FR2983865A1 - Conversion of coal to aromatic compounds involves liquefying, separating light hydrocarbons, hydrocracking, separating naphtha and heavier fractions, reforming naphtha to give hydrogen and reformate with aromatic compounds, and separation - Google Patents

Abstract

Conversion of coal to aromatic compounds involves: (A) coal liquefaction in presence of hydrogen; (B) separating effluent obtained into light hydrocarbons fraction; (C) hydrocracking light hydrocarbons fraction in presence of hydrogen in reactor with fixed-bed hydrocracking catalyst, where conversion at 200[deg] C is greater than 30 wt.%; (D) separating effluent into fraction containing naphtha and fraction heavier than naphtha; (E) catalytic reforming fraction containing naphtha to give hydrogen and reformate containing aromatic compounds; and (F) separating aromatic compounds from reformate. Process for conversion of coal to aromatic compounds involves: (A) coal liquefaction step in the presence of hydrogen; (B) step of separation of the effluent obtained at the end of step (A) into a light fraction of hydrocarbons containing compounds boiling at not > 500[deg] C and a residual fraction; (C) hydrocracking step in the presence of hydrogen of at least a proportion of the light fraction of hydrocarbons obtained at the end of step (B) in at least one reactor containing a fixed-bed hydrocracking catalyst, where the conversion of the fraction at 200[deg] C in the hydrocracking step is greater than 30 wt.%; (D) separation of the effluent obtained at the end of step (C) into at least a fraction containing naphtha and a fraction heavier than the naphtha fraction; (E) catalytic reforming step of the fraction containing naphtha, giving hydrogen and a reformate containing aromatic compounds; and (F) separation step of the aromatic compounds from the reformate.

Description

The present invention relates to a process for producing aromatic compounds from coal. More specifically, the present invention relates to a coal conversion process comprising at least one liquefaction step, followed by a fixed bed hydrocracking step of suitable severity to maximize the production of light aromatic precursors and a step catalytic reforming. The process according to the invention also makes it possible to obtain middle distillates. Aromatic compounds, in particular benzene, toluene and xylenes (BTX), are major intermediates in petrochemicals, for the synthesis of resins, plasticizers and polyester fibers. BTX's main source of production is catalytic reforming, which is used by refiners to improve the octane rating of gasoline by making aromatics. This process produces from the naphtha a fraction rich in aromatic hydrocarbons called reformate from which we can extract, separate and transform the aromatics. The production of aromatics is based today mainly on an oil origin. The current international context is marked by a desire to reduce the dependence on raw materials of petroleum origin. In this context, the search for new charges from non-oil sources is an issue of increasing importance. Given the abundant coal reserves, an attractive alternative for the production of petrochemical intermediates is the liquefaction of coal. The production of fuel bases by direct liquefaction of coal is known. Thus, the application FR 2957607 describes a process for direct liquefaction of coal in two successive stages using boiling bed reactors followed by a hydrocracking step. This process makes it possible to obtain fuel bases (diesel and kerosene) which comply with the required specifications despite a high content of naphtheno-aromatics and more particularly of naphthenes.

The present invention aims at producing BTX-type aromatic compounds from coal using a process comprising a liquefaction step, followed by a hydrocracking and catalytic reforming step. More particularly, the present invention relates to a process for converting coal into aromatic compounds comprising the following steps: a) a step of liquefying coal in the presence of hydrogen, b) a step of separating the effluent obtained at the from step a) into a light hydrocarbon fraction containing boiling compounds at a maximum of 500 ° C and a residual fraction, c) a hydrocracking step in the presence of hydrogen of at least a portion of the so-called light fraction of hydrocarbons obtained at the end of step b) in at least one reactor containing a fixed-bed hydrocracking catalyst, the conversion of the fraction 200 ° C + in the hydrocracking stage being greater than 30% by weight, preferably between 50 and 100% by weight, d) separation of the effluent obtained at the end of stage c) into at least one fraction containing naphtha and a fraction heavier than the naphtha fraction , e) a reforming step ca talytic solution of the naphtha-containing fraction for obtaining hydrogen and a reformate containing aromatic compounds; f) a step of separating the aromatics from the reformate. The research work carried out by the applicant on the liquefaction of coal led him to discover that, surprisingly, this process of liquefying coal followed by a fixed-bed hydrocracking step made it possible to obtain, after separation , a naphtha fraction particularly well adapted by its chemical nature and its very small amount of impurities for the production of aromatic compounds by catalytic reforming. Indeed, the naphtha leaving liquefaction and hydrocracking has a very high content of naphthenes (50 to 90 wt%) and at the same time has very few impurities (generally less than 0.5 ppm sulfur). and less than 0.5 ppm nitrogen) to be fed directly into the catalytic reforming without prior pretreatment. Its unique structure gives it an excellent reactivity with regard to aromatization reactions. The liquefaction stage makes it possible to obtain, at first, hydrocarbons still very heavily loaded with impurities: heteroelements of sulfur, nitrogen and oxygen as well as olefins and polyaromatics. Before sending the light fraction (naphtha) of these hydrocarbons into the catalytic reforming, it is thus necessary to carry out a severe hydrocracking step. This hydrocracking step thus makes it possible to obtain, by cracking, a large naphtha fraction and heavier hydrocarbon fractions (essentially gas oil and vacuum gas oil), but also to remove by deep hydrotreatment all the impurities so as not to poison the sensitive catalysts of the product. catalytic reforming thereafter. Similarly, the severity of the hydrocracking step makes it possible to increase the yield of naphtha fraction (to the detriment of middle distillates, in particular gas oil) and thus ultimately to increase the yield of aromatic compounds and hydrogen produced during the reforming. DETAILED DESCRIPTION The description will be made with reference to FIGS. 1 and 2 without the figures limiting the interpretation. The filler The filler used is coal, preferably of the bituminous or sub-bituminous type. However, lignites can also be used. The liquefaction technology also makes it possible to convert coal into co-treatment with other fillers. The coal may be co-treated with a feedstock selected from petroleum residues, petroleum vacuum distillates, crude oils, synthetic crudes, toasted crudes, deasphalted oils, deasphalting resins, asphalts or deasphalting pitches, derivatives of petroleum conversion processes, aromatic extracts from lubricant base production lines, oil sands or their derivatives, oil shales or their derivatives, or mixtures of such fillers. More generally, hydrocarbon feedstocks containing hydrocarbons containing at least 50% by weight of product distilling above 250 ° C. and at least 25% by weight distilling from above will be grouped together under the term hydrocarbon feedstock. above 350 ° C. The coal may also be co-treated with hydrocarbon waste 5 and / or industrial polymers, organic waste or household plastics, vegetable or animal oils and fats, tars and residues that are not or difficult to recover from the gasification and / or FischerTropsch synthesis of biomass, coal or petroleum residues, lignocellulosic biomass or one or more constituents of the cellulosic biomass selected from the group consisting of cellulose, hemicellulose and / or lignin, algae, charcoal , lignocellulosic biomass pyrolysis oil or algae, pyrolytic lignin, products of hydrothermal conversion of lignocellulosic biomass or algae, activated sludge from water treatment plants, or mixtures of such loads.

