Process for catalytically dehydrogenating a cyclic alkane

Abstract

The invention relates to a process for dehydrogenating a feedstock comprising a cyclic alkane having a 6-carbon ring, in order to produce an effluent comprising hydrogen and an aromatic compound, with a reaction section comprising at least one multitubular reactor which comprises at least two tubes and a shell, said tubes comprising a fixed bed of at least one dehydrogenation catalyst, wherein - the feedstock is supplied to the tubes in gaseous form, with an inlet temperature of the feedstock being greater than or equal to 300°C, an inlet pressure of the feedstock being between 0.1 and 1 MPa, and a weight hourly space velocity (WHSV) of the feedstock at the inlet being between 1 and 15 h-1, - a heat-transfer fluid circulates in the shell, the heat-transfer fluid being introduced at the inlet of the shell in gaseous form and, at the shell outlet, being at least partly in liquid form, - the heat-transfer fluid is introduced into the shell at a flow rate such that the ratio of the flow rate by weight of the heat-transfer fluid at the shell inlet, relative to the flow rate by weight of the feedstock at the inlet of the tubes, is greater than or equal to 1.0, at an inlet temperature of between 320°C and 400°C, and at an inlet pressure of between 0.10 MPa and 1.10 MPa.

Description

[0001] Process for the catalytic dehydrogenation of cyclic alkane

[0002] Technical field

[0003] The present invention relates to a process for the catalytic dehydrogenation of cyclic alkanes, in particular with a 6-carbon cycle, for example mono- or multi-substituted.

[0004] The present invention also relates to a process for the production of hydrogen and hydrocarbons with an unsaturated / aromatic cycle, in particular with a 6-carbon cycle, for example mono- or multi-substituted, which comprises at least one reaction step of catalytic dehydrogenation of cyclic alkanes, in particular with a 6-carbon cycle, for example mono- or multi-substituted.

[0005] An example of application concerns the dehydrogenation of methylcyclohexane to produce hydrogen and toluene, and in the remainder of this text it is this example which will be chosen to illustrate the invention, for the sake of brevity, without the invention being considered as limited to it.

[0006] Prior art

[0007] The dehydrogenation reaction of methylcyclohexane into hydrogen and toluene is a well-known reaction. Indeed, the feedstocks for catalytic reforming of naphthas contain this molecule. For several years, this methylcyclohexane molecule has been considered as a potential vector for transporting hydrogen, what is called LOHC for "Liquid Organic Hydrogen Carrier". It is produced by the hydrogenation reaction of toluene, and the three molecules of dihydrogen that reacted with toluene to produce methylcyclohexane can be released by carrying out the reverse reaction, that is to say the dehydrogenation of said methylcyclohexane. This produces toluene and hydrogen.

[0008] The dehydrogenation reaction of methylcyclohexane has been extensively studied and is well known (see the journal article "A Review of Catalysts for Methylcyclohexane Dehydrogenation" Topics in Catalysis volume 64, pages 509-520 (2021)). It is a highly endothermic reaction (the reaction enthalpy difference AH is approximately 215 kJ / mol). Therefore, a reactor technology that allows adequate thermal compensation is required to achieve complete conversion of methylcyclohexane to hydrogen and toluene.

[0009] More generally, there are several industrial process technologies for the dehydrogenation of petroleum fractions containing methylcyclohexane at varying levels, but generally less than 20% by weight. The main objective of these processes is to produce gasoline or aromatic compounds for petrochemicals. The hydrogen produced is most often used in other units of the refinery or aromatic complex. These processes are characterized by the presence of several reactors in series (often three or four reactors) and by the presence of furnaces between each reactor to raise the temperature of the effluent leaving a given reactor to the required temperature at the inlet of the next reactor. The furnaces compensate for the temperature drop, a consequence of the highly endothermic nature of dehydrogenation reactions.In industrial processes for converting naphtha fractions into gasoline, temperature drops per reactor can reach 80°C per reactor. The presence of secondary paraffin cracking reactions, exothermic reactions, limits temperature drops.

[0010] In the case of dehydrogenation of a feedstock containing only methylcyclohexane, and in a process involving only one reactor, the temperature drop would be equal to nearly 900°C. Even if several reactors were placed in series, it would be necessary to compensate for temperature drops per reactor of more than 100 to 150°C. Conventional catalytic reforming technology is therefore not suitable for carrying out the dehydrogenation reaction of methylcyclohexane under good conditions.

[0011] Furthermore, multi-tubular reactors, also called exchanger reactors, can be used to enable processes to operate under isothermal or pseudo-isothermal conditions. Such reactors have been considered for the dehydrogenation reaction of methylcyclohexane, coupled with the use of heat transfer fluid to provide the calories.

[0012] For example, patents JP2018052768 A2, JP2018012610 A2, JP2018002486 A2, JP2018002487 A2, JP2017065937 A2 and JP2017039632 A2 cite the use of exchanger reactors for the dehydrogenation reaction, with the use of a heat transfer fluid, used in liquid form, to provide the calories necessary for the reaction. The disadvantage of solutions based on heat transfer fluids used in liquid form is that it is very difficult, if not impossible, to maintain isothermal operation of the reactor, unless excessively high flow rates of said fluid are used to provide the calories necessary for the reaction to achieve complete conversion of methylcyclohexane. Summary of the invention

[0013] The invention firstly relates to a process for dehydrogenating a feedstock comprising a cyclic alkane with a 6-carbon ring, in particular methylcyclohexane, in order to produce a dehydrogenation effluent comprising hydrogen and an aromatic compound, in particular toluene, said process implementing a reaction section comprising at least one multitubular reactor which comprises at least two tubes and a calender, said tubes each comprising at least one fixed bed of at least one dehydrogenation catalyst, the process being such that

