FR2983862A1 - 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 biomass. More specifically, the present invention relates to a biomass conversion process comprising at least one liquefaction stage, 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 primarily marked by a desire to reduce dependence on fossil-based raw materials and then by the importance of issues related to global warming and the emission of greenhouse gases. In this context, the search for new loads from non-fossil sources is an issue of increasing importance. Among these fillers, mention may be made, for example, of biomass. Given the abundant and renewable reserves of biomass, an attractive alternative for the production of petrochemical intermediates is the liquefaction of biomass. It is known to produce fuel bases by direct liquefaction. Thus, the application EP 2 348 091 describes a method of direct liquefaction of biomass in two successive stages using boiling bed reactors followed by a hydrocracking step.

The present invention aims to produce BTX-type aromatic compounds from biomass using a process comprising a liquefaction step, followed by a hydrocracking and catalytic reforming step. Biomass in this text refers to lignocellulosic biomass, one or more components of lignocellulosic biomass selected from the group consisting of cellulose, hemicellulose and / or lignin, and / or algae. More particularly, the present invention relates to a process for converting biomass into aromatic compounds comprising the following steps: a) a step of liquefying the biomass in the presence of hydrogen, b) a step of separating the effluent obtained in the presence of hydrogen, from step a) to 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 portion of the so-called light hydrocarbon fraction obtained at the end of step b) in at least one reactor containing a fixed-bed hydrocracking catalyst, the conversion of the 200 ° C + fraction in the hydrocracking step being superior at 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 fraction naphtha, e) a reformation step catalytic fraction of the naphtha-containing fraction for obtaining hydrogen and a reformate containing aromatic compounds; f) a step of separating the aromatic compounds from the reformate. The research work carried out by the applicant on the liquefaction of biomass led him to discover that, surprisingly, this biomass liquefaction process followed by a fixed-bed hydrocracking step made it possible to obtain, after separation, a naphtha fraction particularly well suited 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 of sulfur and less than 0.5 ppm nitrogen) so that it can be sent directly to 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 of hydrogen produced during reforming. .

DETAILED DESCRIPTION The description will be made with reference to FIGS. 1 and 2 without the figures limiting the interpretation. The feedstock The biomass feedstock is selected from algae, lignocellulosic biomass and / or one or more lignocellulosic biomass constituents selected from the group consisting of cellulose, hemicellulose and / or lignin. Lignocellulosic biomass consists essentially of three natural polymers: cellulose, hemicellulose and lignin. It generally contains impurities (sulfur, nitrogen, etc.) and inorganic compounds of various natures (alkali, transition metals, halogen, etc.). Lignocellulosic biomass can consist of wood or plant waste. Other non-limiting examples of lignocellulosic biomass material are farm residues (straw ...), logging residues (first thinning products), logging products, dedicated crops (coppice short rotation), residues from the food industry, household organic waste, waste from wood processing facilities, used timber from construction, paper, recycled or not. Lignocellulosic biomass can also come from by-products of the paper industry such as Kraft lignin, or black liquors from pulp manufacturing. The algae that can be used in liquefaction are macroalgae and / or microalgae. Thus, the charge may consist of prokaryotic organisms such as blue-green algae or cyanobacteria, or eurokaryotic organisms such as groups with unicellular species (Euglenophytes, Cryptophytes, Haptophytes, Glaucophytes, etc.), groups with unicellular species or multicellular like red algae or Rhodophyta, and Stramenopiles including diatoms and brown algae or Phéophycées. Finally the load can also consist of macroalgae such as green algae (causing green tides), kelp or seaweed (also called kelp). The liquefaction technology also makes it possible to convert the biomass into co-treatment with other fillers. The biomass may be co-processed with a feedstock selected from petroleum residues, petroleum vacuum distillates, crude oils, synthetic crudes, toasted crudes, deasphalted oils, deasphalting resins, asphalts, and the like. 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, the term "hydrocarbonaceous feedstock" is intended to include feedstocks containing at least 50% by weight of product distilling above 250.degree. C. and at least 25% by weight distilling above 350.degree. Biomass can 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 hardly recoverable from gasification and / or FischerTropsch synthesis of biomass, coal or petroleum residues, coal, charcoal, lignocellulosic biomass pyrolysis oil or algae, pyrolytic lignin, products of the hydrothermal conversion of lignocellulosic biomass or algae, activated sludge from water treatment plants, or mixtures of such fillers. Pretreatment of the Charge Before liquefaction, the biomass may undergo one or more pretreatment steps. Preferably, the pretreatment comprises a step of partially reducing the water content (or drying), followed by a particle size reduction step to reach the size range suitable for the constitution of the biomass suspension. solvent for treatment in liquefaction reactors. The drying step is carried out at a temperature below 250 ° C, preferably below 200 ° C, preferably for 15 to 200 minutes. The biomass is then sent to a mill which makes it possible to reach the desired particle size. Other pretreatments, adapted to the nature of the feed, can complete or replace the drying and grinding stages, in particular roasting in the case of lignocellulosic biomass or demineralization in the case of algae, these technologies are widely described. in the litterature. In the case of lignocellulosic biomass or one of its constituents, the roasting step is carried out at a temperature of between 200 ° C. and 300 ° C., preferably between 225 ° C. and 275 ° C., absence of air preferably for 15 to 120 minutes, allowing to reach a water content of the biomass to be treated of about 1 to 10%, preferably between 1 and 5%. This step can replace the drying step. The grinding stage is considerably facilitated by the roasting stage which makes it possible to reduce the energy consumption compared to grinding without prior roasting. The pretreatment of the lignocellulosic biomass preferably comprises a roasting treatment. In the case of liquefaction of lignin alone, the roasting step is not necessary. After the pretreatment, biomass particles having a moisture content of from 1 to 50% by weight, preferably from 1 to 35% and more preferably from 1 to 10%, and a particle size of less than 600 are obtained. pm, preferably less than 150 pm. Liquefaction (step a) The biomass, after the possible pretreatment steps described above, is mixed with a solvent, preferably a hydrogen-donor solvent. The biomass / solvent mixture is a suspension of biomass particles dispersed in said solvent which is subsequently sent to the liquefaction stage. To form the suspension, the particle size of the biomass 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 / biomass 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 this region. last, then partial solubilization of the primary products of conversion and transfer of hydrogen to these primary products to allow conversion to 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 totally of a liquid co-charge, such as, for example, 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 may 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 stage is carried out in the presence of a catalyst supported in a bubbling bed, preferably in at least two reactors arranged in series containing a catalyst supported as a bubbling bed.