Pretreatment of the Charge Before liquefaction, the coal may undergo one or more pretreatment steps. The coal undergoes optional pretreatment reducing its ash content, these technologies (washing, extractions ...) are widely described in the literature. With or without center reduction pretreatment, the coal is preferably pretreated to reduce its moisture content (drying), followed by a particle size reduction step (milling). The drying step is carried out at a temperature below 250 ° C, preferably below 200 ° C, preferably for 15 to 200 minutes. The coal is then sent to a mill which makes it possible to reach the desired particle size. After pretreatment, coal particles having a moisture content of 1 to 50%, preferably 1 to 35% and more preferably 1 to 10%, and a particle size of less than 600 μm, preferably less than 600 μm, are obtained. Liquefaction (step a) The coal, after any pretreatment steps described above, is mixed with a solvent, preferably a hydrogen donor solvent. The coal / solvent mixture is a suspension of carbon particles dispersed in said solvent which is subsequently sent to the liquefaction stage. To constitute the suspension, the particle size of the carbon is less than 5 mm, preferably less than 1 mm, preferably less than 600 μm and still more preferably less than 150 μm. The mass ratio solvent / coal is generally from 0.1 to 3, preferably from 0.5 to 2.

The solvent has a triple role: suspending the feedstock upstream of the reaction zone, thus allowing its transport to the latter, and then partial solubilization of the primary products of conversion and transfer of hydrogen to these primary products to allow conversion into liquid by minimizing the amount of solid (coke) and gas formed in said reaction zone.

The solvent may be any type of liquid hydrocarbon known in the art for the preparation of a suspension. The solvent is preferably a hydrogen donor solvent comprising, for example, tetralin and / or naphtheno-aromatic molecules. In the case of a co-treatment with other charges, the solvent may also consist partially or completely of a liquid co-charge, such as, for example, vegetable oils or pyrolysis oils derived from a carbonaceous material ( biomass, coal, oil). According to a preferred variant, the solvent comes from a recycled fraction of the process. This fraction preferably comprises vacuum distillate, and even more preferably vacuum gas oil, resulting from the separation after liquefaction. It is also possible to recycle a portion of the atmospheric distillates, such as diesel, alone or as a mixture with the vacuum distillate fraction. In the present invention, the liquefaction step in the presence of hydrogen can be carried out in the presence of a catalyst supported in a bubbling bed, in the presence of a catalyst dispersed into a bed entrained (also called "slurry" reactor according to the English terminology -saxonne) or without added catalyst (purely thermal conversion). Preferably, the liquefaction step is carried out in the presence of an ebullated bed supported catalyst, preferably in at least two reactors arranged in series containing a bubbling bed supported catalyst. The ebullated bed technology being widely known, only the main operating conditions will be repeated here. Bubbling bed technologies use supported catalysts in the form of extrudates whose diameter is generally of the order of 1 mm or less than 1 mm. The catalysts remain inside the reactors and are not evacuated with the products. The fact of preferably using at least two bubbling bed reactors makes it possible to obtain products of better quality and with a better yield, thus limiting the energy and hydrogen requirements in hydrocracking. In addition, the liquefaction into two reactors makes it possible to have improved operability in terms of the flexibility of the operating conditions and of the catalytic system. It is usually carried out under a pressure of 15 to 25 MPa, preferably 16 to 20 MPa, at a temperature of about 300 ° C to 440 ° C, preferably between 325 ° C to 420 ° C for the first reactor and between 350 ° C and 470 ° C, preferably between 350 and 450 ° C for the second. The hourly mass velocity ((t of feedstock / hr / t of catalyst) is between 0.1 to 5 hr-1 and the amount of hydrogen mixed with the feedstock is usually about 0.1 to 5 normal meters. cubes (Nm3) per kg of filler, preferably from about 0.1 to 3 Nm3 / kg, and most often from about 0.1 to about 2 Nm3 / kg in each reactor. After the first step, the charge conversion is between 30 and 100%, preferably between 50 and 99%, the conversion being able to be defined with respect to insoluble THF for example. The conversion of coal on a dry basis is then all that is not insoluble in THF. Preferably, in the second reactor a temperature at least about 10 ° C higher than that of the reactor of the first stage is used. The reactor pressure in the second liquefaction stage is generally 0.1 to 1 MPa lower than for the first stage reactor to allow for the flow of at least a portion of the effluent from the reactor. of the first step without pumping. Optionally, the effluent obtained at the end of the first liquefaction stage is subjected to a separation of the light fraction and at least a portion, preferably all, of the residual effluent is treated in the second liquefaction stage. This separation is advantageously made in an inter-stage separator described in US Pat. No. 6,270,654 and makes it possible in particular to avoid overcracking of the light fraction in the second liquefaction reactor.

