Method for dehydrogenating ethanol in a multi-tubular reactor

Abstract

The present invention relates to a method for the dehydrogenation of a feedstock containing ethanol, using at least one multi-tubular reactor, which advantageously comprises a plurality of tubes containing at least one dehydrogenation catalyst and a calender, wherein the feedstock is introduced into the tubes in gaseous form, the inlet temperature being greater than or equal to 240°C, the pressure being between 0.1 and 1.0 MPa, and the WWH being between 2 and 15 h. -1 wherein the heat transfer fluid flows through the calender at a flow rate such that the weight ratio of the heat transfer fluid relative to the feedstock is 1.0 or greater, and the heat transfer fluid is introduced into the calender in gas form at an inlet temperature of 260°C or greater and an inlet pressure of 0.10 MPa or greater and 1.10 MPa or less, and exits the calender at least partially in liquid form.

Description

[Technical Field]

[0001] The present invention relates to a process for the catalytic dehydrogenation of ethanol, in particular a process using at least one multi-tubular reactor with a condensing heat transfer fluid. The present invention also relates to a process for the production of butadiene from ethanol, which comprises, as a first reaction stage, a process for catalytic dehydrogenation of ethanol to form acetaldehyde in a multi-tubular reactor with a condensing heat transfer fluid, and then a second reaction stage in which butadiene is produced from a mixture of ethanol and acetaldehyde. [Background technology]

[0002] The dehydrogenation of ethanol to give acetaldehyde can be the first reaction step in several processes, for example, the Ostromislensky process, which is particularly known and is related to the Lebedev process, for the conversion of ethanol to butadiene. The ethanol dehydrogenation reaction is an equilibrium reaction, and the ethanol conversion is usually around 30%. It is a highly endothermic reaction (ΔH reaction = 72.4 kJ / mol). Reactor technology that allows for adequate heat compensation is therefore necessary.

[0003] More generally, several technologies exist for industrial dehydrogenation: ABB's Catofin® process, UOP's Oleflex® process, ThyssenKrupp's Star Process® process, and Linde's PDH process. All of these industrial processes have the objective of dehydrogenating light alkanes to give olefins, specifically dehydrogenating propane and butane to give propylene and butene, respectively. More specifically, the Catofin® and Oleflex® processes employ a series of parallel adiabatic reactors, each with a fixed catalyst bed and a moving catalyst bed. The PDH process uses, for its part, a reactor with a fixed bed externally heated by fuel. The Star Process® process carries out the reaction of a feedstock diluted with steam in tubes arranged within a furnace. The choice of reactor technology appears to depend, at least in part, on the catalyst used for the alkane dehydrogenation to be carried out.

[0004] A multi-tubular reactor, also called a reactor-exchanger, can be used to allow the process to operate under isothermal or quasi-isothermal conditions. For example, this is the case in the process for dehydration of isopropanol described in Patent Document 1 of Mitsui Petrochemical Industries, Ltd. In this patent, isopropanol is dehydrated to give propylene in the presence of alumina at 290°C to 320°C in a vertical tubular reactor having a length of 0.5 m and an internal diameter of 25.4 mm, obtaining a conversion of isopropanol of at least 85% and a selectivity for propylene of 79% or more.

[0005] Patent application WO 02 / 04999 describes, in part, a process for the dehydration of isobutanol to give butenes, comprising simultaneous dehydration and isomerization steps carried out in a multi-tubular reactor under isothermal or quasi-isothermal conditions in the presence of a catalyst containing FER zeolite at a temperature of 300° C. or 350° C. The degree of conversion of isobutanol obtained by this process is approximately 100%, with a selectivity for butenes of at least 97%.

[0006] Finally, the application (Patent Document 3) describes a process for the dehydration of ethanol to give ethylene in a multi-tubular reactor in the presence of ZSM-5 zeolite, at an inlet temperature of the feedstock in the tubes of 420-430°C. To maintain the temperature, a salt melting at 470°C is used as a heat transfer fluid. Under these conditions, the degree of conversion of ethanol achieved is greater than 99%, with a selectivity for ethylene of at least 98%.

[0007] However, none of these documents address the reaction for the dehydrogenation of ethanol to give acetaldehyde. Therefore, the object of the present invention is to provide a process for the dehydrogenation of ethanol to give acetaldehyde, which makes it possible to achieve a satisfactory conversion of ethanol and a satisfactory selectivity for acetaldehyde, while avoiding premature deactivation of the catalyst, limiting the consumption of large amounts of utilities, such as steam, and minimizing capital and operating costs. [Prior art documents] [Patent documents]

[0008] [Patent Document 1] U.S. Patent No. 5,227,563 [Patent Document 2] International Publication No. 2018 / 046515 [Patent Document 3] French Patent Application Publication No. 3089973 Summary of the Invention [Means for solving the problem]

[0009] (Summary of the Invention) The present invention provides a process for the dehydrogenation of ethanol to give acetaldehyde, comprising a step of dehydrogenating a feedstock comprising ethanol, said dehydrogenation step employing a reaction section comprising at least one multi-tubular reactor, said multi-tubular reactor comprising one or more tubes and a shell; the tube(s) each contain at least one fixed bed of at least one dehydrogenation catalyst; The feedstock is in the form of gas, and the inlet temperature of the feedstock is 240°C or higher, the inlet pressure of the feedstock is 0.1 to 1.0 MPa, and the weight hourly space velocity (WWH) of the feedstock at the inlet is 2 to 15 h -1 into said tube(s), a heat transfer fluid is passed through the shell, the heat transfer fluid being introduced into the shell in gaseous form and being at least partially in liquid form at the shell outlet; introducing the heat transfer fluid into the shell at a flow rate such that the ratio of the weight flow rate of the heat transfer fluid at a shell inlet to the weight flow rate of the feedstock at the inlet of a tube(s) is 1.0 or greater; When the heat transfer fluid is introduced into the shell, the inlet temperature of the heat transfer fluid is 260°C or more and 400°C or less, and the inlet pressure of the heat transfer fluid is 0.10 MPa or more and 1.10 MPa or less; The process produces a dehydrogenation effluent containing at least acetaldehyde, hydrogen, and unconverted ethanol. Regarding the method.

[0010] The present invention offers the advantage of compensating for the endothermic nature of the ethanol dehydrogenation reaction (ΔH reaction = 72.4 kJ / mol) by maintaining the temperature of the reaction medium in the temperature range suitable for the dehydrogenation reaction, in particular at values ​​above 230°C, preferably above 240°C, and advantageously above 250°C. The present invention uses a reactor of the reactor-exchanger type, which follows a specific form of condensation of the heat transfer fluid. The phase change enthalpy of the heat transfer fluid can thus be used to provide the heat necessary for the energy requirements of the ethanol dehydrogenation reaction. The present invention also makes it possible to have a uniform temperature on the outer wall of the tubes inside which the reaction takes place.

