FR3012467A1 - OPTIMIZED BIOMASS CONVERSION PROCESS FOR THE PRODUCTION OF PETROCHEMICAL BASES - Google Patents

Abstract

The present invention relates to a process for producing petrochemical bases, in particular olefinic and aromatic compounds, from a liquid feedstock comprising at least one bio-oil, said feed being introduced into at least one hydroreforming stage in the presence of hydrogen and a hydroreforming catalyst comprising at least one transition metal chosen from the elements of groups 3 to 12 of the periodic table and at least one support chosen from active carbons, silicon carbides, silicas, transition aluminas, alumina silicas, zirconium oxide, cerium oxide, titanium oxide and transition metal aluminates, taken alone or as a mixture, to obtain at least one liquid effluent comprising at least one aqueous phase and at least one organic phase, a step of hydrotreating at least a part of the organic phase of the effluent resulting from the first hydroreforming stage, a three i th step of hydrocracking at least a portion of the effluent from the second hydrotreating step, wherein at least a portion of the fraction boiling at a temperature above 160 ° C separated at the end of said third hydrocracking step is recycled in said third hydrocracking step, a separation of the effluent obtained at the end of the third hydrocracking step in at least one fraction containing naphtha and a fraction boiling at a temperature above 160 ° C, a fourth steam cracking step of the fraction containing naphtha from the separation to obtain hydrogen, olefinic compounds having a carbon number of between 2 and 6 and aromatic compounds.

Description

FIELD OF THE INVENTION The present invention relates to a process for the production of petrochemical bases and preferably of olefinic compounds having a carbon number of between 2 and 6, in particular in accordance with the present invention. ethylene, propylene and butadiene, and aromatic compounds from liquids from biomass and in particular from liquefied biomass from the pyrolysis of biomass, also called bio-oil. More specifically, the invention relates to a process for converting bio-oil into an effluent comprising at least olefinic compounds having a carbon number of between 2 and 6 comprising a hydroreforming step, followed by a fixed-bed hydrocracking step of a severity adapted to maximize the production of a naphtha fraction and a steam-cracking step of said naphtha fraction to produce said olefinic compounds having a carbon number of between 2 and 6 and said aromatic compounds (BTX), mixed with hydrogen. Said aromatic compounds are preferably aromatic C 6 -C 8 compounds, and preferably benzene, toluene and xylenes / ethylbenzene, called BTX. State of the art The current international context is primarily characterized by a desire to reduce dependence on fossil-based raw materials and then by the importance of global warming issues and the emission of greenhouse gases. Greenhouse. 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 use of liquids from biomass, usually also called bio-oil or biobrud. Bio-oils are liquid products obtained by thermochemical liquefaction of cellulosic biomass. The method for thermochemical liquefaction of biomass can for example be carried out in the presence or absence of catalyst, in the presence of an inert gas or a gas containing hydrogen and water. Thermochemical processes generally convert biomass into liquid, gaseous and solid products. Among these methods, so-called rapid pyrolysis methods tend to maximize the liquid yield. During rapid pyrolysis, the temperature of the biomass, possibly divided, is rapidly raised to values of greater than about 300 ° C. and preferably of between 450 and 600 ° C. and preferably of between 450 and 550 ° C. Liquid products are condensed in the form of bio-oil. Ringer et al. (Large-Scale Pyrolysis Oit Production: A Technology Assessment and Economic Analysis, M. Ringer, V. Putsche, and J. Scahill, NREL Technical Report NREL / TP-510-37779, November 2006) studied the different technologies used for rapid pyrolysis of biomass on a large scale. They include bubbling fluid beds, circulating fluidized beds, ablative pyrolysis reactors, vacuum pyrolysis and rotary cone pyrolysis.

Bio-oils and in particular bio-oils resulting from the rapid pyrolysis of biomass are complex mixtures of highly oxygenated polar hydrocarbon products obtained from the breaking of biopolymers, and having physical and chemical properties limiting their direct use. Pyrolysis bio-oils have undesirable properties such as: (1) corrosivity due to their high content of water and organic acids; (2) a low specific calorific value due to the high oxygen content, typically of the order of 40% by weight; (3) chemical instability, both storage and thermal, due to the abundance of reactive oxygen functional groups such as the carbonyl group, which can lead to polymerization during storage and subsequently phase separation; (4) a viscosity and a tendency to separate relatively high phases under high shear conditions, in an injector for example; (5) solid particles of coal resulting from pyrolysis, which will always be present in unfiltered bio-oils in a small quantity such that the bio-oil / coal particles ratio is greater than 10: 1 and preferably greater than 50 1. However, said particles can cause plugging of stationary internal combustion engines or poison post-treatment catalysts.

All these aspects combine to make the handling, transportation, storage and implementation of bio-oils difficult and expensive, thus making their integration into current heat and energy systems and technologies very problematic. Over the last twenty years, the approach of direct hydrotreating of bio-oil to stable hydrocarbons or oxygenated liquids has been intensively studied. Elliott has published a detailed study of these many historical developments, including work with model compounds known to be present in bio-oils (Historical Developments in Hydroprocessing Bioils, D.CI Elliott, Energy & Fuels 2007, 21, 17921815). The present invention aims to produce petrochemical bases and preferably olefinic compounds having a carbon number of between 2 and 6, in particular ethylene, propylene and butadiene, and aromatic compounds (BTX), from a liquid feed comprising at least one bio-oil using a process comprising a hydroreforming step, followed by a hydrotreatment step making it possible to obtain a totally deoxygenated organic liquid effluent, and then a step of hydrocracking method for maximizing the production of a naphtha fraction followed by a steam-cracking step of said naphtha fraction obtained to produce said olefinic compounds having a carbon number of between 2 and 6 and said aromatic compounds (BTX) in mix with hydrogen.

