FR3142195A1 - METHOD FOR PRODUCING MIDDLE DISTILLATES BY CO-PROCESSING OF MINERAL FEED WITH A RENEWABLE FEED COMPRISING A CATALYST SEQUENCE INCLUDING A BETA ZEOLITE BASED CATALYST - Google Patents

Abstract

The present invention relates to a process for producing middle distillates from at least one fossil hydrocarbon feed mixed with at least one renewable feed comprising a step of hydrotreating said feeds in the presence of hydrogen and at least one catalyst. hydrotreatment and a step of hydrocracking the hydrotreated feedstock, the hydrocracking step comprising at least one step of bringing said hydrotreated feedstock into contact in a first catalytic zone with at least one first catalyst comprising at least one metal from the group VIB and/or at least one metal from group VIII of the periodic table and a support comprising at least one zeolite having at least one series of channels whose opening is defined by a ring with 12 oxygen atoms (12MR), and at least one binder, followed by bringing into contact in a second catalytic zone all of the effluent from step a), without an intermediate separation step between said first catalytic zone and said second catalytic zone, with a second catalyst comprising at least one metal from group VIB and/or at least one metal from group VIII of the periodic table and a support comprising at least one zeolite with structural code BEA and at least one binder.

Description

PROCESS FOR PRODUCING MIDDLE DISTILLATES BY CO-PROCESSING OF MINERAL FEED WITH A RENEWABLE FEED COMPRISING A CATALYST SEQUENCE INCLUDING A BETA ZEOLITE BASED CATALYST

The present invention relates to a process for producing middle distillates comprising a hydrotreatment step and a hydrocracking step, said hydrocracking step being characterized in that it uses two catalysts in two distinct hydrocracking zones operated in series.

This process converts a set of hydrocarbon feedstocks into products of interest (naphtha, jet, diesel). These hydrocarbon fillers contain, for example, aromatic, and/or olefinic, and/or naphthenic, and/or paraffinic compounds, including so-called fossil fillers, such as distillate vacuum cuts (DSV) or Vaccum Gas Oil according to English terminology. Saxon (VGO) from direct distillation of crude or from conversion units such as FCC, coker, H-Oil or visbreaking, feeds from the Fischer-Tropsch process, feeds from sand transformation bituminous and shale oils, in mixture with at least one renewable filler chosen from vegetable oils, algal oils, cooking oils, animal fats, fresh or used, alone or in mixture, and fillers resulting from the reprocessing of biomass/plastics/tires/and household waste, alone or in mixtures. These fillers are likely to possibly contain metals, and/or nitrogen, and/or oxygen and/or sulfur.

The objective of the process according to the invention is essentially the production of a middle distillate cut, comprising a kerosene cut having initial and final boiling points in a range from 130 to 300°C and a gas oil cut having points initial and final boiling points within a range of 220 to 390°C.

In particular, the present invention relates to a process for producing middle distillates from a hydrocarbon feedstock preferably of the vacuum distillate type, mixed with at least one renewable feedstock chosen from vegetable oils, algal oils, cooking oils, animal fats, fresh or used, alone or in mixture, and loads resulting from the reprocessing of biomass/plastics/tires/and household waste, alone or in mixtures, said process comprising a hydrotreatment step and a hydrocracking stage. The hydrocracking step is carried out in the presence of a sequence of 2 specific catalysts, implemented in two distinct catalytic zones, all of the effluent from the first catalytic zone being directly sent to the second catalytic zone on contact of a second catalyst, without intermediate separation of the gas phase between the 2 catalytic zones of said second hydrocracking stage.

In particular, the first catalyst used in the first catalytic zone of the hydrocracking step comprises at least one metal from group VI and/or at least one metal from group VIII of the periodic table and a support comprising at least one large pore zeolite (12MR), making it possible to convert the hydrocarbon feedstock upstream of the second catalyst to a fairly large extent.

The second catalyst used in the second catalytic zone of the hydrocracking step comprises, according to the invention, at least one metal from group VIB and/or at least one metal from group VIII of the periodic table and a support comprising at least one zeolite of structural type BEA.

The process for producing middle distillates including a step of hydrocracking heavy petroleum cuts is a now essential process in refining which makes it possible to produce, from excess heavy feedstocks that are of little use, lighter fractions such as gasoline, jet fuel and gas oils that the refiner is looking for to adapt its production to the structure of demand. Certain hydrocracking processes also make it possible to obtain a highly purified residue which can provide excellent bases for the production of oils. Compared to catalytic cracking (FCC), the advantage of catalytic hydrocracking is to provide very good quality middle distillates. Conversely, the gasoline produced has a much lower octane number than that produced by catalytic cracking.

This process draws its flexibility from three main elements which are, the operating conditions used, the types of catalysts used and the fact that the hydrocracking of hydrocarbon feedstocks can be carried out in one or two steps. It is thus suitable for treating and transforming any type of hydrocarbon feedstock as long as it is compatible with injection in liquid form into catalytic reactors. In particular, it is likely to use, alone or in mixture, fillers resulting from the distillation of a petroleum crude, but also vegetable oils, animal fats for example.

It can be carried out in one step in one or more reactors in series in the presence of a single catalyst which not only hydrocracks the hydrocarbon feed into lower boiling point products but which also converts the residual organic compounds including sulfur and l nitrogen into hydrogen sulphide and ammonia respectively.

The hydrocracking catalysts used in middle distillate production processes are called bifunctional, that is to say combining an acid function with a hydro-dehydrogenating function. The acid function is provided by oxides whose surface areas generally vary from 150 to 1,000 m ² .g ^-1 such as halogenated aluminas (chlorinated or fluorinated in particular), combinations of boron and aluminum oxides, silica -amorphous alumina and zeolites. The hydro-dehydrogenating function is provided either by one or more metals from group VIB of the periodic table of elements, or by a combination of at least one metal from group VIB of the periodic table and at least one metal from group VIII.

The balance between these two functions is one of the parameters which governs the activity and selectivity of the catalyst. A weak acid function and a strong hydro-dehydrogenating function give catalysts that are not very active, working at a generally high temperature (greater than or equal to 390-400°C), and at a low spatial feed speed (the VVH expressed in volume of load to be treated per unit volume of catalyst and per hour is generally less than or equal to 2 h ^-1 ), but with very good selectivity in middle distillates (jet fuels and gas oils). Conversely, a strong acid function and a weak hydro-dehydrogenating function give active catalysts, but with poorer selectivities in middle distillates.

One type of conventional hydrocracking catalysts is based on moderately acidic amorphous oxides, such as silica-aluminas for example. These systems are used to produce good quality middle distillates, and possibly oil bases. The disadvantage of these catalysts based on amorphous support is their low activity.

Catalysts comprising, for example, zeolite Y, or catalysts comprising, for example, zeolite Beta, have a catalytic activity greater than that of silica-aluminas, but have generally lower selectivities in middle distillates (jet fuels and gas oils).

In general, hydrocracking catalysts which contain only zeolite Y do not make it possible to obtain good cold properties of the gas oil cut (in terms of cloud point, pour point, filterability limit temperature of said cut), but also kerosene cut (in terms of crystal appearance point). These values are highly correlated with the higher boiling point linear paraffins present in these sections. To comply with fuel specifications on cold properties, this may force the refiner to have to lower the final boiling point of the cuts (to reduce the high boiling paraffin content) with the direct effect of reducing their yield. in favor of heavy cuts that are not or less recoverable.

The prior art reports numerous works to improve the middle distillate selectivity of zeolite catalysts in hydrocracking processes. The latter are composed of a hydro-dehydrogenating phase based on transition metals, generally deposited on a support containing a zeolite, most often a USY zeolite. The hydro-dehydrogenating phase is generally in the transition metal sulfide form.

For example, we can cite work relating to the use of catalysts comprising a zeolite Y modified for example by dealumination by steaming or acid attack, the use of composite catalysts, or the use of small crystals of zeolites Y. other patent applications such as patent US7585405 describe the use of a catalyst comprising a mixture of zeolites such as Beta and USY zeolites for improving the performance of hydrocracking catalysts. A fine adjustment of the USY/Beta mass ratio is necessary to maximize the yield of middle distillates, this with a USY zeolite which it is also necessary to precisely select with in particular a mesh parameter between 24.37 and 24.44 Å. No mention is made of the evolution of the relative yields in diesel and kerosene cuts.

Patent application WO08085517, on the contrary, provides for the use of a single source of zeolite for the formulation of the catalyst. The Beta zeolite introduced has a SAR of less than 30, and it is taught that it is preferable for it to undergo little post-treatment after elimination of the structuring agent used in manufacturing. Yield gains in average distillates of the order of 2 to 3 points are demonstrated, compared to other catalytic formulations with Beta zeolite having undergone different heat treatments in the presence of water vapor (steaming), without details on the respective yields in diesel or kerosene cuts.

As a general rule, when the catalyst or the operating conditions of the hydrocracking process are modified, various technical solutions have been proposed and lead to an improvement in the yield of middle distillates, but this increase is generally driven by a gain in yield in gas oil cut. , the gain in kerosene section remaining low or even zero.

An alternative would consist, as described in patent US7749373, in using a first hydrocracking catalyst containing an acid function followed by a second catalyst containing a Beta zeolite, without any intermediate separation of the gas or liquid products between the two. catalytic zones. Preferably in this case, the second catalyst comprises a Beta zeolite having a SAR greater than 25 and very preferably greater than 250. The examples demonstrate that after prior hydrotreatment of a distillate type charge under vacuum, the sequence of, firstly, 75% by volume of a catalyst comprising 10% by weight in its support of a zeolite Y of SAR equal to 30 and having a mesh parameter equal to 24.29 Å, then of a catalyst comprising 3% of Beta in its support, said Beta having a SAR of 300, makes it possible to reduce the operating temperature by 4°C, to reach 87% target conversion, compared to loading carried out with only the catalyst supported on zeolite Y. We then also note a gain of 3.5% by weight on the average distillate yield which is entirely driven by the increase in yield on the diesel cut, the kerosene yield not changing (as well as its cold properties) . In other words, the process implementing said catalyst sequence allows an improvement in the yield of the gas oil cut as well as an improvement in the cold properties of said gas oil cut in terms of pour point but said process does not allow any improvement in the yield of the gas oil. the kerosene cut obtained. It is the same for the cold properties of the kerosene cut which are identical to that obtained compared to a process using only a catalyst comprising a zeolite Y, at the same quantity of catalyst used.

