EP2158303A2

EP2158303A2 - Method for producing middle distillates by hydroisomerisation and hydrocracking of a heavy fraction from a fischer-tropsch effluent

Info

Publication number: EP2158303A2
Application number: EP08805641A
Authority: EP
Inventors: Nathalie Marchal-George; Jean Cosyns; Vincent Coupard; Eric Caprani; Damien Douziech; Aurélie DANDEU; Stéphane FEDOU
Original assignee: IFP Energies Nouvelles IFPEN; Eni SpA
Current assignee: IFP Energies Nouvelles IFPEN; Eni SpA
Priority date: 2007-06-12
Filing date: 2008-06-03
Publication date: 2010-03-03
Anticipated expiration: 2028-06-03
Also published as: FR2917419A1; WO2009004179A2; AU2008270132A1; WO2009004179A3; BRPI0813815A8; CN101730732B; EP2158303B1; RU2469069C2; RU2010100634A; US8709234B2; CN101730732A; BRPI0813815A2; MY157735A; US20100298451A1; CA2689932A1; ZA200908037B; FR2917419B1; CA2689932C

Abstract

The invention relates to a method in which the paraffin effluent from a Fischer-Tropsch synthesis unit is separated in order to obtain a C5+ heavy fraction, said heavy fraction being then hydrogenated in the presence of a hydrogenation catalyst at a temperature of between 80°C and 200°C, at a global pressure of between 0.5 and 6 MPa, at an hourly volumetric rate of between 1 and 10h-1, and at a hydrogen flow rate corresponding to a volumetric hydrogen/hydrocarbon ratio of between 5 and 80 Nl/l/h, the liquid hydrogenated effluent being then contacted with a hydroisomerisation / hydrocracking catalyst without any preliminary separation step, the hydroisomerised/hydrocracked effluent being then distilled for obtaining the middle distillates and optionally oil bases.

Description

PROCESS FOR PRODUCING AVERAGE DISTILLATES BY

HYDROISOMERIZATION AND HYDROCRACKING OF A HEAVY FRACTION

FROM A FISCHER-TROPSCH EFFLUENT

The present invention describes a process for the hydrocracking and hydroisomerization treatment of feedstocks from the Fischer-Tropsch process, making it possible to obtain middle distillates (gas oil, kerosene), ie initial boiling point cuts. at least 150 ° C and final of at most 34O ⁰ C and optionally oil bases.

In the Fischer-Tropsch process, the synthesis gas (CO + H ₂ ) is catalytically converted into oxygenates and substantially linear hydrocarbons in gaseous, liquid or solid form. However, these products, mainly made of normal paraffins, can not be used as such, in particular because of their cold-holding properties that are not very compatible with the usual uses of petroleum fractions. For example, the pour point of a linear hydrocarbon containing 20 carbon atoms per molecule (boiling point equal to about 350 ^° C., that is often included in the middle distillate cut) is + 37 ^° C. about which makes its use impossible, the specification being -15 ° C for diesel. Thus, the hydrocarbons from the Fischer-Tropsch process comprising mainly n-paraffins must be converted into more valuable products such as, for example, gas oil, kerosene, which are obtained, for example, after catalytic hydrocracking / hydroisomerization reactions. These products are generally free of heteroatomic impurities such as sulfur, nitrogen or metals. They contain practically no aromatics, naphthenes and more generally cycles, in particular in the case of cobalt catalysts.

On the other hand, they may have a significant content of unsaturated compounds of olefinic type and oxygenated products (such as alcohols, carboxylic acids, ketones, aldehydes and esters). These oxygenated and unsaturated compounds are more concentrated in the light fractions. Thus in the C5 + fraction corresponding to products boiling at an initial boiling point of between 20 ^° C. and 40 ° C., these compounds represent between 10-20% by weight of olefinic type unsaturated compounds and between 5-10% by weight. of oxygenated compounds.

One of the objectives of the invention is to eliminate, during a hydrotreatment step, upstream of a hydrocracking step, the olefinic type unsaturated compounds, said hydrotreatment step being carried out under less severe conditions. than those of the hydrocracking step. Unsaturated olefinic compounds present in the hydrocracking feeds reduce the service life of a hydrocracking catalyst. Indeed, under the severe hydrocracking / hydroisomerization operating conditions, the hydrogenation of the olefinic type unsaturated compounds being a strongly exothermic reaction, the transformation of the unsaturated compounds can have a negative impact on the hydroisomerization / hydrocracking step and cause, for example, a thermal runaway of the reaction, a large coking of the catalyst or the formation of gum by oligomerization.

One of the advantages of the invention is to provide a method for producing middle distillates from a paraffinic feedstock produced by Fischer Tropsch synthesis in which the hydrocracking step is preceded by a hydrogenation step allowing elimination previously and under less severe conditions than those used in the hydrocracking step, the most reactive elements and in particular unsaturated compounds of olefinic type.

