Y zeolite catalysts in a semi-batch reactor

Catalysis Today 220–222 (2014) 159–167

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Vacuum gas oil hydrocracking performance of bifunctional Mo/Y zeolite catalysts in a semi-batch reactor Reynald Henry a , Melaz Tayakout-Fayolle a,∗ , Pavel Afanasiev a , Chantal Lorentz a , Gregory Lapisardi b , Gerhard Pirngruber b a Institut de Recherches sur la Catalyse et l’Environnement de Lyon (IRCELyon), CNRS UMR 5256, Université Claude Bernard Lyon 1, 2 Avenue Albert Einstein, 69626 Villeurbanne cedex, France b IFP Energies Nouvelles, Rond-point de l’échangeur de Solaize, BP3, 69360 Solaize, France

a r t i c l e

i n f o

Article history: Received 28 March 2013 Received in revised form 14 June 2013 Accepted 24 June 2013 Available online 8 August 2013 Keywords: Hydrocracking Molybdenum sulfide Zeolite Batch reactor

a b s t r a c t A new approach has been developed to characterize bifunctional catalysts in a complex matrix of hydrotreated vacuum gas oil using a batch reactor test. Triphasic reactions were carried out in a reactor equipped with a stationary basket, hydrogen injection and products sampling systems. Bifunctional catalysts containing different relative amounts of alumina-supported NiMo sulfide and zeolite were tested at 400 ◦ C under 120 bars over different reaction times. The repeatability of the test conditions was validated and the lack of mass transfer limitations at phase interfaces was confirmed. Gas and liquid samples were analyzed by one and two-dimensional gas chromatography (GC × GC) respectively to obtain quantitative distributions of linear and branched paraffins, naphthenes and aromatics. The details of the products distribution provided by chromatography were explained using mechanisms of bifunctional catalysis. It has been established that the limiting step defining the total conversion is the scission of the hydrocarbon chains on acid sites of the zeolite. The increase of the molybdenum to zeolite ratio provided an improvement of middle distillate selectivity. © 2013 Elsevier B.V. All rights reserved.

1. Introduction Hydrocracking (HC) in a triphasic reactor is one of the most important processes of the petrochemical industry. The reduction of the average carbon number and the production of aliphatic hydrocarbons are key reactions in catalytic petroleum refining. In the case of hydrocracking of heavy cuts, such as vacuum gas oil, the challenging task is to produce a high quality middle distillate fraction by avoiding overcracking [1]. The main reactions of isomerization and cracking are usually carried out over a bifunctional catalyst [2]. The bifunctional catalysts are characterized by the simultaneous presence of acidic sites which provide the cracking reaction [3], and of hydrogenation–dehydrogenation sites provided by a metal sulfide or a noble metal [4]. Since the real feeds often contain significant amounts of sulfur and nitrogen, transition metal sulfides are generally applied, because they are more resistant to poisoning. By means of careful balancing between the metal sulfide and acidic functions, the catalyst activity and selectivity might be tuned to a specific targeted product, which is usually either gasoline or middle distillate fraction (MD). The cooperation between acid and metal sites has a strong influence on the distribution of

∗ Corresponding author. Tel.: +33 04 72 44 54 16. E-mail address: [email protected] (M. Tayakout-Fayolle). 0920-5861/$ – see front matter © 2013 Elsevier B.V. All rights reserved. http://dx.doi.org/10.1016/j.cattod.2013.06.024