It is also possible to transfer all or part of the used catalyst withdrawn from the reactor of the first liquefaction stage, operating at a lower temperature, directly to the reactor of the second stage, operating at a higher temperature or to transfer in whole or in the used catalyst withdrawn from the reactor of the second stage directly to the reactor of the first stage. This cascade system is described in US Pat. No. 4,816,841. The catalysts used during bubbling bed liquefaction are widely commercialized. These are granular catalysts whose size never reaches those of the catalysts used in bed leads (slurry). The catalyst is most often in the form of extrudates or beads. Typically, they contain at least one hydro-dehydrogenating element deposited on an amorphous support. Generally, the supported catalyst comprises a Group VIII metal selected from the group consisting of Ni, Pd, Pt, Co, Rh and / or Ru, optionally a Group VIB metal selected from the group Mo and / or W, on a amorphous mineral support selected from the group consisting of alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals. The total oxide content of Group VIII and VIB elements is often 5-40 wt% and generally 7-30 wt%. Generally, the weight ratio expressed as Group VI oxide (s) on Group VIII oxide (s) is 1-20 and most often 2-10. For example, a catalyst comprising from 0.5 to 10% by weight of nickel, preferably from 1 to 5% by weight of nickel (expressed as nickel oxide NiO), and from 1 to 30% by weight of molybdenum, Preferably from 5 to 20 wt.% Of molybdenum (expressed as molybdenum oxide MoO 3), on a support. This catalyst may also contain phosphorus (generally less than 20 wt% and most often less than 10 wt%, expressed as phosphorus oxide P2O5). Prior to injection of the feed, the catalysts used in the process according to the present invention are preferably subjected to a sulfurization treatment (in situ or ex situ). The catalysts of the bubbling bed liquefaction steps of the present invention may be the same or different in the reactors. Preferably, the catalysts used are based on CoMo or NiMo on alumina. With reference to FIG. 1, which describes the process according to the invention by liquefaction in two successive boiling bed reactors, the coal (10), preferably pre-treated, and optionally pre-milled so as to facilitate the pretreatment, for reducing its moisture content and ash content is milled in the mill (12) to produce particles of suitable size to form a slurry and be more reactive under the liquefaction conditions. The coal is then contacted with the recycle solvent (15) from the process in the enclosure (14) to form the slurry. If necessary, although rarely necessary, a sulfur-containing compound for maintaining the sulfide catalyst metals may be injected (not shown) into the line emerging from the furnace (14). The suspension is pressurized by the pump (16), preheated in the oven (18), mixed with recycled hydrogen (17) heated in the oven (21), and introduced through the pipe (19) at the bottom of the first reactor (20) bubbling bed operating upflow of liquid and gas and containing at least one hydroconversion catalyst. The reactor (20) usually comprises a recirculation pump (27) for maintaining the bubbling bed catalyst by continuously recycling at least a portion of the liquid withdrawn into the upper part of the reactor and fed back to the bottom of the reactor. The hydrogen can also be introduced with the slurry into the furnace (18) thereby eliminating the furnace (21). The hydrogen supply is supplemented with additional hydrogen (13). The addition of fresh catalyst can be done from the top of the reactor (not shown). The spent catalyst can be withdrawn from the bottom of the reactor (not shown) to either be removed or regenerated to remove carbon and sulfur and / or rejuvenated to remove metals prior to reinjection from the top of the reactor. Optionally, the converted effluent (26) from the first reactor (20) may be separated from the light fraction (71) in an inter-stage separator (70). All or part of the effluent (26) from the first liquefaction reactor (20) is advantageously mixed with additional hydrogen (28), if necessary preheated in the oven (22). This mixture is then injected by the pipe (29) into a second bubbling bed reactor (30) operating at an upward flow of liquid and gas containing at least one hydroconversion catalyst and operating in the same manner as the first reactor . The operating conditions, in particular the temperature, in this reactor are chosen to reach the desired conversion level, as previously described. The reactor (30) usually comprises a recirculation pump (37) for maintaining the bubbling bed catalyst by continuously recycling at least a portion of the liquid withdrawn into the upper part of the reactor and reinjected at the bottom of the reactor.

According to another embodiment, the liquefaction stage can also be carried out in the presence of a bed-dispersed catalysts catalyst. Slurry reactor liquefaction technologies use a dispersed catalyst (also referred to as a slurry catalyst thereafter) in the form of very small particles, the size of which is a few tens of microns or less (typically 0.001 to 100 μm). The catalysts, or their precursors, are injected with the feed to be converted at the inlet of the reactors. The catalysts pass through the reactors with the feedstocks and the products being converted, and then are driven with the reaction products out of the reactors. They are found after separation in the heavy residual fraction.

The slurry catalyst is a sulfurized catalyst preferably containing at least one member selected from the group consisting of Mo, Fe, Ni, W, Co, V, Ru. These catalysts are generally monometallic or bimetallic (by combining, for example, a non-noble group VIIIB element (Co, Ni, Fe) and a group VIB element (Mo, W) .The catalysts used may be heterogeneous solid powders ( such as natural ores, iron sulphate, etc.), dispersed catalysts derived from water-soluble dispersed catalyst ("water soluble dispersed catalyst") such as phosphomolybdic acid, ammonium molybdate, or A mixture of Mo or Ni oxide with aqueous ammonia, the catalysts used are preferably derived from soluble precursors in an organic phase ("oil soluble dispersed catalyst"). The precursors are organometallic compounds such as naphthenates of Mo, Co, Fe, or Ni or such as multi-carbonyl compounds of these metals, for example 2-ethyl hexanoates of Mo or Ni, acetylacetonates of Mo or Ni, C7-C12 fatty acid salts of Mo or W, etc. They can be used in the presence of a surfactant to improve the dispersion of the metals, when the catalyst is bimetallic. Such precursors and catalysts that can be used in the process according to the invention are widely described in the literature.

Additives may be added during the preparation of the catalyst or to the slurry catalyst before it is injected into the reactor. These additives are described in the literature. The operating conditions of the slurry reactor liquefaction stage are identical to those described in the case of bubbling bed liquefaction.

According to another embodiment, the liquefaction stage can also be carried out purely thermally (without added catalyst). The operating conditions of the thermal liquefaction step without added catalyst are identical to those described in the case of bubbling bed liquefaction.

According to a preferred variant, the liquefaction stage is carried out in at least two reactors arranged in series. These reactors may contain either at least one dispersed catalyst, at least one supported catalyst, a mixture of dispersed catalyst (s) and supported, or no added catalyst.

According to one variant, the first reactor contains a dispersed catalyst and the second reactor contains a supported catalyst. According to another variant, the first reactor contains a supported catalyst and the second reactor contains a dispersed catalyst. According to another variant, the first reactor contains no added catalyst and the second reactor contains a dispersed and / or supported catalyst. Separation of the liquefaction effluent (step b) The effluent obtained after the liquefaction is separated (generally in a HPHT high temperature high pressure separator) into a light hydrocarbon fraction containing boiling compounds at most 500 ° C and a residual fraction. The separation is not made according to a precise cutting point, it is more like a flash. Additional separation steps can be provided. The separation step may advantageously be carried out by methods well known to those skilled in the art such as flash, distillation, stripping, liquid / liquid extraction etc. Preferably, the separation is carried out in a fractionation section which may first comprise an HPHT separator, and optionally a low temperature high pressure separator (HPBT), and / or atmospheric distillation and / or vacuum distillation. Preferably, the separation step b) makes it possible to obtain a gaseous phase, at least one atmospheric distillate fraction containing naphtha, kerosene and / or diesel, a vacuum distillate fraction and a vacuum residue fraction. At least a portion and preferably all of the atmospheric distillate fraction, optionally supplemented with at least a portion of the atmospheric residue fraction and / or part of the vacuum distillate fraction and / or other co-charges, is sent in the hydrocracking step. As a co-feed, it is possible to use vacuum distillates of petroleum origin, deasphalted oils, deasphalting resins, derivatives of petroleum conversion processes (heavy or light oils resulting from catalytic cracking, vacuum gas oil from coking ...), aromatic extracts from lubricant base production lines, or mixtures of such fillers. Any other types of non-oil or renewable co-charges mentioned above may also be used in the paragraph "charges".

At least a portion and preferably all of the vacuum distillate fraction is recycled as a solvent in liquefaction step a). The separation step b) can be performed with or without intermediate decompression.

