Improved process for producing kerosene from synthetic hydrocarbons in series with short recycle

Abstract

One objective of the present invention is to minimize the size of the fractionation column located downstream of the hydrocracking section, thereby minimizing the cost of the process without impacting process performance in terms of yield and selectivity of kerosene produced. This is achieved by implementing a separation step at the outlet of hydrotreatment step a) and at the outlet of hydrocracking step c), either in the same separation unit or in separate units. This separation allows at least part, and preferably all, of the hydrocarbon liquid effluent from said separation to be fed directly into hydrocracking step c), bypassing the fractionation step. Figure 2 to be published

Description

Title of the invention: Improved process for producing kerosene from synthetic hydrocarbons in series with short recycle. Field of the invention

[0001] The Fischer-Tropsch process makes it possible to produce synthetic hydrocarbons from a gaseous feed composed mainly of hydrogen and carbon monoxide, also called synthesis gas (CO+H2).

[0002] In particular, in the low-temperature Fischer-Tropsch process, the synthesis gas is catalytically transformed into water, oxygenated products, olefins, and essentially linear paraffins, also called n-paraffins, in gaseous, liquid, or solid form under standard conditions. After water separation, the resulting synthetic hydrocarbons are mainly composed of predominantly linear paraffins but also contain olefins and oxygenated products.

[0003] The effluent from the Fischer-Tropsch process, at the outlet of the Fischer-Tropsch synthesis unit, can be divided into two fractions, a light fraction, called cold condensate, and a heavy fraction, called waxes.

[0004] These synthetic hydrocarbons produced are not directly incorporable into conventional fuel pools or usable as lubricants. For example, the melting point of a normal 14-carbon paraffin with a boiling point of 254°C is approximately 6°C, which makes its direct incorporation into the kerosene fraction impossible given the required specification of -40°C for the disappearance point of crystals for kerosene prepared from paraffins produced by Fischer-Tropsch synthesis according to ASTM D7566. A hydroisomerization step is necessary to lower the pour points of the various hydrocarbon compounds. Furthermore, processing steps are necessary to remove olefinic and oxygenated compounds, increase the yield of middle distillates, and improve the properties of the fractions produced to meet their specifications.To perform the hydroconversion of Fischer-Tropsch synthesis products, so-called bifunctional catalysts are generally used. These are characterized by the combination of an acidic and a hydrogenating function. For the heavy fraction, that is, the fraction with an initial boiling point above the maximum temperature imposed by the kerosene usage specifications, the cracking activity of the aforementioned catalysts is sought in order to convert the molecules into molecules with shorter chains and therefore lower boiling points, until the products formed fall within the desired distillation range. For the light fraction, that is, covering a... With a distillation interval adapted to the target, the sole aim will be to utilize the isomerizing activity of the aforementioned catalysts. To achieve this, one can choose to manipulate the formulation of bifunctional catalysts to make them more or less selective towards cracking or isomerization, but this may prove insufficient given the propensity of so-called isomerizing catalysts to overcrack the products to the point of forming compounds that are too light (gas or gasoline) and therefore no longer usable for the desired purpose, namely kerosene.

[0005] In order to illustrate the problem, we will cite below some results from the prior art which allow us to evaluate it.

[0006] First, patent application FR3084082 describes a process for producing middle distillates, i.e., diesel and kerosene, from a paraffinic feedstock produced by Fischer-Tropsch synthesis, characterized by the treatment of two fractions: a light fraction, called cold condensate, and a heavy fraction, called waxes. In this process, fractionation of the heavy wax fraction yields two other cuts of varying weights. The lighter of the two is then mixed with the cold condensate cut before this mixture is hydrotreated and isomerized. In parallel with this treatment, the heavy wax fraction is hydrocracking to form an effluent, which is then fractionated together with the effluent from the hydroisomerized fraction to obtain at least one gasoline fraction and at least one middle distillate fraction.Although the scheme implements a specific hydroisomerizing catalyst based on at least one IZM-2 zeolite, it appears that approximately 22% by weight of the Fischer-Tropsch synthesis effluent is in the form of naphtha, and it can also be assumed that by specifically aiming to produce kerosene without diesel, the quantity of naphtha would be even higher, which does not allow the problem of maximizing kerosene yield to the desired specifications to be addressed.

[0007] Patent application WO15063213 discloses a process for converting a paraffinic feed, which may be a feed from the Fischer-Tropsch synthesis, in which the feed is transformed in a step a) to obtain a feed at least partially isomerized on a chain of two catalysts in series, a first catalyst 1 and a second catalyst 2 in which catalyst 2 is more isomerizing and less cracking than catalyst 1, then a step of separating the isomerized effluent into one or more middle distillate fractions and a residual fraction.

[0008] The description specifies that the feed transformation step can be carried out in one or more reactors. The description then describes a fractionation step that can lead to a diesel fraction, a kerosene fraction, and a residual fraction boiling above the middle distillates. Catalysts 1 and 2 are preferably composed of an amorphous acid support, preferably of the platinum / silica-alumina type for the first, and of a molecular sieve, preferably of the platinum / zeolite ZSM-12 type, with a silicic binder for the second.

[0009] An optional vacuum distillation step of the residual fraction can be implemented to obtain a base oil distillate fraction and a second residual fraction boiling between 450 and 550°C. Part of said second residual fraction can be recycled to the hydrocracking and hydroisomerization reaction step. The base oil distillate fraction can be sent to a catalytic dewaxing step. It should be noted that if step a) is implemented in two reactors in series, the entire feed from the first reactor is sent to the second reactor. The use of the silica binder for the preparation of the more isomerizing catalyst 2 is presented as essential for achieving selectivity for diesel and / or kerosene.However, the examples show that despite the partial conversion (approximately 20% by weight of base oil is not recycled), 10% by weight of gas is formed and only 67 to 71% of average distillates are produced, with no distinction made between diesel and kerosene. It therefore appears to us that the proposed solution does not address the problem of maximizing kerosene yield to the required specifications.

[0010] Application WO09041478 discloses a process for producing diesel fuel comprising, firstly, a step of fractionating the effluent from the Fischer-Tropsch synthesis into two fractions: an intermediate fraction and a heavier wax fraction. This is followed by a hydroisomerization step of the intermediate fraction and a hydrocracking step of the wax fraction. Finally, the described process includes another fractionation step applied to the mixture of the hydroisomerized intermediate fraction and the hydrocracking heavy wax fraction into two fractions, each comprising a kerosene fraction and a diesel fuel fraction. These fractions are then combined in proportions adjusted to obtain target properties. This process ensures that precise specifications are met, but it does not maximize yield since a fraction of kerosene or diesel fuel must necessarily be discarded.

[0011] Application WO 2021 / 165178 discloses a process for producing naphtha from pyrolysis oil derived from plastic waste. In this process, the hydrogenated effluent can be directly recycled into the selective hydrogenation reaction section or into a hydrotreating step, which allows for the dilution of feed impurities and temperature control, particularly in said reaction section. The target temperature in the reaction section where recycling takes place is between 100 and 250°C, which is specific to the characteristics of the described process, the catalysts used, and the underlying reactions desired and / or undergone. Furthermore, the recycled stream can be drawn from a separation step downstream of the reaction section or of the fractionation without any mention of a technical effect associated with either one, and when the recycle comes from a separation step upstream of the fractionation, this is done in the presence of injected water whose objective is to wash the chloride salts initially present in the feed.

[0012] Application WO 2019 / 134811 discloses a two-stage hydrocracking process in which a high-pressure separation step is carried out after each hydrocracking step and upstream of a fractionation step, said separation step producing a heavy hydrocarbon fraction which is directly processed in the second hydrocracking step. This implementation has three advantages: the presence of dissolved NH3 in the heavy fraction then processed in the second step, which minimizes overcracking during this step; it also limits the formation of heavy polynuclear aromatic molecules because this separation takes place in the presence of hydrogen and their formation in the fractionation column is also avoided; and finally, indirectly, it reduces the proportion of flux to be purged.In the absence of nitrogen in the process according to the invention, as well as of aromatics or naphthenic rings that are precursors of heavy polynuclear aromatic molecules, implementing such a process appears unattractive. Finally, the process seems to operate with a recycle rate of 100%, or at least between 50 and 200%, and with a conversion per pass exceeding 20%, which limits the scope of application of the invention.

[0013] The research work carried out by the applicant led him to discover that in order to selectively produce a kerosene cut with properties conforming to the specifications in force, in particular for the cutting points and cold properties such as the disappearance point of the crystals, it is necessary to combine in a judicious manner at least two hydrocracking and hydroisomerization steps arranged in series while ensuring that operating conditions are maintained such that in these two steps the hydrocracking and hydroisomerization are carried out with a minimum selectivity towards the gas and naphtha cuts.

[0014] An objective of the present invention is to minimize the sizing of the fractionation column located downstream of the hydrocracking section and thus minimize the cost of the process without impacting the performance of the process in terms of yield and selectivity of kerosene produced by implementing a separation step at the outlet of the hydrotreatment step a) and at the outlet of the hydrocracking step c), in the same separation unit or in different units, allowing the direct feeding of at least part and preferably all of the hydrocarbon liquid effluent from said separation to the hydrocracking step c), without going through the fractionation steps b) or d). Object of the invention

[0015] The present invention relates to a process for producing kerosene from a paraffinic feedstock produced by Fischer-Tropsch synthesis, comprising at least the following steps and preferably consisting of:

[0016] a) the hydrotreating of at least part and preferably all of the paraffinic feed produced by Fischer-Tropsch synthesis in the presence of a hydrotreating catalyst and operating at a temperature between 250 and 450°C, at a total pressure between 0.5 and 15 MPa, a hydrogen flow rate adjusted to obtain a ratio between 100 and 3000 normal liters of hydrogen per liter of feed, and at an hourly volumetric rate between 0.1 and 40 h1,

[0017] bl') the separation of the effluent from step a) into at least one hot liquid hydrocarbon fraction and one hot gaseous fraction, said separation being carried out at a temperature above 200°C,

[0018] b2') the separation of the hot gaseous fraction from step bl') into at least one aqueous fraction consisting of at least part of the water formed during said step a) and preferably all of the water formed, a liquid hydrocarbon fraction and a gaseous fraction,

[0019] b) the fractionation in one or more steps of the hydrotreated hydrocarbon liquid fraction from step b2') into at least a gaseous fraction, a naphtha fraction having an initial boiling point between 30 and 60°C and a final boiling point between 100 and 160°C, at least a kerosene fraction having an initial boiling point between 100 and 160°C and a final boiling point between 250 and 320°C, and a heavy fraction having an initial boiling point above 250°C,

[0020] c) hydrocracking all of said heavy fraction having an initial boiling point above 250°C from step b) to produce a hydrocraced effluent, in the presence of a hydrocracking catalyst and operating at a temperature between 250 and 450°C, at a total pressure between 0.2 and 15 MPa, at a space velocity between 0.1 and 10 h1, and at a hydrogen flow rate adjusted to obtain a ratio between 100 and 2000 normal liters of hydrogen per liter of feed, said hydrocracking catalyst comprising at least one noble metal of Group VIII of the periodic table and a support comprising at least one acidic solid, said operating conditions of said step c) being adjusted so that the conversion by passing into products having boiling points above or equal to a temperature Tx into products having boiling points below that same temperature Tx is less than 0.5, of preference less than 0,4 and preferably less than 0.3 and preferably less than 0.2, Tx being inclusive, between 250 and 350°C, preferably between 270 and 340°C and most preferably between 280 and 330°C,

[0021] dl') the separation of the effluent from step c) into at least one hot hydrocracked liquid hydrocarbon fraction and one hot gaseous fraction, said separation being carried out at a temperature above 200°C, said separation steps bl') and dl') advantageously being carried out in the same separation unit or in separate separation units,

[0022] d2') the separation of the hot gaseous fraction from step dl') into at least a liquid hydrocracked hydrocarbon fraction and a gaseous fraction, said separation steps b2') and d2') advantageously being carried out in the same separation unit or in separate separation units,

[0023] d) the fractionation in one or more steps of at least all or part of the hydrocracked hydrocarbon liquid fraction from step d2'), into at least a gaseous fraction, a naphtha fraction having an initial boiling point between 30 and 60°C and a final boiling point between 100 and 160°C, at least a kerosene fraction having an initial boiling point between 100 and 160°C and a final boiling point between 250 and 320°C, and a residual heavy fraction having an initial boiling point above 250°C, said fractionation steps b) and d) advantageously being carried out in the same fractionation unit or in separate fractionation units,

[0024] e) the hydroisomerization of at least a portion of the kerosene fraction having an initial boiling point between 100 and 160°C and a final boiling point between 250 and 320°C from fractionation step b) and of all or part of the kerosene fraction having an initial boiling point between 100 and 160°C and a final boiling point between 250 and 320°C from fractionation step d) where said fractionation steps b) and d) are separate, in the presence of a hydroisomerization catalyst operating at a temperature between 200 and 450°C, a total pressure between 1 and 15 MPa, a space velocity between 0.1 and 10 h⁻¹, and a hydrogen flow rate adjusted to obtain a ratio between 100 and 2000 normal liters of hydrogen per liter of feed, to produce a hydroisomerized effluent,

[0025] f) recycles all or part of the residual heavy fraction having an initial boiling point above 250°C in step c) of hydrocracking,

[0026] g) possibly purging part of the hydrocracked effluent from step c) of hydrocracking,

[0027] h') the separation of the effluent from step e) into at least one liquid fraction hydroisomerized hydrocarbon and a gaseous fraction,

[0028] h) the fractionation of all or part of the hydroisomerized hydrocarbon liquid fraction from step h'), into at least a gaseous fraction, a naphtha fraction having an initial boiling point between 30 and 60°C and a final boiling point between 120 and 160°C, and at least a kerosene fraction having an initial boiling point between 120 and 160°C,

[0029] and wherein at least a part and preferably all of the hot liquid hydrocarbon fraction from bl') and / or at least a part and preferably all of the hot hydrocracked liquid hydrocarbon fraction from dl') are sent to the hydrocracking step c) without going through said fractionation step b) or d).

