dry gas + Y LPG + Y naphtha + Y MD
ð6Þ
where each yield is Yi ¼
mass flow rate of i lump in the product stream LCO mass flow rate in the feed
ð7Þ
The results in Figure 6 show the capacity the catalysts have to produce product streams of commercial interest with a high selectivity of medium distillates, even when the catalyst is significantly deactivated. Deactivation should be palliated for producing naphthas, although operation under high conversion conditions is not convenient, as the selectivity of naphthas decreases under these conditions due to their cracking to LPG and dry gas lumps. This circumstance is enhanced by the support 3394
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Figure 5. Evolution of hydrodesulphurization conversion with time on stream for different catalysts: (a) with supports F and A and (b) with zeolitic supports Y and B (350 °C; 50 bar; nH2, 8.9 molH2/molLCO; WHSV, 4 h1).
acidity, although the results of fast cracking obtained with the support of Hβ zeolite evidence that the cracking is also enhanced by the accessibility to the porous structure. 3.4. CokeAcidity Relationship. As observed in Figure 7, the coke content deposited has a linear relationship with the acidity of the fresh catalysts (Table 3). The effect of the metallic function is also clear, given that the straight lines corresponding to each metallic function are clearly differentiated. The results corresponding to each support are grouped with a continuous line. The dependency of coke content with acidity evidences the catalytic origin of the coke, which is formed within the channels of the support’s porous structure following the well-reported mechanisms of condensation toward polyaromatic structures that are characteristic of acid catalysts.48,49 The coke content corresponds to a limit value characteristic of pseudoequilibrium formation and disappearance (by hydrocracking) in each catalyst, depending mainly on acidity and, to a lesser extent, on the porous structure. Consequently, as fresh catalyst acidity is higher, both its capacity for coke formation and the residual activity are also higher.
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Figure 6. Selectivity evolution with HCO conversion for (a) medium distillate, and (b) naphtha, for the PtPd catalyst on different suports (350 °C; 50 bar; nH2, 8.9 molH2/molLCO; WHSV, 4 h1).
In addition to acidity, the porous structure of the support also influences coke deposition, which explains the higher coke content in the catalysts with support Y12 than those with support B (of similar acidity). This difference is explained by the easier condensation of coke precursors in the cages characteristic of HY zeolite channel intersections HY.43,49 3.5. Deterioration of Catalyst Properties. Table 4 shows that, due to deactivation, the BET surface area and the pore and micropore volume of the catalysts have undergone a serious deterioration compared to the properties corresponding to the fresh catalysts (Table 2). This deterioration is more significant in the catalysts with zeolitic supports, in which the BET surface area decreases by more than 90%, which means a preferential blockage of micropores by coke deposition. An analysis by H2 chemisorption is not possible because the results are masked by coke. Furthermore, and due to the coke of catalytic origin, the catalysts with supports F or A and those with Pd metallic phase (less active) are those that undergo a less severe deterioration of their textural properties. 3395
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Table 4. Physical Properties of the Spent Catalysts (350 °C; 50 bar; nH2, 8.9 molH2 (molLCO)1; WHSV, 4 h1; time on stream, 300 min)a catalyst
Sg (m2 gcat1)
ΔSg (%)
Vp (cm3 gcat1)
ΔVp (%)
Pt/Y12
6.1
99
0.031
94
Pd/Y12
34.1
95
0.078
84
PtPd/Y12
31.7
94
0.095
72
Pt/Y5
18.7
97
0.055
86
Pd/Y5 PtPd/Y5
22.8 18.0
97 97
0.050 0.041
88 90
Pt/B
17.2
97
0.061
91
Pd/B
114.7
81
0.389
45
PtPd/B
70.9
87
0.228
72
Pt/F
71.9
28
0.090
30
Pd/F
23.7
78
0.054
58
PtPd/F
24.4
73
0.095
34
Pt/A Pd/A
150.4 146.2
52 50
0.274 0.276
41 30
PtPd/A
116.0
59
0.229
52
Figure 7. Effect of fresh catalyst acidity on the content of the coke deposited (350 °C; 50 bar; nH2, 8.9 molH2/molLCO; WHSV, 4 h1; time on stream, 300 min).