SUMMARY OF THE INVENTION The present invention relates to a process for producing olefinic compounds having a carbon number of from 2 to 6, in particular olefinic compounds comprising at least ethylene, propylene and butadiene, and of aromatic compounds (BTX), from a liquid feed comprising at least one bio-oil, said process comprising at least the following steps: the introduction of said feedstock in a first hydroreforming step in the presence of hydrogen and a hydroreforming catalyst comprising at least one transition metal chosen from the elements of groups 3 to 12 of the periodic table and at least one support chosen from active carbons, silicon carbides, silicas and transition aluminas , alumina silicas, zirconium oxide, cerium oxide, titanium oxide and transition metal aluminates, taken alone or as a mixture, to obtain at least one liquid effluent comprising at least one aqueous phase and at least one organic phase, said first hydroreforming stage operating at a temperature of between 200 and 380 ° C., at an absolute pressure of between 3 and 30 MPa, at an hourly space velocity comprised between between 0.1 and 5 h -1, the treatment in a second hydrotreatment step of at least a portion of the organic phase of the effluent from the first hydroreforming step in the presence of hydrogen in at least one reactor containing a hydrotreating catalyst, operating at a temperature between 250 and 400 ° C, at a pressure of between 2 MPa and 25 MPa, at a space velocity of between 0.1 1-1-1 and 20 1-1 -1 and with a total amount of hydrogen mixed with the feed such that the volume ratio hydrogen / hydrocarbon is between 100 and 3000 Nm3 / m3, the treatment in a third hydrocracking step of at least a portion of the effluent from the second me hydrotreatment step, in the presence of hydrogen, in at least one reactor containing a hydrocracking catalyst, operating at a temperature between 250 and 480 ° C, at a pressure between 2 and 25 MPa, at a space velocity included between 0.1 and 20 h-1, and with a quantity of hydrogen introduced such that the volume ratio hydrogen on hydrocarbons is between 80 and 5000 Nm3 / m3, in which at least a part of the fraction having a dot of initial boiling between 120 ° C and 200 ° C, separated at the end of said third hydrocracking step is recycled in said third hydrocracking step, a separation of the effluent obtained at the end of the third step d hydrocracking into at least one fraction containing naphtha and a fraction having an initial boiling point of between 120 ° C and 200 ° C, the introduction of at least a portion of the naphtha-containing fraction from step of separation in a fourth th steam-cracking step for producing an effluent comprising at least olefinic compounds having 2 to 6 carbon atoms, aromatic compounds (BTX) and hydrogen, said steam-cracking step operating at a temperature of between 700 and 900 ° C. in the presence of water vapor injected in a weight ratio relative to the fraction containing the naphtha entering said step between 0.2 and 2. An advantage of the present invention is to provide a process, from a load liquid comprising at least one bio-oil, comprising a hydroreforming step followed by a hydrotreatment step and then hydrocracking allowed to obtain, after separation, a naphtha fraction particularly well adapted to its use in a steam cracking step for the production of olefinic and aromatic compounds (BTX) because of its very small amount of impurities. The hydroreforming step makes it possible firstly to obtain a liquid effluent comprising at least one aqueous phase and at least one organic phase still containing significant amounts of impurities: heteroelements in sulfur, nitrogen and oxygen as well as in olefins. and polyaromatic. Before sending the light fraction (naphtha) of the said organic phase into the steam-cracking stage, it is thus necessary to carry out a hydrotreatment stage allowing the removal of nitrogenous, oxygenated and sulfur-containing impurities and then a hydrocracking stage. severe with recycling the fraction distilling above 160 ° C to maximize said naphtha. This hydrocracking step thus makes it possible to obtain a substantial naphtha fraction by cracking. Similarly, the severity of the hydrocracking step makes it possible to increase the yield of naphtha fraction and thus ultimately to increase the yield of olefinic compounds and aromatics produced in the steam-cracking step and preferably in olefinic compounds comprising at least ethylene, propylene and butadiene and aromatic compounds (BTX).

Detailed description of the invention. According to the invention, the filler comprising at least one bio-oil treated in the process according to the invention is a liquid filler. The bio-oil of the filler treated in the process according to the invention may advantageously be a pretreated or non-pretreated biofil oil. Preferably, the bio-oil of the filler treated in the process according to the invention is not pretreated, that is to say that said bio-oil has not undergone any pretreatment step other than pyrolysis. In particular, said bio-oil has not undergone any pretreatment step intended to reduce its metal content, its solids content, its water content or its acid content.

Throughout the rest of the text, the term "liquid filler" is understood to mean a filler comprising less than 2% by weight and preferably less than 1% by weight relative to the bio-oil, of solid particles dispersed in a liquid. This corresponds to a bio-oil / solid particles ratio of greater than 50: 1 and preferably greater than 100: 1. Said charge is therefore not a suspension. A bio-oil is a complex mixture of oxygenates, obtained from breaking biopolymers present in the original biomass. In the case of lignocellulosic biomass, the structures from these three main components, cellulose, hemicellulose and lignin, are well represented by the components of the bio-oil. In particular, a bio-oil is a highly oxygenated polar hydrocarbon product which generally comprises at least 10% by weight of oxygen, preferably from 10 to 60% by weight of oxygen relative to the total weight of said bio-oil. In general, the oxygenated compounds are alcohols, aldehydes, esters, ethers, organic acids and aromatic oxygen compounds. In the case where the bio-oil is derived from the rapid pyrolysis, part of the oxygen is present in the form of free water representing at least 5% by weight, preferably at least 10% by weight and preferably at least 20% by weight. % weight of the bio-oil. These properties make the bio-oil totally immiscible with hydrocarbons, even with aromatic hydrocarbons that typically contain very little or no oxygen. The bio-oils are preferably derived from biomass preferably selected from plants, herbs, trees, wood chips, seeds, fibers, seed husks, aquatic plants, hay and other sources of lignocellulosic materials, such as those from municipal waste, agro-food waste, forest waste, slaughter residues, agricultural and industrial waste (such as sugar cane bagasse, waste from oil palm cultivation, sawdust or straws). Bio-oils can also come from paper pulp and by-products of recycled or non-recycled paper or by-products from paper mills. The bio-oils are advantageously obtained by thermochemical liquefaction of the biomass, preferably by pyrolysis, and preferably by rapid, slow pyrolysis with or without a catalyst (in the presence of a catalyst, this is known as catalytic pyrolysis). Pyrolysis is a thermal decomposition in the absence of oxygen, with thermal cracking of the gas, liquid and solid feedstocks. A catalyst may advantageously be added to enhance the conversion during so-called catalytic pyrolysis. Catalytic pyrolysis generally gives a bio-oil having a lower oxygen content than the bio-oil obtained by thermal decomposition, but the yield of bio-oil is generally lower. The selectivity between the gas, the liquid and the solid is related to the reaction temperature and the residence time of the vapor. Biomass pyrolysis processes, including rapid pyrolysis, are well described in the literature such as, for example, A.V. Bridgewater, H. Hofbauer and S. van Loo, Thermal Biomass Conversion, CPL Press, 2009, 37-78. The slow pyrolysis is advantageously carried out at a temperature of between 350 and 450 ° C. and preferably of the order of 400 ° C. and with a residence time of a few minutes to a few hours which favors the production of a solid product. also called carbonization product or charcoal. In particular, slow pyrolysis generally allows the production of 35% by weight of gas, 30% by weight of liquid and 35% by weight of char. Gasification processes operating at very high temperatures, preferably above 800 ° C, promote the production of gas, and in particular more than 85 g / wt of gas. The intermediate pyrolysis is advantageously carried out at a temperature generally of between 450 and 550 ° C. and with a residence time of the short vapor, preferably between 10 and 20 seconds, which favors the liquid yield. In particular, the intermediate pyrolysis generally allows the production of 30% gas, 50% liquid and 20% char.

The rapid pyrolysis is advantageously carried out at a temperature generally of between 450 and 550 ° C. and with a residence time of the very short vapor, preferably between 0.5 and 2 seconds, which maximizes the liquid yield: in particular, the Fast pyrolysis generally allows the production of 10-20% weight of gas, 60-75% weight of liquid and 10-20% weight of char. The highest liquid yields are therefore obtained in the context of rapid wood pyrolysis processes, with a liquid yield of up to 75% by weight. Part of the oxygen is present in the form of free water representing at least 5% by weight, preferably at least 10% by weight and preferably at least 20% by weight of the bio-oil. Preferably, the bio-oils used in the process according to the present invention are obtained by rapid pyrolysis of the biomass. The bio-oils may also advantageously be obtained by thermochemical liquefaction of the biomass in the broad sense, for example with or without hydrogen, with or without a catalyst or with or without water, such as, for example, hydropyrolysis, catalytic pyrolysis or hydrothermal conversion of biomass.