[0014] - said charge feeds said tubes in gaseous form, at a charge inlet temperature greater than or equal to 300°C, preferably between 320°C and 400°C, preferentially between 320°C and 390°C, preferably between 350°C and 390°C, at an inlet pressure of said charge between 0.1 and 1 MPa, preferably between 0.1 and 0.4 MPa, preferentially between 0.1 and 0.3 MPa, and at a weight hourly space velocity (WPH) of the inlet charge of between 1 and 15 h -1 , preferably between 2 and 10 a.m. -1 ,

[0015] - a heat transfer fluid circulates in said calender, so that said heat transfer fluid is introduced at the inlet of said calender in gaseous form and is, at the outlet of the calender, at least partly in liquid form,

[0016] - said heat transfer fluid is introduced into the calender at a flow rate such that the ratio of the weight flow rate of said heat transfer fluid entering the calender relative to the weight flow rate of said charge entering the tubes is greater than or equal to 1.0

[0017] - said heat transfer fluid is introduced into the calender at an inlet temperature of the heat transfer fluid of between 320°C and 400°C, and at an inlet pressure of the heat transfer fluid of between 0.10 M Pa and 1.10 M Pa.

[0018] Note that the dehydrogenation effluent may also include, in particular, a portion of unreacted cyclic alkane with a 6-carbon cycle.

[0019] The present invention includes a cyclic alkane with a 6-carbon ring, alkyl cyclohexanes, which may be substituted on one or more of the 6 carbons of the ring, by a methyl group or by a longer, linear or branched carbon chain.

[0020] The present invention has the advantage of compensating for the endothermicity of the methylcyclohexane dehydrogenation reaction (AH reaction = 215 kJ / mol) by maintaining the temperature of the reaction medium within a temperature range suitable for the dehydrogenation reaction, in particular at a value greater than or equal to 300°C, preferably greater than or equal to 320°C, more preferably greater than or equal to 350°C.

[0021] The present invention uses a reactor of the reactor-exchanger type and according to a particular mode 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 dehydrogenation reaction of the feedstock, containing in particular methylcyclohexane. The present invention also makes it possible to have a homogeneous temperature on the outer wall of the tubes inside which the reaction is carried out.

[0022] The present invention thus makes it possible to achieve conversion rates of the cyclic alkane with a 6-carbon cycle, in particular methylcyclohexane, close to 100%, and a high selectivity in the aromatic obtained, in particular in toluene when starting from methylcyclohexane, and in hydrogen.

[0023] It is thus possible to achieve methylcyclohexane conversion rates greater than or equal to 95%, preferably of the order of 99%, and a toluene selectivity greater than 90%, preferably of the order of 98%.

[0024] Another advantage of the invention lies in the fact that thermal compensation is obtained without adding a diluent to the feedstock. However, a diluent previously appeared necessary, particularly when adiabatic reactors are used, to "thermally buffer" the temperature drop induced by the endothermicity of the reaction by transmitting its sensible heat to the reaction. The diluent conventionally used until now as a "thermal buffer" was the product of the reaction (here, for example, toluene) which can be recycled at the inlet. This recycling can lead to reductions in selectivity, because the product used as a diluent may be susceptible to parasitic reactions.Furthermore, avoiding the addition of a diluent in the feedstock advantageously makes it possible to limit the size of the equipment and the number of pieces of equipment, and, consequently, to minimize the investment and operating costs (energy costs) of the process: in fact, this avoids, in particular, the use of an additional pump and an additional furnace or heat exchanger to ensure the recirculation and possibly the reheating of the recycle. The total flow rate of fluid circulating in the reactor is not increased, which makes it possible not to increase the size of the reactor, but also that of the lines and any separators and other equipment associated with the reactor on the production line.

[0025] The present invention has yet another advantage: that of avoiding the multiplication of the number of catalytic beds required. Indeed, with adiabatic reactor technology, a succession of fixed catalyst beds, intercalated with heat exchangers, is necessary to compensate for the endothermicity of the reaction and achieve complete conversion of the cyclic alkane with a 6-carbon cycle, such as methylcyclohexane. On the contrary, the present invention, which proposes to use a multitubular reactor coupled with the particular mode of condensation of a suitable heat transfer fluid, makes it possible to achieve the targeted conversion rates, while limiting the investment and operating costs of the process.

[0026] The present invention makes it possible, in particular, to maximize the stability of the catalyst's performance over time, because the embodiment of the invention makes it possible to limit the reactor inlet temperatures, for example between 300 and 400°C, whereas in an implementation with a succession of adiabatic reactors the reactor inlet temperatures would be greater than 470°C.

[0027] The invention also relates to the installation for implementing the method described above.

[0028] The invention also relates to the dehydrogenation effluent obtained with the process or installation mentioned above, which comprises in particular hydrogen (H2) and an aromatic compound, in particular toluene, with at least 87% by weight of aromatic compounds and at least 5% by weight of hydrogen.

[0029] Description of the modes of the invention

[0030] According to the present invention, the expressions "between ... and ..." and "between ... and ..." are equivalent and mean that the limit values ​​of the interval are included in the range of values ​​described. If this were not the case and the limit values ​​were not included in the range described, such clarification will be provided in the present text.

[0031] For the purposes of the present invention, the different parameter ranges / operating conditions for a given step, such as pressure ranges and temperature ranges, may be used alone or in combination. For example, for the purposes of the present invention, a preferred pressure value range may be combined with a more preferred temperature value range.

[0032] In the following, particular and / or preferred embodiments of the invention may be described. They may be implemented separately or combined with each other, without limitation of combination when technically feasible.

[0033] The preferred embodiment of the invention concerns the dehydrogenation of methylcyclohexane into toluene and hydrogen, which will be detailed below, without the invention being limited to it.

[0034] According to a preferred embodiment of the present invention, the hydrocarbon feedstock treated in the dehydrogenation process is a feedstock comprising a cyclic alkane with a 6-carbon ring, in particular methylcyclohexane. Preferably, said feedstock comprises at least 80% by weight of cyclic alkane with a 6-carbon ring, in particular methylcyclohexane, preferably at least 85% by weight of cyclic alkane with a 6-carbon ring, in particular methylcyclohexane, preferably at least 90% by weight or at least 95% by weight of cyclic alkane with a 6-carbon ring, in particular methylcyclohexane.

