Method for catalytic dehydrogenation of cyclic alkanes in a reactor with thermal compensation

Abstract

The invention relates to a method for dehydrogenating a feedstock comprising a cyclic alkane with 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, the tubes comprising a fixed bed of at least one dehydrogenation catalyst comprising an active phase comprising at least platinum and potassium in an amount of between 0.03% and 1% by weight of elemental potassium relative to the total weight of the catalyst, and a support comprising alumina.

Description

[0001] CATALYTIC DEHYDROGENATION PROCESS OF CYCLIC ALKANE IN A THERMALLY COMPENSATED REACTOR

[0002] technical field

[0003] The present invention relates to a process for the catalytic dehydrogenation of cyclic alkanes, in particular those with a 6-carbon ring, for example mono- or multi-substituted. The present invention also relates to a process for the production of hydrogen and unsaturated / aromatic ring hydrocarbons, in particular those with a 6-carbon ring, for example mono- or multi-substituted, which comprises at least one reaction step for the catalytic dehydrogenation of cyclic alkanes, in particular those with a 6-carbon ring, for example mono- or multi-substituted.

[0004] An example of application involves the dehydrogenation of methylcyclohexane to produce hydrogen and toluene, and in the remainder of this text this example will be chosen to illustrate the invention, for the sake of brevity, without the invention being considered as limited to this one.

[0005] Previous technique

[0006] The dehydrogenation of methylcyclohexane to hydrogen and toluene is a well-known reaction. Indeed, the feedstocks for the catalytic reforming of naphthas contain this molecule. For several years, this methylcyclohexane molecule has been considered a potential hydrogen transport carrier, known by the acronym LOHC for "Liquid Organic Hydrogen Carrier." It is produced by the hydrogenation of toluene, and the three dihydrogen molecules that reacted with toluene to form methylcyclohexane can be released by the reverse reaction, that is, the dehydrogenation of methylcyclohexane. This produces toluene and hydrogen.

[0007] The dehydrogenation reaction of methylcyclohexane has been extensively studied and is well understood (see the review article "A Review of Catalysts for Methylcyclohexane Dehydrogenation," Topics in Catalysis, volume 64, pages 509-520 (2021)). It is a highly endothermic reaction (the enthalpy change AH ​​of reaction is approximately 215 kJ / mol). Therefore, reactor technology that provides adequate thermal compensation is necessary to achieve complete conversion of methylcyclohexane to hydrogen and toluene. For example, catalytic reforming is a refining process used to valorize heavy gasoline obtained by distillation. The hydrocarbons in the heavy gasoline feedstock (paraffins and cyclic alkanes), which generally contain approximately 5 to 12 carbon atoms per molecule, are transformed during this process into aromatic hydrocarbons or, failing that, into branched paraffins.The hydrogen produced during the reforming process is most often used in other units of the refinery or aromatics complex. The main reactions involved in catalytic reforming are the dehydrogenation of cyclic alkanes, the dehydrocyclization of paraffins, and the isomerization of paraffins and cyclic alkanes. This process is generally characterized by the presence of several reactors in series, often three or four, and by the presence of furnaces between each reactor to raise the temperature of the effluent leaving a given reactor to the required inlet temperature of the next reactor. This is referred to as an adiabatic reactor because the reforming reaction is carried out without any heat transfer occurring between the reactors and their external environment.In such a process, the furnaces compensate for the temperature drop resulting from the highly endothermic nature of the dehydrogenation reactions. The presence of secondary paraffin cracking reactions, which are exothermic, further limits the temperature drop.

[0008] However, in the case of dehydrogenating a feedstock containing primarily methylcyclohexane, and in a process with only one reactor, the temperature drop would be nearly 900°C. Thus, even with several reactors connected in series, it would be necessary to compensate for temperature drops of around one hundred degrees per reactor.

[0009] Furthermore, the nature of the catalyst impacts the type of catalytic reactions. In the case of the catalytic reforming process, the catalysts used are bifunctional, meaning they consist of two functional groups: a metallic group and an acidic group. Each group plays a well-defined role in the catalyst's activity. The metallic group primarily ensures the dehydrogenation of cyclic alkanes and paraffins, and the hydrogenation of coke precursors. The acidic group ensures the isomerization of cyclic alkanes and paraffins, and the cyclization of paraffins. The acidic group is provided by the support itself, most often pure halogenated alumina. The metallic group is provided by a noble metal from the platinum group and at least one additional metal, mainly tin for the continuous process (moving bed), and rhenium in the semi-regenerative process (fixed bed).Thus, the use of a conventional catalytic reforming catalyst is not optimal in the context of a dehydrogenation process of a feed consisting mainly of methylcyclohexane, because depending on the content of impurities included in the feed to be treated containing methylcyclohexane, so-called "parasitic" reactions may take place, which may lead to the formation of coke deactivating the catalyst, and therefore lead to a drop in the conversion and / or selectivity in toluene, and ultimately a decrease in hydrogen production.

[0010] Conventional catalytic reforming technology is therefore not suitable for carrying out the dehydrogenation reaction of methylcyclohexane under good conditions.

[0011] Furthermore, multitubular reactors, also known as heat exchanger reactors, can be used to enable process operation under isothermal or pseudo-isothermal conditions. Such reactors have been considered for the dehydrogenation reaction of methylcyclohexane, coupled with the use of a heat transfer fluid to supply the necessary heat. For example, patents JP2018052768 A2, JP2018012610 A2, JP2018002486 A2, JP2018002487 A2, JP2017065937 A2, and JP2017039632 A2 describe the use of heat exchanger reactors for the dehydrogenation reaction, employing a liquid heat transfer fluid to provide the heat required for the reaction.

[0012] The disadvantage of solutions based on heat transfer fluids used in liquid form is that it is very difficult, if not impossible, to maintain the most isothermal operation possible of the reactor, unless excessively large flow rates of said fluid are used to provide the calories necessary for the reaction to achieve complete conversion of methylcyclohexane.

[0013] Continuing its research in the field of LOHC, the Applicant has developed a new process dedicated to the dehydrogenation of a feed consisting mainly of cyclic alkanes with a 6-carbon ring, in particular methylcyclohexane, in a system with thermal compensation in the presence of a specific catalyst comprising potassium at a content within a specific range of values.

[0014] Summary of the invention

[0015] The invention relates firstly to a process for dehydrogenating a feed 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 employing a reaction section comprising at least one multitubular reactor which includes at least two tubes and a shell, said tubes each comprising at least one fixed bed of at least one dehydrogenation catalyst, the process being characterized in that:

[0016] - said charge supplies said tubes in gaseous form, at an inlet temperature of the charge greater than or equal to 300°C, preferably between 320°C and 400°C, preferably 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, preferably between 0.1 and 0.3 MPa, and at a weight-per-hour spatial velocity (PPH) of the inlet charge between 0.5 and 15 h -1 preferably between 2 and 10 am -1 ;

[0017] - a heat transfer fluid circulates in said calender, so that said heat transfer fluid is introduced into the inlet of said calender in gaseous form and is, at the outlet of the calender, at least partly in liquid form;

[0018] - said heat transfer fluid is introduced into the shell at a flow rate such that the ratio of the weight flow rate of said heat transfer fluid at the inlet of the shell to the weight flow rate of said charge at the inlet of the tubes is greater than or equal to 1.0;

[0019] - said heat transfer fluid is introduced into the shell at an inlet temperature of the heat transfer fluid between 300°C and 400°C, and at an inlet pressure of the heat transfer fluid between 0.10 MPa and 1.10 MPa;

[0020] - said dehydrogenation catalyst comprises an active phase comprising at least platinum, potassium at a content of between 0.03% and 3% by weight of element potassium relative to the total weight of the catalyst, and a support comprising alumina.

[0021] Note that the dehydrogenation effluent may also include, in particular, a portion of unreacted 6-carbon ring cyclic alkane.