The treated liquid feedstock in the process according to the invention preferably comprises a biofilter content of between 10 and 100% by weight relative to the total weight of said feedstock. The liquid feedstock comprising at least one bio-oil used in the process according to the invention may also advantageously contain other liquid feeds derived from biomass, said liquid feeds being advantageously chosen from vegetable oils, algae or algal oils. , fish oils, fats of vegetable or animal origin, and alcohols derived from the fermentation of the sugars of biomass, or mixtures of such fillers, whether pre-treated or not. Vegetable oils or oils derived from animal fats essentially comprise chemical structures of the triglyceride type which the skilled person also knows under the name tri-ester of fatty acids as well as free fatty acids, the fatty chains of which contain a number of carbon atoms between 9 and 25. The said vegetable oils may advantageously be crude or refined, wholly or in part, and derived from plants selected from rapeseed, sunflower, soybean, palm, olive, coconut, copra, castor oil, cotton, peanut, flax and crambe oils and all oils derived for example from sunflower or rapeseed by genetic modification or hybridization, this list not being limiting. Said animal fats are advantageously chosen from fats composed of residues from the food industry or from the catering industries. Fried oils, various animal oils such as fish oils, tallow, lard can also be used. Said liquid feed comprising at least one bio-oil may also advantageously be treated in admixture with at least one hydrocarbon liquid feed derived from petroleum and / or coal. The hydrocarbon feedstock derived from petroleum can advantageously be selected from direct distillation distillates under vacuum, vacuum distillates from a conversion process, such as those resulting from coking processes, fixed bed hydroconversion or hydrotreating of ebullated bed heavy fractions, products from fluid catalytic cracking units, such as, for example, light catalytic cracked gasoil (LCO) of different origins, heavy catalytic cracked gasoil (HCO) of different origins and any fluid catalytic cracked distillate fraction generally having a distillation range of about 150 ° C to about 370 ° C, the aromatic extracts and paraffins obtained in the preparation of lubricating oils and the solvent deasphalted oils, alone or in mixed. The hydrocarbon feedstock derived from coal may advantageously be chosen from products resulting from the liquefaction of coal and aromatic fractions originating from the pyrolysis or gasification of coal and shale oils or products derived from shale oils, alone or mixed, pretreated or not. Since the hydrogen is produced internally during the step of hydroreforming the liquid feed comprising at least one bio-oil, the hydrogen consumption for the co-treatment mixed with the hydrocarbon feedstocks derived from the oil and / or coal is reduced. Preferably, said liquid feed comprising at least one bio-oil may advantageously be a mixture of any of the abovementioned fillers, said feeds possibly being pretreated or not. In a preferred variant, said liquid feed treated in the process according to the invention consists entirely of bio-oil. Hydroreforming In accordance with the invention, said liquid feedstock comprising at least one bio-oil, optionally at least one liquid hydrocarbonaceous feedstock derived from petroleum and / or coal, optionally other liquid feedstocks derived from biomass, optionally mixed with at least a portion of the organic phase of the effluent from the hydroreforming step which is recycled is introduced into said first hydroreforming step.

The reaction is referred to herein as "hydroreforming" since it brings about a characteristic change in the molecular weight distribution of the feed comprising at least one bio-oil. Preferably, the hydroreforming step is a hydrotreatment / hydro-conversion step in which the liquid feedstock comprising at least the bio-oil and preferably in which the bio-oil is partially deoxygenated and partially hydroconverted. Said hydroreforming step is carried out with a partial production of hydrogen internal to the reaction. Preferably, said hydroreforming step is a partial deoxygenation and partial hydroconversion step. Partial deoxygenation and partial hydroconversion are understood to mean a reaction in which the oxygen content in the organic phase is reduced to a content of between 5 and 25% by weight and makes it possible to obtain less than 20% by weight of a fraction boiling at a temperature greater than 600 ° C. During said first hydroreforming step, the bio-oil is converted in order to stabilize the product, to make it miscible with the hydrocarbons, to cause an easy phase separation of the water in the bio-oil, d lowering its viscosity, lowering its corrosivity, significantly converting high molecular weight fractions to smaller molecules and lowering its oxygen content. The product of the hydroreforming step According to the invention, said first hydroreforming step makes it possible to obtain at least one liquid effluent comprising at least one aqueous phase and at least one organic phase. A gaseous phase is also advantageously obtained. Said gaseous phase containing mainly CO2, CO, and light C 1 -C 4 gases such as, for example, methane. Said gaseous phase may advantageously be separated from said aqueous and organic phases in any separator known to those skilled in the art.

Said aqueous phase and said organic phase are advantageously separated by decantation. The analysis of the liquid effluent shows no sign of polymerization. Said aqueous phase comprises water and may also comprise organic materials, mainly acetic acid and methanol (case of fast pyrolysis bio-oil), dissolved in said aqueous phase. The aqueous phase contains essentially water formed by hydrodeoxygenation and water contained in the unprocessed bio-oil which has not been converted and in general less than about 20% by weight of dissolved organic matter. For pyrolysis bio-oils, the aqueous phase separated in the first reforming step typically contains about 10% by weight of acetic acid and a lower proportion of methanol. Preferably, acetic acid and methanol are recovered in the aqueous phase because they constitute valuable by-products. Acetic acid can advantageously be recovered by various known means such as distillation or evaporation, crystallization in the form of salt, for example an alkaline earth, or by solvent extraction, by means of ion exchangers. liquids for example. Because of its low boiling point and the absence of azeotropic formation with water, the methanol is simply recovered by distillation. Preferably, acetic acid and methanol taken together comprise at least 80% by weight of the organic components in the aqueous phase.

The organic phase contains partially deoxygenated bio-oil and hydrocarbons. The term "partially deoxygenated bio-oil" means a bio-oil containing less than 25% by weight of oxygen and preferably less than 15% by weight of oxygen relative to the total weight of said organic phase. The organic phase therefore comprises less than 25% by weight of oxygen, preferably less than 15% by weight of oxygen, preferably less than 10% by weight of oxygen relative to the total weight of said organic phase. Thus, said organic phase is deoxygenated to a sufficient degree to make it miscible with hydrocarbon feeds such as those used in petroleum refining at high concentrations. In general, said organic phase can mix freely with most hydrocarbon products at concentrations up to at least 30% by weight. On the other hand, since the oxygen content is substantially reduced and the bio-oil is already partially converted, the hydrogen requirements for subsequent refinery stages are greatly minimized. In addition, said organic phase is stabilized so that it can be subsequently treated in a refinery without the risk of formation of solids by polymerization or polycondensation.

The organic phase preferably contains less than 25% by weight of oxygen and preferably less than 15% by weight of oxygen and less than 5% by weight of water, preferably less than 2% by weight of water relative to the mass. total of said organic phase. The hydrogen consumption during said first hydroreforming step is preferably less than about 2% by weight of the bio-oil. Preferably, the total acid number (TAN) is less than about 100 KOH / g oil. The total acid number is expressed in mg KOH / g of oil. This is the amount of potassium hydroxide in milligrams needed to neutralize acids in one gram of oil. There are standard methods for determining the acid number, such as ASTM D 974 and DIN 51558 (for mineral oils, biodiesels), or methods specific to biodiesels using European standard EN 14104 and ASTM D664, widely used worldwide.

The higher heating value (PCS) is greater than about 35 MJ / kg in the organic phase resulting from said first hydroreforming step. Typically, during the first hydroreforming step, the yields in the organic phase relative to the dry biomass, that is to say comprising 0% moisture, introduced into the pyrolysis stage, is between 15 and 45.degree. % by weight, and preferably between 15 and 40% by weight. The yields in aqueous phase relative to the dry biomass are between 10 and 50% by weight, and preferably between 10 and 40% by weight. The gas phase yields relative to the dry biomass are between 5 and 30% by weight. The conversion of the organic fraction of the bio-oil into partially deoxygenated bio-oil represents at least 70% by weight. According to a variant, at least a portion of the organic phase of the effluent from the first hydroreforming step is recycled in said first hydroreforming step. Preferably, the recycle ratio equal to the ratio of the mass flow rate of said organic phase to the mass flow rate of the non-pretreated bio-oil is between 0.05 and 10, preferably between 0.05 and 5, and very preferably between 0.05 and 1.

