PROCESS FOR THE SELECTIVE OXIDATI ON OF CARBON MONOXIDE
Field of the invention
5
The invention relates generally to the field of the
production of hydrocarbon derivatives from hydrocarbons
in the gas phase in the presence of oxygen or of an
oxygen-comprising gas. More specifically, the invention
10 relates to a process for the selective oxidation of
carbon monoxide to give carbon dioxide present in a gas
mixture comprising at least one hydrocarbon or one
hydrocarbon derivative, and to its incorporation in a
process for the production of hydrocarbon derivatives.
15
Prior art and technical problem
Numerous hydrocarbon derivatives are produced
industrially by partial oxidation of an appropriate
20 hydrocarbon in the gas phase in the presence of
molecular oxygen or of a gas comprising molecular
oxygen and of a suitable catalyst.
For example, the main process for the production of
25 acrylic acid is based on the oxidation of propylene
and/or propane. The synthesis of acrylic acid by
oxidation of propylene comprises two stages; the first
is _targeted at the oxidation of the propylene to give
acrolein and the second at the oxidation of the
30 acrolein to give acrylic acid. This synthesis is
generally carried out in two reactors using two
catalytic systems specific for each of the oxidation
stages, the two stages being carried out in the
presence of oxygen or of an oxygen-comprising gas.
35
In the same way, methacrolein and methacrylic acid are
produced industrially by catalytic oxidation of
isobutene and/or tert-butanol.
WO 2011 / 124824 PCT/FR201 1 /050692
2 -
Anhydrides, such as malefic anhydride or phthalic
anhydride, can be produced by catalytic oxidation of
aromatic hydrocarbons, such as benzene or o-xylene, or
5 of straight-chain hydrocarbons, such as n-butane or
butene.
Acrylonitrile is produced by catalytic oxidation in the
gas phase of propylene or propane by air in the
10 presence of ammonia, such a reaction being known under
the name of ammoxidation. Similarly, the ammoxidation
of isobutane/isobutene or methylstyrene results
respectively in methacrylonitrile and atroponitrile.
15 These various processes are carried out in the gas
phase, in the presence of molecular oxygen or of a gas
comprising molecular oxygen, generally in the presence
of air for economic reasons. Secondary reactions also
take place during the catalytic oxidation reactions,
20 including the reaction for the final oxidation of the
reactant, resulting in the formation of carbon oxides
and water.
Thus, these various processes for the manufacture of
25 hydrocarbon derivatives by selective oxidation have in
common the fact that they generate the formation of
carbon oxides, carbon monoxide CO and carbon dioxide
CO2. Carbon monoxide CO is generally produced in excess
with respect to the CO2. They are gaseous compounds
30 which are noncondensable under the temperature and
pressure conditions normally employed in the stages for
the recovery/purification of the hydrocarbon
derivatives. For this reason, they are discharged to
the atmosphere in the form of vents for the
35 incineration of the streams removed during the process,
with the disadvantage of emitting large amounts of CO2
to the atmosphere.
WO 2011/124824 PCT/FR2011/050692
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As in the majority of industrial processes, it is
advantageous to recycle some gas fractions generated
during the manufacture of hydrocarbon derivatives, in
5 particular the gas fraction comprising the unreacted
hydrocarbon, for obvious reasons of productivity and
profitability. In the case of the presence of carbon
monoxide in these recycled gases, there is a risk of
this noncondensable compound accumulating during the
10 process and causing safety problems, in particular
risks of loss of control of the oxidation reaction and
of explosion in the oxidation reactor, carbon monoxide
being flammable. Furthermore, carbon monoxide may have
a harmful effect on the oxidation catalysts; in
15 particular, it is a poison for many catalysts,
including the conventional catalysts for the oxidation
of propylene to give acrolein.
One solution for overcoming these disadvantages was
20 provided in the document EP 484 136, which describes a
process for the production of a hydrocarbon derivative
comprising the recycling of the gas stream resulting
from the reaction, after having successively extracted
or separated, from this stream, the ea(pected
25 hydrocarbon derivative, converted the CO present in
this gas stream into CO2 and then withdrawn a portion
of.the CO2 formed. The recycled stream is thus a stream
depleted in CO and CO2. The reaction for the oxidation
of the hydrocarbon is carried out in the presence of an
30 inert diluent, in particular in the presence of CO2, at
contents sufficiently high to prevent the formation of
a flammable mixture in the system. The catalyst for the
selective oxidation of the CO to CO2 is chosen from
catalysts which are nonoxidizing with respect to the
35 unreacted hydrocarbon present in the gas stream,
generally catalysts based on copper/manganese or
platinum/nickel mixed oxides, optionally supported on
WO 2011/124824 PCT/FR2011/050692
silica or alumina. The examples illustrating this
process are carried out with a CO converter comprising
a fixed catalyst bed. In the case of a process for the
production of acrylic acid from propylene, the tests
5 indicate a degree of 80% for the conversion of the CO
but also a conversion of a portion of the propylene and
propane, thus generating losses of reactant in the
process.
10 Another solution for overcoming these disadvantages was
provided in the document EP 1 007 499 relating to a
high-yield process for the production of maleic
anhydride from n-butane, consisting in bleeding off a
fraction of the recycling gases, so as to prevent inert
15 gases and carbon oxides from accumulating, and in
selectively oxidizing the carbon monoxide present in
the bleed stream to carbon dioxide in the presence of a
catalyst capable of selectively oxidizing the CO to CO2
in a gas having an 02/CO molar ratio of between 0.5 and
20 3. Catalysts which can be used in this case for the
selective oxidation of the CO to CO2 consist, for
example, of supported precious metals, such as
platinum, rhodium, ruthenium or palladium. The CO
converter is preferably of fixed bed tubular type.
