Process and device for selective hydrogenation by catalytic distillation including a reaction zone with an ascending cocurrent flow of liquid and gas

21 ~~~ '~ 6 1 YROCEDE ET DISPO ~ TTTF D ~ HYDRO ~ NATTON S ~ TF (' TTVF BY DISTILLATTON CATALYTTOUE COh: CARRYING A REACTION ZON Ä CO-CURRENT UPWARD LIQUID GAS The invention relates to a reactive distillation device comprising a distillation zone, associated with a reaction zone at least partly internal to said distillation zone and comprising at least one catalytic bed, in which the transformation of the feed is carried out, in the presence of a catalyst and at least one gas stream comprising hydrogen, characterized in that any catalytic bed of the internal part of said reaction zone is passed through in ascending co-current by said gas stream and the liquid. The invention also relates to processes for the selective hydrogenation of light unsaturated hydrocarbons, mainly optional olefins and benzene, included in a mixture consisting mainly of hydrocarbons comprising at least five carbon atoms per molecule, as well as hydroisomerization. at least part of the butene-1 included in a feed mainly comprising olefinic hydrocarbons including isobutene, as well as butene-1 and butenes-2 in a ratio corresponding substantially to thermodynamic equilibrium.

In addition to the selective hydrogenation of the light unsaturated compounds of a reformate, including olefins and benzene, without significant hydrogenation of heavier unsaturated compounds such as toluene and, a fortiori, xylenes, the device according to the invention can be used. apply to various catalytic reactions, balanced or complete, for which it is possible to separate, by distillation, at least one of the products of the reaction, in the pure or diluted state, under temperature and pressure conditions close to those of the reaction, and more particularly to reactions of isomerization of paraffins by reorganization of the backbone, of isomerization of olefins by displacement of the double bond (hydroisomerization) or by reorganization of the backbone , hydrogenation of unsaturated compounds to saturated compounds, dehydrogenation of compounds saturated with unsaturated compounds, all reactions which require the presence of hydrogen.

The hydrogenation catalyst can be placed in the reaction zone according to the various technologies proposed for carrying out catalytic distillations.

These technologies have been developed primarily for etherification reactions, which involve contact between reactants in a homogeneous liquid phase and the solid catalyst.

They are basically of two types.

According to the first type of technology, the reaction and the distillation proceed simultaneously in the same physical space, as taught for example by patent application WO-A-90 / 02,603, patents US-A-4,471,154, US- A-4,475,005,.

2 US-A-4.215.011, US-A-4.307.254, US-A-4.336.407, US-A-4.439.350, US-A-5.189.001, US-A-5.266.546, US- A-5,073.236, US-A-5.215.011, US-A-5.275.790, US-A-5.338.517, US-A-5.308.592, US-A-5.236.663, US-A- 5,338,518, as well as patents EP-B1-0,008,860, EP-B1-0,448,884, EP-B1-0,396,650 and EP-B1-0,494,550 patent application EP-A1-0,559,511.

and the catalyst is then generally in contact with a descending liquid phase, generated by the reflux introduced at the top of the distillation zone, and with an ascending vapor phase, generated by the reboiling vapor introduced at the bottom of the zone.

According to the second type of technology, the catalyst is arranged lo such that the reaction and the distillation generally proceed independently and consecutively, as taught for example by patents US-A-4,847,430, US-A-5,130 .102 and US-A-5,368,691, the vapor from the distillation hardly passing through any catalyst bed in the reaction zone.

For any chemical reaction requiring the addition of a foreign gaseous reagent to the distillation feed, this reagent must be introduced into the reaction zone in a different manner depending on the type of technology chosen to carry out the catalytic distillation.

According to the first type of technology, the gaseous reactant can simply be added to the steam from the distillation at any level, but in any event, before its penetration into the reaction zone, generally substantially at the entrance to minus one catalytic bed of the reaction zone.

According to the second type of technology, the gaseous reactant must be introduced in a manner appropriate to the options chosen to impose the direction of circulation of the liquid and of the gas in the catalyst bed.

It therefore appeared that, for a reaction taking place in the presence of a solid catalyst, between a reagent in liquid phase and a gaseous reagent sparingly soluble in the liquid, such as the hydrogenation of unsaturated hydrocarbons mixed with other hydrocarbons , the option which consists in carrying out the catalytic distillation while avoiding the passage of the distillation vapor over the catalyst and by circulating the liquid and the gaseous reagent in ascending co-current in the catalytic bed is the most effective. It appeared that the pressure drop across the catalytic bed (s) according to the first type of technology does not allow an intimate mixture to be obtained between the liquid phase and the gas stream containing hydrogen.

In fact, according to this type of technology where the reaction and the distillation proceed simultaneously in the same physical space, the liquid phase descends through the catalytic bed in a trickling flow, therefore in streams of liquid.

The gas fraction containing the vaporized feed fraction and the gas stream containing hydrogen rise through the catalytic bed in gas columns.

By this arrangement, the entropy of the system is the highest and the pressure drop across the catalytic bed (s) is the lowest.

Thus, the way of operating according to the first type of technology does not easily make it possible to promote the dissolution of hydrogen in the liquid phase comprising the unsaturated compound (or compounds).

The second type of technology, comprising a specific device for the distribution of the liquid fraction to be hydrogenated and the gas stream containing hydrogen, where said liquid fraction and said gas stream pass through the catalytic bed in ascending co-current, makes it possible to effecting the desired hydrogenation reaction substantially in the absence of the gaseous fraction of the feed and under conditions where the pressure drop across the catalytic bed (s) is greatest.

The increase of ~ 5 the pressure drop, thanks to the specific device according to the invention, makes it possible to increase the solubility of hydrogen in the liquid phase and therefore to promote hydrogenation in the liquid fraction.

On the other hand, in view of the recognized harmfulness of benzene and olefins, unsaturated compounds, the general tendency is to reduce the content of these constituents in gasolines.

Indeed, benzene has carcinogenic properties and it is consequently required to limit as much as possible any possibility of polluting the ambient air, in particular by practically excluding it from automobile fuels.

Reformulated fuels in the United States must not contain more than 1% benzene;

25 in Europe, even if the specifications are not yet so strict, it is recommended to gradually move towards this value.

In addition, olefins have been recognized as being among the most reactive hydrocarbons in the cycle of photochemical reactions with nitrogen oxides, which occurs in the atmosphere and leads to the formation of ozone.

An increase in the concentration of ozone in the air can cause respiratory problems.

Reducing the olefin content of gasolines, and more particularly of the lighter olefins which have the most tendency to volatilize during handling of the fuel, is therefore desirable.

35 The benzene content of a gasoline is very largely dependent on that of the reformate component of this gasoline.

The reformate results from a catalytic 21940'6 4 treatment of naphtha intended to produce aromatic hydrocarbons, comprising mainly 6 to 9 carbon atoms in their molecule and whose very high octane number gives gasoline its properties. anti-knock drugs.

For the reasons of harmfulness described above, it is therefore necessary to reduce the benzene content of the reformate as much as possible.

Several routes are possible.

A first approach consists in limiting, in the naphtha constituting the feed of a catalytic reforming unit, the content of benzene precursors, such as cyclohexane and methylcyclopentane.

This solution effectively makes it possible to significantly reduce the benzene content of the effluent from the reforming unit, but is not sufficient on its own when it comes to reducing contents as low as 1%.

A second route consists in removing, by distillation, a light fraction of the reformate containing the benzene.

This solution leads to a loss of the order of 15 to 20% of hydrocarbons which would be recoverable in gasolines.

A third way consists in extracting the benzene present in the effluent of the reforming unit.