The recycle rates used make it possible to have a sufficiently thermally stable charge overall so that it can be introduced into the preheating furnace as well as into the hydrorefaction reactor without any risk of clogging, to facilitate the conversion and deoxygenation reactions of the biofuel. oil on the catalyst, to improve the dissolution of hydrogen gas in the recycled organic phase from the first hydroreforming step and to better manage the exothermicity generated by the reactions, while minimizing the recycling of all or part of the organic phase resulting from the hydroreforming in order to minimize the investment costs as well as the operating costs. The organic phase recycled in said first hydroreforming stage, and resulting from this stage, also serves to lower the density of the partially deoxygenated bio-oil produced during this first stage, which facilitates the phase clearance of the biofuel. partially deoxygenated oil and the aqueous phase at the hydroreforming outlet. This separation which occurs easily during the cooling of the reaction mixture is another advantage that provides the presence of the organic phase recycled in the reaction mixture of the first hydroreforming step. Another advantage, the presence of the organic phase recycled in the reaction mixture tends to expel water from said organic phase, the water content is thus reduced as much as possible, which is favorable to the use of said organic phase as fuel or as a load for subsequent processing steps. Thus, said phase said organic phase from the first hydroreforming step contains less than 5% by weight of water, preferably less than 2%. According to the invention, the catalyst used in the first hydroreforming step comprises a Group 10 metal, alone or in combination with at least one Group 3 to 12 metal and preferably a Group 10 metal, alone. or in combination with at least one group 5, 6, 7, 8, 9, 10, 11 metal of the Periodic Table. Very preferably, said first supported hydroreforming catalyst comprises nickel, alone or in combination with at least one metal selected from chromium (Cr), molybdenum (Mo), manganese (Mn), tungsten (W) , iron (Fe), cobalt (Co) and copper (Cu). According to the invention, the support of said hydroreforming catalyst is chosen from active carbons, silicon carbides, silicas, transition aluminas, alumina silicas, zirconium oxide, cerium oxide, silicon dioxide, titanium oxide and transition metal aluminates, taken alone or in admixture and preferably from silicas, transition aluminas, silica aluminas, porous carbon and transition metal aluminates. These supports can be taken alone or in mixture. Preferably, said hydroreforming catalyst is a catalyst selected from catalysts comprising nickel (Ni), nickel and chromium (NiCr) or nickel and manganese (NiMn) on porous carbon and catalysts comprising nickel and molybdenum (NiMo) on alumina or nickel aluminate.

Preferably, said hydroreforming catalyst comprises a metal content of group 10 and preferably nickel, of between 1 and 20% by weight, preferably between 5 and 15% by weight, expressed as a percentage by weight of metal oxide of the group Relative to the total mass of said catalyst. In the case where said Group 10 metal is used in combination with at least one Group 3 to 12 metal, said hydroreformed supported catalyst comprises a Group 3 to 12 metal content of from 1 to 20% by weight expressed as a percentage Group 3 to 12 metal oxide weight relative to the total mass of said catalyst. Preferably, said catalyst is prepared by conventional methods such as co-kneading or impregnation followed by one or more heat treatments. Said catalyst is advantageously used in the first hydroreforming step in reduced or sulphurized form. Operating Conditions According to the invention, said first hydroreforming stage operates at a temperature of between 200 and 380 ° C., preferably between 250 and 380 ° C., preferably between 270 and 360 ° C. and very preferably between 270 and 350 ° C at an absolute pressure of between 3 and 30 MPa, preferably between 3.5 and 25 MPa, and preferably between 5 and 20 MPa, at a space velocity of between 0.1 and 5 hours; 1, preferably between 0.2 and 5h-1 and preferably between 0.5 1-1-1 and 1-1-1 relative to the bio-oil, and with a quantity of hydrogen introduced such that that the volume ratio hydrogen on hydrocarbons is between 50 and 2000 Nm3 / m3 and preferably between 100 and 1000 Nm3 / m3. Preferably, the liquid feed comprising at least the bio-oil is not preheated, or only preheated to a temperature below 80 ° C before being introduced into the first hydroreforming step. Indeed, prolonged heating or storage at high temperature can cause deterioration. There are no limitations on the apparatus for carrying out the process. This can be conducted in batch or continuous mode. Nevertheless, for large-scale industrial applications, it will be preferable to operate continuously. The hydroreforming reaction can be carried out in any reactor which facilitates the efficient dispersion of the bio-oil in the reaction mixture. A simple continuous stirred reactor (CSTR) is suitable in batch mode. For continuous applications, downflow, moving bed, bubbling bed or slurry reactors will be more suitable. While keeping the catalyst in the reactor, the moving-bed and ebull-bed reactor technologies have the advantage of allowing easy catalyst replacement during continuous operations, which gives greater flexibility to the operation of the reactor. unit, increases the duty cycle and maintains an activity as a function of time almost constant. These technologies also accept charges containing a few solids or producing solids during reactions, which is not the case for fixed bed reactor technology, except for the use of reactive guards reactors, for example. The technology of the slurry reactor, according to the English terminology, has the same advantages as the moving bed and ebullated bed reactor technologies and the additional advantage of operating with a fresh catalyst, therefore with maximum activity, but the disadvantage to greatly complicate the recovery of the catalyst leaving the reactor with the effluents. Other reactors meeting the principles set forth herein are within the abilities of those skilled in the art of chemical reactor technology and are within the scope of the present invention. Hydrogen consumption During the first hydroreforming step, an internal reforming of the bio-oil into carbon oxides and hydrogen is carried out with a relatively high hydrocarbon yield and a low relative consumption of hydrogen. The consumption of hydrogen is therefore moderate during the first hydroreforming step and the internal production of hydrogen by hydroreforming makes it possible to operate the process with a very limited external hydrogen consumption.

Part of the hydrogen necessary for the reaction in the first hydroreforming stage is generated internally, a priori from the water present in the bio-oil, by hydroreforming a portion of the bio-oil, in such a way that a significant portion of the oxygen is released as carbon dioxide, and the net hydrogen requirement is much lower than in the case of the direct hydrodeoxygenation process. The other part of the hydrogen necessary for the reaction is of external origin. Typically, the hydrogen consumption during the first hydroreforming stage is less than about 2% by weight of the mass of the bio-oil introduced into the reactor, and the corresponding emission of CO2 to produce it is much less than in the case of processes of the hydrodeoxygenation (HDO) type, with a consumption of at least 4 or 5% by weight of external hydrogen (depending on the oxygen content of the feedstock).

Hydrogen of external origin can advantageously be supplied from fossil resources, by gasification / partial oxidation or by steam reforming, by vapor reforming of the bio-oil itself (although this implies a loss of efficiency carbon energy), by steam reforming methane produced and light gas fractions and / or fractions comprising light oxygen compounds from the first hydroreforming step or the subsequent hydrotreatment / hydrocracking step. It is also possible to react the CO recovered with the water produced to obtain hydrogen by means of the known reaction of gas to water conversion by "water gas shift" according to the English terminology (WGS) . Thus, the overall process, from the biomass to the final hydrocarbon product, could be self-sufficient in hydrogen. In the first hydroreforming step, a moderate amount of gaseous hydrocarbon, mainly methane, is formed. This methane fraction can advantageously be reformed using the so-called Steam Methane Reforming SMR process in order to produce bio-hydrogen for the hydroreforming or hydrotreatment and / or hydrocracking steps.

It is also possible to use the bio-coal from the pyrolysis step as feed for the gasification reactor to produce synthesis gas (CO + H2). This synthesis gas can also be used to produce bio-hydrogen for the hydroreforming or hydrotreatment and / or hydrocracking steps.