[0011] Therefore, the present invention makes it possible to achieve a satisfactory degree of conversion of ethanol and a high selectivity for acetaldehyde, in particular a degree of conversion of ethanol of 25% or more, preferably approximately 35%, and a selectivity for acetaldehyde of more than 90%.

[0012] Another advantage of the present invention is that, particularly when an adiabatic reactor is used, thermal compensation can be achieved by transferring its sensible heat to the reaction, without the addition of a diluent to the feedstock, which may otherwise be necessary to "thermally buffer" the temperature drop induced by the endothermic nature of the reaction. A diluent traditionally used as a "thermal buffer" is steam. In fact, too high a water content in the reactor, e.g., more than 20 wt. % relative to the total weight of the feedstock, can be harmful to the dehydrogenation catalyst and cause its premature deactivation. Therefore, the ethanol dehydrogenation process of the present invention does not require the addition of a diluent, thereby limiting the risk of premature deactivation of the dehydrogenation catalyst.

[0013] The present invention also offers the advantage of being able to dehydrogenate ethanol feedstocks of very diverse composition and origin (which may contain, for example, water in a mixture with ethanol). This is because the present invention is particularly applicable to ethanol feedstocks produced from renewable sources derived from biomass, often referred to as "bioethanol." It is also applicable to ethanol-rich effluents, such as those obtained after the treatment of reaction effluents derived in particular from the conversion of ethanol to butadiene and advantageously recycled to the reaction section, which in particular comprises a stage for dehydrogenating ethanol as described in patent FR 3 026 100. More particularly, the present invention allows the dehydrogenation of the ethanol-rich effluent obtained in step E1) of patent FR 3 026 100 and which may contain, for example, up to 18% by weight of water.

[0014] Furthermore, the fact that the addition of diluents to the feedstock is avoided advantageously makes it possible to limit the size of the equipment items and the number of subsequent separation stages, particularly for ethanol dehydrogenation processes that require dilution with steam, and consequently minimize the capital and operating costs of the process.

[0015] The present invention presents yet another advantage: it avoids the need for an increased number of catalyst beds, since adiabatic reactor technology requires a series of fixed catalyst beds interposed with heat exchangers to compensate for the endothermic nature of the reaction and achieve an ethanol conversion of approximately 30%. Conversely, the present invention, which proposes the use of a multi-tubular reactor coupled with a specific form of condensation of an appropriate heat transfer fluid, makes it possible to limit the capital and operating costs of the process while achieving the target conversion (25%-35%, even 30%-35%) and high selectivity for acetaldehyde (over 90%). DETAILED DESCRIPTION OF THE INVENTION

[0016] (Description of the embodiments of the present invention) According to the present invention, the expressions "of between ... and ..." and "between ... and ..." are equivalent and mean that both limits of the interval are included in the stated range of values. If this is not the case and both limits are not included in the stated range, such information will be made clear by the present invention.

[0017] Within the meaning of the present invention, various ranges of parameters for a given stage, such as pressure ranges and temperature ranges, may be used alone or in combination, for example, a range of preferred pressure values ​​may be combined with a range of more preferred temperature values, within the meaning of the present invention.

[0018] Subsequently, specific and / or preferred embodiments of the present invention may be described, which may be used separately or may be combined together without restriction on combinations where this is technically feasible.

[0019] The present invention therefore provides a process for the dehydrogenation of ethanol to give acetaldehyde, comprising, and preferably consisting of, a step of dehydrogenating a feedstock comprising ethanol, preferably at least 50% by weight of ethanol, preferentially at least 70% by weight of ethanol, suitably at least 80% by weight of ethanol, and optionally water, where: - the dehydrogenation step uses a reaction section comprising at least one multi-tubular reactor, the multi-tubular reactor comprising one or more tubes, preferably a plurality of tubes, and a shell; said tube(s) are highly advantageously made from any type of steel, preferably from alloy steel, suitably from stainless steel, and preferably have a length of 1.0 to 6.0 m, preferably 2.0 to 3.0 m, an internal diameter of preferably 30.0 to 60.0 mm, preferentially 40.0 to 50.0 mm, and a wall thickness of the tubes of preferably 1.5 to 5.0 mm, preferentially 2.0 to 4.0 mm, suitably 2.2 to 3.2 mm; each tube contains at least one fixed bed comprising at least one dehydrogenation catalyst, said dehydrogenation catalyst preferably comprising at least elemental copper on an inorganic support, preferably silica, said catalyst being highly advantageously in the form of particles having a mean equivalent diameter of between 0.5 and 10.0 mm, preferably between 1.0 and 5.0 mm; - the feedstock is fed to the tube(s) in gaseous form, wherein the inlet temperature of the feedstock to the tube(s) is 240°C or higher, preferably 240°C to 350°C, preferentially 250°C to 300°C, and suitably 260°C to 290°C, wherein the inlet pressure of the feedstock to the tube(s) is 0.1 to 1.0 MPa, preferably 0.2 to 0.5 MPa, and preferentially 0.3 to 0.4 MPa, wherein the hourly space velocity (WWH) of the feedstock at the inlet of the multi-tubular reactor is 2 to 15 h -1 , preferably 2 to 10 hours -1 and - a heat transfer fluid flows through the shell, preferably co-currently with the feedstock in the tube(s), the heat transfer fluid being introduced into the shell in gaseous form and being at least partially in liquid form at the shell outlet; - the heat transfer fluid is introduced into the shell of the multi-tubular reactor at a flow rate such that the ratio of the weight flow rate of the heat transfer fluid at the shell inlet to the weight flow rate of the feedstock at the tube inlet is greater than or equal to 1.0, preferably greater than or equal to 1.5, and advantageously less than or equal to 10.0, preferably less than or equal to 5.0, and suitably less than or equal to 2.0; - when the heat transfer fluid is introduced into the shell of the multi-tubular reactor, the inlet temperature of the heat transfer fluid is 260°C or more, preferably 270°C or more, and suitably 290°C or more, and 400°C or less, preferably 380°C or less, and the inlet pressure of the heat transfer fluid is 0.10 MPa or more, preferably 0.13 MPa or more, and suitably 0.20 MPa or more, and 1.10 MPa or less, and suitably 0.85 MPa or less; and producing at least a dehydrogenation effluent comprising acetaldehyde, hydrogen, and unconverted ethanol.

[0020] (Feed material) According to the invention, the feedstock to be treated in the dehydrogenation process is an ethanol-containing feedstock, preferably comprising at least 50% by weight of ethanol, preferentially at least 70% by weight of ethanol, and suitably at least 80% by weight of ethanol.

[0021] The feedstock for the dehydrogenation process may further optionally comprise water, preferably in a content of less than 50% by weight, preferentially less than 30% by weight, suitably less than 20% by weight, for example between 1% and 20% by weight, by weight of water relative to the total weight of the feedstock.

