FR2950896A1 - Making middle distillates from paraffin charge produced by Fischer-Tropsch synthesis comprises implementing hydrocracking catalyst comprising hydrodehydrogenating metal and composite support formed by Y-type zeolite and silicon carbide - Google Patents

Abstract

Preparing middle distillates from a paraffin charge produced by Fischer-Tropsch synthesis comprises implementing a hydrocracking and/or hydroisomerization catalyst comprising at least one hydrodehydrogenating metal comprising metals of group VIB and/or VIII of the periodic table and a composite support formed by a Y-type zeolite and silicon carbide, where the: process is operated at a temperature of 270-400[deg] C, a pressure of 1-9 MPa and hourly volume speed of 0.5-5 hours -> 1>, and hydrogen flow is adjusted to obtain a ratio of 400-1500 normal liters of hydrogen per liter of charge.

Description

1

The present invention relates to a method for producing middle distillates from a paraffinic feedstock produced by Fischer-Tropsch synthesis, using a hydrocracking and / or hydroisomerization catalyst comprising at least one hydro-dehydrogenating metal chosen from group formed by Group VIB metals and Group VIII of the Periodic Table, taken alone or as a mixture, and a composite support formed by a zeolite of structural type Y and silicon carbide SiC, said process operating at a temperature between 270 and 400 ° C., a pressure of between 1 and 9 MPa, a hourly space velocity of between 0.5 and 5 hours, a hydrogen flow rate adjusted to obtain a ratio of 400 to 1500 normal liters of hydrogen per second. liter of charge.

PRIOR ART In the low temperature Fischer-Tropsch process, the synthesis gas (CO + H2) is catalytically converted into oxygenates and substantially linear hydrocarbons in gaseous, liquid or solid form. These products are generally free of heteroatomic impurities such as, for example, sulfur, nitrogen or metals. They also contain practically little or no aromatics, naphthenes and more generally cycles especially in the case of cobalt catalysts. On the other hand, they may have a significant content of oxygenated products which, expressed by weight of oxygen, is generally less than about 5% by weight and also an unsaturated content (olefinic products in general) generally less than 10% by weight. However, these products, mainly made of normal paraffins, can not be used as such, in particular because of their cold-holding properties that are not very compatible with the usual uses of petroleum fractions. For example, the pour point of a linear hydrocarbon containing 20 carbon atoms per molecule (boiling point equal to about 340 ° C., that is often included in the middle distillates cut) is + 37 ° C. about which makes its use impossible, the specification being -15 ° C for diesel. The hydrocarbons resulting from the Fischer-Tropsch process comprising mainly n-paraffins must be converted into more valuable products such as, for example, gas oil, kerosene, which are obtained, for example, after catalytic reactions of hydroisomerization and hydrocracking. All catalysts currently used in hydroisomerization / hydrocracking are of the bifunctional type associating an acid function with a hydrogenating function. The acid function is provided by supports with large surface areas (generally 150 to 800 m2.g) having a superficial acidity, such as halogenated aluminas (chlorinated or fluorinated in particular), phosphorus aluminas, combinations of boron oxides and aluminum, and silica-aluminas. The hydrogenating function is provided either by one or more metals of group VIII of the periodic table of the elements, such as iron, cobalt, nickel, ruthenium, rhodium, palladium, osmium, iridium and platinum, or by a combination of at least one Group VI metal such as chromium, molybdenum and tungsten and at least one Group VIII metal.

The equilibrium between the two acid and hydrogenating functions is one of the parameters that govern the activity and the selectivity of the catalyst. A weak acid function and a strong hydrogenating function give

catalysts are weakly active and selective towards isomerization whereas a strong acid function and a low hydrogenating function give very active and cracking-selective catalysts. A third possibility is to use a strong acid function and a strong hydrogenating function to obtain a very active catalyst but also very selective towards isomerization. It is therefore possible, by judiciously choosing each of the functions to adjust the activity / selectivity couple of the catalyst. Conventional catalysts for catalytic hydrocracking are for the most part constituted by weakly acidic supports, such as silica-aluminas, for example. These systems are more particularly used to produce middle distillates of very good quality. Many of the hydrocracking market catalysts are based on Group VIII metal silica-alumina. These systems have a very good selectivity in middle distillates, and the products formed are of good quality. The disadvantage of all these silica-alumina catalyst systems is, as mentioned, their low activity. On the other hand, catalyst systems based on zeolite (in particular zeolite USY or beta) are very active for the hydrocracking reaction but are not very selective. In the field of hydrocracking and hydroisomerization of paraffinic fillers resulting from Fischer-Tropsch synthesis, various catalytic systems are thus proposed. The patent application US2007 / 0131586 A1 teaches the use of a catalyst based on nickel and silica-alumina. Patent EP 0 583 836 B2 describes a process for the preparation of hydrocarbon fuels, characterized in that the hydroconversion step of the feed is carried out on a catalyst containing no crystalline zeolite. The patent application WO2006 / 01912A2 describes a method for reducing the microporous volume of an amorphous solid in order to reduce the selectivity of the bifunctional hydrocracking catalyst resulting in the formation of unwanted light cracking products. US Patent Application 5,968,344 discloses a catalyst for the hydroisomerization of normal long paraffins, composed of a specific amorphous silica-alumina and one or more Group VIIIA metals. The patent application US2003 / 0173253A1 claims an amorphous catalyst for hydrocracking low or no sulfur-containing fillers such as the fillers resulting from the Fischer-Tropsch process, said catalyst consisting of a silica-alumina, a noble metal and a metal oxide of Group V, VI and VII transition. The application US 2005/0145541 A1 describes the use of a hydrocracking catalyst comprising both an amorphous aluminosilicate (silica-alumina) and a crystallized aluminosilicate (zeolite); the examples show that an excessively high proportion of USY zeolite with respect to the silica-alumina leads to a loss of selectivity but also of cold behavior in diesel fuel. On reading the prior art in the field of hydrocracking and hydroisomerization of paraffinic fillers resulting from Fischer-Tropsch synthesis, those skilled in the art would therefore be reluctant to use a catalyst based on zeolite Y .

Patent FR2834655 teaches the synthesis of zeolite / SiC composites and their use in catalysis. This patent does not teach the potential use of a zeolite Y / SiC composite in hydrocracking of

Fischer-Tropsch waxes and does not teach the implementation of post-specific treatments of the composite to adjust the acid level and / or the porosity of the zeolite of said composite. In attempting to develop hydrocracking and paraffinic charge hydroisomerization catalysts such as Fischer-Tropsch synthesis fillers, the Applicant has discovered that a process for producing middle distillates from a paraffinic feedstock produced by Fischer-Tropsch synthesis, using a hydrocracking and / or hydroisomerization catalyst comprising at least one hydro-dehydrogenating metal selected from the group consisting of Group VIB metals and Group VIII of the Periodic Table, taken alone or mixed and a composite support formed by a zeolite type Y and silicon carbide SiC, allowed to increase the yields of middle distillates produced.

OBJECT OF THE INVENTION The present invention thus relates to a process for the production of middle distillates. This process makes it possible to increase the amount of average distillates available by hydrocracking the heavier paraffinic compounds present in the outlet effluent of the Fischer-Tropsch unit, and which have boiling points higher than those of the kerosene cuts. and diesel, for example the 370 ° C + fraction. More specifically, the invention relates to a process for producing middle distillates from a paraffinic feedstock produced by Fischer-Tropsch synthesis employing a particular catalyst as defined in the description which follows.

The present invention relates to a method for producing middle distillates from a paraffinic feedstock produced by Fischer-Tropsch synthesis, using a hydrocracking and / or hydroisomerization catalyst comprising at least one hydro-dehydrogenating metal chosen from group consisting of Group VIB and Group VIII metals of the Periodic Table, alone or as a mixture, and a composite support formed by a Y-type zeolite and SiC silicon carbide, which process operates at a temperature of between 270 and 400 ° C, a pressure of between 1 and 9 MPa, an hourly volume velocity of between 0.5 and 5 h -1, a hydrogen flow rate adjusted to obtain a ratio of 400 to 1500 normal liters of hydrogen per liter charge.

DETAILED DESCRIPTION OF THE INVENTION The present invention relates to a method for producing middle distillates from a paraffinic feedstock produced by Fischer-Tropsch synthesis, using a hydrocracking and / or hydroisomerization catalyst comprising at least one metal hydro-dehydrogenating agent selected from the group consisting of Group VIB metals and Group VIII of the Periodic Table, taken alone or as a mixture and a composite support formed by, preferably comprising, and preferably consisting of, a Y-type zeolite and silicon carbide SiC, said process operating at a temperature between 270 and 400 ° C and preferably between 300 and 390 ° C, a pressure of between 1 and 9 MPa and 4

preferably between 2 and 8 MPa, a hourly space velocity of between 0.5 and 5 h 4 and preferably between 0.8 and 3 h -1, a hydrogen flow rate adjusted to obtain a ratio of 400 to 1500 normal liters. of hydrogen per liter of filler and preferably a ratio of 600 and 1300 normal liters of hydrogen per liter of filler.

Description of the Figures Figure 1 shows the implementation of the method according to embodiment a). Figure 2 shows the implementation of the method according to embodiment b). Figure 3 shows the implementation of the method according to embodiment c).

Figures 4 and 5 show the implementation of the method according to embodiment d). Figures 9, 10 and 11 show the implementation of the method according to embodiment e). Figure 6 shows the XRD diagram of the NaY / SiC composite support. Figure 7 shows a scanning electron micrograph of the NaY / SiC composite. Figure 8 shows the MAS NMR spectrum of 27A1 of the NaY / SiC composite.

The Hydrocracking and Hydroisomerization Catalyst Preferably, said hydrocracking and / or hydroisomerization catalyst used in the process according to the invention comprises a composite support comprising and preferably constituted as a percentage by weight relative to the total mass. of the catalyst: from 70 to 99.5% and preferably from 80 to 99%, preferably from 85 to 99% and very preferably from 90 to 99% of silicon carbide (SiC), the specific surface area of which is measured by the BET method is greater than 5 m 2 / g; From 0.5 to 30%, preferably from 1 to 20%, preferably from 1 to 15% and most preferably from 1 to 10% of a Y type zeolite; the percentages of hydro-dehydrogenating metal being expressed as a percentage by weight relative to the total mass of the catalyst.

a) The Hydro-Dehydrogenating Function According to the invention, the catalyst comprises at least one hydro-dehydrogenating metal selected from the group consisting of Group VIII metals and Group VIB metals, taken alone or as a mixture. Preferably, the group VIII elements are chosen from iron, cobalt, nickel, ruthenium, rhodium, palladium, osmium, iridium or platinum, taken alone or as a mixture. In the case where the elements of group VIII are chosen from the noble metals of group VIII, the elements of group VIII are advantageously chosen from platinum and palladium, taken alone or in admixture.

In the case where the elements of group VIII are chosen from non-noble metals of group VIII, the elements of group VIII are advantageously chosen from iron, cobalt and nickel and preferably cobalt and nickel, taken alone or in mixed. Preferably, the group VIB elements of the catalyst according to the present invention are selected from tungsten and molybdenum, alone or as a mixture. In the case where the hydrogenating function comprises a group VIII element and a group VIB element, the following combinations of metals are preferred: nickel-molybdenum, cobalt-molybdenum, iron-molybdenum, iron-tungsten, nickel-tungsten, cobalt- tungsten, and very preferably: nickel-molybdenum, cobalt-molybdenum, nickel-tungsten. It is also possible to use combinations of three metals such as for example nickel-cobalt-molybdenum. When a combination of Group VIB and Group VIII metals is used, the catalyst is then preferably used in a sulfurized form. In the case where the catalyst comprises at least one Group VIB metal in combination with at least one Group VIII non-noble metal, the Group VIB metal content is advantageously comprised, in oxide equivalent, of 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 non-noble metal content of 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 Group VIB metal in combination with at least one Group VIII non-noble metal, said catalyst may also advantageously comprise at least one doping element selected from the group consisting of silicon, boron and phosphorus, taken alone or as a mixture, the content of doping element being preferably between 0 and 20% by weight of oxide of the doping element, preferably between 0.1 and 15% by weight, very preferably preferred between 0.1 and 10% by weight and even more preferably between 0.5 and 6% by weight relative to the total mass of the catalyst. When the hydro-dehydrogenating element is a noble metal of group VIII, the catalyst preferably contains a noble metal content of between 0.01 and 10% by weight, even more preferably from 0.02 to 5% by weight relative to to the total mass of said catalyst.

The noble metal is preferably used in its reduced and non-sulphurized form. In the case where said hydrocracking and / or hydroisomerization catalyst used in the process according to the invention comprises at least one noble metal of group VIII, said noble metal contained in the catalyst must be reduced before use in the reaction. One of the preferred methods for conducting the metal reduction is the hydrogen treatment at a temperature of from 150 ° C to 650 ° C and a total pressure of from 0.1 to 25 MPa. For example, a reduction consists of a plateau at 120 ° C for one hour then a rise in temperature up to 450 ° C at a rate of 1 ° C / min and then a plateau of two hours at

450 ° C; throughout this reduction step, the flow rate of hydrogen is 1600 normal liters hydrogen / liter catalyst and the total pressure kept constant at 0.1 MPa. Note also that any ex-situ reduction method is suitable. It is also advantageously possible to employ a reduced, non-sulfurized nickel-based catalyst.