Thus, by implementing the SMR process and / or the gasification of bio-coal, the process of conversion and upgrading of the bio-oil can be self-sufficient in hydrogen, without requiring the addition of hydrogen of origin Fossil Hydrotreating step According to the invention, at least a portion and preferably all of the organic phase of the effluent from the first hydroreforming step is sent to said second hydrotreating step in the presence of hydrogen in at least one reactor containing a hydrotreating catalyst, operating at a temperature between 250 and 400 ° C, preferably between 320 and 380 ° C, at a pressure of between 2 MPa and 25 MPa, preferably between 5 MPa and 20 MPa and at a space velocity of between 0.1 hr-1 and 20 hr-1 and preferably between 0.5 hr-1 and 5 hr-1 and with a total amount of hydrogen mixed with the feedstock (y). including chemical consumption and quantity re cycled) such that the volume ratio hydrogen / hydrocarbon is between 100 and 3000 Nm3 / m3 and preferably between 100 and 1000 Nm3 / m3. Said hydrotreatment catalyst advantageously comprises an active phase comprising at least one metal of group 6 chosen from molybdenum and tungsten alone or as a mixture, preferably associated with at least one metal of groups 8 to 10, preferably chosen from nickel and cobalt, alone or in admixture and a support selected from the group consisting of alumina, silica, silica-alumina, magnesia, clays and mixtures of two or more of these minerals. Said support may further contain other compounds such as, for example, oxides selected from the group consisting of boron oxide, zirconia, titanium oxide, phosphoric anhydride. Preferably, said support is made of alumina, and more preferably alumina [eta], [theta], [delta] or [gamma]. The total amount of metal oxides of group 6 and groups 8 to 10 in the hydrotreatment catalyst used is advantageously between 5 and 40% by weight, preferably between 6 and 35% by weight relative to the total mass. of catalyst. In the case where said catalyst comprises nickel, the nickel oxide content is advantageously between 0.5 and 12 ° A> weight and preferably between 1 and 10% by weight relative to the total mass of catalyst. In the case where the catalyst comprises molybdenum, the content of molybdenum oxide is advantageously between 1 and 35% by weight of molybdenum oxide (MoO 3) and preferably between 5 and 30% by weight of molybdenum oxide, these percentages being expressed in% by weight relative to the total mass of catalyst. A preferred catalyst is advantageously chosen from NiMo, NiW or CoMo catalysts on an alumina support. Said catalyst used in the second hydrotreating step may also advantageously contain at least one doping element chosen from phosphorus and boron. Said element may advantageously be introduced into the matrix or, preferably, deposited on the support. It is also possible to deposit silicon on the support, alone or with phosphorus and / or boron. The oxide content in said element is advantageously less than 20%, preferably less than 10% relative to the total mass of catalyst.

The metals of the catalysts used in the second hydrotreatment step of the method according to the invention may be metals or sulphurized metal phases. For maximum efficiency, these metal oxide catalysts are usually converted at least partially to metal sulfides. The catalysts based on metal oxides can be sulphurated by any technique known in the prior art, for example in the reactor (in situ) or ex situ, by contacting the catalyst, at elevated temperature, with a source. of hydrogen sulfide, such as dimethyl disulfide (DMDS). Since biomass normally contains only a very small quantity of sulfur, the use of non-sulphurized catalysts should make it possible to avoid any risk of sulfur contamination in the fuels produced. Among the other catalysts based on metal oxides adapted to the hydrotreatment stage, mention may be made of the metal phases obtained by reduction in hydrogen. The reduction is generally carried out at temperatures from about 150 ° C to about 650 ° C, at a hydrogen pressure of from about 0.1 to about 25 MPa (14.53625 psi). Optional Separation A step of separating at least one organic phase, at least one aqueous phase and at least one gaseous phase may optionally be carried out between said second hydrotreating step and said third hydrocracking step. Said 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.

The organic phase resulting from the second and optionally separated hydrotreatment step is freed from nitrogen, oxygen and sulfur impurities. According to a variant of the invention, said optionally separated organic phase can also be fractionated in a fractionation zone into at least one light fraction having an initial boiling point greater than 70 ° C. or greater than 80 ° C. and a d final boiling below 160 ° C, or below 170 ° C or below 180 ° C and at least one heavy fraction boiling at a temperature greater than 160 ° C or greater than 170 ° C or greater than 180 ° C. In this variant, said heavy fraction boiling at a temperature greater than 160 ° C. or greater than 170 ° C. or greater than 180 ° C., is preferably sent in said third hydrocracking step of the process according to the invention and said light fraction is preferably sent directly into the fourth steam-cracking step, in admixture with the fraction containing naphtha separated after said third hydrocracking step.

The fractionation zone preferably comprises a fractionation section which incorporates a high temperature high pressure separator (HPHT) and then an atmospheric distillation. Hydrocracking step According to the invention, at least a portion of the effluent from the second hydrotreating step, and preferably at least a portion, preferably the entire organic phase of the effluent from said second optionally separated hydrotreating step, mixed with a recycle of at least a portion and preferably all of the fraction boiling at a temperature above 160 ° C separated at the end of a third hydrocracking step , is sent in said third step of hydrocracking in the presence of hydrogen in at least one reactor containing a hydrocracking catalyst, operating at a temperature between 250 and 480 ° C, at a pressure between 2 and 25 MPa, at a space velocity between 0.1 and 20 h -1, and with a quantity of hydrogen introduced such that the volume ratio hydrogen on hydrocarbons is between 80 and 5000 Nm3 / m3. Preferably, the feedstock of said third hydrocracking step is the organic phase of the effluent from the second hydrotreating step, optionally separated and / or fractionated. The sequencing of said second hydrotreatment step followed by the third hydrocracking step of the process according to the invention of said third hydrocracking step makes it possible to carry out on one side a severe hydrocracking in order to obtain a high yield of naphtha cut. 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 of preferably between 320 and 450 ° C., preferably between 350 and 435 ° C., under a pressure of between 3 and 20 MPa, at a space velocity of between 0.1 and 6 h -1, preferably between 0.2 and 3 h -1, and with a quantity of hydrogen introduced such that the volume ratio hydrogen on hydrocarbons is between 100 and 3000 Nm3 / m3. According to the invention, at least a part and preferably all of the fraction having an initial boiling point of between 120 and 200 ° C., preferably between 130 and 200 ° C., and preferably between 140 and 180 ° C. ° C, separated at the end of the third hydrocracking step is recycled in said third step. These operating conditions used in said third hydrocracking step make it possible to obtain a high yield in naphtha section having a final boiling point of between 120 and 200 ° C., preferably between 130 and 200 ° C., and preferably between 140 and 180 ° C. Said operating conditions make it possible to achieve pass conversions, products having initial boiling points of between 120 and 200 ° C., preferably between 130 and 200 ° C., and preferably between 140 and 180 ° C. lighter, greater than 30% by weight, preferably between 50 and 100% by weight and very preferably between 70 and 100% by weight. The pass conversion of the products having initial boiling points of between 120 and 200 ° C. into lighter products and defined as being: (mass of X + in the feedstock entering into said mass step of X + in the effluent exiting from said step) / (mass of X + in the feed entering said step), X + representing the cup whose initial boiling point is between 120 and 200 ° C. Said second hydrotreatment step and said third hydrocracking step can advantageously be carried out in a fixed bed reactor or in a so-called driven bed: moving bed reactor, bubbling bed reactor or suspension reactor. It is possible to use a single catalyst or several different catalysts, simultaneously or successively, in the case of a fixed bed reactor. Said steps can be carried out industrially in one or more reactors, with one or more catalytic beds. The reaction exotherm during the hydrotreatment is limited by any method known to those skilled in the art: recycling of the liquid product, cooling with recycled hydrogen, etc. Said third hydrocracking step according to the invention is preferably carried out in a so-called two-stage scheme, to maximize the yield of naphtha, optionally in combination with possibly a conventional hydrotreatment catalyst located upstream of the hydrocracking catalyst. The so-called two step hydrocracking scheme is widely known in the prior art. The hydrocracking catalysts used in the third hydrocracking step 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 6 of the periodic table of elements, or by a combination of at least one metal of group 6 of the periodic table and at least one metal of groups 8 to 10 Said hydrocracking catalyst may advantageously comprise metals of groups 8 to 10, preferably chosen from nickel and cobalt, preferably in combination with at least one metal of group 6, chosen from molybdenum and tungsten and a support . The element content of groups 8 to 10 is advantageously between 0.5 and 20% by weight of oxide and the content of element of group 6 is advantageously between 1 and 40% by weight of oxide, preferably between 5 and 30% by weight relative to the total mass of said catalyst. The total content of metal oxides of groups 6 and 8 to 10 in the catalyst is generally between 5 and 40% by weight relative to the total weight of said catalyst. In the case where the catalyst comprises at least one group 6 metal in combination with at least one non-noble group 8-10 metal, said catalyst is preferably a sulfurized catalyst.