In order to specifically maximize the production of middle distillates and reduce the carbon footprint of refining processes, it may also prove advantageous to work with renewable feedstocks, in addition to fossil feedstocks, likely to lead to a reduction in the carbon footprint. Among the renewable loads, we can cite for example vegetable oils, animal fats, or used cooking oils which after hydrotreatment generally lead to long linear paraffins comprising mainly nC ₁₅ to nC ₁₈ hydrocarbons with boiling points compatible with the fossil fuel cutting. However, these linear paraffins give the diesel fuel degraded cold properties, which implies having to isomerize these compounds to achieve the required fuel specifications. Another advantageous solution is also to convert these linear paraffins by selective hydrocracking for direct incorporation into the kerosene cut. This also makes it possible to introduce renewable carbon into this fossil cut. The production of Biokerosene, or also called SAF according to Anglo-Saxon terminology (“Sustainable Aviation Fuel”), is particularly sought after to decarbonize the aviation sector, whether civil or military.

Thus, in the context of reducing greenhouse gas emissions, refiners must increase the proportion of components of biological origin in fuels. The public authorities aim to promote the use of biofuels in transport, not only road transport (diesel and gasoline engines) but also air transport. For example, the Ancre (national alliance for coordinating research for energy) roadmap, published in June 2018, assesses the French potential of biofuel production sectors for aviation, taking into account the maturity and development time of these industrial sectors. The following deployment trajectory therefore reflects the French ambition for 2030 of an incorporation of aviation biofuels in France of 5%. In addition, the national low-carbon strategy (SNBC) revised in 2018 sets a long-term orientation for air transport by 2050 aiming at a 50% substitution of conventional fuel of fossil origin by biofuels. Biofuel production has therefore been promoted in recent years and will continue to be greatly strengthened in the years to come.

Numerous patent applications describe the co-treatment of renewable feedstocks with a fossil feedstock with the aim of producing either a diesel cut, a kerosene cut, or both, cuts incorporating a portion of biogenic carbon.

For example, the incorporation of lipidic renewable feedstocks containing triglycerides with gas oil cuts which are petroleum fossil feedstocks in the co-treatment hydrotreatment process has been described in documents FR 2 949 475, EP 1 693 432 and EP 2,046,917 B1. This incorporation makes it possible to produce a renewable diesel cut and integrate it directly into the diesel pool from fossil resources without investing in a dedicated hydrotreatment unit. The advantages of co-processing a feed containing triglycerides with a fossil petroleum cut compared to the treatment of a pure lipid feed are well detailed in document FR 2 949 475 A1, such as for example:

- Limitation of polymerization linked to the thermal instability of renewable loads,

- Reduction by dilution effect of sulfur, nitrogen and aromatic contents,

- Limitation of exotherms linked to the hydrotreatment of biological loads,

- Reduction in the CO and CO ₂ content in the gaseous effluent: limitation of the effect of inhibiting catalytic activity by CO and limitation of the risk of corrosion in the presence of water and CO ₂ ,

- Increase in the solubility of hydrogen in the mixture to be treated.

The hydrotreatment process makes it possible to essentially produce a fraction of renewable gas oil depending on the rate of incorporation at the input to the process, that is to say cuts with an initial boiling point of 250°C or 280°C and final going up to 340°C, or even 370°C.

The hydrocracking process makes it possible to produce, in addition to renewable diesel, lighter biogenic products such as biokerosene and bionaphtha.

The production of biokerosene (or SAF) is attracting more and more attention compared to the production of renewable diesel in the context of the energy transition. On the other hand, the valorization of biological resources into bio naphtha for petrochemicals (with a view to the production of olefins, BTX aromatics, polymers and cosmetic products, etc.) is also increasingly interesting. Refiners are therefore increasingly interested in technical solutions allowing the incorporation of biokerosene into fossil kerosene and the production of bionaphtha for petrochemicals.

It is described in the invention WO 2009/148909 A2, that pure feedstocks containing triglycerides such as palm oil and animal fat can be converted on the sequence of hydrotreatment steps (on a sulfurized NiMo catalyst supported on alumina) and hydrocracking (on a Pt-Pd catalyst supported on amorphous silica-alumina) for the production of bio naphtha, for yields up to 28% by weight.

Patent application WO 2011/012439 A1 describes that lipid feedstocks can be converted into n-paraffins and propane by hydrodeoxygenation and decarbonylation over a conventional hydrotreatment catalyst. These n-paraffins with fossil naphtha are then mixed and injected into the steam cracking unit for the production of olefins, diolefins and aromatics.

Document EP 2 143 777 A1 describes a process for hydrocracking a feedstock consisting of a mixture of a vacuum distillate with vegetable feedstocks or animal fats, with a sequence of hydrotreatment catalyst (NiMo/Al ₂ type O ₃ ), followed by a hydrocracking catalyst (NiMo/zeolite type) then by a post-treatment catalyst (hydrodesulfurization type) making it possible to obtain a gas oil with a very low sulfur content. The incorporation of a biological load with the vacuum distillate, according to this invention, makes it possible to significantly increase the yield in diesel fuel cut but unfortunately this results in a reduction in the yield of kerosene. For example, he notes an increase of 6% in the yield of the diesel cut but a reduction of more than 3% in the kerosene cut with the incorporation of 10% by weight of soybean oil mixed with a DSV. On the other hand, the diesel cut obtained has degraded cold properties: cloud point 3°C higher than the fossil diesel obtained without the incorporation of soybean oil.

It is well known that under hydrotreatment operating conditions, the n-paraffins resulting from the lipid load are poorly converted. At the output of said process, the majority of products found are therefore normal paraffins having an identical or lower number of carbon atoms of only one unit compared to the starting carboxylic acid chain present in the triglycerides. Currently, the most abundant lipid feedstocks for biofuel production are rapeseed, soybean, palm oils and these derivatives, and to a lesser extent used cooking oils and animal fats. These renewable fillers mainly contain triglycerides with fatty chains of 16 and 18 carbon atoms. Thus, the majority of products obtained after hydrotreatment will be normal paraffins containing mainly 15 to 18 carbon atoms. Consequently, the products of these biological loads will be found mainly in the diesel cut, that is to say in the cut having initial and final boiling points within a range of 220 to 390°C.

In this case, obtaining a bio-kerosene cut requires an additional step of selective hydrocracking.

Patent US8039682 describes a process for producing kerosene from a renewable feedstock, possibly in co-processing with a fossil feedstock, comprising a step of hydrotreatment, hydrodeoxygenation, isomerization and selective hydrocracking in the presence of 'a multifunctional catalyst or a series of catalysts, a gas/liquid separation step of the effluent obtained followed by a step of separation of the paraffinic effluent to produce a jet effluent and a light naptha effluent and a more effluent heavier than the jet, then recycling the effluent heavier than the jet in the reaction zone with a recycle rate / volume charge between 0.1 to 8.

Finally, if we had a catalytic system capable of selectively hydroisomerizing and/or hydrocracking compounds from a hydrocarbon gas oil cut to the lighter kerosene cut, it would then be possible to produce large quantities of (Bio-) kerosene from a very wide range of charges without modifying existing processes. However, to our knowledge, and as highlighted by all the examples previously reported, existing knowledge does not allow this type of conversion to be carried out.

The applicant has discovered that the use of a process for producing middle distillates including a hydrocracking step from a fossil hydrocarbon feed mixed with a renewable feed, having previously been hydrotreated, of a sequence of specific catalysts without modification of the fossil feed hydrocracking process allows the production of high yield in kerosene and gas oil cutting including a high incorporation rate of biogenic carbon.

In particular, the implementation of a sequence, in a first catalytic zone, of a first hydrocracking catalyst comprising a zeolite with large pores, preferably having at least one opening of at least 12MR, and metals of the group VIB and Group VIII according to the periodic classification, then, without any separation being made between the first and second catalytic zones, the implementation of a second catalyst, in a second catalytic zone, comprising at least one metal from group VIB and/or at least one metal from group VIII of the periodic table and a support comprising at least one zeolite with structural code BEA, allows:

- to improve the selectivity towards the kerosene cut while simultaneously leading to cold properties of said gas oil and kerosene cuts improved compared to the use of any other catalyst or sequence of catalysts known to those skilled in the art, compared to that obtained in a hydrocracking process using only said first catalyst if the latter was used throughout the hydrocracking stage, i.e. in 100% of the volume devoted to the hydrocracking stage.

The subject of the present invention is a process for producing middle distillates from at least one fossil hydrocarbon feedstock of which at least 50% by weight of the compounds have an initial boiling point greater than 250°C and a final boiling point less than 800°C, said fossil hydrocarbon feedstock being co-treated in a mixture with at least one renewable feedstock chosen from vegetable oils, algal oils, cooking oils and animal fats, fresh or used, alone or in a mixture, and feedstocks from the reprocessing of biomass/plastics/tires/and household waste, alone or in mixtures, said process comprising a step of hydrotreating said feed mixture in the presence of hydrogen and at least one hydrotreatment catalyst and a step of hydrocracking at least part and preferably all of the hydrotreated feedstock, the hydrocracking step being carried out at a temperature between 200°C and 480°C, at a total pressure between 1 MPa and 25 MPa with a ratio volume of hydrogen per volume of hydrocarbon feed of between 80 and 5000 liters per liter and at an Hourly Volume Velocity (VVH) defined by the ratio of the volume flow of liquid hydrocarbon feed to the volume of catalyst loaded in the reactor between 0.1 and 50 h ^-1 , the hydrocracking step of the hydrotreated feed comprising at least:

a) A step of bringing into contact at least part and preferably all of said hydrotreated feed in a first catalytic zone with at least one first catalyst comprising at least one metal from group VIB and/or at least one metal from group VIB group VIII of the periodic table and a support comprising at least one zeolite having at least one series of channels whose opening is defined by a ring with 12 oxygen atoms (12MR), and at least one binder.

b) Followed by bringing into contact in a second catalytic zone all of the effluent from step a), without an intermediate separation step between said first catalytic zone and said second catalytic zone, with a second catalyst comprising at at least one metal from group VIB and/or at least one metal from group VIII of the periodic table and a support comprising at least one zeolite with structural code BEA and at least one binder.

An advantage of the present invention is to provide a process for producing middle distillates using in the hydrocracking step a sequence of specific catalysts making it possible to obtain a high yield of kerosene, as well as cold properties of the kerosene and diesel cuts improved compared to those obtained in the prior art.

In particular, an advantage of the present invention is to provide a hydrocracking process making it possible to mainly improve the selectivity in the kerosene cut, while also making it possible to obtain improved cold properties of the kerosene and gas oil cuts compared to the use of catalysts or sequence of catalysts of the prior art.