25
The document EP 1 434 833 provides, as catalyst for the
selective oxidation of CO to CO2 in a gas stream
comprising at least one alkane, a catalyst substrate,
such as Pt, Pd, Pt-Fe or Pd-Fe, on a silica support,
30 provided with a virtually continuous coating of a
material consisting of a molecular sieve. The selective
oxidation of the CO is carried out in a fixed bed at a
temperature chosen so that the catalyst does not have
any effect on the alkane.
35
In the process for the production of methacrylic acid
from isobutane described in the document EP 495 504,
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use is made, after separation of the condensable
fraction comprising the methacrylic acid, of a stage of
oxidation of the CO present in the noncondensable gas
stream in order to convert it to CO2, followed by a
5 stage of separation of the CO2 in the carbonic
anhydride form by absorption in a liquid comprising,
for example, a solution of potassium carbonate and an
amine. Catalysts which can be used for the conversion
of the CO to C02, without oxidation of the isobutane,
10 for example comprise palladium and/or platinum on a
support, gold on a support, or manganese oxide. In
order to remove the heat of the oxidation reaction, a
reactor of multitubular type is used for the CO
converter.
15
In these various processes of the state of the art, the
selective oxidation of the CO to CO2 is carried out
either, on the one hand, at low space velocities SVs
(generally less than 10 000 h-1), the space velocity
20 being defined as the ratio of the flow rate of
reactants to the volume of catalyst, or as an inverse
of the "contact time", or, on the other hand, for low
concentrations of CO at the inlet; in some cases, it
also results in a simultaneous conversion of the
25 hydrocarbon present in the gas.
Due_ to the risks of flammability associated with the
presence of hydrocarbons and oxygen, processes for the
production of hydrocarbon derivatives by catalytic
30 oxidation in the gas phase of hydrocarbons generally
operate with low contents of hydrocarbons in order to
keep the gaseous reaction mixture outside the
flammability region.
35 In order to increase the productivity of these
processes while observing safety constraints, it is
possible to carry out the reaction of the oxidation of
WO 2011 / 124824 PCT'/ 82011 /050692
the hydrocarbon in the presence of an inert gas with a
high specific heat (also known as specific heat
capacity Cp, the index p indicating a value at constant
pressure), in order to obtain better management of the
5 exothermicity of the reaction and to increase the
concentration of the hydrocarbon in the reaction
mixture, the inert gas constituting a thermal ballast.
Thus, the document EP 293 224 provided for the addition
10 of 5 to 70% by volume of a saturated aliphatic
hydrocarbon of 1 to 5 carbon atoms, such as methane,
ethane or propane, for the reaction of the oxidation of
propylene resulting in acrolein in a two-stage process
for the production of acrylic acid by catalytic
15 oxidation, the second stage consisting in oxidizing the
acrolein to acrylic acid.
Propane is preferred as hydrocarbon which can be used
as thermal ballast in this process. This is because
20 this gas ballast exhibits several advantages with
respect to a ballast formed of inert gases, such as
nitrogen or carbon dioxide. First, it creates a better
thermal ballast as its specific heat (Cp) increases
strongly with the temperature, which is not the case
25 with nitrogen. In addition, it has a degree of chemical
inertia under the conditions under which the reaction
for_, the oxidation of propylene is carried out and its
possible reaction products are very similar in nature
to those of the propylene. Finally, it makes it
30 possible' to more easily meet the constraints of
composition of the mixture related to the question of
the flammability by placing the reaction mixture above
the upper flammability limit. By virtue of the use of
this thermal ballast, the feedstock supplying the
35 reactor for the oxidation of propylene can be greater
as fraction by volume, which increases the productivity
of the conversion while controlling the hot spots in
WO 2011 / 124824 PCT/FR2011 / 050692
7
the bed of catalysts and thus promoting the selectivity
of the reaction.
It is even more necessary to recycle the stream
5 comprising the unreacted propylene (and the propane) in
the case of the use of a thermal ballast such as
propane, due to the cost of the propane, and
consequently the oxidation of the carbon monoxide in a
stream rich in mixtures of hydrocarbons becomes more
10 difficult to carry out selectively.
The problem which the present invention intends to
solve is that of carrying out the conversion of the CO
to CO2 in a stream rich in hydrocarbons or in a mixture
15 of hydrocarbons, without, however; oxidizing these
hydrocarbons. It is possible, in an adiabatic fixed bed
or even in a reactor of the multitubular type with good
heat exchange, for selective conversion of the CO to
start but the reaction can rapidly run away, which
20 greatly increases the temperature of the catalyst and
brings about the conversion of the other compounds
present in the stream. Specifically, in a fixed bed
reactor, the large amounts of CO present lead to an
adiabatic increase in the temperature such that the
25 temperatures for the start of conversion (ignition
temperatures) of all the constituents are exceeded.
The exothermicity of the reaction for the conversion of
CO to CO2 is characterized by AHr of 283 kJ/mol. Even
30 if the exothermicity of the reaction for the selective
oxidation of hydrocarbon is comparable, for example AHr
is 341 kJ/mol for the conversion of propylene to
acrolein, the kinetics of the combustion reaction are
much faster. A "conventional" technology of the
35 multitubular reactor type then does not make it
possible to correctly control the temperature of the
catalyst. Under conditions where the temperature can be
WO 2011/124824 PCT/FR2O11/050692
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controlled, the conversion of the CO to CO2 can be
carried out at a lower temperature than the temperature
for oxidation of the hydrocarbon and thus it can be
carried out selectively.