[0022] The present invention includes a cyclic alkane with a 6-carbon ring, the alkyl cyclohexanes, which can be substituted on one or more of the 6 carbons of the ring, by a methyl group or by a longer carbon chain, linear or branched.

[0023] The present invention has the advantage of compensating for the endothermicity of the methylcyclohexane dehydrogenation reaction (AH reaction = 215 kJ / mol) by maintaining the reaction medium temperature within a 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, and most preferably greater than or equal to 350°C. The present invention employs a reactor-heat exchanger type reactor and a specific method of condensing the heat transfer fluid. The enthalpy of phase change of the heat transfer fluid can thus be used to provide the heat required for the energy demands of the feedstock dehydrogenation reaction, which includes methylcyclohexane. The present invention also ensures a homogeneous temperature on the outer walls of the tubes within which the reaction takes place.

[0024] The present invention thus makes it possible to achieve conversion rates of the cyclic alkane with a 6-carbon ring, in particular methylcyclohexane, close to 100%, and a high selectivity in aromatic obtained, in particular in toluene when starting from methylcyclohexane, and in hydrogen.

[0025] This allows us to achieve methylcyclohexane conversion rates of 95% or higher, preferably around 99%, and toluene selectivity of 90% or higher, preferably around 98%.

[0026] Another advantage of the invention lies in the fact that thermal compensation is achieved without adding a diluent to the feed. However, a diluent had previously been considered necessary, particularly when using adiabatic reactors, to "thermally buffer" the temperature drop induced by the endothermic nature of the reaction by transferring its sensible heat to the reaction. The diluent traditionally used as a "thermal buffer" was the reaction product (here, for example, toluene), which can be recycled at the inlet. This recycling can lead to a decrease in selectivity, as the product used as a diluent may be susceptible to side reactions.Furthermore, avoiding the addition of a diluent to the feedstock significantly reduces the size and number of equipment required, thereby minimizing the investment and operating costs (energy costs) of the process. This eliminates the need for an additional pump, furnace, or heat exchanger to recirculate and potentially reheat the recycled material. The total fluid flow rate in the reactor remains unchanged, thus avoiding an increase in the reactor size, as well as the size of the lines, separators, and other associated equipment on the production line.

[0027] The present invention offers yet another advantage: it avoids the need for multiple catalytic beds. Indeed, with adiabatic reactor technology, a succession of fixed catalyst beds, interspersed with heat exchangers, is required to compensate for the endothermic nature of the reaction and achieve complete conversion of the 6-carbon cyclic alkane, such as methylcyclohexane. In contrast, the present invention, which proposes using a multitubular reactor coupled to a specific condensation method for a suitable heat transfer fluid, makes it possible to achieve the desired conversion rates while limiting the investment and operating costs of the process.

[0028] 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.

[0029] The invention also relates to the installation for implementing the process described above.

[0030] The invention also relates to the dehydrogenation effluent obtained with the process or installation mentioned above, which includes 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.

[0031] According to one or more embodiments of the invention, the charge comprises at least 80% by weight of a cyclic alkane with a 6-carbon ring.

[0032] According to one or more embodiments of the invention, said charge comprises at least 80% by weight of methylcyclohexane.

[0033] According to one or more embodiments of the invention, the filler also comprises 7-carbon hydrocarbon compounds without a 6-carbon ring, 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 filler. According to one or more embodiments of the invention, the filler also comprises 8-carbon hydrocarbon compounds without a 6-carbon ring, 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 filler.

[0034] According to one or more embodiments of the invention, 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 at the calender inlet to the weight flow rate of said charge at the tube inlets is greater than or equal to 5, and preferably less than or equal to 20.0.

[0035] According to one or more embodiments of the invention, the tubes of the multitubular reactor are also supplied with hydrogen, preferably mixed with the feed prior to their supply to said tubes, in particular in a molar ratio between hydrogen and hydrocarbons of the feed of more than 0 and at most 10 mol / mol.

[0036] According to one or more embodiments of the invention, said multitubular reactor comprises at least 100 tubes, preferably at least 1000 tubes, preferably at least 2000 tubes.

[0037] According to one or more embodiments of the invention, the platinum content is between 0.02 and 2% by weight of platinum element relative to the total weight of the catalyst.

[0038] According to one or more embodiments of the invention, said catalyst comprises an element M1 selected from tin, germanium, lead, gallium, indium and thallium.

[0039] According to one or more embodiments of the invention, the content of element M1 is between 0.01 and 10% weight by weight of element M1 relative to the total weight of the catalyst.

[0040] According to one or more embodiments of the invention, said catalyst comprises a halogen element selected from fluorine, chlorine, bromine and iodine.

[0041] According to one or more embodiments of the invention, the halogen content is between 0.75 and 5.5% by weight of the halogen element relative to the total weight of the catalyst. According to one or more embodiments of the invention, said catalyst comprises phosphorus, at a content of between 0.1 and 1% by weight of the phosphorus element relative to the total weight of the catalyst.

[0042] According to one or more embodiments of the invention, said alumina support comprises a specific surface area of ​​between 170 m 2 / g and 250 m 2 / g.

[0043] According to one or more embodiments of the invention, all or part of the charge is obtained by a hydrogenation process of aromatic compounds, including toluene.

[0044] According to one or more embodiments of the invention, hydrogen is separated from the rest of the dehydrogenation effluent.

[0045] Detailed description of the invention

[0046] 1. Definitions

[0047] The present invention includes a cyclic alkane with a 6-carbon ring, the alkyl cyclohexanes, which can be substituted on one or more of the 6 carbons of the ring, by a methyl group or by a longer carbon chain, linear or branched.

[0048] In what follows, the groups of chemical elements are given according to the CAS classification (CRC Handbook of Chemistry and Physics, publisher CRC press, editor-in-chief DR Lide, 81 ème edition, 2000-2001). For example, group VIII according to the CAS classification corresponds to the metals of columns 8, 9 and 10 according to the new IIIPAC classification.

[0049] Specific surface area refers to the specific surface area BET (SBET in m²) 2 / g) determined by nitrogen adsorption in accordance with ASTM D 3663-78 established from the BRUNAUER-EMMETT-TELLER method described in the periodical "The Journal of American Society", 1938, 60, 309.

[0050] The total pore volume of the catalyst or support used for catalyst preparation is the volume measured by intrusion with a mercury porosimeter according to ASTM D4284 at a maximum pressure of 4000 bars (400 MPa), using a surface tension of 484 dyne / cm and a contact angle of 140°, for example with an Autopore III model device from Microméritics®.

[0051] The wetting angle was set at 140° following the recommendations of the book "Techniques de l'ingénieur, traité analyse et caractérisation", pages 1050-1055, written by Jean Charpin and Bernard Rasneur. To obtain greater accuracy, the total pore volume value corresponds to the total pore volume value measured by mercury porosimeter intrusion on the sample at 4000 bar (400 MPa) minus the total pore volume value measured by mercury porosimeter intrusion on the same sample at a pressure of 2 bar (0.2 MPa).

[0052] The packed filling density (PHD) is measured as described in the book "Applied Heating" by G. Martino, J. Miquel, R. Montarnal, A. Sugier, and H. Van Landeghem, Technip, Paris, 1987, chapter 6.2.4, pages 167-168. A graduated cylinder of acceptable dimensions is filled by successive additions, and between each addition, the catalyst is compacted by shaking the cylinder until a constant volume is reached. This measurement is generally performed over 1000 cm³. 3 of catalyst or support material packed into a cylinder with a height-to-diameter ratio close to 5:1. The apparent density of the packed product is calculated by dividing the introduced mass by the volume occupied after packing. The measurement uncertainty is generally on the order of ± 0.01 g / mL. This measurement can preferably be performed on automated devices such as Autotap®, marketed by Quantachrome®.

[0053] The levels of metals, halogens, phosphorus, and potassium are measured by X-ray fluorescence.

