FR2969650A1 - HYDROCARBONATE LOADING CONVERSION METHOD COMPRISING SCHIST HYDROCONVERSION OIL IN BOILING BED, ATMOSPHERIC DISTILLATION FRACTIONATION AND LIQUID / LIQUID EXTRACTION OF HEAVY FRACTION - Google Patents

Abstract

A hydrocarbon feed conversion process comprising a shale oil comprising a boiling bed hydroconversion stage, a fraction fraction fractionated by atmospheric distillation in a light fraction, a naphtha fraction, a gas oil fraction and a heavier fraction than the diesel fraction, a liquid / liquid extraction step of the heavier fraction than the diesel fraction, and a hydrotreatment dedicated to each of the naphtha and diesel fractions. The process aims to maximize the yield of fuel bases.

Description

HYDROCARBONATE LOADING CONVERSION METHOD COMPRISING SCHIST HYDROCONVERSION OIL IN BOILING BED, ATMOSPHERIC DISTILLATION FRACTIONATION AND LIQUID / LIQUID EXTRACTION OF HEAVY FRACTION

The invention relates to a process for converting hydrocarbon feedstocks comprising a shale oil into lighter products, recoverable as fuels and / or raw materials for petrochemicals. More particularly, the invention relates to a hydrocarbon feedstock conversion process comprising a shale oil comprising a step of hydroconversion of the boiling bed feedstock, followed by a step of fractionation by atmospheric distillation into light fraction, naphtha, gas oil and in a heavier fraction than the gas oil fraction, a liquid / liquid exaction step of the heavier fraction than the gas oil fraction, and a hydrotreatment dedicated to each of the naphtha and diesel fractions. This process makes it possible to convert shale oils into fuel base of very good quality and aims in particular at an excellent yield. Given the high volatility of the price of a barrel and a decline in discoveries of conventional oil fields, oil companies are turning to unconventional sources. Alongside oil sands and deep offshore, oil shale, although relatively unknown, is becoming increasingly popular. Oil shales are sedimentary rocks containing an insoluble organic substance called kerogen. By heat treatment in situ or ex-situ ("retorting" in the English terminology) in the absence of air at temperatures between 400 and 500 ° C, these schists release an oil, shale oil , whose general appearance is that of crude oil.

Although of different composition of crude oil, shale oils can be a substitute for the latter and also a source of chemical intermediates. Shale oils can not be a direct substitute for crude oil applications. Indeed, although these oils are similar in some respects to oil (for example by a similar H / C ratio), they are distinguished by their chemical nature and a much higher content of metallic and / or nonmetallic impurities, which makes the conversion of this unconventional resource much more complex than that of oil. Shale oils have higher levels of oxygen and nitrogen than oil. They may also contain higher concentrations of olefins, sulfur or metal compounds (including arsenic). Shale oils obtained by pyrolysis of kerogen contain many olefinic compounds resulting from cracking, which implies an additional demand for hydrogen in refining. Thus, the bromine index, which makes it possible to calculate the concentration by weight of olefinic hydrocarbons (by addition of bromine to the ethylenic double bond), is generally greater than 30 g / 100 g of filler for shale oils, whereas it is between 1 and 5g / 100g of charge for the oil residues. The olefinic compounds resulting from cracking consist essentially of mono olefins and diolefins. The unsaturations present in the olefins are potential source of instability by polymerization and / or oxidation.

The oxygen content is generally higher than in heavy crudes and can reach up to 80% by weight of the filler. Oxygen compounds are often phenols or carboxylic acids. As a result, shale oils may have marked acidity.

The sulfur content varies between 0.1 and 6.50 / 0 by weight, requiring severe desulfurization treatments in order to reach the specifications of the fuel bases. The sulfur compounds are in the form of thiophenes, sulfides or disulfides. In addition, the distribution profile of sulfur in a shale oil may be different from that obtained in conventional oil. The most distinctive feature of shale oils, however, is their high nitrogen content, making them unsuitable as conventional refinery feedstock. The oil generally contains about 0.20% by weight of nitrogen whereas crude shale oils generally contain from about 1 to about 30% by weight or more of nitrogen. In addition, the nitrogen compounds present in the oil are generally concentrated in the higher boiling ranges while the nitrogen of the compounds present in the crude shale oils is generally distributed throughout all the boiling ranges of the material. Nitrogen compounds in petroleum are predominantly non-basic compounds, whereas typically about half of the nitrogen compounds present in crude shale oils are basic in nature. These basic nitrogen compounds are particularly undesirable in the refining feeds because these compounds often act as catalyst poisons. In addition, product stability is a problem that is common to many products derived from shale oil. Such instability, including photosensitivity, appears to result essentially from the presence of nitrogen compounds. Therefore, crude shale oils must generally be subjected to a severe refining treatment (high total pressure) in order to obtain a synthetic crude oil or fuel-based products meeting the specifications in force. Similarly, it is known that shale oils may contain many trace metal compounds, generally present as organometallic complexes. Among the metal compounds, mention may be made of conventional contaminants such as nickel, vanadium, calcium, sodium, lead or iron, but also metal arsenic compounds. In fact, shale oils can contain an amount of arsenic greater than 20 ppm while the amount of arsenic in crude oil is generally in the area of ppb (parts per billion). All these metal compounds are poisons of catalysts. In particular, they irreversibly poison the hydrotreatment and hydrogenation catalysts by settling progressively on the active surface. Conventional metal compounds and some of the arsenic are found mainly in heavy cuts and are removed by settling on the catalyst. On the other hand, when products containing arsenic are able to generate volatile compounds, these can be found partly in lighter cuts and can, therefore, poison the catalysts of the subsequent transformation processes, at the same time. refining or petrochemicals. In addition, shale oils generally contain sandy sediments from oil shale deposits from which shale oils are extracted. These sandy sediments can create clogging problems, especially in fixed bed reactors. Finally, shale oils contain waxes which give them a pour point higher than the ambient temperature, preventing their transport in oil pipelines. Due to considerable resources, and their evaluation as a promising source of oil, there is a real need to convert shale oils into lighter products that can be used as fuels and / or raw materials for petrochemicals. Methods for converting shale oils are known. Conventionally, the conversion is done either by coking, by hydroviscoreduction (thermal cracking in the presence of hydrogen) or by hydroconversion (catalytic hydrogenation). Liquid / liquid extraction processes are also known. Thus, patent US4483763 describes a shale oil conversion process to reduce their nitrogen content. This process comprises a partial hydrogenation step followed by a liquid / liquid extraction step with a mixture of an organic polar solvent, an acid and water. The extraction is done either on a cut distillates means (400-680 ° F = 204-360 ° C), or on all the effluent obtained by hydrogenation. US Pat. No. 5,059,303 describes a shale oil conversion process comprising a boiling bed or fixed bed hydroconversion stage, an optional fractionation step, a liquid / liquid extraction step of a liquid fraction or of all of the liquid effluent before a solvent for extracting the condensed aromatics. The raffinate obtained after evaporation of the solvent is then subjected to fractionation in order to obtain a middle distillate fraction containing up to 1000 ppm of nitrogen and a heavier fraction containing from 500 to 3000 ppm of nitrogen. US Pat. No. 5,059,303 also describes a variant of the process comprising a boiling bed hydroconversion stage, a gas / liquid separation step without pressure reduction, a liquid / liquid phase liquid / liquid extraction step and a hydrotreating stage. the gas phase.