Sg = BET surface area, m2 g1; ΔSg = decrease in the BET surface area, %; Vp = pore volume, cm3 g1; ΔVp = decrease in the pore volume, %.
A comparison of the acid properties of the spent catalysts (Table 5) with those of the fresh catalysts (Table 3) reveals that, in general, there is a major deterioration in acidity. Average acid strength decreases due to coke deposition, although it should be noted that the catalysts with support Y5 maintain a significant value for this property at the pseudostable state. The decrease in acidity for each metallic function occurs according to the following order: for Pt, B > A > F > Y12 > Y5; for Pd, B > Y12 > A > F > Y5; for PtPd: B > A > Y12 > F > Y5. The decrease in acidity for the catalysts with support Y5 follows the order Pd > Pt > PtPd. Overall, the catalyst PtPd/Y5 is the most stable. Figure 8 compares the distribution of acid strength for the fresh catalysts (Figure 8a) with those for the used ones (Figure 8b). As observed, strong acid sites are mostly affected by deactivation, whereas, less affect was shown with medium acid strength. This result is characteristic to the formation of coke on acid catalysts, which occurs according to a mechanism in which the strong acid sites are more active.48,49 Consequently, a significant fraction of weak acid sites remains in the pseudostable state, part of which are probably strong and moderate acid sites in the fresh catalyst that have been blocked by coke, which explains that, in general, the number of weak acid sites increases throughout deactivation. 3.6. CokeResidual Acidity Relationship. Figure 9 shows the relationship between hydrocarbon conversion when the catalysts are at pseudostable state (subsequent to 5 h time on stream) and the acidity remaining in the catalysts. As observed, there is a linear relationship for each metallic phase, which evidences the conversion order PtPd > Pt > Pd. As observed in Figure 9, conversion increases as the remaining total acidity of the catalyst is higher and, consequently, it is the property that conditions its hydrocracking activity. Furthermore, the results corresponding to each support are within zones surrounded by a continuous line. The results evidence the relevance of acidity and the metallic phase on the kinetic performance in the pseudostable state of the catalysts. Thus, for those catalysts with supports B and Y12, Pd
Table 5. Acidity of the Spent Catalysts (350 °C; 50 bar; nH2, 8.9 molH2 (molLCO)1; WHSV, 4 h1; time on stream, 300 min)a
a
(μmolNH3 gcat1) catalyst
R
Rw
Rm Rs
Sa (kJ/molNH3)
ΔR (%)
B/L
Pt/Y12
386 242 143
1
90
22
0.56
Pd/Y12
279 170 107
2
80
54
0.46
PtPd/Y12 357 204 151
2
90
35
0.5
Pt/Y5
551 294 240
17
100
19
0.4
Pd/Y5
557 317 223
17
80
31
0.35
PtPd/Y5 605 336 249
20
90
12
0.32
75
40
1 1
70 90
69 53
0.61
50
31
0.25
1
50
35
Pt/B
351 234 117
Pd/B PtPd/B
192 110 279 180
81 98
Pt/F
42
40
2
Pd/F
44
25
18
47
33
14
50
28
Pt/A
PtPd/F
190 116
72
2
70
36
Pd/A
167 107
59
1
70
53
0.45
PtPd/A
173 138
35
70
50
0.33
a Sa = acid strength of the catalyst, kJ (molNH3)1; ΔR = decrease in the acidity, %; B/L = Br€onsted/Lewis site ratio, molB/molL; R, Rw, Rm, Rs = total, weak, medium, and strong acidity of the catalysts, respectively, μmol NH3 gcatalyst1.
catalysts have the highest residual acidity, whereas conversion is higher for PtPd catalysts. Moreover, there is also a linear relationship between conversion in the pseudostable state and the coke content deposited on the catalysts (Figure 10). It should be noted that the catalysts with support Y have a better performance, given that they maintain higher residual activity in the pseudostable state. This result evidences that acidity has a significant role in hydrocracking and coke formation, according to the reported mechanisms of condensation of aromatics (from LCO heavy components) and 3396
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Figure 9. Relationship between hydrocracking conversion and acidity (residual acidity in the pseudostable state) for the catalysts in pseudostable state (350 °C; 50 bar; nH2, 8.9 molH2/molLCO; WHSV, 4 h1; time on stream, 300 min).
