A method for hydrogenating and hydrogenating vegetable oil in a single step using at least one sulfide catalyst having BEA structured zeolite alone or in combination with FAU structured zeolite.

Abstract

The present invention discloses a method for processing a feedstock derived from renewable resources, comprising: a) a step of hydrogenating the feedstock in the presence of at least one fixed-bed catalyst, wherein the catalyst includes hydrogenation function and an oxide support; b) a step of hydrogenating and isomerizing at least a portion, preferably all, of the hydrocarbon liquid discharge obtained from step a) in the presence of a bifunctional fixed-bed hydrogenation and isomerization catalyst comprising a sulfide phase of a combination of at least one metal from Group VIII of the periodic table and at least one metal from Group VIB, a support containing BEA structure zeolite alone or as a mixture with FAU structure zeolite, and at least one binder; and c) a step of fractionally distilling the discharge obtained from step b) to obtain at least one diesel fraction.

Description

[Technical Field]

[0001] Exploring new renewable energy sources for fuel production presents a significant challenge, both in meeting fuel demand and in addressing environmental and decarbonization-related concerns in the road and air transport sectors.

[0002] In this regard, the upgrading of fuels from renewable resources has attracted considerable new interest in recent years. Among these fuels, those that can be mentioned are, for example, vegetable oils, such as rapeseed or soybean oil, animal fats, used cooking oils, and mixtures of such fuels. These fuels contain chemical structures, such as triglycerides, esters, or fatty acids. Fatty acid chains consist of hydrocarbon-based structures of variable chain length and generally contain mainly 16 to 18 carbon atoms. Other types of fuels containing fatty acids can also be mentioned, for example, tall oil fuel from the paper industry. [Background technology]

[0003] One possible route is the catalytic conversion of these feedstocks from renewable resources, which is carried out by hydrogenating them (in the presence of hydrogen) into deoxygenated paraffinic fuels. Many metal catalysts or sulfides are known to be active in this type of reaction. Depending on the length of the hydrocarbon base chain, the resulting straight-chain paraffins are compatible in terms of boiling point with hydrocarbons present in fossil gas oils and kerosene fuel bases.

[0004] These methods for hydrogenating feedstock derived from renewable resources are already well known and described in numerous patents. For example, the following patents may be mentioned: Patent Documents 1-4.

[0005] Transition metal sulfide catalysts enable the production of linear paraffins, which is achieved by converting oxygenated compounds via the following two reaction pathways: - Hydrodeoxygenation (HDO): Hydrogen is consumed to form water and hydrocarbons (C) with a carbon number equal to the number of carbon atoms in the original fatty acid chain. n ) leads to the formation - Decarboxylation / decarbonylation (DCO): Formation of carbon oxides (carbon monoxide and carbon dioxide: CO and CO2) and hydrocarbons containing one less carbon relative to the original fatty acid chain (C n-1 This leads to the formation of ).

[0006] This conversion also leads to the formation of by-products, such as propane (from the glycerol structure of fatty substances) and methane (through the methanation reaction of carbon oxides under hydrogenation conditions).

[0007] The liquid effluent obtained from these hydrogenation methods, after gas separation, consists essentially of n-paraffins and is substantially free of sulfur-based, nitrogen-based, and oxygen-based impurities. The sulfur content of this effluent is typically 1–20 ppm by weight, the nitrogen content is generally 0.2–30 ppm by weight, and the oxygen content is generally less than 2000 ppm by weight. The paraffins typically have 9–25 carbon atoms, which largely depends on the composition of the fatty acid chain distribution of the renewable feedstock to be hydrogenated.

[0008] However, this liquid spill is generally not incorporated directly into kerosene or gas oil pools because it does not directly meet all fuel specifications, for example, due to insufficient cold properties and / or excessively high boiling points. Specifically, the linear paraffins present have high pour points, and therefore, can freeze when used at low temperatures. For example, eicosane (a linear paraffin with 20 carbon atoms, C 20 H 42The boiling point of eicosane is equal to 340°C, and its melting point is 37°C. The boiling point of eicosane is therefore suitable for its incorporation into gas oil pools, but its melting point may cause freezing problems and limit its applications. For illustrative purposes, the filtration limit temperature for winter gas oil is a maximum of -15°C according to French regulations (standard EN590). Furthermore, the boiling point of eicosane makes its incorporation into kerosene unsuitable, and therefore the final temperature of the D86 distillation curve must be less than 300°C (ASTM standard D1655).

[0009] Depending on the type of fuel targeted (kerosene or gas oil) and the target fuel specifications, additional hydrogenation conversion steps (hydroisomerization and / or hydrocracking reactions) may be required to convert straight-chain paraffins in the hydrogenated liquid effluent. Hydroisomerization allows for the conversion of straight-chain paraffins to branched paraffins while preserving the number of carbon atoms in the molecule. This improves the cold properties of the effluent because branched paraffins have better cold properties and lower boiling points than their corresponding straight-chain paraffins with the same number of carbon atoms. For example, nonadecane has a melting point of 32°C, while one of its single branched isomers, 7-methyloctadecane, has a melting point of -16°C. Hydrocracking allows for the conversion of straight-chain paraffins to straight-chain or branched paraffins with lower molecular weights. This allows for adjustment of the effluent distillation curve as needed to suit kerosene pools with stricter specifications regarding cold properties and maximum boiling points. For example, the hydrocracking of an eicosane molecule may lead to the production of two 2-methylnonane molecules. The boiling point of 2-methylnonane is 167°C, which is suitable for incorporation into the kerosene pool in terms of boiling point. The hydroconversion step is carried out on a bifunctional catalyst having both hydrogenation / dehydrogenation and Brønsted acid functions. The operating conditions may be adapted as needed to promote the hydrogen isomerization or hydrocracking reaction. In all cases, it is desirable to maximize the yields by minimizing the production of decomposition products that are excessively light and cannot be incorporated into the kerosene and gas oil fractions.