OBJECT OF THE INVENTION The specificity of shale oils being to have a certain number of metallic and / or non-metallic impurities makes the conversion of this unconventional resource much more complex than that of oil. The challenge for the industrial development of shale oil conversion processes is therefore the need to develop processes adapted to the load, making it possible to maximize the yield of good quality fuel bases. The conventional refining treatments known from petroleum must therefore be adapted to the specific composition of shale oils. The present invention aims to improve the known processes for converting hydrocarbon feedstocks comprising a shale oil, in particular by increasing the yield of fuel bases for a combination of steps having a specific sequence, and a treatment adapted to each fraction derived from the oils of shale. Likewise, the object of the present invention is to obtain products of good quality having in particular a low content of sulfur, nitrogen and arsenic, preferably respecting the specifications. Another objective is to propose a simple process, that is to say with the least necessary steps, while remaining effective, to limit investment costs. In its broadest form the present invention is defined as a hydrocarbon feed conversion process comprising at least one shale oil having a nitrogen content of at least 0.10 / 0, often at least 10/0 and very often at least 20/0 by weight, characterized in that it comprises the following steps: a) the feed is treated in a hydroconversion section in the presence of hydrogen, said section comprising at least one operating bubbling bed reactor with an upward flow of liquid and gas and containing at least one hydrotreatment catalyst supported, b) the effluent obtained in step a) is sent, at least in part, and often entirely, in a fractionation zone to from which a gaseous fraction, a naphtha fraction, a gas oil fraction and a heavier fraction than the gas oil fraction are recovered by atmospheric distillation, c) the naphtha fraction is treated, at least in part, and often in whole in a hydrotreating section in the presence of hydrogen, said section comprising at least one fixed bed reactor containing at least one hydrotreatment catalyst, d) said diesel fuel fraction is treated, at least in part, and often entirely, in another hydrotreating section in the presence of hydrogen, said section comprising at least one fixed bed reactor containing at least one hydrotreatment catalyst, e) the heavier fraction than the gas oil fraction is subjected to a liquid / liquid extraction to obtain a raffinate and an extract. Usually, the treatment section of step a) comprises from one to three, preferably two reactors in series, and the processing section of steps c) and d) also comprises from one to three reactors in series. The research work carried out by the applicants on the shale oil conversion has led to the discovery that, surprisingly, an improvement of the existing processes in terms of fuel efficiency and purity of the products is possible by a combination of different steps, chained in a specific manner, and a subsequent processing section for each obtained fraction of the process.

The first step includes boiling bed hydroconversion. Bubble bed technology allows, compared to fixed bed technology, to treat highly contaminated loads of metals, heteroatoms and sediments, such as shale oils, while having conversion rates generally greater than 500/0. In fact, in this first step, shale oil is transformed into molecules that make it possible to generate future fuel bases. Most of the metal compounds, sediments and heterocyclic compounds are removed. The effluent exiting the bubbling bed thus contains the most refractory nitrogen and sulfur compounds, and possibly volatile arsenic compounds in lighter components.

The effluent obtained in the hydroconversion stage is then fractionated by atmospheric distillation, making it possible to obtain different fractions for which a treatment specific to each fraction is carried out subsequently. The key step of the process consists in carrying out a fractionation by atmospheric distillation before the liquid / liquid extraction step, in order to separately maximize the lighter fractions (naphtha, gas oil), which subsequently requires a moderate hydrotreatment treatment adapted to each fraction, and to minimize the heavier fraction than the diesel fraction, requiring a more severe treatment by liquid / liquid extraction. Thus, the atmospheric distillation makes it possible, in a single step, to obtain the various desired fractions (naphtha, gas oil), thus facilitating a downstream hydrotreatment adapted to each fraction and, consequently, the direct obtaining of naphtha or diesel fuel base products. respecting the different specifications. Fractionation after hydrotreatment is therefore not necessary. Due to the strong reduction of contaminants in the bubbling bed, the light fractions (naphtha and gas oil) contain less contaminants and can therefore be processed in a fixed bed section generally having improved hydrogenation kinetics compared to the bubbling bed. Similarly, the operating conditions may be milder due to the limited contaminant content. Providing a treatment for each fraction makes it possible to have a better operability according to the desired products. Depending on the operating conditions chosen (more or less severe), it is possible to obtain either a fraction that can be sent to a fuel pool, or a finished product that meets the specifications (sulfur content, smoke point, cetane, aromatic content, etc.). ) in force. The fixed bed hydrotreatment sections preferably comprise, upstream of the hydrotreatment catalytic beds, guard beds specific for the arsenic and silicon compounds optionally contained in the naphtha and / or diesel fractions. The arsenic compounds escaping the ebullating bed (because they are generally relatively volatile) are fixed in the guard beds, avoiding poisoning the downstream catalysts, and making it possible to obtain fuel bases that are heavily depleted of arsenic. The atmospheric distillation also makes it possible to concentrate the most refractory nitrogen compounds in the heavier fraction than the diesel fraction, thus limiting the quantity to be treated by liquid / liquid extraction. The equipment and the amount of solvent required in the liquid / liquid extraction step are thus minimized. The heavier fraction than the diesel fraction resulting from the fractionation step is subjected to a liquid / liquid extraction using a polar solvent. The solvent used is a solvent for extracting aromatic compounds preferentially. Since the remaining refractory nitrogen is usually present in the aromatic compounds, the liquid / liquid extraction step thus makes it possible to reduce nitrogen-containing aromatic compounds refractory to hydrodenitrogenation (denitrogenation by catalytic hydrogenation). It is important to underline that the liquid / liquid extraction is, contrary to the state of the art, only on the heavy fraction, in order to avoid losses of yield of fuel bases during the recovery of the solvent after extraction. . The products which are to be extracted from the heavy fraction preferably have a boiling point higher than the boiling point of the solvent in order to avoid a loss of yield during the separation of the solvent from the raffinate after the extraction. Indeed, during the separation of the solvent from the raffinate, any compound having a boiling point below the boiling point of the solvent will inevitably leave with the solvent and thus lower the amount of raffinate obtained (and thus the yield of fuel bases) . For example, in the case of furfural as an extraction solvent, having a boiling point of 162 ° C, the compounds C10-, compounds representative of the gasoline / naphtha fraction, will be lost. By treating only the heavy fraction containing compounds having boiling points above the boiling point of the extraction solvent, there is no loss of these C10 compounds. In addition, it avoids the contamination of the solvent with the compounds CIO-as well as possible stages of treatment of the solvent for recycling. Solvent recovery is therefore more efficient and economical.