State of the art

The Shell patent application (EP-583,836) describes a process for the production of middle distillates from a filler obtained by Fischer-Tropsch synthesis. In this process, the feedstock resulting from the Fischer-Tropsch synthesis can be treated in its entirety, but preferably the C4- fraction is withdrawn from the feedstock so that only the C5 + fraction boiling at a temperature above 20 ^° C. be introduced in the subsequent step. Said feedstock is subjected to a hydrotreatment to hydrogenate the olefins and alcohols, in the presence of a large excess of hydrogen, so that the conversion of products boiling above 370 ^° C. into products with a lower boiling point is less than 20%. The hydrotreated effluent consisting of high molecular weight paraffinic hydrocarbons is preferably separated from the hydrocarbon compounds having a low molecular weight and in particular the C4- fraction before the second hydroconversion stage. At least part of the remaining C5 + fraction is then subjected to a hydrocracking / hydroisomerization step with a conversion of products boiling above 370 ⁰ C into products with a boiling point of at least 40% by weight.

The present invention provides an alternative process for the production of middle distillates. The advantages of the present invention are: to protect the hydroisomerization / hydrocracking catalyst from the most reactive elements such as unsaturated compounds of olefinic type by the use upstream of the hydroisomerization / hydrocracking stage, a step of hydrogenation of the unsaturated compounds, the elimination of unsaturated compounds of the olefinic type before the hydroisomerization / hydrocracking step making it possible to avoid the formation of coke or gum in the hydroisomerization / hydrocracking zone, to facilitate the control of the temperature profile inside the hydroisomerization / hydrocracking zone by the upstream implementation of the hydroisomerization / hydrocracking step of a hydrogenation step of the unsaturated compounds. The hydrogenation of the olefinic type unsaturated compounds is in fact a strongly exothermic reaction which may have a negative impact on the hydroisomerization / hydrocracking step and cause, for example, a thermal runaway of the reaction in the case where these unsaturated compounds are not not removed upstream of the hydroisomerization / hydrocracking step, - to implement a simplified process in which the amount of hydrogen introduced into the hydrogenation zone corresponds to a quantity of hydrogen slightly in excess of the amount strictly necessary to carry out the hydrogenation reaction of olefinic type unsaturated compounds so that the process does not require the integration of a recycle compressor and that no cracking is carried out in the hydrogenation zone. This allows the direct sending, preferably by pumping, of all the liquid hydrogenated effluent, without intermediate separation step, in the hydroisomerization / hydrocracking zone, as well as the use of a quantity of hydrogen considerably. scaled down.

to greatly improve the cold properties of paraffins resulting from the Fisher-process

Tropsch and having boiling points corresponding to those of diesel and kerosene fractions, (also called middle distillates) and especially to improve the freezing point of kerosene.

to increase the amount of average distillates available by hydrocracking of the heavier paraffinic compounds present in the outlet effluent of the Fischer-Tropsch unit, and which have boiling points higher than those of kerosene and diesel cuts; for example the fraction 370 ° C +.

Figure 1 shows the embodiment of the method according to the largest invention.

More precisely, FIG. 1 represents a process for producing middle distillates from a paraffinic feedstock produced by Fischer-Tropsch synthesis, comprising the following successive stages: a) separation of at least one so-called light C4- gas fraction, with a final boiling point below 20 ^° C., of the effluent from the Fischer Tropsch synthesis unit so as to to obtain a single liquid fraction C5 +, said heavy, with initial boiling point between 20 and 40 ⁰ C ₁ b) hydrogenation of unsaturated compounds of olefinic type of at least a part of said heavy fraction C5 +, in the presence of hydrogen and a hydrogenation catalyst at a temperature of between 100 ^° C. and 180 ^° C., at a total pressure of between 0.5 and 6 MPa ₁ at an hourly space velocity of between 1 and 10 h -1, and at a pressure of hydrogen flow rate corresponding to a volume ratio hydrogen / hydrocarbons of between 5 and 80 Nl / l / h, c) hydroisomerization / hydrocracking of all the liquid hydrogen effluent from step b), without prior separation step , in the presence of hydrogen and a hydroisomerization / hydrocracking catalyst, d) distillation of the hydrocracked / hydroisomerized effluent.

DETAILED DESCRIPTION OF THE INVENTION Throughout the remainder of the description, we will detail the various steps of the method according to the invention with reference to FIGS. 2 and 3 which represent preferred embodiments of the method according to the invention without limiting them. the scope.

Step (a)

Step a) according to the invention, not shown in FIG. 1, is a step of separation of at least one C4- fraction, called light, with a final boiling point of less than 20 ^° C., preferably less than 20 ^° C. 10 ⁰ C and very preferably, below 0 ⁰ C, the effluent from the Fischer Tropsch synthesis so as to obtain a single C5 + fraction, called heavy, initial boiling point between 20 and 40 ⁰ C and preferably having a boiling point greater than or equal to 30 ⁰ C ₁ constituting at least a part of the feedstock of the hydrogenation stage b) according to the invention.

The effluent from the Fischer-Tropsch synthesis unit is, at the output of the Fischer-Tropsch synthesis unit advantageously divided into two fractions, a light fraction, called cold condensate, (line (1)) and a fraction heavy, called waxes, (pipe (3)).