The hydrocracking catalyst may also contain a promoter element chosen from phosphorus, silicon and boron. This element may have been introduced into the matrix or preferably deposited on the support. The concentration of said element is usually less than 20% by weight (on the oxide base) and usually less than 10% relative to the total mass of said catalyst. When boron trioxide (B203) is present, its concentration is less than 10% by weight. Advantageously, the hydrocracking catalyst is chosen from catalysts comprising the following combinations of metals: NiMo, CoMo, NiW, CoW, NiMoP and preferably NiMo, NiW, NiMoWP and NiMoP, more preferably NiMoP.

Said hydrocracking catalyst advantageously comprises a support chosen from the group formed by alumina, silica, silica-aluminas, zeolites, 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.

Preferably, the support of said catalyst comprises at least one zeolite and preferably a zeolite Y of structural type FAU, in order to maximize the yield of hydrocracked naphtha, and ultimately in aromatics after catalytic reforming of said naphtha. The hydrocracking catalyst comprises a zeolite Y content of between 2 and 40% by weight and preferably between 10 and 35% by weight relative to the total weight of said catalyst.

Other preferred catalysts are so-called composite catalysts and comprise at least one hydro-dehydrogenating element chosen from the group formed by the elements of group 6 and groups 8 to 10 and a support based on silico-aluminum matrix and based on at least one zeolite as described in EP1711260. 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). Separation According to the invention, the effluent obtained at the end of the third hydrocracking step undergoes at least one separation step in order to recover at least one fraction containing naphtha and a fraction having an initial boiling point comprised between 120 and 200 ° C, preferably between 130 and 200 ° C, and preferably between 140 and 180 ° C. The naphtha-containing fraction is defined as a fraction having a final boiling point between 120 and 200 ° C, preferably between 130 and 200 ° C, and preferably between 140 and 180 ° C.

According to the invention, at least a part of the fraction having an initial boiling point of between 120 and 200 ° C., preferably between 130 and 200 ° C., and preferably between 140 and 180 ° C., is recycled in said third hydrocracking step. Preferably, all of said fraction is recycled so as to maximize the yield in naphtha section. Said fraction is therefore recycled to extinction in said third hydrocracking step. In the case where all of said fraction is recycled, a purge can advantageously be implemented. 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. A gaseous fraction can also advantageously be separated. Said gaseous fraction can then be advantageously treated in a hydrogen purification unit. The recovered hydrogen is advantageously recycled in said third hydrocracking step and / or in said first hydroreforming step. Gases containing undesirable nitrogen, sulfur and oxygen compounds are removed. Steam cracking step According to the invention, at least a portion and preferably all of the fraction containing the naphtha is sent to the fourth steam cracking step to produce an effluent comprising at least olefinic compounds having from 2 to 6 carbon atoms , aromatic compounds (BTX) and hydrogen. The olefinic compounds having 2 to 6 carbon atoms are preferably ethylene, propylene, and butadiene. The aromatic compounds are preferably the aromatic C 6 -C 8 compounds, and preferably benzene, toluene and xylenes / ethylbenzene, called BTX.

The fraction containing the naphtha obtained after separation of the hydrocracking effluent has a high content of naphthenes and very few impurities thanks to the implementation of the third hydrocracking step under severe conditions. Steam cracking is generally the method of choice for obtaining petrochemical feedstocks such as olefinic compounds having 2 to 6 carbon atoms, such as, for example, ethylene, propylene and butadiene and aromatic compounds ( BTX). Generally, the current steam cracker charges are derived entirely from petroleum gases and liquids ranging from ethane to diesel with variable yields of ethylene and propylene and very little butadiene. The chemical reactions involved in the steam cracking stage are numerous. They are well known; the main ones are thermal cracking reactions but also hydrogenation, dehydrogenation reactions which lead to paraffins, diolefins and acetylenes, and condensation leading to diolefins and aromatics. The fourth steam-cracking stage can advantageously be carried out according to any one of the known methods, and is not limited to a particular process.

According to the invention, said fourth steam cracking step operates at a temperature between 700 and 900 ° C and preferably between 750 and 850 ° C in the presence of water vapor injected in a weight ratio relative to the fraction containing the naphtha entering said step between 0.2 and 2 and preferably between 0.5 and 1.5. The residence time of at least a portion of the fraction containing the naphtha and the water vapor injected is advantageously between 0.05 and 1.5 seconds and preferably between 0.2 and 1.2 seconds. According to a preferred embodiment, said process comprises a step of separation of olefinic compounds having from 2 to 6 carbon atoms, aromatic compounds and hydrogen obtained at the end of the fourth steam-cracking step. Said step of separation of olefinic compounds having 2 to 6 carbon atoms and aromatic compounds (BTX) and hydrogen can advantageously be implemented by any method known to those skilled in the art. Preferably, it is carried out by distillation. The hydrogen produced in said fourth steam cracking stage and advantageously separated during the separation step is preferably recycled in the first hydroreforming step and / or in the second hydrotreating step and the third hydrocracking step.

DESCRIPTION OF THE FIGURE FIG. 1 illustrates a particular embodiment of the method according to the invention. The bio-oil resulting from the rapid pyrolysis of the biomass is introduced via line (1), mixed with hydrogen (3) and the organic phase (8) resulting from the hydroreforming step, in the first hydroreforming zone (2). The liquid effluent from the hydroreforming zone via line (4) is introduced into a separator (5). A gaseous phase containing mainly CO2, H 2 S and light C 1 -C 4 gases such as, for example, methane is separated via line (7) and an aqueous phase is also separated via line (6). An organic phase containing the partially deoxygenated bio-oil and hydrocarbons is withdrawn through line (8) and part of said organic phase is recycled to the hydroreforming zone. The non-recycled portion of said organic phase is then sent to a second hydrotreatment zone (9). The effluent from the second hydrotreatment stage is introduced via the pipe (10) into a separator. A gaseous phase (13), an aqueous phase (12) and an organic phase (14) are separated and the organic phase freed from nitrogen, oxygen and sulfur impurities is sent to a third hydrocracking step in a hydrocracking zone ( 15) in the presence of hydrogen (3).