Several known techniques are applicable in principle: extraction by solvent, extractive distillation, adsorption.

None of these techniques is applied industrially, because none allows to selectively extract benzene in an economical manner.

A fourth way is to chemically transform the 2o benzene to convert it into a constituent not covered by the legal limitations.

Alkylation with ethylene, for example, converts benzene mainly into ethylbenzene.

This operation is however expensive due to the intervention of secondary reactions which require costly separations in energy.

The benzene of a reformate can also be hydrogenated to cyclohexane.

As it is impossible to selectively hydrogenate the benzene from a mixture of hydrocarbons also containing toluene and xylenes, it is therefore necessary to fractionate this mixture beforehand so as to isolate a cut containing only benzene, which can then be hydrogen. A process has also been described in which the benzene hydrogenation catalyst is included in the rectification zone of the distillation column which separates the benzene from the other aromatics (Benzene Reduction - Kerry Rock and Gary Gildert CDTECH - 1994 Conference on Clean Air Act Implementation and Reformulated Gasoline - Oct. 94), which saves equipment costs.

. 'In addition, the isobutene intended for the polymerization must have a level of purity greater than 99% and no longer contain only traces of butene-1 and of butenes2 (a few tens of parts per million by weight, ppm).

Indeed, if the level of impurity in isobutene is too high, the polymers obtained are of poorer quality and the polymerization yield is lower. It is therefore necessary to remove from a hydrocarbon cut containing isobutene the other olefinic hydrocarbons comprising 4 carbon atoms per molecule.

As butene-1 and isobutene have very similar boiling points, it is not possible to separate them by distillation unless considerable means are employed.

The other 1o olefinic hydrocarbons having 4 carbon atoms can be separated from the isobutene by distillation.

The main problem that arises in producing high purity isobutene is therefore the separation of butene-1 from isobutene.

To effect this separation several routes are possible.

~ 5 The first way consists of an extraction with sulfuric acid:

isobutene is selectively hydrated and then regenerated by treatment of the aqueous phase.

If the temperature and the concentration are well controlled, this process makes it possible to obtain isobutene of good purity.

However, the yield usually does not exceed 90%, the extraction is not complete and dimers and oligomers are formed, leading to the formation of toxic acid slurries.

The second route consists of cracking methyl ether of tert-butyl alcohol (MTBE):

isobutene is extracted from the C4 cut by reacting it with methanol to form MTBE.

The MTBE is then cracked into methanol and isobutene over an acid catalyst.

The recovery yield may be at least 96%. The isobutene produced is of good purity but must be freed from dimethyl ether which may form during cracking.

The third possible route is the dehydration of tertiary butyl alcohol (ABT).

In the previous operation, the methanol can be replaced by water, which leads to the production of ABT. The isobutene is then recovered by dehydration of the ABT.

This route is hardly used, mainly because ABT has a strong connection to the propylene oxide market. ABT may, depending on the process, be a by-product of propylene oxide.

US-A-2,403,672 describes a process for separating isobutene from an isobutene-butene-1 mixture which comprises introducing the mixture into an isomerization and fractionation zone in which the catalyst d. The isomerization also acts as a packing body ensuring the distillation function.

This 21940 ~~ 6 solution has the major drawback of not having good distillation efficiency and therefore poor separation capacity of isobutene from butene-1.

According to this technology the reaction and the distillation proceed simultaneously in the same physical space.

The catalyst is in contact with a descending liquid phase, generated by the reflux introduced at the top of the distillation zone, and with an ascending vapor phase, generated by the reboiling vapor introduced at the bottom of the zone.

The invention relates to a reactive distillation device comprising a t0 distillation zone, comprising an exhaustion zone and a rectification zone, associated with a reaction zone, at least partly internal to said distillation zone and comprising at least one catalytic bed, in which the transformation of the feed is carried out in the presence of a catalyst and at least one gas stream comprising hydrogen, said device being characterized in that any catalytic bed of the internal part of said reaction zone is passed through in ascending co-current by said gas flow and the liquid.

The device according to the invention generally comprises o at least one means of distribution of the major part of the liquid from the bottom to the top 2o through the catalyst o at least one means of circulation of the major part of the distillation vapor from the bottom to upwards through the catalytic bed, so that said vapor is practically not in contact with the catalyst, and o at least one means of distributing the major part of the gas flow from the bottom to the top through the catalyst.

The feed which feeds the distillation zone is introduced into said zone generally at least at one level of said zone, preferably mainly at a single level of said zone.

The distillation zone generally comprises at least one column provided with at least one internal distillation unit chosen from the group formed by trays, bulk packings and structured packings, as is known to those skilled in the art, and such that the total overall efficiency is generally at least equal to five theoretical stages.

In the cases known to those skilled in the art where the implementation of a single column poses problems, it is generally preferred to split said 2194U7 ~ zone so as to finally use at least two columns which, placed end to end, produce said zone, that is to say that the rectification, possibly reaction and exhaustion zones are distributed over the columns.

In practice, when the reaction zone is at least partly internal to the distillation zone, the rectification zone or the exhaustion zone, and preferably the exhaustion zone, can generally be located in at least one column different from the column comprising the internal part of the reaction zone.

The means of circulation of the distribution vapor from the bottom to the top through the catalytic bed ~ o passes through the level of the reaction zone where the catalytic bed is located, that is to say it is generally located within of the catalytic bed, but that it may also be located at the periphery of said catalytic bed.

The reaction zone generally comprises at least one catalytic bed, preferably from 2 to 6, even more preferably from 2 to 4 catalytic beds); in the case where at least two catalytic beds are incorporated in the reaction zone, these two beds are optionally separated by at least one internal distillation unit.

2o The device according to the invention is generally such that the flow of the liquid to be transformed is co-current with the flow of the gas flow comprising hydrogen and such that the distillation vapor does not pass practically any catalytic bed of the internal part of the reaction zone (which is translated in practice by the fact that said vapor is separated from said liquid), for any catalytic bed of the internal part of the reaction zone.

In all cases of this second type of technology, any catalytic bed of the part of the reaction zone which is in the distillation zone is generally such that the gas stream comprising hydrogen and the stream of the liquid which will react circulate at co-current, generally ascending, through said bed, although overall, in the catalytic distillation zone, the gas flow comprising hydrogen and the flow of the liquid which will react circulate in countercurrent.

Such systems generally include at least one liquid distribution device which may for example be a liquid distributor, in any catalytic bed of the internal part of the reaction zone. However, insofar as the technologies used in the process according to the invention have been designed for catalytic reactions taking place between liquid reactants, they cannot be suitable without modification for a catalytic reaction, for which one of the CA 02194076 2000 The reactants, hydrogen, is in the gaseous state.

For any catalytic bed in the internal part of the reaction zone, it is therefore generally necessary to add a device for introducing the gas stream comprising hydrogen, for example according to the techniques described below.

Thus, for any catalytic bed of the internal part of the reaction zone, the internal part of the reaction zone comprises at least one liquid distribution device, generally located below said catalytic bed, and at least one device for introducing the liquid. gas stream, generally disposed under the catalytic bed or within the catalytic bed, preferably in the latter case not far from the device for introducing the liquid.

According to one technique, the device for introducing the gas stream into any catalytic bed is identical to the device for distributing liquid in the catalytic bed, that is to say there is a means for introducing the gas into the liquid. upstream of the device for introducing the gas flow is arranged substantially at the level of the liquid distribution device, the gas and the liquid being introduced separately into the catalytic bed.

According to this other technique, the device for introducing the gas stream is placed below or within the catalytic bed, preferably not far from the liquid distribution device.