The two fractions thus defined comprise water, carbon dioxide (CO ₂ ), carbon monoxide (CO) and unreacted hydrogen (H ₂ ). In addition, the light fraction, cold condensate, contains light hydrocarbon compounds C1 to C4, called C4- fraction, in the form of gas. According to a preferred embodiment shown in FIG. 2, the light fraction, called cold condensate (1), and the heavy fraction, called waxes (3), are treated separately in separate fractionation means and then recombined in the pipe ( 5), so as to obtain a single C5 + fraction, so-called heavy, with an initial boiling point of between 20 and 40 ^° C. and preferably having a boiling point greater than or equal to 30 ^° C. The heavy fraction, called waxes, enters a splitting means (4) via the pipe (3). The fractionation means (4) may for example consist of methods well known to those skilled in the art such as rapid expansion (or flash, according to the English terminology), distillation or stripping. Advantageously, a flash or flash tank or a stripper is sufficient to remove most of the water, carbon dioxide (CO ₂ ) and carbon monoxide (CO) through the line (4 ¹ ) of the heavy fraction, called waxes. The light fraction, called cold condensate, enters a fractionation means (2) via the pipe (1). The fractionation means (2) may for example consist of methods well known to those skilled in the art such as a flash or flash tank, distillation or stripping. Advantageously, the fractionation means (2) is a distillation column allowing the elimination of the light and gaseous hydrocarbon compounds C1 to C4, called gas fraction C4-, corresponding to products boiling at a temperature below 20 ^° C., preferably below at 10 ° C and very preferably, below 0 ^° C, by the pipe (2 ¹ ).

The stabilized effluents from the fractionation means (2) and (4) are then recombined in the pipe (5). A stabilized C5 + liquid fraction, corresponding to the products boiling at an initial boiling point of between 20 and 40 ^° C. and preferably having a boiling point greater than or equal to 30 ° C., is thus recovered in the pipe (5) and constitutes the charge of the hydrogenation step b) of the process according to the invention.

According to another preferred embodiment shown in FIG. 3, the light fraction, called cold condensate, leaving the Fischer-Tropsch synthesis unit via line (1) and the heavy fraction, called waxes, coming out of the Fischer-Tropsch synthesis unit via the pipe (3), are recombined in the pipe (18) and treated in the same fractionation means (4). The fractionation means (4) may for example consist of methods well known to those skilled in the art such as flash, distillation or stripping. Advantageously, the fractionation means (4) is a distillation column allowing the elimination of the gas fraction C4-, water, carbon dioxide (CO ₂ ) and carbon monoxide (CO) through the pipe (4 ¹ ).

A stabilized C5 + liquid fraction, corresponding to products boiling at a boiling point of between 20 and 40 ^° C. and preferably having a boiling point greater than or equal to 30 ^° C., is thus recovered at the outlet of the fractionation means (4). ) in the line (5) and constitutes the charge of the hydrogenation step b) of the process according to the invention.

Step b)

Step b) of the process according to the invention is a step of hydrogenation of the olefinic type unsaturated compounds of at least a part and preferably of the whole of the C5 + heavy liquid fraction resulting from step a) of the process according to the invention, in the presence of hydrogen and a hydrogenation catalyst.

Said C5 + liquid heavy fraction is admitted in the presence of hydrogen (line 6) in a hydrogenation zone (7) containing a hydrogenation catalyst which aims to saturate the unsaturated olefinic compounds present in the C5 + heavy liquid fraction. described above.

Preferably, the catalyst used in step (b) according to the invention is a non-crunchy or slightly cracking hydrogenation catalyst comprising at least one metal of group VIIl of the periodic table of the elements and comprising at least one support for refractory oxide base. Preferably, said catalyst comprises at least one metal of the group VH1 chosen from nickel, molybdenum, tungsten, cobalt, ruthenium, indium, palladium and platinum and comprising at least one oxide-based support refractory selected from alumina and silica alumina. Preferably, the Group VIII metal is chosen from nickel, palladium and platinum.

According to a preferred embodiment of step b) of the process according to the invention, the Group VIII metal is chosen from palladium and / or platinum and the content of this metal is advantageously between 0.1% and 5%. % by weight, and preferably between 0.2% and 0.6% by weight relative to the total weight of the catalyst. According to a very preferred embodiment of step b) of the process according to the invention, the Group VIII metal is palladium.

According to another preferred embodiment of step b) of the process according to the invention, the metal of group VIII is nickel and the content of this metal is advantageously between 5% and 25% by weight, preferably between 7%. and 20% by weight based on the total weight of the catalyst.

The catalyst support used in step (b) of the process according to the invention is a refractory oxide-based support, preferably chosen from alumina and silica-alumina.

When the support is an alumina, it has a BET specific surface to limit the polymerization reactions on the surface of the hydrogenation catalyst, said surface being between 5 and 140 m ² / g.

When the support is a silica-alumina, the support contains a percentage of silica of between 5 and 95% by weight, preferably between 10 and 80%, more preferably between 20 and 60% and very preferably between 30 and 50%. , a BET specific surface of between 100 and 550 m ² / g, preferably between 150 and 500 m ² / g, preferably less than 350 m ^z / g and even more preferably less than 250 m ² / g ,

The hydrogenation stage b) of the process according to the invention is preferably carried out in one or more fixed bed reactor (s).