In the sense of the present invention, the different embodiments presented can be used alone or in combination with each other, without limitation of combination.

In the sense of the present invention, the different parameter ranges for a given step such as the pressure ranges and the temperature ranges can be used alone or in combination. For example, within the meaning of the present invention, a preferred range of pressure values can be combined with a more preferred range of temperature values.

In the remainder of the text, the groups of chemical elements are given according to the CAS classification (CRC Handbook of Chemistry and Physics, publisher CRC press, editor-in-chief D.R. Lide, 81st edition, 2000-2001). For example, group VIII according to the CAS classification corresponds to the metals of columns 8, 9 and 10 according to the new IUPAC classification, and group VIB to the metals of column 6.

In the rest of the text, the expressions “between… and…” and “between…. and…” are equivalent and mean that the limit values of the interval are included in the range of values described. If this were not the case and the limit values were not included in the range described, such precision will be provided by the present invention.

In this description, the expression “greater than…” is understood as strictly greater, and symbolized by the sign “>”, and the expression “less than” as strictly inferior, and symbolized by the sign “<”.

In the present description, the overall SiO2/Al2O3 molar ratio of a zeolite is also called SAR or silica-alumina ratio according to Anglo-Saxon terminology. The SiO2/Al2O3 molar ratio is measured by X-ray fluorescence.

The invention relates to a process for producing middle distillates from at least one fossil hydrocarbon feedstock of which at least 50% by weight of the compounds have an initial boiling point greater than 250°C and a final boiling point less than 800°C mixed with at least one renewable load chosen from vegetable oils, algal oils and animal fats, fresh or used, alone or in a mixture, and loads from the reprocessing of plastics/tires/and household waste, alone or in mixtures.

Charges

A wide variety of hydrocarbon feedstocks can be treated by the hydrocracking processes according to the invention. The fossil hydrocarbon feedstock used in the hydrocracking process according to the invention is a hydrocarbon feedstock of which at least 50% by weight of the compounds have an initial boiling point greater than 250°C and a final boiling point less than 800°C, preferably of which at least 60% by weight, preferably of which at least 75% by weight and more preferably of which at least 80% by weight of the compounds have an initial boiling point greater than 250°C and a point final boiling point below 800°C.

The fossil hydrocarbon feedstock is advantageously chosen from gas oils resulting from a catalytic cracking unit or LCO (Light Cycle Oil) according to Anglo-Saxon terminology, atmospheric distillates, vacuum distillates resulting from the direct distillation of crude or conversion units such as FCC, coker, H-Oil or visbreaking, charges coming from units for extracting aromatics from lubricating oil bases or from solvent dewaxing of lubricating oil bases, distillates from fixed bed or ebullated bed desulfurization or hydroconversion processes of RAT (atmospheric residues) and/or RSV (vacuum residues) and/or deasphalted oils. The list above is not exhaustive. Said fillers preferably have a boiling point T5 greater than 200°C, preferably greater than 340°C, that is to say that 95% by weight of the compounds present in the filler have a boiling point greater than 250°C. C, and preferably greater than 340°C.

In accordance with the invention, the fossil hydrocarbon feedstock is treated in the process according to the invention in mixture with at least one renewable feedstock chosen from vegetable oils, cooking oils, algal oils and animal fats, fresh or used, alone. or in a mixture, and loads from the reprocessing of biomass/plastics/tires/and household waste, alone or in mixtures.

Vegetable oils can advantageously be crude or refined, totally or in part, and derived from the following plants: rapeseed, sunflower, soya, palm, palm kernel, olive, coconut, jatropha, this list not being exhaustive. Algal or fish oils are also relevant. Animal fats are advantageously chosen from bacon or fats composed of residues from the food industry or from the catering industries.

These fillers essentially contain chemical structures of the triglyceride type which those skilled in the art also know under the name fatty acid tri ester as well as free fatty acids. A fatty acid tri ester is thus composed of three chains of fatty acids. These fatty acid chains in the form of tri esters or in the form of free fatty acids have a number of unsaturations per chain, also called the number of carbon-carbon double bonds per chain, generally between 0 and 3 but which can be higher especially for oils from algae which generally have a number of unsaturations per chain of 5 to 6.

More generally, the process according to the invention is capable of treating the aforementioned fillers alone or in a mixture whatever the proportions.

Preferably, the content of renewable filler in the filler mixture is between 0.5 and 40% by weight relative to the total mass of the filler treated in the process according to the invention, and preferably between 1 and 35%. by weight, preferably between 2 and 30% by weight and more preferably between 5 and 25% by weight.

These fillers are likely to possibly contain metals, and/or nitrogen, and/or oxygen and/or sulfur. Where appropriate, the process according to the invention may advantageously comprise or not a liquid/gas separation step between the hydrotreatment step of said feedstock and the hydrocracking step.

The nitrogen content of the feeds treated in the processes according to the invention is generally greater than 500 ppm by weight, preferably between 500 and 10000 ppm by weight, more preferably between 700 and 4000 ppm by weight and even more preferably between 1000 and 4000 ppm weight. The sulfur content of the charges treated in the processes according to the invention is advantageously between 0.01 and 5% by weight, preferably between 0.2 and 4% by weight and even more preferably between 0.5 and 3% by weight. % weight.

The filler may possibly contain metals. The cumulative nickel and vanadium content of the charges treated in the processes according to the invention is preferably less than 5 ppm, preferably less than 3 ppm and preferably less than 1 ppm by weight.

The filler may possibly contain asphaltenes. The asphaltene content is generally less than 3000 ppm by weight, preferably less than 1000 ppm by weight, even more preferably less than 300 ppm by weight.

Due to the presence of renewable feedstocks from biomass, said feedstock treated in the process according to the invention may also contain oxygenated compounds. Their content varies greatly depending on the load that one wishes to transform in the process according to the invention. It is generally less than 50% by weight, and preferably is less than 15%, very preferably less than 5% by weight, even more preferably less than 2% by weight.

Method of realization

According to the invention, the process comprises a step of hydrotreating said charges in the presence of hydrogen and at least one hydrotreatment catalyst, said hydrotreatment step preferably operating at a temperature between 200 and 450°C. , under a pressure between 2 and 25 MPa, at a space speed between 0.1 and 6 h ^-1 and at a quantity of hydrogen introduced such that the volume ratio liter of hydrogen/liter of hydrocarbon is between 100 and 5000 NL/L.

The operating conditions such as temperature, pressure, hydrogen recycling rate, hourly space speed, may vary greatly depending on the nature of the load, the quality of the desired products and the facilities available to the refiner.

Preferably, the hydrotreatment step according to the invention operates at a temperature between 250 and 450°C, very preferably between 300 and 430°C, under a pressure between 5 and 20 MPa, at a spatial speed. between 0.2 and 5 h-1, and at a quantity of hydrogen introduced such that the volume ratio liter of hydrogen/liter of hydrocarbon is between 300 and 3000 NL/L.

Conventional hydrotreatment catalysts can advantageously be used in supported or unsupported form, preferably which contain at least one amorphous support and at least one hydro-dehydrogenating element chosen from at least one element of non-noble groups VIB and VIII, and the more often at least one element from group VIB (among molybdenum or tungsten taken alone or in a mixture) and at least one element from non-noble group VIII (among nickel or cobalt taken alone or in a mixture).

Preferably, the amorphous support is alumina or silica-alumina.

Preferred catalysts are chosen from NiMo, NiW, NiMoW or CoMo catalysts supported on alumina and NiMo, NiW or NiMoW supported on silica-alumina.

The effluent resulting from the hydrotreatment step and part of which and preferably all of which enters the hydrocracking step a) generally comprises a nitrogen content preferably less than 300 ppm by weight, less than 200 ppm by weight, preferably less than 100 ppm by weight and preferably less than 50 ppm by weight.

The process according to the invention may advantageously comprise a separation step between the hydrotreatment step and the hydrocracking step.

In one embodiment, the process according to the invention does not include a separation step between the hydrotreatment step and the hydrocracking step.

According to the invention, the process comprises a step of hydrocracking at least a part and preferably all of the hydrotreated feedstock, the hydrocracking step being carried out at a temperature between 200°C and 480°C , at a total pressure between 1 MPa and 25 MPa with a ratio of volume of hydrogen per volume of hydrocarbon feed of between 80 and 5000 liters per liter and at an Hourly Volume Velocity (VVH) defined by the ratio of the volume flow rate of charge liquid hydrocarbon by the volume of catalyst loaded into the reactor between 0.1 and 50 h-1.

Preferably, the hydrocracking step of the process according to the invention is carried out in the presence of hydrogen, between 250 and 480°C, preferably between 320 and 450°C, very preferably between 330 and 435°C. C, under a pressure of between 2 and 25 MPa, preferably between 3 and 20 MPa, at a space speed of between 0.1 and 20 h ^-1 , preferably 0.1 and 6 h ^-1 , preferably between 0.2 and 3 h ^-1 , and the quantity of hydrogen introduced is such that the volume ratio liter of hydrogen/liter of hydrocarbon is between 100 and 3000 L/L.

These operating conditions used in the processes according to the invention generally make it possible to achieve conversions per pass, in products having boiling points lower than 340°C, and better still lower than 370°C, greater than 15% by weight and even more preferably between 20 and 95% by weight.

In addition, the kerosene cut produced by the process according to the invention has a crystal disappearance point lower than -20°C, preferably lower than -30°C and even more preferably lower than -47°C. The gas oil cut produced by the process has a cloud point lower than 15°C, preferably lower than 5°C and very preferably lower than -5°C.

In accordance with the invention, the step of hydrocracking at least part and preferably all of the hydrotreated feed comprises at least:

a) a step of bringing said hydrotreated feedstock into contact in a first catalytic zone with at least one first catalyst comprising at least one metal from group VIB and/or at least one metal from group VIII of the periodic table and a support comprising at least one at least one zeolite having at least one series of channels whose opening is defined by a ring with 12 oxygen atoms (12MR), and at least one binder.

b) followed by bringing into contact in a second catalytic zone all of the effluent from step a), without intermediate separation step between said first and second catalytic zones, with a second catalyst comprising at least one metal from group VIB and/or at least one metal from group VIII of the periodic table and a support comprising at least one zeolite with structural code BEA and at least one binder.

Preferably, the first catalyst used in the first catalytic zone of hydrocracking step a) comprising at least one metal from group VIB and at least one metal from group VIII of the periodic table, taken alone or as a mixture, is preferably a sulfide phase catalyst.

