FR3142196A1 - PROCESS FOR PRODUCING MIDDLE DISTILLATES COMPRISING A CATALYST SEQUENCE INCLUDING A BETA ZEOLITE BASED CATALYST HAVING A SILICA/ALUMINA RATIO LESS THAN 25 - Google Patents

Abstract

The present invention relates to a process for producing middle distillates from at least one hydrocarbon feed comprising a step of hydrotreating said feeds in the presence of hydrogen and at least one hydrotreatment catalyst and a step of hydrocracking the hydrotreated feed, the hydrocracking step comprising at least one step of bringing said hydrotreated feed 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 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, followed by placing in contact in a second catalytic zone with 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 having a SiO2/Al2O3 molar ratio of less than 25, and at least one binder.

Description

PROCESS FOR PRODUCING MIDDLE DISTILLATES COMPRISING A CATALYST SEQUENCE INCLUDING A CATALYST BASED ON BETA ZEOLITE HAVING A SILICA/ALUMINA RATIO OF LESS THAN 25

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 separate hydrocracking zones operated in series.

This process converts a set of hydrocarbon feedstocks into products of interest (naphtha, jet, diesel). These hydrocarbon feedstocks contain, for example, aromatic, and/or olefinic, and/or naphthenic, and/or paraffinic compounds, including so-called fossil feedstocks, such as vacuum distillate cuts (DSV) or Vacuum Gas Oil (VGO) from direct distillation of crude or from conversion units such as FCC, coker, H-Oil or visbreaking, excluding feedstocks from the Fischer-Tropsch process, feedstocks from the transformation of oil sands and shale oils, renewable feedstocks selected from vegetable oils, algal oils, animal fats, fresh or used, alone or in a mixture, and feedstocks from the reprocessing of plastics/tires/and household waste, alone or in a mixture. These charges may 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 diesel cut having initial and final boiling points in a range from 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, comprising a hydrotreatment step and a hydrocracking step. The hydrocracking step is carried out in the presence of a sequence of 2 specific catalysts, implemented in two separate catalytic zones, all of the effluent from the first catalytic zone being directly sent to the second catalytic zone in contact with a second catalyst, without intermediate separation of the gas phase between the 2 catalytic zones of said second hydrocracking step.

In particular, the first catalyst used in the first catalytic zone of the hydrocracking stage 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 stage 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 BEA structural type having a SiO2/Al2O3 molar ratio of less than 25 (also called SAR).

The process for producing middle distillates including a hydrocracking stage of heavy petroleum fractions is now an essential refining process that allows the production, from excess and poorly recoverable heavy feedstocks, of lighter fractions such as gasoline, jet fuels and diesel that the refiner is looking for to adapt its production to the structure of demand. Some hydrocracking processes also allow a highly purified residue to be obtained that can provide excellent bases for the production of oils. Compared to catalytic cracking (FCC), the advantage of catalytic hydrocracking is that it provides very good quality middle distillates. Conversely, the gasoline produced has a much lower octane rating than that resulting from catalytic cracking.

This process derives 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 stages. It is therefore suitable for treating and transforming any type of hydrocarbon feedstock as long as it is compatible with injection in liquid form into catalytic reactors.

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

The hydrocracking catalysts used in this process for producing middle distillates are said to be bifunctional, i.e. 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 aluminium oxides, amorphous silica-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 that 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, generally working at high temperatures (greater than or equal to 390-400°C), and at low feed space velocity (the VVH expressed in volume of feed 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 for middle distillates (jet fuels and diesels). Conversely, a strong acid function and a weak hydro-dehydrogenating function give active catalysts, but with poorer selectivities for middle distillates.

One type of conventional hydrocracking catalyst 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 amorphous-supported catalysts is their low activity.

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

In general, hydrocracking catalysts that only contain zeolite Y do not allow good cold properties to be obtained for the diesel cut (in terms of cloud point, pour point, filterability limit temperature of said cut), but also for the kerosene cut (in terms of crystal appearance point). These values are very correlated with the higher boiling linear paraffins present in these cuts. To comply with fuel specifications on cold properties, this may force the refiner 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 are less recoverable.