[0035] The feedstock for the dehydrogenation process may further optionally comprise other hydrocarbon compounds not comprising a 6-carbon cycle, preferably at a content of less than 20% by weight, preferably less than 15% by weight, preferably less than 10% by weight or 5% by weight relative to the total weight of the feedstock.

[0036] Frequently, the feedstock of the dehydrogenation process may comprise other compounds with 7 carbon atoms without a 6-carbon cycle, such as isoalkanes (3-ethylpentane and dimethylpentanes) and alkylcyclopentanes (ethylcyclopentane and dimethylcyclopentanes), preferably at a content of less than 20% by weight, preferably less than 10% by weight, more preferably less than 5% by weight relative to the total weight of the feedstock. Among these compounds with 7 carbon atoms without a 6-carbon cycle, the feedstock of the dehydrogenation process may in particular comprise n-heptane, since its boiling point is very close to that of methylcyclohexane, preferably at a content of less than 20% by weight, preferably less than 10% by weight, more preferably less than 5% by weight relative to the total weight of the feedstock.

[0037] In addition to any compounds with 7 carbon atoms without a 6-carbon cycle, the feedstock for the dehydrogenation process may also optionally comprise compounds with 8 carbon atoms without a 6-carbon cycle, such as 2,2,4-trimethylpentane and / or dimethylhexanes, the boiling points of which are also very close to that of methylcyclohexane, preferably at a content of less than 20% by weight, preferably less than 10% by weight, more preferably less than 5% by weight relative to the total weight of the feedstock.

[0038] The feedstock treated in the process according to the invention is optionally obtained by a process for the hydrogenation of aromatic compounds, such as toluene or xylenes, and more particularly the aromatic compound such as toluene which has been produced in the dehydrogenation unit implementing the present invention.

[0039] In addition to the hydrocarbon feedstock, the process can also be fed with hydrogen to limit catalyst deactivation.

[0040] Dehydrogenation step

[0041] According to the invention, the dehydrogenation process comprises at least one step of dehydrogenation of the feedstock comprising methylcyclohexane, so as to produce a dehydrogenation effluent comprising at least toluene, hydrogen and optionally unconverted methylcyclohexane.

[0042] This dehydrogenation step uses a reaction section comprising at least one multitubular reactor in which the dehydrogenation reaction takes place. The reaction section may comprise at least two multitubular reactors, and preferably less than ten multitubular reactors.

[0043] Preferably, the reaction section comprises two multi-tubular reactors, one being in operation, i.e. supplied with feed and carrying out the dehydrogenation reaction, the other being in regeneration-replacement mode. The expression "in regeneration-replacement mode" means that the multi-tubular reactor is not supplied with feed comprising methylcyclohexane, and that the catalyst is being regenerated or loaded, or is regenerated and / or loaded and is ready to operate (i.e. waiting to operate).

[0044] Advantageously, each multitubular reactor comprises a plurality of tubes and a shell. The dehydrogenation reaction advantageously takes place in the tubes of the multitubular reactor(s), advantageously in operation. The tubes of the multitubular reactor(s) may also be referred to in this text as reaction tubes.

[0045] The calandria is the typically cylindrical envelope of the reactor inside which the tubes are located, preferably parallel to each other and parallel to the walls of the calandria when there are several tubes, and in which a heat transfer fluid circulates. The calandria may also include one or more baffles or any other system, preferably distributed homogeneously in the calandria, to allow good diffusion and homogenization of the heat transfer fluid, and therefore good heat distribution. The calandria and the tubes may have a particular design / a particular texture to promote condensation and / or evacuation of the heat transfer fluid.

[0046] According to the invention, each tube comprises at least one fixed bed comprising at least one dehydrogenation catalyst. Preferably, each tube contains a fixed bed of a dehydrogenation catalyst.

[0047] In a variant, each tube can be equipped with several catalytic beds, for example two or three beds, separated from each other by inert zones so as to have a larger exchange surface between the beds where the endothermic reaction takes place and the heat transfer fluid which is outside the tubes.

[0048] Preferably, all catalyst beds are loaded with the same catalyst, or mixture of catalysts.

[0049] Preferably, the dehydrogenation catalyst is in the form of particles with an equivalent mean diameter between 0.5 and 10.0 mm, preferably between 1.0 and 5.0 mm. According to the invention, the equivalent mean diameter defines an equivalent mean surface diameter, advantageously determined by the laser diffraction method and the equivalent mean diameter of the particles advantageously corresponds to the mean diameter of spheres having the same specific surface area as said particles.

[0050] Preferably, the dehydrogenation catalyst comprises:

[0051] - a silica and / or alumina support, preferably alumina,

[0052] - at least one noble metal from group VIII, preferably platinum Pt or palladium Pd, and preferably platinum Pt

[0053] - preferably at least one other metal chosen from tin, germanium, rhenium, lead, gallium, indium and thallium, preferably tin Sn

[0054] - and phosphorus P and / or an alkaline earth or an alkali, and preferably either phosphorus P or an alkali, and in the latter case preferably potassium K

[0055] - possibly a halogen, preferably chlorine Cl.

[0056] This may include a catalyst with a dense alumina support and with platinum and tin, such as the catalyst marketed by the company AXENS under the trade name Symphony™ CR157.

[0057] They may also be catalysts comprising an active phase based on at least one metal from group VIII, at least one element M1 chosen from tin, germanium, lead, gallium, indium and thallium, at least one element M2 chosen from alkali or alkaline-earth elements and at least one element X chosen from halogens, and a support comprising at least one refractory oxide, such as:

[0058] - the content of element M2, expressed as an element, is between 0.4 and 8% by weight relative to the total weight of the catalyst,

[0059] - the surface density of the element M2 is between 0.5 and 1.5 atoms M 2 / nm 2 ; And

[0060] - the support is a gamma alumina comprising a specific surface area between 170 and 250 m 2 / g. with M1 preferably tin, M2 preferably potassium, and the group VIII metal preferably platinum.