According to a first embodiment, the effluent resulting from the liquefaction undergoes a separation step with decompression between liquefaction and hydrocracking. This configuration can be qualified as a non-integrated diagram and is illustrated in FIG. 1. Referring to this figure, the effluent treated in the liquefaction reactor (30) is sent via the line (38) into a high pressure high separator. temperature (HPHT) (40) from which a gaseous fraction (41) and a liquid fraction (44) are recovered. The gaseous fraction (41) is sent, optionally mixed with the vapor phase (71) from the optional inter-stage separator (70) between the two liquefaction reactors, generally via an exchanger (not shown) or an air cooler (48). for cooling to a low temperature high pressure separator (HPBT) (72) from which a vapor phase (73) containing the gases (H2, H2S, NH3, H2O, CO2, CO, C1-C4 hydrocarbons, etc.) is recovered. ) and a liquid phase (74). The vapor phase (73) of the low temperature high pressure separator (HPBT) (72) is treated in the hydrogen purification unit (42) from which the hydrogen (43) is recovered for recycling via the compressor. (45) and the line (49) to the reactors (20) and / or (30). The gases containing undesirable nitrogen, sulfur and oxygen compounds are removed from the installation (stream (46)). The liquid phase (74) of the low temperature high pressure separator (HPBT) (72) is expanded in the device (76) and sent to the fractionation system (50). The liquid phase (44) from the high temperature high pressure separation (HPHT) (40) is expanded in the device (47) and sent to the fractionation system (50). Of course, the fractions (74) and (44) can be sent together, after expansion, to the system (50). The fractionation system (50) typically comprises an atmospheric distillation system for producing a gaseous effluent (51), an atmospheric distillate fraction (52) and in particular containing naphtha, kerosene and diesel and an atmospheric residue fraction (55). Part of the atmospheric residue fraction can be sent via line (53) in line (52) to be treated in the hydrocracker. All or part of the atmospheric residue fraction (55) is sent to a vacuum distillation column (56) to recover a vacuum residue fraction (57), unconverted coal and ash, and a vacuum distillate fraction (58). ) containing vacuum gas oil. The vacuum distillate fraction (58) serves at least partially as a solvent for liquefaction and is recycled after pressurization (59) through the line (15) to the vessel (14) to be mixed with the coal. Part of the vacuum distillate fraction (58) not used as a solvent can be introduced via line (54) into the line (52) for further processing in the hydrocracker (80). According to a second embodiment, the effluent resulting from the direct liquefaction undergoes a step of separation without decompression between liquefaction and hydrocracking. This configuration can be described as an integrated diagram and is illustrated in FIG. 2. Referring to this figure, the effluent treated in the second liquefaction reactor (30) is sent via line (38) into a high pressure separator. high temperature (HPHT) (40) from which the so-called light fraction (41) and the residual fraction (44) are recovered. The light fraction (41) is sent directly, optionally mixed with the vapor phase (71) from the optional inter-stage separator (70) between the two reactors, via the line (150) in the hydrocracking reactor.

The residual fraction (44) from the high temperature high pressure separation (HPHT) (40) is expanded in the device (61) and sent to the fractionation system (56). The fractionation system (56) preferably comprises a vacuum distillation system for recovering a vacuum distillate fraction containing the vacuum gas oil (58) and a vacuum residue fraction (57), the unconverted coal and the ashes. . Part of the vacuum distillate (58) may also be fed through line (54) to be treated in the hydrocracker. The vacuum distillate fraction (58) serves at least partially as a solvent for liquefaction and is recycled after pressurization (59) through line (15) to the vessel (14) to be mixed with the coal.

The separation according to the integrated scheme allows a better thermal integration, not to recompress the load sent to the hydrocracking and results in a saving of energy and equipment. This embodiment also makes it possible, by its simplified intermediate splitting, to reduce the consumption of utilities and therefore the investment cost.

The light fractions resulting from the separation steps (whether by the integrated or non-integrated scheme) preferably undergo a purification treatment to recover the hydrogen and recycle it to the liquefaction and / or hydrocracking reactors. The gas phase from the optional inter-stage separator can also be added. The so-called incondensable gases (C1, C2), which can be used either as a fuel used in the furnaces of the different steps of the process scheme, or to be sent to a steam reforming unit for the manufacture of additional hydrogen, or alternatively, can be recovered as well. be sent to a steam cracking furnace to produce olefins and aromatics. Finally, a C3, C4 cut is recovered that can be directly upgraded to liquefied petroleum gas or can be recovered along the same lines as those evoked for incondensable gases. Hydrocracking (step c) The objective of the hydrocracking step is to carry out on one side a rather severe hydrocracking in order to obtain a high yield in naphtha section (and ultimately in aromatic compounds and in hydrogen) at the same time. time that a very deep hydrotreatment to obtain a naphthenic cut sufficiently pure in terms of impurities not to poison catalytic reforming catalysts. Hydrocracking is understood to mean hydrocracking reactions accompanied by hydrotreatment reactions (hydrodenitrogenation, hydrodesulfurization), hydroisomerization, hydrogenation of aromatics and opening of naphthenic rings.

The hydrocracking step according to the invention operates in the presence of hydrogen and of a catalyst at a temperature preferably of between 250 and 480 ° C., preferably between 320 and 450 ° C., very preferably between 380 and 450 ° C. and 435 ° C, at a pressure between 2 and 25 MPa, preferably between 3 and 20 MPa, at a space velocity of between 0.1 and 20 h -1, preferably 0.1 and 6 h -1, of preferred way between 0.2 and 3 h -1, and the amount of hydrogen introduced is such that the volume ratio hydrogen on hydrocarbons is between 80 and 5000 Nm3 / m3 and most often between 100 and 3000 Nm3 / m3. These operating conditions used in the process according to the invention generally make it possible to achieve pass conversions, products having boiling points below 340 ° C., and better still below 370 ° C., greater than 30% by weight and even more preferably between 50 and 100 wt%. The hydrocracking step according to the invention may advantageously be carried out in one or, preferably, several fixed bed catalytic beds, in one or more reactors, in a so-called one-step hydrocracking scheme, with or without separation. intermediate, or, to maximize the yield of naphtha, in a so-called two-step hydrocracking scheme, the so-called one or two-step schemes operating with or without liquid recycling of the unconverted fraction, optionally in combination with a catalyst of conventional hydrotreatment upstream of the hydrocracking catalyst. Such methods are widely known in the prior art.