Preferably, the metals of group VIB of the periodic table are chosen from the group formed by tungsten and molybdenum, taken alone or as a mixture. According to a preferred embodiment, the group VIB metal is molybdenum. According to another preferred embodiment, the group VIB metal is tungsten.

Preferably, the non-noble metals from group VIII of the periodic table are chosen from the group formed by cobalt and nickel, taken alone or as a mixture. According to a preferred embodiment, the non-noble Group VIII metal is cobalt. According to another preferred embodiment, the non-noble Group VIII metal is nickel.

Preferably, said catalyst comprises at least one metal from group VIB in combination with at least one non-noble metal from group VIII, the non-noble metals from group VIII being chosen from the group formed by cobalt and nickel, taken alone or in combination. mixture and the metals of group VIB being chosen from the group formed by tungsten and molybdenum, taken alone or as a mixture.

In one embodiment, the following combinations of metals are used: nickel-molybdenum, cobalt-molybdenum, nickel-tungsten, cobalt-tungsten, and even more advantageously nickel-molybdenum and nickel-tungsten.

In the case where the catalyst comprises at least one metal from Group VIB in combination with at least one non-noble metal from Group VIII, the metal content from Group VIB is advantageously comprised, in oxide equivalent, between 5 and 40% by weight per relative to the total mass of said catalyst, preferably between 10 and 35% by weight and very preferably between 15 and 30% by weight and the content of non-noble metal from group VIII is advantageously included, in oxide equivalent, between 0 .5 and 10% by weight relative to the total mass of said catalyst, preferably between 1 and 8% by weight and very preferably between 1.5 and 6% by weight.

In the case where the catalyst comprises at least one metal from group VIB in combination with at least one non-noble metal from group VIII, said catalyst is a sulfide catalyst.

It is also possible to use combinations of three metals, for example nickel-cobalt-molybdenum, nickel-molybdenum-tungsten, nickel-cobalt-tungsten.

In one embodiment, the following metal combinations are used: nickel-niobium-molybdenum, cobalt-niobium-molybdenum, nickel-niobium-tungsten, cobalt-niobium-tungsten, the preferred combinations being: nickel-niobium-molybdenum, cobalt -niobium-molybdenum. It is also possible to use combinations of four metals, for example nickel-cobalt-niobium-molybdenum.

The catalyst may advantageously also comprise at least one doping element chosen from the group consisting of boron and phosphorus, and preferably phosphorus.

The first catalyst can also advantageously comprise:

- from 0.1 to 15% by weight of oxide, preferably from 0.1 to 10% by weight relative to the total mass of the catalyst of at least one doping element chosen from the group consisting of boron and phosphorus, and preferably phosphorus,

- from 0 to 60% by weight, preferably from 0.1 to 50% by weight, and even more preferably from 0.1 to 40% by weight of oxide relative to the total mass of the catalyst, at least one element chosen from group VB and preferably niobium,

- from 0 to 20% by weight, preferably from 0.1 to 15% by weight and even more preferably from 0.1 to 10% by weight of oxide relative to the total mass of the catalyst of at least an element chosen from group VIIA, preferably fluorine.

In accordance with the invention, the support of the first catalyst used in the first catalytic zone of hydrocracking step a) according to the invention comprises at least one zeolite having at least one series of channels whose opening is defined by a ring with 12 oxygen atoms (12MR), and at least one binder.

The binder used in the support of the first catalyst, also called porous mineral matrix, advantageously consists of at least one refractory oxide, preferably chosen from the group formed by alumina, silica-alumina, clay, titanium oxide, boron oxide and zirconia, taken alone or in mixture. Preferably, the binder is chosen from alumina and silica-alumina, taken alone or as a mixture. More preferably, the binder is alumina. Alumina can advantageously be in all its forms known to those skilled in the art. Very preferably, the alumina is gamma alumina, for example boehmite.

Preferably, said support of said first catalyst comprises from 20 to 99% by weight of binder, preferably from 30% to 99% by weight, and very preferably between 50% and 95% by weight, and very preferably from 60 to 95% by weight relative to the total weight of said support.

The zeolite used in the support of the first catalyst is advantageously chosen from zeolites of structural type FAU, BEA, ISV, IWR, IWW, MEI, UWY, taken alone or as a mixture and preferably from zeolites of structural type FAU and BEA, taken alone or in a mixture.

In a preferred embodiment, the zeolite is chosen from zeolite Y and zeolite beta taken alone or as a mixture and preferably the zeolite is zeolite Y and very preferably dealuminated zeolite USY.

Preferably, said support of said first catalyst comprises from 1 to 80% by weight, preferably from 1 to 70% by weight, even more preferably from 5 to 50% by weight, and very preferably 5 to 40% by weight. weight of at least one zeolite relative to the total mass of said support.

Preferably, said support of said first catalyst comprises and preferably consists of:

- 1 to 80% by weight, preferably 1 to 70% by weight, even more preferably 5 to 50% by weight, and very preferably 5 to 40% by weight of at least one zeolite relative to to the total mass of said support,

- 20 to 99% by weight, preferably from 30 to 99%, preferably from 50 to 95% by weight, and very preferably from 60 to 95% by weight relative to the total mass of said support, of at least minus said binder.

In a preferred embodiment, the first catalyst comprises a Y zeolite alone.

In another preferred embodiment, the first catalyst comprises a USY zeolite and a beta zeolite.

Said zeolites are advantageously defined in the classification “Atlas of Zeolite Framework Types, 6th revised edition”, Ch. Baerlocher, L. B. McCusker, D.H. Olson, 6th Edition, Elsevier, 2007, Elsevier".

The zeolite Y used in the support of the first catalyst advantageously has an initial crystal parameter a0 of the unit cell less than 24.55 Å, preferably less than 24.45 Å and preferably less than 24.40 Å and even more more preferred less than 24.35 Å. Preferably, the initial crystal parameter a0 of the unit cell is greater than 24.24 Å and preferably greater than 24.26 Å.

The zeolite Y used in the support of the first catalyst advantageously has a specific surface area measured by nitrogen physisorption according to the BET method of between 550 and 1200 m ² /g, preferably between 600 and 1100 m ² /g, and preferably between 650 and 1050 m ² /g.

According to a preferred embodiment of the invention, the zeolite Y suitable for the implementation of the catalyst support a) used in the process according to the invention is advantageously prepared from a zeolite Y of structural type FAU having preferably an overall Si/Al atomic ratio after synthesis of between 2.3 and 2.8 and advantageously presenting itself in NaY form after synthesis. Said zeolite Y of structural type FAU advantageously undergoes a step of one or more ionic exchanges before undergoing the dealumination step. The ion exchange(s) make it possible to partially or totally replace the alkaline cations belonging to groups IA and IIA of the periodic classification present in the cationic position in the zeolite Y of structural type FAU as raw from synthesis by NH4+ cations and preferably cations Na+ by NH4+ cations.

By partial or total exchange of alkaline cations with NH4+ cations is meant the exchange of 80 to 100%, preferably 85 to 99.5% and more preferably 88 to 99%, of said alkaline cations with NH4+ cations. At the end of the ion exchange step(s), the remaining quantity of alkaline cations, and preferably the remaining quantity of Na+ cations, in the zeolite Y, relative to the quantity of alkaline cations, preferably Na+, initially present in zeolite Y, is advantageously between 0 and 20%, preferably between 0.5 and 15% and preferably between 1.0 and 12%.

Preferably, this step implements several ion exchanges with a solution containing at least one ammonium salt chosen from chlorate, sulfate, nitrate, phosphate, or ammonium acetate salts, so as to eliminate at least in part, alkaline cations and preferably Na+ cations present in the zeolite. Preferably, the ammonium salt is ammonium nitrate NH4NO3.

Thus, the remaining content of alkaline cations and preferably of Na+ cations in the zeolite Y at the end of the ion exchange(s) step is preferably such that the alkaline cation/aluminum molar ratio and preferably the Na/Al molar ratio is between 0:1 and 0:1, preferably between 0:1 and 0.005:1, and more preferably between 0:1 and 0.008:1.

The desired alkaline cation/aluminum ratio, preferably Na/Al, is obtained by adjusting the NH4+ concentration of the ion exchange solution, the ion exchange temperature and the number of ion exchanges. The concentration of the NH4+ ion exchange solution advantageously varies between 0.01 and 12 mol.L-1, and preferably between 1.00 and 10 mol.L-1. The temperature of the ion exchange step is advantageously between 20 and 100°C, preferably between 60 and 95°C, preferably between 60 and 90°C, more preferably between 60 and 85°C and even more preferably between 60 and 80°C. The number of ion exchanges advantageously varies between 1 and 10 and preferably between 1 and 4.

If necessary, said zeolite Y, preferably of structural type FAU, obtained can then undergo a dealumination treatment step. Said dealumination step can advantageously be carried out by all the methods known to those skilled in the art. Preferably, the dealumination is carried out by a heat treatment optionally in the presence of water vapor (or steaming according to Anglo-Saxon terminology) and/or by one or more acid attacks advantageously carried out by treatment with an aqueous acid solution. mineral or organic.

Preferably, the dealumination step implements a heat treatment followed by one or more acid attacks, or only one or more acid attacks.

Preferably, the heat treatment optionally in the presence of water vapor to which said zeolite Y is subjected is carried out at a temperature between 200 and 900°C, preferably between 300 and 900°C, even more preferably between 400 and 900°C. 750°C. The duration of said heat treatment is advantageously greater than or equal to 0.5 h, preferably between 0.5 h and 24 h, and very preferably between 1 h and 12 h. In the case where the heat treatment is carried out in the presence of water, the volume percentage of water vapor during the heat treatment is advantageously between 5 and 100%, preferably between 20 and 100%, very preferably between 40 and 100%. The volume fraction other than the water vapor possibly present is air. The gas flow rate formed of water vapor and possibly air is advantageously between 0.2 L.h-1.g-1 and 10 L.h-1.g-1 of zeolite Y.

The heat treatment makes it possible to extract the aluminum atoms from the framework of the zeolite Y while maintaining the overall Si/Al atomic ratio of the treated zeolite unchanged.

The heat treatment step in the presence of water vapor can advantageously be repeated as many times as is necessary to obtain the zeolite Y suitable for the implementation of the catalyst support used in the process according to the invention and having a crystal parameter a0 of the unit cell less than 24.55 Å, preferably less than 24.45 Å and preferably less than 24.40 Å and even more preferably less than 24.35 Å.