[0061] Details on the formulations and methods of preparation of such a catalyst can be found in patent application FR23 03931 filed on April 19, 2023.

[0062] They may also be catalysts comprising an active phase based on at least one metal from group VIII, at least one element M1 chosen from tin, germanium, lead, gallium, indium and thallium, at least one element M2 chosen from alkali or alkaline-earth elements, and at least one element X chosen from halogens, and a support comprising at least one refractory oxide, such as:

[0063] - the surface density of element X is between 0.75 and 3.5 X atoms / nm 2 ;

[0064] - the atomic ratio X / M2 is between 1.5 and 3;

[0065] - said support comprises a specific surface area of ​​between 175 and 250 m 2 / g., preferably with a gamma-crystalline structure alumina support, the group VIII metal being platinum, M1 being tin and M2 being potassium.

[0066] Details on the formulations and methods of preparation of such a catalyst can be found in patent application FR23 03932 filed on April 19, 2023.

[0067] Preferably, the multi-tubular reactor(s) comprises a plurality of tubes in the shell, preferably at least 100 tubes, preferably at least 1000 tubes, more preferably at least 2000 tubes. Generally, the multi-tubular reactors comprise up to 20,000 tubes, preferably up to 10,000 tubes. For example, each multi-tubular reactor may contain between 3,000 and 6,000 tubes.

[0068] Advantageously, the reaction tubes have a length preferably between 1.0 and 18.0 m, preferably between 2.0 and 10.0 m, more preferably between 2.0 and 6.0 m.

[0069] When the multitubular reactor comprises a plurality of tubes, all the tubes of said multitubular reactor advantageously have the same length, with the precision due in particular to the manufacture and machining of the tubes.

[0070] The internal diameter of each reaction tube is preferably between 30.0 and 60.0 mm, preferably between 40.0 and 50.0 mm.

[0071] Advantageously, the reaction tubes have a wall thickness preferably between 1.5 and 5.0 mm, preferably between 2.0 and 4.0 mm, and preferably between 2.2 and 3.2 mm.

[0072] Thus, the nominal diameter, or external diameter, of a tube can vary between 33 and 70 mm, preferably between 44 and 57 mm.

[0073] The specific dimensioning of the reaction tubes, in particular the length, the internal diameter and the wall thickness, is advantageously adapted to the pressures exerted on the tube side (i.e. inside the tubes) and on the shell side (i.e. outside the tubes), while advantageously making it possible to limit the pressure losses inside the tubes and therefore to allow the implementation of the dehydrogenation reaction at low pressures. The use of low pressures makes it possible in particular to achieve the almost complete conversion of methylcyclohexane at lower temperatures.

[0074] The size of the multitubular reactor of the dehydrogenation stage, like the diameter of the shell, can be adapted by the person skilled in the art according to his general knowledge, depending in particular on the number of tubes, their length and their diameter.

[0075] Multitubular reactors, particularly industrial ones, and in particular the tubes of said reactors, are conventionally made of a material inert to the reaction, typically steel or nickel. The tubes of the multitubular reactor(s) in the dehydrogenation process are preferably made of steel of any type, preferably alloy steel and preferably stainless steel. Very advantageously, the shell of the multitubular reactor(s), as well as any baffle possibly present in the shell, is made of the same material as the reaction tubes, preferably steel of any type, preferably alloy steel and preferably stainless steel.

[0076] According to the invention, the feedstock is introduced in gaseous form into said tubes of the or each multitubular reactor in operation, preferably at one of the ends of said reaction tubes, and preferably simultaneously into all of the reaction tubes of the multitubular reactor.

[0077] The hydrocarbon feedstock is fed to the reaction tubes at an inlet temperature of said feedstock greater than or equal to 300°C, preferably between 320°C and 400°C, preferentially between 320°C and 390°C and preferably between 350°C and 390°C, at an inlet pressure of said feedstock between 0.1 and 1 MPa, preferably between 0.1 and 0.4 MPa, preferentially between 0.1 and 0.3 MPa, and at a weight hourly space velocity (WPH) of the feedstock at the inlet of the multitubular reactor between 1 and 15 h -1 , preferably between 2 and 10 a.m. 1 , especially between 3 and 7 a.m. -1 .

[0078] Additionally, the reaction tubes can also be supplied with hydrogen to limit catalyst deactivation.

[0079] In this case, the molar ratio between the hydrogen and the hydrocarbons of the feedstock (H2 / HC ratio) at the reactor inlet can generally be between 0 and 10 mol / mol, preferably between 0 and 8 mol / mol, preferentially between 0 and 5 mol / mol and most preferably between 0 and 3 mol / mol. The hydrogen flow is mixed with the hydrocarbon feedstock before entering the reaction tubes and is therefore at the same temperature and pressure conditions as the hydrocarbon feedstock. According to the invention, the weight hourly space velocity, also called weight hourly velocity (WPH, weight per hourly weight), is defined as the ratio of the mass flow rate of the total hydrocarbon feedstock (excluding hydrogen) entering the multitubular reactor to the mass of dehydrogenation catalyst included in all the reaction tubes of said multitubular reactor.

[0080] The inlet temperature of said feedstock into the multitubular reactor(s) may advantageously be gradually increased over time, while advantageously remaining within the inlet temperature range noted above, to at least partially compensate for the deactivation of the dehydrogenation catalyst.

[0081] The dehydrogenation step of the process according to the invention may optionally implement a feedstock heating section, upstream of the reaction section. The heating of the feedstock in the optional 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 be, for example, the heat transfer fluid circulating in the calender.