[0022] The feedstock (especially containing less than 50% by weight of ethanol) may be concentrated prior to the process of the present invention by any means known to those skilled in the art, for example by distillation, absorption, pervaporation, or extraction with a solvent.

[0023] In addition to water, the feedstock may contain impurities, such as butanol, in a content preferably of less than 10% by weight, preferentially less than 5% by weight, and suitably less than 2% by weight, relative to the total weight of the feedstock.

[0024] The feedstocks to be treated in the process according to the invention are optionally obtained by processes for the synthesis of alcohols starting from fossil resources, for example starting from coal, natural gas or carbon-based wastes, etc.

[0025] Preferably, the feedstock is advantageously derived from non-fossil resources. It can be obtained from renewable resources derived from biomass, often referred to as "bioethanol." Bioethanol is a feedstock produced biologically, preferably by fermentation of sugars (e.g., sugar-producing plant crops, such as sugarcane (saccharose, glucose, fructose, and sucrose), beetroot, or also starchy plants (starch)) or lignocellulosic biomass or hydrolyzed cellulose (mainly glucose, xylose, and galactose), containing various amounts of water. For a more complete description of conventional fermentation methods, reference may be made to the publication "Biofuels, Status Quo, Perspectives, and Development Challenges," by Daniel Ballerini, published by Technip, 2006.

[0026] Said feedstock may also advantageously be obtained from synthesis gas.

[0027] Said feedstock may also advantageously be obtained by hydrogenation of the corresponding acid or ester, in which case acetic acid or an acetate ester is advantageously hydrogenated with hydrogen to give ethanol, which may advantageously be obtained by carbonylation of methanol or by fermentation of carbohydrates.

[0028] The feedstock can also be an ethanol effluent obtained after the treatment of an effluent from a process for the conversion of ethanol to butadiene (in particular after the separation and purification steps). More specifically, the feedstock for the dehydrogenation process can be an ethanol effluent obtained after the treatment of a reaction effluent from the conversion of ethanol to butadiene, which mainly comprises ethanol, i.e., at least 50% by weight, preferably at least 70% by weight, and preferentially at least 80% by weight of ethanol, and which is advantageously recycled to the reaction section, which in particular comprises a step for dehydrogenating ethanol. According to a very particular embodiment of the invention, the ethanol-containing feedstock is an ethanol-rich effluent, advantageously obtained at the end of a step for the treatment of an effluent from a process for the conversion of ethanol to butadiene, for example, the ethanol-rich effluent advantageously obtained at the end of step E1) of the process described in patent FR 3 026 100. This ethanol-rich effluent can in particular contain up to 18% by weight of water.

[0029] (Dehydrogenation stage) According to the present invention, the dehydrogenation process comprises at least a step of dehydrogenating the feedstock containing ethanol, resulting in a dehydrogenation effluent containing at least acetaldehyde, hydrogen, and conventional unconverted ethanol. The resulting dehydrogenation effluent may also contain water, especially if the feedstock itself contains water. It may also contain impurities (especially those already present in the feedstock) and / or co-products, especially formed during the dehydrogenation reaction.

[0030] The dehydrogenation step uses a reaction section containing at least one multi-tubular reactor, in which the dehydrogenation reaction occurs. The reaction section may contain at least two multi-tubular reactors, and preferably fewer than 10 multi-tubular reactors. Preferably, the reaction section contains two multi-tubular reactors, one of which is in operation, i.e., is fed with a feedstock and performs the dehydrogenation reaction, and the other is in a regeneration-replacement mode. The expression "regeneration-replacement mode" means that the multi-tubular reactor is not fed with an ethanol-containing feedstock, and the catalyst is in the process of being regenerated or charged, or alternatively, has been regenerated and / or charged and is ready to operate (i.e., awaiting operation).

[0031] Advantageously, each multi-tubular reactor comprises one or more tubes and a shell. The dehydrogenation reaction advantageously occurs in the tube(s) of the multi-tubular reactor(s), advantageously during operation. In the continuation of the disclosure, the tube(s) of the multi-tubular reactor(s) may also be referred to as reaction tube(s). The shell is the reactor casing (typically cylindrical), inside which the tube(s) are located, preferably parallel to each other, if there are several, and to the shell wall, and through which the heat transfer fluid flows. The shell may also comprise one or more baffles or any other system, preferably uniformly distributed within the shell, allowing good diffusion and homogenization of the heat transfer fluid and therefore good distribution of heat. The shell and the tubes may have a specific design or a specific texture that allows promoting condensation and / or drainage of the heat transfer fluid.

[0032] According to the invention, each tube contains at least one fixed bed containing at least one dehydrogenation catalyst. Preferably, each tube contains a fixed bed of dehydrogenation catalyst. Preferably, the dehydrogenation catalyst comprises at least elemental copper and, optionally, elemental chromium, on an inorganic support, preferably silica. Highly advantageously, the dehydrogenation catalyst is in the form of particles having an average equivalent diameter of 0.5 to 10.0 mm, preferably 1.0 to 5.0 mm. According to the invention, the average equivalent diameter defines the average equivalent diameter at the surface, advantageously determined by laser diffraction, and the average equivalent diameter of a particle advantageously corresponds to the average diameter of a sphere having the same specific surface area as the particle. For example, the dehydrogenation catalyst may be the Octolyst® 2001 or Octolyst® 2009 catalyst sold by Evonik.

[0033] Preferably, the multi-tubular reactor(s) contain a plurality of tubes within the shell, preferably at least 100 tubes, preferentially at least 1000 tubes, and suitably at least 2000 tubes. Typically, the multi-tubular reactor contains up to 20,000 tubes, preferably up to 10,000 tubes. For example, each multi-tubular reactor may contain 5,000 to 6,000 tubes.

[0034] Advantageously, the reaction tube(s) preferably have a length of 1.0 to 6.0 m, preferably 2.0 to 3.0 m. When a multi-tubular reactor comprises several tubes, all the tubes of said multi-tubular reactor advantageously exhibit the same length, within the tolerances that are due in particular to the manufacture and machining of the tubes. The internal diameter of each reaction tube is preferably 30.0 to 60.0 mm, preferentially 40.0 to 50.0 mm. Advantageously, the reaction tube(s) preferably exhibit a wall thickness of 1.5 to 5.0 mm, preferentially 2.0 to 4.0 mm, preferably 2.2 to 3.2 mm. The nominal or external diameter of the tubes may therefore vary between 33 and 70 mm, preferably between 44 and 56 mm. The specific dimensions of the reaction tube(s), in particular the length, inner diameter and wall thickness, are advantageously adapted to the pressure applied on the tube side (i.e., inside the tube(s)) and on the shell side (i.e., outside the tube(s)), while advantageously making it possible to limit the pressure drop inside the tube(s) and thus avoid the adverse effects of a drop in pressure on the performance qualities of the dehydrogenation reaction, in particular avoiding a drop in the ethanol conversion.