5
As the differences in ignition temperature between CO
and the various constituents present in a stream of
hydrocarbons are significant, it is theoretically
possible to selectively oxidize CO to CO2 provided that
10 the reaction temperature is maintained between the
ignition temperature of CO and that of the other
constituents.
By way of indication, the increase in the temperature
15 of the gas during the combustion of 1000 ppm of
impurity in air is illustrated in the following
table 1.
Substance Heat of
combustion
kcal/mot
Adiabatic
temperature
increase of the
gas for 1000 ppm
Benzene 750.6 103
Toluene 892,0 123
m-Xylene 1038.9 143
Methyl Ethyl Ketone 540.1 74
Methyl Isobutyl
Ketone
843.0 116
Methanol 149.8 21
Formaldehyde 124.0 17
Acrolein 369.0 51
Acetic Acid 188.3 26
Butyric Acid 482.0 66
Ethyl Acetate 494.7 68
Phenol 690.0 95
m-Cresol 838.0 115
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PC'2/FR2011 /050692
Acetaldehyde 257.8 35
Styrene 1005.0 138
Hydrogen sulfide 122.5 17
Ammonia 76.2 11
Trimethylamine 531.0 73
Larbon Monoxide 67.6 9
Table 1
In an atmosphere rich in gas with a higher specific
heat than that of the air, such as propane, the
5 increase in adiabatic temperature is proportionally
lower.
In this context, the applicant company has discovered
that the use of a fluidized bed to convert the CO to
10 C02 in a stream rich in hydrocarbons makes it possible
to solve the various problems mentioned above and to
optimally meet the requirements of use of thermal
ballast and/or of recycling of reaction gas comprising
an excessively high CO content, in processes for the
15 production of hydrocarbon derivatives by selective
oxidation of hydrocarbons in the gas phase in the
presence of oxygen and more generally in processes for
the production of hydrocarbon derivatives from
hydrocarbons in the gas phase in the presence of
20 oxygen.
The oxidation of CO using a fluidized bed has already
formed the subject of studies, in particular for
studying the influence of the configuration of the
25 reactor comprising a specific catalyst of Pt-Co-Ce/y-
A12O3 type on the conversion of the CO present in a
hydrogen-rich stream (M.P. Lobera et al., Catalysis
Today, 157 (2010), 404-409). In this study, the treated
stream comprises only hydrogen, so that the problems of
30 selective oxidation of CO in a complex mixture, such as
Wa 2011/124824 PCT/FR2011/050692
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a mixture of hydrocarbons, the oxidation of which has
to be avoided, are not posed.
The aim of the present invention is thus to provide a
5 process for the selective oxidation of carbon monoxide
to carbon dioxide which is easily incorporated in
existing industrial processes for the production of
hydrocarbon derivatives.
10 Summary of the invention
A subject matter of the present invention is thus a
process for the selective oxidation of the carbon
monoxide present in a gas mixture comprising at least
15 one hydrocarbon or one hydrocarbon derivative, said
process comprising the stage consisting in bringing
said gas mixture into contact with a solid catalyst
capable of oxidizing the carbon monoxide to carbon
dioxide at a chosen temperature, characterized in that
20 said stage is carried out in a fluidized bed.
The fluidized bed technology provides mixing of the
solid and the gas mixture, and thus homogenization of
the temperature of the catalyst. By obtaining a
25 homogeneous temperature, it is thus possible to control
the selectivity of the reaction for the oxidation of
the, carbon monoxide and to no longer destroy the
molecules of economic value. This homogeneity confers,
on the fluidized bed, an undeniable advantage in
30 comparison with the fixed beds, which are generally
subject to a high temperature gradient. The removal of
the heat of the reaction can be provided by cooling
pins positioned in the fluid bed. The coefficient for
transfer of heat between the suspension and the
35 exchanger tubes is very high, and makes it possible to
efficiently heat or cool the material.
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The process for the selective oxidation of CO according
to the invention can be incorporated in any industrial
process requiring bleeding of CO2 and/or CO for
chemical reasons (inhibition of the reaction) , physical
5 reasons (decrease in the Cp of the reaction gas) or
safety reasons (flammability limits).
Another subject matter of the invention is a process
for the production of a hydrocarbon derivative
10 comprising at least one stage of selective oxidation of
the carbon monoxide present in a gas mixture comprising
at least one hydrocarbon or one hydrocarbon derivative
using a fluidized bed comprising a solid catalyst
capable of oxidizing carbon monoxide to carbon dioxide
15 at a chosen temperature.
Other characteristics and advantages of the invention
will more clearly emerge on reading the detailed
description which follows and the nonlimiting
20 implementational examples of the invention.
Detailed description
The gas mixture subjected to the process for the
25 selective oxidation of CO according to the invention
comprises at least one hydrocarbon or one hydrocarbon
derivative.
The hydrocarbons are saturated or mono- or
30 diunsatu'rated and linear or branched hydrocarbons
comprising from 2 to 6 carbon atoms, or aromatic
hydrocarbons, which can be substituted, comprising from
6 to 12 carbon atoms.
35 Mention may be made, as examples of hydrocarbons, for
example, of ethylene, propane, propylene, n-butane,
isobutane, isobutene, butene, butadiene, isopentene,
WO 2011 / 124824 PCT/F112011/ 050692
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benzene, o-xylene, methylstyrene or naphthalene.
Preferably, the hydrocarbon is chosen from propylene or
propane, alone or as a mixture.