[0010] The appropriate choice of acid phase allows for the isomerization of long, linear paraffins while minimizing decomposition. Therefore, the shape selectivity of zeolites with medium-pore (10MR) or large-pore (12MR) pores makes them particularly suitable for use in obtaining catalysts that are selective for isomerization and hydrogenation. Other types of acid phases, either zeolite or non-zeolite, may be used, such as halogenated (especially chlorinated or fluorinated) alumina, phosphorus-based alumina, silica-alumina, or siliceous alumina.

[0011] However, it is well known that factors other than the acid phase affect the activity and selectivity of bifunctional catalysts. Hydrogenation and hydrocracking of normal paraffins have therefore been the subject of numerous academic studies since the original works of Weisz (Non-Patent Literature 1) or Coonradt and Garwood (Non-Patent Literature 2) in the 1960s.

[0012] The most commonly accepted mechanism involves, firstly, dehydrogenation of the n-paraffin to an n-olefin on the hydrogenation-dehydrogenation site, followed by protonation to a carbenium ion after diffusion to a Brønsted acid site. After structural rearrangement and / or β-cleavage, the carbenium ion desorbs from the acid phase in the form of an olefin after deprotonation. Then, after diffusion to the hydrogenation-dehydrogenation site, the olefin is hydrogenated to form the final reaction product. In this case, it is appropriate to have a sufficiently active hydrogenation / dehydrogenation function on the acid side, firstly, to rapidly supply the olefin to the acid phase, and secondly, to rapidly hydrogenate the olefin intermediate after they have reacted on the acid phase. This makes it possible to maximize the formation of isomerized paraffins by firstly maximizing the activity of the catalyst and secondly limiting excessive decomposition to light hydrocarbons. The use of a sufficiently active hydrogenation function is also desirable to limit the deactivation of the bifunctional catalyst by coking during the hydrogenation conversion of n-paraffin for a certain range of fixed operating conditions (Non-Patent Literature 3).

[0013] Combinations of noble metals (Pt, Pd) or transition metals from Group VIB (Mo, W) with transition metals from Group VIII (Ni, Co) can act as hydrogenation / dehydrogenation catalysts. Noble metals are used in their reduced forms, while transition metals from Group VIB and VIII are used in the form of sulfides.

[0014] The properties of the hydrogenation / dehydrogenation product, and the choice of noble metal or transition metal sulfide types, depend on various criteria, such as economic properties (the price of noble metals is much higher than that of Group VIB and Group VIII transition metals) or the properties of the raw materials to be converted (the impact of the presence of contaminants). Therefore, the hydrogenation / dehydrogenation activity of noble metals is higher than that of transition metal sulfides when the partial pressure of hydrogen sulfide (H2S) in the reaction medium is lower or even zero.

[0015] Patent application Patent Document 5 claims a method for producing paraffinic hydrocarbons from a feedstock containing triglycerides, diglycerides, monoglycerides and / or fatty acids. The method comprises (a) a hydrogenation and deoxygenation step in the presence of hydrogen and a catalyst to obtain an effluent containing water and paraffin; (b) a step of separating the effluent obtained from (a) to obtain a paraffin-rich liquid effluent; and (c) a step of hydrogenating the paraffin-rich effluent in the presence of a catalyst containing hydrogen and nickel sulfide and tungsten and / or molybdenum sulfide as hydrogenation phases and a support containing silica-alumina and / or zeolite. It has been taught that by using sulfide phases instead of noble metals as hydrogenation phases, the need to completely remove impurities from the effluent obtained from step (a) is eliminated.

[0016] The patent (Patent Document 6) describes a method for producing kerosene from renewable feedstock, comprising steps of hydrogenation, isomerization, and selective hydrocracking in the presence of a multifunctional catalyst or a series of catalysts, a gas / liquid separation step of the resulting effluent, followed by a fractionation step to produce a jet effluent, naphtha, and a heavier residual effluent than the jet effluent, which is then recycled to a reaction zone at a recycled / fresh feedstock ratio of 0.1 to 8. The deoxygenation and hydrogenation functions of the catalyst or series of catalysts that may be used in the method according to the present invention can be performed by noble metals, e.g., platinum, palladium, rhodium, and ruthenium, or by a metal sulfide, e.g., NiMo sulfide or NiW sulfide as the active phase. The isomerization and selective hydrocracking functions may be performed by zeolites, e.g., 10-12MR zeolites, e.g., structural zeolites of BEA, MOR, MFI, or FAU, or by amorphous alumina silica. This patent refers to numerous examples of catalysts that can be used in the methods according to the present invention: a Pt-based catalyst dispersed on a support containing Y zeolite; a catalyst containing Pt and Pd on a support containing Y zeolite and amorphous alumina silica. In another embodiment, a catalyst containing Pt and / or Pd on Y or ZSM-5 zeolite, amorphous alumina silica, MOR, SAPO-11 and / or SM3 can be used to catalyze all types of reactions. In another embodiment, a catalyst containing a NiMo sulfide phase on zeolite Y, ZSM-5, amorphous alumina silica, MOR, SAPO-11 and / or SM3 may be used. The catalyst sequence may be an arrangement of a NiMo sulfide catalyst supported on amorphous silica alumina, followed by a Pt-based catalyst supported on amorphous alumina silica. Many other catalysts for selective deoxygenation, isomerization and hydrocracking are also mentioned. For example, a long list of hydrogenocrack / isomerization catalysts is mentioned, comprising a group VIII metal, e.g., Pt and / or Pd, and a support which may be amorphous or crystalline, the support being alumina, amorphous alumina silica, and ferrielite-type zeolites, ALPO-31.It may include SAPO-II, SAPO-31, SAPO-37, SAPO-41, SM-3, MgAPSO-31, FU-9, NU-IO, NU-23, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, ZSM-50, ZSM-57, MeAPO-II, MeAPO-31, MeAPO-41, MeAPSO-11, MeAPSO-31, MeAPSO-41, MeAPSO-46, ELAPO-II, ELAPO-31, ELAPO-41, ELAPSO-II, ELAPSO-31, ELAPSO-41. The patent also teaches that it would be possible to perform the processes of hydroisomerization and hydrocracking without first removing the water and carbon oxides generated during the hydrogenation treatment step. The examples in this patent actually describe the arrangement of the catalysts, but do not specify their properties or the presence of zeolites.