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

For example, we can cite the work relating to the use of catalysts comprising a Y zeolite modified for example by dealumination by steaming or acid attack, the use of composite catalysts, or the use of small crystals of Y zeolites. 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 that it is also necessary to select precisely 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 removal of the structuring agent used in manufacturing. Yield gains in middle 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 the diesel cut, the gain in the kerosene cut remaining low or even zero.

An alternative would consist, as described in patent US7749373, in implementing 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 vacuum distillate type feedstock, the sequence of, firstly, 75% by volume of a catalyst comprising 10% by weight in its support of a Y zeolite with a 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 achieve 87% target conversion, compared with a feedstock carried out with only the catalyst supported on the Y zeolite. A gain of 3.5% by weight in the yield of middle distillates is then also noted, which is entirely carried by the increase in yield on the diesel cut, the kerosene yield not evolving (as well as its cold properties). In other words, the process implementing said catalyst sequence allows an improvement in the yield in diesel cut as well as an improvement in the cold properties of said diesel cut in terms of pour point but said process does not allow any improvement in the yield in the kerosene cut obtained. The same applies to the cold properties of the kerosene cut which are identical to those obtained compared to a process implementing only a catalyst comprising a Y zeolite.

The applicant has discovered that the use, in a hydrocracking step of a process for producing middle distillates from a hydrocarbon feedstock, having previously been hydrotreated, of a sequence, in a first catalytic zone, of a first hydrocracking catalyst comprising a large-pore zeolite, preferably having at least one opening of at least 12MR, and metals from group VI and from group VIII according to the periodic table, then, without any separation being made between the first and a second catalytic zone, of a second catalyst, implemented 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 having a SiO2/Al2O3 molar ratio (SAR) of less than 25, allows,

- to improve not only the selectivity towards the diesel cut, but also the selectivity towards the kerosene cut while simultaneously leading to cold properties of said diesel and kerosene cuts improved compared to the use of any other catalyst or series of catalysts known to those skilled in the art.

The subject of the present invention is a process for producing middle distillates from at least one 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 hydrocarbon feedstock not comprising paraffinic feedstocks from the Fischer-Tropsch process, feedstocks from the transformation of bituminous sands and shale oils, renewable feedstocks chosen from vegetable oils, algal oils and animal fats, fresh or used, alone or in a mixture, and feedstocks from the reprocessing of plastics/tires/and household waste, alone or in a mixture, said process comprising a step of hydrotreatment of said feedstocks 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 of hydrogen volume to hydrocarbon feed volume between 80 and 5000 liters per liter and at an Hourly Volumetric Velocity (HVV) defined by the ratio of the volumetric flow rate of liquid hydrocarbon feed to the volume of catalyst loaded into 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 a portion and preferably all of said hydrotreated feedstock 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 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 contacting 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 having a SiO2/Al2O3 molar ratio of less than 25, 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 allowing a high yield of middle distillate cut to be obtained, as well as cold properties of said middle distillate cut improved compared to those obtained in the prior art.

In particular, an advantage of the present invention is to provide a process for producing middle distillates making it possible to improve not only the selectivity in the diesel cut, but also in the kerosene cut while also making it possible to obtain improved cold properties of these two cuts compared to the use of catalysts or a chain of catalysts of the prior art.

For the purposes of the present invention, the various embodiments presented may be used alone or in combination with each other, without limitation of combination.

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

In the following 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 remainder 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 less, 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 the 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 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.

Charges

A wide variety of hydrocarbon feedstocks can be processed by the processes for producing middle distillates according to the invention including a hydrocracking step. The hydrocarbon feedstock used in the 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 final boiling point less than 800°C.

