Apparatus comprising a catalytic distillation zone comprising a reaction zone with distribution of hydrogen

The invention concerns a reactive distillation apparatus comprising a distillation zone, associated with a reaction zone which is at least in part internal to said distillation zone and comprises at least one catalytic bed in which the feed is transformed in the presence of a catalyst and at least one gas stream containing hydrogen, characterized in that each catalytic bed in the internal portion of said reaction zone is traversed by said gas stream and liquid in ascending co-current mode. The invention also concerns selective hydrogenation processes for light unsaturated hydrocarbons, mainly any olefins and benzene, comprised in a mixture the major portion of which is constituted by hydrocarbons containing at least five carbon atoms per molecule, and the hydroisomerisation of at least a portion of the 1-butene contained in a feed the major portion of which is constituted by olefinic hydrocarbons including isobutene, also 1-butene and 2-butenes in a ratio which substantially corresponds to the thermodynamic equilibrium.

The invention concerns a reactive distillation apparatus comprising a 
distillation zone, associated with a reaction zone which is at least in 
part internal to said distillation zone and comprises at least one 
catalytic bed in which the feed is transformed in the presence of a 
catalyst and at least one gas stream containing hydrogen, characterized in 
that each catalytic bed in the internal portion of said reaction zone is 
traversed by said gas stream and liquid in ascending co-current mode. The 
invention also concerns selective hydrogenation processes for light 
unsaturated hydrocarbons, mainly any olefins and benzene, comprised in a 
mixture the major portion of which is constituted by hydrocarbons 
containing at least five carbon atoms per molecule, and the 
hydroisomerisation of at least a portion of the 1-butene contained in a 
feed the major portion of which is constituted by olefinic hydrocarbons, 
including isobutene, also 1-butene and 2-butenes in a ratio which 
substantially corresponds to the thermodynamic equilibrium. 
In addition to selective hydrogenation of the light unsaturated compounds 
of a reformate including olefins and benzene, without notable 
hydrogenation of heavier unsaturated compounds such as toluene and, a 
fortiori, xylenes, the apparatus of the invention can be applied to 
various catalytic reactions, which may be equilibrated or complete, in 
which at least one of the products of the reaction can be separated in a 
pure or diluted state by distillation under temperature and pressure 
conditions which are close to those of the reaction, and more particularly 
to paraffin isomerisation reactions involving reorganisation of the 
backbone, olefin isomerisation by displacement of the double bond 
(hydroisomerisation) or by reorganisation of the backbone, hydrogenation 
of unsaturated compounds to saturated compounds, dehydrogenation of 
saturated compounds to unsaturated compounds, and any reaction which 
requires the presence of hydrogen. 
The hydrogenation catalyst can be disposed in the reaction zone using the 
different techniques proposed for carrying out catalytic distillation. 
These techniques have principally been developed for etherification 
reactions, which involve contact between the reactants in a homogeneous 
liquid phase and the solid catalyst. They are essentially of two types. In 
the first type of technique, the reaction and the distillation proceed 
simultaneously in the same physical space, as taught, for example, in 
International patent application WO-A-90/02603, U.S. Pat. Nos. 4,471,154, 
4,475,005, 4,215,011, 4,307,254, 4,336,407, 4,439,350, 5,189,001, 
5,266,546, 5,073,236, 5,215,011, 5,275,790, 5,338,517, 5,308,592, 
5,236,663, 5,338,518 and European patents EP-B1-0 008 860, EP-B1-0 448 
884, EP-B1-0 396 650 and EP-B1-0 494 550, also European patent application 
EP-A1-0 559 511. In general, then, 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 vapour phase generated by the 
reboiling vapour introduced at the bottom of the zone. In the second type, 
the catalyst is generally disposed so that the reaction and distillation 
proceed independently and consecutively as taught, for example, in U.S. 
Pat. Nos. 4,847,430, 5,130,102 and 5,368,691, the vapour from the 
distillation step not in practice traversing any catalytic bed in the 
reaction zone. 
For every chemical reaction requiring the addition of a foreign gaseous 
reactant to the distillation feed, this reactant must be introduced into 
the reaction zone in a different manner depending on the type of technique 
selected to carry out the catalytic distillation. In the first type, the 
gaseous reactant can simply be added to the distillation vapour at any 
level, but always before penetration into the reaction zone, generally 
substantially at the inlet to at least one catalytic bed of the reaction 
zone. In the second type, the gaseous reactant must be introduced in a 
manner appropriate to the options selected to impose a circulation 
direction on the liquid and gas in the catalyst bed. 
It thus appears that, for a reaction which is carried out in the presence 
of a solid catalyst between a liquid phase reactant and a gaseous reactant 
which is only slightly soluble in the liquid, such as the hydrogenation of 
unsaturated hydrocarbons mixed with other hydrocarbons, the option which 
consists of carrying out catalytic distillation by avoiding passing the 
distillation vapour over the catalyst and causing the liquid and the 
gaseous reactant to circulate in an ascending co-current in the catalytic 
bed is the most efficient. It appears that the pressure drop across the 
catalytic bed(s) in the first type does not produce an intimate mixture of 
the liquid and gas stream containing hydrogen. In effect, in that type of 
technique, where the reaction and the distillation proceed simultaneously 
in the same physical space, the liquid descends through the catalytic bed 
in rivulets, i.e., in streams of liquid. The gaseous fraction containing 
the vaporised fraction of the feed and the gas stream containing hydrogen 
ascend through the catalytic bed in columns of gas. In this disposition, 
the entropy of the system is greater and the pressure drop across the 
catalytic bed(s) is lower. Thus operating in accordance with the first 
type does not encourage dissolution of hydrogen in the liquid phase 
comprising the unsaturated compound(s). 
The second type, comprising a specific apparatus for distributing the 
liquid fraction to be hydrogenated and the gas stream comprising hydrogen, 
where said liquid fraction and said gas stream traverse the catalytic bed 
in an ascending co-current, can effect the hydrogenation reaction 
essentially in the absence of the gaseous fraction of the feed and under 
conditions under which the pressure drop across the catalytic bed(s) is 
the greatest. Increasing the pressure drop because of the specific 
apparatus of the invention can increase the solubility of the hydrogen in 
the liquid phase and can thus encourage hydrogenation in the liquid 
fraction. 
In addition, because of the recognised toxicity of benzene and olefins, 
which are unsaturated compounds, the general tendency is to reduce the 
amount of these constituents in gasoline. Benzene has cancer-causing 
properties and thus any possibility of it polluting the ambient air has to 
be limited as much as possible, particularly by practically excluding it 
from automobile fuels. In the United States, reformulated gasoline cannot 
contain more than 1% of benzene; in Europe, while the restrictions are not 
yet as severe, it has been recommended that this value be slowly 
approached. Further, olefins have been recognised as being among the most 
reactive hydrocarbons in the cycle of photochemical reactions with oxides 
of nitrogen, which occur in the atmosphere and which lead to the formation 
of ozone. An increase in the ozone concentration in air may be a source of 
respiratory problems. A reduction in the amount of olefins contained in 
gasoline, in particular the lightest olefins which have the greatest 
tendency to volatilise during manipulation of the gasoline, is thus 
desirable. 
The benzene content of a gasoline is largely dependent on that of the 
reformat component of that gasoline. The reformate results from catalytic 
treatment of naphtha for the production of aromatic hydrocarbons which 
principally contains 6 to 9 carbon atoms per molecule and in which the 
very high octane number provides the gasoline with antiknock properties. 
Because of the toxicity described above, it is thus necessary to reduce 
the benzene content of the reformate by as much as possible. A number of 
methods can be envisaged. 
The first method consists of limiting the amount of benzene precursors such 
as cyclohexane and methylcyclopentane in the naphtha constituting the feed 
to a catalytic reforming unit,. This solution substantially reduces the 
benzene content in the effluent from the reforming unit but is not 
sufficient in itself when the content must be reduced to as low as 1%. A 
second method consists of eliminating a light fraction of the reformat 
containing benzene by distillation. This solution produces a loss of the 
order of 15% to 20% of hydrocarbons which could be used in the gasoline. A 
third method consists of extracting the benzene present in the effluent 
from a reforming unit. A number of techniques can, in principle, be 
applied: solvent extraction, extractive distillation, adsorption. None of 
these techniques can be applied on an industrial scale, as none of them 
can selectively extract benzene in an economical manner. A fourth method 
consists of chemically transforming the benzene to convert it to a 
constituent which does not suffer legal limits. Alkylation by ethylene, 
for example, transforms the benzene to mainly ethylbenzene. This 
operation, however, is expensive because of the occurrence of secondary 
reactions which necessitate separation steps which use a great deal of 
energy. 