In this case, the nickel metal content in its oxide form is advantageously between 0.5 and 25% by weight relative to the finished catalyst. Preferably, the catalyst also contains, in addition to the reduced nickel, a group IB metal and preferably copper, or a group IVB metal and preferably tin in proportions such that the mass ratio of the group metal IB or IVB and nickel on the catalyst is advantageously between 0.03 and 1. b) The zeolite composite type Y / silicon carbide The silicon carbide The silicon carbide used in the invention advantageously has a specific surface area ( measured by the BET technique) greater than 5 m 2 / g, preferably between 5 and 300 m 2 / g, and more preferably between 10 and 250 m 2 / g. The pore volume of the support is advantageously between 0.20 cm3 / g and 1.0 cm3 / g, preferably between 0.25 cm3 / g and 0.80 cm3 / g, and more preferably between 0, 25 cm3 / g and 0.75 cm3 / g. BET surface area values were determined by nitrogen adsorption at the temperature of liquid nitrogen according to the BET method known to those skilled in the art.

The manufacture of supports of silicon carbide type that can be used in heterogeneous catalysis is already known and taught in patents such as, in particular, EP 0440569 B1 or US Pat. No. 6,184,178 B1, without this list being limiting. This type of support is industrially manufactured for example by SICAT Sarl.

Synthesis of the zeolite composite of Y type / silicon carbide The composite is obtained by germination and growth of zeolite crystallites on the surface of the already preformed silicon carbide, that is to say already in any form allowing the subsequent use of the prepared catalyst from this composite in an industrial catalytic unit. By way of example, the silicon carbide may be in the form of beads, extrudates or even foam. The silicon carbide may undergo a heat treatment, preferably a calcination, before the zeolitization step. The zeolitization of the silicon carbide can be carried out by any method known to those skilled in the art. For example, it can be zeolitized according to the conventional hydrothermal method: the silicon carbide is immersed in the liquid reaction medium which contains the sources of the elements of the zeolite framework, sources of mineralizing agent (OH "), mineral cations and / or or organic species and a solvent (usually water) .The reaction mixture and the solid are then introduced into a 7

autoclave, then brought to temperature for a time appropriate to the germination and growth of zeolite crystals on the surface of the silicon carbide. The zeolitization of silicon carbide can also be carried out by applying the so-called "Dry Gel Conversion" method. Dry impregnation of the silicon carbide is then carried out by the liquid reaction medium and the solid then undergoes hydrothermal treatment in an autoclave. After zeolitization, the material is then filtered (when the conventional hydrothermal method is used), washed extensively with distilled water, dried and optionally calcined. It can then undergo an ultrasonication step to remove zeolite crystals insufficiently attached to the surface of the silicon carbide. The composite thus obtained may optionally be zeolite again one or more times.

After drying and optional calcination, various characterizations such as X-ray diffraction, scanning electron microscopy or magic angle spinning NMR of aluminum 27 can be performed on the composite in order to highlight the presence of zeolite crystallites. The composite is then advantageously exchanged by at least one treatment with a solution of at least one ammonium salt such as, for example, ammonium nitrate or ammonium chloride, so as to obtain the ammonium form of zeolite Y which once calcined leads to the hydrogen form of said zeolite. After exchange, the atomic ratio Na / Al of the zeolite is generally advantageously less than 0.1 and preferably less than 0.05 and even more preferably less than 0.01. The exchange step can be carried out before or after the deposition of the hydrogenating phase, advantageously before. The composite may also advantageously undergo a dealumination step in order to adjust the level of acidity of the zeolite Y and / or to generate mesoporosity within the zeolite crystals. This dealumination step may be carried out by any method known to those skilled in the art such as hydrothermal treatment, chemical dealumination or a combination of the two methods (see for example "Hydrocracking science and technology" of J. Scherzer and AJ Gruia, Marcel Dekker Inc., 1996). This dealumination step can be carried out before or after the exchange and deposition steps of the hydrogenating phase, advantageously after the exchange step and before the deposition step of the hydrogenating phase. Advantageously, the dealumination step is adjusted to obtain an atomic ratio silicon lattice on aluminum network greater than 5, preferably between 10 and 60, very preferably between 10 and 40. The ratio Si / Al network is calculated from the correlation of Fichtner - Schmittler (H. Fichtner - Schmittler et al., Crystal Research and Technology, 19, 1984, K1-K3) according to the following formula: (Si / Al) network = [192 / (112 Where ao represents the zeolite mesh parameter (in Angstrom) as determined by X-ray diffraction by application of the ASTM D3942 method. 8

Advantageously, the dealumination step is adjusted so as to substantially eliminate from the solid the aluminas extracted from the crystalline lattice, so-called extra-lattice aluminas. This can be achieved by any method known to those skilled in the art. The percentage of extra-lattice aluminum is determined by NMR of aluminum, and corresponds to the percentage of octahedral aluminum (peak resonance at 0 ppm); preferably the composite has a percentage of extra-lattice aluminum less than 40%, and very preferably less than 20%.

c) Deposition of the hydro-dehydrogenating function The hydro-dehydrogenating function can advantageously be introduced at any stage of the preparation, very preferably after the zeolitization, exchange and dealumination steps. The preparation generally ends with calcination of the solid at a temperature of 250 to 600 ° C. In a preferred manner, the composite is impregnated with an aqueous solution. The impregnation of the composite is preferably carried out by the "dry" impregnation method well known to those skilled in the art. The impregnation may advantageously be carried out in a single step by a solution containing all the constituent elements of the hydrogenating function of the final catalyst. The hydro-dehydrogenating function may advantageously be introduced by one or more impregnation operations of the composite, with a solution containing at least one precursor of at least one oxide of at least one metal chosen from the group formed by the metals of the group. VIII and the Group VIB metals, optionally at least one precursor of at least one doping element chosen from boron, silicon and phosphorus and optionally at least one precursor of at least one oxide of a group IB element or group IVB in the case where the catalyst contains reduced nickel, the precursor (s) of at least one oxide of at least one metal of group VIII being preferably introduced after those of group VIB or at the same time as the latter, if the catalyst contains at least one Group VIB metal and at least one Group VIII metal.

In the case where the catalyst advantageously contains at least one element of group VIB, for example molybdenum, it is for example possible to impregnate the composite with a solution containing at least one group VIB element, to dry, to calcine. The impregnation of molybdenum may advantageously be facilitated by the addition of phosphoric acid in the ammonium paramolybdate solutions, which also makes it possible to introduce the phosphorus so as to promote the catalytic activity.

According to another preferred embodiment of the present invention, the deposition of the elements of group IVB or group IB can also be carried out simultaneously using, for example, a solution containing a tin salt or a copper salt. The following doping elements: boron and / or silicon and / or phosphorus can be introduced into the catalyst at any level of the preparation and according to any technique known to those skilled in the art.

A preferred method according to the invention consists in depositing the selected promoter element or elements, for example the boron-silicon pair, on the zeolite composite of type Y / calcined or non-calcined silicon carbide.

preferably calcined. For this purpose, an aqueous solution of at least one boron salt, such as ammonium biborate or ammonium pentaborate, is prepared in an alkaline medium and in the presence of hydrogen peroxide, and a so-called dry impregnation is carried out in which the pore volume of the precursor is filled with the solution containing, for example, boron. In the case where, for example, silicon is also deposited, for example a solution of a silicon-type silicon compound or a silicone oil emulsion is used. The element (s) promoter (s) chosen (s) in the group formed by silicon, boron and phosphorus can advantageously be introduced by one or more impregnation operations with excess solution on the calcined precursor. The boron source may advantageously be boric acid, preferably orthoboric acid H3BO3, ammonium biborate or pentaborate, boron oxide, boric esters. The 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 family of pyridine and quinolines and compounds of the pyrrole family. Boron may be introduced for example by a boric acid solution in a water / alcohol mixture. The preferred phosphorus source is orthophosphoric acid H 3 PO 4, but its salts and esters such as ammonium phosphates are also suitable. The phosphorus may 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 family of pyridine and quinolines and compounds of the pyrrole family. Many sources of silicon can advantageously be employed. Thus, it is possible to use ethyl orthosilicate Si (OEt) 4, siloxanes, polysiloxanes, silicones, silicone emulsions, halide silicates, such as ammonium fluorosilicate (NH4) 2SiF6 or fluorosilicate. sodium Na2SiF6. Silicomolybdic acid and its salts, silicotungstic acid and its salts can also be advantageously employed. Silicon may advantageously be added for example by impregnation of ethyl silicate in solution in a water / alcohol mixture. Silicon can be added, for example, by impregnating a silicon-type silicon compound or silicic acid suspended in water. The noble metals of group VIII of the catalyst of the present invention may advantageously be present in whole or in part in metallic and / or oxide form.

The noble element sources of group VIII which can advantageously be used are well known to those skilled in the art. For the noble metals halides are used, for example chlorides, nitrates, acids such as hexachloroplatinic acid, hydroxides, oxychlorides such as ammoniacal oxychloride ruthenium. It is also advantageous to use cationic complexes such as ammonium salts when it is desired to deposit the metal on the Y-type zeolite by cation exchange.

The sources of group VIB elements and for example 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. Oxides and ammonium salts such as ammonium molybdate, ammonium heptamolybdate and ammonium tungstate are preferably used. 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, use will be made of nitrates, sulphates, hydroxides, phosphates, halides, for example chlorides, bromides and fluorides, carboxylates, for example acetates and carbonates. Group IB source materials that can be used are well known to those skilled in the art. For example, among copper sources, Cu (NO 3) 2 copper nitrate can be used. The sources of Group IVB elements that can be used are well known to those skilled in the art.

For example, among tin sources tin chloride SnCl 2 can be used. The catalysts thus obtained are shaped in the form of grains of different shapes and sizes. They are generally used in the form of cylindrical or multi-lobed extrusions such as bilobed, trilobed, straight-lobed or twisted, but may optionally be manufactured and used in the form of crushed powders, tablets, rings, beads. , wheels. Other techniques than extrusion, such as pelletizing or coating, can advantageously be used.

The present invention therefore relates to a method for producing middle distillates from a paraffinic feedstock produced by Fischer-Tropsch synthesis, using a hydrocracking / hydroisomerization catalyst as described above, said process operating at a temperature between 270 and 400 ° C and preferably between 300 and 390 ° C, a pressure of between 1 and 9 MPa and preferably between 2 and 8 MPa, an hourly space velocity between 0.5 and 5 h-1 and preferably between 0.8 and 3 h -1, a flow rate of hydrogen adjusted to obtain a ratio of 400 to 1500 normal liters of hydrogen per liter of filler and preferably a ratio of 600 and 1300 normal liters of hydrogen per liter charge.

Said method can advantageously be implemented according to the following steps: a) a fractionation of the feedstock, b) a possible hydrotreatment of at least a portion of said feedstock from the fractionation c) a possible step of removal of at least part of the water and possibly CO, CO2, NH 3, H 2 S, d) a passage in the process according to the invention of at least part of said optionally hydrotreated fraction, the conversion on the catalyst according to the invention above describes dot products

boiling point greater than or equal to 370 ° C in products with boiling points below 370 ° C is greater than 40% by weight, e) distillation of the hydrocracked / hydroisomerized fraction to obtain middle distillates, and optionally recycling in step d) of the residual fraction boiling above said middle distillates. In the case where a hydrotreatment step is carried out upstream of said process according to the invention, the hydrotreatment catalysts used in said step are described in the different embodiments.

Embodiments According to the Invention Said method can advantageously be implemented according to the following different embodiments.

a) First Embodiment According to a preferred embodiment of the invention, the process comprises the following stages starting from a feedstock resulting from Fischer-Tropsch synthesis: a) separation of a single so-called heavy fraction at point d initial boiling range between 120 and 200 ° C, b) hydrotreating of at least a part of said heavy fraction, c) fractionation into at least 3 fractions: - at least one intermediate fraction having an initial boiling point Ti between 120 and 200 ° C., and a final boiling point T 2 greater than 300 ° C. and less than 410 ° C., at least one light fraction boiling below the intermediate fraction, at least one heavy fraction boiling above above the intermediate fraction, f) passing at least part of said intermediate fraction over a hydroisomerizing catalyst, g) passing at least part of said heavy fraction in the process according to the invention, f) distillation of the fractions hydrocracked / hydroisomerized to obtain middle distillates, and recycling the residual fraction boiling above said middle distillates in step (e) on the catalyst according to the invention treating the heavy fraction. The description of this embodiment will be made with reference to Figure 1 without Figure 1 limiting the interpretation.