The hydrocracking process may include a first hydrotreating step (also known as hydrorefining) to reduce the heteroatom content before hydrocracking. Such methods are widely known in the prior art. The hydrocracking catalysts used in the hydrocracking processes are all of the bifunctional type associating an acid function with a hydrogenating function. The acid function is provided by supports whose surfaces generally vary from 150 to 800 m 2 / g and having a surface acidity, such as halogenated aluminas (chlorinated or fluorinated in particular), the combinations of oxides of boron and aluminum, amorphous silica-alumina and zeolites. The hydrogenating function is provided either by one or more Group VIB metals of the Periodic Table of Elements, or by a combination of at least one Group VIB metal of the Periodic Table and at least one Group VIII metal. The catalysts may be catalysts comprising Group VIII metals, for example nickel and / or cobalt, most often in combination with at least one Group VIB metal, for example molybdenum and / or tungsten. For example, it is possible to use a catalyst comprising from 0.5 to 10% by weight of nickel (expressed as nickel oxide NiO) and from 1 to 40% by weight of molybdenum, preferably from 5 to 30% by weight of molybdenum (expressed as oxide). of molybdenum MoO3) on an acidic mineral support. The total content of Group VI and VIII metal oxides in the catalyst is generally between 5 and 40 wt%. The weight ratio (based on the metal oxides) between the Group VI metal (metals) and the Group VIII metal (metals) is, in general, about 20 to about 1, and most often about 10 to In the case where the catalyst comprises at least one Group VIB metal in combination with at least one Group VIII non-noble metal, said catalyst is preferably a sulfurized catalyst.

Advantageously, the following combinations of metals are used: NiMo, CoMo, NiW, CoW, NiMoW and even more advantageously NiMo, NiW and NiMoW, more preferably NiMoW. The support will, for example, be 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 boron oxide, zirconia, titanium oxide and phosphoric anhydride. Most often, a support of alumina and preferably alumina ri or y is used.

The catalyst may also contain a promoter element such as phosphorus and / or boron. This element may have been introduced into the matrix or preferably deposited on the support. Silicon may also be deposited on the support, alone or with phosphorus and / or boron. Preferably, the catalysts contain silicon deposited on a support such as alumina, optionally with phosphorus and / or boron deposited on the support, and also containing at least one metal of group VIII (Ni, Co ) and at least one Group VIB metal (Mo, W). The concentration of said element is usually less than 20% by weight (based on oxide) and most often less than 10% by weight. When boron trioxide (B203) is present, its concentration is less than 10% by weight.

Other conventional catalysts comprise zeolite Y of structural type FAU, an amorphous refractory oxide support (alumina most often) and at least one hydro-dehydrogenating element (generally at least one element of groups VIB and VIII, and most often at least one element of group VIB and at least one element of group VIII).

Other catalysts are so-called composite catalysts and comprise at least one hydro-dehydrogenating element chosen from the group consisting of group VIB and group VIII elements and a support based on silicoaluminum matrix and based on at least one zeolite as described in EP1711260. In order to maximize the yield of the hydrocracked naphtha, and ultimately the aromatics after catalytic reforming of said naphtha, the hydrocracking catalyst of step c) preferably comprises a zeolite. Prior to the injection of the feed, the catalysts used in the process according to the present invention are preferably subjected to a sulfurization treatment (in situ or ex situ). With reference to FIGS. 1 and 2, the light fraction resulting from atmospheric distillation (52) according to the non-integrated scheme or from the HPHT separator (150) 10 according to the integrated scheme and containing, in particular, naphtha, kerosene and diesel, optionally supplemented with a portion of the vacuum distillate (54) and / or another co-feed, is fed to the fixed bed hydrocracking reactor (80). It is mixed with recycled hydrogen (66), optionally preheated in the furnace (60) and introduced through line (62) at the top of the fixed bed hydrocracking reactor (80) operating in a downflow of liquid and of gas and containing at least one hydrocracking catalyst. The hydrogen supply is supplemented by additional hydrogen (67). If necessary, the recycled and / or auxiliary hydrogen can also be introduced into the hydrocracking reactor between the various catalytic beds, for example via lines (68) and (69) (quench), in the case of a reactor with 3 catalyst beds. Separation after hydrocracking (step d) The effluent obtained at the end of the hydrocracking step undergoes at least one separation step in order to recover at least one naphtha fraction and send it to catalytic reforming at the end of the hydrocracking step. after. The separation step may advantageously be carried out by methods well known to those skilled in the art such as flash, distillation, stripping, liquid / liquid extraction etc. It preferably comprises a fractionation section which incorporates a high temperature high pressure separator (HPHT) and then an atmospheric distillation. Referring to Figure 1 and preferably, the separation of the effluent (82) is performed in a fractionation section (84) which incorporates a high temperature high pressure separator (HPHT), atmospheric distillation. and optionally vacuum distillation (not shown) which makes it possible to separate a gaseous phase (86), at least a naphtha fraction (88) and a heavier fraction than the naphtha fraction (90). The gaseous fraction (86) is treated in the hydrogen purification unit (106) from which the hydrogen (108) is recovered for recirculation via the compressor (110) and the line (66) to the jet reactors. hydrocracking (80) and / or liquefaction reactors (20) and (30) (not shown). The gases containing undesirable nitrogen, sulfur and oxygen compounds are removed from the installation (stream (112)). Incondensable gases (C1, C2) and liquefied petroleum gas (C3, C4) can be recovered along the same routes as those resulting from liquefaction. According to a variant making it possible to maximize the naphtha fraction, at least part of the heavier fraction than the naphtha fraction (90) is preferably recycled in the hydrocracking step c) (116). In the case of total recycling in particular, a purge (114) is provided. According to another variant (not shown), this heavier fraction than the naphtha fraction (90) can be separated more advantageously, preferably by atmospheric distillation, to obtain at least a fraction of middle distillates (kerosene and / or diesel) and a vacuum distillate fraction containing vacuum gas oil. According to another variant of the process (not shown), the heavier fraction than the naphtha fraction (90) can be at least partly fed to a steam cracker to obtain light olefinic compounds such as ethylene and / or propylene. The heavy fuel oil leaving the steam cracker, generally difficult to recover, can then be advantageously recycled to extinction in the first and / or second liquefaction reactor. It can still be sent to the coal gasification unit, if it exists, to produce hydrogen. According to this variant, the process according to the invention thus makes it possible to maximize the production of aromatics and light olefins from coal.

The naphtha fraction (88) obtained can advantageously be separated (89) into a light (C5-C6) naphtha fraction (96) which is preferably at least partly subjected to an isomerization process (94) to produce isomerate (base for road petrol) (99) and a heavy naphtha fraction (C7-150 at 200 ° C) (98) which is at least partially subjected to the catalytic reforming step (100) to produce reformate (102). ) rich in aromatics. The isomerization processes are widely known in the prior art, the isomerization makes it possible to convert a linear paraffin into isomerized paraffin in order to increase its octane number. The naphtha fraction (88) can also be sent in its entirety in catalytic reforming, without prior separation. Catalytic reforming (step e) The naphtha fraction obtained after separation of the hydrocracking effluent has a high content of naphthenes and very few impurities thanks to severe hydrocracking. It is thus a particularly well suited filler for catalytic reforming. More particularly, the naphtha fraction which is to be sent in the catalytic reforming generally contains between 1 to 50% by weight, preferably between 5 to 30% by weight of paraffins, and between 20 to 100% by weight of naphthenes, preferably between 50 and 90% by weight. and between 0 to 20% aromatic wt. In terms of impurities, it generally comprises a nitrogen content of less than 0.5 ppm and a sulfur content of less than 0.5 ppm. The chemical reactions involved in the reforming process are numerous. They are well known; for the reactions beneficial to the formation of aromatics and to the improvement of the octane number, the dehydrogenation of naphthenes, isomerization of cyclopentanic rings, isomerization of paraffins, dehydrocyclization of paraffins, and for adverse reactions, hydrogenolysis and hydrocracking of paraffins and naphthenes. Also, it is known that catalytic reforming catalysts are particularly susceptible to poisoning which can be caused by metal impurities, sulfur, nitrogen, water and halides.