According to the invention, the hydrocarbon feedstock does not include paraffinic feedstocks from the Fischer-Tropsch process, feedstocks from the processing of bituminous sands and shale oils, feedstocks from biomass chosen from vegetable oils, algal oils and animal fats, fresh or used, alone or in a mixture, and feedstocks from the reprocessing of plastics/tires/and household waste, alone or in a mixture,

The hydrocarbon feedstock is advantageously selected from diesel fuels from a catalytic cracking unit or LCO (Light Cycle Oil) according to the English terminology, atmospheric distillates, vacuum distillates from the direct distillation of crude or from conversion units such as FCC, coker, H-Oil or visbreaking, feedstocks from aromatic extraction units 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. More generally, the process according to the invention is capable of treating the aforementioned feedstocks alone or in a mixture regardless of the proportions. The above list is not limiting. Said fillers preferably have a boiling point T5 greater than 250°C, preferably greater than 340°C, that is to say that 95% of the compounds present in the filler have a boiling point greater than 250°C, and preferably greater than 340°C.

These feeds may 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 feed and the hydrocracking step.

The nitrogen content of the feedstocks treated in the processes according to the invention is generally greater than 500 ppm by weight, preferably between 500 and 10,000 ppm by weight, more preferably between 700 and 4,000 ppm by weight and even more preferably between 1,000 and 4,000 ppm by weight. The sulfur content of the feedstocks 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.

The charge may optionally 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 more preferably less than 1 ppm by weight.

The feedstock may optionally 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 200 ppm by weight.

According to the invention, upstream of the hydrocracking step, the process comprises a step of hydrotreatment of said feeds in the presence of hydrogen and at least one hydrotreatment catalyst, said hydrotreatment step preferably operating at a temperature of between 200 and 450°C, under a pressure of between 2 and 25 MPa, at a space velocity of 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.

Operating conditions such as temperature, pressure, hydrogen recycling rate, hourly space velocity, 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 of between 250 and 450°C, very preferably between 300 and 430°C, under a pressure of between 5 and 20 MPa, at a space velocity of 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 containing at least one amorphous support and at least one hydro-dehydrogenating element chosen from at least one non-noble element from groups VIB and VIII, and most often at least one element from group VIB (from molybdenum or tungsten taken alone or in a mixture) and at least one non-noble element from group VIII (from nickel or cobalt taken alone or in a mixture).

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

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

The effluent from the hydrotreatment stage, part of which enters the hydrocracking stage 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 can advantageously comprise a separation step between the hydrotreatment step and the hydrocracking step.

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

According to the invention, the process comprises a step of hydrocracking at least part and preferably all of the hydrotreated feedstock, the hydrocracking step being carried out at a temperature of between 200°C and 480°C, at a total pressure of between 1 MPa and 25 MPa with a ratio of hydrogen volume to hydrocarbon feedstock volume of between 80 and 5000 liters per liter and at an Hourly Volumetric Velocity (HVV) defined by the ratio of the volumetric flow rate of liquid hydrocarbon feedstock to the volume of catalyst loaded into the reactor of 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, under a pressure between 2 and 25 MPa, preferably between 3 and 20 MPa, at a space velocity between 0.1 and 20 h ^-1 , preferably 0.1 and 6 h ^-1 , preferably between 0.2 and 3 h ^-1 , and the amount of hydrogen introduced is such that the volume ratio liter of hydrogen/liter of hydrocarbon is between 100 and 3000 NL/L.

These operating conditions used in the processes according to the invention generally make it possible to achieve conversions per pass, into 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.

Furthermore, the kerosene (jet fuel) fraction produced by the process according to the invention has a crystal disappearance point of less than -20°C, preferably less than -30°C and even more preferably less than -47°C. The diesel fraction produced by the process has a cloud point of less than 15°C, preferably less than 5°C and very preferably less than -5°C.

According to the invention, the hydrocracking step of at least part and preferably all of the hydrotreated feedstock comprises at least:

a) A step of bringing said hydrotreated charge 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 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 contacting in a second catalytic zone all of the effluent from step a), without an 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 having a SiO2/Al2O3 or SAR molar ratio of less than 25, 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 selected from the group formed by tungsten and molybdenum, taken alone or in mixture. According to a preferred embodiment, the metal of group VIB is molybdenum. According to another preferred embodiment, the metal of group VIB is tungsten.