On the other hand, the device according to the invention is, according to one embodiment of the invention, such that said gas flow mainly comprises hydrogen, the hydrogen coming mainly, preferably almost entirely. , from outside the distillation zone.

CA 02194076 2000-06-22 8a The device according to the invention is generally such that, for the part of the reaction zone internal to the distillation zone, the charge of the reaction zone is taken up to a level of withdrawal and represents at least a part, preferably the major part, of the liquid flowing in the distillation zone, preferably flowing in the rectification zone and even more preferably flowing at an intermediate level of the rectification zone, the effluent from the reaction zone being at least in part, preferably in major part, reintroduced into the distillation zone substantially close, that is to say generally substantially at the same height or substantially above or ~~ 9 ~ 07 ~ 9 substantially below, most often substantially at the same height or substantially above, that is to say located at a distance corresponding to a height between 0 and 4 theoretical plates of a sampling level, preferably of said sampling level, this so as to ensure the continuity of the distillation.

Then, for the part of the reaction zone internal to the distillation zone, the liquid is withdrawn naturally by flow in the part of the reaction zone internal to the distillation zone, and the reintroduction of liquid into the distillation zone is done. also naturally by flowing the liquid from the part of the reaction zone internal to the distillation zone.

Generally, the device according to the invention comprises from 1 to 4 sampling levels which feed) the external part of the reaction zone, when the reaction zone is not entirely internal to the distillation zone.

Generally, the liquid which will react, either partially or totally, first circulates in the external part of the reaction zone and then the internal part of said zone.

So, two cases can arise.

In the first case, the external part of the reaction zone is supplied by a single sampling level, and then, if said part comprises more than two reactors, these are arranged in series or in parallel.

In the second case, preferred according to the present invention, the external part 2o of the reaction zone is supplied by at least two sampling levels.

The device according to the invention is, according to one of the preferred embodiments of the invention, such that the reaction zone is entirely internal to the distillation zone.

25 One of the preferred embodiments of the device according to the invention is such that the catalyst is placed in the reaction zone according to the basic device described in US-A-5,368,691, arranged so that any catalytic bed is fed by the flow. gas comprising hydrogen, regularly distributed at its base, for example according to one of the techniques described above.

According to this 3o technology, if the distillation zone comprises a single column and if the reaction zone is entirely internal to said column, the catalyst included in any catalytic bed, internal to the distillation zone, is then in contact with a liquid phase ascending, generated by the reflux introduced at the top of the distillation zone, and with the hydrogen flowing in the same direction as the liquid; contact with the vapor phase of the distillation is avoided by passing the latter through at least one specially fitted chimney.

The invention also relates to a process for treating a feed, consisting mainly of hydrocarbons comprising at least 5, preferably between 5 and 9, carbon atoms per molecule, and comprising at least one unsaturated compound comprising at most six carbon atoms per molecule including benzene, such as the said feedstock is treated in a distillation zone, comprising an exhaustion zone and a rectification zone, associated with a hydrogenation reaction zone, at least partly internal to said distillation zone, in which the hydrogenation of at least a part, preferably the major part, of the unsaturated compounds comprising at most six carbon atoms per molecule, that is to say comprising up to six (inclusive) carbon atoms per molecule, and contained in the feed, in the presence of a hydrogenation catalyst and at least one gas stream comprising, preferably for the most part, hydrogen, so as to finally leave at the top of the distillation zone an effluent depleted in unsaturated compounds comprising at most six carbon atoms per molecule, and at the bottom of the distillation zone an effluent depleted in unsaturated compounds comprising at most six carbon atoms per molecule , characterized in that any catalytic bed of the internal part of the hydrogenation zone is passed through in ascending co-current by said gas flow and the liquid and is practically not crossed by the distillation vapor.

The hydrogenation reaction zone at least partially carries out the hydrogenation of the benzene present in the feed, generally so that the benzene content of the overhead effluent is at most equal to a certain content, and said reaction zone carries out at least less in part, preferably in major part, the hydrogenation of any unsaturated compound comprising at most six carbon atoms per molecule and different from benzene, optionally present in the feed.

Preferably, the method according to the invention comprises the use of the device according to the invention.

The distillation zone, as well as the characteristics of the gas flow, of the reaction zone, etc. have been described previously in the case of the device according to the invention.

The process according to the invention is, according to one embodiment of the invention, such that the bottom effluent from the distillation zone is mixed with the top effluent from said '~ i ~~ o7s zone.

In such a case, the mixture thus obtained is, after optional stabilization, used as fuel either directly or by incorporation into the fuel fractions.

For carrying out the hydrogenation according to the process according to the invention, the theoretical molar ratio of hydrogen necessary for the desired conversion of benzene is 3.

The quantity of hydrogen injected before or into the hydrogenation zone is optionally in excess with respect to this stoichiometry, and this all the more so as one must hydrogenate in addition to the benzene present in the feed at least partially any compound unsaturated comprising at most six carbon atoms per molecule and present in said filler.

If the conditions are such that there is an excess of hydrogen, the excess hydrogen can be advantageously recovered, for example according to one of the techniques described below.

For example, the excess hydrogen which leaves the head of the distillation zone is recovered, then injected upstream of the compression stages associated with a catalytic reforming unit, mixed with the hydrogen originating from said unit, said unit operating from preferably at low pressure (ie generally a pressure lower than 8 bar).

This excess hydrogen can also be recovered, then compressed and reused in said reaction zone.

The hydrogen used in the reaction zone according to the invention generally originates for the most part, preferably almost entirely, from outside the distillation zone. It can come from any source producing hydrogen at at least 50 volume purity, preferably at least 80% volume purity and more preferably at least 90% volume purity.

For example, we can cite hydrogen from catalytic reforming processes, from PSA

(pressure swing adsorption), electrochemical generation, steam cracking or steam reforming.

The operating conditions of the hydrogenation zone in the case of the process according to the invention are linked to the operating conditions of the distillation.

The distillation is carried out under a pressure generally between 2 and 20 bar, preferably between 4 and 15 bar, even more preferably between 4 and 10 bar (1 bar = 105 Pa), with a reflux rate between 1 and 10, and preferably between 3 and 6.

The zone head temperature is generally between 40 and 180 ° C and the zone bottom temperature is generally between 120 and 280 ° C.

The 219 ~ p76 12 hydrogenation reaction is carried out under conditions which are most generally intermediate between those established at the top and bottom of the distillation zone, at a temperature between 100 and 200 ° C, and preferably between 120 and 180 ° C, and at a pressure between 2 and 20 bar, preferably between 4 and 10 bar.

The liquid subjected to hydrogenation is supplied with hydrogen, the flow rate of which depends on the benzene concentration in said liquid and, more generally, unsaturated compounds comprising at most six carbon atoms per molecule of the charge of the zone of distillation. It is generally at least equal to the flow rate corresponding to the stoichiometry of the hydrogenation reactions 0 involved (hydrogenation of benzene and of other unsaturated compounds comprising at most six carbon atoms per molecule, included in the hydrogenation charge) and at most equal to the flow rate corresponding to 10 times the stoichiometry, preferably less than 6 times the stoichiometry, even more preferably, less than 3 times the stoichiometry.

~ 5 When the hydrogenation zone comprises a part external to the distillation zone, the catalyst disposed in said external part is according to any technology known to those skilled in the art under operating conditions (temperature, pressure, etc.) independent or not, preferably independent, of the operating conditions of the distillation zone.

In the part of the hydrogenation zone external to the distillation zone, the operating conditions are generally as follows.

The pressure required for this hydrogenation step is generally between 1 and 60 bar absolute, preferably between 2 and 50 bar and even more preferably between 5 and 35 bar.