[0054] In this description, the term "include" is synonymous with (means the same as) "include" and "contain," and is inclusive or open-ended and does not exclude other unstated elements. It is understood that the term "include" includes the exclusive and closed term "consist." Furthermore, in this description, a load or effluent consisting essentially, substantially, or exclusively of compound A is defined as a load or effluent consisting of at least 80% by weight, preferably at least 85% by weight, preferably 90% by weight, most preferably 95% by weight, or even 100% by weight of compound A relative to the total weight of the load or effluent.

[0055] In the sense of the present invention, the different parameter ranges / operating conditions for a given step, such as pressure ranges and temperature ranges, can be used alone or in combination. For example, in the sense of the present invention, a preferred range of pressure values ​​can be combined with a more preferred range of temperature values.

[0056] The preferred mode of the invention relates to the dehydrogenation of methylcyclohexane into toluene and hydrogen, and this is the mode which will be detailed below, without the invention being limited to it.

[0057] 2. Charge

[0058] According to a preferred embodiment of the present invention, the hydrocarbon feed treated in the dehydrogenation process is a feed comprising a cyclic alkane with a 6-carbon ring, in particular methylcyclohexane. Preferably, said feed comprises at least 80% by weight of a cyclic alkane with a 6-carbon ring, in particular methylcyclohexane, preferably at least 85% by weight of a cyclic alkane with a 6-carbon ring, in particular methylcyclohexane, and preferably at least 90% by weight or at least 95% by weight of a cyclic alkane with a 6-carbon ring, in particular methylcyclohexane.

[0059] The feed for the dehydrogenation process may also optionally include other hydrocarbon compounds not comprising a 6-carbon ring, 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 feed.

[0060] Frequently, the dehydrogenation process feed may include other 7-carbon compounds without a 6-carbon ring, such as isoalkanes (3-ethylpentane and dimethylpentanes) and alkylcyclopentanes (ethylcyclopentane and dimethylcyclopentanes), and / or n-heptane, as its boiling point is very close to that of methylcyclohexane. The 7-carbon hydrocarbon compounds without a 6-carbon ring are present in the feed preferably at a concentration of at most 20% by weight, at most 15% by weight, preferably at most 10% by weight, and particularly preferably at most 5% by weight relative to the total weight of the feed.

[0061] In addition to possible 7-carbon compounds without a 6-carbon ring, the dehydrogenation process feed may also optionally include 8-carbon compounds without a 6-carbon ring, such as 2,2,4-trimethylpentane and / or dimethylhexanes, whose boiling points are also very close to that of methylcyclohexane, preferably at a content of at most 20% by weight, at most 15% by weight, preferably at most 10% by weight, and particularly preferably at most 5% by weight relative to the total weight of the feed.

[0062] The feed treated in the process according to the invention is preferably obtained by a process of 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.

[0063] 3. Dehydrogenation stage

[0064] According to the invention, the present invention relates to a process for dehydrogenating a feed comprising essentially a cyclic alkane with a 6-carbon ring, in particular methylcyclohexane, so as to produce a dehydrogenation effluent comprising an aromatic compound, in particular toluene, hydrogen and optionally unconverted methylcyclohexane.

[0065] This dehydrogenation step employs a reaction section comprising at least one multitubular reactor in which the dehydrogenation reaction takes place. The reaction section may include at least two multitubular reactors, and preferably fewer than five.

[0066] Preferably, the reaction section comprises two multitubular reactors, one operating (i.e., fed with feedstock and carrying out the dehydrogenation reaction) and the other in regeneration-replacement mode. The term "regeneration-replacement mode" means that the multitubular reactor is not fed with feedstock including methylcyclohexane, and that the catalyst is being regenerated or loaded, or is regenerated and / or loaded and is ready for operation (i.e., awaiting operation).

[0067] Advantageously, each multitubular reactor comprises a plurality of tubes and a shell. The dehydrogenation reaction advantageously takes place within the tubes of the multitubular reactor(s) during operation. The tubes of the multitubular reactor(s) may also be referred to in this text as reaction tubes. The shell is the typically cylindrical casing of the reactor, inside which the tubes are located, preferably parallel to each other and parallel to the shell walls when there are several tubes, and through which a heat transfer fluid circulates. The shell may also include one or more baffles or other systems, preferably distributed homogeneously within the shell, to allow for good diffusion and homogenization of the heat transfer fluid, and thus good heat distribution.The grille and tubes may have a particular design / a particular texture to promote condensation and / or drainage of the heat transfer fluid.

[0068] According to the invention, each tube comprises at least one fixed bed including at least one dehydrogenation catalyst. Preferably, each tube contains a fixed bed of a dehydrogenation catalyst.

[0069] In one 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.

[0070] Preferably, all catalytic beds are loaded with the same catalyst, or the same mixture of catalysts.

[0071] The nature and form of the catalyst is described in a dedicated paragraph below.

[0072] Preferably, the multitubular reactor(s) comprises a plurality of tubes in the shell, preferably at least 100 tubes, preferably at least 1000 tubes, and preferably at least 2000 tubes. Generally, multitubular reactors comprise up to 30,000 tubes, preferably up to 10,000 tubes. For example, each multitubular reactor can contain between 3,000 and 10,000 tubes.

[0073] Advantageously, the reaction tubes preferably have a length between 1.0 and 18.0 m, more preferably between 2.0 and 10.0 m, and more preferably between 2.0 and 6.0 m.

[0074] When the multitubular reactor comprises a plurality of tubes, all the tubes of said multitubular reactor advantageously have the same length, within the precision due in particular to the manufacture and machining of the tubes.

[0075] The internal diameter of each reaction tube is preferably between 30.0 and 60.0 mm, preferably between 40.0 and 50.0 mm. 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.

[0076] Thus, the nominal diameter, or external diameter, of a tube can vary between 33 and 70 mm, preferably between 44 and 57 mm.

[0077] The specific dimensioning of the reaction tubes, in particular the length, internal diameter and 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 limiting the pressure losses inside the tubes and thus allowing the implementation of the dehydrogenation reaction at low pressures.

[0078] The use of low pressures makes it possible to achieve almost complete conversion of methylcyclohexane at lower temperatures.

[0079] The size of the multitubular reactor in the dehydrogenation stage, such as the diameter of the shell, can be adapted by a person skilled in the art according to their general knowledge, depending in particular on the number of tubes, their length and their diameter.

[0080] Industrial multitubular reactors, 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 any type of steel, preferably alloy steel, and most preferably stainless steel. Advantageously, the shell of the multitubular reactor(s), as well as any baffles present in the shell, is made of the same material as the reaction tubes, preferably any type of steel, preferably alloy steel, and most preferably stainless steel.

[0081] According to the invention, the charge 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 the reaction tubes of the multitubular reactor.

[0082] The hydrocarbon feed feeds the reaction tubes at an inlet temperature of said feed greater than or equal to 300°C, preferably between 320°C and 400°C, preferably between 320°C and 390°C and preferably between 350°C and 390°C, at an inlet pressure of said feed between 0.1 and 1 MPa, preferably between 0.1 and 0.4 MPa, preferably between 0.1 and 0.3 MPa, and at a weight-hour spatial velocity (PPH) of the feed inlet to the multitubular reactor between 0.5 and 15 h' 1 preferably between 2 and 10 am 1 , particularly between 3 and 7 am 1 .

[0083] In addition, the reaction tubes can also be supplied with hydrogen to limit catalyst deactivation.

[0084] In this case, the molar ratio of hydrogen to hydrocarbons in the feed (H2 / HC ratio) at the reactor inlet can generally be between 0 and 10 mol / mol, preferably between 0 and 8 mol / mol, preferably between 0 and 5 mol / mol, and most preferably between 0 and 3 mol / mol. The hydrogen stream is mixed with the hydrocarbon feed before entering the reaction tubes and is therefore under the same temperature and pressure conditions as the hydrocarbon feed.

[0085] According to the invention, the hourly spatial velocity in weight, also called hourly weight-per-weight velocity (PPH, weight per hour), is defined as the ratio of the mass flow rate of the total hydrocarbon feed (excluding hydrogen) entering the multitubular reactor to the mass of dehydrogenation catalyst included in all the reaction tubes of said multitubular reactor.