Step (a) The effluent from the Fischer-Tropsch synthesis unit arriving via line 1 is fractionated (for example by distillation) in a separation means (2) in at least two fractions: at least one light fraction and a heavy fraction with an initial boiling point equal to a temperature of between 120 and 200 ° C. and preferably between 130 and 180 ° C. and even more preferably at a temperature of about 150 ° C.,

in other words the cutting point is between 120 and 200 ° C. The light fraction of Figure 1 exits through line (3) and the heavy fraction through line (4). This fractionation can be carried out by methods well known to those skilled in the art such as flash, distillation, etc. As a non-limiting example, the effluent from the Fischer-Tropsch synthesis unit will be subject to flash, decantation to remove water and distillation to obtain at least the two fractions described above. The light fraction is not treated according to the process of the invention but may for example constitute a good load for petrochemicals and more particularly for a steam cracking unit (5). The heavy fraction previously described is treated according to the process of the invention.

Step (b) At least part of said heavy fraction (step a) is allowed in the presence of hydrogen (line 6) in a zone (7) containing a hydrotreatment catalyst which aims to reduce the content of compounds olefinic and unsaturated as well as possibly decompose the oxygenates present in the fraction, as well as possibly break down any traces of sulfur and nitrogen compounds present in the heavy fraction. This hydrotreatment step is non-converting, that is to say that the conversion of the 370 ° C + fraction to 370 ° C fraction is preferably less than 20% by weight, preferably less than 10% by weight, and very preferably less than 5% by weight.

The catalysts used in this step (b) are hydrotreating catalysts which are non-crunchy or slightly crisp and comprise at least one metal of group VIII and / or group VI of the periodic table of elements. The catalyst preferably comprises at least one metal of the metal group formed by nickel, molybdenum, tungsten, cobalt, ruthenium, indium, palladium and platinum and comprises at least one support.

A combination of at least one Group VI metal (especially molybdenum or tungsten) and at least one Group VIII metal (especially cobalt and nickel) of the Periodic Table of Elements may be used. The non-noble group VIII metal concentration, when used, is from 0.01 to 15% by weight of equivalent relative to the finished catalyst and that of the Group VI metal (in particular molybdenum or tungsten). is from 5% to 30% by weight of oxide equivalent relative to the finished catalyst. When a combination of Group VI and Group VIII metals is used, the catalyst is then preferably used in a sulfurized form. Advantageously, at least one element selected from P, B, Si is deposited on the support. This catalyst may advantageously contain phosphorus; indeed, this compound provides two advantages to hydrotreatment catalysts: an ease of preparation, particularly in the impregnation of nickel and molybdenum solutions, and a better hydrogenation activity.

In a preferred catalyst, the total concentration of metals of groups VI and VIII, expressed as metal oxides, is between 5 and 40% by weight and preferably between 7 and 30% by weight and the weight ratio expressed as metal oxide. (or metals) of group VI on metal (or metals) of group VIII is between 1.25 and 20 and preferably between 2 and 10. Advantageously, if there is phosphorus, the concentration of phosphorus oxide P205 will be less than 15% by weight and preferably less than 10% by weight. It is also possible to use a catalyst containing boron and phosphorus; advantageously boron and phosphorus are promoter elements deposited on the support, and for example the catalyst according to patent EP297949. The sum of the amounts of boron and phosphorus, expressed respectively by weight of boron trioxide and phosphorus pentoxide, relative to the weight of support, is about 5 to 15% and the atomic ratio boron on phosphorus is about 1 at 2 and at least 40% of the total pore volume of the finished catalyst is contained in pores with an average diameter greater than 13 nanometers. Preferably, the amount of Group VI metal such as molybdenum or tungsten, is such that the phosphorus atomic ratio on Group VIB metal is about 0.5 to 1.5; the amounts of Group VIB metal and Group VIII metal, such as nickel or cobalt, are such that the atomic ratio of Group VIII metal to Group VIB metal is about 0.3 to 0.7. The quantity of Group VIB metal expressed as weight of metal relative to the weight of finished catalyst is approximately 2 to 30% and the amount of Group VIII metal expressed as weight of metal relative to the weight of finished catalyst is about 0.01 to 15%. Another particularly advantageous catalyst contains promoter silicon deposited on the support. An interesting catalyst contains BSi or PSi. Ni-sulfide catalysts on alumina, NiMo on alumina, NiMo on alumina doped with boron and phosphorus and NiMo on silica-alumina are also preferred. Advantageously, eta or gamma alumina will be chosen as support. In the case of the use of noble metals (platinum and / or palladium), the metal content is preferably between 0.05 and 3% by weight relative to the finished catalyst and preferably between 0.1 and 2% by weight. by weight of the finished catalyst. The noble metal is preferably used in its reduced and non-sulphurized form. It is also possible to use a reduced, non-sulfurized nickel catalyst. In this case the metal content in its oxide form is between 0.5 and 25% by weight relative to the finished catalyst. Preferably, the catalyst also contains a group IB metal such as copper, in proportions such that the mass ratio of the group IB metal and nickel on the catalyst is between 0.03 and 1. These metals are deposited on a support which is preferably an alumina, but which may also be boron oxide, magnesia, zirconia, titania, a clay or a combination of these oxides. These catalysts can be prepared by any method known to those skilled in the art or can be acquired from companies specializing in the manufacture and sale of catalysts.

In the hydrotreatment reactor (7), the feedstock is brought into contact with the catalyst in the presence of hydrogen at operating temperatures and pressures which make it possible to hydrogenate the olefins present in the feedstock. Preferably, the catalyst and the operating conditions chosen will also make it possible to carry out the hydrodeoxygenation, ie the decomposition of the oxygenated compounds (mainly alcohols) and / or the hydrodesulfurization or hydrodenitrogenation of the possible traces of sulfur compounds. and / or nitrogen present in the charge. The reaction temperatures used in the hydrotreatment reactor are between 100 and 400 ° C, preferably between 150 and 350 ° C, even more preferably between 150 and 300 ° C. The total pressure range used varies from 0.5 to 15 MPa, preferably from 1 to 10 MPa and even more preferably from 1 to 9 MPa.

The hydrogen which feeds the hydrotreatment reactor is introduced at a rate such that the volume ratio hydrogen / hydrocarbons is between 50 to 3000 normal liters per liter, preferably between 100 and 2000 normal liters per liter and even more preferably between 150 and 1500 normal liters per liter. The charge rate is such that the hourly space velocity is between 0.1 and 10 h -1, preferably between 0.2 and 5 h -1, and even more preferably between 0.2 and 3 h -1. Under these conditions, the content of unsaturated and oxygenated molecules is reduced to less than 0.5% by weight and to less than 0.1% by weight in general. The hydrotreatment step is conducted under conditions such that conversion to products having boiling points greater than or equal to 370 ° C to products having boiling points below 370 ° C is limited to 20% by weight. weight, preferably less than 10% by weight and even more preferably less than 5% by weight.

Step (c) The effluent from the hydrotreatment reactor is fed via line (8) into a fractionation zone (9) where it is fractionated into at least three fractions: at least one light fraction (exiting via line 10) ) whose compounds have boiling points below a temperature Ti of between 120 and 200 ° C, and preferably between 130 and 180 ° C and even more preferably at a temperature of about 150 ° C. In other words the cutting point is between 120 and 200 ° C; at least one intermediate fraction (line 11) comprising compounds whose boiling points are between the cut point Ti, previously defined, and a temperature T2 greater than 300 ° C, even more preferably greater than 350 ° C and less than 410 ° C or better at 370 ° C; at least one so-called heavy fraction (line 12) comprising the compounds having boiling points higher than the previously defined cutting point T2.

Step (dl

At least a portion of said intermediate fraction is then introduced (line 11), as well as possibly a stream of hydrogen (line 13) into the zone (14) containing a hydroisomerization catalyst. The operating conditions under which this step (d) is carried out are as follows. The pressure is maintained between 0.2 and 15 MPa and preferably between 0.5 and 10 MPa and advantageously from 1 to 9 MPa, the hourly space velocity is between 0.1 h-1 and 10 h-1 and preferably between 0.2 and 7 h -1 and advantageously between 0.5 and 5.0 h -1. The hydrogen flow rate is adjusted to obtain a ratio of 100 to 2000 normal liters of hydrogen per liter of feedstock and preferably between 150 and 1500 liters of hydrogen per liter of feedstock. The temperature used in this step is between 200 and 450 ° C. and preferably from 250 ° C. to 450 ° C., advantageously from 300 to 450 ° C., and even more advantageously above 320 ° C. or for example between 320 and 420 ° C. . The hydroisomerization step (d) is advantageously carried out under conditions such that the pass conversion to products with boiling points greater than or equal to 150 ° C. to products having boiling points below 150 ° C. is the lowest possible, preferably less than 50%, even more preferably less than 30%, and very preferably less than 15% by weight, and allows to obtain middle distillates (gas oil and kerosene) having properties cold (pour point and freezing) sufficiently good to meet the specifications in force for this type of fuel. Thus, in this step (d), it is sought to promote hydroisomerization rather than hydrocracking. The catalysts used are of the bifunctional type, that is to say that they have a hydro / dehydrogenating function and a hydroisomerizing function. The hydro / dehydrogenating function is generally provided either by noble metals (Pt and / or Pd) active in their reduced form or by non-noble metals of group VI (particularly molybdenum and tungsten) in combination with non-noble metals of Group VIII (particularly nickel and cobalt), preferably used in their sulfurized form. The hydroisomerizing function is provided by acidic solids, such as zeolites, halogenated alumina, agile with a pillar, heteropolyacids or sulphated zirconia. An alumina binder may also be used during the catalyst shaping step. The metal function can be introduced on the catalyst by any method known to those skilled in the art, such as for example comalaxing, dry impregnation, exchange impregnation.

In the case where the hydroisomerization catalyst comprises at least one noble metal of group VIII, the noble metal content is advantageously between 0.01 and 5% by weight relative to the finished catalyst, preferably between 0.1 and 4% by weight and very preferably between 0.2 and 2% by weight. Before use in the reaction, the noble metal contained in the catalyst must be reduced. One of the preferred methods for conducting the metal reduction is the hydrogen treatment at a temperature of from 150 ° C to 650 ° C and a total pressure of from 0.1 to 25 MPa. For example, a reduction consists of a plateau at 150 ° C for two hours and a rise in temperature

up to 450 ° C at a rate of 1 ° C / min and then a two-hour plateau at 450 ° C; throughout this reduction step, the hydrogen flow rate is 1000 normal liters hydrogen / liter catalyst and the total pressure kept constant at 0.1 MPa. Note also that any ex-situ reduction method is suitable. In the case where the hydroisomerization catalyst comprises at least one group VI metal in combination with at least one group VIII non-noble metal, the group VI metal content of the hydroisomerization catalyst is advantageously comprised, in oxide equivalent. between 5 and 40% by weight relative to the finished catalyst, preferably between 10 and 35% by weight and very preferably between 15 and 30% by weight and the group VIII metal content of said catalyst is advantageously included, in oxide equivalent, between 0.5 and 10% by weight relative to the finished catalyst, preferably between 1 and 8% by weight and very preferably between 1.5 and 6% by weight. Prior to use in the reaction, group VI and non-noble Group VIII metals must be sulfided. Any in-situ or ex-situ sulphurization method known to those skilled in the art is suitable. The hydro / dehydrogenating metal function can advantageously be introduced on said catalyst by any method known to those skilled in the art, such as, for example, comalaxing, dry impregnation, exchange impregnation. According to step (d) of hydroisomerization of the process according to the invention, the hydroisomerisation catalyst comprises at least one molecular sieve, preferably at least one zeolite molecular sieve and more preferably at least one zeolite molecular sieve. 10 MR one-dimensional as a hydroisomerizing function.

Zeolite molecular sieves are defined in the "Atlas of Zeolite Structure Types" classification, W. M Meier, D. H. Oison and Ch. Baerlocher, 5th revised edition, 2001, which Elsevier also refers to. Zeolites are classified according to the size of their pore openings or channels. One-dimensional 10 MR zeolite molecular sieves have pores or channels whose opening is defined by a ring of 10 oxygen atoms (10 MR aperture). The zeolite molecular sieve channels having a 10 MR aperture are advantageously unidirectional one-dimensional channels that open directly to the outside of said zeolite. The one-dimensional 10 MR zeolite molecular sieves present in said hydroisomerization catalyst advantageously comprise silicon and at least one element T selected from the group formed by aluminum, iron, gallium, phosphorus and boron, preferably aluminum. The Si / Al ratios of the zeolites described above are advantageously those obtained in the synthesis or obtained after post-synthesis dealumination treatments well known to those skilled in the art, such as and not limited to hydrothermal treatments. followed or not by acid attacks or even direct acid attacks by solutions of mineral or organic acids. They are preferably almost completely in acid form, that is to say that the atomic ratio between the monovalent compensation cation (for example sodium) and the element T inserted in the crystalline lattice of the solid is advantageously less than 0.1, preferably less than 0.05 and very preferably less than

0.01. Thus, the zeolites used in the composition of said selective hydroisomerization catalyst are advantageously calcined and exchanged by at least one treatment with a solution of at least one ammonium salt so as to obtain the ammonium form of the zeolites which, once calcined, lead to to the acid form of said zeolites.