The catalytic reforming step may be carried out according to the invention according to any of the known methods using any of the known catalysts, and is not limited to any particular process or catalyst. Numerous patents deal with processes for reforming or producing aromatic compounds with continuous or sequential regeneration of the catalyst. The process schemes generally employ at least two reactors, in which a moving catalyst bed is traversed from top to bottom, traversed by a charge composed of hydrocarbons and hydrogen, charge heated between each reactor. Other process schemes use fixed bed reactors. The continuous process of catalytic reforming of hydrocarbons is a process known to those skilled in the art, it uses a reaction zone comprising a series of 3 or 4 reactors in series, working in a moving bed, and has a regeneration zone of catalyst which itself comprises a number of steps, including a coke combustion step deposited on the catalyst in the reaction zone, an oxychlorination step, and a final step of reducing the catalyst to hydrogen. After the regeneration zone, the catalyst is reintroduced at the top of the first reactor of the reaction zone. This method is for example described in the application FR2801604 or FR2946660. The charge treatment in the reforming reactor (s) is generally carried out at a pressure of 0.1 to 4 MPa and preferably 0.3 to 1.5 MPa at a temperature between 400 and 700. ° C and preferably between 430 and 550 ° C, at a space velocity of 0.1 to 10 h -1 and preferably from 1 to 4 h -1 and with a recycled hydrogen / hydrocarbon ratio (mol) of 0, 1 to 10 and preferably between 1 to 5, and more particularly from 2 to 4 for the process for producing aromatics. The catalyst generally comprises a support (for example formed of at least one refractory oxide, the support may also include one or more zeolites), at least one noble metal (platinum preferably), and preferably at least one promoter metal. (eg, tin or rhenium), at least one halogen and optionally one or more additional elements (such as alkali, alkaline earth, lanthanide, silicon, Group IV B elements, non-noble metals, Group III elements A, etc.). These catalysts are widely described in the literature. Reforming produces a reformate comprising at least 70% aromatics. The conversion is usually above 80%.

The hydrogen (104) produced in the catalytic reforming step e) is preferably recycled to the liquefaction step a) and / or the hydrocracking step c). Separation of aromatic compounds from the reformate (step f) The separation of the aromatic compounds contained in the reformate can advantageously be carried out by any method known to those skilled in the art. Preferably, it is carried out by liquid-liquid extraction, extractive distillation, adsorption and / or crystallization. These methods are known to those skilled in the art.

The liquid-liquid extraction makes it possible to extract the aromatic compounds in the solvent constituting the extract. The paraffinic or naphthenic fractions are insoluble in the solvent. Solvents such as sulfolane, N-methyl-2-pyrrolidone (NMP) or dimethylsulfoxide (DMSO) are generally used. The main extractive agents used in the extractive distillation are N20 methyl-2-pyrrolidone (NMP), n-formylmorpholine (NFM) and dimethylformamide (DMF). Aromatic compounds, essentially of the BTX type (benzene, toluene, xylenes and ethylbenzene) are thus obtained from the coal. EXAMPLES Example 1: Liquefaction steps The two liquefaction stages in bubbling bed reactors are carried out with bituminous type coal, previously ground and dried. The operating conditions for the liquefaction are shown in Table 1, the liquefaction yields in Table 2.

Table 1: Operating conditions for two-stage liquefaction NiMo / Alumina catalyst R1 reactor temperature (° C) 410 Reactor temperature R2 (° C) 440 Pressure, MPa 17 VVH R1 (kg / h dry coal / kg catalyst) 1.2 VVH R2 (kg / h dry coal / kg catalyst) 1.2 H2 input (Nm3 / kg dry coal) 2.8 liquid / coal recycling 1.1 Table 2: Liquefaction yields in two stages (% wt / ashless dry coal, H2 consumption included) Products Yields / coal (% m / m) C1-C4 (gas) 13.53 C5-199 ° C 7.34 199-260 ° C 12.65 260-343 ° C 30.33 343-388 ° C 8.53 388-454 ° C 4.04 454-523 ° C 1.20 523 ° C + 2.41 unconverted coal 13.23 H2O / CO / CO2 / NI-13 / H2S 13.80 C5-388 ° C 58.85 200 ° C + in C5-388 ° C 51.51 Example 2: Hydrocracking step HCK one step maxi distillate moven (HCKO) (non-compliant) This example is an example of a maximum fuel base used to maximize the average distillate yield (kerosene and diesel), gasoline being a priori sent to the catalytic eforming more to make it as fuel as BTX aromatic bases for chemistry. It is not optimized to produce a maximum of gas and therefore to produce a maximum of aromatics. The distillate cuts C5-199 ° C, 199-260 ° C, 260-343 ° C and 343-388 ° C obtained at the liquefaction outlet (Table 2), representing a yield of 58.85% m / m on ashless dry charcoal are mixed (noted C5-388 ° C) to hydrocracking. The portion of the heavy fraction that is not recycled, unconverted coal, and ash are sent to gasification to produce H2. The operating conditions for the hydrocracking are shown in Table 3, the hydrocracking yields in Table 4.