The operating temperature of the hydrogenation zone is generally between 100 and 400 ° C, preferably between 120 and 350 ° C and more preferably between 140 and 320 ° C.

The space velocity within said hydrogenation zone, calculated with respect to the catalyst, is generally between 1 and 50 and more particularly between 1 and 30 h′1 (volume of feed per volume of catalyst and per hour).

The hydrogen flow rate corresponding to the stoichiometry of the hydrogenation reactions involved is between 0.5 and 10 times said stoichiometry, preferably between 1 and 6 times said stoichiometry and even more preferably between 1 and 3 times said said stoichiometry. stoichiometry.

However, the temperature and pressure conditions can also, in the context of the process of the present invention, be between those which are established at the top and at the bottom of the distillation zone.

219 ~ 07fi 13 More generally, regardless of the position of the hydrogenation zone relative to the distillation zone, the catalyst used in the hydrogenation zone according to the invention generally comprises at least one metal chosen from the group formed by nickel and platinum, used as such or preferably deposited on a support.

The metal should generally be in reduced form at least 50% by weight of its totality.

But any other hydrogenation catalyst known to those skilled in the art can also be chosen.

When using platinum, the catalyst may advantageously contain at least one halogen in a proportion by weight relative to the catalyst of between 0.2 and 2%.

Preferably, chlorine or fluorine or a combination of the two is used in a proportion relative to the total weight of catalyst of between 0.2 and 1.5%.

In the case of using a catalyst containing platinum, a catalyst is generally used such that the average size of the platinum crystallites is less than 60.10-10 m, preferably less than 20.10-10 m, even more. preferred less than 10.10-10 m.

In addition, the total proportion of platinum relative to the total weight of catalyst is generally between 0.1 and 1% and preferably between 0.1 and 0.6%.

2o In the case of the use of nickel, the proportion of nickel relative to the total weight of catalyst is between 5 and 70%, more particularly between 10 and 70% and preferably between 15 and 65%.

In addition, a catalyst is generally used such that the average size of the nickel crystallites is less than 100 × 10 × 10 m, preferably less than 80 × 10 −0 m, even more preferably less than 60 × 10 10 m.

The support is generally chosen from the group formed by alumina, silica alumina, silica, zeolites, activated carbon, clays, aluminous cements, rare earth oxides and alkaline earth oxides, alone or as a mixture. .

A support based on alumina or silica is preferably used, with a specific surface area of ​​between 30 and 300 m2 / g, preferably between 90 and 260 m2 / g.

The invention finally relates to a process for the treatment of a feed, mainly comprising olefinic hydrocarbons comprising 4 carbon atoms per molecule, including isobutene, as well as butene-1 and butenes-2 in a ratio corresponding substantially to thermodynamic equilibrium, in which said feed is treated in a distillation zone, comprising an exhaustion zone and a rectification zone, associated with a hydroisomerization reaction zone, is treated ~ 19 ~~ 07 ~ 14, said reaction zone being at least partly internal to said distillation zone and comprising at least one catalytic bed, in which the hydroisomerization of at least a part, preferably most of the butene-1, is carried out in the presence of a hydroisomerization catalyst and a gas stream comprising, preferably for the most part, hydrogen, so as to finally exit at the top of the distillation zone an effluent rich in isobutene, generally of high purity, and at the bottom of the distillation zone an effluent depleted in isobutene, said process being t0 characterized in that any catalytic bed of the internal part of the hydroisomerization zone is passed through in ascending co-current by said gas flow and the liquid and is practically not crossed by the distillation vapor.

Said process allows the production of isobutene of high purity.

Preferably, the method according to the invention comprises the use of the device according to the invention.

The feed which feeds the distillation zone is introduced into said zone generally at least at one level of said zone, preferably mainly at a single level of said zone.

It is in a ratio corresponding substantially to the thermodynamic equilibrium butene-1 - butenes-2 on introduction.

One of the preferred embodiments of the process according to the invention comprises obtaining said feedstock from a cut comprising mainly olefinic hydrocarbons comprising 4 carbon atoms per molecule, including isobutene and butene1, by treatment of said cut in a first hydroisomerization zone, generally independent of the possible external part of the hydroisomerization reaction zone associated with the distillation zone, the major part of the effluent from said first hydroisomerization zone then serving as feed, main or secondary according to the definitions given below in the text, which feeds the distillation zone.

If said feed comprises polyunsaturated compounds, most often diene and / or acetylenic, said compounds are preferably converted into butenes by the first hydroisomerization zone before introduction into the distillation zone.

But any other technique making it possible to obtain, from a cut mainly comprising C4 olefinic hydrocarbons, a feed where the butene-1 and the butenes-2 are in a ratio corresponding substantially to the thermodynamic equilibrium can also be envisaged in within the scope of the invention.

2 ~. ~~~~ 6 t5 The first optional hydroisomerization reaction zone, located upstream of the distillation-reaction zone, carries out at least partially the selective hydrogenation of the polyunsaturated compounds, most often diene such as butadiene, in addition to the hydroisomerization of at least part of butene-1 to butenes2.

It generally comprises at least one hydroisomerization catalytic bed comprising a hydroisomerization catalyst, preferably from 1 to 4 catalytic beds); in the case where at least two catalytic beds are incorporated in said reaction zone, these two beds are preferably distributed in at least two reactors, distributed in series or in parallel, preferably in series.

For example said first ~ 0 reaction zone comprises a single reactor in which there is at least one, and preferably only one, catalytic bed.

One of the preferred implementations of the process of the present invention is such that said first reaction zone comprises two reactors generally distributed in series each comprising at least one, preferably only one, catalytic bed.

When said reaction zone comprises at least two reactors, the optional recycling of at least part of the effluent from at least one of the reactors included in the first reaction zone in said first zone is generally carried out at the inlet of a reactor, preferably of said reactor, preferably before the injection of the gaseous compound comprising hydrogen. It is also possible to recycle around said first zone itself, that is to say generally at the inlet of the first reactor of said zone, preferably before the injection of the gaseous compound comprising. hydrogen; for example, in the case of two reactors, at least part of the effluent from the second reactor is recycled to the inlet of the first reactor.

This advantageously makes it possible to lower the content of polyunsaturated compounds in the effluent from said first reaction zone.

The operating conditions of the first hydroisomerization zone, when it is present, are generally as follows: the catalyst is identical to the catalyst of the hydroisomerization zone which will be described below.

The pressure is generally between 4 and 40 bar (1 bar = 0.1 MPa), preferably between 6 and 30 bar.

The temperature is generally between 10 and 150 ° C, preferably between 20 and 100 ° C.

The H2 / hydrocarbon molar ratio is generally adjusted so as to obtain practically total conversion of polyunsaturated compounds such as butadiene and sufficient isomerization of butene-1 to butenes-2 with limited formation of alkanes.

The hydroisomerization reaction zone associated with the distillation zone generally comprises at least one hydroisomerization catalytic bed comprising a hydroisomerization catalyst, preferably from 2 to 4, even more preferably from 2 to 6 catalytic beds; in the case where at least two catalytic beds are incorporated in said distillation zone, these two beds are preferably separated by at least one internal distillation.

Said hydroisomerization reaction zone at least partially performs the hydroisomerization of at least part, preferably the major part, of butene-1 present in its feed in butenes-2 (cis and trans), generally in such a way that the butene1 content of 1o the overhead effluent from the distillation zone is at most equal to a certain content.

The distillation zone used in the process according to the invention is identical to that described above.

One of the preferred implementations of the process according to the invention comprises supplying the distillation zone, in addition to the supply of the main feed, with a so-called secondary feed (with respect to the main feed), which comes from or not of a hydroisomerization reaction zone such as the first optional hydroisomerization reaction zone 2o, independently or not of the supply of the distillation zone with the main feed.