The said one-dimensional 10 MR zeolite molecular sieve of said hydroisomerization catalyst is advantageously chosen from zeolite molecular sieves of structure type TON (chosen from ZSM-22 and NU-10, taken alone or as a mixture), FER (chosen from ZSM). -35 and ferrierite, alone or as a mixture), EUO (selected from EU-1 and ZSM-50, alone or as a mixture), SAPO-11 or zeolitic molecular sieves ZBM-30 or ZSM-48 , taken alone or in a mixture. Preferably, said one-dimensional 10 MR zeolite molecular sieve is chosen from zeolitic molecular sieves ZBM-30, NU-10 and ZSM-22, taken alone or as a mixture. Very preferably, said one-dimensional 10 MR zeolite molecular sieve is ZBM-30 synthesized with the organic structuring agent triethylenetetramine. Indeed, the use of said ZBM-30 produces much better results in terms of isomerization yield and activity than other zeolites and in particular ZSM-48.

The one-dimensional 10 MR zeolite molecular sieve content is advantageously between 5 and 95% by weight, preferably between 10 and 90% by weight, more preferably between 15 and 85% by weight and very preferably between 20 and 80% by weight. % by weight relative to the finished catalyst. The catalysts obtained are shaped in the form of grains of different shapes and sizes. They are generally used in the form of cylindrical or multi-lobed extrusions such as bilobed, trilobed, straight-lobed or twisted, but may optionally be manufactured and used in the form of crushed powders, tablets, rings, beads. , wheels. The shaping can be carried out with other matrices than alumina, such as, for example, magnesia, amorphous silica-aluminas, natural clays (kaolin, bentonite, sepiolite, attapulgite), silica, titanium, boron oxide, zirconia, aluminum phosphates, titanium phosphates, zirconium phosphates, coal and mixtures thereof. It is preferred to use matrices containing alumina, in all its forms known to those skilled in the art, and even more preferably aluminas, for example gamma-alumina. Other techniques than extrusion, such as pelletizing or coating, can be used.

Step (e) At least part of said heavy fraction is introduced via line (12) into a zone (15) where it is placed, in the presence of hydrogen (25), in contact with a catalyst used in the method according to the present invention and under the operating conditions of the process of the present invention to produce a cut middle distillates (kerosene and diesel) having good cold properties. The catalyst used in the zone (15) of step (e) to carry out the hydrocracking and hydroisomerization reactions of the heavy fraction is the catalyst defined in the first part of the patent application. During this step (e) the fraction entering the reactor is subjected to contact with the catalyst and

presence of hydrogen essentially hydrocracking reactions which, accompanied by hydroisomerization reactions of n-paraffins, will allow to improve the quality of the products formed and more particularly the cold properties of kerosene and diesel, and also of obtain very good yields of middle distillates. The conversion to products having boiling points greater than or equal to 370 ° C. in products with a boiling point below 370 ° C. is greater than 50% by weight, often at least 60% and preferably greater than or equal to at 70%.

Step (f) The effluents leaving the reactors (14) and (15) are sent via lines (16) and (17) to a distillation train, which incorporates an atmospheric distillation and optionally a vacuum distillation, and which has for the purpose of separating on the one hand the light products inevitably formed during steps (d) and (e) for example the gases (C1-C4) (line 18) and a gasoline section (line 19), and distilling at least a diesel cut (line 21) and kerosene (line 20). The gas oil and kerosene fractions can be recycled (line 23) partly, jointly or separately, at the top of the hydroisomerization reactor (14) of step (d). A fraction (line 22) boiling above the diesel fuel is also distilled, that is to say whose compounds which constitute it have boiling points higher than those of the middle distillates (kerosene and diesel). This fraction, called the residual fraction, generally has an initial boiling point of at least 350 ° C., preferably greater than 370 ° C. This fraction is advantageously recycled via line (22) at the top of the hydroisomerization and hydrocracking reactor (15) of the heavy fraction (step e). It may also be advantageous to recycle a portion of the kerosene and / or diesel fuel in step (d), step (e) or both. Preferably, at least one of the kerosene and / or diesel fractions is partially recycled in step (d) (zone 14). It has been found that it is advantageous to recycle a portion of the kerosene to improve its cold properties.

Advantageously and at the same time, the non-hydrocracked fraction is partially recycled in step (e) (zone 15). It goes without saying that the diesel and kerosene cuts are preferably recovered separately, but the cutting points are adjusted by the operator according to his needs. In Figure 1, there is shown a distillation column (24), but two columns can be used to separately treat the sections from areas (14) and (15). In Figure 1, there is shown only the recycling of kerosene on the reactor catalyst (14). It goes without saying that one can also recycle a portion of the gas oil (separately or with kerosene) and preferably on the same catalyst as kerosene. b) Second Embodiment Another embodiment of the invention comprises the following steps:

a) separating at least a light fraction of the feed so as to obtain a single so-called heavy fraction with an initial boiling point of between 120 and 200 ° C, b) hydrotreatment of the said heavy fraction, followed by a step, c) removal of at least a portion of the water and CO, CO2, NH 3, H 2 S, d) passing through the process according to the invention of at least a part of said optionally hydrotreated fraction, the conversion on the catalyst according to the invention described above products with boiling points greater than or equal to 370 ° C in products with boiling points below 370 ° C is greater than 40% by weight, e) distillation of the fraction hydrocracked / hydroisomerized to obtain middle distillates, and recycling in step d) of the residual fraction boiling above said middle distillates. The description of this embodiment will be made with reference to Figure 2 without Figure 2 limiting the interpretation.

Step (a) The effluent from the Fischer-Tropsch synthesis unit arriving via line 1 is fractionated (for example by distillation) in a separation means (2) in at least two fractions: at least one light fraction and an initial boiling point heavy fraction at a temperature between 120 and 200 ° C and preferably between 130 and 180 ° C and even more preferably at a temperature of about 150 ° C, in other words the cutting point is between 120 and 200 ° C. The light fraction of Figure 1 exits through line (3) and the heavy fraction through line (4). This fractionation can be carried out by methods well known to those skilled in the art such as flash, distillation, etc. As a non-limiting example, the effluent from the Fischer-Tropsch synthesis unit will be subject to flash, decantation to remove water and distillation to obtain at least the two fractions described above.

The light fraction is not treated according to the process of the invention but may for example constitute a good load for petrochemicals and more particularly for a steam cracking unit (5). The heavy fraction previously described is treated according to the process of the invention.

Step (b) This fraction is admitted in the presence of hydrogen (line 6) in a zone (7) containing a hydrotreatment catalyst which aims to reduce the content of olefinic and unsaturated compounds as well as to decompose the oxygenated compounds ( mainly alcohols) present in the heavy fraction described above, as well as to decompose any traces of sulfur and nitrogen compounds present in the heavy fraction. This hydrotreatment step is non-converting, that is to say that the conversion of the 370 ° C + fraction to 370 ° C fraction is preferably less than 20% by weight, preferably less than 10% by weight, and very preferably less than 5% by weight. 20

The catalysts used in this step (b) are hydrotreatment catalysts described in step (b) of the first embodiment. In the hydrotreatment reactor (7), the feedstock is contacted in the presence of hydrogen and the catalyst at operating temperatures and pressures to hydrogenate the olefins present in the feedstock. Preferably, the catalyst and the operating conditions chosen will also make it possible to carry out the hydrodeoxygenation, ie the decomposition of the oxygenated compounds (mainly alcohols) and / or the hydrodesulfurization or hydrodenitrogenation of the possible traces of sulfur compounds. and / or nitrogen present in the charge. The reaction temperatures used in the hydrotreatment reactor are between 100 and 400 ° C, preferably between 150 and 350 ° C, even more preferably between 150 and 300 ° C. The total pressure range used varies from 0.5 to 15 MPa, preferably from 1 to 10 MPa and even more preferably from 1 to 9 MPa. The hydrogen which feeds the hydrotreatment reactor is introduced at a rate such that the volume ratio hydrogen / hydrocarbons is between 50 to 3000 normal liters per liter, preferably between 100 and 2000 normal liters per liter and even more preferably between 150 and 1500 normal liters per liter. The charge rate is such that the hourly space velocity is between 0.1 and 10 h -1, preferably between 0.2 and 5 h -1, and even more preferably between 0.2 and 3 h -1. Under these conditions, the content of unsaturated and oxygenated molecules is reduced to less than 0.5% by weight and to less than 0.1% by weight in general. The hydrotreatment step is conducted under conditions such that conversion to products having boiling points greater than or equal to 370 ° C to products having boiling points below 370 ° C is limited to 20% by weight. weight, preferably less than 10% by weight and even more preferably less than 5% by weight.

Step (c) The effluent (line 8) from the hydrotreatment reactor (7) is then introduced into a water removal zone (9) which is intended to remove at least part of the produced water during hydrotreatment reactions. This removal of water can be carried out with or without elimination of the gaseous fraction C4- which is generally produced during the hydrotreating step. The elimination of water is understood to mean the elimination of the water produced by the oxygenation hydrodeoxygenation reactions, but it may also include the elimination at least partly of the water of saturation of the hydrocarbons. The elimination of water can be carried out by all the methods and techniques known to those skilled in the art, for example by drying, passage on a desiccant, flash, decantation ...

Step (d) The hydrotreated heavy fraction thus dried is then introduced (line 10) as well as optionally a stream of hydrogen (line 11) into the zone (12) containing the catalyst used in the process according to the invention. and under the operating conditions of the process of the present invention. Another possibility of

The process also according to the invention consists in sending all the effluent leaving the hydrotreating reactor (without drying) into the reactor containing the catalyst according to the invention and preferably at the same time as a stream of hydrogen. The catalyst used to carry out the hydrocracking and hydroisomerization reactions of the heavy fraction is the catalyst defined in the first part of the patent application. The operating conditions under which this step (d) is carried out are the operating conditions described according to the process according to the invention. The hydroisomerization and hydrocracking step is conducted under conditions such that the pass conversion to products having a boiling point greater than or equal to 370 ° C to products having boiling points below 370 ° C is greater than 40% by weight, and even more preferably at least 50%, preferably greater than 60%, so as to obtain middle distillates (gas oil and kerosene) having sufficiently good cold properties (point of flow, freezing point) to meet the specifications for this type of fuel.

Step (e) The effluent (so-called hydrocracked and hydroisomerized fraction) leaving the reactor (12), step (d), is sent to a distillation train (13), which incorporates an atmospheric distillation and optionally a vacuum distillation, which is intended to separate the conversion products having a boiling point of less than 340 ° C. and preferably less than 370 ° C. and including in particular those formed during step (d) in the reactor (12), and of separating the residual fraction whose initial boiling point is generally greater than at least 340 ° C and preferably greater than or equal to at least 370 ° C. Among the conversion products, in addition to the light gases C1-C4 (line 14), at least one gasoline fraction (line 15) and at least one middle distillate fraction kerosene (line 16) and diesel (line 17) are separated. The residual fraction, whose initial boiling point is generally greater than at least 340 ° C. and preferably greater than or equal to at least 370 ° C., is recycled (line 18) at the top of the hydroisomerization reactor (12). hydrocracking. It may also be advantageous to recycle (line 19) in step (d) (reactor 12) part of the kerosene and / or diesel fuel thus obtained.

c) Third Embodiment Another embodiment of the invention comprises the following steps: a) fractionation of the feedstock in at least three fractions: at least one intermediate fraction having an initial boiling point Ti of between 120 and 200 ° C, and a final boiling point T2 greater than 300 ° C and lower than 410 ° C, - at least one light fraction boiling below the intermediate fraction, - at least one heavy fraction boiling above the intermediate fraction, b) hydrotreatment of at least a portion of said intermediate fraction, then 22

c) removal of at least a portion of the water produced during the hydrotreatment reactions and optionally CO, CO2, NH3, H2S, d) passage of at least a portion of the hydrotreated fraction over a hydroisomerizing catalyst, e) passing through the process according to the invention of at least part of said heavy fraction with a product conversion of 370 ° C + to 370 ° C products greater than 40% by weight, f) distillation of at least a portion of the fractions hydrocracked / hydroisomerized to obtain middle distillates. The description of this embodiment will be made with reference to FIG. 3 without limiting the interpretation.

Step (a) The effluent from the Fischer-Tropsch synthesis unit comprises mainly paraffins, but also contains olefins and oxygenates such as alcohols. It also contains water, CO2, CO and unreacted hydrogen as well as light hydrocarbon compounds CI to C4 in the form of gases, or even possibly sulfur or nitrogen impurities. The effluent from the Fischer-Tropsch synthesis unit arriving via line (1) is fractionated in a fractionation zone (2) in at least three fractions: at least one light fraction (leaving via line 3) whose Component compounds have boiling points below a Ti temperature of between 120 and 200 ° C, and preferably between 130 and 180 ° C and even more preferably at a temperature of about 150 ° C. In other words, the cutting point is situated between 120 and 200 ° C., at least one intermediate fraction (line 4) comprising the compounds whose boiling points are between the cut point Ti, previously defined, and a T2 temperature greater than 300 ° C, even more preferably greater than 350 ° C and less than 410 ° C or more preferably 370 ° C, at least one so-called heavy fraction (line 5) comprising the compounds having boiling points greater than the previously defined cutting point T2. Cutting at 370 ° C makes it possible to separate at least 90% by weight of oxygenates and olefins, and most often at least 95% by weight. The heavy cut to be treated is then purified and removal of the heteroatoms or unsaturated by hydrotreating is then not necessary.