5 The hydrocarbon derivative can be a product of the
partial oxidation of a hydrocarbon and it can then be
chosen from anhydrides, such as phthalic anhydride or
maleic anhydride, aldehydes, such as acrolein or
methacrolein, unsaturated carboxylic acids, such as
10 acrylic acid or methacrylic acid, unsaturated nitriles,
such as acrylonitrile, methacrylonitrile or
atroponitrile, or their mixtures. Preferably, the
hydrocarbon derivative is acrolein and/or acrylic acid.
15 The hydrocarbon derivative can also be a product of the
addition of oxygen or a halogen compound to an
unsaturated hydrocarbon, for example ethylene oxide,
propylene oxide or 1,2-dichloroethane.
20 Mention may be made, as solid catalysts capable of
oxidizing carbon monoxide to carbon dioxide which can
be used in the process according to the invention, of
known catalysts for the selective oxidation of CO, such
as, for example, without this list being limiting,
25 catalysts based on noble metals, such as platinum,
palladium, rhodium or ruthenium, supported on an
inorganic support, such as silica, titanium oxide,
zirconium oxide, alumina or silicalite; or catalysts
based on copper, manganese, cobalt, nickel or iron,
30 optionally in the presence of at least one noble metal,
such as platinum, palladium, rhodium or ruthenium, in
the form of mixed oxides or of alloys optionally
supported on an inorganic support, such as silica,
titanium oxide, zirconium oxide, alumina or silicalite.
35 Highly suitable catalysts are, for example, solids with
a low charge of platinum or palladium (for example of
the order of 2%) on a support of silicalite or sodium
WO 2011 / 124824 PCT/FR2O11/050692
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silicate type.
The. catalyst employed in the process according to the
invention is in the form of solid particles with a
5 particle size ranging from 20 to 1000 microns,
preferably from 40 to 500 microns, more particularly
from 60 to 200 microns. The size distribution of the
particles can be determined according to numerous
methods, in particular according to a simple method,
10 such as sieving with a sequence of sieves of decreasing
mesh sizes, or determination by laser diffraction, for
example with devices of the Malvern brand.
According to the invention, the temperature of the
15 fluidized' bed is between 20°C and 400°C, preferably
between 70°C and 300°C, more preferably between 100°C
and 230°C.
Preferably, a temperature is chosen which is lower than
20 the "ignition temperature of the hydrocarbon and/or of
the hydrocarbon derivative present in the gas mixture,
that is to say lower than the temperature corresponding
to the start of the reaction for the oxidation of the
hydrocarbon and/or hydrocarbon derivative.
25
For example, in the case of propylene, there is
generally at least approximately 20°C and preferably at
least 30°C difference between the ignition temperature
of CO (beginning of combustion of the CO) and the
30 ignition temperature of propylene. This difference is
sufficient to guarantee combustion of the CO without
having combustion of the propylene. If the ignition
temperature of propylene is reached, the reaction
becomes difficult to control due to the high oxygen and
35 propylene content of the gas to be treated.
The fluidized bed can operate batchwise or continuously
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(semibatchwise or open). Preferably, the process
according to the invention is carried out continuously.
This is because, given the ease of withdrawal of solid
particles from the fluidized bed and of addition of
5 solid particles to the fluidized bed during its
operation, the solid phase can be continually replaced
as required. The catalyst bed can maintain an unvarying
activity over time if deactivated catalyst is
continually withdrawn in order to replace it with fresh
10 catalyst. The emptying and the cleaning of the
fluidized beds take place very easily, as for a water
tank. The withdrawal operations can be carried out
continuously or with a degree of periodicity.
Furthermore, the deactivated catalyst can optionally be
15 reactivated ex situ by any appropriate technique, in
order to be subsequently reinjected into the reactor.
Mention may be made, as techniques for reactivation of
the catalyst, without being limiting, of redispersion
of the metals of the catalyst by a reducing treatment,
20 washing the catalyst in order to remove the
contaminants, or reimpregnation of the catalyst with a
fresh charge of the active metals of the catalyst.
Reference may be made to the document on fluidization
25 techniques in Techniques de l'Ingenieur [Techniques of
the Engineer] J3 390 1 to 20.
Advantageously, mild oxidation conditions are used,
that is to say a combination of a relatively inactive
30 catalyst (comprising a low charge of active metal), a
moderate temperature (for example from 80°C to 180°C)
and a fairly short contact time (for example of less
than one second) . Alternatively, a low residence time
makes it possible to envisage a higher temperature for
35 a given catalyst.
Advantageously, use is made of a moderate pressure in
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the fluidized bed reactor, for example of between 1 and
3 bar, and a fairly short residence time, which is
reflected by high space velocities. In the case where
the conditions related to the process require higher
5 pressures, it is then preferable to combine a
relatively inactive catalyst with a temperature which
is as low as possible in order to limit the reactions
for the oxidation of the other constituents of the gas
stream.
10
The linear velocity of the gas mixture in the fluidized
bed can range from 0.1 to 80 cm/s. For fluidized beds
of industrial size, it can range from 50 to 80 cm/s. In
the case of shorter fluidized beds, in particular
15 laboratory fluidized beds, the linear velocity of the
gases is generally between 0.1 and 10 cm/s,.
Generally, the flow rate by volume of the gas stream
will be adjusted to the volume of catalyst and
20 consequently to the size of the reactor, so as to
achieve very high space velocities, SVs, expressed as
hourly flow rate by volume of reactants with respect to
the volume of catalyst.