[0030] The process according to the invention also applies to a situation in which the Fischer-Tropsch synthesis effluent has previously been divided into two fractions, a light fraction, called cold condensate, and a heavy fraction, called waxes, and where these two fractions should be combined beforehand prior to their simultaneous hydrotreatment.

[0031] In general, said process applies to all or part of the effluent from the Fischer Tropsch synthesis.

[0032] An advantage of the present invention is therefore to provide a process for producing kerosene from a paraffinic effluent produced by Fischer-Tropsch synthesis which allows both the maximization of kerosene production and the limitation of the production of light cracked products which cannot be incorporated into a kerosene pool.

[0033] An advantage of the present invention is therefore to provide a process for the production of kerosene from a paraffinic feedstock produced by Fischer-Tropsch synthesis comprising at least one hydrocracking step and at least one hydroisomerization step arranged in series, the operating conditions and the catalysts used in said steps being chosen in a very specific way, to selectively produce a kerosene cut, while minimizing the production of light cuts such as light gases in Ci-C4, naphtha.

[0034] Another advantage of the present invention is to provide a kerosene production process that reduces the investment and operating costs of the process.

[0035] In the following text, and unless otherwise specified, the term "heavy fraction" corresponds to a hydrocarbon effluent obtained at the end of the hydrotreatment and fractionation step(s) after said hydrotreatment step, said hydrocarbon effluent having an initial boiling temperature greater than 250°C, advantageously between 250 and 320°C, preferably between 270 and 310°C and preferably between 280 and 300°C.

[0036] In the following text, and unless otherwise specified, the term "residual heavy fraction" corresponds to a hydrocarbon effluent obtained at the end of the hydrocracking and fractionation step(s) after said hydrocracking step, said hydrocarbon effluent having an initial boiling temperature greater than 250°C, advantageously between 250 and 320°C, preferably between 270 and 310°C and preferably between 280 and 300°C.

[0037] Recycling at least part of the residual heavy fraction to the hydrocracking stage allows all or part of said residual heavy fraction to be cracked and recovered, in particular, into a kerosene fraction having an initial boiling point between 100 and 160°C and a final boiling point between 250 and 320°C. Thus, the non-recycled residual heavy fraction is reduced to the bare minimum, preferably to zero, and failing that, it corresponds to a purge.

[0038] In the following text, the groups of chemical elements are given according to the CAS classification (CRC Handbook of Chemistry and Physics, publisher CRC Press, editor-in-chief DR Lide, 81st edition, 2000-2001). For example, group VIII according to the CAS classification corresponds to the metals of columns 8, 9 and 10 according to the new IUP AC classification, and group VIB to the metals of column 6.

[0039] In the following text, the expressions "between ... and ..." and "between ... and ..." are equivalent and mean that the limit values ​​of the interval are included in the range of values ​​described. If this were not the case and the limit values ​​were not included in the range described, such clarification will be provided by the present invention.

[0040] In this description, the expression "greater than..." is understood as strictly greater than, and symbolized by the sign ">", and the expression "less than" as strictly less than, and symbolized by the sign "<". Detailed description of the invention

[0041] According to the invention, the present invention relates to a process for producing kerosene from a paraffin feedstock produced by Fischer-Tropsch synthesis.

[0042] The Fischer-Tropsch process, well known to those skilled in the art, makes it possible to produce synthetic hydrocarbons from a gaseous feed consisting mainly of hydrogen and carbon monoxide, also called synthesis gas (CO+H2).

[0043] The synthesis gas feeding the Fischer-Tropsch process can advantageously be produced from natural gas, coal, biomass, any source of hydrocarbon compounds, or a mixture of these sources. It can also come from any existing process for supplying CO2, such as the processes of Hydrogen capture is applied to industrial fumes or biogenic CO2, sometimes combined with a CO2-to-CO conversion step. Furthermore, hydrogen can also be obtained from any existing process known to those skilled in the art. Examples include reforming and water electrolysis processes. Hydrogen can also be produced by other methods such as steam reforming of light hydrocarbons or the partial oxidation of various hydrocarbons, including heavy residues. Other hydrogen sources can also be used, such as hydrogen from catalytic cracking gases, which contain significant amounts of CO and CO2. The hydrogen used can also come from the outlet gas of a hydrotreating unit; in this case, the hydrogen may undergo varying degrees of purification to remove impurities such as ammonia (NH3) or hydrogen sulfide (H2S).

[0044] In particular, in the low-temperature Fischer-Tropsch process, the synthesis gas is catalytically transformed into water, oxygenated products, olefins, and essentially linear paraffins, also called n-paraffins, in gaseous, liquid, or solid form under standard conditions. After water separation, the resulting synthetic hydrocarbons are mainly composed of paraffins but also contain olefins and oxygenated products.

[0045] Preferably, after separation of the water, the paraffin effluent from the Fischer-Tropsch process comprises an n-paraffin content greater than 70% by weight and even more preferably greater than 80% by weight in relation to the total mass of said effluent.

[0046] Preferably, the paraffinic effluent from the Fischer-Tropsch process is free from heteroatomic impurities other than oxygen, such as, for example, sulfur, nitrogen or metals.

[0047] Generally, the paraffinic effluent from the Fischer-Tropsch process is, upon exiting the Fischer-Tropsch synthesis unit, divided into two fractions: a light fraction, called cold condensate, and a heavy fraction, called waxes. However, in the case of the present invention, this separation is not mandatory. The process can therefore be applied either to a so-called aggregate effluent combining the cold condensate and the waxes, or to a portion of each of these effluents, or to a reconstituted fraction of these two effluents.

[0048] The light fraction called cold condensate corresponds to hydrocarbons in the gaseous state under the conditions of the Fischer-Tropsch reaction and the heavy fraction called waxes corresponds to hydrocarbons in the liquid state under the conditions of the Fischer-Tropsch reaction.

[0049] The cold condensed light fraction has an initial boiling point T1 between 15 and 40°C and preferably between 20 and 35°C and a final boiling point T2 between 350 and 400°C and preferably between 360 and 380°C and preferably less than 370°C.

[0050] The heavy wax fraction advantageously has an initial boiling point T3 between 100 and 300°C and preferably between 125 and 200°C. Preferably, said heavy fraction contains paraffins having more than 8 carbon atoms.

[0051] This heavy fraction has the particularity, due to its composition, namely essentially n-paraffins having a high number of carbon atoms, which can go up to more than 150, of being solid at room temperature.

[0052] In the light fraction, the paraffin content is greater than 70% by weight, the olefin content is less than 20% by weight, and the oxygenated compound content is less than 10% by weight, the contents being expressed as a percentage by weight relative to the total mass of said light fraction. In the heavy fraction, the paraffin content is greater than 80% by weight, the olefin content is less than 15% by weight, and the oxygenated compound content is less than 5% by weight, the contents being expressed as a percentage by weight relative to the total mass of said heavy fraction. Step a)

[0053] According to the invention, the process comprises a) a hydrotreating step of at least part and preferably all of the effluent from the Fischer-Tropsch synthesis in the presence of a hydrotreating catalyst and operating at a temperature between 250 and 450°C, at a total pressure between 0.5 and 15 MPa, a hydrogen flow rate adjusted to obtain a ratio between 100 and 3000 normal liters of hydrogen per liter of feed, and at an hourly volumetric rate between 0.1 and 40 h1.

[0054] Preferably, said step a) of hydrotreating operates at a temperature between 280 and 420°C and preferably between 330 and 390°C, at a total pressure between 1 and 9 MPa, a hydrogen flow rate adjusted to obtain a ratio between 250 and 2000 normal liters of hydrogen per liter of feed and preferably between 500 and 1500 normal liters of hydrogen per liter of feed, and at an hourly volumetric rate between 0.25 and 20 h1 and preferably between 0.5 and 10 h1.

[0055] Said hydrotreatment step makes it possible to reduce the content of olefinic and unsaturated compounds as well as to decompose the oxygenated compounds present in the effluent from the Fischer-Tropsch synthesis.

[0056] The hydrotreating catalyst implemented in step a) is a conventional hydrotreating catalyst. Said catalyst comprises at least one metal from the group VIII and / or from group VI of the periodic table of elements. Preferably, the catalyst comprises at least one metal from the group of metals formed by nickel, molybdenum, tungsten, cobalt, ruthenium, indium, palladium, and platinum, alone or in mixtures, and comprises at least one support selected from aluminas, boron oxides, magnesia, zirconia, titanium oxides, and clays, or a combination thereof; preferably, said support is an alumina. Said catalysts may advantageously be prepared by any method known to those skilled in the art or may be acquired from companies specializing in the manufacture and sale of catalysts. The shape of the supports and catalysts involved in step a) may be spherical, or extruded in the form of cylinders, trilobes, or quadrilobes.

[0057] When using non-noble metals from Group VIII, a combination of at least one metal from Group VI, preferably molybdenum or tungsten, and at least one metal from Group VIII, preferably cobalt or nickel, is advantageously used. The concentration of the non-noble Group VIII metal, when used, is advantageously from 0.01% to 15% by weight of oxide equivalent relative to the finished catalyst, and that of the Group VI metal is advantageously from 5% to 40% 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 reduced or sulfide form.

[0058] Under these conditions, the content of unsaturated and oxygenated molecules in the effluent from said step a) is reduced to less than 0.5% by weight and to approximately less than 0.1% by weight in general. The hydrotreating step is carried out under conditions such that the conversion of products having boiling points greater than or equal to 300°C into products having boiling points less than 300°C is limited to 20% by weight, preferably less than 10% by weight, and even more preferably less than 5% by weight.

[0059] The effluent from step a) can advantageously be sent to a stage for the removal of at least part of the water formed during said step a) and preferably of all the water formed before being sent to step b) of fractionation according to the invention.

[0060] Said step of removing at least part of the water can advantageously be carried out by all methods and techniques known to those skilled in the art, for example by drying, passing over a desiccant, flash drying or decantation. Step bl')

[0061] The process according to the invention comprises a step of separating the effluent from step a) into at least a hot liquid hydrocarbon fraction and a fraction hot gaseous, said separation being carried out at a temperature above 200°C, preferably above 230°C and more preferably above 250°C and most preferably below 390°C.

[0062] Said separation step bl'), can advantageously be carried out according to all methods and techniques known to those skilled in the art. Preferably, said separation step bl') is carried out in one or more hot separator vessels advantageously operating at a total pressure as close as possible to the total outlet pressure of the hydrotreatment step a), and preferably between 0 and 1 MPa lower than the total outlet pressure of step a), and more preferably between 0 and 0.5 MPa lower than the total outlet pressure of step a).

[0063] Preferably, said gaseous fraction separated in said step bl') advantageously comprises hydrogen and the light gases Ci-C4 and optionally comprises a gaseous fraction heavier than the Ci-C4 fraction and with a boiling point below 300°C.

[0064] Preferably, said gaseous fraction comprises less than 1% by weight of compounds having a boiling point above 300°C, relative to the total weight of said fraction. Step b2')

[0065] According to the invention, the process includes a separation step b2') of the hot gaseous fraction from step bl') into at least an aqueous fraction consisting of at least a part of the water formed during said step a) and preferably all of the water formed, a liquid hydrocarbon fraction and a gaseous fraction.