The effluent from the second hydrocracking stage via line (16) is sent to a separator (17). A fraction having an initial boiling point of between 120 and 200 ° C is separated in line (18) and recycled to the third hydrocracking stage upstream of zone (15). The naphtha fraction having a final boiling point of between 120 and 200 ° C. is separated in the pipe (19) and sent to a fourth steam-cracking step in the zone (23). The mixture comprising the olefinic compounds having from 2 to 6 carbon atoms, and the aromatic compounds (BTX) is obtained at the end of the steam-cracking step in the pipe (24). The examples below illustrate the invention without limiting its scope. Examples: The experiments were conducted with a bio-oil obtained by rapid pyrolysis of hardwood mixtures. The operating conditions of the rapid pyrolysis are a temperature of 500 ° C. and a residence time of the vapors of the order of one second. Table 1 below presents the data obtained following the analysis of this bio-oil. Table 1: Bio-oil analysis Density at 20 ° C, g / cm3 1.21 Cynematous viscosity at 20 ° C, mm2 / s 182.6 pH 2.5 Higher heat value, MJ / kg 18.03 Ability lower heat, MJ / kg b 16.46 Water content,% mass 20.7 Pyrolytic lignin,% mass 19.0 Carbon, mass% 43.9 Hydrogen, mass% 7.39 Nitrogen, mass% <0.05 Oxygen % mass 47.3 Sulfur (ppm) 92 a The higher calorific value is measured with a calorimetric bomb, b the lower calorific value is calculated using the following equation: LHV (J / g) = HHV (J / g EXAMPLE 1 Preparation of the Hydroreforming Catalysts C1 Catalyst C1 is a NiCr / C formulation catalyst. Precursors of nickel and chromium, in their nitrate form (respectively Ni (NO3) 2 and Cr (NO3) 3), are provided by Aldrich. The support is an active carbon in cylindrical extruded form supplied by the company Norit (RX3 Extra). This catalyst C1 is obtained by dry impregnation of an aqueous solution containing the precursors of nickel and chromium. The volume of solution is equal to the volume of water uptake of the support (that is to say the maximum volume of water that can penetrate its porosity). The concentrations of nickel and chromium precursors in solution are determined so as to obtain the target levels on the final catalyst: 10% by weight of nickel and 5% by weight of chromium. After impregnation of this aqueous solution, the catalyst is allowed to mature at room temperature for 4 hours in a chamber saturated with water, then dried in an oven at 70 ° C. under air for 3 hours and finally subjected to a heat treatment at 300 ° C. for 3 hours under nitrogen (at a flow rate of 1.5 L / h / g of catalyst). Table 1. Catalyst Formulation Cl. Catalyst Cl Support Activated Carbon Ni (% wt) 10.2 Cr (% wt) 4.9 Mo (wt%) Example 2: Production of ethylene, propylene, and butadiene This example illustrates the production of an olefinic mixture and in particular ethylene, propylene, and butadiene from the treatment of a non-pretreated bio-oil liquid feed with the characteristics described in Table 1, according to the process according to the invention. For the first hydroreforming step, the non-pre-treated bio-oil is first introduced, mixed with an organic liquid phase resulting from the hydroreforming step at a ratio of 1: 3 (organic liquid phase: bio- oil) simulating a recycling rate of 0.33, in an autoclave reactor with a hydroreforming catalyst. 15 g of the catalyst C1 prepared according to Example 1 are reduced in a reduction cell at 300 ° C. for 3 h under 30 NL / h of hydrogen and then introduced, in an inert glove bag, into a basket which is placed in a the autoclave reactor with 75 g of bio-oil and 25 g of the organic liquid phase from the hydroreforming step. The reactor is sealed and purged with nitrogen and then with hydrogen. The reactor is then pressurized to 4.83 MPa (700 psia) with 18.37 LH H2, then gradually heated with stirring to 330 ° C and maintained at that temperature for 3 hours during which the maximum pressure reached was 13 ° C. , 9 MPa (2016 psia). The equivalent space velocity, defined as the ratio of the volume of bio-oil introduced on the product of the catalyst volume by the duration of the test, is equal to 0.7 h -1. The reactor is then rapidly cooled and cooled to room temperature, which reduces the pressure to 1.17 MPa (170 psia). The reaction gas phase produced, NL 4.9 total, is sent to a collector bottle and analyzed. Gas chromatographic analysis indicated that it contained 63.9% H2, 21.6% 002, 10.6% CH4 and 3.0% CO by volume. The basket containing the catalyst is recovered from the product liquid, which is separated into 50.33 g of a higher homogeneous organic liquid liquid phase of density 980 kg / m 3 and 38.37 g of a lower aqueous phase. No tar deposit is observed in the reactor. The organic liquid phase contains 13.93% by weight of elemental oxygen and 4.68% by weight of water (2.36 g) of water, determined by elemental analysis and Karl-Fischer assay. The overall net hydrogen consumption is 15.26 NL (1.39 g), which corresponds to about 1.85% by weight relative to the mass of bio-oil introduced. From the above data, the organic liquid phase and aqueous phase yields are determined at 38.5% and 29.4% by weight, respectively, relative to the bio-oil, ie 27.8% and 21.2%. weight, respectively, on the basis of dry biomass containing 0% moisture. The organic liquid phase resulting from the hydroreforming step is then sent to a fixed bed reactor traversed in the presence of a conventional hydrotreatment catalyst for a hydrotreating step (HDT). The hydrotreatment catalyst used comprises 4.3% by weight of NiO, 21% by weight of MoO 3 and 5% by weight of P 2 O 5 supported on a gamma alumina. Prior to the test, this catalyst is sulphurized in situ at a temperature of 350 ° C., using a feedstock containing heptane additive of 2% by weight of dimethyl disulphide (DMDS). mL / h of the organic liquid phase resulting from the hydroreforming step is introduced into an isothermal reactor and fixed bed charged with 25 mL of hydrotreatment catalyst. The corresponding space velocity is equal to 1 h-1. 800 Nm3 of hydrogen / m 3 of feed are fed to the reactor maintained at a temperature of 320 ° C and a pressure of 10 MPa (1450 psia). In order to maintain the catalyst in the sulphide state, 50 ppm weight of sulfur in the form of DMDS is added to the feedstock. Under the reaction conditions, the DMDS is completely decomposed to form methane and H2S. After one hour of operation of the test, 19.63 g of a hydrocarbon liquid phase of density 842 kg / m3, containing 0.12% by weight of water, 0.2% by weight of elemental oxygen (limit of quantification of the analyzer) and 99.7 wt% of hydrocarbons was recovered. The net hydrogen consumption corresponds to 2.2% by weight relative to the mass of organic effluent sent to the hydrotreatment stage, which corresponds to a net hydrogen consumption of 0.9% by weight relative to the bio-oil. The overall yield of hydrocarbon liquid is equal to 29.7% by weight relative to the bio-oil, ie 21.4% by weight on the basis of dry biomass.

The hydrocarbon liquid resulting from the hydrotreating step is then sent to a fixed-bed reactor traversed in the presence of a hydrocracking catalyst simulating the conversion reactor for the hydrocracking step. The hydrocracking catalyst used comprises 3.3% by weight of NiO, 16% by weight of MoO 3 and 3.8% by weight of P 2 O 5 supported on a commercial CBV 712 reference zeolite, kneaded with a gamma-alumina in a weight ratio of 1: 4 (zeolite mass). : mass of alumina). Said catalyst is sulphurized. 25 ml / h of hydrocarbon liquid from the hydrotreating step is introduced into an isothermal reactor and fixed bed charged with 25 ml of hydrocracking catalyst. The corresponding space velocity is equal to 1 h-1. 1000 Nm3 of hydrogen / m 3 of feed are introduced into the reactor maintained at a temperature of 380 ° C and a pressure of 12 MPa (1740 psia). The effluent recovered at the outlet of the hydrocracking reactor is cooled, flashed and fractionated into two distillation slices, a fraction whose boiling point is less than 180 ° C., and a fraction whose boiling point is higher. at 180 ° C. The fraction whose boiling point is below 180 ° C called naphtha fraction is recovered and the 180 ° C + cut is recycled continuously to the inlet of the hydrocracking reactor. The pass conversion of products having boiling points above 180 ° C to a product having a boiling point below 180 ° C is estimated at 50%. The overall yield of the naphtha fraction having boiling points between 30 and 180 ° C. is equal to 29.03% by weight relative to the bio-oil, ie 21% by weight on the basis of dry biomass.