Fractionation is obtained here by distillation, but it can be carried out in one or more steps and by other means than distillation. This fractionation can be carried out by methods well known to those skilled in the art such as flash, distillation, etc. As a non-limiting example, the effluent from the Fischer-Tropsch synthesis unit will be subject to flash, decantation to remove water and distillation to obtain at least the three fractions described above. 23

The light fraction is not treated according to the process of the invention but may for example constitute a good load for a petrochemical unit and more particularly for a steam cracker (steam cracking unit 6). The heavier fractions previously described are treated according to the process of the invention. Step (b) Said intermediate fraction is admitted via line (4), in the presence of hydrogen supplied by the tubing (7), into a hydrotreatment zone (8) containing a hydrotreatment catalyst, the purpose of which is to reduce the content of olefinic and unsaturated compounds as well as to possibly decompose the oxygenated compounds (mainly alcohols) present in the intermediate fraction described above, as well as possibly to decompose any traces of sulfur and nitrogen compounds present in the intermediate fraction. This hydrotreating step is non-converting, that is to say that the conversion of the 150 ° C + fraction to 150 ° C fraction is preferably less than 20% by weight, preferably less than 10% by weight, and very preferably less than 5% by weight. The catalysts used in this step (b) are hydrotreatment catalysts described in step (b) of the first embodiment. In the hydrotreatment reactor (8), the feedstock is brought into contact with the catalyst in the presence of hydrogen and at operating temperatures and pressures which make it possible to hydrogenate the olefins present in the feedstock. Preferably, the catalyst and the operating conditions chosen will also make it possible to carry out the hydrodeoxygenation, that is to say the decomposition of the oxygenated compounds (mainly alcohols) and / or the hydrodesulfurization and / or the hydrodenitrogenation of the possible traces. of sulfur compounds and / or nitrogen compounds present in the load. The reaction temperatures used in the hydrotreatment reactor are between 100 and 400 ° C, preferably between 150 and 350 ° C, even more preferably between 150 and 300 ° C. The range of total pressure used varies between 0.5 and 15 MPa, preferably between 1 and 10 MPa and even more preferably between 1 and 9 MPa. The hydrogen that feeds the hydrotreatment reactor is introduced at a rate such that the volume ratio hydrogen / hydrocarbons is between 50 and 3000 normal liters per liter, preferably between 100 and 2000 normal liters per liter and even more preferably between 150 and 1500 normal liters per liter. The charge rate is such that the hourly volume velocity is between 0.1 and 10 h -1, preferably between 0.2 and 5 h -1, and even more preferably between 0.2 and 3 h -1. Under these conditions, the content of unsaturated and oxygenated molecules is reduced to less than 0.5% by weight and to less than 0.1% by weight in general. The hydrotreating step is conducted under conditions such that the conversion to products having boiling points greater than or equal to 150 ° C to products having boiling points below 150 ° C is limited to 20% by weight. weight, preferably less than 10% by weight and even more preferably less than 5% by weight. 24

Step (cI) The effluent from the hydrotreatment reactor is optionally introduced into a water removal zone (9) which is intended to remove at least a portion of the water produced during the hydrotreatment reactions. This removal of water can be carried out with or without elimination of the gaseous fraction C4- which is generally produced during the hydrotreatment stage.The term "removal of water" is understood to mean the elimination of the water produced by the oxygenation hydrodeoxygenation reactions, but it can also include the elimination at least partly of the water of saturation of the hydrocarbons.The elimination of the water can be carried out by all the known methods and techniques of the skilled in the art, for example by drying, passing on a desiccant, flash, decantation ....

Step (d) The fraction thus possibly dried is then introduced (line 10), as well as possibly a stream of hydrogen (line 11), in the zone (12) containing a hydroisomerizing catalyst. Another possibility of the process also according to the invention consists in sending all of the effluent leaving the hydrotreating reactor (without drying) into the reactor containing the hydroisomerizing catalyst and preferably at the same time as a stream of hydrogen. The hydroisomerizing catalysts are as described in step (d) of the first embodiment. The operating conditions under which this step (d) is carried out are as follows. The pressure is maintained between 0.2 and 15 MPa and preferably between 0.5 and 10 MPa and advantageously between 1 and 9 MPa, the hourly space velocity is between 0.1 h -1 and 10 h -1 and preferably between 0.2 and 7 h -1 and advantageously between 0.5 and 5.0 h -1 The flow rate of hydrogen is adjusted to obtain a ratio of 100 to 2000 normal liters of hydrogen per liter of filler and preferably between 150 and 1500 normal liters of hydrogen per liter of charge The temperature used in this stage is between 200 and 450 ° C. and preferably from 250 ° C. to 450 ° C., advantageously from 300 to 450 ° C., and even more advantageously greater than 320 ° C or, for example, between 320 and 420 ° C. The hydroisomerization step (d) is advantageously carried out under conditions such that the pass conversion into products with boiling points greater than or equal to 150 ° C. in products having boiling points below 150 ° C is the lowest possible, preferably less than 50% by weight, even more preferably less than 30%, and makes it possible to obtain middle distillates (gas oil and kerosene) having cold properties (pour point and freezing point) that are sufficiently good to meet the specifications in force for this type of fuel. Thus, in this step (d), it is sought to promote hydroisomerization rather than hydrocracking.

Step (e) Said heavy fraction whose boiling points are higher than the cut point T2, previously defined, is introduced via the line (5) into a zone (13) where it is placed, in the presence of hydrogen (26). ), at

contact of a catalyst according to the invention and under the operating conditions of the process of the present invention to produce a cut distillates medium (kerosene and diesel) having good properties cold. The catalyst used in the zone (13) of step (e) to carry out the hydrocracking and hydroisomerization reactions of the heavy fraction is the catalyst defined in the first part of the patent application. During this step (e) the fraction entering the reactor undergoes in contact with the catalyst and in the presence of hydrogen essentially hydrocracking reactions which, accompanied by hydroisomerization reactions of n-paraffins, will allow to improve the quality of the formed products and more particularly the cold properties of kerosene and diesel, and also to obtain very good yields of middle distillates. The conversion to products having boiling points greater than or equal to 370 ° C. in products with a boiling point of less than 370 ° C. is greater than 40% by weight, often at least 50% and preferably greater than or equal to at 60%. In this step (e), it will therefore seek to promote hydrocracking, but preferably by limiting the cracking of middle distillates.

The choice of operating conditions makes it possible to finely adjust the quality of the products (gas oil, kerosene) and in particular the cold properties of kerosene, while maintaining a good yield of diesel and / or kerosene. The process according to the invention makes it quite interesting to produce both kerosene and diesel fuel which are of good quality while minimizing the production of lighter, unwanted cuts (naphtha, LPG).

Step (f) The effluent leaving the reactor (12), step (d) is sent to a distillation train, which incorporates an atmospheric distillation and optionally a vacuum distillation, and which is intended to separate on the one hand the light products inevitably formed during step (d), for example the gases (C1-C4) (line 14) and a gasoline section (line 15), and distilling at least one gasoil fraction (line 17) and kerosene ( driving 16). The gas oil and kerosene fractions can be recycled (line 25) partly, jointly or separately, at the top of the hydroisomerization reactor (12) of step (d). The effluent at the outlet of step (e) is subjected to a separation step in a distillation train so as to separate, on the one hand, the light products that are inevitably formed during step (e), for example the gases (C1-C4) (line 18) and a petrol cut (line 19), to distil a diesel cut (line 21) and kerosene (line 20) and to distil the fraction (line 22) boiling over diesel fuel, that is, the compounds which constitute it have boiling points higher than those of middle distillates (kerosene + gas oil). This fraction, referred to as the residual fraction, generally has an initial boiling point of at least 350 ° C., preferably greater than 370 ° C. This non-hydrocracked fraction is advantageously recycled at the top of the hydroisomerization and hydrocracking reactor (13) of step (e). 26

It may also be advantageous to recycle a portion of the kerosene and / or diesel in step (d), step (f) or both. Preferably, at least one of the kerosene and / or diesel fractions is recycled in part (line 25) in step (d) (zone 12). It has been found that it is advantageous to recycle a portion of the kerosene to improve its cold properties.

Advantageously and at the same time, the non-hydrocracked fraction is partially recycled in step (f) (zone 13). It goes without saying that the diesel and kerosene cuts are preferably recovered separately, but the cutting points are adjusted by the operator according to his needs. In Figure 3, there is shown two columns (23) and (24) of distillation, but only one can be used to treat all cuts from areas (12) and (13). In Figure 3, there is shown only the recycling of kerosene on the reactor catalyst (12). It goes without saying that one can also recycle a portion of the gas oil (separately or with kerosene) and preferably on the same catalyst as kerosene. Some of the kerosene and / or diesel fuel produced in lines (20) and (21) can also be recycled. d) Fourth Embodiment Another embodiment of the invention comprises the following steps: a) optionally fractionation of the feedstock into at least one heavy fraction with an initial boiling point of between 120 and 200 ° C., and at least one a light fraction boiling below said heavy fraction, b) optional hydrotreatment of at least part of the feedstock or heavy fraction, optionally followed by c) removal of at least part of the water, d) passing at least a portion of the effluent or the fraction possibly hydrotreated in the process according to the invention on a first catalyst according to the invention, e) distillation of the hydroisomerized and hydrocracked effluent to obtain middle distillates (kerosene, gas oil) and a residual fraction boiling above middle distillates, f) passing at least a portion of said residual heavy fraction and / or a portion of said middle distillates in the process according to invention on a second catalyst according to the invention, and distillation of the resulting effluent to obtain middle distillates.

The description of this embodiment will be made with reference to FIGS. 4 and 5, without these figures limiting the interpretation.

Step (1a) When this step is carried out, the effluent from the Fischer-Tropsch synthesis unit is fractionated (for example by distillation) into at least two fractions: at least one light fraction and at least one heavy fraction at least. initial boiling point equal to a temperature between 120 and 200 ° C and 27

preferably between 130 and 180 ° C and even more preferably at a temperature of about 150 ° C, in other words the cutting point is between 120 and 200 ° C. The heavy fraction generally has paraffin contents of at least 50% by weight. This fractionation can be carried out by methods well known to those skilled in the art such as flash, distillation, etc. As a non-limiting example, the effluent from the Fischer-Tropsch synthesis unit will be subject to flash, decantation to remove water and distillation to obtain at least the two fractions described above. The light fraction is not treated according to the process of the invention but may for example constitute a good load for petrochemicals and more particularly for a steam cracking unit. At least one heavy fraction previously described is treated according to the method of the invention.

Step (b) Optionally, this fraction, or at least part of the initial charge, is admitted via line (1) in the presence of hydrogen (supplied via line (2)) to a zone (3) containing a catalytic converter. hydrotreatment which aims to reduce the content of olefinic and unsaturated compounds as well as possibly to decompose the oxygenated compounds (mainly alcohols) present in the heavy fraction described above, as well as possibly to decompose possible traces of sulfur and nitrogen compounds present in the heavy fraction. This hydrotreatment step is non-converting, that is to say that the conversion of the 370 ° C + fraction to 370 ° C fraction is preferably less than 20% by weight, preferably less than 10% by weight, and very preferably less than 5% by weight. The catalysts used in this step (b) are described in step (b) of the first embodiment. In the hydrotreating reactor (3), the feedstock is brought into contact with the catalyst in the presence of hydrogen and at operating temperatures and pressures which make it possible to hydrogenate the olefins present in the feedstock. Preferably, the catalyst and the operating conditions chosen will also make it possible to carry out the hydrodeoxygenation, that is to say the decomposition of the oxygenated compounds (mainly alcohols) and / or the hydrodesulfurization and / or the hydrodenitrogenation of the possible traces of sulfur compounds and / or nitrogen compounds present in the load. The reaction temperatures used in the hydrotreatment reactor are between 100 and 400 ° C, preferably between 150 and 350 ° C, even more preferably between 150 and 300 ° C. The total pressure range used varies from 0.5 to 15 MPa, preferably from 1 to 10 MPa and even more preferably from 1 to 9 MPa. The hydrogen that feeds the hydrotreatment reactor is introduced at a rate such that the volume ratio hydrogen / hydrocarbons is between 50 and 3000 normal liters per liter, preferably between 100 and 2000 normal liters per liter and even more preferably between 150 and 1500 normal liters per liter. The charge rate is such that the hourly volume velocity is between 0.1 and 10 h -1, preferably between 0.2 and 5 h -1 and even more preferably between 0.2 and 3 h -1. In these circumstances, the content of

Unsaturated and oxygenated molecules are reduced to less than 0.5% by weight and to less than 0.1% by weight in general. The hydrotreatment step is conducted under conditions such that the conversion to products having boiling points greater than or equal to 370 ° C to products having boiling points below 370 ° C is limited to 20% by weight. weight, preferably less than 10% by weight and even more preferably less than 5% by weight.