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

Advantageously, 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 content of metal 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 comprised, 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 sulfur 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, with 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 may 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, of 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 one 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 step a) of hydrocracking 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 a 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 as a mixture. Preferably, the binder is chosen from alumina and silica-alumina, taken alone or as a mixture. More preferably, the binder is alumina. The 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 in 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 of at least one zeolite relative to the total mass of said support.

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

- 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 of at least one zeolite relative to the total mass of said support,

- 20 to 99% by weight, preferably from 30 to 99%, more 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 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. Mc Cusker, D.H. Olson, 6th Edition, Elsevier, 2007, Elsevier".

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

The Y zeolite 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 more preferably between 650 and 1050 m ² /g.

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

Partial or total exchange of alkali cations by NH4+ cations is understood to mean the exchange of 80 to 100%, preferably 85 to 99.5% and more preferably 88 to 99%, of said alkali cations by NH4+ cations. At the end of the ion exchange step(s), the remaining amount of alkali cations, and preferably the remaining amount of Na+ cations, in the zeolite Y, relative to the amount of alkali cations, preferably Na+, initially present in the zeolite Y, is advantageously between 0 and 20%, preferably between 0.5 and 15% and more 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 the alkaline cations and preferably the Na+ cations present in the zeolite. Preferably, the ammonium salt is ammonium nitrate NH4NO3.

Thus, the remaining content of alkali cations and preferably of Na+ cations in the Y zeolite at the end of the ion exchange(s) step is preferably such that the alkali cation/aluminium 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 alkali cation/aluminium 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 NH4+ concentration of the 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.

Where appropriate, said Y zeolite obtained can then undergo a dealumination treatment step. Said dealumination step can advantageously be carried out by any 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 the Anglo-Saxon terminology) and/or by one or more acid attacks advantageously carried out by treatment with an aqueous solution of mineral or organic acid.

Preferably, the dealumination step involves 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 of between 200 and 900°C, preferably between 300 and 900°C, even more preferably between 400 and 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 formed of air. The flow rate of gas formed from 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.

Heat treatment allows the extraction of aluminum atoms from the framework of the Y zeolite 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 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 crystalline parameter a0 of the elementary mesh 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 partially or completely eliminate the aluminum debris resulting from the heat treatment step in the presence of water vapor and which partially block the porosity of the dealuminated zeolite; the acid attack therefore makes it possible to unblock the porosity of the dealuminated zeolite.

The acid attack can advantageously be carried out by suspending the Y zeolite, which has optionally previously undergone a 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 can be formic acid, acetic acid, oxalic acid, tartaric acid, maleic acid, malonic acid, malic acid, lactic acid, or any other organic acid soluble in water. The concentration of the mineral or organic acid solution 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 optionally the acid attack step, the method for modifying said zeolite Y advantageously comprises a step of at least a partial or total exchange of the alkaline cations and preferably of the Na+ cations still present in the cationic position in the zeolite Y. The ion exchange step is carried out in a similar manner 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 the acid attack step and optionally the step of partial or total exchange of the alkali cations and preferably of the Na+ cations, the method 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 step of partial or total exchange of the alkali cations. In addition, said calcination step makes it possible to generate the protonated form of zeolite Y and to give it an acidity for its applications.

The calcination can advantageously be carried out in a muffle furnace or a tubular furnace, in dry air or in 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 Y zeolite obtained has an initial crystalline parameter a0 of the elementary mesh of less than 24.55 Å, preferably of less than 24.45 Å and more preferably of less than 24.40 Å and more preferably of 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 more 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 BET method of 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 the 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 (Elsevier, 2001).

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 from 1 to 70%, and preferably from 5 to 50%, and very preferably from 5 to 40% by weight relative to the total weight of said support of a Y zeolite having an initial crystalline parameter a0 of the elementary mesh of less than 24.55 Å;

- 0.1 to 6%, preferably 0.2 to 5%, and more preferably 0.5 to 5% 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 total weight 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 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 having a SiO2/Al2O3 or SAR molar ratio of less than 25, and at least one binder.

Preferably, the effluent from step a) is sent to a second catalytic zone without an 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 sulphide form.