[0017] This effect has been confirmed by Brosius et al. (Non-Patent Document 4), who studied the hydrocracking reaction of n-hexadecane on a Pt catalyst supported on ZSM-5, an MFI-type zeolite, in the presence of a large amount of water (generated in situ by the dehydration of ethanol). It shows that the presence of water greatly reduces the activity of the catalyst, but emphasizes the beneficial effect on selectivity: the secondary cracking observed on the catalyst using ZSM-5 zeolite is suppressed, only primary cracking occurs, and isomerization is also reduced, and thus it is advantageous for the production of linear alkanes. The mode of action caused is the adsorption competition between hydrocarbons and water on the hydrophilic acid sites. Therefore, this study shows a radical modification of activity and selectivity by adding water to the reaction system.

[0018] In attempting to develop a method for treating feedstocks from renewable resources to produce an intermediate distillate, the Applicant has been able to obtain improved activity of the catalyst through the use of a specific hydroisomerization catalyst comprising a sulfide phase combining at least one metal from Group VIB with at least one metal from Group VIII of the Periodic Table, a support comprising at least one zeolite of structure type BEA alone or in combination with a zeolite of structure type FAU, and at least one binder, in an arrangement of at least one hydrotreating step for the feedstock followed by a single-step hydroisomerization step (i.e., without any intermediate separation of the effluent from the hydrotreating step), while demonstrating the maintenance of a high yield of intermediate distillate, preferably the gas oil fraction, relative to the use of conventional catalysts in the prior art in a single-step process.

[0019] The production of a highly catalytically active hydroisomerization catalyst makes it possible, for example, to increase the service life of the catalyst and limit the frequency of fresh catalyst replacement.

[0020] Another advantage of the present invention lies in the single-step process it provides, which includes a single fractionation zone for the effluent from the hydroisomerization step, which is more economical than a two-step process while still achieving a high yield of intermediate distillate fraction, preferably the renewable gas oil fraction.

[0021] Another advantage of the method according to the present invention is that it enables the use of existing hydrotreating units without significant investment.

[0022] For the purposes of the present invention, the various embodiments presented may be used alone or in combination with each other, with no restrictions on the combination.

[0023] For the purposes of the present invention, various ranges of parameters for a given process, such as pressure ranges and temperature ranges, may be used individually or in combination. For example, for the purposes of the present invention, a preferred range of pressure values ​​may be combined with a more preferred range of temperature values.

[0024] In the following text, chemical element groups are given according to the CAS classification (CRC Handbook of Chemistry and Physics, published by CRC press, edited by DR. Lide, 81st edition, 2000-2001). For example, Group VIII in the CAS classification corresponds to the metals from columns 8, 9, and 10 in the new IUPAC classification, and Group VIB corresponds to the metals from column 6.

[0025] In this text, the expressions "of between ... and ..." and "between ... and ..." are synonymous, meaning that both upper and lower limits of the interval are within the range of values ​​stated. If this is not the case, and if both limits are not within the range stated, the present invention will provide such clarification.

[0026] In this document, the expression "...greater than" is understood to mean strictly greater than and is represented by the symbol ">", and the expression "less than" is understood to mean strictly less than and is represented by the symbol "<". [Prior art documents] [Patent Documents]

[0027] [Patent Document 1] U.S. Patent No. 4992605 [Patent Document 2] U.S. Patent No. 5705722 [Patent Document 3] European Patent Application Publication No. 1681337 (Japanese Patent Publication No. 2008-526928) [Patent Document 4] European Patent Application Publication No. 1741768 (Japanese Patent Publication No. 2008-545035) [Patent Document 5] U.S. Patent No. 8809610 [Patent Document 6] U.S. Patent No. 8039682 [Non-patent literature]

[0028] [Non-Patent Document 1] Weisz P., "Adv. Catal.", 1962, 13, 137 [Non-Patent Document 2] HL Coonradt, WE Garwood, "Ind. Eng. Chem.", Process Des, Dev. 3(1), 1964, pages 38-45 [Non-Patent Document 3] Alvarez et al., Journal of Catalysis, 162, 2, 179-189. [Non-Patent Document 4] R. Brosius, PJ Kooyman, JCQ Fletcher, ACS Catal., 2016, 6, 7710 [Overview of the project] [Means for solving the problem]

[0029] (Subject of the invention) More precisely, the present invention relates to a method for processing raw materials obtained from renewable resources, and includes at least the following steps: a) A step of hydrogenating the feed material in a fixed bed in the presence of at least one catalyst; the hydrogenation catalyst comprises a hydrogenation functional material and an oxide support; the temperature during this step is 200 to 450°C, the pressure is 1 MPa to 10 MPa, and the spatiotemporal velocity is 0.1 h -1 ~10h -1 The hydrogen / supply ratio is 70-1700 Nm of hydrogen. 3 / Volume of feedstock (m 3 ) in the presence of the total amount of hydrogen mixed with the feedstock so as to be; b) The step of total hydroisomerization of the hydrocarbon-based liquid effluent from step a); carried out in the presence of a bifunctional fixed bed hydroisomerization catalyst; said catalyst consists of a sulfide phase combining at least one metal from Group VIII of the periodic table with at least one metal from Group VIB, a support containing BEA-structured zeolite alone or in combination with FAU-structured zeolite, and at least one binder; the temperature at which the hydroisomerization step is carried out is 250°C to 500°C, the pressure at that time is 1 MPa to 10 MPa, and the space velocity at that time is 0.1 to 10 h -1 , and the hydrogen / feedstock ratio is 70 to 1500 Nm of hydrogen 3 / Volume of feedstock (m 3 ) in the presence of the total amount of hydrogen mixed with the feedstock so as to be; c) The step of fractionating the effluent from step b); obtaining at least one gas oil fraction.