Benzene of a reformat can also be hydrogenated to cyclohexane. As it is 
impossible to selectively hydrogenate benzene in a mixture of hydrocarbons 
which also contains toluene and xylenes, it is thus necessary first to 
fractionate the mixture to isolate a cut which contains only benzene, 
which can then be hydrogenated. A process has also been described in which 
the benzene hydrogenation catalyst is included in the rectification zone 
of a distillation column which separates benzene from other aromatics 
(Benzene Reduction--Kerry Rock and Gary Gildert, CDTECH--1994--Conference 
on Clean Air Act Implementation and Reformulated Gasoline--October 1994), 
which can save on apparatus. 
Further, isobutene for polymerisation must be more than 99% pure and can 
only contain traces of 1-butene and 2-butenes (several tens of parts per 
million by weight, ppm). If the amount of impurities in the isobutene is 
too high, the polymers obtained are of poorer quality and the 
polymerisation yield is lower. Thus other olefinic hydrocarbons containing 
4 carbon atoms per molecule must be eliminated from a hydrocarbon cut 
containing isobutene. 1-butene and isobutene have very close boiling 
points. Separation by distillation is only possible using drastic 
measures. The other olefinic hydrocarbons containing 4 carbon atoms can be 
separated from the isobutene by distillation. The principal problem in the 
production of high purity isobutene is thus separation of 1-butene from 
isobutene. A number of methods can be used to carry out such a separation. 
The first method consists of extraction using sulphuric acid: isobutene is 
selectively hydrated and then regenerated by treating the aqueous phase. 
If the temperature and concentration are controlled properly, such a 
process can produce isobutene with high purity. However, the yield 
normally does not exceed 90%, extraction is not complete and dimers and 
oligomers are formed which lead to the formation of toxic acid sludge. The 
second method consists of cracking the methyl ether of tertio-butyl 
alcohol (MTBE): isobutene is extracted from the C.sub.4 cut by reacting it 
with methanol to form MTBE. The MTBE is then cracked to methanol and 
isobutene on an acid catalyst. The yield can be at least 96%. The 
isobutene produced is of high purity but the dimethylether which can be 
formed during cracking must be removed. The third possible method is 
dehydration of tertiary butyl alcohol (TBA). In the preceding operation, 
the methanol can be replaced by water, leading to the production of TBA. 
Isobutene is then recovered by dehydration of the TBA. This method is 
practically never used, primarily because TBA is closely linked to the 
propylene oxide market. Using those processes, TBA can be a by-product of 
propylene oxide. 
U.S. Pat. No. 2,403,672 describes a process for the separation of isobutene 
from a mixture of isobutene and 1-butene which comprises introducing the 
mixture into an isomerisation and fractionation zone in which the 
isomerisation catalyst also acts as a packing which carries out the 
distillation function. This solution has the major disadvantage of not 
having good distillation efficiency and thus has a mediocre capacity for 
separating the isobutene from the 1-butene. In this technique, the 
reaction and 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 vapour phase generated by the reboil vapour introduced to the 
bottom of the zone. 
The invention concerns a reactive distillation apparatus comprising a 
distillation zone which comprises a stripping zone and a rectification 
zone associated with a reaction zone, at least a portion of which is 
internal to said distillation zone, and comprising at least one catalytic 
bed in which a feed is transformed in the presence of a catalyst and at 
least one gas stream comprising hydrogen, said apparatus being 
characterized in that each catalytic bed in the internal portion of said 
reaction zone is traversed by an ascending co-current of said gas stream 
and liquid. 
The apparatus of the invention generally comprises: 
at least one means for distributing the major portion of the liquid from 
the bottom to the top through the catalyst; 
at least one means for circulating the major portion of the distillation 
vapour from the bottom to the top through the catalytic bed such that said 
vapour is not in practice in contact with the catalyst; and 
at least one means for distributing the major portion of the gas stream 
from the bottom towards the top through the catalyst. 
The feed supplied to the distillation zone is generally introduced into 
said zone to at least one level in said zone, preferably principally to a 
single level of said zone. 
The distillation zone generally comprises at least one column provided with 
at least one distillation contact means selected from the group formed by 
plates, loose packing and structured packing, as is known to the skilled 
person, and is such that the total overall efficiency is generally at 
least 5 theoretical levels. In known cases in which operation of a single 
column causes problems, it is generally preferable to divide the zone so 
as to use at least two columns which form said zone when placed end to 
end, i.e., the rectification zone, possibly the reaction zone and the 
stripping zone are distributed over the columns. In practice, when the 
reaction zone is at least partially internal to the distillation zone, the 
rectification or stripping zone, preferably the stripping zone, is 
generally in at least one column which is different to the column 
comprising the internal portion of the reaction zone. 
The means for circulating the distributing vapour from the bottom towards 
the top through the catalytic bed passes the reaction zone level where the 
catalytic bed is located, i.e., it is generally located in the catalytic 
bed, but it can also be located at the edge of said catalytic bed. 
The reaction zone generally comprises at least one catalytic bed, 
preferably 2 to 6, and more preferably 2 to 4 catalytic bed(s); when at 
least two catalytic beds are incorporated into the reaction zone, these 
two beds may be separated by at least one distillation contact means. 
The apparatus of the invention is generally such that the flow of liquid to 
be transformed is in a co-current with the gas stream comprising hydrogen 
and such that the distillation vapour does not in practice pass through 
any catalytic bed of the internal portion of the reaction zone (meaning 
that, in practice, said vapour is separated from said liquid), for any 
catalytic bed in the internal portion of the reaction zone. In all cases 
of this second type of technique, each catalytic bed of the portion of the 
reaction zone which is in the distillation zone is generally such that the 
gas stream comprising hydrogen and the liquid stream which will react 
circulates through said bed in a co-current, which is generally ascending, 
even if overall in the catalytic distillation zone the gas stream 
comprising hydrogen and the liquid stream which will react circulates in 
counter-current mode. Such systems generally comprise at least one 
apparatus for distributing liquid which can, for example, be a liquid 
distributor, in each catalytic bed in the internal portion of the reaction 
zone. Nevertheless, provided that the technologies used in the process of 
the invention have been designed for catalytic reactions between liquid 
reactants, without modification they are not suitable for a catalytic 
reaction in which one of the reactants, hydrogen, is in the gaseous state. 
For each catalytic bed of the internal portion of the reaction zone, it is 
thus generally necessary to add an apparatus for introducing a gas stream 
comprising hydrogen, using the techniques described below, for example. 
Thus for each catalytic bed in the internal portion of the reaction zone, 
the internal portion of the reaction zone comprises at least one means for 
distributing liquid, generally located below said catalytic bed, and at 
least one means for introducing a gas stream, generally located below or 
in said catalytic bed, preferably in the latter case close to the liquid 
introduction means. In one technique, the means for introducing a gas 
stream into each catalytic bed is identical to the means for distributing 
liquid in the catalytic bed, i.e., there is a means for introducing gas 
into the liquid upstream of the means for distributing liquid (with 
respect to the direction of circulation of the liquid). In practice and in 
current parlance, this means that gas is injected into the liquid upstream 
of the liquid distribution means. In another technique, the means for 
introducing a gas stream is located substantially at the level of the 
liquid distributing means, the gas and liquid being introduced separately 
into the catalytic bed. In a further technique, the means for introducing 
a gas stream is located below or in the catalytic bed, preferably not far 
from the liquid distribution means. 
Further, in one embodiment of the invention, the apparatus of the invention 
is such that the major portion of said gas stream is hydrogen, the major 
portion of the hydrogen, and preferably almost all thereof, originating 
from external the distillation zone. 
The apparatus of the invention is generally such that, for the portion of 
the reaction zone which is internal to the distillation zone, the feed 
from the reaction zone is drawn off at a draw-off level and represents at 
least a portion, preferably the major portion, of the liquid flowing in 
the distillation zone, preferably flowing in the rectification zone and 
more preferably flowing to an intermediate level of the rectification 
zone, the effluent from the reaction zone being at least in part, 
preferably a major part, re-introduced into the distillation zone 
substantially in the proximity, i.e., generally substantially at the same 
height or just above or just below, usually at the same height or just 
below, i.e., located at a distance corresponding to a height which is in 
the range 0 to 4 theoretical plates from a draw-off, preferably from said 
draw-off, to ensure continuity of distillation. Thus for the portion of 
the reaction zone which is internal to the distillation zone, the liquid 
is drawn off naturally by flow in the portion of the reaction zone which 
is internal to the distillation zone and re-introduction of the liquid to 
the distillation zone also occurs naturally by flow of liquid from the 
portion of the reaction zone which is internal to the distillation zone. 