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

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

Advantageously, 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 second catalyst comprises at least one metal from group VIB in combination with at least one non-noble metal from group VIII, the content of metal 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 comprised, 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 sulfur 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, with 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 may 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, of 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 one element chosen from group VIIA, preferably fluorine.

According to 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 having a SiO2/Al2O3 or SAR molar ratio of less than 25, and preferably less than 24, preferably between 10 and 25, preferentially between 12 and 24, and preferably between 15 and 24 and at least one binder.

The binder used in the support of the second catalyst, also called a 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 as a mixture. Preferably, the binder is chosen from alumina and silica-alumina, taken alone or as a mixture. More preferably, the binder is alumina. The 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.

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 the 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 (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 a size of up to 50 µm, preferably less than 25 µm. On the other hand, the Beta zeolite used in the support of the 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 area 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, of 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 preferably between 21 and 37%.

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

- 1 to 70% by weight, preferably from 1 to 50% by weight, more preferably from 1.5 to 40% by weight, and even more preferably 2 to 30% by weight of said beta zeolite with a SiO2/Al2O3 or SAR molar ratio of less than 25 relative to the total mass of said support,

- 30 to 99% by weight, preferably 50 to 99%, more preferably 60 to 98% by weight, and very preferably 70 to 98% by weight relative to the total mass of said support, of at least 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 as a mixture and preferably from zeolites of structural type FAU and BEA, taken alone or as a mixture. Preferably, the second catalyst may comprise, in addition to the zeolite of structural code BEA having a SAR of less than 25, another zeolite of structural code FAU and preferably a Y zeolite.

Preferably, the zeolite Y which can be used in the second catalyst has an initial crystalline parameter a0 of the elementary mesh of 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 of 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 of structural code BEA with a SAR of less than 25 and at least one zeolite Y, and at least one binder, said support comprising and preferably consisting of, preferably:

- 1 to 50% by weight, preferably from 1 to 30% by weight, more preferably from 1.5 to 20% by weight, and even more preferably 2 to 30% by weight of said beta zeolite having a SAR of less than 25 relative to the total mass of said support,

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

- 30 to 98.5% by weight, preferably from 40 to 98%, more preferably from 70 to 97% by weight, and very preferably from 65 to 96% by weight relative to the total mass of said support, of 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 the whole 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 the whole 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 reactor(s), in a fixed bed or in an ebullated bed. It can be carried out with or without recycle, after said hydrotreatment stage aimed at removing the sulfur, nitrogen or oxygen compounds possibly present in the feeds. In the case where a recycle is carried out, this can be carried out at any location in the process, either in the hydrotreatment stage or in the first catalytic zone of the hydrocracking stage a).

In an embodiment of the so-called 2-stage 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 the hydrotreatment step a).

Preparation of the first and second hydrocracking catalysts according to the invention

The first and second hydrocracking catalysts used respectively in step 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 stage including:

- 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 of structural code BEA with a SAR less than 25,

And,

- the shaping of 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 non-noble elements of group VIII of the periodic table, preferably cobalt, nickel, and their mixtures, and preferably nickel and cobalt, and their mixtures, onto 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 at the end of 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 method comprising the following steps:

- Preparation of the 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,

- Mixing with a porous mineral matrix and shaping 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 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 steps.

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

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

- Mixing with a porous mineral matrix and shaping 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 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 steps.

Preferably, the porous mineral matrix is derived from an alumina gel which has a crystallite size of between 2 and 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% by weight.

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 kneading extrusion, by granulation and by the technique known as oil drop according to Anglo-Saxon terminology.

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

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

Preferably, the support advantageously has a specific surface area greater than 100 ^m2 /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 enclosed in pores with an average diameter 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 rotating plate 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 sizes. They are generally used in the form of cylindrical or polylobe 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. It is however 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 may or may not be hollow), twisted cylindrical, multilobed (2, 3, 4 or 5 lobes for example), rings. Any other shape 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 thus 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 embodiment 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 be advantageously added to a mixture composed of an alumina compound in an acid medium with a totally 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 between 500 and 1000°C, for a duration between 2 and 10 h, in the presence or absence of an air flow containing up to 60% by 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 a Boehmite, said calcination step allows the transition of boehmite to the final alumina.