The secondary feed is generally a C4 cut containing at least isobutene, as well as butene-1 and butenes-2 in a ratio corresponding substantially to thermodynamic equilibrium, and is generally obtained from a steam cracking process, such as the crude C4 cut or raffinate-1, or catalytic cracking;

generally and preferably, the secondary feed is a C4 cut essentially free of polyunsaturated compounds and its butene-1 content is lower than the butene-1 content of the main feed.

If the content of unsaturated compounds in the secondary feed is high, said feed is preferably treated in a selective hydrogenation zone before entering the distillation zone.

When the main feed is introduced in a single level of introduction, the secondary feed is generally introduced into the distillation zone in at least one level of introduction, preferably in a single level of introduction, said dependent level of introduction. of the composition of said secondary filler.

Thus, in a first example the secondary charge may be very rich in isobutene and contain 219 ~~ '~ 6 17 less than 1.5 times of butene-1 than the main charge does not contain, in which case the secondary charge is preferably introduced in a single level generally located above the level of introduction of the main charge.

Or in a second example the secondary charge may be substantially free of butene-1, in which case the secondary charge is preferably introduced in a single level generally located below the introduction level of the main charge. It is also possible to proceed with the possible mixing of the main feed, before it enters the distillation zone, and of the secondary feed.

~ o The hydroisomerization reaction zone associated with the distillation zone generally comprises at least one hydroisomerization catalytic bed, preferably from 2 to 6, even more preferably from 2 to 4 catalytic beds); in the case where at least two catalytic beds are incorporated in the distillation zone, these two beds are optionally separated by at least one internal of ~ 5 distillation.

The hydroisomerization reaction zone at least partially performs the hydroisomerization of at least a part, preferably the major part, of the butene-1 present in the feed to butenes-2 (cis and trans), generally in such a way that the butene-1 content of the overhead effluent is at most equal to a certain content.

The process according to the invention is generally such that the flow of the liquid to be hydroisomerized is co-current with the flow of the gas flow comprising hydrogen, for any catalytic bed of the internal part of the hydroisomerization zone, and such that the distillation vapor does not pass practically any catalytic bed of the internal part of the reaction zone (which results in practice by the fact that said vapor is separated from said liquid to be hydroisomerized).

Any catalytic bed of the part of the reaction zone internal to the distillation zone is generally such that the gas flow comprising hydrogen and the flow of the liquid which will react circulate in co-current, generally ascending, through said bed, 3o even if overall, in the catalytic distillation zone, the gas flow comprising hydrogen and the flow of the liquid which will react flow countercurrently.

Such systems generally include at least one liquid distribution device which may for example be a liquid distributor, for any catalytic bed in the internal part of the reaction zone.

The devices for distributing the gas flow and distributing the liquid have been described previously.

2194076 _ I8 The process according to the invention is generally such that the charge of any part of the hydroisomerization reaction zone, whether internal or possibly external, is taken at the level of a sampling level and represents at least a part, preferably the major part, of the liquid (reflux) flowing in the distillation zone, preferably flowing in the rectification zone and even more preferably flowing at an intermediate level of the rectification zone, the effluent from the hydroisomerization reaction zone being at least in part, preferably in major part, reintroduced into the distillation zone, so as to ensure the continuity of the distillation.

For the possible part of the reaction zone external to the distillation zone, the reintroduction of the effluent into the distillation zone takes place substantially near, that is to say generally substantially at the same height or substantially above. or substantially below, most often substantially at the same height or substantially above, that is to say located at a distance corresponding to a height between 0 and 4 theoretical levels of a sampling level, preferably said sampling level, this to ensure the continuity of the distillation.

For the part of the reaction zone internal to the distillation zone, the liquid withdrawal (reflux) is done naturally by flow into the part of the reaction zone internal to the distillation zone, and the reintroduction of the effluent into the distillation zone. Distillation is also done naturally by 2o flow of the liquid from the reaction zone internal to the distillation zone.

Generally, when the hydroisomerization zone is not entirely internal to the distillation zone, the method according to the invention comprises from 1 to 6, preferably from 2 to 4 levels) of sampling which feed) the external part of the hydroisomerization zone.

In such a case, the liquid to be hydroisomerized, either partially or totally, first circulates in the external part of the hydroisomerization zone and then in the internal part of said zone.

So, two cases can arise.

In the first case, the external part of the hydroisomerization zone is supplied by a single sampling level, and then, if said part comprises more than two reactors, these are arranged in series or in parallel.

In the second case, preferred according to the present invention, the external part of the hydroisomerization zone is supplied by at least two sampling levels.

A part of the external part of said hydroisomerization zone which is fed by a given offtake level, if said outer part comprises at least two offtake levels, generally comprises at least one reactor, preferably a single reactor.

If said part of the external part comprises at least two reactors, each reactor external to the distillation zone is generally supplied by a single level of withdrawal, preferably associated with a single level of reintroduction, said level of withdrawal being distinct from the level of offtake which feeds the other reactor (s).

The process according to the invention is, according to one of the preferred embodiments of the invention, such that the hydrogenation zone is entirely internal to the distillation zone.

The hydrogen used according to the invention for the hydroisomerization of butene-1 generally originates for the most part, preferably almost entirely, from outside the distillation zone. It can be from any source producing hydrogen at at least 50% by volume purity, preferably at least 80% by volume purity, and more preferably at least 90% by volume purity.

For example, we can cite hydrogen from catalytic reforming processes, from PSA

(pressure swing adsorption), electrochemical generation, steam cracking or steam reforming.

2o The operating conditions of the part of the hydroisomerization zone internal to the distillation zone are linked to the operating conditions of the distillation.

The distillation is generally carried out so as to minimize the amount of isobutene in the bottom product in order to maximize the process yield of isobutene and in such a way as to minimize the amount of butenes-2 and butene-1 in the top product, 25 in order to have high purity isobutene in mind.

It is carried out under a pressure generally between 2 and 30 bar, preferably between 4 and 15 bar, even more preferably between 4 and 10 bar, with a reflux rate between 1 and 30, and preferably between 5 and 20.

The zone head temperature is generally between 0 and 200 ° C and the zone bottom temperature is generally between 5 and 250 ° C.

The hydroisomerization reaction is carried out under conditions which are most generally intermediate between those established at the top and at the bottom of the distillation zone, at a temperature between 20 and 150 ° C, and preferably between 40 and 80 ° C. , and at a pressure between 2 and 30 bar, preferably between 4 and 15 bar, even more preferably between 4 and 35 bar.

The liquid subjected to hydroisomerization is supplied by a gas flow comprising, preferably for the most part, hydrogen.

When the hydroisomerization zone comprises a part external to the distillation zone, the catalyst placed in said external part is placed according to any technology known to those skilled in the art under operating conditions (temperature, pressure, etc.) generally. independent of the operating conditions of the distillation zone.

In the optional part of the hydroisomerization zone external to the distillation zone, the operating conditions are generally as follows.

The pressure required for this hydroisomerization step is generally about 1 to 40 bar absolute, preferably about 2 to 30 bar and even more preferably about 4 to 25 bar.

The operating temperature of the hydroisomerization zone is generally about 20 to 150 ° C, preferably about 40 to 100 ° C and more preferably about 40 to 80 ° C.

The space velocity within said hydroisomerization zone, calculated relative to the catalyst, is generally about 1 to 100 and more particularly about 4 to 50 h-1 (volume of feed per volume of catalyst and per hour) .

The corresponding hydrogen flow rate 15 is such that the molar ratio of H2 / hydrocarbons entering the hydroisomerization zone is preferably at least equal to 10-5.