HC products and on the selectivity in the desired intermediate distillate. In the literature, most studies have been done with model molecules, especially linear and branched alkanes [5–9], though there are some studies on the real feedstocks [10,11]. The hydroisomerization of n-alkanes allowed defining the key parameters for the activity and selectivity of HC catalysts. The catalyst acidity has a major effect on the hydrocracking and hydroisomerization performance. Both the density and the strength of acid sites have crucial importance [12]. Other works emphasized the effect of the pore size, since confinement and adsorption play an important role for HC [13,14]. Thus, in the case of ZSM-22 zeolite, Denayer et al. [15] showed that the hydroisomerization and hydrocracking of nalkanes occur exclusively in the pore mouths. To improve real HC catalysts, working on industrial feedstocks is important. However, at a laboratory scale studying of real feedstocks is a challenging task, because of their excessive complexity. Most studies with industrial feedstocks were carried out in continuous flow reactors [16–19]. Few works consider the role of different loadings of metal sulfide and zeolite on the conversion and selectivity [20]. The present work focuses on a comparison between several catalysts with variable amounts of zeolite and metal sites in the conversion of industrial vacuum gas oil (VGO) (boiling range 370–510 ◦ C). In order to obtain new insights into parameters controlling the HC selectivity, a semi-batch catalytic HC test has been

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R. Henry et al. / Catalysis Today 220–222 (2014) 159–167 Table 1 Main properties of the vacuum gasoil (VGO) feed.

Fig. 1. Composition of studied catalysts.

developed. The feed and the effluent have been analyzed in depth by comprehensive 2D Gas Chromatography (GC × GC) [21]. Carbon number distributions were analyzed for five families of hydrocarbons, including linear paraffins, branched paraffins, naphthenes and aromatics [22]. A very small amount of olefins, less than 0.8 wt%, was detected in the naphthenes fraction through NMR analysis of the effluents. The conversion and product distributions were followed in vapor and liquid phase vs. reaction time. 2. Experimental 2.1. Catalysts synthesis Bifunctional catalysts containing supported Ni–Mo and zeolite were supplied by IFP Energies Nouvelles. The zeolite Y was a commercial sample with code name CBV712 from Zeolyst International. Molybdenum and nickel were loaded on alumina support via an incipient wetness impregnation method. The support was in form of cylindrical shape granules with 5 mm length and 1.6 mm diameter. The sulfidation of the catalysts was carried at 400 ◦ C for 4 h under a 4 L/h flow of gas mixture containing 85% of hydrogen (H2 ) and 15% of hydrogen sulfide (H2 S). Several catalysts have been prepared with different ratios between zeolite and metal sites (see Fig. 1). The x-axis corresponds to zeolite weight percentage in the support before impregnation, while the y-axis corresponds to the global MoO3 weight percentage. The catalysts are designated by Cji where the superscript i is the MoO3 concentration, and the subscript j is the zeolite weight concentration. All concentrations (MoO3 and zeolite) are dimensionless and were calculated relative to the amount of zeolite in the medium-loaded catalyst, defined as unity.

Property

Value

Simulated distillation (ASTM) IBP-FBP, ◦ C Number of carbon atoms Aliphatics/aromatics, wt% Density at 20 ◦ C (g/cm3 ) H/C atomic ratio Chemical analysis

370 to 520 22–42 94.8/5.2 0.8589 1.96 C content: 85.7% H content: 14.0% N content: 0.3%

The catalytic tests were conducted in a 300 mL Parr 4842 bench scale pressure reactor equipped with a Robinson–Mahoney stationary catalyst basket [23]. A schematic representation of the experimental setup is shown in Fig. 2. The catalytic tests were carried out at 400 ◦ C and 120 bars, using 10 g of sulfided catalyst. For each experiment 120 g of vacuum gas oil (VGO) were loaded and 1.48 g of dimethyldisulfide was added into the VGO, as well as 0.51 g of aniline. Under the reaction conditions dimethyldisulfide decomposes releasing hydrogen sulfide that maintains the sulfidation state of the catalyst metals. Aniline is decomposed to release ammonia, in order to inhibit partially the acid sites of zeolite. At 50 ◦ C the reactor was filled with hydrogen to a pressure of 72.5 bars. Then the system was heated to 340 ◦ C without adding hydrogen and next, while raising temperature from 340◦ C to 400 ◦ C additional hydrogen was progressively injected to achieve target pressure of 120 bars. The reaction temperature was reached within 22 min under vigorous stirring, ensuring efficient contact between gas, liquid and solid. The zero reaction point is defined as the moment (T0 , PT0 ) when the temperature reached 400 ◦ C. Several reaction times have been investigated: 0, 68, 98, 113, 158, 248 and 338 min. A new experiment, with a fresh feed was performed for each reaction time and for each experiment the product samples were collected. At the end of each test, the reactor was cooled down from 400 ◦ C to ambient temperature in 30 min. During the reaction, hydrogen was injected in the gas phase in order to compensate its consumption and to keep the total pressure at or above 120 bars. The pressure regulation was handled by a West 8800 process controller associated with a Brooks 5850S mass flow controller, while temperature regulation was operated by a Watlow 987 process controller.