Another advantage of the process is the fact that the raffinate from the liquid / liquid extraction step e), after evaporation of the solvent, is preferably sent to a catalytic cracking section [step f)] in which it is treated in conditions for producing a gas fraction, a gasoline fraction, a gas oil fraction and a residual heavy fraction called "slurry" according to the English terminology. This variant makes it possible to maximize the yield of fuel bases. Another advantage is the fact that the extract from the liquid / liquid extraction can be at least partly recycled in the hydroconversion stage a). Recycling makes it possible to increase the yield of fuel bases. Detailed description The hydrocarbon feedstock The hydrocarbon feedstock comprises at least one shale oil. The term "shale oil" is used here in its broadest sense and is intended to include any shale oil or a shale oil fraction, which contains nitrogenous impurities. This includes crude shale oil, whether obtained by pyrolysis, solvent extraction or other means, shale oil that has been filtered to remove solids, or that has been treated by a or more solvents, chemicals, or other treatments and that contains nitrogenous impurities. The term "shale oil" also includes shale oil fractions obtained by distillation or other fractionation technique. The shale oils used in the present invention generally have a Conradson carbon content of at least 0.10% by weight and generally at least 50% by weight, an asphaltene content (IP143 / C7 standard). at least 10%, often at least 20% by weight. Their sulfur content is generally at least 0.10 / 0, often at least 10/0 and very often at least 20/0, or even up to 40/0 or even 70/0 by weight. The amounts of metals they contain is generally at least 5 ppm by weight, often at least 50 ppm by weight, and typically at least 100 ppm by weight or at least 200 ppm by weight. Their nitrogen content is generally at least 0.50%, often at least 1% and very often at least 20% by weight. Their arsenic content is generally greater than 1 ppm by weight, and up to 50 ppm by weight. The process according to the present invention aims at converting shale oils. However, the feed may also contain, in addition to shale oil, other synthetic liquid hydrocarbons, particularly those containing a significant amount of organic cyclic nitrogen compounds. This includes oils derived from coal, oils obtained from heavy tar, tar sands, pyrolysis oils from wood residues such as wood residues, biomass crudes ("biocrudes"), vegetable oils and animal fats. Other hydrocarbon feedstocks can also supplement shale oil. The feeds are selected from the group consisting of vacuum distillates and straight run residues, vacuum distillates and unconverted residues from conversion processes such as, for example, those from distillation to coke ( coking), products resulting from hydroconversion of heavy liquids in fixed bed, products resulting from hydroconversion processes of heavy boilers, and oils désalphatées with solvents (for example the deasphalted oils with propane, with butane and with pentane resulting from the deasphalting of direct distillation vacuum residues or vacuum residues from hydroconversion processes). The fillers may also contain light cutting oil (LCO for "light cycle oil" in English) of various origins, heavy cutting oil (HCO for "heavy cycle oil" in English) of various origins, and also diesel fuel cuts from catalytic cracking generally having a distillation range of about 150 ° C to about 650 ° C. The fillers may also contain aromatic extracts obtained in the context of the manufacture of lubricating oils. The fillers can also be prepared and used in a mixture, in all proportions.

Hydroconversion According to the present invention, the feed is first subjected to a hydroconversion step [step a)] bubbling bed. Hydroconversion is understood to mean hydrogenation, hydrotreatment, hydrodesulfurization, hydrodenitrogenation, hydrodemetallation and hydrocracking reactions. The operation of the bubbling bed catalytic reactor, including the recycle of reactor liquids upwardly through the agitated catalyst bed, is generally well known. Boiling bed technologies use supported catalysts, generally in the form of extrudates whose diameter is generally of the order of 1 ppm or less than 1 mm, for example greater than or equal to 0.7 mm. The catalysts remain inside the reactors and are not evacuated with the products. The catalytic activity can be kept constant by the on-line replacement (addition and withdrawal) of the catalyst. It is therefore not necessary to stop the unit to change the spent catalyst, nor to increase the reaction temperatures along the cycle to compensate for the deactivation. In addition, working under constant operating conditions allows consistent yields and product qualities to be achieved throughout the catalyst cycle. Since the catalyst is kept in agitation by a large recycling of liquid, the pressure drop on the reactor remains low and constant, and the heat of reaction is rapidly averaged over the catalytic bed, which is therefore almost isothermal and does not require cooling. via the injection of quenches. The implementation of boiling bed hydroconversion makes it possible to overcome the problems of catalyst contamination linked to the deposition of impurities present naturally in shale oils.

The conditions of step a) of treating the feedstock in the presence of hydrogen are usually conventional bubbling bed hydroconversion conditions of a liquid hydrocarbon fraction. It is usually carried out under a total pressure of 2 to 35 MPa, preferably 10 to 20 MPa, at a temperature of 300 ° C to 550 ° C and often 400 ° C to 450 ° C. The hourly space velocity (WH) and the hydrogen partial pressure are important factors that are chosen according to the characteristics of the product to be treated and the desired conversion. Most often, the VVH is in a range from 0.2 h -1 to 1.5 h -1 and preferably 0.3 h -1 to 1 h -1. The amount of hydrogen mixed with the feedstock is usually from 50 to 5000 normal cubic meters (Nm3) per cubic meter (m3) of liquid feed, and most often from 100 to 1000 Nm3 / m3, and preferably from 300 to 500 Nm3 / m3; Most often, this hydroconversion stage a) can be carried out under the conditions of the T-STAR® process, as described for example in the article Heavy Oil Hydroprocessing, published by Aiche, March 19-23, 1995, HOUSTON, Texas, paper number 42d. It can also be implemented under the conditions of the H-OIL® process, as described for example in the article published by NPRA, Annual Meeting, March 16-18, 1997, J.J. Colyar and L.I. Wilson under the title THE H-OIL®PROCESS, WORLDWIDE LEADER IN VACUUM RESIDUE HYDROPROCESSING. The hydrogen needed for hydroconversion (and subsequent hydrotreatment) can come from steam reforming hydrocarbons (methane) or gas from oil shale during the production of shale oils.