In general, the apparatus of the invention comprises 1 to 4 draw-off(s) 
which supply the external portion of the reaction zone, when the reaction 
zone is not completely internal to the distillation zone. In general, the 
liquid which will react, either partially or completely, circulates first 
in the external portion of the reaction zone then in the internal portion 
of said zone. Two cases are then possible. In the first case, the external 
portion of the reaction zone is supplied by a single draw-off and then, if 
said portion comprises more than two reactors, these are disposed in 
series or in parallel. In the second case, which is preferred, the 
external portion of the reaction zone is supplied by at least two 
draw-offs. 
In one of the preferred embodiments of the invention, the apparatus of the 
invention is such that the reaction zone is completely internal to the 
distillation zone. 
In a preferred embodiment of the apparatus of the invention, the catalyst 
is disposed in the reaction zone as described in the basic apparatus 
defined in U.S. Pat. No. 5,368,691, and arranged such that each catalytic 
bed is supplied by the gas stream containing hydrogen, regularly 
distributed at its base, using one of the techniques described below, for 
example. Using this technique, if the distillation zone comprises a single 
column and if the reaction zone is completely inside said column, the 
catalyst comprised in each catalytic bed, which is internal to the 
distillation zone, is then in contact with an ascending liquid phase 
generated by the reflux introduced to the top of the distillation zone, 
and with hydrogen which circulates in the same direction as the liquid; 
contact with the vapour phase from distillation is avoided by passing at 
least one specially provided chimney through the distribution zone. 
The invention also concerns a process for the treatment of a feed, a major 
portion of which is constituted by hydrocarbons containing at least 5 and 
preferably 5 to 9 carbon atoms per molecule, and comprising at least one 
unsaturated compound containing at most six carbon atoms per molecule 
including benzene, and treating said feed in a distillation zone 
comprising a stripping zone and a rectification zone, associated with a 
hydrogenation reaction zone which is at least partially internal to said 
distillation zone, in which at least a portion, preferably the major 
portion, of the unsaturated compounds containing at most six carbon atoms 
per molecule, i.e., containing up to and including six carbon atoms per 
molecule and contained in the feed are hydrogenated in the presence of a 
hydrogenation catalyst and at least one gas stream comprising hydrogen, 
preferably as the major portion, to cause an effluent which is depleted in 
unsaturated compounds containing at most six carbon atoms per molecule to 
leave overhead of the distillation zone and an effluent which is depleted 
in unsaturated compounds containing at most six carbon atoms per molecule 
to leave from the bottom of the distillation zone, characterized in that 
each catalytic bed of the internal portion of the hydrogenation zone is 
traversed by an ascending co-current of said gas stream and liquid and the 
catalytic bed is not in practice traversed by the distillation vapour. 
The hydrogenation reaction zone at least partially hydrogenates the benzene 
present in the feed, generally in such a way that the concentration of 
benzene in the overhead effluent is at most equal to a set concentration, 
and said reaction zone hydrogenates at least a portion, preferably the 
major portion, of each unsaturated compound containing at most six carbon 
atoms per molecule (other than benzene) which may be present in the feed. 
The process of the invention preferably includes the use of the apparatus 
of the invention. 
The distillation zone and the characteristics of the gas stream, the 
reaction zone, etc . . . , have been described above with respect to the 
apparatus of the invention. 
In one implementation of the process of the invention, the effluent from 
the bottom of the distillation zone is mixed with the overhead effluent 
from said zone, In this case, after any stabilisation which may be 
necessary, the mixture obtained is used as a fuel either directly or by 
incorporation of fuel fractions. 
To carry out hydrogenation using the process of the invention, the 
theoretical molar ratio of hydrogen which is 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, more so when in addition to the benzene present in the 
feed, each unsaturated compound containing at most six carbon atoms per 
molecule present in the feed must be hydrogenated. If the conditions are 
such that there is an excess of hydrogen, the excess hydrogen can 
advantageously be recovered using one of the techniques described below, 
for example. As an example, the excess hydrogen which leaves the 
distillation zone overhead is recovered then injected upstream of the 
compression steps associated with a catalytic reforming unit, mixed with 
the hydrogen from said unit, said unit preferably operating at low 
pressure (i.e., generally at a pressure of less than 8 bars). This excess 
hydrogen can also be recovered then compressed and used again in the 
reaction zone. 
The major portion, preferably almost all, of the hydrogen used in the 
reaction zone of the invention generally originates from external the 
distillation zone. It can originate from any source which produces 
hydrogen of at least 50% purity by volume, preferably at least 80% purity 
by volume and more preferably at least 90% purity by volume. As an 
example, hydrogen originating from catalytic reforming processes, from PSA 
(pressure swing adsorption), electrochemical generation, steam cracking or 
steam reforming can be used. 
The operating conditions in the hydrogenation zone in the process of the 
invention are linked to the operating conditions used for distillation. 
Distillation is carried out at a pressure which is generally in the range 
2 to 20 bars, preferably in the range 4 to 15 bars, and more preferably in 
the range 4 to 10 bars (1 bar=10.sup.5 Pa), with a reflux ratio in the 
range 1 to 10, preferably in the range 3 to 6. The temperature at the head 
of the zone is generally in the range 40.degree. C. to 180.degree. C. and 
the temperature at the bottom of the zone is generally in the range 
120.degree. C. to 280.degree. C. The hydrogenation reaction is carried out 
under conditions which are most generally intermediate between those 
established overhead and at the bottom of the distillation zone, at a 
temperature which is in the range 100.degree. C. to 200.degree. C., 
preferably in the range 120.degree. C. to 180.degree. C., and at a 
pressure which is in the range 2 to 20 bars, preferably in the range 4 to 
10 bars. The liquid which is hydrogenated is supplied with hydrogen, the 
flow rate of which depends on the concentration of benzene in said liquid 
and, more generally, of the unsaturated compounds containing at most six 
carbon atoms per molecule in the feed from the distillation zone. It is 
generally at least equal to the flow rate which corresponds to the 
stoichiometry of the hydrogenation reactions taking place (hydrogenation 
of benzene and other unsaturated compounds containing at most six carbon 
atoms per molecule comprised in the hydrogenation feed) and at most equal 
to the flow rate which corresponds to 10 times the stoichiometry, 
preferably less than six times the stoichiometry, and more preferably less 
than 3 times the stoichiometry. 
When the hydrogenation zone includes a portion which is external to the 
distillation zone, the catalyst disposed in said external portion can use 
any technique which is known to the skilled person, under operating 
conditions (temperature, pressure . . . ) which are or are not 
independent, preferably independent, of the operating conditions in the 
distillation zone. In the portion of the hydrogenation zone which is 
external to the distillation zone, the operating conditions are generally 
as follows. The pressure required for this hydrogenation step is generally 
in the range 1 to 60 bars absolute, preferably in the range 2 to 50 bars 
and more preferably in the range 5 to 35 bars. The operating temperature 
in the hydrogenation zone is generally in the range 100.degree. C. to 
400.degree. C., preferably in the range 120.degree. C. to 350.degree. C., 
more preferably in the range 140.degree. C. to 320.degree. C. The space 
velocity in said hydrogenation zone, calculated with respect to the 
catalyst, is generally in the range 1 to 50 h.sup.-1, more particularly in 
the range 1 to 30 h.sup.-1 (volume of feed per volume of catalyst per 
hour). The hydrogen flow rate corresponding to the stoichiometry of the 
hydrogenation reactions taking place is in the range 0.5 to 10 times said 
stoichiometry, preferably in the range 1 to 6 times said stoichiometry and 
more preferably in the range 1 to 3 times said stoichiometry. However, the 
temperature and pressure conditions can also be within those established 
at the head and bottom of the distillation zone, without departing from 
the scope of the invention. 
More generally, whatever the position of the hydrogenation zone with 
respect to the distillation zone, the catalyst used in the hydrogenation 
zone of the invention generally comprises at least one metal selected from 
the group formed by nickel and platinum, used as they are or preferably 
deposited on a support. The metal is generally in its reduced form for at 
least 50% by weight of its full quantity. However, any other hydrogenation 
catalyst which is known to the skilled person can also be used. 