Figure 8. Distribution of acidity in the (a) fresh and (b) spent catalysts (350 °C; 50 bar; nH2, 8.9 molH2/molLCO; WHSV, 4 h1; time on stream, 300 min).
oligomerizationaromatizationcondensation of light hydrocracking products.48,49 3.7. Significance of Catalyst Properties by Regression Analysis. A regression analysis has been carried out for the reaction indices (dependent variables) when the catalysts are in pseudostable state concerning their properties (independent variables) to determine whether there are significant relationships between these properties and kinetic performance. The independent variables are the properties of the 15 catalysts studied, i.e., fresh and deactivated (in pseudostable state): BET surface area, pore volume, micropore volume, total acidity, weak acidity, average acidity, strong acidity, medium acidity, Br€onsted/Lewis ratio, metal surface area, metal particle size, and metal content. The dependent variables studied correspond to the pseudostable state of the catalyst and are coke content, hydrocracking conversion, hydrodesulphurization conversion, naphtha selectivity, and medium distillate selectivity. Multilinear regression has been performed following the stepwise procedure (stepwisefit function in Matlab), which is a systematic method for adding and removing terms from a multilinear model based on their statistical significance in the regression. The method proceeds as follows: (1) fitting of the initial model; (2) if any term not in the model has a p-value below the entrance tolerance, add the one with the smallest p value and repeat this step; otherwise, go to step 3; (3) if any term in the model has a p-value higher than the exit tolerance, remove the one with the highest p value and go to step 2; otherwise, end.50 Table 6 shows the p-values calculated for F-statistic, which are the probabilities that Fischer-F values are greater than the calculated F-statistic, i.e., the probabilities for holding the null
Figure 10. Effect of coke content in the hydrocracking conversion (350 °C; 50 bar; nH2, 8.9 molH2/molLCO; WHSV, 4 h1; time on stream, 300 min).
hypothesis. Consequently, a small p-value means that there is a relationship between these dependent and independent variables. Specifically, p-values lower than 0.05 mean the coefficient corresponding to this independent variable is significant in the regression with a 95% confidence interval. Based on the results shown in Table 6, in which the significant relationships are in bold, the relevance of catalyst properties may be assessed. The acidity of the fresh catalyst is the more influential property, given that it conditions: (i) the hydrocracking conversion; (ii) the 3397
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Table 6. p-Values Calculated in the Hypotheses for Testing the Significance of Individual Coefficients in the Multi-Linear Regressions for the Fresh Catalyst (Upper Half) and Spent Catalyst (Lower Half)a Cc
XHC
XHDS
SNaphta
SDM
(Sg)fresh
0.9197
0.0053
0.2679
0.3154
0.2262
(Vp)fresh
0.3332
0.4403
0.1292
0.7678
0.7203
(Vmp)fresh
0.8569
0.6258
0.4335
0.2905
0.1867
Rfresh
0.0001
0.4399
0.0036
0.0032
0.0050
(Rw)fresh
0.3708
0.3788
0.5759
0.6499
0.7828
(Rm)fresh
0.7890
0.8647
0.3475
0.7927
0.9631
(Rs)fresh (Sa)fresh
0.5150 0.1681
0.3634 0.3778
0.4090 0.0000
0.9273 0.9787
0.8130 0.7921
(B/L)fresh
0.0725
0.3252
0.6677
0.4298
0.3640
(Sm)fresh
0.5741
0.2730
0.2555
0.4246
0.5826
(dm)fresh
0.9149
0.1435
0.5749
0.3976
0.4001
(M)fresh
0.9301
0.1392
0.4570
0.5781
0.4786
Cc
XHC
XHDS
Snaphta
SDM
(Sg)spent
0.6375
0.7170
0.2021
0.9372
0.8197
(Vp)spent (Vmp)spent
0.2052 0.3974
0.7945 0.8068
0.3651 0.0522
0.5068 0.0071
0.4659 0.0008
Rspent
0.3944
0.0000
0.4223
0.8960
0.8918
(Rw)spent
0.1944
0.9405
0.8607
0.5887
0.7515
(Rm)spent
0.6762
0.8757
0.2113
0.7807
0.5551
(Rs)spent
0.8743
0.8339
0.1057