Table 3: Operating conditions of hydrocracking HCKO Catalyst NiW / Silica-Alumina Pressure, MPa 16 Temperature (° C) 392 VVH (Nm3 / h C5-388 ° C / m3 of catalyst) 0.5 H2 / HC reactor inlet (Nm3 / h H2 / Nm3 C5-388 ° C) 1300 Recycling residual fraction no20 Table 4: Hydrocracking efficiencies HCKO (`) / 0 wt / dry ashless coal, consumptionH2 included) Products Yield / dry coal (% m / m ) H2S / NH3 / H2O 1.00 C1-C4 0.45 C5-200 ° C 11.83 200-250 ° C 14.78 250-350 ° C 30.77 350 ° C + 1.47 200 ° C + 47.02 Net conversion of 200 ° C + / liquefied 9% Table 5 presents the physicochemical properties of the C5-200 ° C wide naptha section of the ex-liquefied hydrocracked coal effluent, as well as the properties of the 200 ° C + wide gasoil fraction (base for jet fuel and diesel fuel). TABLE 5 Physicochemical properties of HCKO cuts (% wt / liquefied C5-388 ° C entry, including H2 consumption) Cut / method Gas naphtha unit Method of analysis Cutting point ° C C5-200 200+ Yield% m / m 20.11 79.89 ASTM D2892 Density to NF EN ISO g / cm3 0.825 0.880 15 ° C 12185 Hydrogen NMR% m / m 13.70 13.35 ASTM D7171 ppm Nitrogen m / m <0.3 0.4 ASTM4629 ASTM D2622 or PP m Sulfur (UV) m / m <0.5 5 NF EN ISO 20884 Flow ° C -24 ASTM D97 ASTM RON / MON CFR 56/53 D2699 / D2700 Cetane number CFR 51 ASTM D613 / 86 DS 0.5% 73 173 ASTM D2887 DS 5`) / 0 86 213 DS 50 ° / 0 ° C 142 265 DS 95`) / 0 195 350 DS 99.5% 209 430 n-Paraffins% m / m 6.0 6.2 GC * GC IFPEN iso Paraffins% m / m 7.0 4.5 Naphthenes% m / m 78.6 85.0 Monoaromatics% m / m 5.7 2.7 Diaromatics% m / m 2.7 1.6 Example 3: HCK step hydrocracking step (HCK1) (compliant) This example is an example in hydrocracking mode a gasoline max stage without recycling the hydrocracked residual fraction at the entrance hydrocracking, the hydrocracked naphtha being sent to catalytic reforming to essentially make it BTX aromatics for the chemistry. The feedstock sent to the hydrocracking is the same as for example 2: C5-388 ° C cut representing a yield of 58.85% w / w on dry charcoal and ashless. The hydrocracking operating conditions are given in Table 6, the hydrocracking yields in Table 7. Table 6: Operating conditions of hydrocracking HCK1 Catalyst NiW / Silica-Alumina Pressure, MPa 16 Temperature (° C) 410 VVH (Nm3 / h C5-388 ° C / m3 of catalyst) 0.33 H2 / HC reactor inlet (Nm3 / h H2 / Nm3 C5-388 ° C) 2600 Recycling residual fraction no Table 7: HCK1 hydrocracking efficiencies ( % w / ash-free dry coal, H2 consumption included) Products Yield / dry coal (cYom / m) H2S / NH3 / H2O 1.01 C1-C4 2.42 C5-200 ° C 23.37 200-250 ° C 15.16 250-350 ° C 18.12 350 ° C + 0.86 200 ° C + 34.14 Net conversion of 200 ° C + / liquefied 34% 2 983 865 28 Table 8 shows the physico-chemical properties of the C5-200 ° C broad naptha section of the ex-liquefied hydrocracked effluent of coal and the properties of the 200 ° C + wide gasoil fraction (base for jet fuel and diesel fuel).

Table 8: Physicochemical properties of HCK1 cuts (% wt / liquefied entry C5-388 ° C, H2 consumption included) Cut / method Unit naphtha gas oil Method of analysis Cutting point ° C C5-200 200+ Yield% m / m 40.63 59.37 ASTM D2892 Density at g / cm3 0.813 0.857 NF EN ISO 12185 15 ° C NMR hydrogen% m / m 13.80 13.45 ASTM D7171 Nitrogen ppm <0.3 0.5 ASTM4629 m / m Sulfur (UV) ppm <0.5 <5 ASTM D2622 or NF m / m EN ISO 20884 Flow rate ° C <-48 ASTM D97 RON / MY CFR 55/52 ASTM D2699 / D2700 Cetane number CFR 52 ASTM D613 / 86 DS 0.5% 70 193 ASTM D2887 DS 5% 83 209 DS 50% ° C 140 254 DS 95% 194 322 DS 99.5`) / 0 208 371 n-Paraffins% m / m 4.9 6.5 GC * GC IFPEN iso Paraffins% m / m 6.1 7.0 Naphthenes% m / m 85.7 83.5 Monoaromatics% m / m 3.0 Dihydromers% m / m 0.3 0.5 2 983 865 29 Example 4: 2-step hydrocracking step (HCK2) (compliant) This example is an example in two-stage gasoline hydrocracking mode with recycling of the hydrocracked residual fraction 250 ° C + at the Hydrocracked naphtha is fed to catalytic reforming to essentially make it BTX aromatics for the chemistry. The feedstock sent to the hydrocracking is the same as for example 2: C5I-388 ° C cutting representing a yield of 58.85% m / m on dry charcoal and ashless. The hydrocracking operating conditions are shown in Table 9, the yields of hydrocracking in Table 10.

Table 9: Operating conditions of hydrocracking HCK2 Catalyst NiW / Silica-Alumina + Zeolite Pressure, MPa 16 Temperature (° C) 390 VVH (Nm3 / h C5-388 ° C / m3 of catalyst) 0.33 H2 / HC reactor inlet ( Nm3 / h H2 / Nm3 C5-388 ° C) 1300 Recycling residual fraction yes (250 ° C +) Table 10: Hydrocracking efficiencies HCK1 (`) / 0 wt / ash-free dry charcoal, H2 consumption included) Products Yields / dry coal (% m / m) H2S / NH3 / H2O 1.01 C1-C4 8.16 C5-200 ° C 43.56 200-250 ° C 8.77 250 ° C 0 200 ° C + 8.77 Net conversion of 200 ° C + / liquefied 83% The table 11 shows the physicochemical properties of the C5-200 ° C broad naptha section of the ex-liquefied coal hydrocracked effluent as well as the properties of the 200 ° C + wide gasoil fraction.

Table 11: Physico-chemical properties of HCK2 cuts Section / method Unit naphtha gasol Method of analysis C5- Cutting point ° C 200+ 200 Yield% m / m 83.23 16.77 ASTM D2892 Density at 15 ° C g / cm3 0.780 0.840 NF EN ISO 12185 NMR hydrogen% m / m 14.4 14.0 ASTM D7171 Nitrogen ppm <0.3 0.5 ASTM4629 m / m Sulfur (UV) ppm <0.5 <2 ASTM D2622 or NF m / m EN ISO 20884 Pt flow ° C <-48 ASTM D97 RON / MON CFR 59/56 ASTM D2699 / D2700 Crystal Clearance ° C -66 ASTM D7153 Cetane Number CFR 49 ASTM D613 / 86 Smoke Point mm 24 ISO 3014 DS 0.5% 50 180 ASTM D2887 DS 5% 67 209 DS 50% ° C 132 220 DS 95% 192 252 DS 99.5% 205 260 n-Paraffins% m / m 6.9 4.5 GC * GC IFPEN iso Paraffins% m / m 11.4 7.0 Naphthenes% m / m 80.7 86.5 Monoaromatics% m / m 1 2.0 Diaromatics% m / m 0 0 Example 5: Catalytic reforming of naphthas from hydrocracking to aromatics (HCKO, HCK1 and HCK2) Table 12 shows the balances obtained by reforming the C5-200 ° C naphtha in maximum flavor mode The results obtained from Examples 3 (HCK 1) and 4 (HCK 2) compared to the maximum mode Example 2 middle distillates (HCKO). The typical conditions used in reforming are quite mild compared to petroleum naphthas much less rich in naphthenes: 450 to 460 ° C for a level of RON equivalent to 104 (Maxi aromatic mode), a molar ratio H2 / HC of 4 and a Weight Spatial Velocity (PPH) of 2.5h-1.