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 375 ° C for the first reactor and between 350 ° C and 470 ° C, preferably between 350 and 425 ° C for the second. The hourly mass velocity (t / h) / t catalyst is 0.1 to 1-1-1 and the amount of hydrogen mixed with the feed is usually about 0.1 to 2 normal cubic meters (Nm3) per kg of feed, preferably from about 0.1 to 0.5 Nm3 / kg in each reactor. After the first step, the conversion of the charge 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 biomass is then all that is not insoluble in THF. After the first step, the deoxygenation of the charge is between 10 and 100%, preferably between 50 and 99%. 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 the flow of at least a portion of the effluent from the first stage. without pumping is necessary. Optionally, the effluent obtained at the end of the first liquefaction stage is separated from 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 transferring in full or in part 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 Patent 4,816,841. The catalysts used in 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% by weight of molybdenum (expressed as molybdenum oxide MoO 3), on a support. This catalyst may also contain phosphorus (generally less than 20% by weight and most often less than 10% by weight, expressed as phosphorus oxide P2O5).

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). 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. Referring to Figure 1 which describes the process according to the invention by liquefaction in two successive bubbling bed reactors, the biomass (10), preferably pretreated beforehand, is ground in the mill (12) to produce particles. sized to form a suspension and be more reactive under liquefaction conditions. The biomass is then contacted with the recycling solvent (15) from the process in the furnace (14) to form the slurry. If necessary, although rarely necessary, a sulfur compound for maintaining the catalyst metals in sulphide form can 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 Bubbling bed (20) operating at an upward flow 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. 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 one. 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 portion of the reactor and fed back to 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 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 stage 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 and supported catalyst (s), 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 firstly 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 a part of the vacuum distillate fraction and / or other co-charges, is sent to 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 scheme and is illustrated in Figure 1.

Referring to this figure, the effluent treated in the liquefaction reactor (30) is sent via the line (38) into a high temperature high pressure separator (HPHT) (40) from which a gaseous fraction (41) is recovered. ) and a liquid fraction (44). 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) 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 line (15) to the vessel (14) to be mixed with the biomass. 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 schema and is illustrated in Figure 2.