0.7139
0.6370
(Sa)spent
0.0019
0.5963
0.9721
0.8513
0.5915
(B/L)spent
0.8718
0.3969
0.7574
0.5704
0.7294
Sg = BET surface area, m2 g1; Vp = pore volume, cm3 g1; Vmp = micropore volume, cm3 g1; R, Rw, Rm, Rs = total, weak, medium, and strong acidity of the catalysts, respectively, μmol NH3 gcatalyst1; Sa = acid strength of the catalyst, kJ (molNH3)1; B/L = Br€onsted/Lewis site ratio, molB/molL. a
The conversion with fresh catalysts, the level of deactivation by coke, and the residual activity of the catalyst in the pseudostable state depend not only on the porous structure, but mainly on the acidity of the support according to a linear relationship. Furthermore, the residual activity of the catalyst in the pseudostable state also depends linearly on the acidity maintained by the catalyst. The high acid strength of the support is essential for the viability of these catalysts. A Y zeolite with low SiO2/Al2O3 ratio is suitable as a support, which although it enhances coke formation allows maintaining high residual acidity, and therefore significant cracking capacity. Furthermore, the results of coke deposition evidence its catalytic origin and formation by following the wellreported mechanisms for catalytic cracking processes. Moreover, the bimetallic function PtPd maintains higher activity in the pseudostable state of the catalyst than the individual metals, which is due to a more suitable balance between activity and deactivation, and particularly to a more favorable sulphuration equilibrium. The aforementioned results correspond to given operating conditions that are suitable for catalyst discrimination. Nevertheless, an assessment of catalyst viability would require a more profound study than that presented here and be based on a wider range of operating conditions (temperature, pressure, H2/LCO molar ratio, space time) and longer times on stream, with the aim of reducing catalyst deactivation and optimizing the distribution of products for their use as medium distillates and gasoline. Furthermore, it should be taken into account that the results of LCO hydrocracking depend on its composition and, consequently, FCC operating conditions. Nevertheless, the results in this paper are a basis for assessing the effect of catalyst properties on the hydrocracking of LCO and other refinery streams containing heavy heteroatomic components.
’ AUTHOR INFORMATION Corresponding Author
amount of coke deposited; (iii) the hydrodesulphurization conversion, and (iv) the selectivity of naphtha and medium distillates. Furthermore, medium acid strength and the Br€onsted/Lewis ratio are key factors that condition the amount of coke deposited. BET surface area conditions hydrocracking conversion, although shape selectivity conditions the selectivity of products. Likewise, micropore volume is responsible for the selectivity of the naphtha fraction. The catalyst metal properties are significant for the hydrodesulphurization conversion, evidencing that hydrogenolysis occurs mainly on the metal surface, although acid sites are also significant, as indicated by the p-values for the acidity and acid strength of the fresh catalyst.
’ CONCLUSIONS The hydrocracking of LCO (without previous treatment) on supported noble metal catalysts undergoes rapid deactivation in a time on stream of approximately 5 h, with metal sulphuration and coke deposition on the support acid sites taking place simultaneously. This deactivation has a greater effect on the capacity for hydrocracking than for hydrodesulphurization, and it should be noted that the catalyst records a pseudostable performance with significant activity remaining, which is encouraging for the use of these catalysts in the upgrading of LCO to produce medium distillates and naphtha.
*Phone: +34 946 015 341. Fax: +34 946 013 500. E-mail: alazne.
[email protected].