In the hydrogenation zone (7), the feedstock is brought into contact with the hydrogenation catalyst in the presence of hydrogen and at operating temperatures and pressures allowing the hydrogenation of the olefinic unsaturated compounds present in the feedstock. Under these operating conditions, the oxygenated compounds are not converted, the liquid hydrogen effluent from step b) of the process according to the invention therefore does not contain water resulting from the conversion of said oxygenated compounds.

According to the invention, the operating conditions of the hydrogenation stage b) are chosen so that the effluent leaving said hydrogenation zone (7) is in the liquid state: indeed, the amount of hydrogen introduced into the hydrogenation zone (7) corresponds to a quantity of hydrogen in slight excess with respect to the quantity of hydrogen strictly necessary to carry out the hydrogenation reaction of the unsaturated compounds of the type olefin. Thus, no cracking is carried out in the hydrogenation zone (7), and the liquid hydrogenated effluent does not contain hydrocarbon compounds boiling at a temperature below 20 ^° C., preferably below 10 ^° C. and very preferred, less than 0 ⁰ C, corresponding to the gaseous fraction C4-.

The operating conditions of the hydrogenation step b) of the process according to the invention are as follows: the temperature within said hydrogenation zone (7) is between 100 and 180 ^° C. and preferably between 120 and 180 ^° C. and 165 ° C, the total pressure is between 0.5 and 6 MPa, preferably between 1 and 5 MPa and even more preferably between 2 and 5 MPa. The charge flow rate is such that the hourly volume velocity (ratio of the hourly volume flow rate at 15 ° C. of fresh liquid charge to the loaded catalyst volume) is between 1 and 10 h ^-1 , preferably between 1 and 5 h ^-1. and even more preferably between 1 and 4 h ^-1 . The hydrogen which feeds the hydrotreating zone is introduced at a rate such that the hydrogen / hydrocarbon volume ratio is between 5 and 80 Nl / l / h, preferably between 5 and 60, preferably between 10 and 50 Nl / l / h, and even more preferably between 15 and 35 Nl / l / h.

Under these conditions, the olefinic type unsaturated compounds are hydrogenated more than 50%, preferably more than 75% and preferably more than 85%.

The hydrogenation step b) of the process according to the invention is preferably carried out under conditions such as conversion to products having boiling points greater than or equal to 370 ° C in products having lower boiling points. at 370 ⁰ C is zero. Hydrogenated effluent from step b) of the process according to the invention therefore contains no compounds boiling at a temperature below 20 ° C, preferably less than 1O ⁰ C and very preferably less than 0 ⁰ C, corresponding to the gas fraction C4-.

According to a preferred embodiment of step b) of the process according to the invention, use is made of a guard bed (not shown in the figures) containing at least one guard bed catalyst upstream of the hydrogenation zone ( 7) to reduce the content of solid mineral particles and possibly reduce the content of harmful metal compounds for hydrogenation catalysts. The guard bed may advantageously be either integrated in the hydrogenation zone (7) upstream of the hydrogenation catalyst bed or be placed in a separate zone upstream of the hydrogenation zone (7). Indeed, the treated fractions may optionally contain solid particles such as inorganic solids. They may optionally contain metals contained in hydrocarbon structures such as more or less soluble organometallic compounds. By the term fines is meant fines resulting from a physical or chemical attrition of the catalyst. They can be micron or sub-micron. These mineral particles then contain the active components of these catalysts without the following list being limiting: alumina, silica, titanium, zirconia, cobalt oxide, iron oxide, tungsten, rhuthenium oxide, etc. These solid minerals may be present under the calcined mixed oxide form: for example, alumina-cobalt, alumina-iron, alumina-silica, alumina-zirconia, alumina-titanium, alumina-silica-cobalt, alumina-zirconia-cobalt, ....

They may also contain metals within hydrocarbon structures, which may optionally contain oxygen or more or less soluble organometallic compounds. More particularly, these compounds may be based on silicon. It may be for example anti-foaming agents used in the synthesis process. On the other hand, the catalyst fines described above may have a higher silica content than the catalyst formulation resulting from the intimate interaction between the catalyst fines and anti-foaming agents described above.

The guard bed catalysts used may advantageously be in the form of spheres or extrudates. It is however advantageous that the catalyst is in the form of extrudates with a diameter of between 0.5 and 5 mm and more particularly between 0.7 and 2.5 mm. The shapes are cylindrical (which can be hollow or not), cylindrical twisted, multilobed (2, 3, 4 or 5 lobes for example), rings. The cylindrical shape is preferably used, but any other shape may be used.

In order to overcome the presence of contaminants and / or poisons in the feed, the bed bed catalysts may, in another preferred embodiment, have more specific geometric shapes in order to increase their void fraction. The void fraction of these catalysts is between 0.2 and 0.75. Their outer diameter can vary between 1 and 35 mm. Among the particular forms possible without this list being exhaustive: hollow cylinders, hollow rings, Raschig rings, serrated hollow cylinders, crenellated hollow cylinders, pentaring carts, multi-hole cylinders ...