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

The introduction of at least one hydro-dehydrogenating element may advantageously be accompanied by that of at least one promoter element chosen from phosphorus, boron, silicon and preferably phosphorus and optionally the introduction of an element from 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, and preferably calcined, support 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, i.e. 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 of group VIII, the metals of 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 the latter.

The introduction of at least one hydro-dehydrogenating element may then be optionally 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, ammonium tungstate, phosphomolybdic acid, phosphotungstic acid and their salts, silicomolybdic acid, silicotungstic acid and their salts. Preferably, ammonium oxides and salts such as ammonium molybdate, ammonium heptamolybdate and ammonium tungstate are used.

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

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. Tungstophosphoric or tungstomolybdic acids can be used.

The phosphorus content is adjusted, without limiting the scope of the invention, so 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 may be Anderson heteropolyanions, for example.

The boron source may be boric acid, preferably orthoboric acid H ₃ BO ₃ , ammonium biborate or pentaborate, boron oxide, boric esters. Boron may 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 of the pyridine and quinoline family and compounds of the pyrrole family. Boron may be introduced for example by a solution of boric acid in a water-alcohol mixture.

The sources of group VB elements that can be used are well known to those skilled in the art. For example, among the sources of niobium, it is possible to use oxides, 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 that 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 hydrofluoric acid. It is also possible to use hydrolyzable compounds that can release fluoride anions in 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 of an aqueous solution of hydrofluoric acid or ammonium fluoride.

An organic additive may 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 feedstock, the catalysts used in the processes according to the present invention are subjected to a sulfurization treatment making it possible to transform, at least in part, the metal species into sulfide before they are brought into contact with the feedstock to be treated. This sulfurization activation treatment 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, i.e. in the reactor, or ex situ. This treatment can be carried out simultaneously for the two catalysts used in the process or successively.

A conventional sulfurization 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) at 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 a support S1 comprising a zeolite Y and a zeolite beta

The S1 support is prepared by kneading-extrusion of 15% by weight of commercial USY zeolite having a mesh parameter of 24.30 Å, a molar SiO ₂ /Al ₂ O ₃ ratio (SAR) of 30, a specific surface area measured by nitrogen physisorption according to the BET method of 890 m ² /g, and 5% by 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 and then calcined at 600 ° C in humid air (5% by 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 kneading-extrusion of 5% by 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% by 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 kneading-extrusion of 20% by weight of commercial USY zeolite having a mesh parameter of 24.30 Å, a SiO ₂ /Al ₂ O ₃ molar 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% by 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 (non-compliant) comprising a beta zeolite with SAR > 200

The S2 support is prepared by kneading-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% by 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 S5 aluminum support

The S2 support is prepared by kneading-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 contain 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.

Catalysts C1, C2, C3 and C4 are prepared in the same way by dry impregnation of supports S1, S2, S3 and S4 respectively using an aqueous solution containing the elements Ni, Mo. This solution is obtained by dissolving 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 steps. A catalyst C5_int is first prepared from support S5 in a manner similar to catalysts C1 to C4, but with different target contents and similar drying. Then this catalyst C5_Int is dry impregnated with a triethylene glycol solution such that the TEG/Mo molar ratio is 0.5 mol/mol and the volume balance of the solvent is 50% v/v 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: Comparison of the performance of the catalyst sequence C1 then C2 (compliant) to C1 then C4 (non-compliant) and C3 then C2 (compliant) to C3 then C4 (non-compliant) in hydrocracking of a vacuum distillate

The preparation and characteristics of the catalysts are described in Example 6). Below, Example 7A will be compared with Examples 7C, 7D, and Example 7B will be compared with Examples 7E, 7F. For all the catalyst sequences, whether compliant or non-compliant, the hydrotreatment catalyst C5 is added upstream of the hydrocracking catalyst(s).

- example 7A (non-compliant, comparative): sequence of a C5 hydrotreatment catalyst, followed by the C1 hydrocracking catalyst.