When platinum is used, the catalyst can advantageously contain at least one 
halogen in a proportion which is in the range 0.2% to 2% by weight with 
respect to the catalyst. Chlorine or fluorine is preferably used, or a 
combination of the two in a proportion which is in the range 0.2% to 1.5% 
with respect to the total catalyst weight. When a catalyst containing 
platinum is used, a catalyst is generally used in which the average size 
of the platinum crystallites is below 60.times.10.sup.-10 m, preferably 
less than 20.times.10.sup.-10 m and more preferably less than 
10.times.10.sup.-10 m. Further, the overall proportion of the platinum is 
generally in the range 0.1% to 1% with respect to the total catalyst 
weight, preferably in the range 0.1% to 0.6%. 
When nickel is used, the proportion of nickel is in the range 5% to 70% 
with respect to the total weight of the catalyst, more particularly in the 
range 10% to 70% and preferably in the range 15% to 65%. Further, a 
catalyst is generally used in which the average size of the nickel 
crystallites is less than 100.times.10.sup.-10 m, preferably less than 
80.times.10.sup.-10 m, and more preferably less than 60.times.10.sup.-10 
m. 
The support is generally selected from the group formed by alumina, 
silica-aluminas, silica, zeolites, activated charcoal, clays, aluminous 
cements, rare earth oxides and alkaline-earth oxides, used alone or as a 
mixture. Preferably, a support based on alumina or silica is used, with a 
specific surface area which is in the range 30 to 300 m.sup.2 /g, 
preferably in the range 90 to 260 m.sup.2 /g. 
Finally, the invention concerns a process for the treatment of a feed 
comprising, as its major portion, olefinic hydrocarbons containing 4 
carbon atoms per molecule, including isobutene, also 1-butene and 
2-butenes in a ratio which substantially corresponds to the thermodynamic 
equilibrium, in which said feed is treated in a distillation zone 
comprising a stripping zone and a rectification zone associated with a 
hydroisomerisation reaction zone, said reaction zone being at least 
partially internal to said distillation zone and comprising at least one 
catalytic bed, in which hydroisomerisation of at least a portion and 
preferably the major portion of 1-butene is carried out in the presence of 
a hydroisomerisation catalyst and a gas stream comprising hydrogen, 
preferably as its major portion, such that an effluent which is rich in 
isobutene, generally of high purity, leaves the distillation zone overhead 
and an effluent which is depleted in isobutene leaves the bottom, said 
process being characterized in that each catalytic bed in the internal 
portion of the hydroisomerisation zone is traversed by an ascending 
co-current of said gas stream and liquid and is not in practice traversed 
by distillation vapour. The process can be used to produce high purity 
isobutene. 
The process of the invention preferably includes the use of the apparatus 
of the invention. 
The feed supplied to the distillation zone is generally introduced into 
said zone to at least one level of said zone, preferably principally to a 
single level of said zone. It is in a ratio which substantially 
corresponds to the 1-butene/2-butenes thermodynamic equilibrium on 
introduction. In a preferred implementation of the process of the 
invention, the feed is obtained from a cut with a major portion comprised 
of olefinic hydrocarbons containing 4 carbon atoms per molecule, including 
isobutene and 1-butene, by treatment of said cut in a first 
hydroisomerisation zone, which is generally independent of the optional 
portion of the hydroisomerisation reaction zone which is external to the 
associated distillation zone, the major portion of the effluent from said 
first hydroisomerisation zone acting as the feed, which is principal or 
secondary according to the definitions given below in the text, which 
supplies the distillation zone. If the feed includes polyunsaturated 
compounds, usually dienes and/or acetylenes, the compounds are preferably 
transformed to butenes in the first hydroisomerisation zone before 
introduction into the distillation zone. However, any other technique 
which can produce a feed in which the 1-butene and 2-butenes are in a 
ratio which substantially corresponds to the thermodynamic equilibrium 
from a cut with olefin C.sub.4 hydrocarbons as its major portion is also 
within the scope of the invention. 
The first optional hydroisomerisation reaction zone located upstream of the 
distillation-reaction zone effects at least partial selective 
hydrogenation of the polyunsaturated compounds, usually dienes such as 
butadiene, in addition to hydroisomerisation of at least a portion of the 
1-butene to 2-butenes. It generally comprises at least one catalytic 
hydroisomerisation bed comprising a hydroisomerisation catalyst, 
preferably with 1 to 4 catalytic beds; when at least two catalytic beds 
arc incorporated into the reaction zone, these two beds are preferably 
distributed over at least two reactors, in series or in parallel, 
preferably in series. As an example, said first reaction zone comprises a 
single reactor containing at least one catalytic bed, preferably only one. 
In a preferred implementation of the process of the present invention, 
said first reaction zone comprises two reactors which are generally in 
series each comprising at least one catalytic bed, preferably only one. 
When the reaction zone comprises at least two reactors, any recycling of 
at least a portion of the effluent from at least one of the reactors in 
the first reaction zone to the first zone is generally made to the inlet 
of one reactor, preferably to said reactor, preferably before injection of 
the gaseous compound comprising hydrogen. It is also possible to recycle 
around the first zone itself, i.e., generally to the inlet of the first 
reactor to said zone, preferably before injection of the gaseous compound 
comprising hydrogen; as an example, with two reactors, at least a portion 
of the effluent from the second reactor is recycled to the inlet to the 
first reactor. This can advantageously reduce the concentration of 
polyunsaturated compounds in the effluent from the first reaction zone. 
The operating conditions of the first hydroisomerisation zone, when 
present, are generally as follows: the catalyst is identical to the 
catalyst from the hydroisomerisation zone which will be described below. 
The pressure is generally in the range 4 to 40 bar (1 bar=0.1 MPa), 
preferably in the range 6 to 30 bar. The temperature is generally in the 
range 10.degree. C. to 150.degree. C., preferably in the range 20.degree. 
C. to 100.degree. C. The H.sub.2 /hydrocarbons molar ratio is generally 
adjusted so as to obtain practically complete conversion of the 
polyunsaturated compounds such as butadiene and sufficient isomerisation 
of 1-butene to 2-butenes with limited alkane formation. 
The hydroisomerisation reaction zone associated with the distillation zone 
generally comprises at least one catalytic hydroisomerisation bed 
comprising a hydroisomerisation catalyst, preferably 2 to 4, and more 
preferably 2 to 6 catalytic beds; when at least two catalytic beds are 
incorporated into said distillation zone, these two beds are preferably 
separated by at least one distillation contact means. The 
hydroisomerisation reaction zone at least partially hydroisomerises at 
least a portion, preferably the major portion, of the 1-butene present in 
the feed to 2-butenes (cis and trans), generally such that the 1-butene 
concentration in the overhead effluent from the distillation zone is a 
maximum of a certain value. 
The distillation zone used in the process of the invention is identical to 
that described above. 
In a preferred implementation of the process of the invention, in addition 
to supplying the distillation zone with the. principal feed, it is 
supplied with a secondary feed (secondary with respect to the principal 
feed) which may or may not originate from a hydroisomerisation reaction 
zone such as the first, optional, hydroisomerisation reaction zone, and 
may or may not be independent of the supply of principal feed to the 
distillation zone. The secondary feed is generally a C.sub.4 cut 
containing at least isobutene, also 1-butene and 2-butenes in a ratio 
which substantially corresponds to the thermodynamic equilibrium, and 
generally originates from a steam cracking process such as a crude C.sub.4 
cut or the 1-raffinate, or from catalytic cracking; generally and 
preferably, the secondary feed is a C.sub.4 cut which is essentially free 
of polyunsaturated compounds and the 1-butene content is lower than the 
1-butene content in the principal feed. If the amount of unsaturated 
compounds in the secondary feed is high, the feed is preferably treated in 
a selective hydrogenation zone before its entry into the distillation 
zone. 
When the principal feed is introduced at a single introduction level, the 
secondary feed is generally introduced into the distillation zone to at 
least one introduction level, preferably to a single introduction level, 
said introduction level depending on the composition of the secondary 
feed. Thus in a first example, the secondary feed can be very rich in 
isobutene and contain less than 1.5 times the 1-butene contained in the 
principal feed, in which case the secondary feed is preferably introduced 
at a single level generally located above the level at which the principal 
feed is introduced. In a second example, the secondary feed is practically 
free of 1-butene, in which case the secondary feed is preferably 
introduced to a single level generally located below the level at which 
the principal feed is introduced. It is also possible to mix the principal 
feed before its entry into the distillation zone with the secondary feed. 