The heat treatment step possibly in the presence of water vapor is advantageously followed by an acid attack step. Said acid attack makes it possible to eliminate partially or entirely the aluminum debris resulting from the heat treatment step in the presence of water vapor and which partially blocks the porosity of the dealuminated zeolite; the acid attack therefore makes it possible to unclog the porosity of the dealuminated zeolite.

The acid attack can advantageously be carried out by suspending the zeolite Y, which has optionally previously undergone heat treatment, in an aqueous solution containing a mineral or organic acid. The mineral acid can be nitric acid, sulfuric acid, hydrochloric acid, phosphoric acid or boric acid. The organic acid may be formic acid, acetic acid, oxalic acid, tartaric acid, maleic acid, malonic acid, malic acid, lactic acid, or any other acid. organic soluble in water. The concentration of the solution of mineral or organic acid in the solution advantageously varies between 0.01 and 2.0 mol.L-1, and preferably between 0.5 and 1.0 mol.L-1. The temperature of the acid attack step is advantageously between 20 and 100°C, preferably between 60 and 95°C, preferably between 60 and 90°C and more preferably between 60 and 80°C. The duration of the acid attack is advantageously between 5 minutes and 8 hours, preferably between 30 minutes and 4 hours, and preferably between 1 hour and 2 hours.

At the end of the heat treatment step(s) optionally in the presence of water vapor and possibly the acid attack step, the process of modifying said zeolite Y advantageously comprises a step of at least one partial exchange or total of the alkaline cations and preferably the Na+ cations still present in the cationic position in the zeolite Y. The ion exchange step is carried out in a manner similar to the ion exchange step described above.

At the end of the heat treatment step(s) optionally in the presence of water vapor and optionally of the acid attack step and optionally of the step of partial or total exchange of alkaline cations and preferably cations Na+, the process for modifying said zeolite Y may include a calcination step. Said calcination makes it possible to eliminate the organic species present within the porosity of the zeolite, for example those provided by the acid attack step or by the partial or total exchange step of the alkaline cations. In addition, said calcination step makes it possible to generate the protonated form of zeolite Y and to give it acidity for its applications.

The calcination can advantageously be carried out in a muffle furnace or in a tubular furnace, under dry air or under an inert atmosphere, in a licked bed or in a crossed bed. The calcination temperature is advantageously between 200 and 800°C, preferably between 450 and 600°C, and preferably between 500 and 550°C. The duration of the calcination stage is advantageously between 1 and 20 hours, preferably between 6 and 15 hours, and preferably between 8 and 12 hours.

Thus, said zeolite Y obtained has an initial crystal parameter a0 of the unit cell less than 24.55 Å, preferably less than 24.45 Å and preferably less than 24.40 Å and even more preferably less than 24 , 35 Å and a specific surface area measured by nitrogen physisorption according to the BET method of between 550 and 1200 m ² /g, preferably between 600 and 1100 m ² /g, and preferably between 650 and 1050 m ² /g .

In the preferred embodiment where the support of the first catalyst also comprises, in addition to the USY zeolite, a Beta zeolite, the Beta zeolite preferably has an overall SiO2/Al2O3 or SAR molar ratio of between 10 and 100, preferably between 20 and 50, and preferably between 20 and 30. The Beta zeolite used in the support of the first catalyst according to the invention advantageously has a specific surface area measured by nitrogen physisorption according to the B.E.T. method. between 400 and 800 m2/g, preferably between 500 and 750 m2/g, and preferably between 550 and 700 m2/g.

Beta zeolite is generally synthesized from a reaction mixture containing a structuring agent. The use of structuring agents is well known to those skilled in the art: for example, US patent 3,308,069 describes the use of tetraethylammonium hydroxide, and US patent 5,139,759 describes the use of tetraethylammonium cation derived from a tetraethylammonium halide compound. Another standard method for preparing Beta zeolite is given in the book Verified Synthesis of Zeolitic Materials.

In said preferred embodiment where the support of the first catalyst comprises a USY zeolite and a Beta zeolite, said first catalyst comprises and preferably consists of:

- 0.9 to 79.9%, preferably 1 to 70%, and preferably 5 to 50%, and very preferably 5 to 40% by weight relative to the total weight of said support of a zeolite Y having an initial crystal parameter a0 of the unit cell less than 24.55 Å;

- 0.1 to 6%, preferably 0.2 to 5%, and preferably 0.5 to 5% weight by weight relative to the total weight of said support of a Beta zeolite; And

- from 20 to 99% by weight, preferably from 30% to 99% by weight, and very preferably from 50% to 95% by weight, and even more preferably from 60 to 95% by weight relative to the weight total of said support of at least one binder.

In accordance with step b) of the process according to the invention, all of the effluent from step a) is brought into contact in a second catalytic zone, without an intermediate separation step between the first and the second catalytic zone. , with a second catalyst comprising at least one metal from group VIB and/or at least one metal from group VIII of the periodic table and a support comprising at least one zeolite with structural code BEA and at least one binder.

Preferably, the effluent from step a) is sent to a second catalytic zone without intermediate gas/liquid separation step.

Preferably, the second catalyst used in the second catalytic zone of hydrocracking step b) comprising at least one metal from group VIB and at least one metal from group VIII of the periodic table, taken alone or as a mixture, is preferably a catalyst in sulfide form.

Preferably, the metals of group VIB of the periodic table are chosen from the group formed by tungsten and molybdenum, taken alone or as a mixture. According to a preferred embodiment, the group VIB metal is molybdenum. According to another preferred embodiment, the group VIB metal is tungsten.

Preferably, the non-noble metals from group VIII of the periodic table are chosen from the group formed by cobalt and nickel, taken alone or as a mixture. According to a preferred embodiment, the non-noble Group VIII metal is cobalt. According to another preferred embodiment, the non-noble Group VIII metal is nickel.

Preferably, said second catalyst comprises at least one metal from group VIB in combination with at least one non-noble metal from group VIII, the non-noble metals from group VIII being chosen from the group formed by cobalt and nickel, taken alone or in mixture and the metals of group VIB being chosen from the group formed by tungsten and molybdenum, taken alone or in mixture.

In one embodiment, the following metal combinations are used: nickel-molybdenum, cobalt-molybdenum, nickel-tungsten, cobalt-tungsten, the preferred combinations are: nickel-molybdenum, cobalt-molybdenum, cobalt-tungsten, nickel-tungsten and even more advantageously nickel-molybdenum and nickel-tungsten.

In the case where the second catalyst comprises at least one metal from Group VIB in combination with at least one non-noble metal from Group VIII, the metal content from Group VIB is advantageously comprised, in oxide equivalent, between 5 and 40% by weight. relative to the total mass of said catalyst, preferably between 10 and 35% by weight and very preferably between 15 and 30% by weight and the content of non-noble metal from group VIII is advantageously included, in oxide equivalent, between 0.5 and 10% by weight relative to the total mass of said catalyst, preferably between 1 and 8% by weight and very preferably between 1.5 and 6% by weight.

In the case where the second catalyst comprises at least one metal from group VIB in combination with at least one non-noble metal from group VIII, said catalyst is a sulfide catalyst.

It is also possible to use combinations of three metals, for example nickel-cobalt-molybdenum, nickel-molybdenum-tungsten, nickel-cobalt-tungsten.

In one embodiment, the following metal combinations are used: nickel-niobium-molybdenum, cobalt-niobium-molybdenum, nickel-niobium-tungsten, cobalt-niobium-tungsten, the preferred combinations being: nickel-niobium-molybdenum, cobalt -niobium-molybdenum. It is also possible to use combinations of four metals, for example nickel-cobalt-niobium-molybdenum.

The catalyst may advantageously also comprise at least one doping element chosen from the group consisting of boron and phosphorus, and preferably phosphorus.

The second catalyst can also advantageously contain:

- from 0.1 to 15% by weight of oxide, preferably from 0.1 to 10% by weight relative to the total mass of the catalyst of at least one doping element chosen from the group consisting of boron and phosphorus, and preferably phosphorus,

- from 0 to 60% by weight, preferably from 0.1 to 50% by weight, and even more preferably from 0.1 to 40% by weight of oxide relative to the total mass of the catalyst, at least one element chosen from group VB and preferably niobium,

- from 0 to 20% by weight, preferably from 0.1 to 15% by weight and even more preferably from 0.1 to 10% by weight of oxide relative to the total mass of the catalyst of at least an element chosen from group VIIA, preferably fluorine.

In accordance with the invention, the support of the second catalyst used in the second catalytic zone of hydrocracking step b) according to the invention comprises at least one zeolite with structural code BEA and at least one binder.

Preferably, the zeolite with structural code BEA is a Beta zeolite and preferably a beta zeolite having an overall silica to alumina SAR atomic ratio of between 10 and 300, preferably between 10 and 100, preferably between 10 and 50, preferably between 10 and 30, more preferably between 10 and 25, even more preferably between 12 and 24 and even more preferably between 15 and 24.

Beta zeolite is generally synthesized from a reaction mixture containing a structuring agent. The use of structuring agents is well known to those skilled in the art: for example, US patent 3,308,069 describes the use of tetraethylammonium hydroxide, and US patent 5,139,759 describes the use of tetraethylammonium cation derived from a tetraethylammonium halide compound. Another standard method for preparing Beta zeolite is indicated in the work Verified Synthesis of Zeolitic Materials (Elsevier, 2001).

The Beta zeolite used in the support of the second catalyst according to the invention advantageously has an average crystal size of less than 500 nm, preferably less than 350 nm. Advantageously, these crystals can be in the form of agglomerates of size up to 50 μm, preferably less than 25 μm. On the other hand, the Beta zeolite used in the support of catalyst a) according to the invention has a specific surface area measured by nitrogen physisorption according to the BET method of between 500 and 800 m ² /g, preferably between 500 and 750 m ² /g, and preferably between 550 and 720 m ² /g, a mesoporous volume greater than 0.2 ml/g, preferably greater than 0.40 ml/g determined by BJH and an external surface of at least 120 m ² /g, preferably at least 140 m ² /g, very preferably at least 150 m ² /g and even more preferably, between 150 and 300 m ² /g. Finally, preferably, the proportion of external surface area relative to the total BET surface area of the Beta zeolite included in the support of catalyst b) is between 15 and 50%, very preferably between 18 and 40% and even more preferred, between 21 and 37%.

The binder used in the support of the second catalyst, also called porous mineral matrix, advantageously consists of at least one refractory oxide, preferably chosen from the group formed by alumina, silica-alumina, clay, titanium oxide, boron oxide and zirconia, taken alone or in mixture. Preferably, the binder is chosen from alumina and silica-alumina, taken alone or as a mixture. More preferably, the binder is alumina. Alumina can advantageously be in all its forms known to those skilled in the art. Very preferably, the alumina is gamma alumina, for example boehmite.