[0086] The inlet temperature of said charge in the multitubular reactor(s) can advantageously be gradually increased over time, while remaining advantageously within the range of inlet temperatures noted above, to compensate at least in part for the deactivation of the dehydrogenation catalyst.

[0087] The dehydrogenation step of the process according to the invention may optionally incorporate a feed heating section upstream of the reaction section. The feed in this optional heating section can be heated by any method known to those skilled in the art, for example, by heat exchange with a fluid, which could be, for instance, the heat transfer fluid circulating in the calender, in a heat exchanger, or in a preheating furnace.

[0088] The charge flow can be upward or downward, preferably downward, in each tube, when the furnace is in its operating position, such that the tubes are all oriented vertically or substantially vertically. According to the invention, a heat transfer fluid circulates in the shell of the operating multitubular reactor(s), particularly 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 partially in liquid form (i.e., in liquid form or as 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 vapor phase at the bubble point) and condenses at least partially in contact with the tubes inside which the dehydrogenation reaction takes place, which is an endothermic reaction.

[0089] Preferably, the heat transfer fluid is introduced into the shell of said multitubular reactor at a fluid inlet temperature between 300°C and 400°C, preferably between 375°C and 400°C, preferably less than or equal to 390°C, and / or at a heat transfer fluid inlet pressure 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.

[0090] The heat input required 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, most preferably greater than or equal to 350°C, is therefore very advantageously ensured by the enthalpy of phase change, in particular condensation, of the heat transfer fluid used.

[0091] The heat transfer fluid is chosen to be thermally stable under the operating conditions described above. The choice of 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 furnace operation). Preferably, the heat transfer fluid does not induce corrosion of equipment, such as the multitubular reactor or the feed / outlet ducts with which it is fitted.

[0092] Advantageously, the heat transfer fluid can 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 vapor pressures) compatible with the dehydrogenation reaction and / or such that its enthalpy of phase change from the gaseous to the liquid state covers the energy requirement of the dehydrogenation reaction.

[0093] The heat transfer fluid can be a pure substance or a eutectic mixture of several compounds.

[0094] Preferably, the heat transfer fluid is an oil comprising a eutectic mixture of organic compounds, preferably two organic compounds having very advantageously similar boiling points. Preferably, these organic compounds have saturated vapor pressures such that the difference between the saturated vapor pressures of the organic compounds in the oil, at a given temperature, is less than or equal to 50 Pa, preferably less than or equal to 20 Pa, and preferably less than or equal to 10 Pa.

[0095] More specifically, the heat transfer fluid comprises a mixture of biphenyl and diphenyl oxide. For example, the heat transfer fluid is the oil marketed by DOW CHEMICAL under the name DOWTHERM™ A.

[0096] The temperature and / or inlet pressure of the heat transfer fluid in the grille can be gradually increased over time, advantageously within the temperature and inlet pressure ranges noted above, to compensate at least in part for the deactivation of the dehydrogenation catalyst.

[0097] 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 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.

[0098] Under these conditions, and because the heat transfer coefficient on the shell side (i.e., the condensation side) is much higher than the heat transfer coefficient within the tubes, the temperature remains advantageously 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 reactor design because, with 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 reactor operation, leading to a reduction in equipment cost and a simplification of its design.Advantageously, the dehydrogenation process may include a heat transfer fluid conditioning step comprising a phase of recovering the heat transfer fluid from the shell outlet of the multitubular reactor of the dehydrogenation step, followed by a phase of compressing and / or heating the heat transfer fluid to obtain a gaseous heat transfer fluid at the temperature and pressure at which the heat transfer fluid inlet the shell of the dehydrogenation step. This conditioning step may include other phases, intermediate or not, between recovery and compression / heating.

[0099] In such a reactor and with the particular operating conditions of the process according to the invention, in particular by using the enthalpy of condensation of the heat transfer fluid introduced into the calender 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 takes place advantageously 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.

[0100] Advantageously, under such conditions, the dehydrogenation effluent obtained from the operating multitubular reactor 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.

[0101] 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 makes it very advantageous to obtain methylcyclohexane conversion rates of at least 95%, preferably at least 99%, and a high toluene selectivity, in particular at least 90% wt%, preferably at least 98%.

[0102] These performance levels are achieved without the addition of a thermal diluent, which could adversely affect the activity and selectivity of the dehydrogenation catalyst, and without increasing the number of catalytic beds and / or reactors as in a process involving a series of adiabatic reactors: this limits investment and operating costs. The heat supply method proposed by the invention also allows for easy adaptation of the operating conditions to the potential evolution or deactivation of the catalyst.

[0103] During the dehydrogenation step, the transformation of the feedstock may be accompanied by the deactivation of the dehydrogenation catalyst, for example by coking, adsorption of inhibitory compounds, and / or sintering. The dehydrogenation catalyst can therefore advantageously undergo periodic regeneration or replacement. Thus, in a particular embodiment of the invention, the process includes a regeneration-replacement step.

[0104] 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, to alternate the reaction (or operating) phases and the regeneration and / or replacement phases of the 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, that is, the one that has been used in at least one dehydrogenation step, to be replaced with "fresh" dehydrogenation catalyst, that is, one that has not yet been used.

[0105] The regeneration of the dehydrogenation catalyst can advantageously be achieved by oxidizing the coke and inhibitor compounds under an air stream or an air / nitrogen mixture, for example, by using combustion air recirculation, 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 supplementing the air supply. Regeneration preferably takes place at a pressure between atmospheric pressure and the reaction pressure.

[0106] According to a particular embodiment of the invention, the regeneration-replacement step comprises:

[0107] - the replacement of the dehydrogenation catalyst, in particular the replacement of the worn catalyst with a fresh catalyst, or

[0108] - the regeneration of the dehydrogenation catalyst. This regeneration is conventional, for example in the field of reforming. It may include, for example, at least five phases, including at least a first phase of purging with an inert or neutral gas, preferably nitrogen, at a temperature preferably between 150 and 350°C, preferably between 200 and 300°C, thus enabling the stripping of hydrocarbons trapped in the porosity of the catalyst; at least a second phase of purging with a gas comprising oxygen, preferably comprising nitrogen and oxygen, and preferably a mixture comprising nitrogen and oxygen with an oxygen content of less than 7% vol, thus enabling controlled combustion in an oxidizing environment 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;sign of complete coke combustion, at least a third chlorination phase, and more specifically oxychlorination, intended to improve the dispersion of the metallic active phase by the simultaneous injection of a gas comprising oxygen, preferably comprising nitrogen and oxygen and preferably a mixture comprising nitrogen and oxygen with an oxygen content between 5 and 15% vol, and a chlorinated compound, at a temperature between 400 and 600°C, preferably between 450°C and 550°C, preferably between 500°C and 530°C, possibly followed by calcination using a dry air purge 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 nitrogen purge phase at a temperature preferably between 150 and 350°C,preferably between 200 and 300°C and at least a fifth phase of hydrogen scavenging to reduce the metallic phase present in the catalyst.

[0109] Advantageously, the process according to the invention can include the separation of hydrogen from the rest of the dehydrogenation effluent: the dehydrogenation effluent obtained at the end of the reaction section of the dehydrogenation step, which includes at least one aromatic of the toluene type, hydrogen, and optionally the starting unconverted cyclic alkane of the methyl cyclohexane type, can be sent to a separation section to separate the hydrogen generated during the dehydrogenation reaction from the rest of the effluent.

[0110] 4. Catalyst

[0111] The catalyst used in the process according to the invention comprises, preferably, at least one platinum-based active phase, and a support comprising alumina, said catalyst further comprising potassium at a content of between 0.03% and 3% by weight of element potassium relative to the total weight of the catalyst.