[0066] Said separation step b2'), can advantageously be carried out according to all methods and techniques known to those skilled in the art. Preferably, said separation step b2') is carried out in one or more cold separator vessels advantageously operating at a total pressure as close as possible to the outlet pressure of the hydrotreatment step a) and preferably between 0 and 0.5 MPa lower than the total outlet pressure of step bl').

[0067] Preferably said step b2') is carried out at a temperature between 20 and 80°C and preferably between 30 and 70°C and preferably between 40 and 60°C.

[0068] Preferably, said gaseous fraction separated in said step b2') advantageously comprises hydrogen and light gases Ci-C4 and optionally comprises a gaseous fraction heavier than the Ci-C4 fraction and with a boiling point below 100°C.

[0069] Preferably, said gaseous fraction comprises less than 1% by weight of compounds having a boiling point above 100°C, relative to the total weight of said fraction.

[0070] Preferably, said steps b1') and b2') are carried out respectively in a hot flask and in a cold flask. Step b)

[0071] The process according to the invention comprises a step b) of fractionating in one or more steps the hydrotreated hydrocarbon liquid fraction from step b2') into at least one gaseous fraction advantageously comprising the light C1-C4 gases, a naphtha fraction having an initial boiling point between 30 and 60°C and a final boiling point between 100 and 160°C, at least one kerosene fraction having an initial boiling point between 100 and 160°C and a final boiling point between 250 and 320°C, and a heavy fraction having an initial boiling point greater than 250°C.

[0072] Said fractionation step can advantageously be carried out by atmospheric distillation and in some cases by the combination of atmospheric distillation and vacuum distillation.

[0073] Another embodiment of this fractionation step consists of carrying out an additional stripping of the effluent from step a) prior to atmospheric fractionation in order to purge the gaseous fractions at the inlet of said atmospheric fractionation column. Step c)

[0074] According to the invention, the process comprises a step c) of hydrocracking at least a part and preferably all of the heavy fraction from step b) of fractionation to produce a hydrocraced effluent, in the presence of a hydrocracking catalyst and operating at a temperature between 250 and 450°C, at a total pressure between 0.2 and 15 MPa, at a space velocity between 0.1 and 10 h1, and with a hydrogen flow rate adjusted to obtain a ratio of between 100 and 2000 normal liters of hydrogen per liter of feed, said hydrocracking catalyst comprising at least one noble metal of group VIII and a support comprising an acidic solid, said operating conditions of said step c) being adjusted so that the conversion by passing from products having boiling points greater than or equal to a temperature Tx to products having boiling points less than this same temperature Tx is less than 0.5,preferably less than 0.4, preferably less than or equal to 0.3 and preferably less than 0.2, with Tx between 250 and 350°C, preferably between 270 and 340°C and most preferably between 280 and 330°C.

[0075] Preferably, step c) operates at a temperature between 280 and 450°C, and even more preferably between 320 and 420°C, at a total pressure between 0.5 and 10 MPa, more preferably between 1 and 9 MPa, at a space velocity between 0.2 and 7 h1, and more preferably between 0.5 and 5 h1, and with a hydrogen flow rate adjusted to obtain a ratio between 150 and 1500 normal liters of hydrogen per liter of charge and more preferably between 300 and 1500 normal liters of hydrogen per liter of charge.

[0076] The charge of step c) advantageously has an initial boiling temperature greater than 250°C, preferably between 250 and 320°C, preferably between 270 and 310°C and more preferably between 280 and 300°C.

[0077] According to the invention, the hydrocracking catalyst implemented in step c) comprises at least one noble hydro-dehydrogenating metal from group VIII of the periodic table and a support comprising at least one acid solid and preferably a Brønsted acid solid, and optionally a binder.

[0078] Preferably, said hydrocracking catalyst comprises at least one noble metal from group VIII selected from platinum and palladium, taken alone or in mixture, preferably active in their reduced form.

[0079] The noble metal content of said catalyst is advantageously between 0.01 and 5% by weight relative to the finished catalyst, preferably between 0.02 and 4% by weight, most preferably between 0.03 and 3% by weight, even more preferably between 0.05 and 2% by weight and most preferably between 0.05 and 1% by weight.

[0080] The metallic function is advantageously introduced into the catalyst by any method known to those skilled in the art, such as co-mixing, dry impregnation or exchange impregnation.

[0081] Advantageously, the Brønsted acid solid comprises and preferably consists of silica alumina and / or zeolite.

[0082] In the case where the Brønsted acid solid comprises a zeolite, the latter is preferably a zeolite of structural type FAU, *BEA, ISV, IWR, IWW, MEI, UWY, taken alone or in a mixture, and preferably selected from zeolites of structural type FAU and *BEA, taken alone or in a mixture. In a preferred embodiment, the zeolite is selected from zeolite Y and beta zeolite taken alone or in a mixture, and preferably the zeolite is zeolite Y, and most preferably dealuminated zeolite USY.

[0083] Preferably the Brønsted acid solid comprises and is preferably made of silica-alumina.

[0084] Optionally, said hydrocracking catalyst support may also include a binder. Preferably said support includes a binder when said The support comprises a zeolite. The binder is advantageously selected from silica (SiO2), alumina (Al2O3), clays, titanium dioxide (TiO2), boron dioxide (B2O3), and zirconia (ZrO2), alone or in mixtures. Preferably, the binder is selected from silica and alumina, and even more preferably, the binder is alumina in all its forms known to those skilled in the art, such as, for example, gamma alumina.

[0085] A preferred hydrocracking catalyst according to the invention advantageously comprises at least one noble metal, said noble metal being platinum, and a silica-alumina as a Brønsted acid solid, without any other binder.

[0086] The silica content of the silica-alumina, expressed as a percentage by weight, is advantageously between 1% and 95%, advantageously between 5% and 95%, preferably between 10% and 80%, most preferably between 20% and 70%, and even more preferably between 22% and 45%. This silica content is perfectly measured using X-ray fluorescence.

[0087] A preferred hydrocracking catalyst used in the process according to the invention comprises a particular silica-alumina. Preferably, said catalyst comprises 0.05% to 10% by weight, preferably between 0.1% and 5% by weight of at least one noble metal from Group VIII, preferably selected from platinum and palladium (preferably platinum), deposited on a silica-alumina support, without any other binder, containing an amount of silica (SiO2) of between 1% and 95%, expressed as a percentage by weight, preferably between 5% and 95%, preferably between 10% and 80%, and most preferably between 20% and 70%, and even more preferably between 22% and 45%, said catalyst having:

[0088] - a specific BET surface area of ​​100 to 500 m2 / g, preferably between 200 and 450 m² / g and preferably between 200 and 300 m² / g,

[0089] - an average mesopore diameter measured by mercury porosimetry including between 4 and 12 nm, preferably between 4 and 11 nm and most preferably between 5 and 11 nm,

[0090] - a total porous volume measured by mercury porosimetry between 0.2 and 1.2 ml / g, preferably between 0.3 and 1.0 ml / g and most preferably between 0.3 and 0.9 ml / g,

[0091] - a volume of macropores measured by mercury porosimetry, the diameter of which is greater than 50 nm, less than 0.02 ml / g.

[0092] - an alkali or alkaline-earth compound content of less than 300 ppm by weight and preferably less than 200 ppm by weight.

[0093] The average diameter of the mesopores is defined as the diameter corresponding to the cancellation of the curve derived from the volume of mercury intrusion obtained by mercury porosimetry for pore diameters between 3.7 and 50 nm.

[0094] Preferably, the dispersion of the metal of said preferred catalyst is advantageously between 5% and 100%, preferably between 5% and 90%, and most preferably between 10% and 90%. The dispersion, representing the fraction of metal accessible to the reagent relative to the total amount of metal in the catalyst, is advantageously measured, for example, by H2 / O2 titration or by transmission electron microscopy.

[0095] Preferably, the noble metal distribution coefficient of said preferred catalyst is greater than 0.1, preferably greater than 0.2, and most preferably greater than 0.4. The noble metal distribution represents the distribution of the metal within the catalyst grain, the metal being either well or poorly dispersed. Thus, it is possible to obtain platinum that is poorly distributed (for example, detected in a ring whose thickness is significantly less than the radius of the grain) but well dispersed, meaning that all the platinum atoms located in the ring will be accessible to the reactants. The noble metal distribution coefficient can be measured by a Castaing microprobe.

[0096] The noble metal salt is advantageously introduced by one of the usual methods used to deposit the metal on the surface of a solid. One preferred method is dry impregnation, which consists of introducing the metal salt into a volume of solution equal to the pore volume of the mass of solid to be impregnated. Before the reduction operation, the catalyst can advantageously undergo calcination, for example, by treatment under dry air at a temperature of 300 to 750°C, and preferably at 450°C, for 0.25 to 10 hours, and preferably for 2 hours.

[0097] During this step c) the feed 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 improve the quality of the products formed and more particularly the cold properties of kerosene, and also to selectively produce kerosene with respect to naphtha which requires maintaining low conversion rates per pass.

[0098] According to the invention, the operating conditions of said step c) are adjusted so that the conversion by passing into products having boiling points greater than or equal to a temperature Tx into products having boiling points less than this same temperature Tx is less than 0.5, preferably less than 0.4, preferably less than 0.3 and preferably less than 0.2, Tx being between 250 and 350°C, preferably between 270 and 340°C and most preferably between 280 and 330°C.

[0099] The temperature Tx, and the resulting conversion rate, is (are) preferably chosen so as to maximize the kerosene yield of the process according to the invention while respecting the specifications of aviation kerosene.

[0100] The expected effect of the applied operating conditions and the pass conversion they generate is to limit overcracking into light products such as gas or naphtha, the latter not being desired.

[0101] When the temperature increase required to preserve the target pass conversion becomes too great and the selectivity in light products, gas and / or gasoline increases, the catalytic cycle can be interrupted and the catalyst is discharged and then replaced with a new fresh catalyst with which the desired effect will again be found.

[0102] The conversion per pass is defined as:

[0103] Conversion of Tx+ to Tx = [(% mass of effluent Tx) - (% mass of feed Tx)] / [100 - (% mass of feed Tx)],

[0104] with:

[0105] - % mass of effluent Tx: mass percentage of compounds having points boiling points lower than Tx in the hydrocracking effluent,

[0106] - % mass of Tx charge: mass percentage of compounds having points boiling points lower than Tx in the hydrocracking feed.

[0107] It is advantageously observed that the fraction of compounds having a boiling point in the naphtha range, therefore lower than those incorporable into the kerosene product, which is found in the hydrocracked effluent from step c) is then greatly reduced by limiting the conversion of the heavy fraction corresponding to the compounds with boiling points above Tx to levels below 0.5, preferably below 0.4, preferably less than or equal to 0.3 and preferably less than 0.2. Step dl')

[0108] According to the invention, the process according to the invention comprises a separation step dl') of the effluent from step c) into at least a hot hydrocracked hydrocarbon liquid fraction and a hot gaseous fraction.

[0109] Said step dl') is advantageously implemented in the same ranges of operating conditions as step bl').

[0110] Preferably, said gaseous fraction separated in said step bl') advantageously comprises hydrogen and the light gases Cl-C4 and optionally comprises a gaseous fraction heavier than the Ci-C4 fraction and with a boiling point below 300°C.

[0111] Preferably, said gaseous fraction comprises less than 1% by weight of compounds having a boiling point above 300°C, relative to the total weight of said fraction.

[0112] According to the invention, said separation steps bl') and dl') can advantageously be carried out in the same separation unit or in separate separation units, and preferably in the same separation unit.

[0113] In the preferred embodiment where the separation steps bl') and dl') are carried out in the same separation unit, all or part of the effluent from step a) and all or part of the effluent from step c) are treated jointly in said same separation unit.

[0114] According to the invention, at least a part and preferably all of the hot liquid hydrocarbon fraction from bl') and / or at least a part and preferably all of the hot hydrocracked liquid hydrocarbon fraction from dl') are sent to the hydrocracking step c) without going through said fractionation step b) or said fractionation step d) in the case where step b) and step d) are carried out in the same unit).

[0115] This implementation according to the invention makes it possible to minimize the sizing of the fractionation column(s) of step(s) b) and d) and to minimize the operating cost of the process.