2.2. Catalysts characterization Transmission electron microscopy (TEM) was done on a JEOL 2010 device using a 200 kV accelerating voltage. Ethanol suspension of an analyzed solid was put on a copper grid sample holder, covered with holey carbon film. The images were treated with Digital Micrograph software (Gatan). Specific surface area and pore volume of the solids were determined by nitrogen adsorption–desorption at 77 K using Micromeritics ASAP 2010 device. Before measurements, the samples were evacuated at 300 ◦ C for 3 h. Specific areas and pore volumes were calculated from isotherms by applying the BET and BJH methods for mesopores and with density functional theory (DFT) method for micropores. 2.3. Hydrocracking experiments Hydroconversion was performed with a hydrotreated vacuum gas oil (VGO) feedstock. The main properties of the applied VGO are summarized in Table 1.

Fig. 2. Experimental setup scheme.

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Fig. 3. Total pressure and temperature (a), volume of injected hydrogen (b) and vapor phase composition versus reaction time (c).

Injected hydrogen, pressure and temperature were constantly monitored with a Grant 2010 Squirrel Data Logger. Fig. 3 gives an example of the time evolution of the total pressure, the temperature, the flow rate of injected hydrogen and the composition of the vapor phase. Fig. 3c shows that the amount of light hydrocarbons in the vapor phase gradually increases during the reaction, causing the pressure to increase beyond 120 bars after 250 min. That is also why no hydrogen was injected beyond this time, and the amount of hydrogen in the reactor constantly decreased during the reaction. Knowing that the partial pressure of hydrogen has an influence on the hydrogenation–dehydrogenation equilibrium, all catalysts were tested with the reaction times limited at 360 min. Gas samples were analyzed by a micro GC (SRA instruments) coupled to an Agilent 5975C Mass spectrometry detector. The device was equipped with three different modules. Module 1 was a 10 m, 3 ␮m film thickness alumina column, used to analyze 3–6