[0030] According to a preferred embodiment, the hydrotreating step a) and the hydroisomerization step b) are carried out in a single step, and the method according to the invention does not include an intermediate separation step between step a) and step b).

Mode for Carrying Out the Invention

[0031] (Detailed Description of the Invention) (Feedstock) The present invention is particularly dedicated to the preparation of gas oil fuel bases and optionally kerosene fuel bases that meet new environmental standards, starting from feedstocks obtained from renewable resources.

[0032] The feedstocks used in the methods according to the present invention, obtained from renewable resources, are advantageously selected from oils and fats of plant or animal origin, or mixtures of such feedstocks, and contain triglycerides and / or free fatty acids and / or esters. The vegetable oils may advantageously be crude or may be refined whole or partially, and may be obtained from the following plants: rapeseed, sunflower, soybean, palm, palm kernel, olive, coconut, and tung; this list is not limiting. Seaweed or fish oils are also relevant. The animal fats may advantageously consist of residues from the food industry or may also be selected from lard or fats obtained from the catering industry.

[0033] These raw materials essentially contain triglyceride-type chemical structures, known to those skilled in the art under the name fatty acid triesters, as well as free fatty acids. Fatty acid triesters, therefore, consist of three fatty acid chains. These fatty acid chains are in the form of triesters or free fatty acids, and the number of unsaturated atoms per chain, also called the number of carbon-carbon double bonds per chain, is generally 0-3, although this can be higher, especially in the case of oils derived from algae, and generally the number of unsaturated atoms per chain is 5-6.

[0034] The number of molecular unsaturations present in the feedstock obtained from renewable resources used in the present invention is therefore expressed per triglyceride molecule and is preferably 0 to 18. In these feedstocks, the degree of unsaturation is expressed as the number of unsaturations per hydrocarbon-based fatty chain and is preferably 0 to 6.

[0035] Raw materials obtained from renewable resources generally contain various impurities, particularly heteroatoms, such as nitrogen. The nitrogen content in vegetable oils or animal fats is generally about 1 ppm by weight to 100 ppm by weight, depending on their properties.

[0036] Other types of raw materials containing fatty acids may also be mentioned, such as tall oil from the paper industry.

[0037] (Methods and catalysts) Advantageously, the raw material may undergo a pretreatment or pre-refinement step prior to step a) of the method according to the present invention, and through appropriate treatment, contaminants, such as metals, such as alkali compounds, such as alkaline earth metal compounds and phosphorus, may be removed, for example, on an ion exchange resin. Appropriate treatments may be, for example, heat treatment and / or chemical treatments well known to those skilled in the art.

[0038] According to step a) of the method according to the present invention, optionally pretreated feedstock is placed in contact with at least one hydrogenation catalyst on a fixed bed, at a temperature of 200 to 450°C, preferably 220 to 350°C, preferably 220 to 320°C, and more preferably 220 to 310°C. The pressure is 1 MPa to 10 MPa, preferably 1 MPa to 6 MPa, and more preferably 1 MPa to 4 MPa. The spatiotemporal velocity, i.e., the volume of feedstock per unit volume of catalyst and per unit time, is 0.1h -1 ~10h -1 The feedstock is placed in contact with the catalyst in the presence of hydrogen. The total amount of hydrogen mixed with the feedstock is such that the hydrogen / feedstock ratio is 70-1700 Nm³. 3 Volume of raw materials supplied (m³) 3 ), preferably hydrogen 150-1500 Nm 3 Volume of raw materials supplied (m³) 3 It will be made to be ).

[0039] In step a) of the method according to the present invention, the fixed-bed hydrogenation catalyst preferably comprises, alone or in a mixture, at least one metal from Group VIII and / or Group VIB, and a support selected from the group formed by alumina, silica, silica-alumina, magnesia, clay, and at least two of these minerals. The support may preferably contain other compounds, such as oxides selected from the group consisting of boron oxide, zirconia, titanium oxide, and phosphorus pentoxide. A preferred support is an alumina support, more preferably η, δ, or γ alumina.

[0040] The hydrogenation catalyst is advantageously a catalyst comprising a metal from Group VIII, preferably in combination with at least one metal from Group VIB, wherein the metal from Group VIII is preferably selected from nickel and cobalt, either alone or in a mixture, and the metal from Group VIB is preferably selected from molybdenum and tungsten, either alone or in a mixture.

[0041] The content of metal oxides from Group VIII, preferably nickel oxide, is advantageously 0.5% to 10% by weight of nickel oxide (NiO), preferably 1% to 5% by weight of nickel oxide, and the content of metal oxides from Group VIB, preferably molybdenum trioxide, is advantageously 1% to 35% by weight of molybdenum oxide (MoO3), preferably 5% to 30% by weight, and the percentages are expressed as weight percentages relative to the total mass of the catalyst.

[0042] In step a), the total content of metal oxides from Group VIB and Group VIII in the catalyst is preferably 5% to 45% by weight, and more preferably 6% to 35% by weight, relative to the total mass of the catalyst.

[0043] The hydrogenation catalyst used in step a) of the method according to the present invention may be selected to direct the reaction selectivity toward a hydrogenation pathway that preserves as many carbon atoms as possible in the fatty chain, i.e., a hydrogenation deoxygenation (HDO) pathway, thereby upgrading the renewable feedstock to the fuel fraction and maximizing the yield of hydrocarbons that fall within the distillation range of kerosene and / or gas oil, thereby limiting the loss of carbon in the form of carbon oxides and methane. For this purpose, the method is preferably carried out at a relatively low temperature. By maximizing the hydrogenation function, it is also possible to limit polymerization and / or condensation reactions that lead to the formation of coke, which degrades the stability of the catalyst performance.

[0044] The catalyst used in hydrogenation step a) of the method according to the present invention may, advantageously, contain a doping element, which is selected from phosphorus and boron alone or in a mixture, and is preferably phosphorus. The doping element may be introduced into a matrix or preferably deposited on a support. Silicon may be deposited on the support alone or together with phosphorus and / or boron and / or fluorine.