[0112] In one embodiment, the catalyst comprises between 0.03% and 3% by weight of potassium relative to the total weight of the catalyst, preferably between 0.03% and 2.5% by weight, preferably between 0.03% and 2% by weight, more preferably between 0.03% and 1.8% by weight, even more preferably between 0.03% and 1.7% by weight, and even more preferably between 0.03% and 1.6% by weight.

[0113] Advantageously, the platinum content, expressed as an element, is between 0.02 and 2% by weight relative to the total weight of the catalyst, preferably between 0.05 and 1.5% by weight, even more preferably between 0.1 and 0.8% by weight, and even more preferably between 0.15 and 0.5% by weight.

[0114] Platinum can be present in the final catalyst as an oxide, sulfide, halide, oxyhalide, in chemical combination with one or more of the other components of the catalyst, or as an elemental metal.

[0115] Preferably, the active phase of the catalyst may also include at least one M1 element selected from tin, germanium, lead, gallium, indium and thallium.

[0116] Advantageously, the content of element M1, selected from tin, germanium, lead, gallium, indium and thallium, expressed as an element, in the catalyst according to the invention is between 0.01 and 10% by weight relative to the total weight of the catalyst, more preferably between 0.05 and 5% by weight, and very preferably between 0.1 and 1% by weight, and even more preferably between 0.15 and 0.5% by weight.

[0117] Preferably, the M1 element is tin (Sn).

[0118] Element M1 can exist in the catalyst as an oxide, sulfide, halide, oxyhalide, in chemical combination with one or more of the other components of the catalyst, or as an elemental metal.

[0119] The catalyst may further comprise a halogen element (hereinafter referred to as X). The halogen content is advantageously between 0.75 and 5.5% by weight relative to the total weight of the catalyst, more preferably between 0.85 and 4% by weight, most preferably between 0.9 and 3.5% by weight, and even more preferably between 1 and 1.5% by weight. Preferably, the halogen element is selected from the group consisting of fluorine, chlorine, bromine, and iodine. Preferably, the halogen element is chlorine (Cl).

[0120] In one embodiment according to the invention, the catalyst may further comprise phosphorus. The phosphorus may be present in the final catalyst as an oxide or mixed oxide compound, phosphate, polyphosphate, sulfide, halide, oxyhalide, hydride, in chemical combination with one or more of the other components of the catalyst.

[0121] In this embodiment, the phosphorus content, expressed as an element, in the catalyst is between 0.1 and 1% by weight relative to the total weight of the catalyst, preferably between 0.2 and 0.8% by weight.

[0122] The specific surface area of ​​the catalyst is advantageously between 170 m 2 / g and 250 m 2 / g, preferably between 180 m 2 / g and 220 m 2 / g, preferably between 185 m 2 / g and 220 m 2 / g, more preferentially between 185 m 2 / g and 210 m 2 / g, and even more preferentially between 185 m 2 / g and 195 m 2 / g.

[0123] The catalyst advantageously has a total pore volume measured by mercury porosimetry of between 0.1 cm 3 / g and 1.5 cm 3 / g, preferably between 0.4 cm 3 / g and 0.8 cm 3 / g, and preferably between 0.4 cm 3 / g and 0.7 cm 3 / g.

[0124] Advantageously, the catalyst has a packed filling density (DRT) value between 0.40 and 0.90 g / mL, preferably between 0.50 and 0.85 g / mL.

[0125] 5. Support

[0126] The catalyst support comprises alumina, preferably is made of alumina. The alumina(s) may be of type r|, y or 5. Preferably, they are of type y or 5. Even more preferably, they are of type y.

[0127] The specific surface area of ​​the support is advantageously between 170 m 2 / g and 250 m 2 / g, preferably between 180 m 2 / g and 220 m 2 / g, preferably between 185 m 2 / g and 220 m 2 / g, more preferentially between 185 m 2 / g and 210 m 2 / g, and even more preferentially between 185 m 2 / g and 195 m 2 / g.

[0128] The substrate advantageously exhibits a total pore volume measured by mercury porosimetry of between 0.1 cm 3 / g and 1.5 cm 3 / g, preferably between 0.4 cm 3 / g and 0.8 cm 3 / g, and preferably between 0.4 cm 3 / g and 0.7 cm 3 / g. Advantageously, the support has a packed filling density (DRT) value between 0.50 and 0.70 g / mL, preferably between 0.60 and 0.65 g / mL.

[0129] Preferably, the support is in the form of particles, preferably spheres, with an equivalent average diameter between 0.5 and 10.0 mm, preferably between 1.0 and 5.0 mm. According to the invention, the equivalent average diameter defines an equivalent average surface diameter, advantageously determined by the laser diffraction method, and the equivalent average diameter of the particles advantageously corresponds to the average diameter of spheres having the same specific surface area as said particles.

[0130] The substrate can be obtained by any technique known to those skilled in the art. The shaping can be carried out, for example, by extrusion, by pelletizing, by the oil-drop coagulation method, by rotary plate granulation, or by any other method well known to those skilled in the art.

[0131] 6. Process for preparing the catalyst

[0132] The catalyst used in the process according to the invention can be prepared using any technique known to a person skilled in the art.

[0133] Platinum

[0134] Advantageously, platinum is supplied to the support in any suitable manner, such as coprecipitation, ion exchange, or impregnation. Preferably, it is introduced by impregnation of the support, for example, by excess or dry impregnation (the volume of solution containing the element to be introduced corresponding to the porosity of the support), and preferably by excess impregnation. For this purpose, the support is impregnated with an impregnation solution, aqueous or organic, or alternatively consisting of a mixture of water and at least one organic solvent, comprising at least one platinum precursor.

[0135] In general, hydrogen chloride or another similar acid may also be added to the impregnation solution to further facilitate the incorporation or fixation of platinum to the substrate surface and promote a uniform distribution of platinum in the substrate.

[0136] Preferably, platinum precursors belong to the following group, but this list is not exhaustive: hexachloroplatinic acid, bromoplatinic acid, ammonium chloroplatinate, platinum chlorides such as PtCh or PtCl, platinum dichlorocarbonyl dichloride, platinum tetraamine chloride, or dihydroxyplatinemia. Organic platinum complexes, such as platinum(II) diacetylacetonate, may also be used. These precursors can be used alone or in mixtures. Preferably, the precursor used is hexachloroplatinic acid.

[0137] Element M1

[0138] The element M1 chosen from tin, germanium, lead, gallium, indium and thallium can be supplied in any suitable way, such as coprecipitation, ion exchange or impregnation, and this at any stage of the catalyst preparation process.

[0139] According to one variant, element M1 can be introduced into the support, for example, during support synthesis or during support shaping. While not exhaustive, techniques involving addition before or during the dissolution of the support's oxide precursors during support synthesis, with or without curing, are suitable. The introduction can therefore be simultaneous with or subsequent to the mixing of the support precursors. Element M1 can be introduced during support synthesis using a sol-gel technique or added to an alumina sol. Element M1 can also be introduced during support processing using prior art shaping techniques such as extrusion or oil-drop forming.

[0140] According to a second variant, element M1 can be introduced onto the substrate, for example, by impregnating the previously prepared substrate. Impregnation of the substrate with a solution, aqueous or organic, or consisting of a mixture of water and at least one organic solvent, comprising one or more precursors of elements M1, can be carried out by excess solution or dry impregnation. Impregnation of the substrate with a solution containing one or more precursors of element M1 can be carried out before, after, or simultaneously with the impregnation of the Group VIII metal. Impregnation can be carried out in the presence of species that influence the interaction between the precursor of element M1 and the substrate.These species can be, for example, and without limitation, mineral acids (HCl, HNO3) or organic acids (carboxylic or polycarboxylic acid types), and organic compounds of the complexing type, as described for example in US patents 6,872,300 and 6,291,394. Preferably, the impregnation is carried out according to any technique known to those skilled in the art allowing a homogeneous distribution of the element M1 within the support.

[0141] According to a third variant, element M1 can also be introduced partly during the synthesis or shaping of the support and partly by deposition onto the shaped support. Preferably, element M1 is introduced into the support, i.e., during the synthesis or shaping of the support. In the case of an alumina-based support in the form of beads prepared by the draining technique, the precursor of element M1 is introduced into the suspension to be drained.