[0082] The flow of the charge may be in an upward or downward mode, preferably downward, in each tube, when the furnace is, in the operating position, in a configuration such that the tubes are all oriented vertically or substantially vertically.

[0083] According to the invention, a heat transfer fluid circulates in the shell of the multitubular reactor(s) in operation, in particular between said reaction tubes, co-currently or counter-currently, preferably co-currently, with the flow circulating inside the reaction tubes. The heat transfer fluid is introduced into the shell of the multitubular reactor(s) in gaseous form and is, at the shell outlet, at least partly in liquid form (i.e. in liquid form or in the form of a gas-liquid mixture). In other words, the heat transfer fluid is advantageously introduced into the shell in the form of saturated vapor (i.e. in the vapor phase at the bubble point) and condenses at least partly in contact with the tubes inside which the dehydrogenation reaction takes place, which is an endothermic reaction.

[0084] Preferably, the heat transfer fluid is introduced into the shell of said multitubular reactor at an inlet temperature of the heat transfer fluid greater than or equal to 350°C, preferably between 375°C and 400°C, preferably less than or equal to 390°C, and / or at an inlet pressure of the heat transfer fluid greater than or equal to 0.13 MPa, preferably between 0.20 MPa and 1.10 MPa, preferably less than or equal to 0.85 MPa.

[0085] The heat input necessary to maintain the temperature in the tubes within a temperature range compatible with the dehydrogenation reaction, in particular at least equal to 300°C, preferably greater than or equal to 320°C, very preferably greater than or equal to 350°C, is therefore very advantageously ensured by the enthalpy of phase change, in particular of condensation, of the heat transfer fluid used.

[0086] The heat transfer fluid is chosen so as to be thermally stable under the operating conditions described above. The choice of the heat transfer fluid may also be guided by other constraints. The heat transfer fluid may be chosen to be inert with respect to the reactants and products of the dehydrogenation reaction (even if, in fact, it does not come into contact with them during normal operation of the furnace). Preferably, the heat transfer fluid does not induce corrosion of equipment, such as the multitubular reactor or the feed / outlet conduits with which it is equipped.

[0087] Advantageously, the heat transfer fluid may have a single boiling point at a given pressure. In particular, it is advantageously chosen so that it has a boiling point or a range of saturated vapor temperatures (which is a function of the vapor pressures) compatible with the dehydrogenation reaction and / or such that its phase change enthalpy from the gaseous state to the liquid state covers the energy requirement of the dehydrogenation reaction.

[0088] The heat transfer fluid can be a pure body or a eutectic mixture of several compounds.

[0089] Preferably, the heat transfer fluid is an oil comprising a eutectic mixture of organic compounds, preferably of two organic compounds having very advantageously close boiling points. Preferably, these organic compounds have saturated vapor pressures such that the difference between the saturated vapor pressures of the organic compounds of the oil, at a given temperature, is less than or equal to 50 Pa, preferably less than or equal to 20 Pa, more preferably less than or equal to 10 Pa. More particularly, the heat transfer fluid comprises, preferably consists of, a mixture of biphenyl and diphenyl oxide. For example, the heat transfer fluid is the oil marketed by the company DOW CHEMICAL under the name DOWTHERM™ A.

[0090] The inlet temperature and / or pressure of the heat transfer fluid in the shell may be gradually increased over time, advantageously in the inlet temperature and pressure ranges noted above, to at least partially compensate for the deactivation of the dehydrogenation catalyst.

[0091] The mass flow rate of said heat transfer fluid in the shell is advantageously adjusted so that the ratio of the mass flow rate of said heat transfer fluid in the shell relative to the mass flow rate of the charge introduced into the tubes is greater than or equal to 1.0, preferably greater than or equal to 5.0, and advantageously less than or equal to 20.0.

[0092] Under these conditions, and since the transfer coefficient on the shell side (i.e. condensation side) is much higher than the transfer coefficient within the tubes, the temperature advantageously remains constant along each of the reactor tubes and equal to the condensation temperature of the heat transfer fluid. The temperature is thus homogeneous for all the tubes and very close to the shell temperature. This is advantageous for the design of the reactor since, the temperature being homogeneous throughout the reactor, the expansion of the reactor material is the same between the tubes and between the tubes and the shell during operation of the reactor, leading to a reduction in the cost of the equipment / a simplification of its design.

[0093] Advantageously, the dehydrogenation process may comprise a step of conditioning the heat transfer fluid comprising a phase of recovering the heat transfer fluid at the outlet of the shell of the multitubular reactor of the dehydrogenation step, followed by a phase of compressing and / or heating the heat transfer fluid to obtain a heat transfer fluid in gaseous form at the temperature and pressure at which the heat transfer fluid enters the shell of the dehydrogenation step. This conditioning step may comprise other phases, intermediate or not, between recovery and compression / heating.

[0094] In such a reactor and with the particular operating conditions of the process according to the invention, in particular by using the condensation enthalpy of the heat transfer fluid introduced into the calandria at specific temperatures and pressures and at a flow rate adjusted relative to that of the feed in the reaction tubes, the dehydrogenation reaction of methylcyclohexane into toluene and hydrogen advantageously takes place under isothermal or pseudo-isothermal conditions: the temperature of the reaction medium at the reactor outlet (i.e. of the dehydrogenation effluent at the reactor outlet) is similar to the temperature at the feed inlet or has a difference of less than 30°C, preferably less than 15°C, relative to the temperature of the feed at the reactor inlet.

[0095] Advantageously, under such conditions, the dehydrogenation effluent obtained at the end of the multitubular reactor in operation has a temperature preferably greater than or equal to 300°C, preferably greater than or equal to 320°C, preferably greater than or equal to 350°C.

[0096] The specific conditions of the process according to the invention thus make it possible to achieve the desired performance. In particular, the use of a reactor-exchanger under the specific operating conditions of the invention very advantageously makes it possible to obtain methylcyclohexane conversion rates of at least 95%, preferably at least 99% and a high selectivity for toluene, in particular at least 90% by weight, preferably at least 98%.