The catalyst of step a) is preferably a conventional granular hydroconversion catalyst comprising on an amorphous support at least one metal or metal compound having a hydro-dehydrogenating function. In general, a catalyst is used whose porous distribution is suitable for the treatment of metal-containing fillers. The hydro-dehydrogenating function may be provided by at least one Group VIII metal selected from the group consisting of nickel and / or cobalt, optionally in combination with at least one Group VIB metal selected from the group consisting of molybdenum and / or tungsten. For example, a catalyst comprising from 0.5 to 10% by weight of nickel and preferably from 1 to 50% by weight of nickel (expressed as nickel oxide NiO) and from 1 to 300% by weight of nickel may be used. molybdenum, preferably from 5 to 200/0 by weight of molybdenum (expressed as molybdenum oxide MoO3), on an amorphous mineral support. The total content of Group VIB and VIII metal oxides is often from 5 to 400% by weight and in general from 7 to 300% by weight and the weight ratio expressed as metal oxide between group VI metal (or metals). on metal (or metals) of Group VIII is in general from 20 to 1 and most often from 10 to 2. The support of the catalyst will for example be chosen from the group formed by alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals. This support may also contain other compounds, for example oxides chosen from the group formed by boron oxide, zirconia, titanium oxide and phosphoric anhydride. Most often an alumina support is used, and very often a support of alumina doped with phosphorus and possibly boron. In this case, the P 2 O 5 phosphoric anhydride concentration is usually less than about 20% by weight and most often less than about 10% by weight and at least about 0.001% by weight. The concentration of boron trioxide B 2 O 3 is usually from about 0 to about 10% by weight. The alumina used is usually y (gamma) or n (eta) alumina. This catalyst is most often in the form of extruded. Preferably, the catalyst of step a) is based on nickel and molybdenum, doped with phosphorus and supported on alumina. For example, it is possible to use a catalyst of the HTS 458 type marketed by AXENS. Prior to the injection of the feedstock, the catalysts used in the process according to the present invention are preferably subjected to a sulphurization treatment making it possible, at least in part, to convert the metal species into sulphides before they come into contact with the feedstock. treat. This activation treatment by sulfurization is well known to those skilled in the art and can be performed by any method already described in the literature either in-situ, that is to say in the reactor, or ex-situ. The spent catalyst is partly replaced by fresh catalyst by withdrawal at the bottom of the reactor and introduction to the top of the fresh or new catalyst reactor at regular time interval, for example by spot addition, or almost continuously.

For example, fresh catalyst can be introduced every day. The replacement rate of spent catalyst with fresh catalyst can be, for example, from about 0.05 kg to about 10 kg per m 3 of feedstock. This withdrawal and replacement are performed using devices for the continuous operation of this hydroconversion step. The unit usually comprises a recirculation pump for maintaining the bubbling bed catalyst by continuously recycling at least a portion of the liquid withdrawn at the top of the reactor and reinjected at the bottom of the reactor. It is also possible to send the spent catalyst withdrawn from the reactor into a regeneration zone in which the carbon and the sulfur contained therein are eliminated, and then to return this regenerated catalyst to the hydroconversion stage a). The operating conditions coupled with the catalytic activity make it possible to obtain conversion rates of the feedstock which may range from 50 to 950/0, preferably from 70 to 950/0. The conversion rate mentioned above is defined as the mass fraction of the feed at the inlet of the reaction section minus the mass fraction of the heavy fraction having a boiling point greater than 343 ° C. at the outlet of the reaction zone. reaction section, all divided by the mass fraction of the feed at the inlet of the reaction section.

The bubbling bed technology makes it possible to treat highly contaminated loads of metals, sediments and heteroatoms, without encountering problems of loss of pressure or clogging known in the case of use of fixed bed. Metals such as nickel, vanadium, iron and arsenic are largely removed from the charge by settling on the catalysts during the reaction. The remaining (volatile) arsenic will be removed during the hydrotreatment stages by specific guard beds. The sediments contained in the shale oils are also removed by replacing the catalyst in the bubbling bed without disturbing the hydroconversion reactions. These steps also make it possible to remove most of the nitrogen by hydrodenitrogenation, leaving only the most refractory nitrogen compounds. The hydroconversion of step a) makes it possible to obtain an effluent containing at most 3000 ppm, preferably at most 2000 ppm by weight of nitrogen.