[0045] The weight content of the oxide of the doping element is, advantageously, less than 20%, preferably less than 10%, and advantageously at least 0.001%.

[0046] The metal catalyst used in hydrogenation step a) of the method according to the present invention is a metal sulfide or a metal phase, preferably a metal sulfide.

[0047] Using a single catalyst or several identical or different catalysts simultaneously or sequentially in step a) of the method according to the present invention would not constitute a departure from the context of the present invention. This step may be carried out industrially in one or more reactors having one or more catalyst beds.

[0048] The hydrogenation treatment step a) enables hydrogenation deoxygenation, hydrogenation denitrification, and hydrogenation desulfurization of the supplied raw material.

[0049] According to step b) of the method according to the present invention, at least a portion, preferably all, of the effluent from step a) of the method according to the present invention is converted in the presence of a bifunctional fixed-bed hydrogenation isomerization catalyst, the catalyst comprising a sulfide phase comprising at least one metal from Group VIII of the periodic table combined with at least one metal from Group VIB, a support comprising BEA-structured zeolite alone or in combination with FAU-structured zeolite, and at least one binder, the temperature during which the hydrogenation isomerization step is performed is 250 to 500°C, the pressure is 1 MPa to 10 MPa, and the spatiotemporal velocity is 0.1 to 10 h -1 The hydrogen / supply ratio is 70-1500 Nm of hydrogen. 3 Volume of raw materials supplied (m³) 3 This is carried out in the presence of the entire amount of hydrogen that is mixed with the feedstock so that it becomes ).

[0050] The operating conditions for step b) of the hydrogen isomerization process are adjusted to promote the hydrogen isomerization and / or hydrocracking reactions. Preferably, the temperature during step b) of the hydrogen isomerization process of the method according to the present invention is 250°C to 450°C, more preferably 250°C to 400°C, the pressure is 2 MPa to 10 MPa, more preferably 1 MPa to 9 MPa, and the spatiotemporal velocity is advantageously 0.2 to 7 h -1 , preferably 0.5 to 5 hours -1 Therefore, the hydrogen flow rate in that case is advantageous if the hydrogen / supply volume ratio is the volume of the supply material (m³). 3 ) Hydrogen per unit: 100-1000 standard m 3 Preferably, the volume of the raw material to be supplied (m³ 3 ) Hydrogen per unit: 150-1000 standard m 3 It will be done in such a way.

[0051] According to the present invention, the hydrogenation isomerization catalyst is a bifunctional catalyst comprising a sulfide phase comprising at least one metal from Group VIII of the periodic table and at least one metal from Group VIB, a support comprising at least one zeolite of the BEA structure type, either alone or in combination with a zeolite of the FAU structure type, and at least one oxide binder.

[0052] The hydrogenation isomerization catalyst is advantageously a catalyst comprising a sulfide phase comprising at least one group VIII metal and at least one group VIB metal, wherein the group VIII metal is preferably selected from nickel and cobalt alone or as a mixture, and the group VIB metal is preferably selected from molybdenum and tungsten alone or as a mixture. Preferably, the hydrogenation isomerization catalyst comprises a nickel-molybdenum sulfide phase, a nickel-molybdenum-tungsten sulfide phase, or a nickel-tungsten sulfide phase. Preferably, the hydrogenation isomerization catalyst comprises a nickel-tungsten sulfide phase.

[0053] The content of a group VIB metal, preferably tungsten and / or molybdenum, is advantageously 5% to 45% by weight, preferably 10% to 40% by weight, and very preferably 15% to 35% by weight, relative to the finished catalyst, in terms of oxide equivalents, and the content of a group VIII metal, preferably nickel and / or cobalt, is advantageously 0.5% to 10% by weight, preferably 1% to 8% by weight, and very preferably 1.5% to 6% by weight, relative to the finished catalyst, in terms of oxide equivalents, in the catalyst. According to the present invention, the catalyst is used in the form of its sulfide.

[0054] The catalyst used in step b) of the hydrogenation conversion step of the method according to the present invention may, advantageously, contain a doping element, which is selected from phosphorus and boron alone or in a mixture, and is preferably phosphorus. The doping element may be introduced into the matrix or, preferably, deposited on a support. Silicon may be deposited on the support alone or together with phosphorus and / or boron and / or fluorine.

[0055] The weight content of the oxide of the doping element is, advantageously, less than 20%, preferably less than 10%, and advantageously at least 0.001%.

[0056] The metal is advantageously introduced into the catalyst by any method known to those skilled in the art, such as co-melting, dry impregnation, over-impregnation, or exchange impregnation.

[0057] According to the present invention, the hydrogenation isomerization catalyst also advantageously comprises either a BEA-structured zeolite alone or in a mixture with a FAU-structured zeolite, and at least one oxide binder.

[0058] Preferably, the BEA-structured zeolite is a beta-zeolite, and the FAU-structured zeolite is a Y-zeolite.

[0059] Preferably, the hydrogenation isomerization catalyst comprises beta-zeolite alone or in a mixture with Y-zeolite.

[0060] In a preferred embodiment, the hydrogenation isomerization catalyst comprises Y zeolite and beta zeolite.

[0061] In another preferred embodiment, the hydrogenation isomerization catalyst comprises only beta-zeolite.

[0062] The binder is advantageously selected from silica (SiO2), alumina (Al2O3), clay, titanium oxide (TiO2), boron oxide (B2O3), and zirconia (ZrO2), either alone or in mixtures. Preferably, the binder is selected from silica, silica-alumina, and alumina, and more preferably, the binder is alumina in all its forms known to those skilled in the art, such as gamma alumina.

[0063] A suitable hydrogenation isomerization catalyst comprises a nickel-tungsten sulfide phase, at least one Y-zeolite and at least one beta-zeolite, preferably a support made of these, and at least one alumina binder, preferably composed of these.

[0064] Preferably, the carrier contains 5% to 50% by weight, preferably 7% to 45% by weight, and most preferably 10% to 40% by weight of zeolite relative to the total weight of the catalyst.