25 The process according to the invention is
advantageously carried out with high space velocities
SVsr, for example ranging from 1000 h-1 to 30 000 h-1,
preferably greater than 10 000 h-1, or better still
ranging from 10 000 h_1 to 30 000 h-1. The process
30 according to the invention is particularly well suited
to gas streams comprising more than 0.5 mol% of carbon
monoxide at the inlet and preferably more than 1 mol%
of carbon monoxide.
35 Another subject matter of the invention is a process
for the production of a hydrocarbon derivative
comprising at least the following stages:
WO 2011 / 124824 PCT/FR2011/050592
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a) at least one hydrocarbon and oxygen or an oxygencomprising
gas are brought into contact with an
appropriate catalyst, resulting in a gas mixture
5 comprising at 'Least one hydrocarbon derivative,
unconverted hydrocarbon, oxygen and carbon monoxide,
b) the hydrocarbon derivative is separated or extracted
from the reaction stream resulting from stage a),
10
c) the carbon monoxide present in the gas stream is
then converted to carbon dioxide using a fluidized bed
comprising a solid catalyst capable of oxidizing carbon
monoxide to carbon dioxide at a chosen temperature,
15 producing a gas stream depleted in carbon monoxide,
d) said stream depleted in carbon monoxide is recycled
to reaction stage a).
20 It is clearly understood that this process can comprise
other stages preliminary, intermediate and/or
subsequent to those mentioned above which are well
known to a person skilled in the art.
25 According to the process of the invention, stage a) is
carried out under appropriate conditions, in particular
regarding the nature of the catalyst, the temperature
and the optional presence of an inert gas as thermal
ballast, according to processes known to a person
30 skilled in the art which make it possible to
manufacture the desired hydrocarbon derivative.
The reaction carried out can be an oxidation reaction
or a reaction for the addition of oxygen to an
35 unsaturated hydrocarbon.
In an alternative form of the invention, stage a) is
WO 2011/124824 PCT/FR2011/050692
carried out in the presence of a thermal ballast which
is inert under the conditions of the reaction carried
out.
5 Stage b) consists in recovering the hydrocarbon
derivative according to conventional methods using
techniques such as absorption in a solvent followed by
extraction, distillation, crystallization,
condensation.
10
On conclusion of this stage b), the gas stream, freed
from most of the hydrocarbon derivative, generally
comprising unconverted hydrocarbon, oxygen, water
vapor, inert gases, such as nitrogen and argon, carbon
15 monoxide and carbon dioxide, is brought into contact
with a solid catalyst capable of oxidizing the carbon
monoxide to carbon dioxide at a chosen temperature, in
a fluidized bed (stage c)), resulting in a stream
depleted in carbon monoxide, which can be recycled,
20 according to stage d), to the reaction stage a),
optionally after having bled it of a portion of the
carbon dioxide formed.
In a specific embodiment of the invention, the process
25 for the production of a hydrocarbon derivative
incorporating a stage c) of partial oxidation of CO to
C02 -using a fluidized bed relates to the manufacture of
acrylic acid by catalytic oxidation of propylene using
oxygen or an oxygen-comprising mixture and in the
30 presence: of propane as thermal ballast.
This reaction, widely operated industrially, is
generally carried out in the gas phase and generally in
two stages:
35
The first stage carries out the substantially
quantitative oxidation of the propylene to give an
wo 2011/124824 PCT/FR2011/050692
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acrolein-rich mixture, in which the acrylic acid is a
minor component, and then during the second stage
carries out the selective oxidation of the acrolein to
give acrylic acid.
5
The reaction conditions of these two stages, carried
out in two reactors in series or in a single reactor
comprising the 2 reaction stages in series, are
different and require catalysts appropriate to the
10 reaction; however, it is not necessary to isolate the
first-stage acrolein during this two-stage process.
The reaction conditions of these two stages, carried
out in two reactors in series or in a single reactor
15 comprising the 2 reaction stages in series, are
different and require catalysts appropriate to the
reaction; however, it is not necessary to isolate the
first-stage acrolein during this two-stage process.
20 The reactor can be supplied with a propylene feedstock
of low purity, that is to say comprising propane, such
that the propane/propylene ratio by volume is at least
equal to 1. As the large gas ballast formed by the
propane results in a better management of the
25 exothermicity of the reaction, the reactor can be
supplied with a feedstock more concentrated in
propylene in order to increase the productivity of the
process. The other components of the reactive stream
can be inert compounds, such as nitrogen or argon,
30 water and oxygen.
The gas mixture resulting from the reaction for the
oxidation of acrolein consists, apart from acrylic
acid:
35
of light compounds which are noncondensable under
the temperature and pressure conditions normally
WO 2011 / 124824 PCT/FR2011/ 050692
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employed: nitrogen, unconverted oxygen,
unconverted propylene, propane present in the
propylene or added as thermal ballast, carbon
monoxide and carbon dioxide, which are formed in a
5 small amount by final oxidation or going around in
circles, by recycling, in the process,
of condensable light compounds: in particular,
water generated by the reaction for the oxidation
10 of propylene, unconverted acrolein, light
aldehydes, such as formaldehyde and acetaldehyde,
acids, such as acetic acid, the main impurity
generated in the reaction section,
15 - of heavy compounds: furfuraldehyde, benzaldehyde,
maleic acid, maleic anhydride, benzoic acid, 2-
butenoic acid, phenol, and the like.
The second stage of the manufacture, corresponding to
20 stage b) of the process according to the invention,
consists in recovering the acrylic acid present in the
gaseous effluent stream resulting from the oxidation
reaction.