The hydroisomerisation reaction zone associated with the distillation zone 
generally comprises at least one catalytic hydroisomerisation bed, 
preferably 2 to 6 and more preferably 2 to 4 catalytic beds; when at least 
two catalytic beds are incorporated into the distillation zone, these two 
beds are optionally separated by at least one distillation contact means. 
The hydroisomerisation reaction zone at least partially hydroisomerises at 
least a portion, preferably the major portion, of the 1-butene present in 
the feed to 2-butenes (cis and trans), generally such that the 1-butene 
concentration in the overhead effluent from the distillation zone is a 
maximum of a pre-set value. 
The process of the invention is generally such that the flow of the liquid 
to be hydroisomerised is in a co-current with the flow of the gas stream 
comprising hydrogen in each catalytic bed of the internal portion of the 
hydroisomerisation zone, and such that the distillation vapour does not in 
practice traverse any catalytic bed in the internal portion of the 
reaction zone (meaning that in practice, the vapour is separated from the 
liquid to be hydroisomerised). Each catalytic bed in the portion of the 
reaction zone which is internal to the distillation zone is generally such 
that the gas stream comprising hydrogen and the liquid stream which is to 
be reacted circulate in a co-current, generally ascending, across the bed, 
even if overall in the catalytic distillation zone, the gas stream 
comprising hydrogen and the liquid stream to be reacted circulate in a 
counter-current. Such systems generally comprise at least one liquid 
distribution apparatus which can, for example, be a liquid distributor, 
for each catalytic bed in the internal portion of the reaction zone. The 
distribution apparatus for the gas streams and for liquid distribution 
have been described above. 
The process of the invention is generally such that in all parts of the 
hydroisomerisation reaction zone, whether internal or, optionally, 
external, the feed is drawn off at the height of a draw-off and represents 
at least a portion, preferably the major portion, of the liquid (reflux) 
flowing in the distillation zone, preferably flowing in the rectification 
zone and more preferably flowing in an intermediate level of the 
rectification zone, the effluent from the hydroisomerisation reaction zone 
being at least in part, preferably the major part, re-introduced into the 
distillation zone, so as to ensure continuity of distillation. For the 
optional portion of the reaction zone which is external of the 
distillation zone, re-introduction of the effluent from the distillation 
zone is effected substantially in the proximity, i.e., generally 
substantially at the same height or just above or just below, generally at 
the same height or just above, i.e., located at a distance corresponding 
to a height which is in the range 0 to 4 theoretical plates from a 
draw-off, preferably from said draw-off, to ensure continuity of 
distillation. For the portion of the reaction zone which is internal to 
the distillation zone, liquid (reflux) draw-off occurs naturally by flow 
in the portion of the reaction zone which is internal to the distillation 
zone, and re-introduction of the effluent to the distillation zone also 
occurs naturally by flow of liquid from the internal reaction zone to the 
distillation zone. 
In general, when the hydroisomerisation zone is not completely internal to 
the distillation zone, the process of the invention comprises 1 to 6, 
preferably 2 to 4 draw-offs, which supply the external portion of the 
hydroisomerisation zone. In such a case, the liquid to be hydroisomerised, 
either partially or completely, circulates first in the external portion 
of the hydroisomerisation zone then in the internal portion of that zone. 
Two cases are then possible. In the first case, the external portion of 
the reaction zone is supplied by a single draw-off and thus, if said 
portion comprises more than two reactors, these are disposed in series or 
in parallel. In the second case, which is preferred, the external portion 
of the hydroisomerisation zone is supplied by at least two draw-offs. A 
portion of the external portion of the hydroisomerisation zone which is 
supplied by a given draw-off, if the external portion comprises at least 
two draw-offs, generally comprises at least one reactor, preferably a 
single reactor. If said portion of the external portion comprises at least 
two reactors, each reactor which is external to the distillation zone is 
generally supplied by a single draw-off, preferably associated with a 
single re-introduction level, said draw-off being distinct from the 
draw-off which supplies the other reactor(s). 
In preferred implementation of the invention, the process of the invention 
is such that the hydrogenation zone is completely internal to the 
distillation zone. 
The major portion of the hydrogen used for hydroisomerisation of 1-butene, 
preferably almost all thereof, originates from external the distillation 
zone. It can originate from any source which produces hydrogen in at least 
50% purity by volume, preferably at least 80% purity by volume and more 
preferably at least 90% purity by volume. As an example, hydrogen 
originating from catalytic reforming processes, from PSA (pressure swing 
adsorption), electrochemical generation, steam cracking or steam reforming 
can be used. 
The operating conditions in the portion of the hydroisomerisation zone 
internal to the distillation zone are linked to the operating conditions 
used for distillation. Distillation is generally carried out in a manner 
which minimises the quantity of isobutene in the bottom product to 
maximise the yield of isobutene from the process and minimise the quantity 
of 2-butenes and 1-butene in the overhead product to produce high purity 
isobutene overhead. It is carried out at a pressure which is generally in 
the range 2 to 30 bars, preferably in the range 4 to 15 bars, and more 
preferably in the range 4 to 10 bars, with a reflux ratio in the range 1 
to 30, preferably in the range 5 to 20. The temperature at the head of the 
zone is generally in the range 0.degree. C. to 200.degree. C. and the 
temperature at the bottom of the zone is generally in the range 5.degree. 
C. to 250.degree. C. The hydroisomerisation reaction is carried out under 
conditions which are most generally intermediate between those established 
overhead and at the bottom of the distillation zone, at a temperature 
which is in the range 20.degree. C. to 150.degree. C., preferably in the 
range 40.degree. C. to 80.degree. C., and at a pressure which is in the 
range 2 to 30 bars, preferably in the range 4 to 15 bars, and more 
preferably in the range 4 to 10 bars. The liquid which is hydroisomerised 
is supplied with a gas stream comprising hydrogen, preferably as the major 
portion. 
When the hydroisomerisation zone includes a portion which is external to 
the distillation zone, the catalyst disposed in said external portion can 
use any technique which is known to the skilled person, under operating 
conditions (temperature, pressure . . . ) which are generally independent 
of the operating conditions in the distillation zone. In the optional 
portion of the hydroisomerisation zone which is external to the 
distillation zone, the operating conditions are generally as follows. The 
pressure required for this hydroisomerisation step is generally in the 
range of about 1 to 40 bars absolute, preferably in the range of about 2 
to 30 bars and more preferably in the range of about 4 to 25 bars. The 
operating temperature in the hydroisomerisation zone is generally in the 
range of about 20.degree. C. to 150.degree. C., preferably in the range of 
about 40.degree. C. to 100.degree. C., more preferably in the range of 
about 40.degree. C. to 80.degree. C. The space velocity in said 
hydroisomerisation zone, calculated with respect to the catalyst, is 
generally in the range of about 1 to 100 h.sup.-1, more particularly in 
the range of about 4 to 50 h.sup.-1 (volume of feed per volume of catalyst 
per hour). The corresponding hydrogen flow rate is such that the H.sub.2 
/hydrocarbons molar ratio on entering the hydroisomerisation zone is 
preferably at least about 10.sup.-5. This ratio is usually about 10.sup.-5 
to about 3 and often about 10.sup.-3 to about 1. However, the temperature 
and pressure conditions can also be in the range which is established at 
the head and bottom of the distillation zone, without departing from the 
scope of the invention. 
In order to carry out hydroisomerisation using the process of the 
invention, the theoretical molar ratio of hydrogen necessary for the 
desired conversion of 1-butene in the reaction zone associated with the 
distillation zone is such that the H.sub.2 /hydrocarbons molar ratio on 
entering said zone is at least 10.sup.-5. This molar ratio can be 
optimised so that all the hydrogen is consumed in the hydroisomerisation 
reaction to avoid the need for a hydrogen recovery apparatus at the outlet 
to the reaction zone, and so that the secondary hydrogenation reactions 
can be minimised to maximise the isobutene yield from the process and 
finally, such that there is sufficient hydrogen along the whole length of 
the reaction zone so that the hydroisomerisation reaction of 1-butene to 
2-butenes can take place. However, if these conditions are such that there 
is an excess of hydrogen, the excess hydrogen can advantageously be 
recovered using one of the techniques described above, for example. As an 
example, the excess hydrogen leaving the distillation zone overhead is 
recovered, then injected upstream of the compression stages associated 
with a catalytic reforming unit, mixed with the hydrogen from said unit, 
said unit preferably operating at low pressure (i.e., generally a pressure 
of less than 8 bar). This excess hydrogen can also be recovered then 
compressed and used again in the reaction zone. 