Preferably, said guard bed catalysts used are not impregnated with an active phase. Guard beds can be marketed by Norton-Saint-Gobain, for example example MacroTrap® guard beds. Guard beds can be marketed by Axens in the ACT family: ACT077, ACT935, ACT961 or HMC841, HMC845, HMC941 or HMC945. It may be particularly advantageous to superpose these catalysts in at least two different beds of varying heights. Catalysts having the highest void content are preferably used in the first catalytic bed or first catalytic reactor inlet. It may also be advantageous to use at least two different reactors for these catalysts. These guard bed catalysts used may advantageously have macroporosity. In a preferred embodiment, the macroporous volume for a mean diameter at 50 nm is greater than 0.1 cm ³ / g and a total volume greater than 0.60 cm ³ / g. In another embodiment, the mercury volume for a pore diameter greater than 1 micron is greater than 0.5 cm ³ / g and the mercury volume for a pore diameter greater than 10 microns is greater than 0.25 cm ³ /boy Wut. These two embodiments can advantageously be associated in a mixed bed or a combined bed. The preferred guard beds according to the invention are HMC and IΑCT961. After passing over the guard bed, the solids content is advantageously less than 20 ppm, preferably less than 10 ppm and even more preferably less than 5 ppm. The soluble silicon content is advantageously less than 5 ppm, preferably less than 2 ppm and even more preferably less than 1 ppm.

At the end of step b) of the process according to the invention, all the liquid hydrogenated effluent is directly sent to a hydrocracking / hydroisomerization zone (10).

Step c) According to step c) of the process according to the invention, all the liquid hydrogenated effluent from step b) of the process according to the invention is directly sent, without prior separation step, in the zone hydroisomerization / hydrocracking system (10) containing the hydroisomerization / hydrocracking catalyst and preferably at the same time as a hydrogen stream (line 9).

The operating conditions in which the hydroisomerization / hydrocracking step (c) of the process according to the invention is carried out are preferably as follows:

The pressure is generally maintained between 0.2 and 15 MPa and preferably between 0.5 and 10 MPa and advantageously from 1 to 9 MPa, the space velocity is generally comprised between 0.1 h -1 and 10 h -1 and preferably between 0.2 and 7 h -1 is advantageously between 0.5 and 5.0 h -1. The hydrogen content is generally between 100 and 2000 normal liters of hydrogen per liter of filler and per hour and preferably between 150 and 1500 liters of hydrogen per liter of filler.

The temperature used in this step is generally between 200 and 450 ^° C. and preferably from 250 ^° C. to 450 ° C., advantageously from 300 to 450 ° C., and still more advantageously greater than 32 ^° C. or for example between 320-420 ^° C. vs.

Step c) of hydroisomerization and hydrocracking of the process according to the invention is advantageously carried out under conditions such that the pass conversion into products with boiling points greater than or equal to 37O ⁰ C into products having points. boiling point below 370 ^° C. is greater than 80% by weight, and even more preferably at least 85%, preferably greater than 88%, so as to obtain middle distillates (gas oil and kerosene) having sufficiently good cold (pour point, freezing point) to meet the specifications in force for this type of fuel.

Hydroisomerization / hydrocracking catalysts

The majority of catalysts currently used in hydroisomerization / hydrocracking are of the bifunctional type associating an acid function with a hydrogenating function. The acid function is generally provided by supports with large surface areas (150 to 800 m2.g-1 generally) having a surface acidity, such as halogenated aluminas (chlorinated or fluorinated in particular), phosphorus aluminas, combinations of oxides of boron and aluminum, silica aluminas. The hydrogenating function is generally provided either by one or more metals of group VIII of the periodic table of the elements, such as iron, cobalt, nickel, ruthenium, rhodium, palladium, osmium, iridium and platinum, or by an association of at least a group VI metal such as chromium, molybdenum and tungsten and at least one group VIIl metal.

In the case of bi-functional catalysts, the equilibrium between the two acid and hydrogenating functions is the fundamental parameter which governs the activity and the selectivity of the catalyst. A weak acidic function and a strong hydrogenating function give catalysts which are not very active and selective towards isomerization whereas a strong acid function and a low hydrogenating function give very active and cracking-selective catalysts. A third possibility is to use a strong acid function and a strong hydrogenating function to obtain a very active catalyst but also very selective towards isomerization. It is therefore possible, judiciously choosing each of the functions to adjust the activity / selectivity of the catalyst.

Advantageously, the hydroisomerization-hydrocracking catalysts are bifunctional catalysts comprising an amorphous acid support (preferably a silica-alumina) and a hydro-dehydrogenating metal function preferably provided by at least one noble metal. The support is said to be amorphous, that is to say devoid of molecular sieves, and in particular of zeolite, as well as the catalyst. The amorphous acidic support is advantageously a silica-alumina but other supports are usable. When it is a silica-alumina, the catalyst preferably does not contain added halogen, other than that which could be introduced for the impregnation of the noble metal, for example.

More generally, and preferably, the catalyst does not contain added halogen, for example fluorine. In general, and preferably, the support has not been impregnated with a silicon compound.