It is thus seen that in the gasoline maximum modes (HCK1 and HCK2), and considering that C9 and C7 can be converted to C8 via the aromatic complex chain, it is possible to obtain very high C6 - C8 aromatics relative to initial char on dry basis, up to about 15% in one step mode and up to about 27.5% in two step mode.

Table 12: Yields of reforming of naphtha C5-200 ° C ex hydrocracking Yield / coal case HCK 1 step case HCK 1 step case HCK 2 `Yom / m (HCKO) mode (HCK1) steps (HCK2) maxi distillates max mode mode maxi means gasoline C5 + 11.22 22.09 40.67 H2 0.47 1.03 1.91 Cl + C2 0.04 0.08 0.33 C3 + C4 0.09 0.17 0.72 Benzene (C6) 0.88 2.05 3.76 Toluene (C7) 2.29 5.36 9.80 Aromatic in C8 (1) 2.01 4.43 8.08 Aromatic in C9 2.39 3.99 7.48 C10 aromatics 1.77 2.28 4.06 Total aromatics C6-C9 7.58 15.83 29.12 (1) Approximate distribution of aromatics in C8: 30% ethyl-benzene, 35% meta xylene, 15% para-xylene and 20% ortho-xylene

Claims (15)

REVENDICATIONS1. A process for converting coal into aromatic compounds comprising the following steps: a) a step of liquefying the coal in the presence of hydrogen, b) a step of separating the effluent obtained at the end of step a) into a light hydrocarbon fraction containing boiling compounds at not more than 500 ° C and a residual fraction; c) a hydrocracking step in the presence of hydrogen of at least a part of the so-called light hydrocarbon fraction obtained at from step b) in at least one reactor containing a fixed-bed hydrocracking catalyst, the conversion of the fraction 200 ° C + in the hydrocracking step being greater than 30% wt, d) separation of the effluent obtained at the end of step c) in at least one fraction containing naphtha and a heavier fraction than the naphtha fraction, e) a step of catalytic reforming of the fraction containing naphtha making it possible to obtain hydrogen and a reformate containing com aromatic substances, f) a step of separating the aromatics from the reformate. 20 2. Method according to the preceding claim wherein the liquefaction step a) in the presence of hydrogen is carried out in the presence of a catalyst supported bubbling bed, in the presence of a catalyst dispersed in bed leads or without added catalyst. 25 3. Method according to one of the preceding claims wherein the liquefaction step a) is carried out in at least two reactors arranged in series each containing a catalyst supported bubbling bed. 30 4. Method according to the preceding claim wherein the liquefaction step a) operates at a temperature between 300 ° C and 440 ° C for the first reactor and a temperature between 350 ° C and 470 ° C for the second reactor, then at a pressure of between 15 and 25 MPa, at an hourly mass velocity ((t of feed / h) / t of catalyst) of between 0.1 and 5 h -1 and at a hydrogen / charge ratio of between 0, 1 and 5 Nm3 / kg in each reactor. 5. Method according to one of the preceding claims wherein the hydrocracking step c) operates at a temperature between 250 and 480 ° C, at a pressure between 2 and 25 MPa, at the space velocity between 0.1 and 20 h-1 and the amount of hydrogen introduced is such that the volume ratio hydrogen on hydrocarbons is between 80 and 5000 Nm3 / m3. 6. Method according to one of the preceding claims wherein the hydrocracking catalyst of step c) comprises a zeolite. 7. Method according to one of the preceding claims wherein the catalytic reforming step operates at a pressure of 0.1 to 4 MPa, at a temperature between 400 and 700 ° C, at a space velocity of 0.1 to 10 h-1 and with a ratio of recycled hydrogen / hydrocarbons (mol.) of 0.1 to 10. 8. Process according to one of the preceding claims wherein the hydrogen produced in the catalytic reforming step e) is recycled in the liquefaction step a) and / or in the hydrocracking step c). 9. Method according to one of the preceding claims wherein the separation step b) provides a gaseous phase, at least a fractiondistillat atmospheric containing naphtha, kerosene and / or diesel, a vacuum distillate fraction and a residue fraction under vacuum. 10. Method according to the preceding claim wherein at least a portion and preferably all of the atmospheric distillate fraction, optionally supplemented with at least a portion of the vacuum distillate fraction and / or other co-charges, is sent to the hydrocracking step c) and at least a part and preferably all of the vacuum distillate fraction is recycled as a solvent in the liquefaction step a). 11. Process according to the preceding claims wherein the fraction containing naphtha from step d) is separated into a light naphtha fraction and a heavy naphtha fraction, the light naphtha fraction is at least partly subjected to an isomerization process. the heavy naphtha fraction is at least partly subjected to the catalytic reforming step e). 12.Procédé according to the preceding claims wherein the heavier fraction heavier than the naphtha fraction obtained in step d) is at least partly recycled in the hydrocracking step c). 13.Procédé according to the preceding claims wherein the heavier fraction heavier than the naphtha fraction obtained in step d) is at least partly sent to a steam cracker to obtain light olefinic compounds. 14. Process according to the preceding claims, in which the step of separating the aromatic compounds from the reformate is carried out by liquid-liquid extraction, extractive distillation, adsorption and / or crystallization. 25 30 15. Method according to one of the preceding claims wherein said coal is co-treated with a filler selected from petroleum residues, vacuum distillates of petroleum origin, crude oils, synthetic crudes, crude oils with head, deasphalted oils, deasphalting resins, asphalts or deasphalting pitches, derivatives of petroleum conversion processes, aromatic extracts from lubricant base production lines, oil sands or derivatives thereof, oil shales or their derivatives , hydrocarbon waste and / or industrial polymers, organic waste or household plastics, vegetable or animal oils and fats, tars and residues that are not or difficult to recover from gasification and / or Fischer-Tropsch synthesis of biomass, coal or petroleum residues, lignocellulosic biomass or one or more constituents cellulosic biomass selected from the group consisting of cellulose, hemicellulose and / or lignin, algae, charcoal, pyrolysis oil of lignocellulosic biomass or algae, pyrolytic lignin, products of hydrothermal conversion of lignocellulosic biomass or algae, activated sludge from water treatment plants, or mixtures of such loads.