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carbon atoms compounds. The carrier gas was hydrogen; injection temperature was maintained at 100 ◦ C, while column temperature was kept at 90 ◦ C. Module 2 was a 8 m, 30 ␮m ft. Poraplot U column, used to analyze light hydrocarbons and hydrogen sulfide. The carrier gas was hydrogen; backflush injector temperature was maintained at 90 ◦ C and the column temperature was kept at 80 ◦ C. Module 3 was a 10 m, 12 ␮m ft., 5 A˚ molecular sieve column, used to analyze hydrogen. Carrier gas was argon, backflush injector temperature was maintained at 80 ◦ C and column temperature was kept at 90 ◦ C. All three modules were equipped by TCD detectors to get quantitative data and simultaneous Mass spectrometry detection allows us gas identification. Gas volume compositions were obtained from ␮GC analysis results. Knowing the total pressures in the vapor condenser and in the reactor, partial pressures of hydrogen, hydrogen sulfide and hydrocarbons lighter than hexane were directly calculable through the ideal gas law. Ammonia volume fraction was not measurable and was estimated from initial aniline weight. The gas was typically composed of CH4 , C2 H6 , C3 H8 , C4 H10 , C5 H12 , H2 S and H2 (with, for example, volume percentage of 9.0, 0.3, 0.5, 1.0, 0.3, 4.9 and 77.1, respectively). Most methane is produced by decomposition of disulfide dimethyl during the heating step. Liquid samples, prepared in carbon disulfide, were characterized by two-dimensional gas chromatography. The GC × GC apparatus consists of a modified Agilent 7890A gas chromatograph [24] equipped with a multimode injector and an Agilent capillary flow modulator (Agilent G3486A CFT Modulator). Injection (1 ␮L) was performed at 250 ◦ C with a temperature program of 500 ◦ C/min to 350 ◦ C with a split ratio of 125:1. The detector used was a standard flame ionization detector (FID) at 360 ◦ C, allowing acquisition of quantitative data. Two columns were applied for GC × GC with hydrogen as carrier gas. The first column was a Phenomenex ZEBRON DB1 HT, 15 m × 0.1 mm × 0.1 ␮m. The second column was a Phenomenex ZEBRON ZB35 HT, 5 m × 0.25 mm × 0.1 ␮m. A flow rate of 0.15 mL/min for the first column and 28 mL/min for the second column in constant flow mode allow us to set the modulation time at 8.965 s, with 8.765 s of sampling time and 1.2 s of injection time. The oven temperature for the first column was programmed from 40 ◦ C (8 min) to 360 ◦ C (5 min) at 2 ◦ C/min and a secondary oven for the second column was programmed from 50 ◦ C (8 min) to 360 ◦ C (27 min) at 2.25 ◦ C/min. Acquisition frequency of the FID detector was set to 100 Hz. The carbon number distributions in the feed and in the liquid products were defined with the GCImage and GCProject softwares from Zoex Corporation, for five families of hydrocarbons: linear paraffins, branched paraffins, naphthenes and aromatics. The liquid density was determined at 20 ◦ C with a Mettler Toledo DR40 apparatus. Simulated distillation (SIMDIS) curves were produced from twodimensional gas chromatography data with the GCProject software which follows the ASTM D2887 standards. Global conversion (vacuum gasoil 370 ◦ C+ cut to 370 ◦ C− cut) as well as selectivity in the middle distillate fraction (boiling range 150–370 ◦ C) were calculated from these SIMDIS data with the following expressions: Conversion270◦ C− =

Fractionfeedstock 370◦ C+ − Fractioneffluent 370◦ C+ Fractionfeedstock 370◦ C+

SlectivityMiddle Distillate =

Fractioneffluent MD − Fractionfeedstock MD Fractionfeedstock370◦ C+ − Fractioneffluent370◦ C+

The average mass balance was 97.5 wt% because some volatile products as C5 H12 and C6 H14 were partially lost upon opening the

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Fig. 4. TEM images of bifunctional catalyst: general view (a); zeolite crystal (b) and alumina supported sulfide (c).

reactor. Moreover, all the liquid was difficult to recover from the impeller and from the top cover of reactor.

3. Results and discussion 3.1. Catalysts characterization Transition electron microscopy gives insight into composition and morphology of the catalysts. In the TEM images, we identified finely divided particles of alumina-supported molybdenum sulfide mixed with relatively large crystallites of zeolite (Fig. 4a). Zeolite crystals are not compact but are traversed by multiple irregular channels formed by dealumination (Fig. 4b). Alumina forms loosely packed agglomerates of particles on which stacks of sulfide slabs of several nm length are supported (Fig. 4c). Sulfide slabs were observed mostly on alumina but not on zeolite. This picture was more or less similar whatever the catalyst composition. However, for the highest zeolite contents, crystals of zeolite tend to agglomerate forming blocks, sometimes tens of microns size (not shown). Nitrogen adsorption isotherms (Fig. 5) show the presence of two types of porosity, obviously related to the presence of zeolite (micropores and part of mesopores) and of the supported sulfide catalyst (mesopores). Textural properties correlate with the catalysts’ composition. A higher fraction of mesopores and a lower

0.7 0.6 Fig. 5. Nitrogen adsorption isotherms for the C0.5 and C1.5 catalysts.

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Table 2 Chemical composition and textural properties of the catalysts under study.