Preferably, said support of said second catalyst comprises from 20 to 99% by weight of binder, preferably from 30% to 99% by weight, and very preferably between 50% and 95% by weight, and very preferably from 60 to 95% by weight relative to the total weight of said support.

Preferably, said support of the second catalyst comprises and preferably consists of:

- 1 to 70% by weight, preferably 1 to 50% by weight, preferably 1.5 to 40% by weight, and even more preferably 2 to 30% by weight of said beta zeolite relative to the total mass of said support,

- 30 to 99% by weight, preferably from 50 to 99%, preferably from 60 to 98% by weight, and very preferably from 70 to 98% by weight relative to the total mass of said support, of at least minus said binder.

The second catalyst used in step b) of the process according to the invention may advantageously contain at least one other zeolite in the support of said second catalyst. Where appropriate, the other zeolite(s) are advantageously chosen from zeolites having at least one series of channels whose opening is defined by a ring with 12 oxygen atoms (12MR), in particular from zeolites of structural type FAU, BEA, ISV, IWR, IWW, MEI, UWY, MFI, MTW or even zeolite IZM-2, taken alone or in a mixture and preferably among zeolites of structural type FAU and BEA, taken alone or mixed. Preferably, the second catalyst may comprise, in addition to the zeolite with structural code BEA, another zeolite with structural code FAU and preferably a zeolite Y.

Preferably, the zeolite Y which can be used in the second catalyst has an initial crystal parameter a0 of the unit cell less than 24.55 Å, preferably less than 24.45 Å and preferably less than 24.40 Å and even more preferably less than 24.35 Å and a specific surface area measured by nitrogen physisorption according to the B.E.T. method. between 550 and 1200 m2/g, preferably between 600 and 1100 m2/g, and preferably between 650 and 1050 m2/g.

Preferably, said zeolite Y can be prepared according to a process identical to that described for zeolite Y used in step a) of the process according to the invention.

In this case, the support of the second catalyst used in the process according to the invention comprises at least one zeolite with structural code BEA and at least one zeolite Y, and at least one binder, said support comprising and preferably consisting of, preference :

- 1 to 50% by weight, preferably 1 to 30% by weight, preferably 1.5 to 20% by weight, and even more preferably 2 to 30% by weight of said beta zeolite relative to the total mass of said support,

- 0.5 to 50% by weight, preferably 1 to 30% by weight, preferably 1.5 to 10% by weight, and even more preferably 2 to 5% by weight of said zeolite Y per relative to the total mass of said support,

- 30 to 98.5% by weight, preferably 40 to 98%, preferably 70 to 97% by weight, and very preferably 65 to 96% by weight relative to the total mass of said support, d 'at least said binder.

In this case, the second catalyst used in step b) of the process according to the invention is preferably different from the first catalyst used in step a) of said process.

In the hydrocracking step, the volume fraction of the first catalyst used in step a) relative to all of the two catalysts used in steps a) and b) is preferably between 50 and 95% vol and preferably between 60 and 95% vol, very preferably between 65 and 90% vol and even more preferably between 65 and 85% vol.

Preferably, the volume fraction of the second catalyst used in step b) relative to all of the two catalysts used in steps a) and b) is between 5 and 50% vol and preferably between 5 and 40%. vol, very preferably, between 10 and 35% vol and even more preferably between 15 and 35% vol.

The process for producing middle distillates according to the invention can advantageously be implemented according to all the embodiments known in the prior art.

Said process can advantageously be implemented in one or two stages, in one or more reactors, in a fixed bed or in an ebullated bed. It can be carried out with or without recycling, after said hydrotreatment step aimed at eliminating sulfur, nitrogen or oxygen compounds possibly present in the feeds. In the case where recycling is carried out, this can be carried out at any location in the process, either at the hydrotreatment stage or in the first catalytic zone of hydrocracking stage a).

In one embodiment of the so-called 2-step hydrocracking step, the effluent from step b) of the hydrocracking step is sent to a fractionation column to recover at least one middle distillate cut and at least one heavy cut comprising the unconverted compounds and said heavy cut is advantageously sent either to a second hydrocracking step in the presence of hydrogen and a hydrocracking catalyst or to hydrotreatment step a).

Preparation of the first and second catalysts according to the invention

The first and second catalysts used respectively in steps a) and b) of the hydrocracking step of the process according to the invention are advantageously prepared according to the conventional methods used in the prior art.

In particular, the catalyst(s) is(are) prepared according to a preparation process comprising:

- a support preparation step comprising:

- the mixture of at least one porous mineral matrix with:

In the case of the first catalyst, at least one zeolite having at least one series of channels whose opening is defined by a ring with 12 oxygen atoms (12MR),

In the case of the second catalyst, at least one zeolite with structural code BEA with a SAR less than 25,

And,

- shaping said mixture;

- the introduction of at least one hydro-dehydrogenating element chosen from the group formed by the elements of group VIB of the periodic table, preferably molybdenum and tungsten, the elements of non-noble group VIII of the periodic table, preferably cobalt, nickel, and their mixtures, and preferably nickel and cobalt, and their mixtures, on the support by:

- addition of at least one precursor of said element during shaping so as to introduce at least part of said element, and/or

- impregnation of the support with at least one precursor of said element,

- possibly a drying and/or calcination step following the preparation of the support and/or the step of introducing at least one hydro-dehydrogenating element.

More particularly, the first catalyst is prepared according to a preparation process comprising the following steps:

- Preparation of zeolite having at least one series of channels whose opening is defined by a ring with 12 oxygen atoms (12MR) according to the methods known from the prior art, preferably, a Y zeolite,

- Mixed with a porous mineral matrix and shaped to obtain the support,

- Introduction of at least one hydro-dehydrogenating element onto the support by at least one of the following methods: (i) addition of at least one precursor of said element during shaping so as to introduce at least one part of said element, (ii) impregnation of the support with at least one precursor of said hydro-dehydrogenating element,

Possibly drying and/or calcination of the products obtained at the end of each of the aforementioned preparation stages.

Analogously, more particularly, the second catalyst is prepared according to a preparation process comprising the following steps:

- Preparation of zeolite with structural code BEA with a SAR less than 25 according to the processes known from the prior art,

- Mixed with a porous mineral matrix and shaped to obtain the support,

- Introduction of at least one hydro-dehydrogenating element onto the support by at least one of the following methods: (i) addition of at least one precursor of said element during shaping so as to introduce at least one part of said element, (ii) impregnation of the support with at least one precursor of said hydro-dehydrogenating element,

Possibly drying and/or calcination of the products obtained at the end of each of the aforementioned preparation stages.

Preferably, the porous mineral matrix comes from an alumina gel which has a crystallite size of between 2 to 35 nm.

Preferably, the alumina gel used comprises a sulfur content of between 0.001% and 1% by weight, preferably between 0.001 and 0.40% by weight, very preferably between 0.003 and 0.33% by weight, and more preferably between 0.005 and 0.25% by weight.

Preferably, the alumina gel used comprises a sodium content of between 0.001% and 1% by weight, preferably between 0.001 and 0.15% by weight, very preferably between 0.0015 and 0.10% by weight, and 0.002 and 0.040 wt.%.

The alumina gel used advantageously has a high dispersibility rate which makes it possible to facilitate the step of shaping said gel according to all the methods known to those skilled in the art and in particular by extrusion kneading, by granulation and by the technique known as oil drop according to Anglo-Saxon terminology.

Said alumina has a specific surface area and a porous distribution calibrated and adapted to its use in a hydrocracking process of said hydrocarbon feedstock.

Preferably, the alumina is purely mesoporous and devoid of micropores.

Preferably, the support advantageously has a specific surface area greater than 100 m ² /g, and a mesoporous volume greater than or equal to 0.5 ml/g, preferably greater than or equal to 0.6 ml/g.

The mesoporous volume of the support is defined as the volume included in the pores having an average diameter of between 2 and 50 nm and is measured using the mercury intrusion method.

Preferably, the alumina thus prepared and used in the invention is a non-mesostructured alumina.

The support can advantageously be shaped by any technique known to those skilled in the art. The shaping can be carried out for example by extrusion, by pelletizing, by the oil-drop coagulation method, by granulation on a turntable or by any other method well known to those skilled in the art.

The support is preferably shaped in the form of grains of different shapes and dimensions. They are generally used in the form of cylindrical or polylobed extrudates such as trilobes, quadrilobes or polylobes of straight or twisted shape, but can optionally be manufactured and used in the form of crushed powders, tablets, rings, balls, wheels. However, it is advantageous for the catalyst to be in the form of extrudates with a diameter of between 0.5 and 5 mm and more particularly between 0.7 and 3 mm and even more particularly between 1.0 and 2.5 mm. The shapes are cylindrical (which can be hollow or not), twisted cylindrical, multilobed (2, 3, 4 or 5 lobes for example), rings. Any other form can be used.

One of the preferred shaping methods consists of co-kneading said zeolites with the binder, preferably alumina, in the form of a wet gel for a few tens of minutes, preferably between 10 and 40 minutes, then passing the paste as well. obtained through a die to form extrudates with a diameter preferably between 0.5 and 5 mm.

According to another of the preferred shaping methods, said zeolites can be introduced during the synthesis of the porous mineral matrix. For example, according to this preferred mode of the present invention, said zeolites, preferably Y, and optionally Beta, are added during the synthesis of a porous mineral matrix, such as for example a silico-aluminum matrix: in this case , said zeolites can advantageously be added to a mixture composed of an alumina compound in an acid medium with a completely soluble silica compound.

In accordance with the invention, the raw material obtained at the end of the shaping step then undergoes a calcination step at a temperature of between 500 and 1000°C, for a period of between 2 and 10 hours, in presence or absence of air flow containing up to 60% volume of water.

Preferably, said calcination step operates at a temperature between 540°C and 850°C.

Preferably, said calcination step operates for a period of between 2 hours and 10 hours.

When the binder is an alumina gel otherwise identified as Boehmite, said calcination step allows the transition from boehmite to the final alumina.

The introduction of the elements of group VIB and/or VIII can optionally take place during the shaping step, by addition of at least one compound of said element, so as to introduce at least a part of said element.

The introduction of at least one hydro-dehydrogenating element can advantageously be accompanied by that of at least one promoter element chosen from phosphorus, boron, silicon and preferably phosphorus and optionally by the introduction of an element from the group VIIA and/or VB. The shaped solid is optionally dried at a temperature of between 60 and 250°C and optionally calcined at a temperature of 250 to 800°C for a period of between 30 minutes and 6 hours.