[0116] In a highly preferred embodiment of the invention, at least part and preferably all of the hot liquid hydrocarbon fraction from bl') and at least part and preferably all of the hot hydrocracked liquid hydrocarbon fraction from dl') are sent to hydrocracking step c) without going through said fractionation step b) or through fractionation step d) (or said fractionation step d) in the case where step b) and step d) are carried out in the same unit).

[0117] In a highly preferred embodiment of the invention, only at least a part and preferably all of the hot hydrocracked liquid hydrocarbon fraction from dl') is sent to the hydrocracking step c) without going through said fractionation step d) (or said fractionation step b in the case where step b) and step d) are carried out in the same unit).

[0118] In this case where fractionation b) and d) are common, at least a part and preferably all of the hot liquid hydrocarbon fraction from bl') is mixed with the hydrocracked liquid hydrocarbon fraction from step d2'), and the mixture is sent to fractionation step b). Step d2')

[0119] The process according to the invention includes a separation step d2') of the hot gaseous fraction from step d1') into at least a liquid hydrocarbon fraction and a gaseous fraction.

[0120] Said step d2') is advantageously implemented in the same ranges of operating conditions as step b2').

[0121] Preferably, said gaseous fraction separated in said step d2') advantageously comprises hydrogen and light gases Ci-C4 and optionally comprises a gaseous fraction heavier than the Ci-C4 fraction and with a boiling point below 100°C.

[0122] Preferably, said gaseous fraction comprises less than 1% by weight of compounds having a boiling point above 100°C, relative to the total weight of said fraction.

[0123] Preferably, said steps dl') and d2') are carried out respectively in a hot flask and in a cold flask.

[0124] According to the invention, said separation steps b2') and d2') can advantageously be carried out in the same separation unit or in separate separation units, and preferably in the same separation unit.

[0125] According to one variant of the process, the 4 steps bl'), b2'), dl') and d2') are carried out in different separation units.

[0126] According to another variant of the process, steps bl') and dl') are carried out in different units and steps b2') and d2') are carried out in the same separation unit.

[0127] According to another variant of the process, steps bl') and dl') are carried out in the same separation unit and steps b2') and d2') are carried out in the same separation unit. Step d)

[0128] The process according to the invention includes a step d) of fractionating at least all or part of the hydrocracked hydrocarbon liquid fraction from step d2').

[0129] Preferably, the entire hydrocracked hydrocarbon liquid fraction from step d2') is fractionated in step d).

[0130] In another embodiment, part of the hydrocracked hydrocarbon liquid fraction from step d2') is purged and the other part is fractionated in step d).

[0131] Step d) enables the fractionation of said hydrocracked hydrocarbon liquid fraction into at least one gaseous fraction advantageously comprising light Ci-C4 gases, at least one gasoline or naphtha fraction having an initial boiling point between 30 and 60°C and a final boiling point between 100 and 160°C, at least one kerosene fraction having an initial boiling point between 100 and 160°C and a final boiling point between 250 and 320°C, and a residual heavy fraction having an initial boiling point above 250°C, said fractionation steps b) and d) advantageously being carried out in the same fractionation unit or in separate units.

[0132] Said fractionation step can advantageously be carried out by atmospheric distillation and in some cases by the combination of atmospheric distillation and vacuum distillation.

[0133] Another embodiment of this fractionation step consists of carrying out additional stripping of the effluent from step c) prior to atmospheric fractionation in order to purge the gaseous fractions at the inlet of said atmospheric fractionation column.

[0134] Said steps b) and d) of fractionation may advantageously be carried out in the same fractionation unit or in separate units. Preferably, steps b) and d) of fractionation are carried out in the same unit.

[0135] In the preferred embodiment where fractionation steps b) and d) are carried out in the same fractionation unit, all or part of the hydrotreated hydrocarbon liquid fraction from step b2') and all or part of the hydrocracked hydrocarbon liquid fraction from step d2') are treated jointly in said same fractionation unit.

[0136] In a preferred embodiment in which fractionation steps b) and d) are carried out in the same unit, the hydrocracked hydrocarbon liquid fraction from step d2') undergoes purging before being fractionated.

[0137] In a highly preferred embodiment, the separation steps bl') and dl') are carried out in the same separation unit, the steps b2') and d2') are carried out in the same separation unit and the fractionation steps b) and d) are carried out in the same fractionation unit. Step e)

[0138] According to the invention, the process comprises a step e) of hydroisomerizing at least a portion of the kerosene fraction having an initial boiling point between 100 and 160°C and a final boiling point between 250 and 320°C from fractionation step b) and all or part of the kerosene fraction having an initial boiling point between 100 and 160°C and a final boiling point between 250 and 320°C from fractionation step d) in the case where said fractionation steps b) and d) are separate. Step e) is carried out in the presence of a hydroisomerization catalyst operating at a temperature between 200 and 450°C, a total pressure between 1 and 15 MPa, and a spatial velocity between between 0.1 and 10 h1, a hydrogen flow rate adjusted to obtain a ratio of between 100 and 2000 normal liters of hydrogen per liter of feed to produce a hydroisomerized effluent.

[0139] Preferably, said hydroisomerization step e) operates at a temperature between 250 and 450°C and preferably between 280 and 430°C, a total pressure between 1 and 10 MPa and preferably between 1 and 9 MPa, a space velocity between 0.2 and 7 h₁ and preferably between 0.5 and 5 h₁, a hydrogen flow rate adjusted to obtain a ratio between 150 and 2000 normal liters of hydrogen per liter of charge and preferably between 150 and 1500 normal liters of hydrogen per liter of charge.

[0140] According to a first variant, all the kerosene fractions obtained after step(s) b) and d) are sent to the hydroisomerization step e).

[0141] The catalyst used in step e) of hydroisomerization according to the invention is advantageously of the bifunctional type, that is to say that it has a hydro / dehydrogenating function and a hydroisomerizing function.

[0142] Preferably, the hydroisomerization catalyst comprises and is preferably made of at least one noble metal from Group VIII of the periodic table (which acts as a hydro / dehydrogenating function), and a support comprising and preferably made of at least one binder and at least one zeolite selected from TON structural type zeolites, preferably selected from ZSM-22 and NU-10, taken alone or in mixtures, FER structural type zeolites, preferably selected from ZSM-35 and ferrierite, taken alone or in mixtures, EUO structural type zeolites, preferably selected from EU-1 and ZSM-50, taken alone or in mixtures, AEL structural type zeolites, preferably SAPO-11, *MRE structural type zeolites, preferably selected from ZSM-48, ZBM-30, EU-2 and EU-11, taken alone or in mixtures, structural MTW-type zeolites preferably selected from ZSM-12, TPZ-12, Theta-3, NU-13, CZH-5,taken alone or in mixtures, MOR structural type zeolites preferably selected from mordenite and LZ-211, taken alone or in mixtures, BEA structural type zeolites preferably selected from Beta and Tschernichite) and IZM-2 zeolite, taken alone or in mixtures.

[0143] The hydro / dehydrogenating function

[0144] Preferably the noble metal of group VIII of the catalyst used in step e) is chosen from palladium and platinum and is preferably platinum.

[0145] Advantageously, the hydro / dehydrogenating (metallic) element, preferably platinum, can be introduced onto the catalyst support by any known method of the man of the trade, such as comalaxage, dry impregnation, impregnation by exchange.

[0146] According to one or more embodiments, the content of Group VIII noble metal, and preferably the platinum content, in the catalyst used in step e) is between 0.01% and 4% by weight, preferably between 0.05% and 2% by weight, relative to or total weight of said catalyst.

[0147] The catalyst used in step e) may also advantageously comprise at least one additional metal selected from the group formed by the metals of groups IIIA, IVA and VIIB of the periodic table of elements and preferably selected from gallium, indium, tin and rhenium. Said additional metal is preferably selected from indium, tin and rhenium.

[0148] Preferably, the content of at least one additional metal in the catalyst used in step e) is between 0.01% and 2% by weight, preferably between 0.05% and 1% by weight, relative to the total weight of said catalyst.

[0149] In one embodiment, the sulfur content in the hydroisomerization catalyst is such that the ratio of the number of moles of sulfur to the number of moles of at least one metal from Group VIIIB is between 0.3 and 3. In one or more embodiments, the presence of sulfur in the catalyst results from an optional sulfidation step of the hydroisomerization catalyst. In one or more embodiments, the presence of sulfur in the catalyst results from potentially present impurities, such as, for example, in the alumina binder. In another embodiment, the catalyst does not contain sulfur.

[0150] The hydroisomerizing function.

[0151] Preferably, the catalyst contains at least one zeolite of the following structural types: TON (e.g., selected from ZSM-22 and NU-10, taken alone or in a mixture) and / or FER (e.g., selected from ZSM-35 and ferrierite, taken alone or in a mixture) and / or EUO (e.g., selected from EU-1 and ZSM-50, taken alone or in a mixture) and / or AEL (e.g., SAPO-11) and / or *MRE (e.g., selected from ZSM-48, ZBM-30, EU-2 and EU-11, taken alone or in a mixture) and / or MTW (e.g., selected from ZSM-12, TPZ-12, Theta-3, NU-13, CZH-5, taken alone or in a mixture) and / or MOR (e.g., selected from mordenite or LZ-211, taken alone or in a mixture) and / or *BEA (e.g., selected from the Beta or Tschernichite). The structural codes are defined in the International Zeolite Association (IZA: http: / / www.iza-structure.org / databases / ) classification. The zeolite may also be IZM-2, whose structural code is unknown.

[0152] Preferably, the catalyst contains at least one zeolite selected from IZM-2 zeolite and one MTW structural type zeolite preferably selected from ZSM-12, TPZ-12, Theta-3, NU-13, CZH-5, taken alone or in mixture.

[0153] Preferably, the catalyst contains at least one zeolite selected from IZM-2 and ZSM-12, taken alone or in mixture.

[0154] The zeolites are preferably in their acidic form, that is to say, the atomic ratio between the monovalent compensating cation (for example, sodium) and the aluminum inserted in the crystal lattice of the solid is advantageously less than 0.1, preferably less than 0.05, and most preferably less than 0.01. According to one or more embodiments, the zeolites entering into the composition of said hydroisomerization catalyst are advantageously calcined. According to one or more embodiments, said zeolites are 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, leads to the acidic form of said zeolites.

[0155] According to a preferred embodiment, the catalyst comprises a support comprising IZM-2 zeolite alone or a support comprising ZSM-12 zeolite alone.

[0156] Zeolite IZM-2 is a crystalline microporous solid whose crystal structure and preparation process are described in patent application FR2918050A1. The structural code of zeolite IZM-2 is not yet known. Zeolite ZSM-12 is a crystalline microporous solid whose crystal structure is described on the website of the International Zeolyst Association (http: / / www.iza-structure.org / ). It is a one-dimensional zeolite with 12 MRs, and its structural code is MTW. A preparation process for this zeolite is described, for example, in the article "Synthesis of zeolite ZSM-12 in the System (MTEA)2O-Na2O-SiO2-Al2O3-H2O" by S. Ernst et al. (Zeolites, 7, 5, 458-462, DOI10.1016 / 0144-2449(87)90015-7).

[0157] Preferably, the catalyst used in step e) comprises a zeolite content of between 1% and 90% by weight, preferably between 3% and 80% by weight, and more preferably between 4% and 60% by weight, preferably between 4% and 30% by weight and even more preferably between 4% and 20% by weight relative to or total weight of said catalyst.

[0158] The binder.

[0159] Preferably, the binder of the catalyst support in step e) is amorphous or crystalline. Preferably, the binder used in the catalyst support in step e) is advantageously chosen from the group consisting of alumina, silica, silica-alumina, clays, titanium oxide, boron oxide, zirconia, and aluminates, alone or in mixtures. Preferably, the binder is alumina. Preferably, said binder may contain alumina in all its forms known to those skilled in the art, such as, for example, alpha, gamma, eta, and delta aluminas.

[0160] Preferably, the catalyst used in step e) comprises a binder content of between 10% and 99% by weight, relative to or total weight of said catalyst, i.e., of in order to ensure that the elements constituting the catalyst used in step e) are complemented to 100% by weight.

[0161] According to the invention, the catalyst support comprises the zeolite mixed with a binder. The shaping of the support in the form of a mixture is preferably carried out by co-mixing, extrusion, and then heat treatment of the zeolite with the binder or a binder precursor, such as, for example, boehmite, which, upon heat treatment, is transformed into alumina.

[0162] A preferred catalyst for step e) comprises and is preferably made of platinum, and a support comprising and preferably made of a ZSM-12 zeolite and an alumina binder.

[0163] Another preferred catalyst for step e) comprises and is preferably made of platinum, and a support comprising and preferably made of an IZM-2 zeolite and an alumina binder.