2

Smicroporous (m /g) Smesoporous (m2 /g) Vmesoporous (mL/g) dp micropores (nm) dp mesopores (nm)

0.6 C0.5

0.7 C0.5

0.6 C1.0

0.7 C1.0

0.8 C1.0

0.6 C1.5

0.7 C1.5

0.8 C1.5

82 180 0.61 1.1 14.1

84 123 0.49 1.1 15.2

139 212 0.55 1.1 9.5

149 144 0.47 1.1 10.3

171 110 0.42 1.1 11.0

257 115 0.48 1.1 11.2

219 145 0.50 1.1 10.2

237 73 0.39 1.1 12.0

fraction of micropores were expectedly observed in the catalysts with lower zeolite loadings, as can be inferred from comparison of two curves in Fig. 5, corresponding to the lowest and the highest zeolite loadings, respectively (Table 2).

under our operation conditions was confirmed for the vapor–liquid and liquid–solid phase interfaces.

3.2. Reactor characterization

At the initial point (T0 , PT0 ) the composition of the liquid mixture is already different from the initial feed, because some hydroconversion occurred during the heating period. For each catalyst, the composition of the liquid and of the gas phase was determined at the point (T0 , PT0 ), in order to calculate the conversion and the selectivity. Fig. 6 compares the carbon number distributions of the initial feed and at the point (T0 , PT0 ) for an example of the middle-loaded catalyst. Table 3 gives the values of conversion and selectivity for all catalysts at the point (T0 , PT0 ). It can be seen from Table 3 that the cracking activity increased with an increasing amount of zeolite. Fig. 6b shows that the primary products of hydroconversion are iso-paraffins and some naphthenes. Lesser amount of n-paraffins

Prior to comparing the catalysts performance, we checked the influence of the reaction conditions relevant to the macroscopic mass transfer, namely the stirring rate and the catalyst grain size. 0.7 For the C1.0 catalyst, the influence of the stirring speed on the gasliquid mass transfer was studied. For three different stirring speeds (800, 1000 and 1200 RPM), the conversion, the selectivity and the carbon number distributions were not affected. Experiments with crushed catalysts were also carried out. The grinded catalyst (125–350 ␮m) showed the same performance as the original catalyst pellets. Therefore, the absence of mass transfer limitations

3.3. The initial point conditions

0.7 Fig. 6. Carbon number distributions corresponding to the initial feed and to the zero point of catalytic test (T0 , PT0 ) with catalyst C1.0 .

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Table 3 Conversion and selectivity for all catalysts at point (T0 , PT0 ). 0.6 C0.5 ◦

−

Conversion (370 C ) at (T0 , PT0 ) Selectivity in middle distillate at (T0 , PT0 )

0.7 C0.5

0.6 C1.0

0.7 C1.0

0.8 C1.0

0.6 C1.5

0.7 C1.5

0.8 C1.5

8.8

8.2

12.7

12.5

12.5

18.9

18.1

17.4

83.6

84.0

72.4

72.4

72.4

72.5

72.1

71.7

0.7 Fig. 7. Carbon number distributions corresponding to 68 min (a), 158 min (b) and 338 min (c) of reaction with catalyst C1.0 .

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0.7 Fig. 8. Evolution of the paraffins (a) and unsaturated hydrocarbons families in the middle distillate cut as a function of following total conversion, for the catalysts C1.0 and 0.7 . C0.5

than of iso-paraffins was detected. This observation is expected, since the cracking reactivity of alkanes increases with increasing number of branching in the chain [25]. 3.4. Time evolution of the reaction products distribution Fig. 7 shows the carbon number distributions of n-paraffins, iso-paraffins, naphthenes and aromatics for different reaction 0.7 catalyst. All distributions times, given for the example of the C1.0 moved toward the lower carbon numbers and the products formed were mostly iso-paraffins. Fig. 7(b) shows a very sharp rise of distribution maximum from C22 to C7 for 158 min. The naphthenes and the aromatics present in the initial feed have completely disappeared after 158 min. As expected, the distribution of paraffin products shifted from the gasoil fraction (C13 –C22 ) to the gasoline fraction (C7 –C12 ) with increasing reaction time. From 158 min, the yield of light iso-paraffins and aromatics increased rapidly. The observed increase of aromatics production probably occurs because of the lack of H2 at the end of reaction (cf. Fig. 3). In fact, after 158 min, the lightest products continued to crack to produce light hydrocarbons (C1–C6). The total pressure therefore increased and hydrogen was not introduced anymore because this pressure went over the regulation pressure (120 bars). These