Step (c) The effluent (line 4) from the hydrotreatment reactor (3) is optionally introduced into a zone (5) of water removal which aims to eliminate at least partly the water produced during hydrotreatment reactions. This removal of water can be carried out with or without elimination of the gaseous fraction C4- which is generally produced during the hydrotreating step. The elimination of water is understood to mean the elimination of the water produced by the oxygenation hydrodeoxygenation reactions, but it may also include the elimination at least partly of the hydrocarbon saturation water. The removal of water can be carried out by all the methods and techniques known to those skilled in the art, for example by drying, passage on a desiccant, flash, decantation ....

Step (d) At least part and preferably all of the hydrocarbon fraction (at least part of the feed or at least part of the heavy fraction of step a) or at least part of the fraction or the hydrotreated and possibly dried feedstock) is then introduced (line 6) as well as possibly a stream of hydrogen (line 7) into the zone (8) containing the catalyst according to the invention. Another possibility of the process also according to the invention consists in sending part or all of the effluent leaving the hydrotreating reactor (without drying) into the reactor containing the catalyst according to the invention and preferably at the same time as a flow of hydrogen.

Step (e) The hydroisomerized and hydrocracked effluent leaving the reactor (8), step (d), is sent to a distillation train (9) which incorporates an atmospheric distillation and optionally a vacuum distillation which is intended to separate conversion products having a boiling point of less than 340 ° C. and preferably less than 370 ° C. and including in particular those formed during step (d) in the reactor (8), and of separating the residual fraction of which the initial boiling point is generally greater than at least 340 ° C and preferably greater than or equal to at least 370 ° C. Among the conversion products and hydroisomerized it is separated, in addition to the light gases C1-C4 (line 10) at least one gasoline fraction (line 11), and at least one middle distillates fraction kerosene (line 12) and diesel (line 13) .

Step (f) 29

The process according to the invention uses a second zone (16) containing a hydrocracking and hydroisomerization catalyst described in the first part of the patent. It passes on this catalyst, in the presence of hydrogen (line 15), an effluent chosen from a part of the product kerosene (line 12), a part of the gas oil (line 13) and the residual fraction and preferably the residual fraction of which the initial boiling point is generally greater than at least 370 ° C. During this step, the fraction entering the reactor (16) undergoes, in the presence of hydrogen, hydroisomerization and / or hydrocracking reactions which will make it possible to improve the quality of the products formed and more particularly the properties cold kerosene and diesel, and obtain improved middle distillate yields over the prior art.

The choice of operating conditions makes it possible to finely adjust the quality of the products (middle distillates) and in particular the cold properties. The operating conditions in which this step (f) is carried out are the operating conditions in accordance with the process according to the invention. The operator will adjust the operating conditions on the first and second hydrocracking and hydroisomerization catalyst so as to obtain the desired product qualities and yields. Thus, generally, on the first catalyst, the pass conversion to products with boiling points greater than or equal to 150 ° C in products with boiling points below 150 ° C is less than 50% by weight, preferably less than 30% by weight. These conditions make it possible in particular to adjust the kerosene / diesel ratio produced as well as the cold properties of middle distillates, and more particularly of kerosene. Also generally, on the second catalyst, when the residual fraction is treated, the conversion per pass to products with boiling points greater than or equal to 370 ° C in products with boiling points below 370 ° C is greater 40% by weight, preferably greater than 50%, more preferably 60%. It may even be advantageous to have conversions of at least 80% by weight.

When a part of the kerosene and / or diesel fuel is treated on the second catalyst, the pass conversion into products with boiling points greater than or equal to 150 ° C in products with boiling points below 150 ° C is lower. at 50% by weight, preferably less than 30%. In general, the operating conditions applied in the reactors (8) and (16) may be different or identical. Preferably, the operating conditions used in the two hydroisomerization and hydrocracking reactors are chosen to be different in terms of operating pressure, temperature, hourly volume rate and H2 / filler ratio. This embodiment allows the operator to adjust the qualities and / or yields of kerosene and diesel. The effluent from the reactor (16) is then sent via line (17) in the distillation train so as to separate the conversion products, gasoline, kerosene and diesel.

FIG. 4 shows an embodiment with the residual fraction (line 14) passing through the hydroisomerization and hydrocracking zone (16) (step f), the effluent obtained being sent (line 17) in the zone (9) of separation. Advantageously, at the same time, the kerosene and / or diesel can be partly recycled (line 18) in the zone (8) of hydroisomerization and hydrocracking (step d) on the first catalyst. In FIG. 5, a portion of the kerosene and / or diesel fuel produced passes into the hydroisomerization and hydrocracking zone (16) (step f), the effluent obtained being sent (line 17) to the zone (9). of seperation. At the same time, the residual fraction (line 14) is recycled to the hydroisomerization and hydrocracking zone (8) (step d) on the first catalyst. It has been found that it is advantageous to recycle a portion of the kerosene on a hydrocracking and hydroisomerization catalyst to improve its cold properties. In the figures, only the recycling of kerosene has been shown. It goes without saying that one can also recycle a portion of the gas oil (separately or with kerosene) and preferably on the same catalyst as kerosene.

d) Fifth Embodiment Another embodiment of the invention comprises the following steps: a) separation of at least one C4- gas fraction, called light, with a final boiling point of less than 20 ° C., of effluent from the Fischer Tropsch synthesis unit so as to obtain a single heavy C5 + liquid fraction with an initial boiling point of between 20 and 40 ° C, b) hydrogenation of the unsaturated compounds of the olefinic type from at least a portion of said C5 + heavy fraction, in the presence of hydrogen and a hydrogenation catalyst at a temperature between 80 ° C and 200 ° C, at a total pressure of between 0.5 and 6 MPa, at a hourly volume velocity between 1 and 10 hl, and at a hydrogen flow rate corresponding to a volume ratio hydrogen / hydrocarbons of between 5 and 80 normal liters of hydrogen per liter of feed, c) passage of the entire effluent hydrogenated liquid from step b) without prior separation step, in the process according to the invention in the presence of hydrogen and a catalyst according to the invention, d) distillation of the hydrocracked / hydroisomerized effluent.

Step (a) Step a), not shown in FIG. 9, is a step of separation of at least one so-called light C4 fraction with a final boiling point of less than 20 ° C., preferably less than 20 ° C. 10 ° C and very preferably, below 0 ° C, the effluent from Fischer-Tropsch synthesis so as to obtain a single so-called heavy C5 + fraction, initial boiling point between 20 and 40 ° C and preferably having a temperature

boiling point greater than or equal to 30 ° C, constituting at least part of the feedstock of the hydrogenation stage b) according to the invention. The effluent from the Fischer-Tropsch synthesis unit is at the outlet of the Fischer-Tropsch synthesis unit advantageously divided into two fractions, a light fraction, called cold condensate, and a heavy fraction, called waxes. The two fractions thus defined comprise water, carbon dioxide (CO2), carbon monoxide (CO) and unreacted hydrogen (H2). In addition, the light fraction, cold condensate, contains light hydrocarbon compounds C, to C4, called C4- fraction, in the form of gas. According to a preferred embodiment shown in FIG. 10, the light fraction, called cold condensate (1), and the heavy fraction, called waxes (3), are treated separately in separate fractionation means and then recombined in the line ( 5), so as to obtain a single C5 + fraction with an initial boiling point of between 20 and 40 ° C and preferably having a boiling point greater than or equal to 30 ° C. The heavy fraction, called waxes, enters a fractionation means (4) via line (3). The fractionation means (4) may for example consist of methods well known to those skilled in the art such as rapid expansion (or flash, according to the English terminology), distillation or stripping. Advantageously, a flash or flash balloon or a stripper is sufficient to remove most of the water, carbon dioxide (CO2) and carbon monoxide (CO) through the line (4 ') of the heavy fraction, called waxes. The light fraction, called cold condensate, enters a fractionation means (2) via the pipe (1). The fractionation means (2) may for example consist of methods well known to those skilled in the art such as a flash or flash tank, a distillation or a stripping. Advantageously, the fractionation means (2) is a distillation column allowing the elimination of light and gaseous hydrocarbon compounds CI to C4, called gas fraction C4-, corresponding to products boiling at a temperature below 20 ° C, preferably below at 10 ° C and very preferably, below 0 ° C, through the pipe (2 '). The stabilized effluents from the fractionation means (2) and (4) are then recombined in the pipe (5). A stabilized C5 + liquid fraction, corresponding to products boiling at an initial boiling point of between 20 and 40 ° C. and preferably having a boiling point greater than or equal to 30 ° C., is thus recovered in the pipe (5) and constitutes the charge of the hydrogenation step b) of the process according to the invention. According to another preferred embodiment shown in FIG. 11, the light fraction, called cold condensate, leaving the Fischer-Tropsch synthesis unit via line (1) and the heavy fraction, called waxes, coming out of the Fischer-Tropsch synthesis unit via the pipe (3), are recombined in the pipe (18) and treated in the same fractionation means (4). The fractionation means (4) may for example consist of methods well known to those skilled in the art such as flash, distillation or stripping. Advantageously, the fractionation means (4) is a distillation column 32

allowing the removal of the gas fraction C4-, water, carbon dioxide (CO2) and carbon monoxide (CO) through the pipe (4 '). A stabilized C5 + liquid fraction, corresponding to the products boiling at a boiling point of between 20 and 40 ° C. and preferably having a boiling point greater than or equal to 30 ° C., is thus recovered at the outlet of the fractionation means (4). ) in the pipe (5) and constitutes the charge of the hydrogenation step b) of the process according to the invention.

Step (bl Step b) is a step of hydrogenation of the olefinic type unsaturated compounds of at least a portion and preferably all of the C5 + heavy liquid fraction from step a) of the process according to in the presence of hydrogen and a hydrogenation catalyst. Preferably, the catalyst used in step (b) is a non-cracking or slightly cracking hydrogenation catalyst comprising at least one metal of group VIII of the periodic table and comprising at least one oxide-based support. refractory.

Preferably, said catalyst comprises at least one group VIII metal chosen from nickel, molybdenum, tungsten, cobalt, ruthenium, indium, palladium and platinum and comprising at least one oxide-based support refractory selected from alumina and silica alumina. Preferably, the group VIII metal is chosen from nickel, palladium and platinum. According to a preferred embodiment of step b) of the process according to the invention, the group VIII metal is chosen from palladium and / or platinum and the content of this metal is advantageously between 0.1% and 5%. % by weight, and preferably between 0.2% and 0.6% by weight relative to the total weight of the catalyst. According to a very preferred embodiment of step b) of the process according to the invention, the Group VIII metal is palladium. According to another preferred embodiment of step b) of the process according to the invention, the metal of group VIII is nickel and the content of this metal is advantageously between 5% and 25% by weight, preferably between 7%. and 20% by weight based on the total weight of the catalyst. The catalyst support used in step (b) of the process according to the invention is a refractory oxide-based support, preferably chosen from alumina and silica-alumina. When the support is an alumina, it has a BET specific surface to limit the polymerization reactions on the surface of the hydrogenation catalyst, said surface being between 5 and 140 m 2 / g. When the support is a silica-alumina, the support contains a percentage of silica of between 5 and 95% by weight, preferably between 10 and 80%, more preferably between 20 and 60% and very preferably between 30 and 50%. a BET specific surface area of between 100 and 550 m 2 / g, preferably between 150 and 500 m 2 / g, preferably less than 350 m 2 / g and even more preferably less than 250 m 2 / g.

The hydrogenation step b) is preferably carried out in one or more fixed bed reactor (s). In the hydrogenation zone (7), the feedstock is brought into contact with the hydrogenation catalyst in the presence of hydrogen and at operating temperatures and pressures allowing the hydrogenation of the olefinic unsaturated compounds present in the feedstock. Under these operating conditions, the oxygenated compounds are not converted, the liquid hydrogen effluent from step b) of the process according to the invention therefore does not contain water resulting from the conversion of said oxygenated compounds. The operating conditions of the hydrogenation step b) are chosen so that the effluent leaving said hydrogenation zone (7) is in the liquid state: indeed, the quantity of hydrogen introduced into the zone hydrogenation (7) via line (6) corresponds to a quantity of hydrogen in slight excess with respect to the amount of hydrogen strictly necessary to carry out the hydrogenation reaction of the olefinic type unsaturated compounds. Thus, no cracking is carried out in the hydrogenation zone (7), and the liquid hydrogenated effluent does not contain hydrocarbon compounds boiling at a temperature below 20 ° C., preferably below 10 ° C. and very preferred, lower than 0 ° C, corresponding to the gaseous fraction C4-.