[0065] The hydrogenation process a) and the hydrogenation isomerization process b) may be advantageously carried out in a single reactor or in different reactors, preferably in a single reactor.

[0066] In cases where steps a) and b) are carried out in a single reactor, one or more catalyst beds containing at least one hydrogenation catalyst may be used. Similarly, one or more catalyst beds containing at least one hydrogenation isomerization catalyst according to the present invention may be used.

[0067] In the hydrogenation process a), the proportion of the hydrogenation catalyst is advantageously 50% to 90%, preferably 55% to 85%, of the total catalyst volume.

[0068] In step b) hydrogenation isomerization, the proportion of the hydrogenation isomerization catalyst is advantageously 10% to 50%, preferably 15% to 45%, of the total catalyst volume.

[0069] The term "total catalyst volume" means the sum of the volume of hydrogenation catalyst and the volume of hydrogenation isomerization catalyst contained in each of steps a) and b), regardless of whether steps a) and b) are performed in a single reactor or in several reactors.

[0070] According to step c) of the method according to the present invention, the effluent from step b) undergoes a fractionation step that enables the recovery of at least one gas-oil fraction.

[0071] Preferably, step c) includes a gas-liquid separation step and a subsequent water removal step.

[0072] Step c) may advantageously include an atmospheric distillation step and optionally a vacuum distillation step to obtain at least one intermediate distillate fraction.

[0073] The objective of step c) is to separate the gas from the liquid, remove water, and, in particular, recover a hydrogen-rich gas which may contain a light fraction, for example, a C1-C4 fraction and at least one gas-oil fraction, may contain at least one kerosene fraction, and may contain at least one naphtha fraction.

[0074] In a preferred embodiment, the method according to the present invention does not include any steam decomposition step directed at at least one of the effluents from step c).

[0075] (Examples) Examples 1 to 5 describe the preparation of catalysts C1 to C5.

[0076] Examples 6 to 9 describe the evaluation of hydrogenation treatment and hydrogenation isomerization of feedstock obtained from renewable resources using catalyst combinations in which C1 is present together with one of catalysts C2 to C5.

[0077] (Example 1: Preparation of hydrogenation catalyst (C1)) This catalyst is based on nickel, molybdenum, and phosphorus on alumina. The molybdenum oxide (MoO3) content is 22% by weight, the nickel oxide (NiO) content is 4% by weight, and the phosphorus oxide (P2O5) content is 5% by weight, relative to the total weight of the finished catalyst supported on γ-alumina. Shaping is performed using a die with an orifice having a diameter of 1.85 mm. The catalyst is obtained by dry impregnation with an aqueous solution containing metal precursors (molybdenum trioxide and nickel hydroxycarbonate) and orthophosphoric acid. This catalyst then undergoes a calcination process.

[0078] (Example 2: Preparation of a hydrogenation catalyst (C2) that does not conform to the present invention) (Preparation of silica-alumina gel SA1) Boehmite gel G1 is prepared according to Example 1 of Patent US4154812 and spray-dried. This has the following characteristics:

[0079] [Table 1]

[0080] 126 g of this boehmite gel is dispersed in 1450 g of water acidified with 3.9 g of 68% nitric acid. The resulting suspension is stirred at ambient temperature using a mechanical stirrer. A sodium silicate solution is diluted to the concentration required to obtain an amount equivalent to 60 g of SiO2 in the sol, and 1 liter of silica sol is prepared by passing this diluted sodium silicate solution through an ion exchange resin (pre-acidified). The resulting silica sol is added to the boehmite suspension at a flow rate of 22 mL / min using a peristaltic pump. The mixture is then heated to 60°C and aged at this temperature for 1 hour with stirring. The suspension is then filtered using a sintered Buchner device to obtain silica-alumina gel SA1.

[0081] X-ray fluorescence measurement of silica-alumina gel SA1 indicates that the silica content is 32.2% by weight, expressed as the weight percentage of SiO2 relative to the total oxide content (SiO2 + Al2O3).

[0082] The ignition loss for this gel SA1 is 70.8%.

[0083] Loss on ignition corresponds to the water content of the material; this is measured through the loss of mass after heat treatment at 1000°C for 4 hours.

[0084] (Zeolite Z1) USY type Z1 zeolite (Si / Al ratio 15 at / at, lattice constant 24.28 A) is used.

[0085] (Molding of carrier S1) 218 g of silica-alumina gel SA1, 8 g of zeolite Z1, and 11.7 g of boehmite gel G1 were mixed in a Z-arm kneader and kneaded at 50 rpm. The resulting paste was then extruded through a 2.5 mm diameter trilobed die. The amount of zeolite Z1 added corresponds to a 10% by weight mass content of Z1 relative to the total weight of the dry carrier.

[0086] (Hydrothermal treatment of carrier S1) After drying in a ventilated oven at 80°C for 20 hours, the extruded material is subjected to hydrothermal treatment at 450°C for 2 hours under an airflow containing less than 40 g of water per kilogram of dry air, and then at 800°C for 2 hours in the presence of water vapor containing 200 g of water per kilogram of dry air.

[0087] (Preparation of catalyst C2 (not conforming to the present invention)) The extruded carrier S1 is then subjected to a dry impregnation process with aqueous solutions of ammonium metatungstate and nickel nitrate, followed by static maturation in a water maturator at room temperature for 24 hours, and then drying at 120°C for 5 hours. The weight content of the completed dried catalyst is 21% tungsten oxide WO3 and 3.5% nickel oxide NiO.

[0088] (Example 3: Preparation of a hydrogenation-conversion catalyst (C3) conforming to the present invention) To prepare the catalyst support, a mixture of alumina gel and beta zeolite (Si / Al ratio 13 at / at) is shaped by kneading and extrusion through a die equipped with a 2 mm diameter trilobed orifice, dried at 80°C, and then calcined at 550°C. The beta zeolite content in the support is 10% by weight.