operating conditions corresponding to a lack of hydrogen have been explored in order to have the distribution evolution, particularly for the aromatics. In this case, the aromatics seem to be yielded by the dehydrogenation of naphthenes rather than by the direct dehydrogenation and cyclization of paraffins. At the beginning, the cracking provided high selectivity for C13 –C22 production, particularly for n- and iso-paraffins. After some reaction time, the C13 –C22 selectivity decreased at the expense of C7 –C12 production with high iso-paraffins and aromatics content. All distributions continued to move toward the left but the ratio between the families of the products significantly changed. The amounts of aromatics and iso-paraffins have significantly increased in comparison with the other products. Fig. 8 shows evolution of the hydrocarbon families in the middle 0.7 0.7 distillate fraction versus total conversion for C1.0 and C0.5 catalysts. The amounts of n-paraffin and iso-paraffin families increased linearly with the total conversion. The amount of naphthenes and aromatics increased with the conversion, and then started to decrease above 50%. However above 57% of conversion, fast increase of the proportion of aromatics was observed, corresponding to the beginning of the lack of hydrogen. This result is consistent with what was observed in Fig. 7 for total conversion. In comparison 0.7 0.7 catalyst, the C0.5 catalyst produced more n-paraffins with the C1.0

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Fig. 9. Total conversion (370 ◦ C+ to 370 ◦ C− ) versus contact time at 400 ◦ C for all studied catalysts.

and naphthenes, but less iso-paraffins. There are more n-paraffins in the products of cracking for a low content of zeolite, probably due to a lower isomerization activity. The aromatics were produced in equal quantities for both catalysts. 3.5. HC activity and selectivity as a function of catalyst composition In this section we investigated the parameters defining the catalysts performance, i.e. what kind of site represents a limiting step reaction for the whole process. For the ensemble of this work’s experimental results, we correlated the total conversion values to the contact time between the feed and the zeolite. The contact time was defined by the following expression: yc =

zeolite weight tR feed weight

where tR is the reaction time. Fig. 9 attests that for all catalysts and all reaction times, the total conversion is a monotonously increasing function of the contact time with zeolite. As with the influence of the metal content, its impact on the total activity is much lesser than that of zeolite. Indeed, the points in Fig. 8 curve, corresponding to the catalysts with strongly different metals loading but the same zeolite content, lie in close proximity. It is generally accepted that hydrocracking proceeds as scission of carbenium species derived from the dehydrogenation of alkanes on metal sites and further protonation of alkenes on the Brönsted acidic sites of zeolite [26]. Whatever the details of the mechanism, in our case, the total hydrocracking activity of the catalysts containing either low or high zeolite content does not seem to be limited by the amount of metal. Being rather minor, the metal influence on the overall activity is not negligible. Obviously, the contact time calculated from the weight loading of zeolite gives a good first approximation, but is not an exact descriptor for the specific conversion. While the points on the curve in Fig. 9 follow a clear general trend, some systematic 0.7 deviations can be noticed. In comparison with the C1.0 catalyst with middle zeolite loading, the specific conversions were somewhat higher for the catalysts with the lower zeolite content catalysts (0.5), and somewhat lower for the catalysts with higher zeolite content (1.5). The relative position of the points corresponding to different zeolite content in Fig. 9 suggests that the metal sulfide contributes positively to the overall feed conversion. Alternatively speaking, in the catalysts with lower zeolite loading, zeolite is more efficiently used for the catalytic reaction. By this reason in Fig. 9 we see three well distinguished groups of points corresponding to different zeolite loadings. Fig. 10 shows the evolution of hydrogen consumption versus global conversion (370 ◦ C+ to 370 ◦ C− ). For all catalysts, the quantity of consumed hydrogen increased with the conversion. The quantity of consumed hydrogen was linear until 50% of conversion. Common