[0035] The size of the multi-tubular reactor of the dehydrogenation stage c), for example the diameter of the shell, can be adapted by the person skilled in the art according to general knowledge, depending inter alia on the number of tubes, their length and their diameter.

[0036] Multi-tubular reactors, especially industrial ones, and especially the tubes of said reactors, are conventionally made from materials inert to the reaction, typically made from steel or nickel. The tube(s) of the multi-tubular reactor(s) in the dehydrogenation process are preferably made from any type of steel, preferentially from alloy steel, and preferably from stainless steel. Highly advantageously, the shell of the multi-tubular reactor(s), as well as any baffles optionally present in the shell, are of the same material as the reactor tubes, preferably made from any type of steel, preferentially from alloy steel, and preferably from stainless steel.

[0037] According to the present invention, the feedstock is introduced in gas form into the tube(s) of each multi-tubular reactor, advantageously during operation, preferably at one end of the reaction tube(s), and suitably simultaneously into all reaction tubes of the multi-tubular reactor if the multi-tubular reactor contains a plurality of tubes. The inlet temperature of the feedstock when being fed into the reaction tube(s) is 240°C or higher, preferably 240°C to 350°C, preferentially 250°C to 300°C, suitably 260°C to 290°C, and the inlet pressure of the feedstock is 0.1 to 1.0 MPa, preferably 0.2 to 0.5 MPa, preferentially 0.3 to 0.4 MPa, and the weight hourly space velocity (WWH) of the feedstock at the inlet of the multi-tubular reactor is 2 to 15 h -1 , preferably 2 to 10 hours -1 , especially 3 to 7 hours -1 , for example 5h -1 Therefore, the flow rate of the feedstock at the inlet of the multi-tubular reactor (containing, for example, 4,000 kg to 30,000 kg, preferably 14,000 kg to 17,000 kg of dehydrogenation catalyst) can vary between 20,000 kg / h and 150,000 kg / h, preferentially between 50,000 kg / h and 100,000 kg / h, and preferably between 70,000 kg / h and 85,000 kg / h. According to the present invention, the weight hourly space velocity (WWH) can be defined as the ratio of the weight flow rate of the total feedstock entering the multi-tubular reactor to the weight of the dehydrogenation catalyst contained in all of the reaction tubes of the multi-tubular reactor.

[0038] The inlet temperature of the feedstock to the multi-tubular reactor(s) may advantageously be increased gradually, while advantageously remaining within the inlet temperature ranges mentioned above, so as to at least partially compensate for the deactivation of the dehydrogenation catalyst.

[0039] The dehydrogenation stage of the process according to the invention may optionally use a section for heating the feedstock upstream of the reaction section. Heating of the feedstock in any heating section may be carried out by any method known to those skilled in the art, for example by heat exchange with a fluid which may very particularly be a heat transfer fluid circulating in the shell.

[0040] The flow of the feedstock may be in an ascending or descending manner in each tube, preferably in a descending manner.

[0041] According to the invention, a heat transfer fluid flows through the shell(s) of the multi-tubular reactor(s), advantageously during operation, in particular between said reaction tubes, advantageously in co-current or counter-current flow with respect to the flow inside the reaction tubes, preferably in co-current flow. The heat transfer fluid is introduced into the shell(s) of the multi-tubular reactor(s) in gaseous form and is at least partially in liquid form (i.e., in liquid form or as a gas-liquid mixture) at the shell outlet. In other words, the heat transfer fluid is advantageously introduced into the shell in saturated vapor form (i.e., in the vapor phase at the bubble point) and partially (or at least partially) condenses upon contact with the tubes inside which the endothermic dehydrogenation reaction takes place. The heat contribution necessary to maintain the temperature in the tube(s) in the temperature range compatible with the dehydrogenation reaction, in particular at least equal to 230°C, preferably above 240°C, and highly preferably above 250°C, is thus very advantageously guaranteed by the phase change enthalpy, in particular the enthalpy of condensation, of the heat transfer fluid used.

[0042] The heat transfer fluid is selected so that it is thermally stable under the operating conditions described above. The choice of heat transfer fluid may also be guided by other constraints: preferably, the heat transfer fluid is inert to the reactants and products of the dehydrogenation reaction; preferably, the heat transfer fluid does not induce corrosion of items of equipment, such as the shell-and-tube reactor or conduits. Highly advantageously, the heat transfer fluid exhibits a single boiling point at a given pressure. It is particularly advantageously selected so that it exhibits a range of boiling points or saturated vapor temperatures (which depends on the vapor pressure) compatible with the dehydrogenation reaction and / or so that its enthalpy of phase change from the gaseous state to the liquid state covers the energy requirements of the dehydrogenation reaction. Preferably, the heat transfer fluid is an oil, preferably comprising an organic compound, preferably a eutectic mixture of two organic compounds, highly advantageously with close boiling points and preferably with saturated vapor pressures such that the difference between the saturated vapor pressures of the organic compounds of the oil at a given temperature is 50 Pa or less, preferably 20 Pa or less, and suitably 10 Pa or less. More specifically, the heat transfer fluid comprises, and preferably consists of, a mixture of biphenyl and diphenyl oxide, for example, the oil sold by Dow under the name DOWTHERM® A.

[0043] Advantageously, the inlet temperature of the heat transfer fluid when it is introduced into the shell of the multi-tubular reactor, advantageously during operation, is at least 260°C, preferably at least 270°C, suitably at least 290°C, and at most 400°C, preferably at most 380°C, and the inlet pressure of the heat transfer fluid is at least 0.10 MPa, preferably at least 0.13 MPa, suitably at least 0.20 MPa, and at most 1.10 MPa, preferably at most 1.06 MPa, suitably at most 0.85 MPa. More specifically, the heat transfer fluid is introduced into the shell, advantageously in the form of saturated vapor, at a temperature preferably of 260°C or higher and a pressure of 0.10 MPa or higher, preferably a temperature of 270°C or higher and a pressure of 0.13 MPa or higher, preferably a temperature of 290°C or higher and a pressure of 0.20 MPa or higher, preferably a temperature of 400°C or lower and a pressure of 1.10 MPa or lower, and suitably a temperature of 380°C or lower and a pressure of 0.85 MPa or lower.

[0044] The inlet temperature and / or pressure of the heat transfer fluid in the shell may be gradually increased, advantageously within the inlet temperature and pressure ranges described above, to at least partially compensate for deactivation of the dehydrogenation catalyst.

[0045] The weight flow rate of said heat transfer fluid in the shell is advantageously adjusted so that the ratio of the weight flow rate of said heat transfer fluid in the shell to the weight flow rate of the feedstock introduced into the tube(s) is greater than or equal to 1.0, preferably greater than or equal to 1.5, and advantageously less than or equal to 10.0, preferably less than or equal to 5.0, and suitably less than or equal to 2.0.