25 This stage can be carried out by countercurrentwise
absorption. For this, the gas resulting from the
reactor is introduced at the bottom of an absorption
column, where it encounters, countercurrentwise, a
solvent introduced at the column top. The light
30 compounds, under the temperature and pressure
conditions normally employed (respectively more than
50°C and less than 2 x 105 Pa), are removed at the top
of this absorption column. The solvent employed in this
column is water. The water could be replaced by a
35 hydrophobic solvent having a high boiling point, as is
described, for example, in the BASF patents FR
WO 2011/124824
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PCT/FR2011 /050692
2 146 386 or US 5 426 221, and in the patent FR
96/14397.
The operating conditions of this absorption stage are
5 as follows:
The gaseous reaction mixture is introduced at the
column bottom at a temperature of between 130°C and
250°C. The water is introduced at the column top at a
10 temperature of between 10°C and 60°C. The respective
amounts of water and gaseous reaction mixture are such
that the water/acrylic acid ratio by weight is between
1/1 and 1/4. The operation is carried out at
atmospheric pressure.
15
An aqueous mixture of acrylic acid in water (ratio by
weight from 1/1 to 4/1) freed from most of the
unconverted acrolein and light compounds, in particular
noncondensable light compounds, including CO, is thus
20 obtained, this mixture generally being referred to as
"crude acrylic acid".
This crude acrylic acid is
combination of stages which
then subjected to a
can differ in their
25 sequence according to the process: dehydration, which
removes the water and the formaldehyde (dehydrated
acrylic acid), removal of the light products (in
particular the acetic acid), removal of the heavy
products, optionally removal of certain impurities by
30 chemical' treatment.
The gas stream exiting from the preceding stage of
extraction of the acrylic acid by countercurrentwise
absorption, which consists mainly of unconverted
35 propylene, unconverted oxygen, propane, CO and CO2, and
other minor inert gases or light impurities, is brought
into contact with a solid catalyst capable of oxidizing
WO 2011/124824 PC'y/FR20111 /050592
- 21 -
carbon monoxide to carbon dioxide at a chosen
temperature, in a fluidized bed, resulting in a stream
depleted in carbon monoxide which can be recycled to
the reaction stage, optionally after having bled it of
5 a portion of the carbon dioxide formed.
In an alternative form of the process of the invention,
all or part of the gas stream exiting from the unit for
the conversion of carbon monoxide to carbon dioxide is
10 sent to a selective permeation unit in order to
separate a first stream predominantly comprising the
inert compounds, such as CO, C02, nitrogen and/or
argon, and a second stream predominantly comprising
propylene and propane. The permeation unit employs one
15 or more semipermeable membranes having the property of
separating the inert compounds from the hydrocarbons.
This separation generally takes place at a pressure of
the order of 10 bar and at a temperature of
approximately 50°C. Use may be made of membranes based
20 on hollow fibers composed of a polymer chosen from:
polymides, polymers of cellulose derivatives type,
polysulfones, polyamides, polyesters, polyethers,
polyetherketones, polyetherimides, polyethylenes,
polyacetylenes, polyethersulfones, polysiloxanes,
25 polyvinylidene fluorides, polybenzimidazoles,
polybenzoxazoles, polyacrylonitriles, polyazoaromatics
and-the copolymers of these polymers.
Said second stream enriched in propylene and propane is
30 advantageously recycled to the reaction stage without
accumulation of CO2 and other inert gases, such as
argon, in the recycling loop.
In a second alternative form of the process of the
35 invention, all or part of the gas stream entering the
unit for the conversion of carbon monoxide to carbon
dioxide is sent beforehand to a selective permeation
WO 2011 / 124624 PCT/FR2011/050692
- 22 runit,
such as that described above, to separate at
least a portion of the 002r the gas stream entering the
CO converter then being depleted in 002 and unconverted
oxygen.
5
Other characteristics and advantages of the invention
will become apparent in the experimental part which
will follow.
10 EXPERIMENTAL PART
Example 1
Oxidation tests were carried out on pure compounds
15 using a catalyst from Johnson Matthey (2% Pt on Ce02)
in a reactor comprising a bath of molten salt with an
internal diameter of 25.4 mm and with a catalyst height
of 30 cm (i.e., 164 g).
20 The test consisted in monitoring the conversion of the
pure compound (conversion test where each constituent
is tested individually) in a mixture of nitrogen and
oxygen (3 mol%). For some tests, a portion of the
nitrogen was replaced by water (20 mol%) . The following
25 concentrations (which represent the order of magnitude
of the concentrations expected for each of the
reactants in a real stream) were tested under SV
conditions of 25 000 h-1, the conversion of the compound
being determined as a function of the temperature:
30
- CO: 2.8 mol%
- Propylene: 0.75 mol%
- Acrolein: 0.75 mol%
- Propane: 50 mol%
35 - optionally water
WO 2011/124824 PCT/FR,2011/050092
- 23 -
These conditions correspond to conditions for the
oxidation of propylene in the presence of propane as
thermal ballast.
5 The change in the oxidation of the pure compounds as a
function of the temperature, respectively in the
presence of water and in the absence of water, is
reproduced in figures 1 and 2. In figure 1, the
conversion of the CO is 100% for a temperature of 180°C
10 and the oxidation of the propylene begins significantly
at a temperature of 235°C. In figure 2, curve 1
corresponds to acrolein, curve 2 to CO and curve 3 to
propylene.
15 The "ignition" temperatures corresponding to the start
of the oxidation reaction for the pure compounds are
collated in table 2 below. In this table, the
temperatures shown correspond to the temperatures
applied to the reactor and not those of the catalyst.