[0097] And these performances are achieved without adding a thermal diluent which could have unfavorable consequences on the activity and selectivity of the dehydrogenation catalyst, and without multiplying the number of catalytic beds and / or reactors as in a process comprising a chain of adiabatic reactors: this limits the investment and operating costs. The heat supply method proposed by the invention also makes it possible to easily adapt the operating conditions to the possible evolution / deactivation of the catalyst.

[0098] During the dehydrogenation step, the transformation of the feedstock may be accompanied by deactivation of the dehydrogenation catalyst, for example by coking, by adsorption of inhibiting compounds and / or by sintering. The dehydrogenation catalyst may therefore advantageously undergo periodic regeneration or be replaced. Thus, in a particular embodiment of the invention, the method comprises a regeneration-replacement step. In this particular embodiment, the reaction section preferably comprises at least two multitubular reactors, to adopt an alternating mode: in this case, the multitubular reactors are used in an alternating mode, also called swing mode, in order to alternate the reaction (or operation) phases and the regeneration and / or replacement phases of said dehydrogenation catalyst.The objective of regeneration is to burn off the organic deposits contained on the surface and within the dehydrogenation catalyst. Replacement allows the spent catalyst, i.e. the one that has been used in at least one dehydrogenation step, to be replaced with "fresh" dehydrogenation catalyst, i.e. one that has not yet been used.

[0099] The regeneration of the dehydrogenation catalyst can advantageously be carried out by oxidation of the coke and the inhibiting compounds under an air flow or under an air / nitrogen mixture, for example by using recirculation of the combustion air, in order to dilute the oxygen and control the regeneration exotherm. In this case, the oxygen content is advantageously adjusted at the reactor inlet by an air make-up. The regeneration preferably takes place at a pressure between atmospheric pressure and the reaction pressure.

[0100] According to a particular embodiment of the invention, the regeneration-replacement step comprises:

[0101] - replacement of the dehydrogenation catalyst, in particular replacement of the used catalyst with fresh catalyst, or

[0102] - regeneration of the dehydrogenation catalyst.

[0103] This regeneration is conventional, for example in the field of reforming. It may comprise, for example, at least five phases, including at least a first phase of a sweep by an inert or neutral gas, preferably by nitrogen at a temperature preferably between 150 and 350°C, preferably between 200 and 300°C, thus allowing the stripping of hydrocarbons trapped in the porosity of the catalyst, at least a second phase of a sweep by a gas comprising oxygen, preferably comprising nitrogen and oxygen and preferably a mixture comprising nitrogen and oxygen whose oxygen content is less than 7% vol, thus allowing controlled combustion to be carried out in an oxidizing medium at a temperature between 300 and 650°C, preferably between 350°C and 600°C, preferably between 380°C and 530°C, advantageously until there is no more oxygen consumption, a sign of complete combustion of the coke,at least a third phase of chlorination, and more precisely of oxychlorination, intended to improve the dispersion of the active metallic phase by the simultaneous injection of a gas comprising oxygen, preferably comprising nitrogen and oxygen and preferentially a mixture comprising nitrogen and oxygen whose oxygen content is between 5 and 15% vol, and of a chlorinated compound, at a temperature between 400 and 600°C, preferably between 450°C and 550°C, preferentially between 500°C and 530°C, optionally followed by calcination using a sweep with dry air intended to improve the fixation of the active phase, dry the catalyst and fix its chlorine content to the value required for optimal catalytic performance, at least a fourth phase of a nitrogen sweep at a temperature preferably between 150 and 350°C,preferably between 200 and 300°C and at least a fifth phase of a hydrogen sweep to reduce the metallic phase present in the catalyst.,

[0104] Advantageously, the process according to the invention may comprise the separation of hydrogen from the remainder of the dehydrogenation effluent: the dehydrogenation effluent obtained at the end of the reaction section of the dehydrogenation step, which comprises at least one aromatic of the toluene type, hydrogen, and optionally the starting unconverted cyclic alkane of the methyl cyclohexane type, may be sent to a separation section to separate the hydrogen generated during the dehydrogenation reaction from the remainder of the effluent.

[0105] Examples

[0106] The following examples illustrate the invention without limiting its scope.

[0107] Example 1 illustrates the process of dehydrogenation of methylcyclohexane according to an implementation with adiabatic reactor technology.

[0108] The charge to be treated comprises at least 99.9% by weight of methylcyclohexane.

[0109] The dehydrogenation reaction is implemented with three adiabatic reactors R1, R2, R3 in series.

[0110] At the inlet of each reactor, a preheating furnace is required to raise the temperature of the reaction fluid to the desired inlet temperature. Each adiabatic reactor includes a fixed bed of catalyst marketed by AXENS under the trade name Symphony™ CR157, which is a high-density catalyst with platinum and tin.

[0111] Table 1 below summarizes all reactor parameters and operating conditions that remain constant during operation.

[0112] Table 1

[0113] Table 2 below summarizes the performance, as well as the operating conditions that had to be modified in order to maintain performance over the cycle time. Table 2 It is noted that the methylcyclohexane dehydrogenation reactors implemented with adiabatic reactor technology have a temperature drop of at least 100°C per reactor at the beginning of the cycle. As the catalyst deactivates during operation, the loss of catalyst activity is compensated by increasing the reactor inlet temperature. Thus, at the end of the cycle, the inlet temperatures had to be increased by 46°C compared to the start-up conditions, in order to maintain the conversion of methylcyclohexane at a constant value, which in this example is approximately 95.7%. A very high deactivation of the catalyst in the first reactor is also noted, illustrated by the sharp decrease in the temperature drop.

[0114] The dehydrogenation effluent is recovered at the outlet of the last reactor at a temperature between 364°C and 367°C throughout the cycle.