[0164] According to a preferred embodiment, the catalyst used in step e) more particularly comprises, and preferably consists of:

[0165] - from 1% to 90% by weight, preferably from 3% to 80% by weight and even more preferred 4% to 60% zeolite weight;

[0166] - from 0.01% to 4% by weight, preferably from 0.05% to 2% by weight of at least one metal of the group VIIIB, preferably platinum;

[0167] - possibly from 0.01% to 2% by weight, preferably from 0.05% to 1% by weight of less one additional metal chosen from the group formed by the metals of groups IIIA, IVA and VIIB; - possibly a sulfur content, preferably such that the ratio of the number of moles of sulfur to the number of moles of group VIIIB metal(s) is between 0.3 and 3; and - possibly at least one binder, preferably alumina, ensuring the complement to 100% in the catalyst, relative to the total weight of the catalyst in step e).

[0168] Preferably, the catalyst used in step e) is shaped into cylindrical or multilobed extrudates such as bilobed, trilobed, or multilobed straight or twisted shapes. According to one or more embodiments, the catalyst used in step e) is shaped into crushed powders, tablets, rings, balls, or wheels. Techniques other than extrusion, such as pelletizing or coating, may advantageously be used. The shaping may also advantageously be carried out in the presence of the various constituents of the catalyst and extrusion of the resulting mineral paste by pelletizing, shaping into balls using a rotary dripper or drum, drop coagulation, oil-drop, oil-up, or any other known agglomeration process. powder containing alumina and possibly other ingredients chosen from those mentioned above.

[0169] Furthermore, the supports used in the process according to the present invention may advantageously have been treated, as is well known to those skilled in the art, with additives to facilitate shaping and / or improve the final mechanical properties of the supports. Examples of additives include, in particular, cellulose, carboxymethyl cellulose, carboxyethyl cellulose, tall oil, xanthan gums, surfactants, flocculating agents such as polyacrylamides, carbon black, starches, stearic acid, polyacrylic alcohol, polyvinyl alcohol, biopolymers, glucose, polyethylene glycols, etc.

[0170] Extrusion can advantageously be carried out using any commercially available conventional tool. The dough resulting from the mixing process is advantageously extruded through a die, for example using a piston or a single or double screw extrusion die. This extrusion step can advantageously be carried out by any method known to those skilled in the art.

[0171] The catalyst support implemented in step e) according to the present invention is then advantageously subjected to a drying step carried out according to any technique known to a person skilled in the art.

[0172] Preferably, drying is carried out under an airflow. Said drying may also advantageously be carried out under a flow of any oxidizing, reducing, or inert gas. Preferably, drying is advantageously carried out between 50 and 180°C, most preferably between 60 and 150°C, and most preferably between 80 and 130°C.

[0173] Said support, possibly dried, then preferably undergoes a calcination step.

[0174] The calcination step is advantageously carried out in the presence of molecular oxygen, for example by performing an air purge, at a temperature advantageously above 200°C and less than or equal to 1100°C. The calcination step can advantageously be carried out in a flow bed, a lick bed, or in a static atmosphere. For example, the furnace used can be a rotary kiln or a vertical radial flow bed kiln. Preferably, the calcination step is carried out for more than one hour at 200°C and less than one hour at 1100°C. The calcination can advantageously be carried out in the presence of steam and / or in the presence of an acidic or basic vapor. For example, the calcination can be carried out under partial pressure of ammonia.

[0175] Post-calcination treatments may optionally be carried out in order to improve the properties of the support, for example textural properties.

[0176] Preferably, the noble metal contained in said catalyst used in step e) can advantageously be reduced. One preferred method for carrying out the reduction of the metal is treatment at a temperature between 150 and 650°C and a total pressure between 0.1 and 25 MPa. For example, a reduction may consist of a two-hour holding period at 150°C followed by a temperature ramp-up to 450°C at a rate of 1°C / min and then a two-hour holding period at 450°C; during the reduction step, the hydrogen flow rate may be 1000 normal m³ hydrogen / m³ catalyst and the total pressure may be maintained constant at 0.1 MPa. Any reduction method can advantageously be considered, either in situ (the reduction of the catalyst is carried out in the unit itself where the catalytic reaction is carried out), or ex situ (the reduction is carried out outside the unit where the catalytic reaction is carried out, before loading the catalyst into the unit).

[0177] The operating conditions implemented in step e) make it possible to obtain an isomerized effluent containing little or no olefinic compounds and oxygenated compounds.

[0178] The isomerized effluent from step e) is then sent to step h'). Step f)

[0179] According to the invention, the process includes recycling all or part of the residual heavy fraction having an initial boiling point above 250°C in step c) of hydrocracking.

[0180] In the variant where steps b) and d) are carried out in a single fractionation unit, the hydrocracked effluent from step d2'), part of which may be purged, is advantageously recycled in said common fractionation unit mixed with all or part of the effluent from step a) of hydrotreatment.

[0181] In this case, said common fractionation unit makes it possible to obtain at least one heavy fraction which includes the heavy fraction of the hydrotreated effluent and the residual heavy fraction of the hydrocracked effluent.

[0182] In said variant, said heavy fraction which is sent to the hydrocracking step c) comprises the heavy fraction of the hydrotreated effluent and the residual heavy fraction of the hydrocracking effluent.

[0183] In the variant where steps b) and d) are carried out in separate fractionation units, only the hydrocracked effluent from step d2'), part of which may be purged, is fractionated in said fractionation step d) to obtain at least one residual heavy fraction having an initial boiling point above 250°C.

[0184] Preferably, all or part of said residual heavy fraction, possibly purged, is advantageously recycled directly into said hydrocracking step c).

[0185] Recycling the residual heavy fraction is a key step in the invention because this recycling step allows at least a portion of the residual heavy fraction to be converted into at least a kerosene fraction. This increases the kerosene yield of the process. Furthermore, it is more advantageous to operate with a low conversion rate per pass and with recycling than to attempt to convert more of the heavy fraction in a single pass without recycling, since the latter would lead to overcracking the product and thus the formation of a significant quantity of gas and / or gasoline.

[0186] Said recycle is advantageously operated with a rate corresponding to 1 to 19 times the mass flow rate of fresh feed from the hydrotreatment step a) and separated in step b). Preferably, the recycle rate is between 2 and 9 and even more preferably, the recycle rate is operated between 3 and 6.

[0187] The recycle rate is defined as the ratio of the mass flow rate of feedstock entering the hydrocracking step c), i.e., the cumulative flow rate of the heavy fraction and the residual heavy fraction to the flow rate of the heavy fraction. It can be directly calculated from the per-pass conversion previously defined according to the following formula:

[0188] Recycle Tx = 1 / (Conversion of Tx+ to Tx)

[0189] The recycle rate is then advantageously between 2 and 20, preferably it is between 3 and 10 and even more preferably it is between 4 and 6. Step g)

[0190] An optional purging of a portion of the liquid hydrocarbon fraction from step dl') may be carried out. The purging constitutes a fraction that will not be recycled in any step of the process according to the invention and in particular neither to the hydrocracking step c) nor to the hydroisomerization step e).

[0191] Generally it is a part of the residual heavy fraction and advantageously the heaviest fraction of said residual heavy fraction.

[0192] The implementation of the purging makes it possible to maximize the amount of residual heavy fraction recycled to the hydrocracking step c) while maintaining the performance of the process, the possible heavy compounds refractory to hydrocracking that may have accumulated during the recycle steps can be eliminated by the purging.

[0193] Said purging can advantageously be carried out upstream or downstream of the fractionation step and preferably downstream of the fractionation step d) when steps b) and d) are carried out in separate fractionation units.

[0194] In another embodiment, the purging is carried out during, upstream or downstream of the fractionation step and preferably downstream of the fractionation step when steps b) and d) are carried out in the same fractionation unit.

[0195] Preferably the purge corresponds to less than 10% by weight of the total incoming load in the hydrotreating step a), very preferably less than 7% by weight, very preferably less than 5% by weight and even more preferably less than 3% by weight.

[0196] The purging can be carried out by all methods known to a person skilled in the art and can, for example, consist of a draw-off or a simple flash, the separation temperature then being chosen as high as possible, preferably above 300°C, preferably above 370°C, most preferably above 450°C and even more preferably above 540°C.

[0197] The effluent from step c), purge deducted is then separated and then fractionated. Step h')

[0198] According to the invention, the process according to the invention comprises a separation step h') of the effluent from step e) into at least a hydroisomerized hydrocarbon liquid fraction and a gaseous fraction.

[0199] Said separation step h'), can advantageously be carried out according to all methods and techniques known to those skilled in the art. Preferably, said separation step h') is carried out in one or more cold separator vessels advantageously operating at a total pressure as close as possible to the outlet pressure of the hydroisomerization step e), and preferably between 0 and 1 MPa lower than the total outlet pressure of step e).

[0200] Preferably said step h') is carried out at a temperature between 20 and 80°C and preferably between 30 and 70°C and preferably between 40 and 60°C.

[0201] Preferably, said gaseous fraction separated in said step h') advantageously comprises hydrogen and the light gases Ci-C4 and optionally comprises a gaseous fraction heavier than the Ci-C4 fraction and with a boiling point below 100°C.

[0202] Preferably, said gaseous fraction comprises less than 1% by weight of compounds having a boiling point above 100°C, relative to the total weight of said fraction.

[0203] Preferably, said separation step h') is carried out in separation units distinct from said separation steps bl'), b2') and dl') and d2'). Step h)

[0204] The process according to the invention comprises a step h) of fractionating all or part of the hydroisomerized hydrocarbon liquid fraction from step h') to obtain at least one gaseous fraction advantageously comprising light C1-C4 gases, at least one naphtha fraction having an initial boiling point between 30 and 60°C and a final boiling point between 120 and 160°C, and at least one kerosene fraction having an initial boiling point between 120 and 160°C.

[0205] At the end of this fractionation step h), the kerosene obtained has all the specifications required to be used as aviation fuel according to current standards.

[0206] A mass balance demonstrates excellent selectivity of the process towards this final kerosene fraction, with unwanted fractions produced in reduced proportions compared to other implementations described in the literature to our knowledge. In particular, the yield of fractions with a boiling point higher than the target end point of kerosene has been reduced to a minimum, as required by the invention, with the purge being less than 10% by weight of the incoming feed mass in step a). Furthermore, the yield of light fractions from cracking during one or more of steps a), c), or e) is also extremely reduced due to the serial treatment of the heavy fraction of the incoming feed, first on a hydrocracking catalyst with a low conversion rate per pass, then on a hydroisomerization catalyst selective towards isomerization at the expense of cracking. List of figures

[0207] [Fig.1]

[0208] Fig. 1 illustrates the implementation of the process according to the prior art in which all the liquid hydrocarbon fractions from the separation steps bl'), b2'), d1') and d2') are sent to the fractionation step b).

[0209] [Fig.2]

[0210] Fig. 2 illustrates the implementation of the process according to a preferred embodiment of the invention in which all of the hot liquid hydrocarbon fraction from bl') and all of the hot hydrocracked liquid hydrocarbon fraction from dl') are sent directly into the hydrocracking step c).

[0211] [Fig.3]

[0212] Fig. 3 illustrates the implementation of the process according to a preferred embodiment of the invention in which only the entire hot hydrocracked liquid hydrocarbon fraction from dl') is sent directly into the hydrocracking step c).

[0213] [Fig.4]

[0214] Figure 4 illustrates the implementation of the process according to an embodiment of the invention in which the separation steps bl') and dl') are carried out in the same unit and the steps b2') and d2') are carried out in the same separation unit.

[0215] In [Fig. 1] illustrating the prior art, the synthesis gas 1, composed mainly of carbon monoxide and hydrogen, is directed to a Fischer-Tropsch synthesis unit (A'). At the outlet of unit (A'), the effluent is divided into two streams: the hydrocarbon fraction 2, which includes the cold condensate and the heavy fraction corresponding to the waxes, and the non-cold-condensable gaseous fraction 3, which can be treated directly or after possible reprocessing in the Fisher-Tropsch synthesis unit (recycle). Fraction 2 is sent to the hydrotreating unit (A), whose effluent 4 is separated into two streams in a hot tank (B1'). The gaseous fraction 5 is sent to a cold tank (B2'). The hot liquid hydrocarbon fraction 6 from Bl' is fractionated in the fractionation unit (B / D) into a gaseous fraction 7, a naphtha fraction 8, a kerosene fraction 9 and a heavy fraction 10.In the cold flask B2', 3 streams are separated: the aqueous fraction 11 which is not treated in the process, the gaseous fraction 12 and the liquid hydrocarbon fraction 13 which is fractionated in the fractionation unit (B / D), together with fraction 6, into a gas fraction 7, a naphtha fraction 8, a kerosene fraction 9 and a heavy fraction 10. .