A preferred hydroisomerization / hydrocracking catalyst used in step c) of the process according to the invention comprises up to 3% by weight of metal of at least one hydro-dehydrogenating element chosen from noble metals of group VIII, preferably deposited on the support, and very preferably, the noble metal of group VIII being platinum and a support comprising (or preferably consisting of) at least one silica-alumina, said silica-alumina having the following characteristics: a weight content silica SiO ₂ between 5 and 95%, preferably between 10 and 80%, more preferably between 20 and 60% and even more preferably between 30 and 50% by weight. an Na content of less than 300 ppm by weight and preferably less than 200 ppm by weight, a total pore volume of between 0.45 and 1.2 ml / g measured by mercury porosimetry, the porosity of said silica-alumina being the next :

The volume of the mesopores whose diameter is between 40 ° and 150 ° and whose mean diameter varies between 80 and 140 Å and preferably between 80 and 120 Å, represents between 20 and 80% of the total pore volume measured by porosimetry. Mercury, ii / The volume of the macropores, whose diameter is greater than 500 Å, and preferably between 1000 A and 10000 A represents between 20 and 80% of the total measured pore volume by mercury porosimetry,

a specific surface area of between 100 and 550 m ² / g, preferably between 150 and 500 m ² / g, preferably less than 350 m ² / g and even more preferably less than 250 m ² / g .

A second preferred hydroisomerization / hydrocracking catalyst used in stage c) of the process according to the invention comprises up to 3% by weight of metal of at least one hydro-dehydrogenating element chosen from the noble metals of group VIII of the periodic classification and preferably the noble metal of group VIII being platinum, from 0.01 to 5.5% by weight of oxide of a doping element selected from phosphorus, boron and silicon and a non-zeolitic support to silica-alumina base containing an amount greater than 15% by weight and less than or equal to 95% by weight of silica (SiO ₂ ), said silica-alumina having the following characteristics: a mean pore diameter, measured by mercury porosimetry, including between 20 and 140 Å, a total pore volume, measured by mercury porosimetry, of between 0.1 ml / g and 0.5 ml / g, a total pore volume, measured by nitrogen porosimetry, of between 0.1 ml / g and 0.6 ml / g,

a BET specific surface area of between 100 and 550 m ² / g, a pore volume, measured by mercury porosimetry, included in pores with a diameter greater than 140 Å of less than 0.1 ml / g, a pore volume, measured by mercury porosimetry, included in pores with a diameter greater than 160 Å less than 0.1 ml / g, a pore volume, measured by mercury porosimetry, contained in pores with diameters greater than 200 Λ, less than 0.1 ml g, a pore volume, measured by mercury porosimetry, included in pores with diameters greater than 500 Å less than 0.1 ml / g. an X-ray diffraction pattern which contains at least the principal characteristic lines of at least one of the transition aluminas included in the group consisting of alpha, rho, chi, eta, gamma, kappa, theta and delta alumina. a packed packing density of the catalysts greater than 0.55 g / cm ³ . Advantageously, the characteristics associated with the corresponding catalyst are identical to those of the silica alumina described above. The two stages b) and c) of the process according to the invention, hydrogenation and hydroisomerization-hydrocracking, can advantageously be carried out on the two types of catalysts in two or more different reactors, and / or in the same reactor.

A third preferred hydroisomerization / hydrocracking catalyst used in step c) of the process according to the invention comprises at least one hydro-dehydrogenating element chosen from the non-noble metals of group VIII and the metals of group VIB of the periodic table, preferably between 2.5 and 5% by weight of Group VIII non-noble elemental oxide and between 20 and 35% by weight of Group VIB element oxide with respect to the weight of the final catalyst and, preferably, the non-noble metal of group VIIl is nickel and the metal of group VIB is tungsten, optionally from 0.01 to 5.5% by weight of oxide of a doping element chosen from phosphorus, boron and silicon and preferably, from 0.01 to 2.5% by weight of oxide of a doping element and a non-zeolitic support based on silica-alumina containing an amount greater than 15% by weight and less than or equal to 95% by weight of silica (SiO _2), preferably an amount greater than 15% by weight and less than or equal to 50% by weight of silica, said silica-alumina having the following characteristics: a mean pore diameter, measured by mercury porosimetry, of between 20 and 140 Λ, - a total pore volume, measured by mercury porosimetry, between 0.1 ml / g and 0.5 ml / g, a total pore volume, measured by nitrogen porosimetry, of between 0.1 ml / g and 0.6 ml / g, a specific surface area BET between 100 and 550 m ² / g, - a pore volume, measured by mercury porosimetry, included in pores with a diameter greater than 140 Λ less than 0.1 ml / g, a pore volume, measured by mercury porosimetry , comprised in pores with a diameter greater than 160 Å of less than 0.1 ml / g, a pore volume, measured by mercury porosimetry, contained in pores with diameters greater than 200 Å, of less than 0.1 ml / g, a porous volume, measured by mercury porosimetry, included in the pores of diameter re greater than 500 Å less than 0.1 ml / g. an X-ray diffraction pattern which contains at least the principal characteristic lines of at least one of the transition aluminas included in the group consisting of alpha, rho, chi, eta, gamma, kappa, theta and delta alumina. a packed packing density of the catalysts greater than 0.55 g / cm ³ . Advantageously, the characteristics associated with the corresponding catalyst are identical to those of the silica alumina described above.

When the third preferred hydroisomerization / hydrocracking catalyst is used in step c) of the process according to the invention, said catalyst is sulphurized.

According to a first preferred embodiment of the process according to the invention, a palladium-containing catalyst is used in the hydrogenation step b) and in the hydroisomerization / hydrocracking step c), a platinum-containing catalyst.