[0115] In this operation, we note that the selectivity towards toluene is initially very high, at 94.5% at the beginning of the cycle, but this selectivity deteriorates throughout the cycle and reaches 84.6% at the end of the cycle. As a result, hydrogen production is not constant. Indeed, at the beginning of the cycle, the hydrogen yield is 5.81% by weight, which corresponds to a hydrogen production of 1634 kg / h. At the end of the cycle, however, the hydrogen yield is only 4.78% by weight, which corresponds to a hydrogen production of 1344 kg / h.

[0116] The sharp increase in operating temperature throughout the cycle also induces a very sharp increase in losses of the hydrogen carrier (methyl cyclohexane) due to the formation of permanent gases due to cracking reactions. If these losses are limited at the start of the cycle to approximately 44 kg / h, they correspond to 3034 kg / h at the end of the cycle, or 10.78% by weight of the fresh charge.

[0117] Table 3 below gives the composition of the liquid product. Table 3

[0118] After the dehydrogenation reaction, the liquid fraction contains not only toluene and unconverted methylcyclohexane, but also almost 5% by weight of other impurities, including about 3.8% by weight of alkylcyclopentanes with 7 carbon atoms. There is also production of about 0.5% by weight of light products such as benzene and other compounds with 5 and 6 carbon atoms. Like the permanent gases, these therefore represent a loss of compounds with 7 carbon atoms.

[0119] Example 2 (in accordance with the invention)

[0120] Example 2 illustrates a dehydrogenation process according to the invention.

[0121] The charge to be treated comprises at least 99.9% by weight of methylcyclohexane.

[0122] The dehydrogenation reaction is carried out in a multi-tubular reactor made of alloy steel, the tubes of which include a fixed bed of the same catalyst as that used in Example 1. The heat transfer fluid used is the oil marketed by the company DOW under the trade name Dowtherm™ A. It is introduced into the calender in gaseous form, in particular in the form of saturated vapor.

[0123] Table 4 below summarizes all reactor parameters and operating conditions that remain constant during operation. Table 4

[0124] Table 5 below summarizes the performance, as well as the operating conditions that had to be modified in order to maintain performance over the cycle time.

[0125] Table 5

[0126] The methylcyclohexane dehydrogenation reactors according to an implementation with a multitubular reactor technology coupled with the particular mode of condensation of a suitable heat transfer fluid, make it possible to work with a virtually isothermal temperature profile. Thanks to this implementation, similar performances are obtained at much lower inlet temperatures than in a methylcyclohexane dehydrogenation process according to an implementation with adiabatic reactor technology. It is noted that, in this example according to the invention, the inlet temperatures are approximately 100°C lower than in the previous example. Nevertheless, the conversion of methylcyclohexane is also 95.7% by weight.

[0127] The dehydrogenation effluent is recovered at the reactor outlet at a temperature of approximately 377°C at the start of the cycle and 397°C at the end of the cycle. At the calandria outlet, a gas-liquid mixture of Dowtherm™ A oil is recovered.

[0128] As the catalyst deactivates during operation, the loss of catalyst activity is compensated by increasing the reactor inlet temperature. Thus, at the end of the cycle, the inlet temperatures had to be increased by 20°C compared to the start-up conditions, in order to maintain the conversion of methyl cyclohexane at a constant value. This is a much lower value than in the case of implementation in adiabatic reactors, where the inlet temperatures had to be increased by 46°C between the start and the end of the cycle. Indeed, thanks to operation at a lower temperature, coke formation, and therefore catalyst deactivation, is much lower in an isothermal implementation.

[0129] In this operation, we note that the selectivity towards toluene remains constant throughout the cycle at a value of approximately 95.4%. In comparison with an implementation in adiabatic reactors (example 1), the selectivity is therefore approximately 1 point higher at the start of the cycle, approximately 3 points higher in the middle of the cycle and approximately 11 points higher at the end of the cycle.

[0130] Thanks to the isothermal mode operation, not only constant conversion is maintained, but also constant selectivity. As a result, hydrogen production remains constant throughout the cycle at a value of approximately 1650 kg / h. Compared to implementation in adiabatic reactors, the quantity of hydrogen produced is slightly higher (by 16 kg / h) at the beginning of the cycle but a hydrogen production difference of more than 300 kg / h, or 22%, is obtained at the end of the cycle. The operation at lower temperature, as well as the low increase in operating temperature throughout the cycle, make it possible to limit losses of hydrogen carrier (methyl cyclohexane) due to the formation of permanent gases due to cracking reactions.At the start of the cycle, these losses are only 9 kg / h and they only rise to 32 kg / h (i.e. 0.11% weight of the fresh load) at the end of the cycle, which corresponds to a reduction of more than 98% compared to implementation with adiabatic reactors.

[0131] Table 6 below gives the composition of the liquid product.

[0132] Table 6

[0133] After the dehydrogenation reaction, the liquid fraction contains more than 96% by weight of unconverted toluene and methylcyclohexane, thus limiting the impurity content to less than 4% by weight, including less than 3.5% by weight of alkylcyclopentanes with 7 carbon atoms. The production of light products, such as benzene and other compounds with 5 and 6 carbon atoms, remains limited to less than 0.3% by weight relative to the fresh feedstock throughout the cycle.

[0134] At the start of the cycle, the performances obtained from the process according to the invention are similar to those obtained from the process with adiabatic reactors since these processes make it possible to achieve a conversion of methyl cyclohexane greater than 95% by weight with a selectivity to toluene of approximately 95%. On the other hand, the use of lower temperatures makes it possible to reduce the preheating of the fresh charge, to significantly reduce the deactivation of the catalyst throughout the cycle, to limit losses of light products, and above all to maintain constant hydrogen production throughout the cycle.