The operating conditions of the hydrogenation step b) are as follows: the temperature within said hydrogenation zone (7) is between 80 ° C. and 200 ° C., preferably between 100 ° C. and 180 ° C., and preferably, between 120 and 165 ° C, the total pressure is between 0.5 and 6 MPa, preferably between 1 and 5 MPa and even more preferably between 2 and 5 MPa. The charge flow rate is such that the hourly volume velocity (ratio of the hourly volume flow rate at 15 ° C. of fresh liquid charge to the volume of catalyst charged) is between 1 and 10 hours, preferably between 1 and 5 hours. and even more preferably between 1 and 4 hours. The hydrogen which feeds the hydrotreating zone is introduced at a rate such that the volume ratio hydrogen / hydrocarbons is between 5 and 80 normal liters of hydrogen per liter of filler, preferably between 5 and 60, more preferably between 10 and 50, and even more preferably between 15 and 35.

Under these conditions, the olefinic type unsaturated compounds are hydrogenated more than 50%, preferably more than 75% and preferably more than 85%. The hydrogenation step b) is preferably conducted under conditions such that the conversion to products having boiling points greater than or equal to 370 ° C to products having boiling points below 370 ° C is zero. . The hydrogenated effluent from step b) of the process according to the invention therefore does not contain compounds boiling at a temperature below 20 ° C., preferably below 10 ° C. and very preferably below 0 ° C. C, corresponding to the gas fraction C4-. According to a preferred embodiment of step b), a guard bed (not shown in the figures) containing at least one guard bed catalyst upstream of the hydrogenation zone (7) is used in order to reduce the solid mineral particle content and possibly reduce the content of harmful metal compounds for hydrogenation catalysts. The guard bed can advantageously be integrated in

the hydrogenation zone (7) upstream of the hydrogenation catalyst bed is placed in a separate zone upstream of the hydrogenation zone (7). Indeed, the treated fractions may optionally contain solid particles such as inorganic solids. They may optionally contain metals contained in hydrocarbon structures such as more or less soluble organometallic compounds. By the term fines is meant fines resulting from a physical or chemical attrition of the catalyst. They can be micron or submicron. These mineral particles then contain the active components of these catalysts without the following list being limiting: alumina, silica, titanium, zirconia, cobalt oxide, iron oxide, tungsten, rhuthenium oxide, etc. These solid minerals may be present under calcined mixed oxide form: for example, alumina-cobalt, alumina-iron, alumina-silica, alumina-zirconia, alumina-titanium, alumina-silicobobalt, alumina-zirconia-cobalt, .... They may also contain metals within hydrocarbon structures, possibly containing oxygen or more or less soluble organometallic compounds. More particularly, these compounds may be based on silicon. It may be, for example, anti-foaming agents used in the synthesis process. On the other hand, the catalyst fines described above may have a higher silica content than the catalyst formulation resulting from the intimate interaction between the catalyst fines and anti-foaming agents described above. The guard bed catalysts used may advantageously be in the form of spheres or extrudates. It is however advantageous that the catalyst is in the form of extrudates with a diameter of between 0.5 and 5 mm and more particularly between 0.7 and 2.5 mm. The shapes are cylindrical (which can be hollow or not), cylindrical twisted, multilobed (2, 3, 4 or 5 lobes for example), rings. The cylindrical shape is preferably used, but any other shape may be used. In order to overcome the presence of contaminants and / or poisons in the feed, the bed bed catalysts may, in another preferred embodiment, have more specific geometric shapes in order to increase their void fraction. The void fraction of these catalysts is between 0.2 and 0.75. Their outer diameter can vary between 1 and 35 mm. Among the particular forms possible without this list being exhaustive: hollow cylinders, hollow rings, Raschig rings, serrated hollow cylinders, crenellated hollow cylinders, pentaring carts, multi-hole cylinders ...

Preferably, said guard bed catalysts used are not impregnated with an active phase. The guard beds can be marketed by Norton-Saint-Gobain, for example the MacroTrap® guard beds. Guard beds can be marketed by Axens in the ACT family: ACT077, ACT935, ACT961 or HMC841, HMC845, HMC941 or HMC945. It may be particularly advantageous to superpose these catalysts in at least two different beds of variable height. The catalysts having the highest void content are preferably used in the first catalytic bed or first catalytic reactor inlet. It can also be advantageous to use at least two different reactors for

these catalysts. These guard bed catalysts used may advantageously have macroporosity. In a preferred embodiment, the macroporous volume for a mean diameter at 50 nm is greater than 0.1 cm3 / g and a total volume greater than 0.60 cm3 / g. In another embodiment, the mercury volume for a pore diameter greater than 1 micron is greater than 0.5 cm3 / g and the mercury volume for a pore diameter greater than 10 microns is greater than 0.25 cm3 / g . These two embodiments can advantageously be associated in a mixed bed or a combined bed. The preferred guard beds according to the invention are HMC and ACT961. After passing over the guard bed, the solids content is advantageously less than 20 ppm, preferably less than 10 ppm and even more preferably less than 5 ppm. The soluble silicon content is advantageously less than 5 ppm, preferably less than 2 ppm and even more preferably less than 1 ppm. At the end of step b), all of the liquid hydrogen effluent is directly sent to a hydrocracking / hydroisomerization zone (10).

Step (c) In accordance with step c), all the liquid hydrogenated effluent from step b) is sent directly, without prior separation step, to the hydroisomerization / hydrocracking process (10) according to US Pat. invention containing the hydroisomerization / hydrocracking catalyst described in the first part of the patent application and preferably together with a stream of hydrogen (line 9).

The operating conditions in which the hydroisomerization / hydrocracking step (c) is carried out are the operating conditions described according to the process according to the invention.

Step (d) The effluent (so-called hydrocracked / hydroisomerized fraction) leaving the hydroisomerization / hydrocracking zone (10) from step (c) is sent, according to step d), in a train distillation system (11), which incorporates atmospheric distillation and optionally vacuum distillation, the purpose of which is to separate the conversion products having a boiling point of less than 340 ° C. and preferably less than 370 ° C., and especially including those formed in step (c) in the hydroisomerization / hydrocracking reactor (10), and to separate the residual fraction whose initial boiling point is generally greater than at least 340 ° C and preferably greater than or equal to at least 370 ° C. Among the conversion products and hydroisomerized, it is separated in addition to the light gases C1-C4 (line 12) at least one gasoline fraction (or naphtha) (line 13), and at least one middle distillate fraction kerosene (line 14) and diesel (line 15). Preferably, the residual fraction, whose initial boiling point is generally greater than at least 340 ° C. and preferably greater than or equal to at least 370 ° C., is recycled (line 16) in step c) at the top of the zone (10) for hydroisomerization and hydrocracking. It may also be advantageous to recycle (line 17) at least in part and preferably in whole,

in step (c) (zone 10) at least one of the kerosene and diesel fuel cuts thus obtained. The gas oil and kerosene cuts are preferably recovered separately or mixed, but the cutting points are adjusted by the operator according to his needs. It has been found that it is advantageous to recycle a portion of the kerosene to improve its cold properties.

The invention is not limited to these five embodiments.

The products obtained The obtained gas oil (s) has a pour point of at most 0 ° C, generally below -10 ° C and often below -15 ° C. The cetane number is greater than 60, generally greater than 65, often greater than 70. The resulting kerosene (s) have a freezing point of not more than -35 ° C, generally less than -40 ° C. The smoke point is greater than 25 mm, usually greater than 30 mm. In this process, the production of gasoline (not sought) is as low as possible. The gasoline yield will always be less than 50% by weight, preferably less than 40% by weight, advantageously less than 30% by weight or even 20% by weight or even 15% by weight.

EXAMPLE 1 Preparation of the Hydrotreating Catalyst (Cl) The catalyst is an industrial catalyst based on palladium on alumina noble metal with a palladium content of 0.3% by weight relative to the total weight of the finished catalyst, provided by the company 20 AXENS.

EXAMPLE 2 Preparation of a Hydroisomerization and Hydrocracking Catalyst According to the Invention (C2) a) Preparation of the Y-type / Silicon Carbide Composite Silicon Carbide Silicon Carbide is supplied by the Company SICAT, in the form of cylindrical extrusions with a diameter of 1 mm. Its main characteristics are provided in Table 1 below. Table 1: main characteristics of the silicon carbide used. Crystallized phase (DRX) f3-SiC BET surface area (N2 porosimetry) 25 m2 / g microporous volume (N2 porosimetry) no median pore diameter (Hg porosimetry) 40 nm total pore volume (Hg porosimetry) 0.35 cm 2 / g crush in Shell type bed (Shell SMS standard> 3 MPa 1471-74) impurities Fe <0.05% by weight, Al <0.1% by weight, Ca <0.01 ° A, weight 37

Synthesis of the zeolite composite of Y / SiC type. 100 grams of silicon carbide are calcined in air in a muffle furnace in a thin layer at 900 ° C for 2 hours. The zeolitization of the support is then carried out by hydrothermal treatment. The preparation of the gel used for the synthesis of the Y-type zeolite was carried out in three stages, using as aluminum source sodium aluminate (50-56 wt.% Al 2 O 3, 40-45 wt.% Na 2 O , Sigma-Aldrich) and as a silicon source a solution of sodium silicate (27% by weight SiO 2, 14% by weight of NaOH, Sigma-Aldrich). 20.4 grams of sodium hydroxide and 10.6 grams of sodium aluminate are first dissolved in 100 ml of distilled water at room temperature; after complete dissolution, 113.6 grams of sodium silicate are added. The mixture is then stirred until the appearance of a dense gel and then allowed to mature for 24 hours. In parallel a second gel is prepared by dissolving 0.7 grams of sodium hydroxide and 67.4 grams of sodium aluminate in 655 ml of distilled water; after complete dissolution, 712.2 grams of sodium silicate are added. The mixture is then stirred vigorously until a dense gel appears. The final gel is obtained by adding this second gel to 82.6 grams of the first gel with vigorous stirring for 20 minutes in order to obtain a homogeneous final gel. Finally this final gel is mixed with 100 grams of silicon carbide in a Teflon autoclave made of stainless steel. The assembly is heated in a muffle furnace at 100 ° C. under autogenous pressure for 22 hours in order to crystallize the zeolite. The mixture thus obtained is composed of bulk crystals of NaY zeolite and NaY / SiC composite extrudates which are easily recovered by sieving. The extrudates are then rinsed with distilled water until the pH of the wash water is less than 9, and then undergo a 30-minute ultrasonication step to remove the NaY crystals not or poorly fixed on the surface. silicon carbide. The extrudates are finally dried in an oven at 110 ° C. for 24 hours. After drying, different characterizations are performed on the NaY / SiC composite. X-ray diffraction shows the presence of NaY crystallites, which is confirmed by the scanning electron microscopy characterization of the solid, which shows the formation of crystallites of about 400 nm on the surface of silicon carbide extrusions (Figure 7). The composite was also analyzed by magic angle spinning NMR (MAS) of aluminum 27, in MQMAS probe 4mm on a Bruker Avance 400 spectrometer with a sample rotation speed of 14 kHz: only one peak corresponding to the aluminum + IV engaged in the zeolite structure is observed (FIG. 8), which confirms that the crystallites observed by scanning electron microscopy are crystallites of NaY. The presence of zeolite in the composite is also confirmed by the nitrogen porosimetry results since the initial solid changes from a BET surface area of 25 m 2 / g and a microporous volume of zero to a BET specific surface area of 42 m 2. / g (of which 25 m2 / g of microporous surface) and a microporous volume of 0.01 cm3 / g. The amount of NaY zeolite in the composite was evaluated by solubilizing it in an aqueous solution of hydrofluoric acid at 48% by weight for one hour at room temperature and measuring the

resulting weight loss of the solid. With this method, the zeolite content of the composite is evaluated at 50% by weight. The aluminum and sodium contents in the composite measured by the Induced Plasma Mass Spectrometry (ICP-MS) technique after mineralization of the sample are equal to 0.61% by weight and 0.52% by weight, respectively.

The solid then undergoes successive exchange steps in order to obtain zeolite Y in its ammonium NH4Y form. The protocol used is similar to that proposed by Gola et al. (Gola et al Microporous and Mesoporous Materials 40, 2000, 73-83): for each exchange, the composite is brought into contact with an ammonium nitrate solution (10 M, 5 cm 3 of solution per gram of composite) then refluxed for 4 hours. The composite is then rinsed with distilled water and then dried overnight in an oven at 120 ° C. After five successive exchanges, the solid is calcined in air in muffle furnace at 550 ° C for 12 hours, with a temperature rise ramp of 5 ° C / min. The sodium content measured by ICP-MS on the composite exchanged 5 times is less than 40 ppm. This corresponds to an atomic ratio Na / Al of less than 0.01. The composite then undergoes a dealumination step to adjust the acidity and porosity of the zeolite.