According to a second preferred embodiment of the process according to the invention, a palladium-containing catalyst is used in the hydrogenation step b) and in the hydroisomerization / hydrocracking step c), a sulphurized catalyst containing at least one hydro-dehydrogenating element selected from Group VIII non-noble metals and Group VIB metals.

According to a third preferred embodiment of the process according to the invention, in the hydrogenation step b), a catalyst containing at least one non-noble group VIII hydroxide dehydrogenating element and in the hydroisomerisation stage c) is used. hydrocracking, a sulphurized catalyst containing at least one hydro-dehydrogenating element chosen from Group VIII non-noble metals and Group VIB metals.

Step (d) The effluent (so-called hydrocracked / hydroisomerized fraction) leaving the hydroisomerization / hydrocracking zone (10), resulting from step (c) of the process according to the invention, is sent, in accordance with the step d) of the process according to the invention, in a distillation train (11), which incorporates an atmospheric distillation and optionally a vacuum distillation, which aims to separate the conversion products of boiling point below 34O ⁰ C and preferably less than 370 ⁰ C and including in particular those formed during step (c) in the hydroisomerization / hydrocracking reactor (10), and to separate the residual fraction whose initial boiling point is generally higher at least 340 ^° C. and preferably greater than or equal to at least 370 ^° C. Among the conversion and hydroisomerized products, at least one gasoline (or naphtha) fraction is separated from the light C1-C4 gases (line 14) (conduct 13), and at least one middle distillate fraction kerosene (line 14) and diesel (line 15). Preferably, the residual fraction, whose initial boiling point is generally greater than at least 340 ^° C. and preferably greater than or equal to at least 37 ^° C. C is recycled (line 16) in step c) of the process according to the invention at the head of the zone (10) for hydroisomerization and hydrocracking. According to another embodiment of step d) of the process according to the invention, said residual fraction can provide excellent bases for the oils.

It may also be advantageous to recycle (line 17) at least partly and preferably completely, in step (c) (zone 10) at least one of the kerosene and diesel fuel cuts thus obtained. The gas oil and kerosene cuts are preferably recovered separately or mixed, but the cutting points are adjusted by the operator according to his needs. It has been found that it is advantageous to recycle a portion of the kerosene to improve its cold properties.

The products obtained

The gas oil (s) obtained has a pour point of at most 0 ^° C., generally below -10 ^° C. and often below -15 ° C. The cetane number is greater than 60, generally greater than 65, often greater than 70.

The kerosene (s) obtained has a freezing point of not more than -35 ° C, generally less than -40 ^° C. The smoke point is greater than 25 mm, generally greater than 30 mm. In this process, the production of gasoline (not sought) is as low as possible.

The yield of gasoline is always less than 50% by weight, preferably less than 40% by weight, advantageously less than 30% by weight or 20% by weight or even 15% by weight.

Claims

A process for producing middle distillates from a paraffinic feedstock produced by Fischer-Tropsch synthesis, comprising the following steps:

e) separation of at least one C4- gas fraction, called light, with a final boiling point of less than 20 ^° C., from the effluent from the Fischer Tropsch synthesis unit so as to obtain a single C5 + liquid fraction , said heavy, initial boiling point between 20 and 40 ⁰ C, f) hydrogenation of olefinic unsaturated compounds of at least a portion of said C5 + heavy fraction, in the presence of hydrogen and a catalyst d hydrogenation at a temperature of between 100 ° C. and 180 ^° C., at a total pressure of between 0.5 and 6 MPa, at an hourly space velocity of between 1 and 10 h -1, and at a hydrogen flow rate corresponding to a hydrogen / hydrocarbon volume ratio of between 5 and 80 Nl / l / h, g) hydroisomerization / hydrocracking of all the liquid hydrogenated effluent from step b), without prior separation step, in the presence of hydrogen and a hydroisomerization / hydrocracking catalyst, h) distillation of the hydrocracked / hydroisomerized effluent.

2. Process according to claim 1, in which said effluent from the Fischer Tropsch synthesis unit is, at the outlet of the Fischer-Tropsch synthesis unit divided into two fractions, a light fraction, called cold condensate, and a fraction heavy, called waxes.

3. Method according to claim 2 wherein the light fraction, called cold condensate, and the heavy fraction, called waxes, are treated separately in separate fractionation means and then recombined, so as to obtain a single C5 + fraction, called heavy, with an initial boiling point of between 20 and 40 ^° C.

4. The method of claim 2 wherein the light fraction, called cold condensate and the heavy fraction, called waxes, are recombined and processed in the same fractionation means.

5. Method according to one of claims 1 to 4 wherein said hydrogenation catalyst comprises at least one metal of group VIII of the periodic table of elements and comprising at least one support based on refractory oxide.

6. Method according to one of claims 1 to 5 wherein the group VIII metal is palladium.

7. Method according to one of claims 1 to 6 wherein the hydrogenation of olefinic type unsaturated compounds of at least a portion of said heavy fraction, operates at a volume ratio hydrogen / hydrocarbons of between 10 and 50 Nl / l. / h.

8. The method of claim 8 wherein the hydrogenation of olefinic unsaturated compounds of at least a portion of said heavy fraction, operates at a volume ratio hydrogen / hydrocarbons of between 15 and 35 Nl / l / h.