Claims

Claims 1. Process for the dehydrogenation of a feedstock comprising a cyclic alkane with a 6-carbon ring, in particular methylcyclohexane, in order to produce a dehydrogenation effluent comprising hydrogen and an aromatic compound, in particular toluene, said process implementing a reaction section comprising at least one multitubular reactor which comprises at least two tubes and a calender, said tubes each comprising at least one fixed bed of at least one dehydrogenation catalyst, characterized in that - said charge feeds said tubes in gaseous form, at a charge inlet temperature greater than or equal to 300°C, at an inlet pressure of said charge between 0.1 and 1 MPa, and at a weight hourly space velocity (WPH) of the input charge between 1 and 15 h -1 , preferably between 2 and 10 a.m. -1 , - a heat transfer fluid circulates in said calender, so that said heat transfer fluid is introduced at the inlet of said calender in gaseous form and is, at the outlet of the calender, at least partly in liquid form, - said heat transfer fluid is introduced into the calender at a flow rate such that the ratio of the weight flow rate of said heat transfer fluid entering the calender relative to the weight flow rate of said charge entering the tubes is greater than or equal to 1.0 - said heat transfer fluid is introduced into the calender at an inlet temperature of the heat transfer fluid of between 320°C and 400°C, and at an inlet pressure of the heat transfer fluid of between 0.10 M Pa and 1.10 M Pa.

2. Method according to the preceding claim, characterized in that the heat transfer fluid is an oil, comprising a eutectic mixture of organic compounds having saturated vapor pressures such that the difference between the saturated vapor pressures of the organic compounds of the oil at a given temperature is less than or equal to 50 Pa, preferably less than or equal to 20 Pa, more preferably less than or equal to 10 Pa.

3. Method according to one of the preceding claims, characterized in that the heat transfer fluid circulates in said calender in co-current with the circulation of the charge in the tubes of the multi-tubular furnace.

4. Method according to one of the preceding claims, characterized in that said heat transfer fluid is introduced into the calender at a flow rate such that the ratio of the weight flow rate of said heat transfer fluid entering the calender relative to the weight flow rate of said charge entering the tubes is greater than or equal to 5, and preferably less than or equal to 20.

0.

5. Method according to one of the preceding claims, characterized in that the feed comprises at least 80% by weight, in particular at least 85% or 90% or 95% by weight, of cyclic alkane with a 6-carbon cycle, in particular methylcyclohexane.

6. Method according to the preceding claim, characterized in that the feedstock also comprises hydrocarbon compounds with 7 carbon atoms without a 6-carbon cycle, in particular isoalkanes and / or alkylcyclopentanes, and / or n-heptane, preferably in a content of at most 20% or 15% by weight, in particular at most 10% by weight, and preferably at most 5% of the feedstock.

7. Process according to claim 5 or 6, characterized in that the feedstock also comprises 8-carbon hydrocarbon compounds without a 6-carbon cycle, in particular 2,2,4-trimethylpentane and / or dimethylhexanes, preferably in a content of at most 20% or 15% by weight, in particular at most 10% by weight, and preferably at most 5% of the feedstock.

8. Method according to one of the preceding claims, characterized in that all or part of the feedstock is obtained by a process of hydrogenation of aromatic compounds, in particular comprising toluene.

9. Method according to one of the preceding claims, characterized in that the tubes of the multitubular reactor are also supplied with hydrogen, preferably mixed with the feedstock prior to their supply to said tubes, in particular in a molar ratio between the hydrogen and the hydrocarbons of the feedstock of between more than 0 and at most 10 mol / mol.

10. Method according to one of the preceding claims, characterized in that said multitubular reactor comprises at least 100 tubes, preferably at least 1000 tubes and preferably at least 2000 tubes, and preferably less than 20,000 tubes, preferably less than 10,000 tubes, for example between 3,000 and 6,000 tubes.

11. Method according to one of the preceding claims, characterized in that the tubes of the multitubular reactor have: - a length between 1.0 and 18.0 m, preferably between 2.0 and 10.0 m, more preferably between 2.0 and 6.0 m, - an internal diameter between 30.0 and 60.0 mm, preferably between 40.0 and 50.0 mm, - and preferably a tube wall thickness between 1.5 and 5.0 mm, in particular between 2.0 and 4.0 mm, for example between 2.2 and 3.2 mm.

12. Method according to one of the preceding claims, characterized in that the dehydrogenation catalyst comprises - a silica and / or alumina support, preferably alumina, - at least one noble metal from group VIII, preferably platinum Pt or palladium Pd, in particular Pt - preferably at least one other metal chosen from tin, germanium, rhenium, lead, gallium, indium and thallium, preferably tin Sn - and phosphorus P and / or an alkaline earth or an alkali, and preferably either phosphorus P or an alkali, and in the latter case preferably potassium K - possibly a halogen, preferably chlorine Cl.

13. Method according to one of the preceding claims, characterized in that said method comprises a step of conditioning the heat transfer fluid comprising a sub-step of recovering the liquid heat transfer fluid at the outlet of the shell of the multitubular reactor of the dehydrogenation step followed by a step of compressing and / or heating the heat transfer fluid to obtain a heat transfer fluid in gaseous form at the temperature and pressure of the heat transfer fluid entering the shell of the dehydrogenation step.

14. Method according to one of the preceding claims, characterized in that the inlet temperature of the charge in the (or the) is between 320°C and 390°C, in particular between 350°C and 390°C.

15. Method according to one of the preceding claims, characterized in that the inlet pressure of the charge in the tube(s) is at an inlet pressure of between 0.1 and 0.4 MPa, in particular between 0.1 and 0.3 MPa.

16. Method according to one of the preceding claims, characterized in that the heat transfer fluid is introduced into the shell of said multitubular reactor at an inlet temperature of the heat transfer fluid greater than or equal to 350°C, preferably between 375°C and 400°C, preferably less than or equal to 390°C, and / or at an inlet pressure of the heat transfer fluid greater than or equal to 0.13 MPa, preferably between 0.20 MPa and 1.10 MPa, preferably less than or equal to 0.85 MPa.

17. Method according to one of the preceding claims, characterized in that the hydrogen is separated from the remainder of the dehydrogenation effluent.

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