[0142] The precursors of element M1 can be mineral or organometallic, possibly water-soluble organometallic. The precursor of element M1 can be selected from the group consisting of halogenated compounds, hydroxides, carbonates, carboxylates, sulfates, tartrates, and nitrates. These forms of element M1 can be introduced into the catalyst preparation medium as is or generated in situ (for example, by the introduction of tin and carboxylic acid). When the element M1 is tin, tin-based organometallic precursors can be chosen, for example, from the following list: SnF t, where R represents an alkyl group, for example the butyl group, MeaSnCl, MeaSnCh, EtaSnCl, EtaSnCh, EtSnCh, iPrSnCh and the hydroxides MeaSnOH, Me2Sn(OH)2, EtaSnOH, Et2Sn(OH)2, the oxides (BuaSnO, acetate BuaSnOC(O)Me. Preferably, tin halogenated species, particularly chlorinated ones, are used.Even more preferably, the precursor of element M1 is SnCh or SnCk.

[0143] Potassium

[0144] Potassium can be supplied in any suitable manner, such as coprecipitation, ion exchange, or impregnation, at any stage of the catalyst preparation process. In particular, it can be introduced according to the three variants described for element M1. Preferably, it is introduced by impregnation, either dry or in excess; most preferably, it is introduced in excess; and most preferably, it is introduced after the platinum has been introduced as described above.

[0145] Preferably, the precursors can be chosen from KOH, KCl, KNO3, K2CO3, or K2PtCl6, taken alone or in mixture.

[0146] Element X

[0147] Element X, chosen from among the halogens, can be supplied by any suitable method, such as coprecipitation, ion exchange, or impregnation, at any stage of the catalyst preparation process. In particular, it can be introduced according to the three variants described for element M1.

[0148] In a preferred embodiment, element X is chlorine. Chlorine is introduced by impregnation, and particularly preferably, it is introduced by impregnation simultaneously with platinum or prior to platinum impregnation. As a possible counter-anion of compounds containing element M1, platinum, and potassium, chlorine can also be introduced simultaneously with these elements, for example, by using precursors such as H2PtCl, SnCh, or KCl, or by treating the support with hydrochloric acid in one of the preparation steps. Chlorine can also be supplied in the preparation process by an additional oxychlorination step.

[0149] Phosphorus can be supplied in any suitable manner, such as coprecipitation, ion exchange, or impregnation, at any stage of the catalyst preparation process. In particular, it can be introduced according to the three variants described for element M1.

[0150] According to one variant, phosphorus is introduced into the support, i.e. during its shaping, for example simultaneously with element M1.

[0151] According to another variant, phosphorus is introduced by impregnation, and particularly preferably it is introduced by impregnation at the same time as the group VIII metal. In this case, the impregnation solution contains the precursor of the group VIII metal and the precursor of phosphorus.

[0152] Phosphorus precursors can be acids or salts, and can be chosen from the following compounds: H3PO4, H3PO3, H3PO2, NH4H2PO4, or (NhL^HPCU.

[0153] Order of introduction of precursors

[0154] As described above, the catalyst preparation process involves several implementation methods, distinguished primarily by the order in which platinum, element M1, potassium, element X, and optionally phosphorus are introduced onto and / or into the prepared support—that is, during support synthesis or during support preparation. The catalyst preparation process includes the simultaneous or successive introduction, in any order, of platinum, element M1, potassium, and optionally phosphorus and / or element X onto and / or into the prepared support.

[0155] When element M1 and / or potassium and / or optionally phosphorus and / or element X are introduced into the support, i.e., during support synthesis or support shaping, the preparation process generally includes a drying step and a calcination step before platinum deposition. Drying is generally carried out at a temperature of 250°C or lower, preferably between 50°C and 250°C, and more preferably between 70°C and 200°C, under air or an inert atmosphere. Calcination is preferably carried out at a temperature between 350°C and 750°C, and more preferably between 400°C and 600°C, and even more preferably between 400°C and 550°C. The temperature ramp-up may be continuous or include intermediate temperature plateaus, these plateaus being reached at fixed or variable rates of temperature ramp-up.These temperature increases can therefore be identical or differ in their rate (in degrees per minute or per hour). The gas atmosphere used during calcination contains oxygen. Air can therefore also be used during this calcination stage. The calcination gas may also contain water.

[0156] When one or more elements among platinum, element M1, potassium, element X and / or optionally phosphorus, are introduced onto the shaped support, preferably by dry or excess impregnation, the introduction of said elements may be simultaneous by a single impregnation solution or take place separately by several impregnation solutions containing one or more of the components and this in any order.

[0157] Any impregnation solution described in the present invention may comprise any polar solvent known to those skilled in the art. The polar solvent used is advantageously chosen from the group consisting of methanol, ethanol, water, phenol, and cyclohexanol, alone or in mixtures. The polar solvent may also advantageously be chosen from the group consisting of propylene carbonate, DMSO (dimethyl sulfoxide), N-methylpyrrolidone (NMP), or sulfolane, alone or in mixtures. Preferably, a polar protic solvent is used. A list of common polar solvents and their dielectric constants can be found in the book "Solvents and Solvent Effects in Organic Chemistry," C. Reichardt, Wiley-VCH, 3rd edition, 2003, pages 472-474. Most preferably, the solvent used is water or ethanol, and particularly preferably, the solvent is water.

[0158] After each impregnation step, the resulting catalyst precursor is preferably dried to remove all or part of the solvent introduced during impregnation, preferably at a temperature below 250°C, more preferably between 50°C and 250°C, and most preferably between 70°C and 200°C. Drying is advantageously carried out for a period of between 1 and 24 hours, preferably between 1 and 20 hours. Drying is performed under air or under an inert atmosphere (e.g., nitrogen). After the drying step, the catalyst is preferably calcined, generally under air. The calcination temperature is generally between 350°C and 650°C, and preferably between 400°C and 650°C, and even more preferably between 450°C and 550°C. The temperature ramp may optionally include temperature plateaus. The calcination time is generally between 0.5 hours and 16 hours, preferably between 1 hour and 5 hours.The gas atmosphere used during calcination contains oxygen. Therefore, air can also be used in this calcination stage. The calcination gas may also contain water.

[0159] In a particular embodiment, the catalyst according to the invention is prepared by a preparation process comprising the following successive steps: a) a support comprising the element M1 selected from tin, germanium, lead, gallium, indium and thallium is prepared to obtain a first catalyst precursor; b) the first catalyst precursor obtained in step a) is dried under a flow of a neutral gas or under a flow of a gas containing oxygen at a temperature less than or equal to 250°C, then calcined at a temperature between 350°C and 750°C to obtain a first dried and calcined catalyst precursor; c) the first dried and calcined catalyst precursor obtained in step b) is impregnated with an impregnation solution comprising at least one platinum precursor, and optionally at least one precursor of an element X, and optionally phosphorus to obtain a second catalyst precursor;d) the second catalyst precursor obtained in step c) is dried under a flow of neutral gas or under a flow of oxygen-containing gas at a temperature of 250°C or less, and then calcined at a temperature between 350°C and 650°C to obtain a second dried and calcined catalyst precursor; e) the second dried and calcined catalyst precursor obtained in step d) is impregnated with an impregnation solution comprising a potassium precursor, and optionally at least one precursor of element X, to obtain a third catalyst precursor; f) the third catalyst precursor obtained in step e) is dried under a flow of neutral gas or under a flow of oxygen-containing gas at a temperature of 250°C or less, and then calcined at a temperature between 350°C and 650°C;(g) Optionally, an oxychlorination step is carried out at a temperature between 350°C and 550°C and under a pressure between 0.1 MPa and 1.5 MPa, followed by calcination at a temperature between 350°C and 650°C; it being understood that element X is supplied at least once in step (c) and / or step (e), or possibly in step (g). In an embodiment according to the invention, when element X is chlorine, said chlorine element is supplied at least once in either step (c) and / or step (e), and / or step (g).