In a first step, the solid undergoes heat treatment in a bed traversed in the presence of steam under the following operating conditions. Typically 5 grams of composite are placed in a quartz reactor in a tube furnace and then mounted at 750 ° C under nitrogen flow (0.4 I nitrogen / h / gram of solid) with a rise ramp of 5 ° C / min. Once the temperature of the plateau reached, the nitrogen passes beforehand into a water saturator at 95 ° C before arrival in the quartz reactor. At the end of the stage, the water saturator is passed and it returns to room temperature under nitrogen flow only. Various tests have made it possible to set the dwell time at one hour in order to obtain a ratio (Si / Al) network equal to 15. The extra-network aluminas generated by the heat treatment are then removed by contacting the solid with a aqueous solution of Na2H2-EDTA according to a protocol similar to that proposed by Gola et al. (Gola et al Microporous and Mesoporous Materials 40, 2000, 73-83) namely 20 cm3 of aqueous solution per gram of composite, and a molar ratio of Na2H2-EDTA to the aluminum present in the composite of 2. The mixture is heated in a refluxing flask for 4 hours, then the solid is rinsed with distilled water and finally exchanged twice with ammonium nitrate solution according to the protocol described in the preceding paragraph, in order to eliminate the sodium ions. introduced by Na2H2-EDTA. Finally, the solid is calcined in air in muffle furnace at 550 ° C for 12 hours, with a temperature ramp of 5 ° C / min. Finally, the composite thus obtained has an Na content of less than 50 ppm, a ratio (Si / Al) network of 17 and a percentage of extra-network aluminum measured by NMR less than 20%.

b) deposition of the hydro-dehydrogenating function (C2) The extrusions of the composite are subjected to a step of dry impregnation with an aqueous solution of platinum nitrate tetramine Pt (NH 3) 4 (NO 3) 2, left to mature in maturation water for 24 hours at

room temperature and then calcined at 450 ° C (ramp rise of 5 ° C / min) for two hours in bed traversed in dry air (2 I air / h / gram of solid). The weight content of platinum of the finished catalyst after calcination is 0.29%. Its dispersion measured by H2-02 titration is 37% and the distribution coefficient of platinum measured by microprobe of Castaing is equal to 0.88.

Example 3 Treatment of a Fischer-Tropsch Cargo According to Embodiment (b) of the Process According to the Invention A Fischer-Tropsch synthesis feedstock over a cobalt catalyst is separated into two fractions, the most Table 2: Characteristics of the heavy fraction Simulated distillation T (5% by weight): 175 ° C (25% by weight): 246 ° C (50% by weight): 346 ° C (75% by weight) % wt): 444 ° C (95% wt): 570 ° C compounds 370 ° C + (by GC) 43% wt% at 15 ° C 0.797 nitrogen content 7 ppm sulfur content <limit detection detailed analysis of C30 fraction (GC) 82% by weight n-paraffins 6% by weight i-paraffins 11% by weight olefins 1% by weight oxygenated This heavy fraction is treated in a hydrogen-traversed bed lost on the hydrotreating catalyst Cl under operating conditions which allow the elimination of olefinic and oxygen compounds as well as traces of nitrogen. The operating conditions selected are as follows: hourly volume velocity VVH (load volume / catalyst volume / hour) = 2 h -1 total working pressure: 5 MPa hydrogen / feed ratio: 200 normal liters / liter 20 temperature: 270 ° VS

After this hydrotreatment, the contents of olefins, oxygenates and nitrogenous compounds in the effluent fall below the detection thresholds, whereas the conversion of the 370 ° C + fraction to the 370 ° C fraction is negligible (less than 5% by weight). ; see Table 3. Carbon monoxide and / or carbon dioxide and / or water and / or ammonia formed during hydrotreatment are removed by a flash and decant step. Table 3: Characteristics of the heavy fraction after hydrotreatment. Simulated T distillation (5% weight): 172 ° CT (25% weight): 242 ° CT (50% weight): 343 ° CT (75% weight): 441 ° CT (95% weight): 568 ° C compounds 370 ° C + (by GC) 41% weight density at 15 ° C 0.797 nitrogen content <limit detection sulfur content <limit detection detailed analysis of the fraction Cao (GC) 91% weight n-paraffins 9% weight i-paraffins <limit detection of olefins <oxygenated detection limit The hydrotreated effluent constitutes the hydrocracking feedstock sent to the hydroisomerization and hydrocracking catalyst C2 according to the invention. Before testing, the catalyst C2 undergoes a reduction step under the following operating conditions: hydrogen flow rate: 1600 normal liters per hour and per liter of catalyst, rise in ambient temperature to 120 ° C.: 10 ° C./min, plateau one hour at 120 ° C, rise from 120 ° C to 450 ° C at 5 ° C / min, two-hour stage at 450 ° C, pressure: 0.1 MPa After reduction, the catalytic test is carried out in the following conditions: total pressure of 7 MPa, - hydrogen load ratio of 600 normal liters / liter, hourly volume velocity (VVH) equal to 1.5 h4. The conversion of the 370 ° C + fraction is taken as: C (370 ° C +) = [(% of 370 ° C "effluents) - (% of 370 ° C-load) j 1 [100 - (% 370 ° C) C- load)] with% 370 ° C-effluents = mass percentage of compounds with boiling points below 370 ° C in the effluents, and% of 370 ° C-load = mass percentage of compounds with dots boiling below 370 ° C in the hydrocracking feed.

The reaction temperature is adjusted to 330 ° C so as to obtain a conversion level of the 370 ° C + fraction equal to 70% by weight. The gas chromatographic analyzes make it possible to obtain the distribution of the different cuts in the hydrocracked effluent (Table 4): C1-C4 cut: hydrocarbons of 1 to 4 carbon atoms, including C5-C9 cut: hydrocarbons of 5 to 9 carbon atoms included (naphtha cut) - C10-C14 cut: hydrocarbons with 10 to 14 carbon atoms inclusive (kerosene cut) - C15-C22 cut: hydrocarbons of 15 to 22 carbon atoms inclusive (gas oil cut) - C22 + cut : hydrocarbons with more than 22 carbon atoms (370 ° C + cut).

Table 4: Cross-sectional distribution of the hydrocracked effluent (GC analysis). % cut weight C1-C4 1.7 cut C5-C9 16.2 cut C10-C14 36.2 cut C15-C22 33.6 cut C22 + 12.3 These results show (Table 4) that the use of a catalyst hydroisomerization and hydrocracking according to the invention and in a process according to the invention allows by hydrocracking and hydroisomerization of a paraffinic feedstock from the Fischer-Tropsch synthesis process to produce middle distillates (kerosene and diesel).

Claims (11)

REVENDICATIONS1. Process for the production of middle distillates from a paraffinic feedstock produced by Fischer-Tropsch synthesis, using a hydrocracking and / or hydroisomerization catalyst comprising at least one hydro-dehydrogenating metal chosen from the group formed by metals group VIB and group VIII of the Periodic Table, taken alone or in admixture and a composite support formed by a Y-type zeolite and silicon carbide SiC, said process operating at a temperature between 270 and 400 ° C at a temperature of pressure between 1 and 9 MPa, at a hourly space velocity of between 0.5 and 5 h -1, a hydrogen flow rate adjusted to obtain a ratio of 400 to 1500 normal liters of hydrogen per liter of feedstock. 2. Process according to claim 1, in which the said hydrocracking and / or hydroisomerization catalyst comprises a composite support consisting, in weight percentage, of: 70 to 99.5% of silicon carbide (SiC), the specific surface area of which is measured by the BET method is greater than 5 m 2 / g, -0.5 to 30%, of a type Y zeolite. 3. Method according to one of claims 1 or 2 wherein the elements of group VIII of said hydrocracking catalyst and / or hydroisomerisation are selected from the noble metals of group VIII, and are selected from platinum and palladium, taken alone or mixed. 4. Method according to one of claims 1 to 3 wherein the noble metal content of said hydrocracking catalyst and / or hydroisomerisation is between 0.01 and 10% by weight relative to the total weight of said catalyst. 5. Method according to one of claims 1 to 4 wherein the pore volume of the support of said hydrocracking catalyst and / or hydroisomerisation is between 0.20 cm3 / g and 1.0 cm3 / g. 6. Method according to one of claims 1 to 5 wherein said process is carried out according to the following steps: a) a fractionation of the feed, b) a possible hydrotreatment of at least a portion of said feedstock from fractionation, c) a possible step of removing at least a portion of the water and optionally CO, CO2, NH3, H2S, d) a passage in the process according to one of claims 1 to 5 of at least a part of said optionally hydrotreated fraction, conversion to the catalyst according to the invention above describes products with a boiling point greater than or equal to 370 ° C. in products with a boiling point of less than 370 ° C. is greater than 40 ° / U by weight, e) a distillation of the hydrocracked / hydroisomerized fraction to obtain middle distillates, and optionally recycling in step d) of the residual fraction boiling above said middle distillates. 7. Process according to claim 6, in which the process comprises the following steps: a) separation of a single so-called heavy fraction with an initial boiling point of between 120 and 200 ° C., b) hydrotreatment of at least a portion of said heavy fraction, c) fractionation into at least 3 fractions: - at least one intermediate fraction having an initial boiling point Ti of between 120 and 200 ° C, and a final boiling point T 2 of greater than 300 ° C and lower at 410 ° C, at least one light fraction boiling below the intermediate fraction, at least one heavy fraction boiling above the intermediate fraction, d) at least a portion of said intermediate fraction passing through hydroisomérisant catalyst, e) passage of at least a portion of said heavy fraction in the process according to one of claims 1 to 5, f) distillation of hydrocracked / hydroisomerized fractions to obtain middle distillates, and rec yclage of the residual fraction boiling above said middle distillates in step (e) on the catalyst according to one of claims 1 to 5 treating the heavy fraction. 8. The method of claim 6 wherein the method comprises the following steps: a) separation of at least a light fraction of the feed so as to obtain a single so-called heavy fraction initial boiling point between 120 and 200 ° C, b) hydrotreating said heavy fraction, followed by a step c) removing at least a portion of the water and CO, CO2, NH3, H2S, d) passing through the process according to one of Claims 1 to 5 of at least a portion of said optionally hydrotreated fraction, conversion to the catalyst according to the invention above describes products with a boiling point greater than or equal to 370 ° C in dot-matrix products. boiling below 370 ° C is greater than 40% by weight, e) distillation of the hydrocracked / hydroisomerized fraction to obtain middle distillates, and recycling in step d) of the residual fraction boiling above said middle distillates. 9. Process according to claim 6, in which the process comprises the following steps: a) fractionation of the feedstock in at least three fractions: at least one intermediate fraction having an initial boiling point Ti of between 120 and 200 ° C .; and a final boiling point T 2 greater than 300 ° C and lower than 410 ° C, at least one light fraction boiling below the intermediate fraction, at least one heavy fraction boiling above the intermediate fraction, b) hydrotreatment of at least a portion of said intermediate fraction, then c) removal of at least a portion of the water produced during the hydrotreatment reactions and optionally CO, CO2, NH3, H2S, d) passage of at least a portion of the hydrotreated fraction on a hydroisomerizing catalyst, e) passing through the process according to one of claims 1 to 5 of at least a portion of said heavy fraction with a product conversion of 370 ° C + to 370 ° products. VS- greater than 40% by weight, f) distilling at least a portion of the hydrocracked / hydroisomerized fractions to obtain middle distillates. The process according to claim 6 wherein the process comprises the following steps: a) optionally fractionation of the feedstock into at least one heavy fraction having an initial boiling point of between 120 and 200 ° C, and at least one boiling light fraction below said heavy fraction, b) optional hydrotreatment of at least part of the feedstock or heavy fraction, optionally followed by c) removal of at least part of the water, d) transfer of at least a part of the effluent or fraction possibly hydrotreated in the process according to one of claims 1 to 5 on a first catalyst according to the invention, e) distillation of the hydroisomerized and hydrocracked effluent to obtain distillates means (kerosene, diesel) and a residual fraction boiling over middle distillates, f) passing at least part of said residual heavy fraction and / or a part of said middle distillates in the selective process n one of claims 1 to 5 on a second catalyst according to the invention, and distillation of the resulting effluent to obtain middle distillates. 11. The method of claim 6 wherein the process comprises the following steps: a) separation of at least one gas fraction C4 ", called light, with a final boiling point below 20 ° C, of the effluent from the Fischer-Tropsch synthesis unit so as to obtain a single C5 + heavy liquid fraction with an initial boiling point of between 20 and 40 ° C; and b) hydrogenation of the olefinic unsaturated compounds of at least a part of said C5 + heavy fraction, in the presence of hydrogen and a hydrogenation catalyst at a temperature of between 80 ° C and 200 ° C, at a total pressure of between 0.5 and 6 MPa, at a mean hourly velocity including between 1 and 10 h -1, and at a hydrogen flow rate corresponding to a volume ratio of hydrogen / hydrocarbons of between 5 and 80 normal liters of hydrogen per liter of feedstock, c) passage of the totality of the liquid hydrogenated effluent from step b), without a step of preliminary separation, in the process according to one of claims 1 to 5 in the presence of hydrogen and a catalyst according to one of claims 1 to 5, d) distillation of the hydrocracked / hydroisomerized effluent.