[0089] The carrier extrudes are then subjected to a dry impregnation process using aqueous solutions of molybdenum trioxide, nickel hydroxycarbonate, and phosphoric acid, matured in a water maturer at room temperature for 24 hours, and dried at 120°C for 5 hours. After drying, the finished catalyst has a molybdenum oxide (MoO3) weight content of 19%, a nickel oxide (NiO) content of 3.8%, and a phosphorus (P2O5) content of 4.3%.

[0090] (Example 4: Preparation of a hydrogenation-conversion catalyst (C4) conforming to the present invention) To prepare the catalyst support, a mixture of alumina gel and beta-zeolite (Si / Al ratio 13 at / at) and USY-zeolite (Si / Al ratio 15 at / at, lattice constant 24.28 A) was shaped by kneading and extrusion through a die equipped with a 2 mm diameter trilobed orifice, dried at 80°C, and then calcined at 550°C. The beta-zeolite content in the support is 3% by weight. The USY-zeolite content in the support is 10% by weight.

[0091] The carrier extrudes are then subjected to a dry impregnation process using aqueous solutions of ammonium metatungstate and nickel nitrate, matured in a water maturer at room temperature for 24 hours, and dried at 120°C for 5 hours. After drying, the final catalyst has a weight content of tungsten oxide WO3 of 28% and a nickel oxide NiO content of 3.6%.

[0092] (Example 5: Preparation of a hydrogenation catalyst (C5) conforming to the present invention) To prepare the catalyst support, a mixture of alumina gel and beta-zeolite (Si / Al ratio 13 at / at) and USY-zeolite (Si / Al ratio 15 at / at, lattice constant 24.28 A) was shaped by kneading and extrusion through a die equipped with a 2 mm diameter trilobed orifice, dried at 80°C, and then calcined at 550°C. The beta-zeolite content in the support is 5% by weight. The USY-zeolite content in the support is 5% by weight.

[0093] The carrier extrudes are then subjected to a dry impregnation process using aqueous solutions of molybdenum trioxide, nickel hydroxycarbonate, and phosphoric acid, matured in a water maturer at room temperature for 24 hours, and dried at 120°C for 5 hours. After drying, the weight content of the finished catalyst is 19% molybdenum oxide (MoO3), 3.8% nickel oxide (NiO), and 4.3% phosphorus (P2O5).

[0094] (Examples 6-9: Evaluation of hydrogenation treatment and hydrogenation conversion of raw materials supplied from renewable resources using the method according to the present invention, in combination of catalyst C1 with catalysts C2-C6) In a temperature-regulated reactor designed to ensure isothermal operation, the fixed bed is packed with 60% by volume of hydrogenation catalyst C1 and 40% by volume of hydrogenation-conversion catalyst. Catalyst C1 is placed at the reactor inlet, and the hydrogenation-conversion catalyst is placed at a second location (at the reactor outlet). The catalysts are pre-sulfurized, and the hydrogenation and hydrogenation-conversion of pre-refined rapeseed oil is carried out. The density of this rapeseed oil is 920 kg / m³.3 The oxygen content is 11% by weight. Table 2 shows the detailed fatty acid distribution of rapeseed oil. Prior to the hydrogenation process, dimethyl disulfide is added to the feedstock to adjust its sulfur content to 50 ppm by weight.

[0095] [Table 2]

[0096] Before hydrogenating the raw material, the catalyst was added to the unit with 2% by weight of dimethyl disulfide, and the hydrogen / supplemented gas oil was sulfurized in-situ under a total pressure of 7 MPa. 3 / m 3 The volume of sulfurization feedstock per unit volume of catalyst and per unit time is set to 0.41. Sulfurization is carried out at 350°C for 12 hours, with a temperature rise rate of 10°C per hour.

[0097] After sulfidation, the operating conditions of the unit are adjusted to perform hydrogenation and hydrogenation of the feedstock: - HSV (Supply material volume / Total catalyst volume / Time): 0.41h -1 , - Total operating pressure: 7 MPa, - Hydrogen / supply ratio: 1000 Nm 3 Hydrogen / m 3 feedstock.

[0098] The hydrogen used is supplied by Air Products, and its purity exceeds 99.999% by volume.

[0099] The severity of the hydrogenation conversion is varied by applying a stable temperature range of 330-360°C.

[0100] At the unit's outlet, in-line analysis using gas chromatography and a gas counter can be used to calculate the mass of light hydrocarbons generated and present in the hydrogen stream.

[0101] The liquid effluent is accumulated over 12 hours. The liquid effluent is then quantified and analyzed by simulated distillation (ASTM D2887) to determine the yield of the intermediate distillate ("120°C+ fraction," corresponding to hydrocarbons present in the gaseous and liquid fractions with boiling points above 120°C).

[0102] The yield of the intermediate distillate is calculated as follows: Yield (intermediate distillate) = [(mass of liquid effluent)] * % fraction (120℃+ / 100+ mass [C8-C13] gas) / (mass of liquid effluent + mass of light hydrocarbons (gas) + mass of water + mass of COx) * 100 The mass of the liquid spill corresponds to the mass of the liquid collected over 12 hours.

[0103] % fraction 120°C+ is obtained by simulated distillation: the mass fraction of the liquid effluent with a boiling point above 120°C.

[0104] Mass [C8-C13] gas is obtained by online gas chromatography analysis of the hydrogen stream at the unit outlet. This corresponds to the mass of hydrocarbon-based compounds containing 8 to 13 carbon atoms.

[0105] The mass of water is the sum of the mass of precipitated water present in the liquid effluent and the mass of water present in the gas, and is analyzed online by gas chromatography.

[0106] The mass of COx is the mass of carbon oxides (CO and CO2) present in the gas phase and is determined by online analysis using gas chromatography.

[0107] Furthermore, the cloud point is determined using the ASTM method D5773.

[0108] According to Example 6 (which is not consistent with the present invention), catalyst C1 is packed into the first region (60% of the volume), and then catalyst C2 is packed into the second region (40% of the volume).

[0109] According to Example 7 (which conforms to the present invention), catalyst C1 is packed into the first region (60% of the volume), and then catalyst C3 is packed into the second region (40% of the volume).