[0046] Under these conditions, the transfer coefficient on the shell side (i.e., on the condensation side) is much higher than the transfer coefficient within the tubes, so the temperature advantageously remains constant along each of the reactor tubes and equal to the condensation temperature of the heat transfer fluid. Therefore, the temperature is uniform for all tubes and very close to the temperature of the shell, which is advantageous from the reactor design point of view, because the temperature is uniform throughout the reactor and, during operation, the expansion of the reactor material will be the same between the tubes and between the tubes and the shell, leading to a reduction in the cost of the item of equipment.

[0047] Advantageously, the dehydrogenation process may comprise a stage of conditioning of the heat transfer fluid, comprising a phase of recovery of the heat transfer fluid at the shell outlet of the multi-tubular reactor of the dehydrogenation stage, followed by a phase of compression and / or heating of the heat transfer fluid to obtain a heat transfer fluid in gaseous form at the temperature and pressure of the heat transfer fluid inlet in the shell of the dehydrogenation stage.

[0048] In such a reactor, depending on the specific operating conditions of the process according to the invention, in particular by using the enthalpy of condensation of a heat transfer fluid introduced into the shell at a specific temperature and pressure and with a flow rate adjusted relative to that of the feedstock in the reaction tube, the reaction for dehydrogenation of ethanol to give acetaldehyde advantageously takes place under isothermal or quasi-isothermal conditions, i.e. so that the temperature of the reaction medium at the reactor outlet (i.e. the dehydrogenation effluent at the reactor outlet) is similar to the inlet temperature of the feedstock or exhibits a difference of less than 30°C, preferably less than 15°C, relative to the temperature of the feedstock at the reactor inlet. Advantageously, under these conditions, the dehydrogenation effluent obtained at the end of the multi-tubular reactor advantageously exhibits, during operation, a temperature preferably greater than or equal to 230° C., preferentially greater than or equal to 240° C., suitably greater than or equal to 250° C., suitably greater than or equal to 260° C., and preferably less than or equal to 350° C., preferentially less than or equal to 300° C., suitably less than or equal to 290° C., and a pressure at the reactor outlet of, for example, between 0.1 and 0.5 MPa, preferably between 0.2 and 0.4 MPa.

[0049] Therefore, the specific conditions of the process according to the invention make it possible to achieve the desired performance qualities. In particular, the use of a reactor-exchanger under the specific operating conditions of the invention makes it possible to obtain, highly advantageously, ethanol conversions of at least 25%, preferably at least 30%, and even 35% ethanol conversion, and high selectivity for acetaldehyde, particularly at least 90% by weight. These performance qualities are achieved, as in processes involving an adiabatic reactor arrangement, without the addition of thermal diluents, which may have adverse effects on the activity of the dehydrogenation catalyst, and without increasing the number of catalyst beds and / or reactors, thus limiting capital and operating costs. The heat supply provided by the invention also makes it possible to easily adapt the operating conditions to changes and possible deactivation of the catalyst.

[0050] According to the present invention, the conversion of the ethanol feedstock, as a weight percentage, is defined by the following formula: [1 - (weight per hour of ethanol at outlet / weight per hour of ethanol at inlet)] x 100 The weight of ethanol per hour at the inlet or outlet corresponds to the weight flow rate at the inlet or outlet of the multi-tubular reactor and can be determined conventionally, for example by gas chromatography.

[0051] During the dehydrogenation step, the conversion of the feedstock may be accompanied by deactivation of the dehydrogenation catalyst, for example, by coking, adsorption of inhibitory compounds, and / or sintering. Therefore, the dehydrogenation catalyst can advantageously be periodically regenerated or replaced. Therefore, in a particular embodiment of the present invention, the method includes a regeneration-replacement step. In this particular embodiment, the reaction section preferably comprises at least two multi-tubular reactors. Preferably, the multi-tubular reactors are used in an alternating mode, also called a swing mode, alternating between a reaction (or operation) phase and a regeneration and / or replacement phase of the dehydrogenation catalyst. The purpose of regeneration is to burn off organic deposits and nitrogen- and sulfur-containing substances contained on and within the dehydrogenation catalyst. Replacement allows the replacement of a used catalyst, i.e., a catalyst used during at least one dehydrogenation step, with a fresh dehydrogenation catalyst, i.e., a catalyst that has not yet been used.

[0052] The regeneration of the dehydrogenation catalyst can be advantageously carried out by oxidizing coke and inhibitory compounds under a stream of air or an air / nitrogen mixture, for example, by using recirculation of combustion air with or without water to dilute the oxygen and control the exothermic heat of regeneration. In this case, the oxygen content is advantageously adjusted by the contribution of air at the reactor inlet. Regeneration preferably occurs at a pressure between atmospheric pressure and the reaction pressure.

[0053] According to a very particular embodiment of the invention, the regeneration-exchange stage comprises: - replacement of the dehydrogenation catalyst, in particular replacement of a used catalyst with a fresh catalyst; the catalyst preferably comprises at least elemental copper on an inorganic support, preferentially silica; or - regeneration of a dehydrogenation catalyst, preferably comprising at least three phases; at least a first phase is flushed with nitrogen, preferably at a temperature of 200-350°C, preferably 250-300°C; at least a second phase is flushed with oxygen, preferably with a gas containing nitrogen and oxygen, advantageously until no more oxygen is consumed, which is a sign of complete coke combustion, at a temperature of 300-650°C, preferably 350-600°C; and at least a third phase is flushed with nitrogen, preferably at a temperature of 200-350°C. The regeneration may optionally further comprise a redispersion phase, preferably comprising at least elemental copper and optionally elemental chromium, on an inorganic support, preferably silica.

[0054] Highly advantageously, the dehydrogenation effluent obtained at the end of the reaction section of the dehydrogenation stage comprises at least acetaldehyde, hydrogen, optionally water and unconverted ethanol, which can be sent to a separation section to optionally separate the hydrogen produced at least in part during the dehydrogenation reaction.

[0055] The dehydrogenation effluent obtained at the end of the reaction section or the effluent obtained at the end of the separation section can also be subjected, either directly or indirectly, to a treatment for separating a stream containing unconverted ethanol, which can be recycled to the reaction section and fed to the multi-tubular reactor(s), optionally in admixture with the feedstock.

[0056] Advantageously, the dehydrogenation method according to the invention can be integrated as a reaction step into more conventional methods for the conversion of ethanol. In particular, the dehydrogenation method according to the invention can be integrated into a method for the production of butadiene from ethanol as the first reaction step for the conversion of ethanol to acetaldehyde, advantageously followed by a second reaction step for the conversion of the ethanol-acetaldehyde mixture to butadiene. Such a method for the production of butadiene from ethanol in two reaction steps can be, for example, the method described in Patent FR 3 026 100. More specifically, step A) of the method described in Patent FR 3 026 100 is replaced by the dehydrogenation method described herein above, while steps B), C1), D1), D2), D3), E1), E2), and optional steps C2), D2bis), F) of the method described in Patent FR 3 026 100 remain the same.