’ ACKNOWLEDGMENT This work was carried out with the financial support of the Ministry of Science and Education of the Spanish Government (Project CTQ2006-03008/PPQ) and of the Basque Government (Project GIC07/24-IT-220-07). ’ NOTATION B/L = Br€onsted/Lewis site ratio, molB/molL Cc = coke content in the catalyst, wt % dm = mean diameter of metal crystals, nm HCO = mass flow rate of Heavy Cycle Oil, gHCO h1 Kp = sulfur adsorption equilibrium constant, eq 5 M = metal content in the catalyst, wt% m1 = metallic function in the catalyst mi = mass flow rate of i lump, g h1 nH2 = H2/LCO molar ratio, molH2 (molHC)1 PH2, PH2S = partial pressure of hydrogen and hydrogen sulfur, bar Sa = acid strength of the catalyst, kJ (molNH3)1 Si, So = mass flow rate of sulfur at the reactor inlet and outlet, g h1 Si = selectivity of i lump, eq 6 3398
dx.doi.org/10.1021/ef200523g |Energy Fuels 2011, 25, 3389–3399
Energy & Fuels Sg = BET surface area, m2 g1 Sm = metal surface area, m2 g1 Vp, Vmp = pore and micropore volume, cm3 g1 XHC = hydrocracking conversion, eq 1 XHDS = hydrodesulphurization conversion, eq 2 Yi = yield of i lump, eq 7
’ SUPERSCRIPTS i,o = inlet and outlet ’ GREEK LETTERS R, Rw, Rm, Rs = total, weak, medium, and strong acidity of the catalysts, respectively, μmol NH3 gcatalyst1 ΔSg, ΔVp, ΔR = decrease in the BET surface area, pore volume, and acidity, respectively, % ’ REFERENCES (1) Huber, G. W.; Corma, A. Angew. Chem, Int. Ed. 2007, 46, 7184–7201. (2) Corma, A.; Huber, G. W.; Sauvanaud, L.; O’Connor, P. J. Catal. 2007, 247, 307–327. (3) Rana, M. S.; Samano, V.; Ancheyta, J.; Diaz, J. A. I. Fuel 2007, 86, 1216–1231. (4) Arandes, J. M.; Torre, I.; Casta~no, P.; Olazar, M.; Bilbao, J. Energy Fuels 2007, 21, 561–569. (5) Corma, A.; Huber, G. W.; Sauvanaud, L.; O’Connor, P. J. Catal. 2008, 257, 163–171. (6) Sims, R.; Taylor, M.; Saddler, J.; Mabee, W. From 1st- to 2nd Generation Biofuel Technologies. An Overview of Current Industry and RD&D Activities; OECD/IEA: Paris Cedex, 2008 (www.iea.org). (7) Shimada, H.; Sato, K.; Honna, K.; Enomoto, T.; Ohshio, N. Catal. Today 2009, 141, 43–51. (8) Ahmad, M. I.; Zhang, N.; Jobson, M. J. Cleaner Prod. 2010, 18, 889–899. (9) Castelo Branco, D. A.; Gomes, G. L.; Szklo, A. S. Energy Policy 2010, 38, 3098–3105. (10) Gayubo, A. G.; Valle, B.; Aguayo, A. T.; Olazar, M.; Bilbao, J. Ind. Eng. Chem. Res. 2010, 49, 123–131. (11) Ren, T.; Patel, M.; Blok, K. Energy 2006, 31, 425–451. (12) Corma, A.; Sauvanaud, L. In Fluid Catalytic Cracking VII: Materials, Methods and Process Innovations; Occelli, M.L., Ed.; Elsevier B.V.: Amsterdam, 2007; Ch. 4. (13) Corma, A.; Martínez, C.; Sauvanaud, L. Catal. Today 2007, 127, 3–16. (14) Ancheyta, J.; Rodríguez, S. Energy Fuels 2002, 16, 718–723. (15) Babich, I. V.; Moulijn, J. A. Fuel 2003, 82, 607–631. (16) Speight, J. G. Catal. Today 2004, 98, 55–60. (17) Ancheyta, J.; Rana, M. S.; Furimsky, E. Catal. Today 2005, 109, 3–15. (18) Leyva, C.; Rana, M. S.; Trejo, F.; Ancheyta, J. Ind. Eng. Chem. Res. 2007, 46, 7448–7466. (19) Albertazzi, S.; Baraldini, I.; Busca, G.; Finocchio, E.; Lenarda, M.; Storaro, L.; Talon, A.; Vaccari, A. Appl. Clay Sci. 2005, 29, 224–234. (20) Laredo, G. C.; Saint-Martin, R.; Martinez, M. C.; Castillo, J.; Cano, J. L. Fuel 2004, 83, 1381–1389. (21) Nylen, U.; Sassu, L.; Melis, S.; J€aras, S.; Boutonnet, M. Appl. Catal., A 2006, 299, 1–13. (22) Nylen, U.; Pawelec, B.; Boutonnet, M.; Fierro, J. L. G. Appl. Catal., A 2006, 299, 14–29. (23) Ancheyta-Juarez, J.; Aguilar-Rodríguez, E.; Salazar-Sotelo, D.; Betancourt-Rivera, G.; Leiva-Nuncio, M. Appl. Catal., A 1999, 180, 195–205. (24) Fujikawa, T.; Idei, K.; Ohki, K.; Mizuguchi, H.; Usui, K. Appl. Catal., A 2001, 205, 71–77.