This ratio is most often about 10'5 to about 3 and very frequently about 10-3 to about 1.

But the temperature and pressure conditions can also, in the context of the process of the present invention, be between those which are established at the top 2o and at the bottom of the distillation zone.

For carrying out the hydroisomerization according to the process of the invention, the theoretical molar ratio of hydrogen necessary for the desired conversion of butene1 in the reaction zone associated with the distillation zone is such that the molar ratio of H2 / incoming hydrocarbons. in said zone is at least equal to 10-5.

This molar ratio can be optimized in such a way that on the one hand all the hydrogen is consumed in the hydroisomerization reaction in order to avoid a device for recovering the hydrogen at the outlet of said reaction zone, on the other hand to minimize the side reactions of hydrogenation of isobutene in order to maximize the 3o yield of the isobutene process and finally in such a way that there is sufficient hydrogen all along said reaction zone so that the hydroisomerization reaction of butene-1 to butenes-2 can take place.

However, if the conditions are such that there is an excess of hydrogen, the excess hydrogen can advantageously be recovered, for example according to one of the techniques described below.

35 For example, the excess hydrogen which leaves the head of the distillation zone is recovered, then injected upstream of the compression stages associated with a catalytic reforming unit, mixed with the hydrogen coming from said unit, said unit preferably operating at low pressure (ie generally a pressure of less than 8 bar).

This excess hydrogen can also be recovered, then compressed and reused in said reaction zone.

In the case of the presence of an external part of the hydroisomerization zone associated with the distillation zone, the process according to the invention makes it possible to isomerize a large part of the butene-1 into butenes-2 outside of the distillation zone, optionally under pressure and / or temperature conditions different from those used in the column.

Preferably. the temnératurP at the PntrPa (respectively at the outlet) of the sampling level which feeds a catalytic bed of the part of the hydroisomerization zone located outside the column, is substantially similar, that is to say that the difference is appreciably less than 10 ° C compared to the temperature at the height of the sampling level (respectively of the reintroduction level).

Likewise, the hydroisomerization reaction can advantageously be carried out in the part of the reaction zone located outside the column at a higher pressure than that used inside the distillation zone.

This pressure increase also allows an increased dissolution of the gas stream containing hydrogen in the liquid phase containing the butene1 to be isomerized.

In such a case, the method according to the invention comprises the use of the so-called "pump-around" technique, that is to say of pumping in a loop, which consists in passing outside the zone. distillation a part, preferably the major part, of the liquid (reflux) by a factor preferably greater than 1, i.e. the flow rate of a catalytic bed from the external part of the hydroisomerization zone associated with the distillation zone, said bed being supplied to a sampling level with part of the liquid effluent (reflux) flowing on the distillation plate associated with said sampling level (ie on which said portion of liquid effluent is taken) and by at least a portion of the liquid corresponding to the recycling of the effluent from said bed substantially above or substantially below or substantially at the same height as said sampling level, is greater than 1 times the flow rate of liquid flowing on said plate, for example equal to 1.5 times.

More generally, the catalyst used in the hydroisomerization zone according to the process of the present invention generally comprises at least one metal chosen 219 ~ 07 ~ 22 from the group formed by the noble metals of group VIII of the Periodic Table of the Elements and nickel, that is to say chosen from the group formed by ruthenium, rhodium, palladium, osmium, iridium, platinum, preferably palladdium, or nickel, used as such or preferably deposited on a support.

The metal should generally be in reduced form for at least 50% by weight of its totality.

The noble metal content of the catalyst is usually about 0.01 to about 2% by weight.

In the case of the use of nickel, the proportion of nickel relative to the total weight of catalyst is between 5 and 70%, and preferably between 10 and 70%, and a catalyst such as size 10 is generally used. average nickel crystallites is less than 10 nm, preferably less than 8 nm, even more preferably less than 6 nm.

But any other hydroisomerization catalyst known to those skilled in the art can also be chosen.

The catalyst is usually treated with a sulfur compound and then with hydrogen before it is used.

The catalyst is generally sulfurized in situ or ex situ such that sulfur is chemisorbed on at least part of the metal.

The chemisorbed sulfur has the effect of promoting the isomerization reaction of butene-1 to butenes-2 over the hydrogenation reaction of isobutene and therefore of maximizing the isobutene yield of the process.

2o The support of the hydroisomerization catalyst is generally chosen from the group formed by alumina, silica-aluminas, silica, zeolites, activated carbon, clays, aluminous cements, rare earth oxides and oxides alkaline earth, singly or as a mixture.

A support based on alumina or silica, with a specific surface area of ​​between 10 and 300 m2 / g, preferably between 30 and 70 m2 / g, is preferably used.

By way of nonlimiting example which can be used in the context of the present invention, mention may be made of commercial catalysts such as that sold by the company Catalysts and Chemicals under the reference C-31, that sold by the company Girdler Corporation under the reference G- 55 or preferably those sold by the company Procatalyse under the reference LD-265, LD-265S, LD-267 and LD-267R.

Examples Examples 1 and 2 which follow show the use of a zone for the hydrogenation of unsaturated compounds comprising at most six carbon atoms 21940 "~~ 23 per molecule including benzene, according to the invention (example 2), and the use of a hydrogenation zone which is not according to the invention (example 1), with a catalyst placed in bulk on distillation trays crossed by the liquid which circulates in the descending direction and by the vapor which circulates in the ascending direction of said distillation zone.

Example 1 (Comparative) A metal distillation column with a diameter of 50 mm is used, made to adiabatic by heating jackets, the temperatures of which are regulated so as to reproduce the temperature gradient which is established in the column.

On a height of 4.5 m, the column comprises, from the head to the foot: a rectification zone composed of 11 perforated plates with overflow and descent, a hydrogenating catalytic distillation zone and a depletion zone composed of 63 ~ 5 perforated trays.

The hydrogenating catalytic distillation zone consists of three reactive trays, which are here perforated distillation trays, with weir and descent, the weirs of which are raised by 3.5 cm and the volume of which lies between the level of the top of the weir and the tray can be lined with catalyst.

A metal screen placed at the top of the weir acts as a filter to prevent catalyst particles 2o from being evacuated with the liquid leaving the tray.

Each of the three cells is packed with 36 g of nickel catalyst sold by the company PROCATALYSE under the reference LD 746.

260 g / h of a reformate comprising essentially hydrocarbons having at least 5 carbon atoms in their molecule, the composition of which is presented in Table 1, are introduced on the 37th plate of the column, starting from the bottom.

A flow rate of 18 NI / h of hydrogen is also introduced at the base of each cell.

The column is brought into operation by establishing a reflux rate equal to 5 and by regulating the bottom temperature at 176 ° C. and the pressure at 7 bar.

Under a stabilized regime, a residue and a liquid distillate, the compositions of which are given in Table 1, are collected at a rate of 138 g / h and 113 g / h, respectively.

A small part of the distillate, consisting of the lighter hydrocarbons, is discharged from the column with the excess hydrogen and is not counted.

From analyzes of the effluents it can be deduced that the hydrogenation rates of the olefins and benzene in the feed are 100 and 55%, respectively, while the toluene is not affected.

Example 2: (according to the invention) The same apparatus as that described in Example 1 is used, but with a catalytic distillation zone of different design.

The hydrogenating catalytic distillation zone here consists of three catalytic distillation doublets, each doublet itself being constituted by a catalytic cell surmounted by three perforated trays.

The construction detail of a catalytic cell as well as its arrangement in the column are shown schematically in the figure.

The catalytic cell 1 consists of a cylindrical container with a flat bottom, with an outside diameter less than 2 mm than the inside diameter of the column.