Fractionation by atmospheric distillation The effluent obtained in the hydroconversion stage is sent at least partly, and preferably entirely, into a fractionation zone from which a gaseous fraction, a naphtha fraction, is recovered by atmospheric distillation. a diesel fraction and a heavier fraction than the diesel fraction. Preferably, the effluent obtained in step a) is fractionated by atmospheric distillation into a gaseous fraction having a boiling point of less than 50 ° C., a naphtha fraction boiling between about 50 ° C. and 150 ° C., a fraction gas oil boiling between about 150 ° C and 370 ° C, and a heavier fraction than the gas oil fraction generally boiling above 340 ° C, preferably above 370 ° C. The naphtha and diesel fractions are then sent separately to hydrotreatment sections. The heavy fraction undergoes liquid / liquid extraction. The gaseous fraction contains gases (H2, H2S, NH3, H2O, CO2, CO, C1-C4 hydrocarbons, etc.). It may advantageously undergo a purification treatment to recover the hydrogen and recycle it in the hydroconversion section of step a) or in the hydrotreatment sections of steps c) and d). The C3_ and C4 hydrocarbons can, after purification treatments, be used to constitute LPG (liquefied petroleum gas) products. Incondensable gases (C1-C2) are generally used as internal fuel for heating furnaces of hydroconversion and / or hydrotreating reactors. Hydrotreatment of the naphtha fraction and the diesel fraction The naphtha and gasol fractions are then subjected separately to a fixed bed hydrotreatment [steps c) and d]. By hydrotreating is meant hydrodesulfurization, hydrodenitrogenation and hydrodemetallation reactions. The objective is, according to the operating conditions chosen in a more or less severe way, to bring the different cuts to the specifications (sulfur content, smoke point, cetane, aromatic content, etc.) or to produce an oil synthetic crude. The fact of treating the naphtha fraction in a hydrotreatment section and the gas oil fraction in another hydrotreatment section makes it possible to have a better operability in the operating conditions, in order to be able to bring each cut to the required specifications with a maximum yield. and this in one step by cutting. Thus, fractionation after hydrotreatment is not necessary. The difference between the two hydrotreating sections is based more on differences in operating conditions than on the choice of catalyst. The fixed-bed hydrotreatment sections preferably comprise, upstream of the hydrotreatment catalytic beds, specific guard beds for the arsenic (arsenic) compounds and silicon optionally contained in the naphtha and / or diesel fractions. The arsenic compounds having escaped the bubbling bed (because they are generally relatively volatile) are fixed in the guard beds, thus avoiding poisoning the catalysts downstream, and making it possible to obtain fuel bases that are heavily depleted of arsenic. Guard beds for removing arsenic and silicon from naphtha or diesel cuts are known to those skilled in the art. They include for example an absorbent mass comprising nickel deposited on a suitable support (silica, magnesia or alumina) as described in FR2617497, or an absorbent mass comprising copper on a support, as described in FR2762004. Mention may also be made of guard beds marketed by AXENS: ACT 979, ACT989, ACT961, ACT981. The operating conditions of each hydrotreatment section are adapted to the batch to be treated. The operating conditions for hydrotreatment of the naphtha fraction are generally milder than those of the diesel fraction. In the hydrotreatment step of the naphtha fraction [step c)], it is usually carried out under an absolute pressure of 4 to 15 MPa, often 10 to 13 MPa. The temperature in this step c) is usually 280 ° C to 380 ° C, often 300 ° C to 350 ° C. This temperature is usually adjusted according to the desired level of hydrodesulfurization. Most often, the hourly space velocity (WH) is in the range of 0.1 hr-1 to hr-1, and preferably 0.5 hr-l to hr-1. The amount of hydrogen mixed with the feedstock is usually from 100 to 5000 normal cubic meters (Nm3) per cubic meter (m3) of liquid feed, and most often from 200 to 1000 Nm3 / m3, and preferably from 300 to 500 Nm3 / m3. Use is advantageously carried out in the presence of hydrogen sulphide (for the sulphidation of the catalyst) and the partial pressure of the hydrogen sulphide is usually 0.002 times to 0.1 times, and preferably 0.005 times to 0.05 times the total pressure. . In the step of hydrotreating the gas oil fraction [step d)], it is usually carried out under an absolute pressure of 7 to 20 MPa, often 10 to 15 MPa. The temperature in this step c) is usually 320 ° C to 450 ° C, often 340 ° C to 400 ° C. This temperature is usually adjusted according to the desired level of hydrodesulfurization. The hourly mass velocity ((t of feedstock / h) / t of catalyst) is between 0.1 and 1 hr-1. Most often, the hourly space velocity (WH) is in a range of from 0.2 h -1 to 1 h -1, and preferably from 0.3 h -1 to 0.8 h -1. The amount of hydrogen mixed with the feedstock is usually from 100 to 5000 normal cubic meters (Nm3) per cubic meter (m3) of liquid feed, and most often from 200 to 1000 Nm3 / m3, and preferably from 300 to 500 Nm3 / m3. Useful operation is carried out in the presence of hydrogen sulfide, and the partial pressure of hydrogen sulfide is usually 0.002 times to 0.1 times, and preferably 0.005 times to 0.05 times the total pressure.

In the hydrotreatment sections, the ideal catalyst must have a high hydrogenating power, so as to achieve a deep refining of the products, and to obtain a significant lowering of the sulfur and nitrogen content. In the preferred embodiment, the hydrotreatment sections operate at a relatively low temperature, which is in the sense of deep hydrogenation and coking limitation of the catalyst. It is not within the scope of the present invention to use, in the hydrotreatment sections, simultaneously or successively, a single catalyst or several different catalysts.

Usually, the hydrotreatment of steps c) and d) is carried out industrially, in one or more liquid downflow reactors. In both hydrotreating sections [step c) and d)] the same type of catalyst is used, the catalysts in each section being identical or different. At least one fixed bed of conventional hydrotreatment catalyst is used, comprising on an amorphous support at least one metal or metal compound having a hydro-dehydrogenating function.

The hydro-dehydrogenating function may be provided by at least one Group VIII metal selected from the group consisting of nickel and / or cobalt, optionally in combination with at least one Group VIB metal selected from the group consisting of molybdenum and / or tungsten. For example, a catalyst comprising from 0.5 to 10% by weight of nickel and preferably from 1 to 5% by weight of nickel (expressed as nickel oxide NiO) and from 1 to 30% by weight of molybdenum can be used. preferably from 5 to 20% by weight of molybdenum (expressed as molybdenum oxide MoO 3), on an amorphous mineral support. The total metal oxide content of Groups VI and VIII is often from about 5 to about 40% by weight, and generally from about 7 to 30% by weight, and the weight ratio of metal oxide to metal (or metals) of group VIB on metal (or metals) of group VIII is generally from about 20 to about 1, and most often from about 10 to about 2. The carrier is for example selected from the group consisting of alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals. This support may also contain other compounds and for example oxides chosen from the group formed by boron oxide, zirconia, titanium oxide and phosphoric anhydride. Most often an alumina support is used and very often a support of alumina doped with phosphorus and possibly boron. In this case, the P 2 O 5 phosphoric anhydride concentration is usually less than about 20% by weight and most often less than about 10% by weight and at least about 0.001% by weight. The concentration of boron trioxide B 2 O 3 is usually from about 0 to about 10% by weight. The alumina used is usually y (gamma) or n (eta) alumina. This catalyst is most often in the form of beads or extrudates. Before the charge is injected, the catalysts used in the process according to the present invention are preferably subjected to a sulphurization treatment making it possible, at least in part, to transform the metallic species into sulphide before they come into contact with the charge. treat. This activation treatment by sulfurization is well known to those skilled in the art and can be performed by any method already described in the literature either in-situ, that is to say in the reactor, or ex-situ. The hydrotreatment of stage c) of the naphtha section makes it possible to obtain a section containing at most 1 ppm by weight of nitrogen, preferably at most 0.5 ppm of nitrogen and at most 5 ppm by weight of sulfur, of preferably at most 0.5 ppm of sulfur. The hydrotreatment of step d) of the diesel fraction makes it possible to obtain a cut containing at most 100 ppm of nitrogen, preferably at most 200 ppm of nitrogen and at most 50 ppm of sulfur, preferably at most 10 ppm. ppm of sulfur.