20
Compounds CO Propylene Acrolein Propane
Ignition
temperature 225°C 265°C 285°C
without water
Ignition
temperature with < 180°C 235°C > 300°C
water
Table 2
With the pure compounds, the ignition temperature of
the reaction is significantly different (difference
25 > 30°C), showing an advantage for sufficient evacuation
of the energy to keep the reaction temperature at
temperatures where the selectivity is good.
Example 2 (comparative)
30
WO 2011/124824 PCT/FR2011/050692
- 24 -
Example 1 is reproduced with the fixed bed of catalyst
and a stream comprising the mixture of the compounds
CO,.C02, propane, propylene, acrolein, water and oxygen
with the following concentrations, in which it is
5 desired to selectively oxidize the CO to C02:
- CO: 2.8 mol%
- Propylene: 0.75 mol%
- Acrolein: 0.75 mol%
10 - Water: 20 mol%
- Oxygen: 3 mol%
- Propane: 50 mol%
The tests carried out show that, in the presence of the
15 mixture of compounds, complete combustion of the
reactants is observed as long as the oxygen is
available, in the following order of reactivity:
CO > propylene > acrolein > propane
20
The temperature is not controlled: a hot spot can be
measured where the temperature reached in the catalytic
bed is greater by 150°C at least than that of the oven.
Consequently, this difference being much greater than
25 that measured between the ignition temperatures of the
pure substances, all the oxidation reactions are
stressed at the same time, increasing even more the
overall exothermicity. The oxidation reactions are
found to be slowed down only when the oxygen has been
30 consumed.
The test was reproduced with dilution of the catalyst
by 90% by weight with inert materials (alumina/silica
beads originating from Saint-Gobain) . The diluting of
35 the catalyst has the object of reducing the catalytic
activity of the reactor in order to exert better
control over its operating temperature. Thus, the
WO 2011/124824 PCT/FR2011/050692
- 25 -
reactor was filled with 10% by weight of the initial
catalyst and 90% by weight of inert materials for an
equivalent volume of catalyst + inert materials. The
aim is also to distribute the heat from the oxidation
5 reaction over a greater reaction volume. The heat from
the reaction is removed by heat transfer by the wall;
the inert solid provides a greater number of points of
contact between the catalyst and the wall and should
thus make possible better control of the temperature in
10 the catalytic bed. Despite high dilution of the
catalyst (which amounts to a strong increase in the
space velocity), the difference in temperature between
the hot spot of the catalytic bed and the temperature
of the oven (or the ignition temperature of the
15 reaction of the oxidation of CO) remains greater than
the difference in ignition temperature of CO and
propylene.
Under these conditions, the oxidation reaction is not
20 always controlled and the diluting of the catalyst with
inert materials (90% by weight) does not make it
possible to solve the problem. Here again, the catalyst
is inadequate for the reactor technology used.
25 Example 3: Preparation of catalysts for the selective
oxidation of carbon monoxide
Preparation of a catalyst A
30 200 g of porous silica spheres with a diameter of 80
microns are impregnated, by nascent humidity
impregnation, with a solution comprising 10.2 g of
citric acid, 20.2 g of tetraammineplatinum(II)
hydrogencarbonate (comprising 50.6% of platinum), 8 g
35 of iron(III) nitrate nonahydrate and 103.5 g of
demineralized water. Gentle heating with stirring is
used to evaporate the excess water, the solid being
WO 2011 / 124824 PCT/FI12011/050692
- 26 -
kept rotated in a rotating oven in order to prevent
agglomeration. Finally, the powder is dried at 105°C
and then calcined under air at 500°C for 2 hours.
5 Preparation of a catalyst B
A catalyst is prepared by impregnation of Puralox@ SCCA
5-150 alumina from Sasol according to the following
protocol:
10
300 g of alumina are introduced into a 3 1 jacketed
reactor heated to 100°C and flushing with air is
carried out in order to fluidize the alumina. A
solution of 15.3 g of citric acid, 30.3 g of
15 tetraammineplatinum(II) hydrogencarbonate (comprising
50.6% of platinum), 12 g of iron nitrate nonahydrate
and 155 g of demineralized water is then continuously
injected using a pump. The ratio targeted (weight of
metal / weight of final catalyst) being 0.5% Pt-0.5% Fe
20 % by weight, the duration of addition of the solution
is 2h. The catalyst is subsequently left at 105°C in
an oven for 16 h and then calcined at 500°C for
2 hours.
25 This alumina has, at the start, grains having a median
diameter of approximately 85 μm and exhibits the
surface and porosity characteristics indicated below:
BET surface (m2/g) 148
30 Hg total'' pore volume (cm3/g) 0.87
D50: apparent mean diameter of 50% of the population of
the particles: 85 μm
Porosity peak (nm): 9
35 Example 4
Example 2 is reproduced but using a fluidized bed.
WO 2011/124824 PCT/F122011/050692
- 27 -
The catalysts A and B are employed in a fluidized bed
supplied with a stream having the composition described
in table 3 below and preheated to 100°C. This stream
5 provides for the fluidization of the catalyst. The
total pressure in the reactor is 2.2 bar, with a linear
velocity of the gases of 10 cm/s.