It is provided at its lower part, above the bottom, with a grid 2 which serves both as a support for the catalyst and as a distributor for the hydrogen, and at its upper part, with a grid for retaining the catalyst. catalyst 3, the height of which can be varied.

Catalyst 4 fills the entire volume between these two grids.

The catalytic cell receives the liquid coming from the upper distillation plate 5, by the descent 6.

After having traversed the cell in the ascending direction, the liquid is discharged by overflow 20 by the descent 7 and flows on the lower distillation plate 8.

The steam coming from the lower plate 8 passes through the central chimney 9 integral with the cell, entering through orifices 10 (only one visible in the figure) and leaving under the upper plate 5 through orifices 11 (only one visible in the figure). ).

The hydrogen is introduced at the foot of the catalytic cell through the pipe 12, then through the orifices 13 (six in total) distributed over the periphery of the cell, in the immediate vicinity of the bottom.

Seals 14 prevent any leakage of hydrogen before it reaches the catalytic bed.

Each of the three cells is packed with 36 g of nickel catalyst sold by the company PROCATALYSE under the reference LD 746.

On the 37th plate of the column, starting from the bottom, 260 g / h of the same load as that used in Example 1 and the composition of which is presented in the second column of the table are introduced.

A flow rate of 6 Nl / h of hydrogen is also introduced at the base of each cell.

The column is brought into operation by establishing a reflux rate equal to 5 and by regulating the bottom temperature to 176 ° C. and the pressure to 7 bar.

In a stabilized regime, a residue and a liquid distillate, the compositions of which are given in Table 1, are collected at a rate of 143 g / h and 106 g / h, respectively.

A small part of the distillate, made up of the lightest hydrocarbons, is discharged from the column with the excess hydrogen and is not counted.

From analyzes of the effluents it can be deduced that the hydrogenation rates of the olefins and of the benzene in the feed are respectively 100 and 87%, while the toluene is not affected.

Table 1: compositions of the feed and of the effluents from the catalytic column, in% by weight, example 1 example 2 feedstock residue of the distillate residue of the liquid distillate. liq.

C5 and lighter 7.65 10.22 7.36 of which: olefins 0.11 0 0 C6 44.839.55 89.78 12.4 92.59 of which: olefins 0.13 0 0 benzne 6.07 0.63 5 , 45 0.07 1.84 cyclohexane 1.1 8.34 0.34 12.16 0.73 C7: 42.5580.72 78.27 0.05 of which: tolune 4.78 9.1 8.87 C8 and heavier 4.97 9.73 9.33 conversion of 100% 100% olefins conversion of 55% 87% benzne conversion of 15% 70% hydrogen to II is observed that the process according to the present invention allows better conversion of benzene and better conversion of hydrogen.

. , t 26 Examples 3 and 4 which follow illustrate the case of a process according to the invention for the treatment of a feed mainly comprising olefinic hydrocarbons comprising 4 carbon atoms per molecule, including isobutene, including butene. -1 and butenes-2 in a ratio corresponding substantially to thermodynamic equilibrium.

Example 3 The operations of hydroisomerization of a C4 cut and of distillation were carried out successively in a batch manner.

The charge was hydroisomerized for the first time. The effluent from this first test was distilled; the distillation head, representative of an intermediate withdrawal, was hydroisomerized. The hydroisomerization effluent, representative of what would be reinjected into the column, was distilled.

The head of this second distillation was hydroisomerized, and the effluent from this third hydroisomerization was distilled.

The hydroisomerization operations are carried out in a pilot unit having an adiabatic reactor.

The reactor is filled with 1.5 l of the LD265 catalyst sold by the company Procatalyse.

The catalyst is sulfurized and activated in situ according to a procedure recommended by the supplier of the catalyst.

The distillation operations were carried out in an adiabatic column with an internal diameter of 163 mm and a height of 10 m.

The column consists of 4 beds 1.78 m high above the feed injection, filled with a packing sold by the company Sulzer under the name M550Y and 2 beds 1 m high below. injection of the charge, filled with Pall rings.

- First hydroisomerization.

3o The average operating conditions during the test are as follows ~ Reactor temperature: 80 ° C ~ Reactor pressure: 20 bar ~ Residence time: 0.25 h.

~ H2 / feed molar ratio: 3 ~ '~ ~' ~ 6 27 Table 2 below shows the compositions of the feed and the effluent of the hydroisomerization reactor operating under the conditions described above.

TABLE 2 Load (% weight Effluent% oids <C4 0.25 0.23 i C4 2.98 3.10 i C4 = 44.90 44.42 C4 = 1 26.95 4.26 C4 == 1.3 0 , 13 0.00 n C4 11.72 14.41 C = 2 trans 8.73 21.37 No C 0.24 0.23 Me C clo C 0.06 0.06 C4 = 2 cis 4.03 11, 92> C 0.01 0.00 with the following legend for this table and for the following tables <C4: composed of less than 4 (4 excluded) carbon atoms per molecule (or C3-) t0 iC4 v isobutane i C4 =: isobutne C4 = 1: butne-1 C4 == 1,3: butadine-1,3 n C4: normal-butane t5 C4 = 2 trans: butne-2 trans No C5: nopentane (or dimthyl propane) Me Cyclo C3: methyl cyclopropane C4 = 2 cis: butne-2 cis C4: compound more than 4 (4 excluded) carbon atoms per 20 molecule (or C5 +) 219 4 0 '~ 6' - First distillation.

The distillation of the effluent from the test presented above was carried out with the following operating conditions ~ Column pressure: 4 bar ~ Reflux rate (R / D): 20 ~ Temperature of the charge: 33 ° C ~ Reflux temperature: 32 ° C lo ~ Temperature at the top of the column: 57 ° C ~ Temperature at the bottom of the column: 63 ° C.

Table 3 below shows the compositions of the feed and of the overhead effluent from the distillation column operating under the conditions described above.

TABLE 3 Char e% oids Tte% oids <C4 0.23 0.44 i C4 3.10 6.71 i C4 = 44.42 83.35 C4 = 1 4.26 7.39 C4 == 1.3 ~ 0.00 0.00 n C 14.41 1.62 C4 = 2 trans 21.37 0.44 No C 0.23 Me C clo C 0.06 C4 = 2 cis 11.92 0.05> C4 _ _ - Second hydroisomerization.

The average operating conditions during the test are as follows ~ Reactor temperature: 65 ° C 29 ~ Reactor pressure: 20 bar ~ Residence time: 0.25 h ~ H2 / feed molar ratio: 0.6 The table 4 below shows the compositions of the feed and of the effluent from the hydroisomerization reactor operating under the conditions described above.

TABLE 4 Char e% oids Effluent% oids <C 0.44 0.39 i C4 6.71 6.91 i C4 = 83.35 82.94 C4 = 1 7.39 0.81 C4 == 1 ~ 3 - n C4 1.62 2.09 C4 = 2 trans 0.44 4.44 No C - _ Me C clo C - _ C4 = 2 cis 0.05 2.42> C4 - _ 10. Second distillation.

The distillation of the effluent from the test presented above was carried out with the following operating conditions ~ Column pressure: 4 bar ~ Reflux rate (R / D): 13.5 ~ Temperature of the charge: 36 ° C ~ Reflux temperature: 41 ° C ~ Temperature at the top of the column: 51 ° C ~ Temperature at the bottom of the column: 55 ° C.

Table 5 below shows the compositions of the feed and overhead effluent from the distillation column operating under the conditions described above.