- example 7B (non-compliant, comparative): sequence of a C5 hydrotreatment catalyst, followed by the C3 hydrocracking catalyst.

- example 7C (compliant): the sequence of catalysts: hydrotreatment (C5), followed by the first hydrocracking catalyst C1, followed by the second hydrocracking catalyst C2. In the hydrocracking step, the respective proportion of catalysts C1 and C2 is as follows: 75/25% by volume.

- example 7D (non-compliant): sequence of hydrotreatment catalyst (C5) followed by the first hydrocracking catalyst C1, followed by the second hydrocracking catalyst C4. The respective proportion of catalysts C1 and C4 is as follows: 75/25% by volume.

- example 7E (compliant): sequence of hydrotreatment catalyst (C5), followed by the first hydrocracking catalyst C3, followed by the second hydrocracking catalyst C2. The respective proportion of catalysts C3 and C2 is as follows: 75/25% by volume.

- example 7F (non-compliant): sequence of hydrotreatment catalyst (C5) followed by hydrocracking catalyst C3, followed by the second hydrocracking catalyst C4. The respective proportion of catalysts C3 and C4 is as follows: 75/25% by volume.

The characteristics of the vacuum distillate (VDS) are given in Table 1.

In all examples, the vacuum distillate, before being injected into the hydrocracking catalyst beds, is pretreated with the hydrotreatment catalyst C5, at a total pressure of 140 MPa, and a VVH and temperature pair such that it makes it possible to obtain an organic nitrogen content of 10 ppm in the hydrotreated DSV, i.e., at the inlet of the hydrocracking catalytic beds.

The performances of the catalysts for the different examples are compared at the same total pressure of 14 MPa, and at the same VVH.

Catalyst bed temperatures are adjusted to target a total net conversion of the 370°C+ cut of 82% (i.e. conversion of the cut boiling above 370°C in the charge).

The operating conditions and process performances according to the examples are compared in Table 3. The process performances are evaluated on the following points:

- the converting activity, via the weighted average temperature of hydrocracking sections 1 and 2 combined, necessary to reach 82% of the total conversion of the 370°C+ cut, or WABT (Weighted Average Bed Temperature according to Anglo-Saxon terminology),

- the yield of middle distillates 150-370°C (consisting of the kerosene cut 150-280°C and the diesel cut 280-370°C),

- and cold flow properties ( via the crystal disappearance point for the kerosene cut and the cloud point for the diesel cut) (Table 3).

The yields of main products given in Table 3 are calculated relative to the liquid feed, i.e. the vacuum distillate before the hydrotreatment step. It should be noted that the sum of the yields given in Table 3 is not 100% (but 100.2%) because the yields of co-products (H2S and NH3) and the hydrogen consumption are not reported in Table 3, these remaining identical between the different systems.

Vacuum Distillate (VDD) Feed Characteristics

Name of the charge DSV Density at 15°C g/cm3 0.9284 Refractive index at 70°C 1,4972 Sulfur % m/m 1.89 Nitrogen ppm m/m 1395 Carbon % m/m 85.6 Hydrogen % m/m 12.32 UV-Monoaromatics % m/m 23.15 UV-Diaromatics % m/m 10.92 UV-Triaromatics % m/m 12.92 UV-Total Aromatics % m/m 46.99 Aromatic Carbon % 17.8 DS initial point °C 316.1 DS 5% weight °C 394.2 DS 10% weight °C 413 DS 20% weight °C 435.6 DS 30% weight °C 449.8 DS 40% weight °C 466.4 DS 50% weight °C 479.8 DS 60% weight °C 498.2 DS 70% weight °C 514.7 DS 80% weight °C 537.2 DS 90% weight °C 562.5 DS 95% weight °C 580.8 DS full stop °C 624.6

Operating conditions and process performance (example 7)

In Examples 7A and 7B (non-compliant), we used a single hydrocracking catalyst corresponding to the formulations known from the prior art: catalyst comprising USY and Beta zeolites (catalyst C1, example 7A) and catalyst comprising USY zeolite alone (catalyst C3, example 7B).