The step of introducing at least one hydro-dehydrogenating element is advantageously carried out by a method well known to those skilled in the art, in particular by one or more operations of impregnation of the shaped and calcined or dried support, and preferably calcined, with a solution containing the precursors of the elements of group VIB and/or VIII, optionally the precursor of at least one promoter element and optionally the precursor of at least one element of group VIIA and/or group VB .

Preferably, said introduction is carried out by a dry impregnation method with a solution containing the precursors of the hydro/dehydrogenating function, that is to say elements of group VIB and/or VIII, optionally followed by a drying step and preferably without a calcination step.

In the case where the catalyst of the present invention contains a non-noble metal from group VIII, the metals from group VIII are preferably introduced by one or more operations of impregnation of the shaped and calcined support, after those of group VIB or at the same time as these.

The introduction of at least one hydro-dehydrogenating element can then optionally be followed by drying at a temperature between 60 and 250°C and optionally by calcination at a temperature between 250 and 800°C.

The sources of molybdenum and tungsten are advantageously chosen from oxides and hydroxides, molybdic and tungstic acids and their salts, in particular ammonium salts such as ammonium molybdate, ammonium heptamolybdate, tungstate d ammonium, phosphomolybdic acid, phosphotungstic acid and their salts, silicomolybdic acid, silicotungstic acid and their salts. Ammonium oxides and salts such as ammonium molybdate, ammonium heptamolybdate and ammonium tungstate are preferably used.

The sources of non-noble group VIII elements which can be used are well known to those skilled in the art. For example, for non-noble metals we will use nitrates, sulfates, hydroxides, phosphates, halides such as chlorides, bromides and fluorides, carboxylates such as acetates and carbonates.

The preferred source of phosphorus is orthophosphoric acid H ₃ PO ₄ , but its salts and esters such as ammonium phosphates are also suitable. Phosphorus can for example be introduced in the form of a mixture of phosphoric acid and a basic organic compound containing nitrogen such as ammonia, primary and secondary amines, cyclic amines, compounds of the pyridine and quinoline family and compounds of the pyrrole family. Tungsto-phosphoric or tungsto-molybdic acids can be used.

The phosphorus content is adjusted, without limiting the scope of the invention, in such a way as to form a mixed compound in solution and/or on the support, for example tungsten-phosphorus or molybdenum-tungsten-phosphorus. These mixed compounds may be heteropolyanions. These compounds can be Anderson heteropolyanions, for example.

The source of boron can be boric acid, preferably orthoboric acid H ₃ BO ₃ , ammonium biborate or pentaborate, boron oxide, boric esters. Boron can for example be introduced in the form of a mixture of boric acid, hydrogen peroxide and a basic organic compound containing nitrogen such as ammonia, primary and secondary amines, cyclic amines, compounds from the pyridine and quinoline family and compounds from the pyrrole family. Boron can be introduced for example by a solution of boric acid in a water-alcohol mixture.

The sources of group VB elements which can be used are well known to those skilled in the art. For example, among the sources of niobium, oxides can be used, such as diniobium pentaoxide Nb ₂ O5, niobic acid Nb ₂ O ₅ .H ₂ O, niobium hydroxides and polyoxoniobates, niobium alkoxides of formula Nb(OR1) ₃ where R1 is an alkyl radical, niobium oxalate NbO(HC ₂ O ₄ ) ₅ , ammonium niobate. Niobium oxalate or ammonium niobate are preferably used.

The sources of group VIIA elements which can be used are well known to those skilled in the art. For example, fluoride anions can be introduced in the form of hydrofluoric acid or its salts. These salts are formed with alkali metals, ammonium or an organic compound. In the latter case, the salt is advantageously formed in the reaction mixture by reaction between the organic compound and the hydrofluoric acid. It is also possible to use hydrolyzable compounds that can release fluoride anions into water, such as ammonium fluorosilicate (NH ₄ ) ₂ SiF ₆ , silicon tetrafluoride SiF ₄ or sodium tetrafluoride Na ₂ SiF ₆ . Fluorine can be introduced for example by impregnation with an aqueous solution of hydrofluoric acid or ammonium fluoride.

An organic additive can optionally be added at any stage of preparation of said catalyst and preferably in the impregnation solution alone or with the different metal precursors.

Prior to the injection of the charge, the catalysts used in the processes according to the present invention are subjected to a sulfidation treatment making it possible to transform, at least in part, the metal species into sulphide before they are brought into contact with the charge to be treated. . This activation treatment by sulfurization is well known to those skilled in the art and can be carried out by any method already described in the literature either in-situ, that is to say in the reactor, or ex-situ. This treatment can be carried out simultaneously for the two catalysts used in the process or successively.

A classic sulfidation method well known to those skilled in the art consists of heating the catalyst in the presence of hydrogen sulfide (pure or for example under a flow of a hydrogen/hydrogen sulfide mixture) to a temperature between 150 and 800°C. , preferably between 250 and 600°C, generally in a crossed-bed reaction zone.

The invention is illustrated by the following examples which are in no way limiting.

Example 1: Preparation of an S1 support comprising a Y zeolite and a beta zeolite

The support S1 is prepared by mixing-extrusion of 15% by weight of commercial USY zeolite having a mesh parameter of 24.30 Å, a SiO ₂ /Al ₂ O ₃ molar ratio (SAR) of 30, a specific surface area measured by physisorption of nitrogen according to the BET method of 890 m ² /g, and 5% weight of commercial Beta zeolite having a molar SiO ₂ /Al ₂ O ₃ ratio of 24, a specific surface area measured by nitrogen physisorption according to the BET method of 670 m ² /g, in the presence of commercial boehmite (Pural SB3 from SASOL). The extrudates obtained are dried at 80°C then calcined at 600°C in humid air (5% weight of water per kg of dry air). The calcined support comprises, on a dry basis, 15% by weight of USY zeolite, 5% by weight of Beta zeolite and 80% by weight of alumina.

Example 2: Preparation of an S2 support comprising a beta zeolite with SAR <25

The S2 support is prepared by mixing-extrusion of 5% by weight of commercial Beta zeolite having a molar SiO2/Al2O3 ratio of 24, a specific surface area measured by nitrogen physisorption according to the B.E.T. method. of 670 m2/g, in the presence of commercial boehmite (Pural SB3 from SASOL). The extrudates obtained are dried at 80°C then calcined at 600°C in humid air (5% weight of water per kg of dry air). The calcined support comprises, on a dry basis, 5% by weight of Beta zeolite, and 95% by weight of alumina.

Example 3: Preparation of an S3 support comprising a Y zeolite

The S3 support is prepared by mixing-extrusion of 20% by weight of commercial USY zeolite having a lattice parameter of 24.30 Å, a molar SiO ₂ /Al ₂ O ₃ ratio of 30, a specific surface area measured by nitrogen physisorption according to the BET method of 890 m ² /g, in the presence of commercial boehmite (Pural SB3 from SASOL). The extrudates obtained are dried at 80°C then calcined at 600°C in humid air (5% weight of water per kg of dry air). The calcined support comprises, on a dry basis, 20% by weight of USY zeolite, and 80% by weight of alumina.

Example 4: Preparation of an S4 support (compliant) comprising a beta zeolite with SAR > 200

The support S2 is prepared by mixing-extrusion of 5% by weight of commercial Beta zeolite having a molar SiO ₂ /Al ₂ O ₃ ratio of 250, a specific surface area measured by nitrogen physisorption according to the BET method of 580 m ² /g , in the presence of commercial boehmite (Pural SB3 from SASOL). The extrudates obtained are dried at 80°C then calcined at 600°C in humid air (5% weight of water per kg of dry air). The calcined support comprises, on a dry basis, 5% by weight of Beta zeolite, and 95% by weight of alumina.

Example 5: Preparation of an aluminum S5 support

Support S2 is prepared by mixing-extrusion of commercial boehmite (Pural SB3 from SASOL). The extrudates obtained are dried at 80°C then calcined at 600°C in humid air (5% weight of water per kg of dry air). The calcined support therefore does not include zeolite.

Example 6: Preparation of catalysts:

- C1: NiMoP supported on S1 (USY+Beta),

- C2: NiMoP supported on S2 (Beta only whose SAR = 24)

- C3: NiMoP supported on S3 (USY),

- C4: NiMoP supported on S4 (Beta only whose SAR = 250),

- C5: NiMoP supported on S5 alumina without zeolite.

The catalysts C1, C2, C3 and C4 are prepared in the same way by dry impregnation of the supports S1, S2, S3 and S4 respectively using an aqueous solution containing the elements Ni, Mo. This solution is obtained by dissolution of the following precursors in water: nickel nitrate, and ammonium heptamolybdate. The quantity of precursors in solution is adjusted according to the concentrations targeted on the final catalyst. After dry impregnation, the catalyst is dried at 120°C in air.

Catalyst C5 is prepared in two stages. A C5_int catalyst is first prepared from the S5 support in a manner analogous to catalysts C1 to C4, but with different target contents and similar drying. Then this C5_Int catalyst is dry impregnated with a triethylene glycol solution such that the TEG/Mo molar ratio is 0.5 mol/mol and the additional volume composed of the solvent consists of 50% v/v of water and 50% v/v ethanol. The final catalyst C5 is obtained after a final drying step at 100°C.

The mass percentages in the catalyst are those indicated in the table below:

Catalysts C1 C2 C3 C4 C5 MoO3 (% dry weight) 17.1 17.2 17.1 17.3 23.6 NiO (% dry weight) 3.8 3.7 3.7 3.9 4.6

Example 7 (comparative) : hydrocracking of a mixture comprising 83% by weight of vacuum distillate and 17% by weight of animal fat, on the sequence of hydrotreatment catalyst followed by the hydrocracking catalyst based on a zeolite USY and a Beta zeolite.

Example 8 (comparative): hydrocracking of a mixture comprising 83% by weight of vacuum distillate and 17% by weight of animal fat, on the sequence of hydrotreatment catalyst followed by the hydrocracking catalyst based on a zeolite USY alone.

Example 9 (compliant): hydrocracking of a mixture comprising 83% by weight of vacuum distillate and 17% by weight of animal fat, on the sequence of hydrotreatment catalyst followed by the first hydrocracking catalyst based on a USY zeolite and a Beta zeolite, followed by a second hydrocracking catalyst based on a Beta zeolite having the SAR ratio of 250. The volume proportion of the first hydrocracking catalyst and the second hydrocracking catalyst is respectively 75 and 25%vol.