[0057] Therefore, the present invention further relates to a process for the production of butadiene from an ethanol feedstock containing a minimum of 80 wt.% ethanol, comprising at least: A) the step of converting ethanol to acetaldehyde using the method for dehydrogenation of ethanol described above, wherein the feedstock fed to the tubes of the multi-tubular reactor is at least part of the ethanol-rich effluent advantageously obtained from step E1) to produce a dehydrogenated effluent, Optionally, a separation section is used to treat the dehydrogenation effluent to at least separate a hydrogen effluent in gas form and an ethanol / acetaldehyde effluent in liquid form; B) a stage for conversion to butadiene, comprising at least a reaction section B to which at least part or all of the dehydrogenation effluent obtained from stage A) or any ethanol / acetaldehyde effluent obtained from any separation section of stage A) is fed, optionally with an ethanol-rich liquid effluent advantageously obtained from any stage C1), and part or all of the acetaldehyde-rich effluent advantageously obtained from stage E1), and which operates in the presence of a catalyst, which preferably comprises elemental tantalum and an inorganic support, which inorganic support is Preferably containing silica, operating at a temperature of 300°C to 400°C, preferably 320°C to 370°C, and at a pressure of 0.1 to 1.0 MPa, preferably 0.1 to 0.5 MPa, and suitably 0.1 to 0.3 MPa, and adjusting the feed flow rate so that the molar ratio of ethanol to acetaldehyde at the inlet of the reaction section is 1 to 5, preferably 1 to 3.5, preferably 2 to 3, and highly suitably 2.4 to 2.7, and using a separation section to treat the effluent from the reaction section B to separate at least a gaseous effluent from a liquid effluent; C1) an optional hydrogen treatment stage, comprising at least one compression section for compressing the hydrogen effluent obtained from stage A) to a pressure of 0.1 to 1.0 MPa, advantageously 0.1 to 0.7 MPa, preferably 0.4 to 0.68 MPa, and a gas-liquid scrubbing section, to which a portion of the ethanol-rich effluent advantageously obtained from stage E1) and a portion of the ethanol / acetaldehyde effluent obtained from stage A) are fed at a temperature of 15°C to -30°C, preferably 0°C to -15°C, and to which the compressed hydrogen effluent is fed at a temperature of 25°C to 60°C, preferentially 30°C to 40°C, to produce at least an ethanol-rich liquid effluent and a purified hydrogen effluent; D1) Stage of butadiene extraction: including at least the following: (i) a compression section in which the gas effluent obtained from stage B) is compressed to a pressure of 0.1 to 1.0 MPa, preferably 0.1 to 0.7 MPa, and suitably 0.2 to 0.5 MPa, and optionally the compressed gas effluent obtained from stage B) is subsequently cooled to a temperature of 25°C to 60°C, preferentially 30°C to 40°C, (ii) a gas-liquid scrub wash section, comprising a scrub wash column, into which is fed at the top an ethanol stream consisting of the ethanol feedstock of the process and optionally a portion of the ethanol-rich effluent advantageously obtained from step E1), at a temperature between 20 and -20°C, preferentially between 15 and 5°C, and at the bottom the gaseous effluent obtained from step B) compressed and optionally cooled in section (i), to produce at least a liquid scrub wash effluent and a gaseous by-product effluent; and (iii) a distillation section, operating at a pressure of 0.1 to 1 MPa, preferably 0.2 to 0.5 MPa, to which at least the liquid effluent obtained from step B) and the liquid effluent from the gas-liquid scrubbing section are fed, producing at least a crude butadiene effluent and an ethanol / acetaldehyde / water effluent; D2) a stage for the first purification of butadiene, comprising at least a gas-liquid scrubbing section to which is fed, at the bottom, the crude butadiene effluent obtained from D1) and, at the top, a water stream which may be a water stream of external origin relative to the process for the production of butadiene and / or a portion of the aqueous effluent advantageously obtained from step E1), the water stream preferably cooled to a temperature below 25°C, preferably below 20°C, preceding the gas-liquid scrubbing section, the gas-liquid scrubbing section advantageously operating at a pressure of 0.1 to 1 MPa, which gas-liquid scrubbing section gives rise to a pre-purified butadiene effluent at the top and a wastewater effluent at the bottom; D3) a subsequent stage of butadiene purification, to which at least the pre-purified butadiene effluent obtained from stage D2) is fed to at least produce purified butadiene, said subsequent purification stage advantageously using a section for drying the pre-purified butadiene effluent obtained from stage D2), preferably in the presence of at least one adsorbent, followed by at least a cryogenic distillation section or at least a distillation and extractive distillation section; E2) a step for removing impurities and brown oil, wherein the ethanol / acetaldehyde / water effluent obtained from step D1) and at least a portion of the water-rich effluent advantageously obtained from step E1) are fed to produce at least a water / ethanol / acetaldehyde raffinate, a light brown oil effluent and a heavy brown oil effluent, Said stage of impurity and brown oil removal preferably uses at least: E2i) a scrubbing / back-scrubbing section, operating at a pressure of 0.1 to 0.5 MPa, preferentially 0.2 to 0.4 MPa, to which is fed the ethanol / acetaldehyde / water effluent obtained from stage D1), preferentially at the bottom a hydrocarbon effluent and at least a portion of the water-rich effluent advantageously obtained from stage E1) at the top, said hydrocarbon and water-rich effluents preferably being at a temperature of 10 to 70°C, preferentially 45 to 55°C, and which gives rise to said water / ethanol / acetaldehyde raffinate and hydrocarbon extract, E2ii) a section for distillation of light brown oil; feeding the hydrocarbon extract and producing said light brown oil effluent and hydrocarbon residue; and E2iii) a section for the distillation of heavy brown oil; feeding hydrocarbon residues to produce said heavy brown oil effluent and hydrocarbon distillate, which advantageously constitutes, at least in part, the hydrocarbon effluent from the scrubbing / back-scrubbing section; E1) a stage for treating the effluent; the water / ethanol / acetaldehyde raffinate advantageously obtained from stage E2) is fed, preferably using at least two distillation sections, in particular at least one section for the distillation of acetaldehyde and at least one section for the distillation of water and ethanol, said stage for treating the effluent producing at least an ethanol-rich effluent, preferably comprising at least 80% by weight of ethanol, an acetaldehyde-rich effluent, preferably comprising at least 80% by weight of acetaldehyde, and a water-rich effluent, preferably comprising at least 80% by weight of water.

[0058] In this particular embodiment of the invention, the ethanol-rich effluent obtained in step E1 and recycled as feedstock for step A) in which the dehydrogenation of ethanol takes place may conventionally contain up to 18% by weight of water.