ARTICLE
(25) Santikunaporn, M.; Herrera, J. E.; Jongpatiwut, S.; Resasco, D. E.; Alvarez, W. E.; Sughrue, E. L. J. Catal. 2004, 228, 100–113. (26) Ding, L.; Zheng, Y.; Zhang, Z.; Ring, Z.; Chen, J. Catal. Today 2007, 125, 229–238. (27) Ding, L.; Zheng, Y.; Zhang, Z.; Ring, Z.; Cheng, J. Appl. Catal., A 2007, 319, 25–37. (28) Ding, L.; Zheng, Y.; Yang, H.; Parviz, R. Appl. Catal., A 2009, 353, 17–23. (29) Huang, L.; Huang, Q.; Xiao, H.; Eic, M. Microporous Mesoporous Mater. 2008, 111, 404–410. (30) Huang, L.; Huang, Q.; Xiao, H.; Eic, M. Microporous Mesoporous Mater. 2008, 114, 121–128. (31) Yoshimura, Y.; Toba, M.; Matsui, T.; Harad, M.; Ichihashi, Y.; Bando, K. K.; Yasuda, H.; Ishihara, H.; Morita, Y.; Kameoka, T. Appl. Catal., A 2007, 322, 152–171. (32) Tailleur, R. G. Fuel 2008, 87, 2551–2562. (33) Tailleur, R. G. Catal. Today 2008, 130, 492–500. (34) Tailleur, R. G. Energy Fuels 2008, 22, 1509–1518. (35) Casta~ no, P.; Pawelec, B.; Fierro, J. L. G.; Arandes, J. M.; Bilbao, J. Fuel 2007, 86, 2262–2274. (36) Peeters, E.; Cattenot, M.; Geantet, C.; Breysse, M.; Zotin, J. L. Catal. Today 2008, 133, 299–304. (37) de Oliveira, A. M.; Crizel, L. E.; da Silveira, R. S.; Pergher, S. B. C.; Baibich, I. M. Catal. Commun. 2007, 8, 1293–1297. (38) Aguayo, A. T.; Gayubo, A. G.; Vivanco, R.; Olazar, M.; Bilbao, J. Appl. Catal., A 2005, 283, 197–207. (39) Gayubo, A. G.; Aguayo, A. T.; Olazar, M.; Vivanco, R.; Bilbao, J. Chem. Eng. Sci. 2003, 58, 5239–5249. (40) Emeis, C. A. J. Catal. 1993, 141, 347–354. (41) Casta~ no, P.; Arandes, J. M.; Pawelec, B.; Fierro, J. L. G.; Gutierrez, A.; Bilbao, J. Ind. Eng. Chem. Res. 2007, 46, 7417–7425. (42) Gutierrez, A.; Casta~ no, P.; Azkoiti, M. J.; Bilbao, J.; Arandes, J. M. Chem. Eng. J. 2011(doi: 10.1016/j.cej.2011.04.061). (43) Elordi, G.; Olazar, M.; Lopez, G.; Casta~ no, P.; Bilbao, J. Appl. Catal., B 2011, 102, 224–231. (44) Martínez, A.; Prieto, G.; Arribas, M. A.; Concepci on, P. Appl. Catal., A 2006, 339, 224–236. (45) Bartholomew, C. H. Appl. Catal., A 2001, 212, 17–60. (46) Fujikawa, T.; Idei, K.; Ohki, K.; Mizuguchi, H.; Usui, K. Appl. Catal., A 2001, 205, 71–77. (47) Yasuda, H.; Yoshimura, Y. Catal. Lett. 1997, 46, 43–48. (48) Guisnet, M.; Magnoux, P. Appl. Catal., A 2001, 212, 83–96. (49) Guisnet, M.; Costa, L.; Ram^oa Ribeiro, F. J. Mol. Catal., A 2009, 305, 69–83. (50) Draper, N.; Smith, H. Applied Regression Analysis, 2nd ed.; John Wiley and Sons, 1981; pp 307312.
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