Fig. 10. Hydrogen consumption versus conversion (370 ◦ C+ to 370 ◦ C− ) for all studied catalysts.

trend of hydrogen consumption vs. conversion was observed for all catalysts, showing that the reaction pathway is the same whatever zeolite content. Over a rate of conversion corresponding to 50%, the increase of H2 versus conversion consumption slightly deviates from linearity and diminishes. This may be due to the lack of hydrogen. Changing the sulfide to zeolite ratio in the catalysts affects the ratios between the products families and influences the boiling 0.7 0.7 and C0.5 catalysts, the point range of the products. Thus, for C1.0 ◦ diesel (250–370 C boiling point range) to kerosene (150–250 ◦ C) ratios were respectively equal to 0.723 and 0. 689, while the gasoil (200–370 ◦ C) to gasoline (200 ◦ C− ) ratios were respectively equal to 0.938 and 1.190. The diesel to kerosene ratio is similar for both catalysts, which means the middle distillate composition remains similar. On the other hand, the gasoil selectivity is greatly improved 0.7 to the detriment of gasoline for the C0.5 catalyst at similar conversion. The selectivity in the desired middle distillate (MD) products is achieved by means of adjusting the cracking and the hydrogenating functions in a bifunctional catalyst. Fig. 11 shows that for all catalysts studied in this work, the MD selectivity decreased with the rise of conversion. On the other hand, the selectivity in MD increases with the increase of the ratio metal/zeolite. Fig. 11 clearly demonstrates the higher selectivity of the low zeolite catalysts in comparison with other catalysts. In agreement with the literature, when the desired product is middle distillate, moderate acidity and high ratio of hydrogenation to acid functionality are favorable [27]. Indeed, in Fig. 11 we see a systematic upward shift

Fig. 11. Selectivity of middle distillate versus total conversion (370 ◦ C+ to 370 ◦ C− ).

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of the conversion–selectivity curves when metal to zeolite ratio increases. 4. Conclusions In this work, hydrocracking (HC) products distribution of a real complex feed was studied as a function of metal sulfide to zeolite ratio in a newly developed semi-batch reactor. Under our operating conditions, the limiting reaction step defining the total HC conversion is hydrocarbon chain scission on zeolite sites. Within the whole range of compositions studied in this work, the total specific conversion is weakly affected by the amount of metal, at a fixed amount of zeolite. Only at the highest zeolite loading a slight decrease of specific activity was observed, suggesting the first signs of deficiency of hydrogenating functionality. The amount of metal sulfide significantly influences catalysts selectivity. Gradual increase of the fraction of desired middle distillate is in step with the increase of the metal to zeolite ratio. The selectivity toward middle distillate attains the highest values at low zeolite contents, corresponding to the lowest acidity. Analysis of the HC products distributions shows that the reaction follows the mechanistic network generally acknowledged in the literature, in which the n- and iso-paraffins give lighter n- and iso-paraffins, with olefins as intermediates. Further work is directed toward understanding of the reaction kinetics and building of quantitative mechanistic network, describing the transformation between different families of the reaction products. References [1] C. Marcilly, Oil & Gas Science and Technology – Revue de l’IFP 56 (5) (2001) 499–514. [2] G.F. Froment, Catalysis Today 1 (1987) 455–473. [3] C. Leyva, M.S. Rana, F. Trejo, J. Ancheyta, Industrial & Engineering Chemistry Product Research 46 (23) (2007) 7448–7466. [4] M. Steijns, G.F. Froment, Industrial & Engineering Chemistry Product Research and Development 20 (4) (1981) 660–668.

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