[0059] The dehydrogenation process according to the invention is therefore considered to be particularly advantageous in this butadiene production process, insofar as the ethanol-rich effluent fed to the dehydrogenation reaction section, which usually contains up to 18% by weight of water, makes it possible to envisage the addition of steam as a thermal diluent to the feedstock of the dehydrogenation stage without premature deactivation of the dehydrogenation catalyst, the performance quality of which generally decreases when the feedstock contains more than 20% by weight of water.

[0060] The following examples illustrate the present invention but do not limit the scope of the invention.

[0061] (Example) (Example 1: Consistent with the present invention) Example 1 illustrates a dehydrogenation process consistent with the present invention.

[0062] The feedstock to be treated contains 82% by weight ethanol and 18% by weight water.

[0063] The dehydrogenation reaction is carried out in a multi-tubular reactor made of alloy steel, the tubes of which contain a fixed bed of Octolyst® 2001 catalyst sold by Evonik. The feedstock is simultaneously introduced into the tubes in gaseous form. The heat transfer fluid used is Dowtherm® A oil from Dow, which is introduced into the shell in gaseous form, in particular in saturated vapor form.

[0064] All the parameters of the reactor and operating conditions used are summarized in Table 1.

[0065] [Table 1]

[0066] The dehydrogenation effluent is collected at the reactor outlet at a flow rate of 78,346 kg / h, a temperature of approximately 277°C, and a pressure of approximately 0.29 MPa (i.e., approximately 0.6 bar, i.e., a pressure drop of approximately 0.06 MPa). At the shell outlet, a gas-Dowtherm® A oil liquid mixture is collected at 290°C.

[0067] The dehydrogenation effluent obtained is analyzed by gas chromatography and shows the following composition: 57% by weight of ethanol, 22% by weight of acetaldehyde, 18% by weight of water, 2% by weight of other compounds, in particular: ethyl acetate, acetic acid and butanol, Approximately 1% hydrogen by weight.

[0068] The performance quality of the process obtained is satisfactory, since it makes it possible to achieve a conversion of 35% by weight of ethanol with a selectivity for acetaldehyde of 92%.

[0069] (Example 2: Not in accordance with the present invention) Example 2 illustrates a method for carrying out a dehydrogenation reaction in an adiabatic reactor.

[0070] A feedstock identical to that of Example 1 is processed by the method of Example 2: it contains 18% by weight water and 82% by weight ethanol.

[0071] The same Octolyst® 2001 catalyst from Evonik is used as the dehydrogenation catalyst.

[0072] The dehydrogenation reaction is carried out in a series of 11 axial adiabatic reactors, each containing a fixed bed of dehydrogenation catalyst (Octolyst® 2001), with heat exchangers inserted between them to heat the liquid flow between each bed. The reaction unit therefore contains 11 adiabatic reactors in series and 10 heat exchangers.

[0073] A feedstock containing 82 wt. % ethanol and 18 wt. % water is introduced into the first reactor at an inlet temperature of 275°C, an inlet pressure of 0.57 MPa, and a flow rate of 78346 kg / h, which corresponds to a WWH2h of ethanol. -1 Corresponds to.

[0074] The parameters of the adiabatic reactor with axial fixed bed, the operating conditions and the degree of conversion of ethanol obtained are presented in Table 2. The outlet pressure of the 11th reactor is 0.25 MPa.

[0075] [Table 2]

[0076] At the outlet of the reaction unit, a conversion of 34% of ethanol is achieved, with a selectivity towards acetaldehyde of 92%.

[0077] (Example 3: Not in accordance with the present invention) Example 3 illustrates a method for conducting a dehydrogenation reaction in an adiabatic reactor in the presence of a thermal diluent.

[0078] The same Octolyst® 2001 catalyst from Evonik as used in the processes described in Examples 1 and 2 is used as the dehydrogenation catalyst.

[0079] A portion of the feedstock to be processed by the method of Example 3 is diluted with steam to a 60% by weight water / 40% by weight ethanol mixture. This feedstock is introduced into a series of radial adiabatic reactors. Each radial adiabatic reactor contains a fixed bed of dehydrogenation catalyst, with a heat exchanger inserted between them to heat the liquid flow between each bed. The feedstock, containing 40% by weight ethanol and 60% by weight water, is introduced into the first reactor at an inlet temperature of 275°C, an inlet pressure of 0.37 MPa, and a flow rate of 78,346 kg / h.

[0080] The parameters of the adiabatic reactor with radial fixed bed, the operating conditions and the degree of conversion of ethanol obtained are presented in Table 3. The outlet pressure of reactor No. 4 is 0.20 MPa.

[0081] [Table 3]

[0082] Dilution of ethanol by 60 wt. % with steam makes it possible to achieve a conversion of 35% of ethanol after only four adiabatic reactors.

[0083] However, under the conditions of the process of Example 3, faster deactivation of the dehydrogenation catalyst is observed than in the process of Example 2, especially with an ethanol feedstock diluted to 60 wt % water.

Claims

1. 1. A process for the dehydrogenation of ethanol to give acetaldehyde, comprising a step of dehydrogenating a feedstock comprising ethanol, said dehydrogenation step employing a reaction section comprising at least one multi-tubular reactor, said multi-tubular reactor comprising one or more tubes and a shell; the one or more tubes each contain at least one fixed bed of at least one dehydrogenation catalyst; The feedstock is fed to the one or more pipes in the form of a gas, wherein the inlet temperature of the feedstock is 240° C. or higher, the inlet pressure of the feedstock is 0.1 to 1.0 MPa, and the weight hourly space velocity (WWH) of the feedstock at the inlet is 2 to 15 h －１ and a heat transfer fluid is passed through the shell, the heat transfer fluid being introduced into the shell in gaseous form and being at least partially in liquid form at the shell outlet; introducing the heat transfer fluid into the shell at a rate such that the ratio of the weight flow rate of the heat transfer fluid at a shell inlet to the weight flow rate of the feedstock at the inlet of one or more tubes is equal to or greater than 1.0; When the heat transfer fluid is introduced into the shell, the inlet temperature of the heat transfer fluid is 260°C or more and 400°C or less, and the inlet pressure of the heat transfer fluid at this time is 0.10 MPa or more and 1.10 MPa or less; The process produces a dehydrogenation effluent comprising at least acetaldehyde, hydrogen, and unconverted ethanol. method.

2. 10. The method of claim 1, wherein the heat transfer fluid is an oil containing a eutectic mixture of organic compounds, the saturated vapor pressures being such that the difference between the saturated vapor pressures of the organic compounds of the oil at a given temperature is 50 Pa or less, or 20 Pa or less, or 10 Pa or less.

3. 3. The method of claim 1 or 2, wherein the feedstock comprises at least 50 wt% ethanol, or at least 70 wt% ethanol, or at least 80 wt% ethanol.

4. The method described in claim 3, wherein the feedstock contains water at a content of less than 50 wt%, or less than 30 wt%, or less than 20 wt%, based on the weight of water relative to the total weight of the feedstock.