Referring to this figure, the effluent treated in the second liquefaction reactor (30) is sent via line (38) into a high temperature high pressure separator (HPHT) (40) from which the so-called light fraction is recovered. (41) and the residual fraction (44). The light fraction (41) is sent directly, optionally mixed with the vapor phase (71) from the optional interstage 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). 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 biomass. 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 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 fairly severe hydrocracking in order to obtain a high yield in naphtha (and ultimately in aromatic compounds and hydrogen). at the same time as a very deep hydrotreatment making it possible to obtain a sufficiently pure naphthenic cut in terms of impurities so as not to poison the catalysts of the catalytic reforming. Hydrocracking is understood to mean hydrocracking reactions accompanied by hydrotreating reactions (hydrodenitrogenation, hydrodesulfurization), hydroisomerization, aromatic hydrogenation and opening of the 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, preferably between 0.2 and 11-1, 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 metals of group VIB of the periodic table of elements, or by a combination of at least one metal of group VIB of the periodic table and at least one metal of group VIII. 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, 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 molybdenum oxide) can be used. Mo03) 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, an alumina support is used and preferably rt or 7 alumina. 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 (on the oxide base) and usually less than 10%. When boron trioxide (B203) is present, its concentration is less than 10% by weight. Other conventional catalysts include zeolite Y of FAU structural type, an amorphous refractory oxide support (usually alumina) 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 hydrocracking naphtha, and ultimately in 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 (150) separator 20 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 the 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 for sending it to catalytic reforming as a result .

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), an atmospheric distillation and optionally a distillation under vacuum (not shown) which allows 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 the 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 fraction heavier 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 sent 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 biomass. 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 from 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 fed into catalytic reforming generally contains between 1 to 50 wt%, preferably between 5 and 30 wt% of paraffins, between 20 and 100 wt% of naphthenes, preferably between 0 and 90% and between 0 to 20% aromatics 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 reactions beneficial to aromatics formation and improvement of the octane number, the dehydrogenation of naphthenes, isomerization of cyclopentanic rings, isomerization of paraffins, dehydrocyclization of paraffins, and for the adverse reactions, hydrogenolysis and hydrocracking of paraffins and naphthenes. Similarly, 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 bed of catalyst, traversed by a charge composed of hydrocarbons and hydrogen, flows heated up between each reactor. Other process schemes use fixed bed reactors. The continuous process for the 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 process is for example described in the application FR 2 801604 or FR2946660.

The treatment of the feedstock in the reforming reactor (s) is generally carried out under a pressure of 0.1 to 4 MPa and preferably of 0.3 to 1.5 MPa at a temperature of between 400 and 700 ° C. C and preferably between 430 and 550 ° C, at a space velocity of 0.1 to 1-1-1 and preferably from 1 to 4 11-1 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 ( for example tin or rhenium), at least one halogen and optionally one or more additional elements (such as alkaline, alkaline earth, lanthanide, silicon, group IV B elements, non-noble metals, group III A elements, etc.). These catalysts are widely described in the literature.

The reforming makes it possible to obtain 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 N-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 biomass.

Claims (15)

REVENDICATIONS1. Process for converting biomass into aromatic compounds comprising the following steps: a) a step of liquefying the biomass in the presence of hydrogen, b) a step of separating the effluent obtained at the end of step a) into a light fraction of hydrocarbons 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 hydrocarbon fraction obtained at from step b) in at least one reactor containing a fixed bed hydrocracking catalyst, the conversion of the 200 ° C + fraction in the hydrocracking step being greater than 30 wt%, d) separation of the hydrocracking catalyst. effluent obtained at the end of step c) in at least one fraction containing naphtha and a fraction heavier 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 s aromatic compounds, f) a step of separating aromatic compounds from the reformate. 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. 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. 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. Method 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. 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. 14. Method according to one of the preceding claims wherein biomass is selected from algae, lignocellulosic biomass and / or one or more lignocellulosic biomass constituents selected from the group consisting of cellulose, hemicellulose and / or lignin. 15.Procédé according to one of the preceding claims wherein said biomass is co-treated with a filler selected from petroleum residues, vacuum distillates of petroleum origin, crude oils, synthetic crudes, crude oils 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 oil residues, coal, charcoal, bioma pyrolysis oil lignocellulosic or algae ss, pyrolytic lignin, products of hydrothermal conversion of lignocellulosic biomass or algae, activated sludge from water treatment plants, or mixtures of such fillers.