When a portion of the hydroisomerisation zone associated with the 
distillation zone is external to the distillation zone, the process of the 
invention can isomerise a large portion of the 1-butene to 2-butenes 
external the distillation zone, optionally under different temperature 
and/or pressure conditions to those used in the column. The inlet 
temperature (and, respectively, the outlet temperature) at the draw-off 
which supplies a catalytic bed of the portion of the hydroisomerisation 
zone which is external to the column is preferably substantially similar, 
i.e., the difference is substantially less than 10.degree. C. with respect 
to the temperature at the height of the draw-off (with respect to the 
re-introduction level). Similarly, the hydroisomerisation reaction can 
advantageously be carried out in the portion of the reaction zone which is 
external the column at a pressure which is higher than that used internal 
to the distillation zone. This pressure increase can thus increase 
dissolution of the gas stream containing hydrogen in the liquid phase 
containing 1-butene to be isomerised. 
In such a case, the process of the invention comprises the use of a 
technique known as "pumparound" which consists of passing a portion, 
preferably the major portion, of the liquid (reflux) outside the 
distillation zone in an amount which is preferably a factor of more than 
1, i.e., the flow rate of a catalytic bed in the external portion of the 
hydroisomerisation zone associated with the distillation zone, said bed 
being supplied at a draw-off with a portion of the liquid effluent 
(reflux) flowing on the distillation plate associated with said draw-off 
(i.e., from which said portion of liquid effluent is drawn off) and with 
at least a portion of the liquid corresponding to recycling the effluent 
from said bed just above or just below or substantially at the same level 
as said draw-off, is more than once the flow rate of the liquid flowing on 
said plate, for example 1.5 times. 
More generally, the catalyst used in the hydroisomerisation zone of the 
process of the invention generally comprises at least one metal selected 
from the group formed by noble metals from group VIII of the periodic 
classification of the elements and nickel, i.e., selected from the group 
formed by ruthenium, rhodium, palladium, osmium, iridium and platinum, 
preferably palladium, or nickel, used as it is or, preferably, deposited 
on a support. At least 50% by weight of the metal is generally in its 
reduced form. The noble metal content of the catalyst is generally about 
0.01% to about 2% by weight. When nickel is used, the proportion of nickel 
with respect to the total catalyst weight is in the range 5% to 70%, 
preferably in the range 10% to 70%, and in general, a catalyst is used in 
which the average size of the nickel crystallites is less than 10 nm, 
preferably less than 8 nm, more preferably less than 6 nm. However, any 
other hydroisomerisation catalyst which is known to the skilled person can 
also be selected. Before use, the catalyst is normally treated with a 
sulphur compound then by hydrogen. The catalyst is generally sulphurated 
in situ or ex situ such that sulphur is chemisorbed onto at least a 
portion of the metal. The chemisorbed sulphur encourages the isomerisation 
of 1-butene to 2-butenes over the isobutene hydrogenation reaction and 
thus maximises the isobutene yield of the process. 
The hydroisomerisation catalyst support is generally selected from the 
group formed by alumina, silica-aluminas, silica, zeolites, activated 
charcoal, clays, aluminous cements, rare earth oxides and alkaline-earth 
oxides, used alone or as a mixture. A support based on alumina or silica 
is preferably used, with a specific surface area which is in the range 10 
to 300 m.sup.2 /g, preferably in the range 30 to 70 m.sup.2 /g.

By way of non limiting example, according to the present invention, 
commercial catalysts can be used, such as those sold by the company 
Catalysts and Chemicals under reference C-31 or those sold by the Girdler 
Corporation under reference G-55 or, preferably, those sold by Procatalyse 
under references LD-265, LD-265S, LD-267 and LD-267R. 
EXAMPLES 
Examples 1 and 2 show the operation of a zone for hydrogenation of 
unsaturated compounds comprising at most six carbon atoms per molecule, 
including benzene, in accordance with the invention (Example 2) and the 
operation of a hydrogenation zone which was not in accordance with the 
invention (Example 1), with a catalyst which was loosely packed on 
distillation plates traversed by the liquid which circulated downwardly 
and by a vapour which circulated upwardly in the distillation zone. 
Example 1 (comparative) 
A metal distillation column with a diameter of 50 mm was used, rendered 
adiabatic with heating envelopes in which the temperatures were regulated 
to produce the temperature gradient established in the column. Over a 
height of 4.5 m, the column comprised, from head to foot: a rectification 
zone composed of 11 sieve plates with downcomers, a catalytic 
hydrogenating distillation zone and a stripping zone composed of 63 
perforated plates. The catalytic hydrogenating distillation zone was 
constituted by three reactive plates which in this instance were sieve 
plates with downcomers, with the weirs raised by 3.5 cm and in which the 
volume between the top of the weir and the plate could be packed with 
catalyst. A metal screen placed at the top of the overflow acted as a 
filter to prevent catalyst particles from being evacuated with the liquid 
leaving the plate. 
Each of the three cells was packed with 36 g of nickel catalyst sold by 
PROCATALYSE with the reference LD 746. 260 g/h of a reformat comprising 
essentially hydrocarbons containing at least 5 carbon atoms per molecule 
was introduced to the 37.sup.th plate in the column counting from the 
bottom. The reformat composition is shown in Table 1. 18 Nl/h of hydrogen 
was also introduced to the base of each cell. The column was started by 
establishing a reflux ratio of 5, regulating the bottom temperature to 
176.degree. C. and the pressure at 7 bars. 
At steady state, 138 g/h and 113 g/h respectively of residue and a liquid 
distillate were recovered. The compositions are given in Table 1. A small 
portion of the distillate, constituted by the lightest hydrocarbons, was 
evacuated from the column with excess hydrogen and was not accounted for. 
Analysis of the effluents led to the deduction that the degree of 
hydrogenation of the olefins and benzene in the feed were respectively 
100% and 55% while toluene was unaffected. 
Example 2: (in accordance with the invention) 
The apparatus of Example 1 was used but the catalytic distillation zone was 
of a different design. The catalytic hydrogenating distillation zone in 
this instance was constituted by three catalytic distillation doublets, 
each doublet being itself constituted by a catalytic cell over which three 
perforated plates were mounted. Construction details of a catalytic cell 
and its disposition in the column are shown schematically in the figure. 
Catalytic cell 1 consists of a cylindrical container with a flat base with 
an external diameter which is 2 mm smaller than the internal diameter of 
the column. At its lower portion above the base, a screen 2 is provided 
which acts both as a support for the catalyst and as a distributor for the 
hydrogen. A catalyst retaining screen 3 is provided at the top; the height 
thereof can be varied. Catalyst 4 fills the entire volume between these 
two screens. The catalytic cell receives the liquid from upper 
distillation plate 5 via downcomer 6. After passing through the cell in 
the upward direction, the liquid is evacuated by overflowing via downcomer 
7 and flows onto lower distillation plate 8. The vapour from lower plate 8 
passes into the central chimney 9, which is solid with the cell, by 
penetrating via orifices 10 (only one shown in the figure) and leaving 
below upper plate 5 via orifices 11 (only one shown in the figure). 
Hydrogen is introduced to the base of the catalytic cell via conduit 12 
then via orifices 13 (six in total) which are distributed around the 
periphery of the cell, in the immediate vicinity of the base. Seals 14 
prevent any hydrogen from escaping prior to its arrival at the catalytic 
bed. 
Each of the three cells was packed with 36 g of nickel catalyst sold by 
PROCATALYSE with the reference LD 746. 260 g/h of the same feed as that 
used in Example 1 was introduced to the 37.sup.th plate in the column 
counting from the bottom. The reformat composition is shown in the second 
column of the Table. 6 Nl/h of hydrogen was introduced to the base of each 
cell. The column was started by establishing a reflux ratio of 5, 
regulating the bottom temperature to 176.degree. C. and the pressure at 7 
bars. 
At steady state, 143 g/h and 106 g/h respectively of residue and a liquid 
distillate were recovered. The compositions are given in Table 1. A small 
portion of the distillate, constituted by the lightest hydrocarbons, was 
evacuated from the column with excess hydrogen and was not accounted for. 
Analysis of the effluents led to the deduction that the degree of 
hydrogenation of the olefins and benzene in the feed were respectively 
100% and 87% while toluene was unaffected. 