5. The method of any one of claims 1 to 4, wherein the at least one multi-tubular reactor comprises a plurality of tubes.

6. The method of claim 5, wherein the at least one multi-tubular reactor comprises at least 100 tubes and less than 20,000 tubes.

7. The method according to any one of claims 1 to 6, wherein the length of one or more tubes of the multi-tubular reactor is 1 to 6 m.

8. 8. The process according to any one of claims 1 to 7, wherein the inner diameter of one or more tubes of the multi-tubular reactor is from 30.0 to 60.0 mm, or from 40.0 to 50.0 mm, and the wall thickness of the tube is from 1.5 to 5.0 mm, or from 2.0 to 4.0 mm, or from 2.2 to 3.2 mm.

9. The process of any one of claims 1 to 8, wherein the dehydrogenation catalyst comprises at least elemental copper on an inorganic support.

10. The method of claim 9, wherein the inorganic support of the dehydrogenation catalyst is silica.

11. The method of any one of claims 1 to 10, wherein the dehydrogenation catalyst is in the form of particles having an average equivalent diameter of from 0.5 to 10.0 mm, or from 1.0 to 5.0 mm.

12. 12. The method of any one of claims 1 to 11, wherein the inlet temperature of the feedstock in one or more tubes is from 240°C to 350°C, or from 250°C to 300°C, or from 260°C to 290°C.

13. 13. The method of any one of claims 1 to 12, wherein the inlet pressure of the feedstock in the one or more tubes is 0.2 to 0.5 MPa, or 0.3 to 0.4 MPa.

14. The weight hourly space velocity (WWH) of the feedstock at the inlet of the reactor is 2 to 10 h －１ The method according to any one of claims 1 to 13, wherein

15. 15. The method according to any one of claims 1 to 14, wherein the heat transfer fluid is introduced into the shell of the multi-tubular reactor at a flow rate such that the ratio of the weight flow rate of the heat transfer fluid at the shell inlet to the weight flow rate of the feedstock at the inlets of one or more tubes is 1.5 or more and 10.0 or less, or 5.0 or less, or 2.0 or less.

16. The method according to any one of claims 1 to 15, wherein the heat transfer fluid has an inlet temperature when introduced into the shell of the multi-tubular reactor of 260°C or more, or 270°C or more, or 290°C or more, and 400°C or less, or 380°C or less, and an inlet pressure of the heat transfer fluid at that time is 0.10 MPa or more, or 0.13 MPa or more, or 0.20 MPa or more, and 1.10 MPa or less, or 0.85 MPa or less.

17. 17. The process according to any one of claims 1 to 16, comprising a stage of conditioning of the heat transfer fluid, which stage comprises a sub-stage of recovery of liquid heat transfer fluid at the shell outlet of the multi-tubular reactor of the dehydrogenation stage and a stage of compression and / or heating of the heat transfer fluid to obtain a heat transfer fluid in gaseous form at the temperature and pressure of the heat transfer fluid inlet in the shell of the dehydrogenation stage.

18. 1. A process for the production of butadiene from an ethanol feedstock containing a minimum of 80 wt. % ethanol, the process comprising at least: A) a step of converting ethanol to acetaldehyde using a method for dehydrogenation of ethanol according to any one of claims 1 to 17; the feedstock fed to the tubes of a multi-tubular reactor is at least partly a part of the ethanol-rich effluent, producing a dehydrogenated effluent; and using a separation section to treat the dehydrogenation effluent to separate at least a hydrogen effluent in gas form and an ethanol / acetaldehyde effluent in liquid form; B) a stage for conversion to butadiene, comprising at least a reaction section B, to which reaction section B is fed at least part or all of the dehydrogenation effluent from step A) or the ethanol / acetaldehyde effluent from the separation section of step A), a liquid ethanol-rich effluent, a liquid acetaldehyde-rich effluent, or both, and which operates in the presence of a catalyst, the operation being carried out at a temperature of 300°C to 400°C and at a pressure of 0.1 to 1.0 MPa, the feed flow rate being adjusted so that the molar ratio of ethanol to acetaldehyde at the inlet of the reaction section is between 1 and 5, and a separation section is used to treat the effluent from reaction section B, separating at least a gaseous effluent from a liquid effluent; D1) Stage of butadiene extraction, comprising at least: (i) a compression section, in which the gaseous effluent obtained from step B) is compressed to a pressure of 0.1 to 1.0 MPa; (ii) a gas-liquid scrub washing section, comprising a scrub washing column, at the top of which is fed an ethanol stream consisting of the ethanol feedstock of the process and a portion of the ethanol-rich effluent obtained from step E1) at a temperature of 20 to -20°C, and at the bottom of which is fed the compressed gaseous effluent obtained from step B), to produce at least a liquid scrub washing effluent and a gaseous by-product effluent; and (iii) a distillation section, operating at a pressure of 0.1 to 1 MPa, to which is fed at least the liquid effluent obtained from step B) and the liquid effluent from the gas-liquid scrubbing section, and which produces at least a crude butadiene effluent and an ethanol / acetaldehyde / water effluent; D2) a stage for the first purification of butadiene, comprising at least a gas-liquid scrubbing section to which is fed at the bottom the crude butadiene effluent obtained from D1) and at the top a water stream which is a water stream of external origin to the process for the production of butadiene and / or a portion of the aqueous effluent obtained from step E1), said gas-liquid scrubbing section giving rise to a pre-purified butadiene effluent at the top and a wastewater effluent at the bottom; D3) a further stage of butadiene purification, wherein at least the pre-purified butadiene effluent obtained from step D2) is fed to produce at least a purified butadiene effluent; E2) a step for removing impurities and brown oil, wherein the ethanol / acetaldehyde / water effluent obtained from step D1) and at least a portion of the water-rich effluent obtained from step E1) are fed to produce at least a water / ethanol / acetaldehyde raffinate, a light brown oil effluent and a heavy brown oil effluent; E1) A stage of effluent treatment, in which at least the water / ethanol / acetaldehyde raffinate obtained from stage E2) is fed to produce at least an ethanol-rich effluent, an acetaldehyde-rich effluent and a water-rich effluent.

19. A method for the production of butadiene according to claim 18, comprising the steps of: C1) A stage for treating hydrogen, comprising at least a compression section for compressing the hydrogen effluent obtained from stage A) to a pressure of 0.1 to 1.0 MPa, and a gas-liquid scrubbing section, to which a portion of the ethanol-rich effluent and a portion of the ethanol / acetaldehyde effluent obtained from stage A) are fed at a temperature of 15°C to -30°C, and to which the compressed hydrogen effluent is fed at a temperature of 25°C to 60°C, to produce at least an ethanol-rich liquid effluent and a purified hydrogen effluent.

20. A method for producing butadiene according to claim 18 or 19, wherein the catalyst in step B) comprises elemental tantalum and an inorganic support, the inorganic support comprising silica.

Images