Liquid / liquid extraction The heavier fraction than the diesel fraction from the fractionation fraction by atmospheric distillation is then directed to a liquid / liquid type extraction step [step e)]. The purpose of this step is to extract the aromatic compounds, including the refractory nitrogen from the heavy fraction, to obtain a raffinate that can be used as a filler for catalytic cracking in a conventional fluid bed catalytic cracking unit. This makes it possible to maximize the yield of fuel bases. Thus, the liquid / liquid extraction makes it possible to value a fraction that is classically too refractory for a hydrotreatment. The extraction is carried out using a known solvent to preferentially extract aromatic compounds. As a solvent, it is possible to use furfural, N-methyl-2-pyrrolidone (NMP), sulfolane, dimethylformamide (DMF), dimethylsulfoxide (DMSO), phenol, or a mixture of these solvents in equal proportions or different. The liquid / liquid extraction can be carried out by any means known to those skilled in the art. The extraction is generally carried out in a mixer-settler or in an extraction column. Preferably, the extraction is carried out in an extraction column. The operating conditions are in general a solvent / filler ratio of 1/1 to 3/1, preferably of 1/1 to 1.8 / 1, a temperature profile of between room temperature and 150.degree. C., preferably between 50.degree. ° C and 150 ° C. The pressure is between atmospheric pressure and 2 MPa, preferably between atmospheric pressure and 1 MPa.

The chosen solvent has a sufficiently high boiling point in order to be able to thin the heavy fraction resulting from the fractionation without vaporizing, the heavy fraction being typically conveyed at temperatures of between 200 ° C. and 300 ° C. After contact of the solvent with the heavy fraction, two phases are formed: (i) the extract, consisting of the parts of the heavy fraction insoluble in the solvent (and highly concentrated in aromatics containing refractory nitrogen), and (ii) ) the raffinate, consisting of the solvent and the soluble portions of the heavy fraction, which constitutes a catalytically crackable feedstock for increasing the yield of fuel bases. The solvent is distilled off from the soluble parts and recycled internally to the liquid / liquid extraction process; the management of the solvent being known to those skilled in the art. The extraction makes it possible to obtain a raffinate containing at most 1500 ppm, preferably at most 1000 ppm of nitrogen. At least a portion, and preferably all of the raffinate resulting from the liquid / liquid extraction, is preferably sent to a catalytic cracking step. According to a preferred variant, at least a portion and preferably all of the extract obtained in step e) of liquid / liquid extraction is recycled to the inlet of step a). According to another variant, the extract is sent to an oxyvapogazeification section in which it is converted into a gas containing hydrogen and carbon monoxide. This gaseous mixture can be used for the synthesis of methanol or for the synthesis of hydrocarbons by the Fischer-Tropsch reaction. This mixture, in the context of the present invention, is preferably sent to a steam conversion section ("shift" conversion) in which, in the presence of water vapor, it is converted into hydrogen and dioxide. of carbon. The hydrogen obtained can be used in steps a), c) and d) of the process according to the invention. The extract obtained in step e) can also be used as solid fuel, or, after fluxing, as a liquid fuel, or in the composition of bitumens and / or heavy fuels. The liquid / liquid extraction of the heavy fraction thus makes it possible to extract the refractory aromatic compounds containing nitrogen and the contaminants (metals). The fact of carrying out the extraction only on the heavy fraction makes it possible to avoid pressure losses for the catalytic cracking and thus to increase the overall yield of the process. Recycling the extract in step a) of hydroconversion also makes it possible to increase the yield. Catalytic cracking Finally, according to a variant mentioned above in a catalytic cracking step [step f)], at least a portion, and preferably all of the raffinate obtained in step e), can be sent after evaporation of the solvent, in a conventional catalytic cracking section, in which it is conventionally treated under conditions well known to those skilled in the art, to produce a gaseous fraction, a gasoline fraction, a gas oil fraction and a heavy fraction (called "slurry"). according to the English terminology). The diesel fraction will be, for example, at least partly sent to the fuel tanks (pools) and / or recycled, at least partially, or in whole, in step d) hydrotreating diesel. The heavy fraction will be, for example, at least partly, if not all, sent to the heavy fuel tank (pool) and / or recycled at least partly, if not all, to the catalytic cracking step fl. In the context of the present invention, the conventional catalytic cracking expression includes cracking processes comprising at least one partial combustion catalyst regeneration stage, and those comprising at least one total combustion catalyst regeneration stage, and / or those comprising both at least one partial combustion step and at least one total combustion step. For example, a brief description of catalytic cracking (the first industrial implementation of which dates back to 1936 (HOUDRY process) or 1942 for the use of a fluidized bed catalyst) can be found in ULLMANS ENCYCLOPEDIA OF INDUSTRIAL CHEMISTRY VOLUME A 18, 1991, pages 61 to 64. A conventional catalyst comprising a matrix, optionally an additive and at least one zeolite is usually used. The amount of zeolite is variable but usually from about 3 to 600% by weight, often from about 6 to 500% by weight and most often from about 10 to 450% by weight. The zeolite is usually dispersed in the matrix. The amount of additive is usually from about 0 to about 300% by weight. The amount of matrix is 1000% by weight. The additive is generally chosen from the group formed by the metal oxides of group IIA of the periodic table of elements, such as, for example, magnesium oxide or calcium oxide, rare earth oxides and titanates. Group IIA metals. The matrix is most often a silica, an alumina, a silica-alumina, a silica-magnesia, a clay or a mixture of several of these products. The most commonly used zeolite is zeolite Y. The cracking is carried out in a substantially vertical reactor, either in ascending or descending mode. The choice of the catalyst and the operating conditions depend on the desired products as a function of the feedstock treated, as described, for example, in the article by M.Marcilly, pages 990-991, published in the review of the Institut Français du Pétrole nov. .-Dec. 1975 pages 969-1006. The operation is usually at a temperature of 450 ° C to 600 ° C and reactor residence times of less than 1 minute, often from about 0.1 to about 50 seconds. The catalytic cracking step f 1 can also be a catalytic cracking step in a fluidized bed, for example by the process known as R2R. This step can be carried out in a manner conventionally known to those skilled in the art under suitable cracking conditions in order to produce lower molecular weight hydrocarbon products. Functional descriptions and catalysts for use in fluidized bed cracking in this step are described, for example, in US-A-5286690, US-A-65324696 and EP-A-699224.