In examples 7C and 7D, we replaced 25% by volume of catalyst C1 with the catalyst comprising zeolite Beta alone whose SAR is equal to 24 (catalyst C2, example 7C in accordance with the invention) or with the catalyst comprising zeolite Beta whose SAR is equal to 250 (catalyst C4, example 7D not in accordance with the invention).

- Example 7C: the use of catalyst C2 comprising Beta zeolite with SAR of 24 allowed us to have several advantages: increase in the yield of middle distillates (+1.7%), in particular kerosene (+1.2%) and diesel (+0.5%), and improvement in the cold properties 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 4°C respectively) compared to example 7A.

- Example 7D: the use of catalyst C4 comprising high SAR Beta zeolite (250) did not improve the yield of the middle distillate or that of the Kerosene cut alone, despite a gain in the cold properties of the kerosene and diesel cuts compared to example 7A.

In examples 7E and 7F, we replaced 25% by volume of catalyst C3 with the catalyst comprising zeolite Beta alone whose SAR is equal to 24 (catalyst C2, example 7E in accordance with the invention) or the catalyst comprising zeolite Beta whose SAR is equal to 250 (catalyst C4, example 7F not in accordance with the invention).

- Example 7E: the use of catalyst C2 comprising the SAR 24 Beta zeolite makes it possible to increase the yield of middle distillate (+1.4%), in particular of the kerosene cut (+1.1%), slightly of the diesel cut (+0.3% by weight) and to slightly improve the cold properties of the kerosene and diesel cuts compared to example 7B (reduction in the crystal disappearance point of the kerosene cut and the cloud point of the diesel cut by 3 and 4°C respectively).

- Example 7F: the use of catalyst C4 comprising Beta zeolite with high SAR (250) did not improve the yield of the middle distillate or that of the Kerosene cut alone despite a gain in the cold properties of the kerosene and diesel cuts compared to example 7B.

These examples justify that the use of a supported catalyst comprising a Beta zeolite whose SiO ₂ /Al ₂ O ₃ (SAR) molar ratio is less than 25, downstream of a first hydrocracking catalyst, is very advantageous for jointly improving the yield in kerosene cut, in diesel cut, and the cold properties of the kerosene and diesel cuts.

Claims (11)

Process for producing middle distillates from at least one 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 hydrocarbon feedstock not comprising paraffinic feedstocks from the Fischer-Tropsch process, feedstocks from the transformation of bituminous sands and shale oils, feedstocks from biomass chosen from vegetable oils, algal oils and animal fats, fresh or used, alone or in a mixture, and feedstocks from the reprocessing of plastics/tires/and household waste, alone or in a mixture, said process comprising a step of hydrotreatment of said feedstocks 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 of between 1 MPa and 25 MPa with a ratio of hydrogen volume to hydrocarbon feed volume of between 80 and 5000 liters per liter and at an Hourly Volumetric Velocity (HVV) defined by the ratio of the volumetric flow rate 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 a portion and preferably all of said hydrotreated feedstock 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 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 contacting 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 having a SiO2/Al2O3 molar ratio of less than 25, and at least one binder. Process according to claim 1, in which the hydrocarbon feedstock is selected 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, feedstocks 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. Process according to one of claims 1 or 2, in which the hydrocracking step 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, under a pressure between 2 and 25 MPa, preferably between 3 and 20 MPa, at a space velocity 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 NL/L. 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 in 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 crystalline parameter a0 of the elementary mesh of less than 24.55 Å, preferably less than 24.45 Å and more 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 has a SiO2/Al2O3 or SAR molar ratio of less than 25, preferably between 10 and 25, preferentially between 12 and 24, and more preferably between 15 and 24. Process according to claim 7, in which the second catalyst used in step b) may 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 oxygen atoms (12MR), preferably chosen from zeolites of structural type FAU, BEA, ISV, IWR, IWW, MEI, UWY, MFI, MTW or even zeolite IZM-2, taken alone or as a mixture and preferably from zeolites of structural type FAU and BEA, taken alone or as 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 the set 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 the whole 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 the hydrotreatment step a).