The C5-180 ° C naphtha section is then sent to a steam cracking furnace in the presence of steam in a weight ratio of H 2 O / feed entering the furnace equal to 1 and operating at a temperature of 800 ° C. The residence time in the oven is 0.5 seconds. The effluent comprising at least olefinic compounds having 2 to 6 carbon atoms, aromatic compounds (BTX) and hydrogen at the reactor outlet is cooled and recovered.

The yields of ethylene, propylene, and butadiene relative to the bio-oil and relative to the dry biomass are given in the table below. The yield of ethylene, propylene and butadiene relative to the bio-oil is 12.9% and 9.3% compared to dry biomass. The yield of benzene, toluene and C8 aromatics relative to the bio-oil is 4.17% and 3.02% relative to the dry biomass Hydrogen and ethylene, propylene, and butadiene are separated by distillation. Table 2: Efficiency of the steam-cracking furnace% wt / Bio-oil rdt / Dry biomass H2 0.9 0.26 0.19 Methane 15.2 4.41 3.19 Ethylene 20 5.81 4.19 Propylene 18.3 5.31 3.84 Butadiene 6 1.74 1.26 Other C4 7.2 2.09 1.51 C5-200 including BTX 22.6 6.56 4.74 Benzene 8.10 2.35 1.70 Toluene 4.11 1.19 0.86 C8 aro 2.18 0.63 0.46 non aro 8.22 2.39 1.72 Fuel 9.8 2.85 2.05

Claims (15)

REVENDICATIONS1. A process for producing olefinic compounds having a carbon number of 2 to 6 and aromatic compounds from a liquid feedstock comprising at least one bio-oil, said process comprising at least the following steps; the introduction of said feedstock in a first hydroreforming step in the presence of hydrogen and a hydroreforming catalyst comprising at least one transition metal selected from the elements of groups 3 to 12 of the periodic table and at least one support chosen from activated carbons, silicon carbides, silicas, transition aluminas, alumina silicas, zirconium oxide, cerium oxide, titanium oxide and transition metal aluminates, taken alone or as a mixture, to obtain at least one liquid effluent comprising at least one aqueous phase and at least one organic phase, said first hydroreforming stage operating at a temperature of between 200 and 380 ° C., at an absolute pressure of between 3 and and 30 MPa, at a space velocity of between 0.1 and 50, the treatment in a second hydrotreatment step of at least a portion of the organic phase of the effluent from the first stage. ape for hydroreforming in the presence of hydrogen in at least one reactor containing a hydrotreatment catalyst, operating at a temperature of between 250 and 400 ° C., at a pressure of between 2 MPa and 25 MPa, at an hourly space velocity comprised between between 0.1 h'1 and 20 h-1 and with a total amount of hydrogen mixed with the feed such that the volume ratio hydrogen / hydrocarbon is between 100 and 3000 Nm3 / m3, the treatment in a third step of hydrocracking at least a portion of the effluent from the second hydrotreating step, in the presence of hydrogen, in at least one reactor containing a hydrocracking catalyst, operating at a temperature of between 250 and 480 ° C, at a pressure between 2 and 25 MPa, at a space velocity of between 0.1 and 20 h -1, and with a quantity of hydrogen introduced such that the volume ratio hydrogen on hydrocarbons is between 80 and 5000 Nm3 / m3, in which at least u part of the fraction having an initial boiling point between 120 ° C and 200 ° C separated at the end of said third hydrocracking step is recycled to said third hydrocracking step, separation of the effluent obtained at the end of the third hydrocracking stage in at least one fraction containing naphtha and a fraction having an initial boiling point of between 120 ° C and 200 ° C, the introduction of at least a portion of the fraction containing naphtha from the separation step in a fourth steam-cracking step to produce an effluent comprising at least olefinic compounds having 2 to 6 carbon atoms, aromatic compounds and hydrogen, said steam-cracking step operating at a temperature between 700 and 900 ° C in the presence of water vapor injected in a weight ratio relative to the fraction containing the naphtha entering said step between 0.2 and The method according to claim 1, wherein said first hydroreforming step operates at a temperature of between 270 and 360 ° C. 3. Method according to one of claims 1 or 2, wherein said first hydroreforming step I0 operates at a temperature between 270 and 350 ° C. 4. Method according to one of claims 1 to 3, wherein said first supported hydroreforming catalyst comprises nickel, alone or in combination with at least one metal selected from chromium (Cr), molybdenum (Mo), manganese (Mn), tungsten (W), iron (Fe), cobalt (Co) and copper (Cu) and a support selected from silicas, transition aluminas, silica aluminas, porous carbon and transition metal aluminates, alone or in admixture. The method according to claim 4, wherein said first hydroreformed supported catalyst is a catalyst selected from catalysts comprising nickel (Ni), nickel and chromium (NiCr) or nickel and manganese (NiMn) on porous carbon and the catalyst comprising nickel and molybdenum (NiMo) on alumina or nickel aluminate. 20 6. Method according to one of claims 1 to 5, wherein a step of separating at least one organic phase, at least one aqueous phase and at least one gaseous phase can be implemented between said second hydrotreating step and said third hydrocracking step. The process according to claim 6, wherein said organic phase is fractionated in a fractionation zone into at least one light fraction boiling at a temperature between 80 and 160 ° C and at least one heavy fraction boiling at a higher temperature. at 160 ° C. 8. A process according to claim 7, wherein said heavy booitlant fraction at a temperature above 160 ° C is fed into said third hydrocracking step and said light fraction is sent directly to the fourth steam cracking step, in admixture with the fraction containing naphtha separated at the end of said third hydrocracking step. 9. Process according to one of claims 1 to 8, wherein said hydrocracking catalyst comprises metals of groups 8 to 10 selected from nickel and cobalt, in combination with at least one metal of group 6, chosen from molybdenum and tungsten and a support selected alumina, silica, silica-aluminas, zeolites, magnesia, clays and mixtures of at least two of these minerals. The process according to claim 9, wherein the support of said hydrocracking catalyst comprises at least one zeolite Y of FAU structural type. 11. Process according to one of claims 1 to 10, in which the totality of the fraction having an initial boiling point between 120 ° C and 200 ° C, separated at the end of the third hydrocracking step is recycled in said third step. 12. Method according to one of claims 1 to 11, wherein said fourth steam-cracking step operates at a temperature between 750 and 850 ° C in the presence of water vapor injected in a weight ratio relative to the fraction containing the naphtha entering said step between 0.5 and 1.5. The process according to claim 12, wherein the residence time of at least a portion of the naphtha-containing fraction and the injected water vapor is between 0.05 and 1.5 seconds. 14. Process according to one of Claims 1 to 13, in which the said process comprises a step of separation of the olefinic compounds having from 2 to 6 carbon atoms and the aromatic compounds of hydrogen obtained at the end of the fourth stage. steam cracking. The process according to claim 14, wherein the hydrogen produced in said fourth steam cracking step and separated during the separation step is recycled to the first hydroreforming step and / or the second hydrotreatment step and third hydrocracking step.