The fluidized catalytic cracking reactor is operable with upflow or downflow. Although this is not a preferred embodiment of the present invention, it is also conceivable to perform catalytic cracking in a moving bed reactor. Particularly preferred catalytic cracking catalysts are those containing at least one zeolite, usually in admixture with a suitable matrix such as, for example, alumina, silica, silica-alumina. Figure 1 shows schematically the method according to the present invention. Figure 2 schematically shows a variant of the process including the catalytic cracking step.

According to FIG. 1, the feedstock comprising the shale oil (1) to be treated enters through the line (21) in the bubbling bed hydroconversion section (2), in the presence of hydrogen (3), hydrogen (3) being introduced by the line (33). The effluent from the boiling bed hydroconversion section (2) is sent via line (23) to an atmospheric distillation column (4), at the outlet of which a gaseous fraction (30), a naphtha fraction is recovered. (25), a gas oil fraction (27) and a heavier fraction than the gas oil fraction (29). The gaseous fraction (30) containing hydrogen can be purified (not shown) to recycle hydrogen and reinject it into the bubbling bed hydroconversion section (2) via line (33) and / or hydrotreatment (6) and / or (8) via lines (35) and (37). The naphtha fraction (25) is sent to a fixed-bed hydrotreatment section (6) at the outlet from which a naphtha fraction (13) depleted of impurities is recovered. The gas oil fraction (27) is sent to a fixed bed hydrotreatment section (8) at the outlet of which a diesel fuel fraction (15) depleted of impurities is recovered. The two hydrotreatment sections (6) and (8) are fed with hydrogen via lines (35) and (37). The heavier fraction than the gas oil fraction (29) is directed to a liquid / liquid extraction step (10) to extract the aromatics. This extraction step is carried out using a solvent (not shown) and produces a raffinate (17) and an extract (19). The extract (19), via line (39), can be used as fuel or feed a gasification unit to produce hydrogen and energy. It can also be recycled within the hydroconversion section (2) via the line (31).

In FIG. 2, the steps (and reference signs) of hydroconversion, separation and hydrotreatments are identical to FIG. 1. The raffinate (17) leaving the liquid / liquid extraction step can be directed into a catalytic cracking section (12). The effluent from this section is directed via line (43) to a fractionation section (14), preferably an atmospheric distillation, from which a fraction of fuels or middle distillates, including at least one gasoline fraction (45 ), a gas oil fraction (47) and a heavy fraction (51). The diesel fraction (47) is sent, at least in part, to fuel tanks (pools) and / or is recycled, at least in part, or in whole, to step d) of hydrotreating diesel (8) via the line (49). The heavy fraction ("slurry") (51) is, for example, at least partly, if not all, sent to the heavy fuel tank and / or is recycled, at least in part, or in whole, to the catalytic cracking step (12) via the line (53). Example A shale oil is processed, the characteristics of which are shown in Table 1.

Table 1: Characteristics of shale oil feed Density 15/4 - 0.951 Hydrogen% by weight 10.9 Sulfur% by weight 1.9 Nitrogen% by weight 1.8 Oxygen% by weight 2.7 Asphaltenes% by weight 3, Conradson Carbon% wt. 4, 5 Metals ppm 236 The shale oil is processed in a bubbling bed reactor containing commercial HOC 458 catalyst from Axens. The operating conditions are as follows: - Temperature in the reactor: 425 ° C - Pressure: 195 bar (19.5 MPa) - Hydrogen / charge ratio: 400 Nm3 / m3 - WH global: 0.3 h The liquid products from the reactor are fractionated by atmospheric distillation into a naphtha fraction (C5 + - 150 ° C), a gas oil fraction (150-370 ° C) and a residual fraction 370 ° C +. The naphtha fraction is subjected to a fixed bed hydrotreatment using a NiMo catalyst on alumina. The operating conditions are the following: - Temperature in the reactor: 320 ° C - Pressure: 50 bar (5 MPa) - Hydrogen / charge ratio: 400 Nm3 / m3 - global WH: 1 h-1 The diesel fraction is subjected to a fixed bed hydrotreatment using a NiMo catalyst on alumina. The operating conditions are the following: - Temperature in the reactor: 350 ° C - Pressure: 120 bar (12 MPa) - Hydrogen / charge ratio: 400Nm3 / m3 - WH overall: 0.6 h-1 The residual fraction is subjected to a liquid / liquid furfural extraction with a solvent / charge ratio of 1.8 / 1; at a temperature of 100 ° C, and at atmospheric pressure. A raffinate and an extract are obtained. The raffinate is then catalytically cracked using a catalyst containing 200% by weight of zeolite Y and 800% by weight of a silica-alumina matrix. This preheated charge at 135 ° C. is brought into contact at the bottom of a vertical reactor with a hot regenerated catalyst from a regenerator. The catalyst inlet temperature in the reactor is 720 ° C. The ratio of catalyst flow to charge flow is 6.0. The heat input of the catalyst at 720 ° C allows the vaporization of the feedstock and the cracking reaction, which are endothermic. The average residence time of the catalyst in the reaction zone is about 3 seconds. The operating pressure is 1.8 bar absolute. The temperature of the catalyst measured at the outlet of the riser fluidized bed reactor is 525 ° C. The cracked hydrocarbons and the catalyst are separated by cyclones located in a stripper zone where the catalyst is stripped. The catalyst, which is coked during the reaction and then stripped into the disengaging zone, is then fed into the regenerator. The coke content of the solid (delta coke) at the inlet of the regenerator is 0.850 / 0. This coke is burned by air injected into the regenerator. The highly exothermic combustion raises the solid temperature from 525 ° C to 720 ° C. The regenerated and hot catalyst leaves the regenerator and is returned to the bottom of the reactor. The hydrocarbons separated from the catalyst leave the zone of disengagement. They are directed to a main fractionation tower from which the gases and gasoline cuts are leading, then in order of increasing boiling point, the LCO, HCO and slurry sections (370 ° C +) at the bottom of the tower. Table 2 gives the properties of the different loads of each step as well as the yields obtained in the different units and the overall yield. It is then observed that starting from 100% by weight of shale oil, 87.20% by weight of products (LPG, naphtha, middle distillates) are obtained at the commercial specifications Euro V.