[0110] According to Example 8 (which conforms to the present invention), catalyst C1 is packed into the first region (60% of the volume), and then catalyst C4 is packed into the second region (40% of the volume).

[0111] According to Example 9 (which conforms to the present invention), catalyst C1 is packed into the first region (60% of the volume), and then catalyst C5 is packed into the second region (40% of the volume).

[0112] Catalyst C2 is defined as the standard. The performance criteria are as follows: - Conversion activity: Expressed as the temperature increase relative to the standard required to achieve a cloud point of -7°C in the liquid effluent. Negative values ​​reflect an increase in activity. - Yield of intermediate distillate: Obtained for a cloud point of -7°C in the liquid effluent. This is expressed as a deviation relative to the baseline. Negative values ​​reflect yield loss.

[0113] The main characteristics of the resulting spillage and the associated operating conditions are reported in Table 3.

[0114] [Table 3]

[0115] The compliant catalyst C3 enables the production of a gas-oil fraction with a 21°C increase in conversion activity compared to the basic case, while simultaneously limiting the decrease in intermediate distillate yield.

[0116] The compliant catalyst C4 enables the production of a gas-oil fraction with a 19°C increase in conversion activity compared to the basic case, while simultaneously allowing for a +3 increase in the yield of the intermediate distillate compared to the basic case.

[0117] The compliant catalyst C5 enables the production of a gas-oil fraction with a 14°C increase in conversion activity compared to the basic case, while simultaneously limiting the decrease in the yield of the intermediate distillate.

Claims

1. A method for processing raw materials obtained from renewable resources, comprising at least the following steps: a) A step of hydrogenating the feed material in a fixed bed in the presence of at least one catalyst; the hydrogenation catalyst comprises a hydrogenation functional material and an oxide support; the temperature during this step is 200 to 450°C, the pressure is 1 MPa to 10 MPa, and the spatiotemporal velocity is 0.1 h －１ ~10h －１ The hydrogen / supply ratio is 70-1700 Nm of hydrogen. ３ / Volume of raw materials supplied (m³) ３ This is performed in the presence of the entire amount of hydrogen mixed with the supply material. b) A step of hydrogenation isomerization of the entire carbon hydrocarbon-based liquid effluent from step a); carried out in the presence of a bifunctional fixed-bed hydrogenation isomerization catalyst; the catalyst comprises a sulfide phase comprising at least one metal from Group VIII of the periodic table combined with at least one metal from Group VIB, a support containing BEA-structured zeolite alone or in combination with FAU-structured zeolite, and at least one binder; the temperature during the hydrogenation isomerization step is 250°C to 500°C, the pressure is 1 MPa to 10 MPa, and the spatiotemporal velocity is 0.1 to 10 h －１ The hydrogen / supply ratio is 70-1500 Nm of hydrogen. ３ / Volume of raw materials supplied (m³) ３ This is performed in the presence of the entire amount of hydrogen mixed with the supply material. c) Fractionation of the effluent from step b); obtain at least one gas-oil fraction.

2. The method according to claim 1, wherein the raw material obtained from renewable resources is selected from oils and fats of plant or animal origin, or mixtures of such raw materials, and contains triglycerides and / or free fatty acids and / or esters.

3. In step a), the feedstock is placed in contact with the catalyst in a fixed bed, wherein the temperature at that time is 220 to 350 °C, the pressure at that time is 1 MPa to 6 MPa, and the space velocity at that time is 0.1 h －１ to 10 h －１ and the feedstock is placed in contact with the catalyst in the presence of hydrogen, and the method according to claim 1 or 2 is carried out in the presence of the total amount of hydrogen mixed with the feedstock such that the ratio of hydrogen / feedstock is 150 to 750 Nm ３ of hydrogen per volume (m ３ ) of the feedstock.

4. The method according to any one of claims 1 to 3, wherein the fixed-bed hydrogenation catalyst used in step a) preferably comprises, alone or in a mixture, at least one metal from Group VIII and / or Group VIB, and a support selected from the group formed by alumina, silica, silica-alumina, magnesia, clay, and mixtures of at least two of these minerals.

5. The temperature during the hydrogenation isomerization step b) is 250°C to 450°C, more preferably 250°C to 400°C, the pressure at which time is 2 MPa to 10 MPa, more preferably 1 MPa to 9 MPa, and the spatiotemporal velocity at which time is 0.2 to 7 h －１ , more preferably 0.5 to 5h －１ Therefore, the hydrogen flow rate in that case is advantageous if the hydrogen / supply volume ratio is the volume of the supply material (m³). ３ ) 100-1000 standard m of hydrogen per unit ３ Preferably, the volume of the raw material to be supplied (m³ ３ ) Hydrogen per unit: 150-1000 standard m ３ The method according to any one of claims 1 to 4, wherein the method is performed in such a manner.

6. The method according to any one of claims 1 to 5, wherein the hydrogenation isomerization catalyst comprises a sulfide phase comprising at least one group VIII metal and at least one group VIB metal, the group VIII metal is preferably selected from nickel and cobalt alone or as a mixture, and the group VIB metal is preferably selected from molybdenum and tungsten alone or as a mixture.

7. The method according to claim 6, wherein the hydrogenation isomerization catalyst comprises a nickel-molybdenum sulfide phase, a nickel-molybdenum-tungsten sulfide phase, or a nickel-tungsten sulfide phase, preferably a nickel-tungsten sulfide phase.

8. The method according to any one of claims 1 to 7, wherein the hydrogenation isomerization catalyst comprises beta zeolite alone or in a mixture with Y zeolite.

9. The method according to claim 8, wherein the hydrogenation isomerization catalyst comprises Y zeolite and beta zeolite.

10. The method according to any one of claims 1 to 9, wherein the hydrogenation treatment step a) and the hydrogenation isomerization step b) are carried out in a single reactor or in different reactors, preferably in a single reactor.

11. The method according to any one of claims 1 to 10, wherein an intermediate separation step is not included between step a) and step b).

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