9. Method according to one of claims 1 to 8 wherein using a guard bed containing at least one guard bed catalyst upstream of the hydrogenation zone, said guard bed being either integrated in the zone d Hydrogenation upstream of the hydrogenation catalyst bed is placed in a separate zone upstream of the hydrogenation zone.

10. Method according to one of claims 1 to 9 wherein said step (c) hydroisomerization / hydrocracking is carried out at a pressure between 0.2 and 15 MPa, at a space velocity is between 0.1 h -1 and 10 h -1, at a hydrogen content between 100 and 2000 normal liters of hydrogen per liter of filler and per hour and at a temperature between 200 and 45O ⁰ C.

11. Method according to one of claims 1 to 10 wherein said hydroisomerization / hydrocracking catalyst comprises up to 3% by weight of metal of at least one hydro-dehydrogenating element selected from noble metals of group VIII and a carrier. comprising (or preferably consisting of) at least one silica-alumina, said silica-alumina having the following characteristics: a weight content of silica SiO 2 of between 5 and 95% an Na content of less than 300 ppm by weight, - a porous volume total between 0.45 and 1.2 ml / g measured by mercury porosimetry, the porosity of said silica-alumina being as follows:

i / The volume of mesopores whose diameter is between 4θA and 150Â °, and whose mean diameter varies between 80 and 140 Λ, represents between 20 and 80% of the total pore volume measured by mercury porosimetry, ii / The volume of macropores , whose diameter is greater than 500 Å, and preferably between 1000 Å and 10000 Å, represents between 20 and 80% of the total pore volume measured by mercury porosimetry,

A specific surface area of between 100 and 550 m 2 / g.

12. Method according to one of claims 1 to 10 wherein said hydroisomerization / hydrocracking catalyst comprises up to 3% by weight of metal of at least one hydro-dehydrogenating element selected from the noble metals of Group VIlI of the periodic classification, from 0.01 to 5.5% weight of oxide of a doping element chosen from phosphorus, boron and silicon and a non-zeolitic support based on silica-alumina containing an amount greater than 15% by weight and less than or equal to 95% by weight of silica (SiO2), said silica-alumina having the following characteristics: a mean pore diameter, measured by mercury porosimetry, of between 20 and 140 A, a total pore volume, measured by porosimetry at mercury, between 0.1 ml / g and 0.5 ml / g, a total pore volume, measured by nitrogen porosimetry, of between 0.1 ml / g and 0.6 ml / g,

a BET specific surface area of between 100 and 550 m ² / g, a pore volume, measured by mercury porosimetry, included in pores with a diameter greater than 140 Å of less than 0.1 ml / g, a porous volume, measured by porosimetry to mercury, included in pores with a diameter greater than 160 Å less than 0.1 ml / g, a pore volume, measured by mercury porosimetry, contained in pores with diameters greater than 200 Å, less than 0.1 ml / g, a pore volume, measured by mercury porosimetry, included in pores with diameters greater than 500 Å less than 0.1 ml / g. an X-ray diffraction diagram which contains at least the principal characteristic lines of at least one of the transition aluminas included in the group composed of alpha, rho, chi, eta, gamma, kappa, theta and delta alumina. a packed packing density of the catalysts greater than 0.55 g / cm ³ .

13. Process according to one of claims 1 to 10 wherein said hydroisomerization / hydrocracking catalyst comprises between 2.5 and 5% by weight of Group VIII elemental oxide and between 20 and 35% by weight of oxide. of group VIB element relative to the weight of the final catalyst, optionally from 0.01 to 5.5% by weight of oxide of a doping element chosen from phosphorus, boron and silicon and a non-zeolitic support based on silica - alumina containing an amount greater than 15% by weight and less than or equal to 95% by weight of silica (SiO2), said silica - alumina having the following characteristics: a mean pore diameter, measured by mercury porosimetry, of between 20 and 140 Λ, a total pore volume, measured by mercury porosimetry, of between 0.1 ml / g and 0.5 ml / g, a total pore volume, measured by nitrogen porosimetry, of between 0.1 ml / g and 0 , 6 ml / g, a BET specific surface area of between 100 and 550 m ² / g, a volume in When measured by mercury porosimetry, in pores with a diameter greater than 140 Å less than 0.1 ml / g, a porous volume, measured by mercury porosimetry, included in pores with a diameter greater than 160 A less than 0 , 1 ml / g, a pore volume, measured by mercury porosimetry, included in pores with a diameter greater than 200 Λ, less than 0.1 ml / g, - a pore volume, measured by mercury porosimetry, included in the Pore diameter greater than 500 Å less than 0.1 ml / g. an X-ray diffraction pattern which contains at least the principal characteristic lines of at least one of the transition aluminas included in the group consisting of alpha, rho, chi, eta, gamma, kappa, theta and delta alumina. a packed packing density of the catalysts greater than 0.55 g / cm ³ .

The process of claim 13 wherein said catalyst is sulfided.

15. Method according to one of claims 1 to 14 wherein a residual fraction whose boiling point is greater than at least 340 ° C is recycled in step c).

16. Method according to one of claims 1 to 15 wherein at least one kerosene and diesel fuel cuts from step d) is recycled at least in part in step c).