[0216] The heavy fraction from the fractionation unit (B / D) is sent to a hydrocracking unit (C) where a partially converted effluent 14 is produced. This effluent is then separated into two streams in a separation unit D1' comprising a hot tank. The hot hydrocraced liquid hydrocarbon fraction 16 is then sent together with the liquid hydrocarbon fraction 13 from the separation unit B2' and with the fraction 6 from the separation unit B1' to the fractionation unit B / D.The gaseous fraction from Dl' 15 is sent to a separation unit D2' comprising a cold flask generating two streams: a gaseous fraction 17, a liquid hydrocracked hydrocarbon fraction 18 which is then fractionated in the fractionation unit B / D jointly with the liquid hydrocarbon fraction 13 from the separation section B2', with the fraction 6 from the separation section Bl' and with the fraction 16 from the separation section Dl', into a gas fraction 7, a naphtha fraction 8, a kerosene fraction 9 and a heavy fraction 10. Before their fractionation, the liquid hydrocarbon fraction 13 from the separation section B2', the fraction 6 from the separation section Bl', the fraction 16 from the separation section Dl' and the fraction 18 from the separation section D2' are combined into a fraction 19.

[0217] The kerosene hydrocarbon cut 9 produced by the fractionation unit (B / D) is then sent to a hydroisomerization unit (E) whose effluent 20 is separated into two streams in a separation unit H'): a gaseous fraction 21 and a hydroisomerized hydrocarbon liquid fraction 22 which is then fractionated in a unit (H) fractionation at the end of which three fractions are produced: a kerosene fraction 23, a naphtha fraction 24 and finally a gas fraction 25.

[0218] Figure 2, illustrating one embodiment of the invention, is similar to the [Fig.1], with the following exceptions.

[0219] The hot liquid hydrocarbon fraction 6 from Bl' and the hot hydrocracked liquid hydrocarbon fraction 16 from Dl' are combined into a stream 26 which is then sent directly into the hydrocracking unit (C) and not into the fractionation unit (B / D).

[0220] The fractionation unit (B / D) is therefore fed by a flow 19 consisting of the grouping of only the liquid hydrocarbon fractions 13 and 18 respectively from the separation sections B2' and D2'.

[0221] Fig. 3 illustrating a second embodiment of the invention is similar to Fig. 1, except for the following: the hot hydrocracked liquid hydrocarbon fraction 16 from D1' is sent directly into the hydrocracking unit (C) and not into the fractionation unit (B / D).

[0222] Figure 4 illustrates yet another embodiment of the invention. similar to [Fig.2], with the following exceptions.

[0223] The separation steps bl') and dl') are carried out in the same separation unit B17D1' and the separation steps b2') and d2') are carried out in the same separation unit B27D2'.

[0224] The fraction 14 from the hydrocracking unit (C) is then combined with the stream 4 from the hydrotreatment to form the stream 27 which feeds the separation unit B17 Dl'.

[0225] Fractions 6 and 16 are no longer isolated, but it is directly the liquid hydrocarbon stream 26 which is produced at the exit of the separation section B17D1' which feeds the hydrocracking unit (C).

[0226] Fractions 5 and 15 are no longer isolated, but only the gaseous flow 28 from the common separation section B17D1' feeds the separation section B27D2'.

[0227] The aqueous fractions 11 and 17 are no longer isolated, but only the aqueous flow 29 comes from the separation section B27D2'.

[0228] Finally, the fluxl9 is directly from the separation section B2' / D2' without prior mixing.

[0229] The examples illustrate the invention without limiting its scope. Examples

[0230] Example 1: preparation of the hydrotreating catalyst Cl (compliant).

[0231] The catalyst is an industrial catalyst based on nickel, molybdenum and phosphorus on alumina with molybdenum oxide (MoO3) contents of 22 wt%, in oxide of nickel NiO of 4% by weight and of phosphorus oxide P2O5 of 5% by weight relative to the total weight of the finished catalyst.

[0232] Example 2: preparation of the hydrocracking catalyst C2 (compliant).

[0233] The silica-alumina powder is prepared according to the synthesis protocol described in patent EPI 415 712A. The quantities of orthosilicic acid and aluminum hydrate are chosen so as to have a composition of 70% by weight of alumina A12O3 and 30% by weight of silica SiO2 in the final solid.

[0234] This mixture is rapidly homogenized in a commercial colloidal mill in the presence of nitric acid so that the nitric acid content by weight of the suspension exiting the mill is 8% relative to the silica-alumina solid. The suspension is then conventionally dried in a spray dryer from 300°C to 60°C. The powder thus prepared is shaped in a Z-arm in the presence of 8% nitric acid relative to the anhydrous product. Extrusion is carried out by passing the paste through a die with 1.4 mm diameter orifices. The extrudates thus obtained are oven-dried at 140°C and then calcined under a flow of dry air at 550°C and finally calcined at 850°C in the presence of steam.

[0235] The characteristics of the support thus prepared are as follows:

[0236] - an average mesopore diameter measured by mercury porosimetry of 8.0 nm,

[0237] - a total pore volume measured by mercury porosimetry of 0.49 ml / g,

[0238] - a mesoporous volume measured by mercury porosimetry of 0.48 ml / g,

[0239] - a volume of macropores, the diameter of which is greater than 50 nm less than 0.02 ml / g,

[0240] - a BET surface of 235 m2 / g.

[0241] 50 grams of silica-alumina extrudates are then subjected to a step The catalysts were dry-impregnated with an aqueous solution of platinum tetramine nitrate, left to mature in a water-cooled maturing tank for 24 hours at room temperature, and then calcined for two hours under dry air in a flow bed at 450°C (temperature ramp of 5°C / min). The platinum content by weight of the finished catalyst after calcination was 0.16%, its dispersion measured by H2 / O2 titration was 33%, and its distribution coefficient measured by Castaing microprobe was 0.96.

[0242] Example 3: Preparation of a C3 hydroisomerization catalyst (conforming). Synthesis of ZEOLITH IZM-2.

[0243] Zeolite IZM-2 was synthesized in accordance with the teachings of French patent FR 2 918 050 B. A colloidal silica suspension known by the trade name Ludox HS-40, marketed by Aldrich, is incorporated into a solution composed of sodium hydroxide (Prolabo), a dibromide structuring agent The mixture consists of 1,6-bis(methylpiperidinium)hexane, aluminum hydroxide (Aldrich), and deionized water. The molar composition of the mixture is as follows: 1 SiO2; 0.0060 Al2O3; 0.1666 Na2O; 0.1666 1,6-bis(methylpiperidinium)hexane; 33.3333 H2O. The mixture is vigorously stirred for half an hour. After homogenization, the mixture is transferred to a PARR-type autoclave. The autoclave is heated for 5 days at 170°C with agitation using a rotisserie (30 rpm). The resulting product is filtered, washed with deionized water to achieve a neutral pH, and then dried overnight at 100°C in an oven. The solid is then placed in a muffle furnace for calcination to remove the structuring agent. The calcination cycle includes a temperature rise to 200°C, a two-hour plateau at this temperature, a temperature rise to 550°C followed by an eight-hour plateau at this temperature, and finally a return to ambient temperature.The temperature ramps are carried out at a rate of 2°C / min. The resulting solid is then refluxed for 2 hours in an aqueous ammonium nitrate solution (10 mL of solution per gram of solid, 3 M ammonium nitrate concentration) to exchange the alkali sodium cations for ammonium ions. This reflux step is repeated four times with fresh ammonium nitrate solution. The solid is then filtered, washed with deionized water, and dried overnight in an oven at 100°C. Finally, to obtain the zeolite in its acidic form (protonated H+), a calcination step is performed at 550°C for ten hours (temperature ramp of 2°C / min) in a flow-through bed under dry air (2 normal liters per hour per gram of solid). The resulting solid was analyzed by X-ray diffraction and identified as IZM-2 zeolite.Characterizations using 27A1 isotope NMR, X-ray fluorescence and ICP methods provide the following results for 1'IZM-2: .

[0244] - weight percentage of A1VI hexacoordinate aluminum atoms: 5%,

[0245] - ratio of the number of moles of silicon divided by the number of moles of aluminum lattice density, in mole / mole, Si / Al: 72,

[0246] - ratio of the number of moles of sodium divided by the number of moles of aluminium network, in mole / mole, Na / Al: 0.03. Preparation of the IZM-2 / alumina support.

[0247] The IZM-2 / alumina support is obtained by mixing and extruding IZM-2 zeolite with a Pural SB3 type alumina gel. The mixed paste is extruded through a 1.8 mm diameter quadrilobe die. After drying overnight in an oven at 110°C, the extrudates are calcined at 500°C for two hours (temperature ramp of 5°C / min) in a flow bed under dry air (2 normal liters per hour per gram of solid). The weight content of IZM-2 zeolite in the support after calcination is 13% by weight.

[0248] Platinum impregnation on the IZM-2 / alumina support.

[0249] Platinum impregnation is carried out by dry impregnation of the support with an aqueous solution containing platinum tetramine nitrate Pt(NH3)4(NO3)2. 50 grams of support are used and dry-impregnated in a dripping container. After impregnation, the solid is left to mature for at least five hours in laboratory air and then dried overnight in an oven at 110°C. Finally, a calcination step is carried out under a flow of dry air (2 normal liters per hour per gram of solid) in a tubular furnace under the following conditions:

[0250] - temperature rise from ambient to 450°C at 5°C / min;

[0251] - two-hour plateau at 450°C;

[0252] - descent to ambient temperature.

[0253] The Pt content measured by FX on the calcined C3 catalyst is 0.29% by weight, its dispersion measured by H2 / O2 titration is 52%, its distribution coefficient measured by Castaing microprobe is 0.91.

[0254] Example 4: Preparation of a C4 hydroisomerization catalyst (compliant). Zeolite ZSM-12.

[0255] ZSM-12 zeolite is supplied by Zeolyst. The solid was analyzed by X-ray diffraction and identified as being composed of ZSM-12 zeolite. Characterizations using 27Al isotope NMR, X-ray fluorescence, and ICP methods yielded the following results for ZSM-12:

[0256] - weight percentage of A1VI hexacoordinated aluminium atoms: 0%;

[0257] - ratio of the number of moles of silicon divided by the number of moles of aluminum network, in mole / mole, Si / Al: 43;

[0258] - ratio of the number of moles of sodium divided by the number of moles of aluminium network, in mole / mole, Na / Al: 0.009. Preparation of the ZSM-12 / alumina support.

[0259] The ZSM-12 / alumina support is obtained by mixing and extruding ZSM-12 zeolite with a Pural SB3 type alumina gel. The mixed paste is extruded through a 1.8 mm diameter quadrilobe die. After drying overnight in an oven at 110°C, the extrudates are calcined at 500°C for two hours (temperature ramp of 5°C / min) in a flow bed under dry air (2 normal liters per hour per gram of solid). The weight content of ZSM-12 zeolite in the support after calcination is 6 wt%.

[0260] Platinum impregnation on the ZSM-12 / alumina support.

[0261] Platinum impregnation is carried out by dry impregnation of the support with an aqueous solution containing platinum tetramine nitrate Pt(NH3)4(NO3)2. 50 grams of support are used and dry-impregnated in a dripping container. After impregnation, the solid is left to mature for at least five hours in laboratory air and then dried overnight in an oven at 110°C. Finally, a calcination step is carried out under a flow of dry air (2 normal liters per hour per gram of solid) in a tubular furnace under the following conditions:

[0262] - temperature rise from ambient to 450°C at 5°C / min;

[0263] - two-hour plateau at 450°C;

[0264] - descent to ambient temperature.

[0265] The Pt content measured by FX on the calcined C4 catalyst is 0.31% by weight, its dispersion measured by H2 / O2 titration is 38%, its distribution coefficient measured by Castaing microprobe is 0.87.

[0266] Example 5: Process for producing kerosene according to prior art.

[0267] Example 5 is illustrated in [Fig.1].

[0268] Step a) hydrotreating of the paraffinic feedstock. (Catalytic unit A).

[0269] A paraffinic filler from the Fischer-Tropsch synthesis has the characteristics given in Table 1. This filler includes cold condensates and waxes.