Table 2 Unit Bed Extraction Raffinat Extract Total Bubbling FCC Charge Properties Fraction Oil Raffinate Extract 1 shale heavy bed charge bubbling point of ° C C5 + 360+ 360+ 360+ initial cut (° C) Yield% in 100 15.0 8 , 1 6.9 on shale weight oil Density - 0.951 0.926 0.899 0.960 15/4 Sulfur% in 1.9 0.25 0.12 0.40 Weight Total nitrogen% in 1.8 0.60 0.11 1, 14 weight Yield of each unit (G azde% in 2,4 10,0 oil liquefied weight, LPG) Naphtha% in 23,0 55,0 weight Distillates% in 55,5 14,0 means weight Oil no% in 15, 0 converted weight Yield% on oil shale weight (G azde% in 2,4 0,8 3,2 oil liquefied weight, LPG) Naphtha% in 23,0 4,4 27,4 weight Distillates% in 55,5 1.1 56.6 average weight Total bases% in 80.9 6.3 87.2 fuels liquid weights

Claims (15)

REVENDICATIONS1. A hydrocarbon feed conversion process comprising at least one shale oil having a nitrogen content of at least 0.10 / 0, often at least 10/0 and very often at least 20/0 by weight, characterized in that it comprises the following steps: a) The feed is treated in a hydroconversion section in the presence of hydrogen, said section comprising at least one bubbling bed reactor operating at an upward flow of liquid and gas and containing at least one minus a supported catalyst, b) The effluent obtained in step a) is sent, at least in part, and often entirely in a fractionation zone from which a gaseous fraction, a naphtha fraction, is recovered by atmospheric distillation. , a gas oil fraction and a heavier fraction than the gas oil fraction, c) said naphtha fraction is treated, at least in part, and often entirely in a hydrotreating section, in the presence of hydrogen, said section comprising at least one fixed bed reactor containing at least one hydrotreatment catalyst, d) Said gas oil fraction is treated, at least in part, and often completely in another hydrotreatment section, in the presence of hydrogen, said section comprising at least one fixed bed reactor containing at least one hydrotreatment catalyst; e) The heavier fraction than the gas oil fraction is subjected to liquid / liquid extraction in order to obtain a raffinate and an extract. 2. Method according to claim 1, wherein the effluent obtained in step a) is fractionated by atmospheric distillation into a gaseous fraction having a boiling point below 50 ° C, a naphtha fraction boiling between about 50 ° C and 150 ° C, a gas oil fraction boiling between about 150 ° C and 370 ° C, and a heavier fraction than the diesel type fraction boiling generally above 370 ° C. 3. Method according to one of the preceding claims, wherein the solvent of the liquid / liquid extraction step e) is chosen from the group formed by furfural, N-methyl-2-pyrrolidone, sulfolane, dimethylformamide, dimethylsulfoxide, phenol, or a mixture of these solvents in equal or different proportions. 4. Method according to one of the preceding claims, wherein the step e) of liquid / liquid extraction is carried out with a solvent / charge ratio of 1/1 to 3/1, preferably 1/1 to 1.8. / 1, at a temperature between room temperature and 150 ° C, and at a pressure between atmospheric pressure and 2 MPa, preferably between atmospheric pressure and 1 MPa. 5. Method according to one of the preceding claims, wherein the fixed bed hydrotreating sections of steps c) and / or d) comprise, upstream of the hydrotreatment catalytic beds, guard beds specific to the arsenic compounds and of silicon. 6. Method according to one of the preceding claims, wherein at least a portion of the raffinate obtained in step e) liquid / liquid extraction is sent, after evaporation of the solvent, in a catalytic cracking section, called step f ), in which it is treated under conditions making it possible to produce a gas fraction, a gasoline fraction, a gas oil fraction and a heavy fraction. 7. The method of claim 6, wherein at least a portion of the heavy fraction obtained in step f) of catalytic cracking is recycled to the input of this step f). 8. The process according to claim 6, wherein at least a portion of the gas oil fraction obtained in the catalytic cracking step fl is recycled to the diesel fuel hydrotreating step d). 9. Method according to one of the preceding claims, wherein at least a portion of the extract obtained in step e) liquid / liquid extraction is recycled to the input of step a). 10. Method according to one of the preceding claims, wherein the hydroconversion step a) operates at a temperature between 300 ° C and 550 ° C, preferably between 400 ° C and 450 ° C, at a total pressure between 2 and 35 MPa, preferably between 10 and 20 MPa, at an hourly mass velocity (t of feedstock / h) / t of catalyst) of between 0.2 and 1.5 h -1, preferably between 0 and , 3 h-1 and 1 h-1, and a hydrogen / charge ratio of between 50 and 5000 Nm3 / m3, preferably between 100 and 1000 Nm3 / m3. 11. Method according to one of the preceding claims, wherein step c) of hydrotreating the naphtha fraction operates at a temperature between 280 ° C and 380 ° C, preferably between 300 ° C and 350 ° C, at a total pressure of between 4 and 15 MPa, preferably between 10 and 13 MPa, at an hourly mass velocity (t of feedstock / h) / t of catalyst) of between 0.1 and 5 h -1, preferably between 0.5 h -1 and 1 h -1, and at a hydrogen / charge ratio of between 100 and 5000 Nm 3 / m 3, preferably between 200 and 1000 Nm 3 / m 3. 12. Method according to one of the preceding claims, wherein the step d) hydrotreating of the gas oil fraction operates at a temperature between 320 ° C and 450 ° C, preferably between 340 ° C and 400 ° C, at a total pressure of between 7 and 20 MPa, preferably between 10 and 15 MPa, at an hourly mass velocity (t of load / h) / t of catalyst) of between 0.1 and 1 h -1, preferably between 0.3 h -1 and 0.8 h -1, and at a hydrogen / charge ratio of between 100 and 5000 Nm 3 / m 3, preferably between 200 and 1000 Nm 3 / m 3. The process according to one of the preceding claims, wherein the catalyst of hydroconversion step a) comprises a Group VIII metal selected from the group consisting of Ni and / or Co, optionally a Group VIB metal selected from the group formed by Mo and / or W, on an amorphous support chosen from the group formed by alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals. The process according to one of the preceding claims, wherein the catalyst of hydrotreating steps c) and d) comprises a Group VIII metal selected from the group consisting of Ni and / or Co, optionally a selected Group VIB metal. in the group formed by Mo and / or W, on an amorphous support selected from the group formed by alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals. The process according to one of the preceding claims, wherein the hydrocarbon feedstock also comprises a feedstock selected from the group consisting of coal derived oils, heavy tar oils and oil sands, vacuum distillates and the like. direct distillation residues, vacuum distillates and unconverted residues from residue conversion process, deasphalted oils to solvents, light cutting oils, heavy cutting oils, diesel fuel cuts from catalytic cracking having in particular Generally a distillation range of about 150 ° C to about 650 ° C, aromatic extracts obtained in the manufacture of lubricating oils, or mixtures of such fillers.