[0160] Preferably, potassium is introduced on a catalyst precursor containing platinum, which has been previously dried and calcined. Introducing potassium after platinum prevents platinum from being leached out during platinum impregnation.

[0161] In step a), a substrate comprising element M1, preferably tin, is prepared. Element M1, preferably tin, can be introduced at any time during substrate preparation, and preferably during shaping, or by impregnation onto an already formed substrate. Preferably, element M1 is introduced during substrate shaping.

[0162] Similarly, phosphorus can be introduced at any stage of substrate preparation, and preferably during shaping, or by impregnation onto an already formed substrate. According to one variant, phosphorus is introduced into the substrate, i.e., during substrate shaping, preferably with element M1, preferably tin. According to another variant, phosphorus is introduced by impregnation, and particularly preferably it is introduced by impregnation simultaneously with platinum.

[0163] The introduction of platinum can advantageously be carried out by one or more excess impregnations of solution on the support, or by one or more dry impregnations, and, preferably, by a single excess impregnation of said support (preferably containing element M1, preferably tin, and possibly phosphorus), using solution(s), preferably aqueous, containing the platinum precursor and preferably the phosphorus precursor (when the support does not contain or only partially contains phosphorus).

[0164] In step e), the second dried and calcined catalyst precursor obtained in step d) is impregnated with an impregnation solution comprising at least one potassium precursor. The introduction of potassium can advantageously be carried out by one or more excess impregnations of the solution onto the support, or by one or more dry impregnations, and preferably by a single dry or excess impregnation of said precursor, using solution(s), preferably aqueous, containing at least one potassium precursor.

[0165] In the optional step g), chlorine is supplied by means of an oxychlorination treatment. Such a treatment can, for example, be carried out at a temperature between 350°C and 550°C and under a pressure between 0.1 MPa and 1.5 MPa, for a duration preferably between 30 minutes and 10 hours, and under an air flow containing the desired quantity of chlorine and possibly containing water.

[0166] According to another embodiment, the catalyst according to the invention can be prepared by preparing a support comprising tin by introducing the tin precursor during the shaping of the support, followed by one or more excess impregnations of solution on the support, or by one or more dry impregnations, using solution(s), preferably aqueous, containing a platinum precursor, a phosphorus precursor and a potassium precursor, alone or in mixture, then drying and calcining under the conditions described above.

[0167] When the various precursors used in the preparation of the catalyst according to the invention do not contain halogens or contain halogens in insufficient quantities, it may be necessary to add a halogenated compound during preparation. Any compound known to those skilled in the art can be used and incorporated into any of the steps in the preparation of the catalyst according to the invention. In particular, it is possible to use organic compounds such as methyl or ethyl halides, for example dichloromethane, dichloroethane, dichloropropane, chloroform, methylchloroform, or carbon tetrachloride.

[0168] The halogen can also be added by impregnation with an aqueous solution of the corresponding acid, for example, hydrochloric acid, at any point during the preparation. A typical protocol involves impregnating the solid to introduce the desired amount of halogen. The catalyst is then kept in contact with the aqueous solution for a sufficient time to deposit this amount of halogen.

[0169] In one embodiment of the invention, prior to using the catalyst in the catalytic reactor and implementing a dehydrogenation process, a reduction treatment step is carried out in the presence of a reducing gas to obtain a catalyst comprising platinum and said element M1 at least partially in metallic form. This step is advantageously carried out in-situ, that is, after loading the catalyst into a dehydrogenation reactor. Performing the reduction treatment of the catalyst in-situ eliminates the need for an additional step of passivating the catalyst with an oxygenated compound or with CO2, which is necessarily required when the catalyst is prepared by performing an ex-situ reduction treatment, that is, outside the reactor used for dehydrogenation.Indeed, when the reduction treatment is carried out ex-situ, it is necessary to carry out a passivation step in order to preserve the metallic phase of the catalyst in the presence of air (during the transport and loading operations of the catalyst in the hydrogenation reactor), and then to carry out a new step of reducing the catalyst.

[0170] The reducing gas is preferably hydrogen. Hydrogen can be used pure or in a mixture (for example, a hydrogen / nitrogen, hydrogen / argon, or hydrogen / methane mixture). When hydrogen is used in a mixture, any proportion is possible.

[0171] Preferably, the reduction treatment is carried out at a temperature between 100°C and 600°C, and more preferably between 200°C and 580°C, under a stream of pure or diluted hydrogen, up to the maximum reduction temperature, followed by holding at that temperature for, for example, 30 minutes to 6 hours. The temperature rise to the desired reduction temperature is generally slow, for example, set between 0.1 and 10°C / min, preferably between 0.3 and 7°C / min.

[0172] The hydrogen flow rate, expressed in L / hour / gram of catalyst precursor, is between 0.01 and 100 L / hour / gram of catalyst, preferably between 0.05 and 10 L / hour / gram of catalyst precursor, even more preferably between 0.1 and 5 L / hour / gram of catalyst precursor.

[0173] Passivation step (optional)

[0174] The preparation process may advantageously include a passivation step using a sulfur compound, which improves the selectivity of the catalysts and prevents thermal runaway during the start-up of new catalysts (known as "run-away" in English). The passivation step is carried out using methods known to those skilled in the art.

[0175] The passivation step with a sulfur compound is generally carried out at a temperature between 20 and 350°C, preferably between 40 and 200°C, typically for 10 to 240 minutes. The sulfur compound is chosen, for example, from the following: thiophene, thiophane, alkyl monosulfides such as dimethyl sulfide, diethyl sulfide, dipropyl sulfide, and propylmethyl sulfide, or an organic disulfide with the formula HO-R1-S-S-R2-OH, such as dithiodiethanol with the formula HO-C2H4-SS-C2H4-OH (often called DEODS). The sulfur content is generally between 0.1 and 2% by weight of the compound relative to the total weight of the catalyst. Examples

[0176] The following examples illustrate the invention without limiting its scope.

[0177] Example 1 illustrates the dehydrogenation process of methylcyclohexane according to an implementation with adiabatic reactor technology.

[0178] The feedstock to be treated comprises at least 99.9% by weight of methylcyclohexane.

[0179] The dehydrogenation reaction is implemented with three adiabatic reactors R1, R2, R3 in series.

[0180] 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, referred to as catalyst A hereafter.

[0181] Catalyst A is prepared on a gamma alumina support, aiming for a deposition of 0.25 wt% platinum, 0.4 wt% phosphorus, and 1.1 wt% chlorine on the final catalyst. The gamma alumina support is in the form of beads with a diameter of 1.7 mm. The properties of this support are as follows: Specific surface area (SBET) = 200 m² 2 .g' 1 Total pore volume (TPV) = 0.65 cm 3 . g _1, packed filling density (DRT) = 0.62 g. cm' 1 tin content 0.28% by weight.

[0182] To 100 grams of alumina support containing tin, add 400 cm 3 of an aqueous solution of hexachloroplatinic acid, phosphoric acid and hydrochloric acid. It is left in contact for 4 hours and then wrung out. It is dried at 120°C for 15 hours and then calcined at 500°C under an air flow of 100 liters per hour for 3 hours, with a temperature rise rate of 7°C per minute.

[0183] Chlorine content exceeding 1.1% by weight after calcination is adjusted to 1.1% by weight by a partial dechlorination heat treatment at 520°C under dry air and with an additive of 8000 ppm vol of water for 2.5 hours.

[0184] The catalyst A obtained after dechlorination contains 0.23% by weight of platinum, 0.28% by weight of tin, 0.39% by weight of phosphorus and 1.13% by weight of chlorine.

[0185] Table 1 below summarizes all the parameters of the reaction system and the operating conditions that remain constant during operation. Table 1

[0186] Table 2 below summarizes the performance obtained with catalyst A. The conversion of methyl cyclohexane is 99.43% and the selectivity for toluene is 98.48%.