[0270] Table 1: Characteristics of the charge from the Fischer-Tropsch synthesis

[0271] [Tables 1] Simulated Distillation Temperature (5% wt.): 120 °C (25% wt.): 215 °C (50% wt.): 318 °C (75% wt.): 421 °C (95% wt.): 546 °C Compounds 300°C+ (by GC) 55% wt. Density at 15°C 0.7679. Nitrogen content <2 ppm. Sulfur content <2 ppm. Total atomic oxygen content (by IR) 8700 ppm. Mass flow rate 17 tons / h

[0272] The total atomic oxygen content in said fraction is measured by the infrared absorption technique described in patent application US2009 / 0018374A1.

[0273] The oxygen content represents the content of molecules comprising at least one oxygen atom present in the heavy fraction and is expressed as a weight percentage relative to the total pass of said fraction. The oxygen content is measured by gas chromatography.

[0274] The paraffinic charge is treated in a flow bed, on the hydrotreating catalyst Cl under operating conditions which allow the elimination of olefinic and oxygenated compounds as well as traces of nitrogen.

[0275] The selected operating conditions are as follows: - hourly volumetric rate WH (charge volume / catalyst volume / hour) = 2 h - Total working pressure: 7 MPa, - hydrogen / charge ratio: 700 normal liters / liter, - Temperature: 330°C.

[0276] Before testing, the catalyst Cl undergoes a reduction step under the following operating conditions: - pure hydrogen flow rate: 1600 normal liters per hour per liter of catalyst, - rising ambient temperature to 120°C: 10°C / min, - one-hour plateau at 120°C, - temperature rise from 120°C to 450°C at 5°C / min, - two-hour plateau at 450°C, - Total pressure: 7 MPa

[0277] Hot separation step bl') (separation unit Bl').

[0278] The effluent from step a) is sent to a separation unit Bl', comprising a separator vessel operating at a pressure of 6.9 MPa and a temperature of 330°C. The separation unit Bl' generates the following two streams: - A hot liquid hydrocarbon stream collected in liquid form and sent to step b) of fractionation. - A hot gas stream is collected and directed to step b2') after partial condensation.

[0279] Step b2') of cold separation (separation unit B2').

[0280] The hot gas stream from step bl') is partially condensed and sent to a separation unit B2', comprising a separator vessel operating at a pressure of 6.9 MPa and a temperature of 48°C. The separation unit B2' generates the following three streams: - An aqueous flow consisting of at least part of the water formed during catalytic unit A and preferably of all the water formed. - A liquid hydrocarbon stream collected in liquid form and sent to step b) of fractionation. - A gaseous flow.

[0281] The liquid hydrocarbon flows from steps bl') and b2') are directed to the fractionation step b).

[0282] Step b) Fractionation of the hydrotreated and separated paraffinic filler. (Fractioning Unit B / D)

[0283] The liquid hydrocarbon streams from steps bl') and b2)' are fractionated by atmospheric distillation into different hydrocarbon cuts, namely light gases from C1 to C4>, a naphtha cut with an initial boiling point of 35°C and a final boiling point of 120°C, a kerosene cut with an initial boiling point of 120°C and a final boiling point of 300°C, and a heavy fraction with an initial boiling point greater than 300°C.

[0284] Step c) hydrocracking of the heavy fraction from step b). (Catalytic unit C)

[0285] The heavy fraction from step b) mixed with the residual heavy fraction from step d) below constitutes the hydrocracking feed sent to the hydrocracking catalyst C2.

[0286] Before testing, the C2 catalyst undergoes a reduction step under the following operating conditions:

[0287] - pure hydrogen flow rate: 1600 normal liters per hour per liter of catalyst,

[0288] - rise from ambient temperature to 120°C: 10°C / min,

[0289] - one-hour plateau at 120°C,

[0290] - temperature rise from 120°C to 450°C at 5°C / min,

[0291] - two-hour plateau at 450°C,

[0292] - pressure: 6.5 MPa

[0293] After reduction, the catalytic test is carried out under the following conditions:

[0294] - total pressure of 6.5 MPa,

[0295] - hydrogen-to-charge ratio of 800 normal litres / litre,

[0296] - hourly volumetric velocity (WH) equal to 1 h1.

[0297] The conversion of the fraction 300°C+ is taken as equal to:

[0298] C(300°C+) = [ (% of 300°C effluents ) - (% of 300°C load) ] / [ 100 - (% of 300°C load)]

[0299] with

[0300] % of 300°C effluents = mass percentage of compounds having points boiling points below 300°C in effluents,

[0301] and

[0302] % of 300°C charge = mass percentage of compounds having points boiling points below 300°C in the hydrocracking feed.

[0303] The reaction temperature is adjusted to obtain a conversion level of the 300°C+ fraction equal to 0.65. This corresponds to a recycle rate equal to 1.5.

[0304] Hot separation step dl') (separation unit Dl').

[0305] The effluent from step c) is sent to a separation unit Dl', comprising a separator vessel operating at a total pressure of 6.4 MPa and a temperature of 330°C. The separation unit Dl' generates the following two streams: - A hot liquid hydrocarbon stream collected in liquid form and sent to step d) of fractionation. - A hot gas stream is collected and directed towards step d2').

[0306] Cold separation step d2'l (separation unit D2'l

[0307] The gaseous effluent from step d1') is sent, after partial condensation, to a separation unit D2' comprising a separator vessel operating at a pressure of 6.1 MPa and a temperature of 48°C. The separation unit D2' generates the following two streams: - A liquid hydrocarbon stream collected in liquid form and sent to step d) of fractionation. - A gaseous flow.

[0308] The hydrocarbon flows from d1') and d2') are directed to the fractionation step d). The fractionation step d) and the fractionation step b) are carried out in the same fractionation unit B / D.

[0309] Step d) Fractionation of the hydrocracking effluent. (Fractioning Unit B / D)

[0310] The hydrocracking effluent obtained from steps d1') and d2') is fractionated by atmospheric distillation into different hydrocarbon fractions, namely light gases from C1 to C4, a naphtha fraction with an initial boiling point of 35°C and a final boiling point of 120°C, a kerosene fraction with an initial boiling point of 120°C and a final boiling point of 300°C, and a residual heavy fraction with an initial boiling point above 300°C. This residual heavy fraction is then recycled to the hydrocracking step c). The fractionation of the hydrocracking effluent is carried out in the same distillation column as that used in step b). Therefore, there is a joint fractionation of the hydrocracking effluent and the liquid hydrocarbon effluents from steps b1') and b2').

[0311] Step e) Hydroisomerization of the kerosene fraction from step b). (Catalytic unit E)

[0312] The kerosene fraction from step b) constitutes the hydroisomerization feed supplied to the hydroisomerization catalyst C3. Before testing, catalyst C3 undergoes a reduction step under the following operating conditions:

[0313] - pure hydrogen flow rate: 1600 normal liters per hour per liter of catalyst,

[0314] - rise from ambient temperature to 120°C: 10°C / min,

[0315] - one-hour plateau at 120°C,

[0316] - temperature rise from 120°C to 450°C at 5°C / min,

[0317] - two-hour plateau at 450°C,

[0318] - pressure: 6.5 MPa.

[0319] After reduction, the catalytic test is carried out under the following conditions:

[0320] - total pressure of 6.5 MPa,

[0321] - hydrogen-to-charge ratio of 800 normal litres / litre,

[0322] - hourly volumetric velocity (WH) equal to 2 h1.

[0323] The temperature is adjusted so as to obtain a value of -40°C for the disappearance temperature of the crystals for the kerosene obtained after the fractionation step h).

[0324] Step h') Separation of the hydroisomerization effluent. (Separation Unit H')

[0325] The hydroisomerized effluent then undergoes a flash separation and settling step. Following this separation, carbon monoxide and / or carbon dioxide, water, ammonia, and a gaseous fraction comprising hydrocarbons with fewer than five atoms formed during hydroisomerization are separated. A hydroisomerized hydrocarbon liquid fraction is also separated.

[0326] Step h) Fractionation of the separated hydroisomerization effluent. (Fractioning Unit H)

[0327] The hydroisomerization effluent obtained at the end of step h') is fractionated by atmospheric distillation into different hydrocarbon cuts, namely light gases from C1 to C4, a naphtha cut with an initial boiling point of 35°C and a final boiling point of 120°C, and a kerosene cut with an initial boiling point of 120°C and a final boiling point of 300°C. This fractionation is carried out in an atmospheric distillation unit H.

[0328] Material balances performed over 24 hours allow the calculation of yields in different cuts: - yield in light gases: mass of gas Ci to C4 / mass of paraffinic filler * 100, - Naphtha cutting yield: mass of liquid effluent with an initial boiling point of 35°C and a final boiling point of 120°C / mass of paraffinic feedstock * 100, - Kerosene cutting yield: mass of liquid effluent with an initial boiling point of 120°C and a final boiling point of 300°C / mass of paraffinic feedstock * 100,

[0329] where - the mass of C4 Ci gas corresponds to the quantity of C4 Ci gas generated during 24 hours at the end of step a) hydrotreating the paraffin feed from the Fischer-Tropsch synthesis, and steps b) and d) of joint fractionation of the hydrotreatment effluent and separation of the paraffin charge from the Fischer-Tropsch synthesis and the hydroconversion effluent and step h) of fractionation of the hydroisomerization effluent; - The mass of paraffinic feed corresponds to the quantity of paraffinic feed from the Fischer-Tropsch synthesis consumed during 24 hours; the mass of liquid effluent with an initial boiling point of 35°C and a final boiling point of 120°C corresponds to the quantity of liquid effluent with an initial boiling point of 35°C and a final boiling point of 120°C generated during 24 hours following steps b) and d) of joint fractionation of the hydrotreating effluent and separation of the paraffinic feed from the Fischer-Tropsch synthesis and the hydrocracking effluent, and step h) of fractionation of the hydroisomerization effluent - The mass of liquid effluent with an initial boiling point of 120°C and a final boiling point of 300°C corresponds to the quantity of liquid effluent with an initial boiling point of 120°C and a final boiling point of 300°C generated during 24 hours following h) fractionation of the hydroisomerization effluent. The disappearance point of the crystals of said kerosene cut is then measured according to ASTM D5972.

[0330] Example 6: Kerosene production process not according to the invention with a high conversion level equal to 0.65 on step c)

[0331] Example 6 is illustrated in [Fig. 3]. In particular, the liquid hydrocarbon flow

[0332] hot from step dl') is directly directed to step c) of hydroconversion without going through the fractionation step d), which reduces the external energy input required to separate the effluents from step c). Unlike the invention, the conversion rate set for step c) is 0.65

[0333] Step a) hydrotreatment of the paraffin filler.

[0334] Identical to example 5.

[0335] Step b) fractionation of the hydrotreated and separated paraffinic filler.

[0336] Identical to example 5.

[0337] Step c) hydrocracking of the heavy fraction from step b).

[0338] Identical to example 5.

[0339] Step d) fractionation of the hydrocracking effluent.

[0340] Identical to example 5.

[0341] Step e) hydroisomerization of the kerosene fraction from step b).

[0342] Identical to example 5.

[0343] Step h') separation of the hydroisomerization effluent.

[0344] Identical to example 5.

[0345] Step h) fractionation of the separated hydroisomerization effluent.

[0346] Identical to example 5. Steps bl'), b2') and steps dl') and d2')

[0347] Identical to example 5, except that the hot liquid hydrocarbon flow from dl') is directed to the hydroconversion step c) without going through the fractionation step d).

[0348] Example 7: Process for producing kerosene according to prior art with a low conversion level equal to 0.2 in step c)

[0349] Example 7 is illustrated in [Fig. 1]. This example is identical to example 5, except that the conversion level of the 300°C+ fraction in step c) is 0.2, compared to 0.65 in example 5.

[0350] Step a) hydrotreatment of the paraffin filler.

[0351] Identical to example 5.

[0352] Step b) fractionation of the hydrotreated and separated paraffinic filler.

[0353] Identical to example 5.

[0354] Step c) hydrocracking of the heavy fraction from step b).

[0355] Identical to example 5 except that the reaction temperature is adjusted in such a way to obtain a conversion level of the 300°C+ fraction equal to 0.2. This corresponds to a recycle rate equal to 5.

[0356] Step d) fractionation of the hydrocracking effluent.

[0357] Identical to example 5.

[0358] Step e) hydroisomerization of the kerosene fraction from step b).

[0359] Identical to example 5. Step bl'), b2') and Step dl') and d2')

[0360] Identical to example 5.

[0361] Step h') separation of the hydroisomerization effluent.

[0362] Identical to example 5.

[0363] Step h) fractionation of the separated hydroisomerization effluent.

[0364] Identical to example 5.