TABLE 1 
______________________________________ 
compositions of feed and effluents in the catalytic column 
composition, in % by weight 
example 1 example 2 
liquid liquid 
feed residue distillate 
residue 
distillate 
______________________________________ 
C5 and lighter 
7.65 10.22 7.36 
of which: olefins 
0.11 0 0 
C6 44.83 9.55 89.78 12.4 92.59 
of which: olefins 
0.13 0 0 
: benzene 
6.07 0.63 5.45 0.07 1.84 
: cyclohexane 
1.1 8.34 0.34 12.16 0.73 
C7: 42.55 80.72 78.27 0.05 
of which: toluene 
4.78 9.1 8.87 
C8 and heavier 
4.97 9.73 9.33 
olefin conversion 
100% 100% 
benzene conversion 
55% 87% 
hydrogen conversion 
15% 70% 
______________________________________ 
It can be seen that the process of the present invention produces better 
conversion of benzene and better hydrogen conversion. 
Examples 3 and 4 illustrate the case of a process of the invention for 
treatment of a feed the major portion of which is comprised by olefinic 
hydrocarbons containing 4 carbon atoms per molecule, including isobutene, 
1-butene and 2-butenes in a ratio which substantially corresponds to the 
thermodynamic equilibrium. 
Example 3 
Hydroisomerisation operations were carried out successively and 
discontinuously on a C.sub.4 distillation cut. The feed was 
hydroisomerised a first time. The effluent from the first test was 
distilled: the distillation head, representing an intermediate extraction, 
was hydroisomerised. The hydroisomerisation effluent, representing what 
would be re-injected into the column, was distilled. The head from the 
second distillation was hydroisomerised and the effluent from this third 
hydroisomerisation step was distilled. 
The hydroisomerisation operations were carried out in a pilot unit provided 
with an adiabatic reactor. The reactor was filled with 1.5 l of catalyst 
LD-265 from PROCATALYSE. The catalyst was sulphurated and activated in 
situ using the procedure recommended by the supplier of the catalyst. 
Distillation operations were carried out in an adiabatic column with an 
internal diameter of 163 mm and a height of 10 m. the column was 
constituted by 4 beds which were 1.78 m high above the feed injection 
point, filled with a packing sold by SUIZER under the trade name M550Y and 
2 beds 1 m high below the feed injection point, filled with Pall rings. 
First Hydroisomerisation 
The average operating conditions during the test were as follows: 
Reactor temperature: 80.degree. C. 
Reactor pressure: 20 bar 
Residence time: 0.25 h 
H.sub.2 /feed molar ratio: 3 
Table 2 below shows the compositions of the feed and effluent in the 
hydroisomerisation reactor operating under the conditions described above. 
TABLE 2 
______________________________________ 
Feed (weight %) 
Effluent (weight %) 
______________________________________ 
&lt;C.sub.4 0.25 0.23 
iC.sub.4 2.98 3.10 
iC.sub.4.sup..dbd. 
44.90 44.42 
C.sub.4.sup..dbd. 1 
26.95 4.26 
C.sub.4.sup..dbd..dbd. 1,3 
0.13 0.00 
nC.sub.4 11.72 14.41 
C.sub.4.sup..dbd. 2trans 
8.73 21.37 
Neo C.sub.5 0.24 0.23 
Me cyclo C.sub.3 
0.06 0.06 
C.sub.4.sup..dbd. 2cis 
4.03 11.92 
&gt;C.sub.4 0.01 0.00 
______________________________________ 
The legend for the table and the following tables is as follows: 
&lt;C.sub.4 : compounds with less than 4 (4 excluded) carbon atoms per 
molecule (or C.sub.3.sup.-) 
iC.sub.4 : isobutene 
iC.sub.4.sup.= : isobutene 
C.sub.4.sup.= 1: 1-butene 
C.sub.4.sup.== 1,3: 1,3-butadiene 
nC.sub.4 : normal butane 
C.sub.4.sup.= 2trans: trans 2-butene 
Neo C.sub.5 : neopentane (or dimethylpropane) 
Me cyclo C.sub.3 : methyl cyclopropane 
C.sub.4.sup.= 2cis: cis 2-butene 
&gt;C.sub.4 : compounds containing more than 4 (4 excluded) carbon atoms per 
molecule (or C.sub.5.sup.+) 
First Distillation 
Distillation of the effluent from the above test was carried out under the 
following operating conditions: 
Column pressure: 4 bar 
Reflux ratio (R/D): 20 
Feed temperature: 33.degree. C. 
Reflux temperature: 32.degree. C. 
Column head temperature: 57.degree. C. 
Column bottom temperature: 63.degree. C. 
Table 3 below shows the compositions of the feed and the overhead effluent 
from the distillation column operating under the conditions described 
above. 
TABLE 3 
______________________________________ 
Feed (weight %) 
Head (weight %) 
______________________________________ 
&lt;C.sub.4 0.23 0.44 
iC.sub.4 3.10 6.71 
iC.sub.4.sup..dbd. 
44.42 83.35 
C.sub.4.sup..dbd. 1 
4.26 7.39 
C.sub.4.sup..dbd..dbd. 1,3 
0.00 0.00 
nC.sub.4 14.41 1.62 
C.sub.4.sup..dbd. 2trans 
21.37 0.44 
Neo C.sub.5 0.23 -- 
Me cyclo C.sub.3 
0.06 -- 
C.sub.4.sup..dbd. 2cis 
11.92 0.05 
&gt;C.sub.4 -- -- 
______________________________________ 
Second Hydroisomerisation 
The average operating conditions during the test were as follows: 
Reactor temperature: 65.degree. C. 
Reactor pressure: 20 bar 
Residence time: 0.25 h 
H.sub.2 /feed molar ratio: 0.6 
Table 4 below shows the compositions of the feed and effluent in the 
hydroisomerisation reactor operating under the conditions described above. 
TABLE 4 
______________________________________ 
Feed (weight %) 
Effluent (weight %) 
______________________________________ 
&lt;C.sub.4 0.44 0.39 
iC.sub.4 6.71 6.91 
iC.sub.4.sup..dbd. 
83.35 82.94 
C.sub.4.sup..dbd. 1 
7.39 0.81 
C.sub.4.sup..dbd..dbd. 1,3 
-- -- 
nC.sub.4 1.62 2.09 
C.sub.4.sup..dbd. 2trans 
0.44 4.44 
Neo C.sub.5 -- -- 
Me cyclo C.sub.3 
-- -- 
C.sub.4.sup..dbd. 2cis 
0.05 2.42 
&gt;C.sub.4 -- -- 
______________________________________ 
Second Distillation 
Distillation of the effluent from the above test was carried out under the 
following operating conditions: 
Column pressure: 4 bar 
Reflux ratio (R/D): 13.5 
Feed temperature: 36.degree. C. 
Reflux temperature: 41.degree. C. 
Column head temperature: 51.degree. C. 
Column bottom temperature: 55.degree. C. 
Table 5 below shows the compositions of the feed and the overhead effluent 
from the distillation column operating under the conditions described 
above. 
TABLE 5 
______________________________________ 
Feed (weight %) 
Head (weight %) 
______________________________________ 
&lt;C.sub.4 0.39 0.65 
iC.sub.4 6.91 13.71 
iC.sub.4.sup..dbd. 
82.94 84.82 
C.sub.4.sup..dbd. 1 
0.81 0.51 
C.sub.4.sup..dbd..dbd. 1,3 
-- -- 
nC.sub.4 2.09 0.14 
C.sub.4.sup..dbd. 2trans 
4.44 0.12 
Neo C.sub.5 -- -- 
Me cyclo C.sub.3 
-- -- 
C.sub.4.sup..dbd. 2cis 
2.42 0.05 
&gt;C.sub.4 -- -- 
______________________________________ 
Third Hydroisomerisation 
The average operating conditions during the test were as follows: 
Reactor temperature: 60.degree. C. 
Reactor pressure: 20 bar 
residence time: 0.25 to 0.1 h 
H.sub.2 /feed molar ratio: 1 
Table 6 below shows the compositions of the feed and effluent in the 
hydroisomerisation reactor operating under the conditions described above. 
TABLE 6 
______________________________________ 
Feed (weight %) 
Effluent (weight %) 
______________________________________ 
&lt;C.sub.4 0.65 0.57 
iC.sub.4 13.71 14.55 
iC.sub.4.sup..dbd. 
84.82 84.07 
C.sub.4.sup..dbd. 1 
0.51 0.03 
C.sub.4.sup..dbd..dbd. 1,3 
-- -- 
nC.sub.4 0.14 0.22 
C.sub.4.sup..dbd. 2trans 
0.12 0.38 
Neo C.sub.5 -- -- 
Me cyclo C.sub.3 
-- -- 
C.sub.4.sup..dbd. 2cis 
0.05 0.18 
&gt;C.sub.4 -- -- 
______________________________________ 
Third Distillation 
Distillation of the effluent from the above test was carried out under the 
following operating conditions: 
Column pressure: 4 bar 
Reflux ratio (R/D): 13.5 
Feed temperature: 36.degree. C. 