TABLE 5 to Char e% oids Tte% oids <C4 0.39 0.65 i C4 6.91 13.71 i C4 = 82.94 84.82 C4 = 1 0.81 0.51 C4 == 1 ~ 3 - n C4 2.09 0.14 C4 = 2 trans 4.44 0.12 No C - _ Me C clo C - C4 = 2 cis 2.42 0.05> C4 _ _ Third hydroisomerization.

The average operating conditions during the test are as follows ~ Reactor temperature: 60 ° C ~ Reactor pressure: 20 bar ~ Residence time: 0.25 to 0.1 h ~ H2 / feed molar ratio: 1 Le Table 6 below shows the compositions of the feed and of the effluent from the hydroisomerization reactor operating under the conditions described above.

'- 31 TABLE 6 Load% oids Effluent% oids <C4 0.65 0.57 i C4 13.71 14.55 i C4 = 84.82 84.07 C4 = 1 0.51 0.03 C4 == 1 ~ 3 - _ n C4 0.14 0.22 C4 = 2 trans 0.12 0.38 No C _ _ Me C clo C - _ C = 2 cis 0.05 0.18> C4 _ _ - Third distillation.

The distillation of the effluent from the test presented above was carried out with the following operating conditions ~ Column pressure: 4 bar lo ~ Reflux rate (R / D): 13.5 ~ Temperature of the charge: 36 ° C ~ Reflux temperature: 41 ° C ~ Temperature at the top of the column: 53 ° C ~ Temperature at the bottom of the column: 55 ° C.

Table 7 below shows the compositions of the feed and of the overhead effluent from the distillation column operating under the conditions described above.

..

32 TABLE 7 Char e% oids tte% oids <C4 0.57 0.57 i C 14.55 14.66 i C4 = 84.07 84.69 C4 = 1 0.03 0.03 C4 == 1 ~ 3 - n C4 0.22 0.01 C4 = 2 trans 0.38 0.04 No C - _ Me C clo C - C = 2 cis 0.18> C4 - _ These successive and discontinuous hydroisomerization and distillation operations represent the operation of separating butene-1 from isobutene which is carried out continuously in the case of the process according to the invention.

Example 4 1o Pilot hydroisomerization tests were carried out using raffinate-1 on the hydroisomerization catalyst LD267R marketed by the company Procatalyse which lines each of the catalytic beds.

The results of these tests are shown in Table 8 below; they made it possible to determine the calculation parameters which make it possible to simulate the method according to the invention by means of suitable software.

The software t5 used for this simulation is marketed under the name Pro2 by the company SIMCI.

. ~ ~~~ 1 ~~! '~ 6 33 TABLE 8: pilot test results TC 40 60 90 50 50 50 50 50 50 50 VVH 30 30 30 30 30 30 30 30 20 40 h-1 P 10 10 10 6.5 10 15 1 10 10 10 bar 0 H2 / H 0.17 0.17 0.17 0.170.17 0.17 _ 0.19 0.17 0.17 C 0.1 m / m char effl effl effl effleffl effl effl effl effl effl <0.14 0.11 0.12 0 , 11 0.100.10 0.10 0.10 0.11 0.10 0.09 C4 iC4 5.69 5.75 5.75 5.73 5.715.76 5.75 _ 5.76 5.75 5.74 5.7 2 iC4 = 78.6 78.7 78.7 78.7 78.778.7 78.7 _ 78.7 78.7 78.7 7 1 2 3 8 2 3 78.7 2 1 4 4 1- 3.66 1.30 0.91 0.75 1.151.01 1.18 1.13 1.00 0.8 1.32 C4 = n-C4 7.16 7.19 7.17 7.14 7.147.18 7.19 7.16 7.20 7.19 7.18 tr2- 4.36 5 , 40 5.48 5.40 5.465.49 5.41 5.46 5.48 _ 5.37 C4 = 5.59 cs2- 0.32 1.54 1.85 2.14 1.661.74 1.64 1 , 69 1.75 1.86 1.56 C4 = or cnar = cnarge and ettt = etnuent The hydroisomerizing catalytic distillation zone comprises 2 or 3 catalytic distillation doublets, each of its doublets being of the type shown in the figure, each doublet being itself constituted by a catalytic cell surmounted by three perforated plates.

Two examples thus simulated by the calculation were carried out.

They are described below.

Example 4A The configuration of the unit, comprising three t5 hydroisomerization catalytic beds located inside the column, called reactive trays, is as follows ~ column of 130 theoretical trays, numbered from top to bottom, ~ tray of power supply n ° 90, ~ the reagent trays are trays 10, 25 and 39.

They each contain 7.5 m3 of catalyst.

- 34 Operating conditions liquid flow rate of the column feed: 292.9 kmol / h reflux rate: 12, pressure at the top of the column: 6.2 bar absolute, pressure at the bottom of the column: 7 bar absolute, temperature of l '' column feed: 59 ° C, temperature at the top of the column: 52 ° C, temperature at the bottom of the column: 64.5 ° C, temperature of reagent No.10: 53 ° C, lo pressure of reagent N ° 10: 6.6 bar absolute, liquid flow through reagent plate 10: 1660 kmole / h, temperature of reagent plate 25: 54 ° C, pressure of reagent plate N ° 25: 6.6 bar absolute, liquid flow through reagent plate 25: 1660 kmole / h, temperature of reagent plate N ° 39: 54 ° C, pressure of reagent plate 39: 6.7 bar absolute, liquid flow rate through reagent plate 39: 1660 kmole / h.

With this configuration and under these operating conditions, the simulation led to the following results Column head supply Column bottom column kmole / hkmole / h kmole / h C4 1.12 1.12 0.00 i C4 4.46 5.58 0 , 00 i C = 110.08 108.07 0.89 C4 = 1 7.53 0.02 0.17 n C4 55.27 0.13 55.23 C4 = 2 tr 79.76 0.03 84.56 C4 = 2 cis 33.49 0.00 35.91 1.21 0.00 0.00 Total 292.92 114.95 176.76 Yield of isobutene at the top of the column: 98.2 Mole ratio butene-1 / isobutene at the head of the column: 1.85 x 10-4 - _ 35 Example 4B The configuration of the unit, comprising two hydroisomerization catalytic beds located inside the column, called reactive trays, is as follows ~ column of 130 theoretical trays, numbered from top to bottom, ~ feed tray n ° 90, ~ the reagent trays are trays 10 and 39.

They each contain 7.5 m3 of catalysts.

t0 Operating conditions column feed liquid flow rate: 292.9 kmol / h reflux rate: 12, column head pressure: 6.2 bar absolute, column bottom pressure: 7 bar absolute, IS temperature of l '' column feed: 59 ° C, temperature at the top of the column: 52 ° C, temperature at the bottom of the column: 64.5 ° C, temperature of reagent No. 10 plate: 53 ° C, pressure of reagent No. 10: 6.6 bar absolute, 2nd liquid flow through the reagent tray 10: 1660 kmole / h, temperature of reactive plate 39: 54 ° C, pressure of reactive plate 39: 6.7 bar absolute, liquid flow through reagent 39 plate: 1660 kmol / h.

25 With this configuration and under these operating conditions, the simulation led to the following results ~~~~ o ~~ 36 Column head power Column head column kmole / hkmole / h kmole / h <C4 1.12 1.12 0, 00 i C4 4.46 5.17 0.00 i C4 = 110.08 108.48 0.89 C = 1 7.53 0.09 0.17 n C4 55.27 0.13 55.20 C4 = 2 tr 79.76 0.06 84.50 C = 2 cis 33.49 0.00 35.90 H2 1.21 0.44 0.00 Total 292.92 115.49 176.66 Isobutene yield at the top of the column : 98.6 Butene-1 / isobutene molar ratio at the top of the column: 8.30 x 10-4