[0365] Example 8: Kerosene production process according to the invention with a low conversion level equal to 0.2 in step c)

[0366] Example 8 is illustrated in [Fig. 3]. In particular, the hot liquid hydrocarbon stream from step dl') is directed directly to hydroconversion step c) without passing through fractionation step d). This example is identical to example 6, except that the conversion level of the 300°C+ fraction in step c) is 0.2, compared to 0.65 in example 6.

[0367] Step a) hydrotreatment of the paraffin filler.

[0368] Identical to example 6.

[0369] Step b) fractionation of the hydrotreated and separated paraffinic filler.

[0370] Identical to example 6.

[0371] Step c) hydrocracking of the heavy fraction from step b).

[0372] Identical to example 6 except that the reaction temperature is adjusted in such a way to obtain a conversion level of the 300°C+ fraction equal to 0.2. This corresponds to a recycle rate equal to 5.

[0373] Step d) fractionation of the hydrocracking effluent.

[0374] Identical to example 6.

[0375] Step e) hydroisomerization of the kerosene fraction from step b).

[0376] Identical to example 6.

[0377] Step h') separation of the hydroisomerization effluent.

[0378] Identical to example 6.

[0379] Step h) fractionation of the separated hydroisomerization effluent.

[0380] Identical to example 6. Steps bl'), b2') and steps dl') and d2')

[0381] Identical to example 6.

[0382] Example 9: Kerosene production process according to the invention

[0383] Example 9 is illustrated in [Fig.2]. In particular, according to the invention, the liquid hydrocarbons from bl') and dl') are directed to the hydroconversion step c) without going through the fractionation steps b) and d), thereby reducing the external energy input required to separate the effluents from steps a) and c).

[0384] Step a) hydrotreatment of the paraffin filler.

[0385] Identical to example 7.

[0386] Step b) fractionation of the hydrotreated and separated paraffinic filler.

[0387] Identical to example 7.

[0388] Step c) hydrocracking of the heavy fraction from step b).

[0389] Identical to example 7.

[0390] Step d) fractionation of the hydrocracking effluent.

[0391] Identical to example 7.

[0392] Step e) hydroisomerization of the kerosene fraction from step b).

[0393] Identical to example 7.

[0394] Step h') separation of the hydroisomerization effluent.

[0395] Identical to example 7.

[0396] Step h) fractionation of the separated hydroisomerization effluent.

[0397] Identical to example 7.

[0398] Step bl'). b2') and Step dl') and d2') Separation without energy input

[0399] Identical to example 7, except that the liquid hydrocarbons from bl') and dl') are directed to the hydroconversion step c) without going through the fractionation steps b) and d).

[0400] Example 10: Comparison of different processes using the pinch analysis method

[0401] The different processes were compared according to pinch analysis (Ian C. Kamp, Pinch Analysis and Process Integration: A User Guide on Process Integration for the Efficient Use of Energy, 2nd Edition - December 5, 2006, Elsevier), but also according to the yields of products of interest obtained. Table 2 below shows all the relevant data. It should be noted that the yields include consideration of hydrogen consumption (equal to the difference between the sum of the yields minus 100%).

[0402] The pinch method is a methodology for optimizing the energy consumption of a refining process based on the calculation of the minimum energy consumption achievable according to the laws of thermodynamics.

[0403] The use of this method allows a reliable comparison of the minimum energy consumption potential of the processes and confirms the intrinsic reduced energy requirements according to the invention.

[0404] Table 2: Yields in different cuts, properties of kerosene cuts and utility consumption for the different examples.

[0405] [Tables2] Comparative Example 5 Comparative Example 6 Comparative Example 7 Comparative Example 8 according to the invention Example 9 according to the invention Operating temperature β1', °C 330 330 330 330 330 Content at 300°C- in the hot liquid effluent at step β1' 30% 30% 30% 30% 30% Operating temperature dβ1', °C 330 330 330 330 330 Content at 300°C- in the hot liquid effluent at step dβ1' 32% 32% 13% 13% 13% Hot utility consumption at the fractionation furnace b) / d), GCal / h 1.87 1.03 2.60 1.91 1.44 Hot utility consumption at the inlet furnace of stage c), GCal / h 0.90 0.42 1.51 0.79 0.58 Estimated hot utility consumption using the pinch-off method at a minimum approach of (DTmin) 30°C BASE BASE-5 0% BASE BASE-3 5% BASE-5 0% Conversion rate / pass on C2, % 65.3 65.1 20 20 20 Light gas yield Ci to C4 3.4 3.8 2.9 3.0 3.0 Naphtha cut yield 27.2 28.8 24.2 24.0 24.3 Kerosene cut yield 69.7 67.7 73.2 73.3 73.0

[0406] Comparison of example 5 according to the prior art and comparative example 6 demonstrates that implementing the process according to the invention, without respecting the conversion criterion in step c), makes it possible to significantly reduce the external energy input required to separate the effluents in step a), but leads to a clear degradation of the yield in Kerosene cutting (loss of two points of yield), making the process uninteresting.

[0407] Comparison of example 7 according to the prior art and examples 8 and 9 demonstrates that the process according to the invention makes it possible to significantly reduce the external energy input required to separate the effluents from steps a) and c) and this without loss of kerosene cutting yield.

Claims

1. Demands A process for producing kerosene from a paraffinic feedstock produced by Fischer-Tropsch synthesis comprising at least the following steps and preferably consisting of: step a) hydrotreating at least part and preferably all of the paraffinic feedstock produced by Fischer-Tropsch synthesis in the presence of a hydrotreating catalyst and operating at a temperature between 250 and 450°C, at a total pressure between 0.5 and 15 MPa, a hydrogen flow rate adjusted to obtain a ratio between 100 and 3000 normal liters of hydrogen per liter of feedstock, and at an hourly volumetric rate between 0.1 and 40 h1, step b1') separating the effluent from step a) into at least a hot liquid hydrocarbon fraction and a hot gaseous fraction, said separation being carried out at a temperature above 200°C,step b2') the separation of the hot gaseous fraction from step b1') into at least an aqueous fraction consisting of at least some of the water formed during said step a) and preferably all of the water formed, a liquid hydrocarbon fraction and a gaseous fraction; step b) the fractionation in one or more steps of the hydrotreated liquid hydrocarbon fraction from step b2') into at least a gaseous fraction, a naphtha fraction having an initial boiling point between 30 and 60°C and a final boiling point between 100 and 160°C, at least a kerosene fraction having an initial boiling point between 100 and 160°C and a final boiling point between 250 and 320°C, and a heavy fraction having an initial boiling point above 250°C; step c) the hydrocracking of all of said heavy fraction having an initial boiling point above 250°C from step b) to produce a hydrocracked effluent,in the presence of a hydrocracking catalyst and operating at a temperature between 250 and 450°C, at a total pressure between 0.2 and 15 MPa, at a space velocity between 0.1 and 10 h1, and at a hydrogen flow rate adjusted to obtain a ratio between 100 and 2000 normal liters of hydrogen per liter of feed, said catalyst, hydrocracking comprising at least one noble metal of group VIII of the periodic table and a support comprising at least one acidic solid, said operating conditions of said step c) being adjusted so that the conversion by passing into products having boiling points greater than or equal to a temperature Tx into products having boiling points less than that same temperature Tx is less than 0.5, preferably less than 0.4 and preferably less than 0.3 and preferably less than 0.2, Tx being between 250 and 350°C, preferably between 270 and 340°C and most preferably between 280 and 330°C, step dl ') the separation of the effluent from step c) into at least one hot hydrocraced liquid hydrocarbon fraction and one hot gaseous fraction, said separation being carried out at a temperature greater than 200°C,said separation steps b1) and d1) advantageously being carried out in the same separation unit or in separate separation units, step d2') the separation of the hot gaseous fraction from step d1) into at least one liquid hydrocracked hydrocarbon fraction and one gaseous fraction, said separation steps b2') and d2') advantageously being carried out in the same separation unit or in separate separation units, step d) the fractionation in one or more steps of at least all or part of the liquid hydrocracked hydrocarbon fraction from step d2') into at least one gaseous fraction, a naphtha fraction having an initial boiling point between 30 and 60°C and a final boiling point between 100 and 160°C, at least one kerosene fraction having an initial boiling point between 100 and 160°C and a final boiling point between 250 and 320°C,and a residual heavy fraction having an initial boiling point above 250°C, said fractionation steps b) and d) advantageously being carried out in the same fractionation unit or in separate fractionation units, step e) the hydroisomerization of at least a portion of the kerosene fraction having an initial boiling point between 100 and 160°C and a final boiling point between 250 and 320°C from fractionation step b) and of all or part of the kerosene fraction having an initial boiling point between 100 and

2.

3. 160°C and a final boiling point between 250 and 320°C from the fractionation step d) in the case where said fractionation steps b) and d) are separate, in the presence of a hydroisomerization catalyst operating at a temperature between 200 and 450°C, a total pressure between 1 and 15 MPa, a space velocity between 0.1 and 10 h1, a hydrogen flow rate adjusted to obtain a ratio between 100 and 2000 normal liters of hydrogen per liter of feed, to produce a hydroisomerized effluent, step f) recycles all or part of the residual heavy fraction having an initial boiling point above 250°C in the hydrocracking step c), step g) optionally purging part of the hydrocracking effluent from the hydrocracking step c, step h') the separation of the effluent from step e) into at least one hydroisomerized hydrocarbon liquid fraction and one gaseous fraction, step h) the fractionation of all or part of the hydroisomerized hydrocarbon liquid fraction from step h'), into at least one gaseous fraction, a naphtha fraction having an initial boiling point between 30 and 60°C and a final boiling point between 120 and 160°C, and at least one kerosene fraction having an initial boiling point between 120 and 160°C, and wherein at least part and preferably all of the hot liquid hydrocarbon fraction from bl') and / or at least part and preferably all of the hot hydrocracked liquid hydrocarbon fraction from dl') are sent directly to hydrocracking step c) without passing through said fractionation step b) or d). A method according to claim 1 wherein said separation step bl') is carried out at a temperature above 230°C and more preferably above 250°C) most preferably below 390°C, in one or more hot separator vessels operating at a total pressure between 0 and 1 MPa lower than the total outlet pressure of step a) and preferably between 0 and 0.5 MPa lower than the total outlet pressure of step a). A method according to claim 1 or 2, wherein said separation step b2') is carried out in one or more cold separator vessels operating at a total pressure between 0 and 0.5 MPa lower than the total outlet pressure of step bl') and at a temperature between 20 and 80°C and preferably between 30 and 70°C and preferably between 40 and 60°C.

4. A method according to any one of the preceding claims wherein step c) operates at a temperature between 280 and 450°C, and more preferably between 320 and 420°C, at a total pressure between 0.5 and 10 MPa, more preferably between 1 and 9 MPa, at a space velocity between 0.2 and 7 h1, and more preferably between 0.5 and 5 h*, and with a hydrogen flow rate adjusted to obtain a ratio between 150 and 1500 normal liters of hydrogen per liter of charge and more preferably between 300 and 1500 normal liters of hydrogen per liter of charge.

5. A process according to any one of claims 1 to 4, wherein said hydroisomerization catalyst used in step e) comprises at least one noble metal of Group VIII of the periodic table and a support comprising at least one binder and at least one zeolite selected from the structural-type zeolites TON, FER, EUO, AEL, *MRE, MTW, MOR, *BEA and zeolite IZM-2, taken alone or in mixtures

6. A process according to claim 5 wherein the catalyst of step e) comprises and is preferably made of platinum, and a support comprising and preferably made of a ZSM-12 zeolite and an alumina binder.

7. A process according to claim 5 wherein the catalyst of step e) comprises and is preferably made of platinum, and a support comprising and preferably made of an IZM-2 zeolite and an alumina binder.

8. A method according to any one of claims 1 to 7 in which said step dl') is carried out under the same ranges of operating conditions as step bl').

9. A method according to any one of claims 1 to 8 in which said separation steps bl') and dl') are carried out in the same separation unit.

10. A method according to any one of claims 1 to 9 in which said step d2') is carried out under the same ranges of operating conditions as step b2').

11. A method according to any one of the preceding claims wherein said separation steps b2') and d2') are carried out in the same separation unit.

12. A method according to any one of claims 1 to 7 in which the 4 steps bl'), b2'), dl') and d2') are carried out in different separation units.

13. A method according to any one of claims 1 to 7 wherein steps bl') and dl') are carried out in different units and steps b2') and d2') are carried out in the same separation unit.

14. A method according to any one of claims 1 to 7 wherein steps bl') and dl') are carried out in the same separation unit and steps b2') and d2') are carried out in the same separation unit.

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