Reflux temperature: 41.degree. C. 
Column head temperature: 53.degree. C. 
Column bottom temperature: 55.degree. C. 
Table 7 below shows the compositions of the feed and the overhead effluent 
from the distillation column operating under the conditions described 
above. 
TABLE 7 
______________________________________ 
Feed (weight %) 
Head (weight %) 
______________________________________ 
&lt;C.sub.4 0.57 0.57 
iC.sub.4 14.55 14.66 
iC.sub.4.sup..dbd. 
84.07 84.69 
C.sub.4.sup..dbd. 1 
0.03 0.03 
C.sub.4.sup..dbd..dbd. 1,3 
-- -- 
nC.sub.4 0.22 0.01 
C.sub.4.sup..dbd. 2trans 
0.38 0.04 
Neo C.sub.5 -- -- 
Me cyclo C.sub.3 
-- -- 
C.sub.4.sup..dbd. 2cis 
0.18 -- 
&gt;C.sub.4 -- -- 
______________________________________ 
These successive and discontinuous hydroisomerisation and distillation 
operations represent the separation of 1-butene from isobutene which is 
carried out continuously in the process of the invention 
Example 4 
Pilot hydroisomerisation tests were carried out using a 1-raffinate using 
the hydroisomerisation catalyst LD267R from PROCATALYSE which packed each 
of the catalytic beds. The results of these tests are shown in Table 8 
below: they allowed computation parameters to be determined which allowed 
the process of the invention to be simulated using suitable software. The 
software used for this simulation is sold by SIMCI under the trade name 
Pro2. 
TABLE 8 
__________________________________________________________________________ 
pilot test results 
__________________________________________________________________________ 
T.degree. C. 
40 80 90 50 50 50 50 50 50 50 
HSV h.sup.-1 
30 30 30 30 30 30 30 30 20 40 
P bar 10 10 10 6.5 
10 15 10 10 10 10 
H.sub.2 /HC m/m 
0.17 
0.17 
0.17 
0.17 
0.17 
0.17 
0.1 
0.19 
0.17 
0.17 
feed 
effl 
effl 
effl 
effl 
effl 
effl 
effl 
effl 
effl 
effl 
&lt;C4 0.14 
0.11 
0.12 
0.11 
0.10 
0.10 
0.10 
0.10 
0.11 
0.10 
0.09 
iC4 5.69 
5.75 
5.75 
5.73 
5.71 
5.76 
5.75 
5.72 
5.76 
5.75 
5.74 
iC4 = 78.67 
78.71 
78.72 
78.73 
78.78 
78.72 
78.73 
78.74 
78.72 
78.71 
78.74 
1-iC4 = 
3.66 
1.30 
0.91 
0.75 
1.15 
1.01 
1.18 
1.13 
1.00 
0.8 
1.32 
n-C4 7.16 
7.19 
7.17 
7.14 
7.14 
7.18 
7.19 
7.16 
7.20 
7.19 
7.18 
tr2-C4 = 
4.36 
5.40 
5.48 
5.40 
5.46 
5.49 
5.41 
5.46 
5.48 
5.59 
5.37 
cs2-C4 = 
0.32 
1.54 
1.85 
2.14 
1.66 
1.74 
1.64 
1.69 
1.75 
1.86 
1.56 
__________________________________________________________________________ 
where effl = effluent. 
The catalytic hydroisomerising distillation zone comprised 2 or 3 catalytic 
distillation doublets, each of the doublets being of the type shown in the 
figure, each doublet being itself constituted by a catalytic cell over 
which three perforated plates were mounted. 
Two examples which were simulated using the calculation were carried out. 
They are described below. 
Example 4A 
The configuration of the unit, comprising three catalytic 
hydroisomerisation beds located inside the column, termed reactive plates, 
was as follows: 
column with 130 theoretical plates, numbered from top to bottom; 
supply to plate no. 90; 
the reactive plates were plates 10, 25 and 39. Each contained 7.5 m.sup.3 
of catalyst. 
Operating conditions: 
Flow rate of liquid supply to column: 292.9 kmole/h; 
Reflux ratio: 12; 
Column head pressure: 6.2 bars absolute; 
Column bottom pressure: 7 bars absolute; 
Temperature of supply to column: 59.degree. C.; 
Column head temperature: 52.degree. C.; 
Column bottom temperature: 64.5.degree. C.; 
Temperature of reactive plate no. 10: 53.degree. C.; 
Pressure of reactive plate no. 10: 6.6 bars absolute; 
Flow rate of liquid traversing reactive plate no. 10: 1660 1kmoluh; 
Temperature of reactive plate no. 25: 54.degree. C.; 
Pressure of reactive plate no. 25: 6.6 bars absolute; 
Flow rate of liquid traversing reactive plate no. 25: 1660 kmole/h; 
Temperature of reactive plate no. 39: 54.degree. C.; 
Pressure of reactive plate no. 39: 6.7 bars absolute; 
Flow rate of liquid traversing reactive plate no. 39: 1660 kmole/h. 
With this configuration and under those operating conditions, the 
simulation produced the following results: 
______________________________________ 
Column supply 
Column head 
Column bottom 
(kmole/h) (kmole/h) (kmole/h) 
______________________________________ 
&lt;C4 1.12 1.12 0.00 
iC4 4.46 5.58 0.00 
iC4 = 110.08 108.07 0.89 
C4 = 1 7.53 0.02 0.17 
nC4 55.27 0.13 55.23 
C4 = 2tr 79.76 0.03 84.56 
C4 = 2cis 
33.49 0.00 35.91 
H.sub.2 1.21 0.00 0.00 
Total 292.92 114.95 176.76 
______________________________________ 
Yield of isobutene at column head: 98.2% 
1-butene/isobutene molar ratio at column head: 1.85.times.10.sup.-4. 
Example 4B 
The configuration of the unit, comprising two catalytic hydroisomerisation 
beds located inside the column, termed reactive plates, was as follows: 
column with 130 theoretical plates, numbered from top to bottom; 
supply to plate no. 90; 
the reactive plates were plates 10 and 39. Each contained 7.5 m.sup.3 of 
catalyst. 
Operating conditions: 
Flow rate of liquid supply to column: 292.9 kmole/h; 
Reflux ratio: 12; 
Column head pressure: 6.2 bars absolute; 
Column bottom pressure: 7 bars absolute; 
Temperature of supply to column: 59.degree. C.; 
Column head temperature: 52.degree. C.; 
Column bottom temperature: 64.5.degree. C.; 
Temperature of reactive plate no. 10: 53.degree. C.; 
Pressure of reactive plate no. 10: 6.6 bars absolute; 
Flow rate of liquid traversing reactive plate no. 10: 1660 kmol/h; 
Temperature of reactive plate no. 39: 54.degree. C.; 
Pressure of reactive plate no. 39: 6.7 bars absolute; 
Flow rate of liquid traversing reactive plate no. 39: 1660 kmole/h. 
With this configuration and under those operating conditions, the 
simulation produced the following results: 
______________________________________ 
Column supply 
Column head 
Column bottom 
(kmole/h) (kmole/h) (kmole/h) 
______________________________________ 
&lt;C4 1.12 1.12 0.00 
iC4 4.46 5.17 0.00 
iC4 = 110.08 108.48 0.89 
C4 = 1 7.53 0.09 0.17 
nC4 55.27 0.13 55.20 
C4 = 2tr 79.76 0.06 84.50 
C4 = 2cis 
33.49 0.00 35.90 
H.sub.2 1.21 0.00 0.00 
Total 292.92 115.49 176.66 
______________________________________ 
Yield of isobutene at column head: 98.6% 
1-butene/isobutene molar ratio at column head: 8.30.times.10.sup.-4. 
The preceding examples can be repeated with similar success by substituting 
the generically or specifically described reactants and/or operating 
conditions of this invention for those used in the preceding examples. 
The entire disclosures of all applications, patents and publications, cited 
above and below, and of corresponding application French No. 95/15.530, 
filed Dec. 27, 1995, are hereby incorporated by reference. 
From the foregoing description, one skilled in the art can easily ascertain 
the essential characteristics of this invention, and without departing 
from the spirit and scope thereof, can make various changes and 
modifications of the invention to adapt it to various usages and 
conditions.