Example 10 (compliant) : hydrocracking of a mixture comprising 83% by weight of vacuum distillate and 17% by weight of animal fat, on the sequence of hydrotreatment catalyst followed by the first hydrocracking catalyst based on a USY zeolite alone, followed by a second hydrocracking catalyst based on a Beta zeolite having the SAR ratio of 250. The volume proportion of the first hydrocracking catalyst to the second hydrocracking catalyst is 75 and 25% vol.

Example 11 (compliant) : hydrocracking of a mixture comprising 83% by weight of distillate and 17% by weight of animal fat, on the sequence of hydrotreatment catalyst followed by the first hydrocracking catalyst based on a zeolite USY and a Beta zeolite, followed by a second hydrocracking catalyst based on a Beta zeolite having the SAR ratio of 24. The volume proportion of the first hydrocracking catalyst to the second hydrocracking catalyst is 75 and 25%vol.

Example 12 (compliant): hydrocracking of a mixture comprising 83% by weight of vacuum distillate and 17% by weight of animal fat, on the sequence of hydrotreatment catalyst followed by the first hydrocracking catalyst based on a USY zeolite alone, followed by a second hydrocracking catalyst based on a Beta zeolite having the SAR ratio of 24. The volume proportion of the first hydrocracking catalyst to the second hydrocracking catalyst is 75 and 25% vol.

For Examples 7 to 12, the characteristics of the mixture charge comprising 83% weight of vacuum distillate and 17% weight of animal fat are given in Table 2.

The feedstock, before being injected into the hydrocracking stage, is pretreated with a hydrotreatment catalyst, at a total pressure of 14 MPa and at a temperature and VVH pair making it possible to obtain an organic nitrogen content. of 10 ppm in the hydrotreated feed, that is to say, at the entrance to the hydrocracking catalytic beds.

All of the hydrotreated effluent from the hydrotreating section is directly sent to the hydrocracking section without an intermediate separation step. The temperatures of the hydrocracking catalysts are controlled to achieve a total conversion of the 370°C+ cut of 86% by weight.

In examples 9 to 12, the volume ratio of the first hydrocracking catalyst relative to the second hydrocracking catalyst is set at 75/25% vol. There is no intermediate separation of the effluent between these two hydrocracking catalysts.

All examples are compared at iso total pressure of 14 MPa, at iso total volume of catalysts and at iso feed rate.

The performance of the process is evaluated on the points below:

- the converting activity via the weighted average temperature of hydrocracking sections 1 and 2 combined necessary to reach 86% of the total conversion of the cut 370°C+, or WABT (Weighted Average Bed Temperature according to Anglo-Saxon terminology) ( Table 3),

- kerosene yield 150-280°C (Table 2).

- and the cold flow properties ( via the disappearance point of kerosene crystals and the cloud point of diesel) (Table 2).

The yields of main products given in Table 2 are calculated in relation to the liquid feed, that is to say the mixture comprising 83% by weight of vacuum distillate and 17% by weight of animal fat before the step of hydrotreatment. It should be noted that the sum of the yields given in Table 3 is not completed at 100% (but at 98.7%) because the yields of co-products (H2S, NH3, CO, CO2 and water) and the hydrogen consumption are not reported in Table 3, these being not different from one system to another because produced during the identical hydrotreatment step for all the systems compared.

Characteristics of the mixture of DSV and animal fat

Operating conditions and performance of the process of Examples 7-12

In Example 9 in accordance with the invention using the sequence of a catalyst comprising a USY zeolite and a Beta zeolite followed by a catalyst comprising a Beta zeolite having a SAR = 250, we were able to obtain an improvement in the properties at cold of the kerosene and diesel cuts (reduction in the crystal disappearance point of the kerosene cut and the cloud point of the diesel cut by 5 and 9°C respectively) as well as an improvement in kerosene yield (+0.1%) compared to Example 7 using a single catalyst comprising the USY and Beta hydrocracking zeolites over the entire catalytic zone.

In Example 10 in accordance with the invention using the sequence of a catalyst comprising a USY zeolite alone followed by a catalyst comprising a Beta zeolite having a SAR = 250, we were able to obtain an improvement in the cold properties of the cuts kerosene and diesel (reduction in the crystal disappearance point of the kerosene cut and the cloud point of the diesel cut by 7 and 10°C respectively) as well as an improvement in kerosene yield (+0.2%) compared to Example 8 using a single hydrocracking catalyst over the entire USY catalytic zone alone.

In Example 11 (in accordance with the invention), the sequence of a supported hydrocracking catalyst comprising a mixture of USY and Beta zeolites followed by the supported catalyst comprising the Beta zeolite with an SAR equal to 24 (according to the invention), allows us at the same time to obtain an improvement in the cold properties of diesel and kerosene (reduction in the disappearance point of crystals from the kerosene cut and the cloud point of the diesel cut by respectively 4 and 8°C in Example 11 according to the invention) and to maximize the kerosene yield (+1.3%) compared to Example 7.

In example 12 (in accordance with the invention), the sequence of a supported hydrocracking catalyst comprising a USY zeolite alone followed by the catalyst supported on the Beta zeolite whose Si/Al ratio of 24, allows us to improve the kerosene yields (+0.6%) compared to Example 8. Likewise, the cold properties of kerosene and diesel are also improved compared to Example 8 (reduction in the disappearance point crystals of the kerosene cut and the cloud point of the gas oil cut of respectively 4 and 8°C), showing that the invention also works with a first catalyst consisting solely of USY, without Beta.

Claims (11)

Process for producing middle distillates from at least one fossil hydrocarbon feedstock of which at least 50% by weight of the compounds have an initial boiling point greater than 250°C and a final boiling point less than 800°C, said fossil hydrocarbon feedstock being co-treated in a mixture with at least one renewable feedstock chosen from vegetable oils, algal oils, cooking oils and animal fats, fresh or used, alone or in a mixture, and feedstocks resulting from the reprocessing of biomass/plastics/tires/and household waste, alone or in mixtures, said process comprising a step of hydrotreating said mixture of feed in the presence of hydrogen and at least one hydrotreatment catalyst and a step of hydrocracking at least part and preferably all of the hydrotreated feedstock, the hydrocracking step being carried out at a temperature between 200°C and 480°C, at a total pressure between 1 MPa and 25 MPa with a ratio volume of hydrogen per volume of hydrocarbon feed of between 80 and 5000 liters per liter and at an Hourly Volume Velocity (VVH) defined by the ratio of the volume flow of liquid hydrocarbon feed to the volume of catalyst loaded into the reactor of between 0, 1 and 50 h ^-1 , the hydrocracking step of the hydrotreated feed comprising at least:
a) A step of bringing into contact at least part and preferably all of said hydrotreated feed in a first catalytic zone with at least one first catalyst comprising at least one metal from group VIB and/or at least one metal from group VIB group VIII of the periodic table and a support comprising at least one zeolite having at least one series of channels whose opening is defined by a ring with 12 oxygen atoms (12MR), and at least one binder.
b) Followed by bringing into contact in a second catalytic zone all of the effluent from step a), without an intermediate separation step between said first catalytic zone and said second catalytic zone, with a second catalyst comprising at at least one metal from group VIB and/or at least one metal from group VIII of the periodic table and a support comprising at least one zeolite with structural code BEA and at least one binder. Process according to claim 1 in which the hydrocarbon feedstock is chosen from light gas oils from a catalytic cracking unit, atmospheric distillates, vacuum distillates from direct distillation of crude or from conversion units such as FCC , coker, H-Oil or visbreaking, feeds from units for extracting aromatics from lubricating oil bases or from solvent dewaxing of lubricating oil bases, distillates from desulfurization processes or fixed bed or ebullated bed hydroconversion of RAT (atmospheric residues) and/or RSV (vacuum residues) and/or deasphalted oils. Process according to one of claims 1 or 2 in which the vegetable oils are crude or refined, totally or in part, and derived from the following plants: rapeseed, sunflower, soya, palm, palm kernel, olive, coconut, jatropha, and animal fats are chosen from bacon or fats composed of residues from the food industry or from the catering industries. Process according to one of the preceding claims in which the zeolite used in the support of the first catalyst is chosen from zeolites of structural type FAU, BEA, ISV, IWR, IWW, MEI, UWY, taken alone or as a mixture and preferably from zeolites of structural type FAU and BEA, taken alone or in a mixture. Process according to claim 4 in which the zeolite used in the support of the first catalyst is chosen from zeolite Y and zeolite beta taken alone or as a mixture and preferably the zeolite is zeolite Y and very preferably dealuminated zeolite USY. Process according to claim 5 in which the zeolite Y used in the support of the first catalyst has an initial crystal parameter a0 of the unit cell less than 24.55 Å, preferably less than 24.45 Å and preferably less than 24, 40 Å and even more preferably less than 24.35 Å. Process according to one of the preceding claims in which the zeolite with structural code BEA used in the support of the second catalyst is a beta zeolite preferably having a SiO2/Al2O3 or SAR molar ratio of between 10 and 300, preferably between 10 and 100, preferably between 10 and 50, preferably between 10 and 30, more preferably between 10 and 25, even more preferably between 12 and 24 and even more preferably between 15 and 24. Process according to claim 7 in which the second catalyst used in step b) can contain at least one other zeolite chosen from zeolites having at least one series of channels whose opening is defined by a ring with 12 atoms of oxygen (12MR), preferably chosen from zeolites of structural type FAU, BEA, ISV, IWR, IWW, MEI, UWY, MFI, MTW or even the zeolite IZM-2, taken alone or in mixture and preferably from zeolites of structural type FAU and BEA, taken alone or in a mixture. Process according to one of the preceding claims in which the volume fraction of the first catalyst used in step a) of the hydrocracking step relative to all of the two catalysts used in steps a) and b) is preferably between 50 and 95% vol and preferably between 60 and 95% vol, very preferably between 65 and 90% vol and even more preferably between 65 and 85% vol. Process according to one of the preceding claims in which the volume fraction of the second catalyst used in step b) relative to all of the two catalysts used in steps a) and b) is between 5 and 50% vol and preferably between 5 and 40% vol, very preferably between 10 and 35% vol and even more preferably between 15 and 35% vol. Process according to one of the preceding claims in which the effluent from step b) of the hydrocracking step is sent to a fractionation column to recover at least one middle distillate cut and at least one heavy cut comprising the unconverted compounds and said heavy cut is sent either to a second hydrocracking step in the presence of hydrogen and a hydrocracking catalyst or to hydrotreatment step a).