[0187] Table 2

[0188] Example 2 (according to the invention)

[0189] Example 2 illustrates a dehydrogenation process according to the invention. The feed to be treated comprises at least 99.9% by weight of methylcyclohexane.

[0190] The dehydrogenation reaction is carried out in a multi-tube alloy steel reactor whose tubes include a fixed bed with a catalyst. The heat transfer fluid used is oil marketed by Dow under the trade name Dowtherm™ A. It is introduced into the shell in gaseous form, specifically as saturated vapor. Table 3 below summarizes all the reactor parameters and operating conditions, which remain constant during operation.

[0191] Table 3

[0192] In this example, catalysts B, C and D were used.

[0193] Catalyst B: Pt-PK / Al2O3-Sn-Cl (according to the invention)

[0194] Catalyst B is prepared from catalyst A aiming for a deposition of 0.06 wt% potassium.

[0195] To 100 g of this calcined solid, add 440 cm 3of a potassium carbonate solution by excess impregnation. It is dried at 120°C for 15 h then calcined at 500°C under an air flow of 100 litres per hour for 3 hours, with a temperature rise rate of 7°C per minute.

[0196] The chlorine content is adjusted to 1.1% by weight by a partial chlorination heat treatment at 520°C under dry air and with the addition of 4000 ppm vol of hydrochloric acid for 2 hours.

[0197] The catalyst B obtained after calcination contains 0.23% by weight of element platinum, 0.28% by weight of element tin, 0.39% by weight of phosphorus, 1.13% by weight of chlorine and 0.061% by weight of potassium.

[0198] Catalyst C: Pt-PK / AhOs-Sn-Cl (according to the invention)

[0199] Catalyst C is prepared from catalyst A by aiming for a deposition of 0.25% weight of potassium.

[0200] To 100 g of this calcined solid, add 440 cm 3of a potassium carbonate solution by excess impregnation. It is dried at 120°C for 15 h then calcined at 500°C under an air flow of 100 litres per hour for 3 hours, with a temperature rise rate of 7°C per minute.

[0201] The chlorine content is adjusted to 1.1% by weight by a partial chlorination heat treatment at 520°C under dry air and with the addition of 4000 ppm vol of hydrochloric acid for 2 hours.

[0202] The catalyst B obtained after calcination contains 0.23% by weight of element platinum, 0.28% by weight of element tin, 0.39% by weight of phosphorus, 1.13% by weight of chlorine and 0.247% by weight of potassium.

[0203] Catalyst D: Pt-PK / AhCh-Sn-Cl (according to the invention)

[0204] A catalyst D is prepared on the same support as that of catalyst A, aiming for a deposition of 0.25% by weight of platinum, 0.4% by weight of phosphorus, 1.5% by weight of chlorine and 1.5% by weight of potassium.

[0205] To 100g of alumina support containing tin, add 400 cm 3 of an aqueous solution of hexachloroplatinic acid, phosphoric acid and hydrochloric acid. It is left in contact for 4 hours and then wrung out. It is dried at 120°C for 15 hours and then calcined at 500°C under an air flow of 100 liters per hour for 3 hours, with a temperature rise rate of 7°C per minute.

[0206] To 100 g of this calcined solid, add 75 cm 3 of a potassium carbonate solution at a concentration of 22.3 g / L by dry impregnation. It is allowed to mature, dried at 120°C for 15 h then calcined at 500°C under an air flow of 100 litres per hour for 3 hours, with a temperature rise rate of 7°C per minute.

[0207] The catalyst D obtained after calcination contains 0.27% by weight of the element platinum, 0.28% by weight of the element tin, 0.38% by weight of the element phosphorus, 1.48% by weight of the element chlorine and 1.53% by weight of the element potassium.

[0208] Table 4 below summarizes all the reactor parameters and operating conditions that remain constant during operation. The same operating conditions are chosen for all catalysts.

[0209] For catalyst A used in adiabatic mode, the inlet temperature of each reactor was chosen to achieve the same level of conversion as in the other cases. Table 4

[0210] Table 5 below summarizes the performance obtained with the different catalysts. The presence of potassium increases toluene selectivity. Catalyst A exhibits the lowest selectivity, while catalyst D has the highest. The conversion remains stable because, at thermodynamic equilibrium, increasing the conversion requires raising the reactor temperature. Hydrogen production increases as the potassium content increases until it reaches a plateau.

[0211] Table 5

Claims

DEMANDS 1. A process for dehydrogenating a feed comprising a cyclic alkane with a 6-carbon ring in order to produce a dehydrogenation effluent comprising hydrogen and an aromatic compound, said process employing a reaction section comprising at least one multitubular reactor comprising at least two tubes and a shell, said tubes each comprising at least one fixed bed of at least one dehydrogenation catalyst, characterized in that: - said charge supplies said tubes in gaseous form, at an inlet temperature of the charge greater than or equal to 300°C, at an inlet pressure of said charge between 0.1 and 1 MPa, and at an hourly spatial velocity in weight of the inlet charge between 0.5 and 15 h' 1 preferably between 2 and 10 am -1 ; - a heat transfer fluid circulates in said calender, so that said heat transfer fluid is introduced into 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 shell at a flow rate such that the ratio of the weight flow rate of said heat transfer fluid at the inlet of the shell to the weight flow rate of said charge at the inlet of the tubes is greater than or equal to 1.0; - said heat transfer fluid is introduced into the shell at an inlet temperature of the heat transfer fluid between 300°C and 400°C, and at an inlet pressure of the heat transfer fluid between 0.10 MPa and 1.10 MPa; - said dehydrogenation catalyst comprises an active phase comprising at least platinum, potassium at a content of between 0.03% and 3% by weight of element potassium relative to the total weight of the catalyst, and a support comprising alumina.

2. A method according to claim 1, wherein the charge comprises at least 80% by weight of a cyclic alkane with a 6-carbon ring.

3. A method according to any one of claims 1 or 2, wherein said charge comprises at least 80% by weight of methylcyclohexane.

4. A process according to any one of the preceding claims, wherein the feed also comprises 7-carbon hydrocarbon compounds without a 6-carbon ring, in particular iso-alkanes and / or alkyl-cyclopentanes, 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 feed.

5. A process according to any one of the preceding claims, wherein the feed also comprises 8-carbon hydrocarbon compounds without a 6-carbon ring, 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 feed.

6. A method according to any 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 at the calender inlet to the weight flow rate of said charge at the tube inlets is greater than or equal to 5, and preferably less than or equal to 20.

0.

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

8. A method according to any one of the preceding claims, characterized in that said multitubular reactor comprises at least 100 tubes, preferably at least 1000 tubes, preferably at least 2000 tubes.

9. A process according to any one of the preceding claims, wherein the platinum content is between 0.02 and 2% by weight of element platinum relative to the total weight of the catalyst.

10. A method according to any one of the preceding claims, wherein said catalyst comprises an element M1 selected from tin, germanium, lead, gallium, indium and thallium.

11. A method according to the preceding claim, wherein the content of element M1 is between 0.01 and 10% weight by weight of element M1 relative to the total weight of the catalyst.

12. A method according to any one of the preceding claims, wherein said catalyst comprises a halogen element selected from fluorine, chlorine, bromine and iodine.

13. A method according to the preceding claim, wherein the halogen content is between 0.75 and 5.5% by weight of halogen element relative to the total weight of the catalyst.

14. A process according to any one of the preceding claims, wherein said catalyst comprises phosphorus, at a content of between 0.1 and 1% weight of phosphorus element relative to the total weight of the catalyst.

15. A method according to any one of the preceding claims, wherein said alumina support comprises a specific surface area of ​​between 170 m² 2 / g and 250 m 2 / g.

16. A process according to any one of the preceding claims, characterized in that all or part of the feedstock is obtained by a process of hydrogenating aromatic compounds, in particular including toluene.

17. A process according to any one of the preceding claims, characterized in that hydrogen is separated from the rest of the dehydrogenation effluent.

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