Molar composition %
Acetaldehyde 0.11%
Acrolein 0.24%
H2O 2.43%
02 10.26%
Argon 1.99%
CO 3.95%
C02 23.90%
Propane 53.87%
Propylene 1.67%
Acetic acid 0.00%
Acrylic acid 0.01%
Nitrogen 1.56%
Pressure (bar) 2.2
Temperature (°C) 40
Table 3
10
Thef products, are collected at the reactor outlet and
are analyzed
the water.
by chromatography after having condensed
15 The results obtained appear in table 4 below:
Catalyst Conversion Conversion Conversion Conversion
of the CO of the of the of the
%) propylene acrolein propane
(%) (%) (%)
WO 2011/124824
-- 28 -
PCT/fl2011/050692
A
- -----------------
90 5 3
B 100 4 1 na
Table 4
Example 5
The catalysts C, D and E defined below are obtained in
the form of granules and were milled to a particle size
of less than 315 microns. 150 g of this powder are
sieved in order to select the fraction between 80 and
160 microns.
10
Catalyst C: ND-520 catalyst from N.E. Chemcat,
Pt/alumina, in the form of 3 mm beads, with a density
of 0.74 kg/1.
15 Catalyst D: DASH 220 catalyst from N.E. Chemcat, 0.5%
Pt/alumina, in the form of 3.2 mm granules, with a
specific surface of 100 m2/g.
Catalyst E: ND103 catalyst from N.E. Chemcat, 0.5%
20 Pd/alumina, in the form of beads with a diameter of
3 mm, with 250 m2/g.
150 g of solid catalyst were placed in a fluidized bed.
The fluidized, bed consists of a stainless steel tube
25 with a diameter of 41 mm and a total height of 790 mm.
The fluidized bed is immersed in a fluidized sand bath
heated by electrical elements installed inside the
bath. Three thermocouples recorded the temperature
gradient along the tube. A stream with the molar
30 composition described in table 5 below was supplied at
a flow rate of 1760 ml/min (standard conditions), i.e.
a linear velocity of the gases of approximately
2.2 cm/s, below a porous metal plate which distributes
the gas across the diameter of the reactor.
35
WO 2011/124824
11 29 -
The total pressure in the fluidized bed is 1 bar and
the temperature is maintained at approximately 100°C.
Molar composition of
the gas entering the
fluid bed
%
Acetaldehyde 0.120
Acrolein 0.31%
H2O 2.68%
02 4.39%
Argon 1.95%
CO 5.31%
CO2 13.34%
Propane 68.73%
Propylene 1.93%
Acetic acid 0.00%
Acrylic acid 0.01%
Nitrogen 1.24%
Pressure (bar) 1
Temperature (°C) 50
Table 5
5
The products are collected at the fluidized bed outlet
and, after having condensed the water, are analyzed by
liquid chromatography.
10 The results obtained appear in table 6 below:
Catalyst Conversion Conversion Conversion Conversion
of the CO of the of the of the
%) propylene acrolein propane
(%) (%) (%)
C 82 3 1 0
D 90 4 1 0
E 65 3 1 0
WO 2011/124824 PCT/FR2011/050692
- 30 -
Table 6
Example 6: Preparation of catalysts
5 Three series of tests were carried out starting from
the following catalysts:
- 0.5Pt catalyst (corresponding to 0.5%Pt and
0.5%Fe/on alumina A1203), bulk density
10 0.78 g/ml, mean particle size 114 μm.
15
l.5Pt catalyst (corresponding to 1.5%Pt and
l.5%Fe/on alumina A1203), bulk density
0.79 g/ml.
2Pd catalyst (corresponding to 2% Pd/on zeolite
D/UR from Engelhard/BASF), bulk density
0.55 g/ml, mean particle size 144 μm.
20 The first twocatalysts were prepared according to the
protocols below and the third catalyst is a catalyst
sold by BASF. The latter was used as is without
specific modification apart from a predrying.
25 The catalysts for the selective oxidation of CO were
prepared by impregnation of the precursors in solution
(Pt-and Fe salts) on the Sasol alumina Puralox SCCA 5-
150.
30 1) In 'a first step, the catalysts for micro-fluidized
bed tests were prepared by nascent humidity
impregnation.
The platinum and iron salts and also citric acid were
35 weighed out and dissolved in a beaker with water. The
0.5%Pt/0.5% Fe/alumina catalyst was prepared by mixing
0.049 g of citric acid, 0.0987 g of
WO 2011 / 124824 PC,T/FR2011 / 050692
31 -
tetraammineplatinum(II) hydrogencarbonate (comprising
50.6% of Pt), 0.36 g ,of ferric nitrate hydrate, 8.10 g
of water and 9.9 g of alumina. The total volume of the
solution was calculated in order to be equal to the
5 total pore volume of the support (0.87 ml/g). The
solution was subsequently gradually added to the
support while stirring vigorously. It was completely
incorporated by the support and no liquid was
detectable at the surface or between the particles. The
10 resulting paste was dried at 110-120°C for 24 h and
then calcined in the air at 485-520°C for 2 hours.
2) For the tests in a larger reactor, another
technique was employed where the impregnation took
15 place in a fluidized bed, the method described above
not being very applicable to 200 g of catalyst, in
particular for reasons of homogeneity. The following
method was used to prepare 200 g of 0.5%Pt/0.5%Fe/Al203.
The same precursor salts and the same concentrations
20 were used as for the preparations described above.
The solution of precursor and of citric acid is
supplied dropwise above the support placed in a
cylindrical quartz tube heated to 105-125°C and
25 fluidized with air. The liquid flow rate was calculated
and maintained at 0.5-1 ml/min. The injection nozzle
was- placed at approximately 10 cm above the bed. The
air was used as fluidization gas and the flow rate was
adjusted in order to maintain an ebullating bed state
30 and also in order to prevent pneumatic transportation
of the particles. The solid thus obtained was dried at
120°C for 24 h and then calcined at 500°C.
Two levels of charging were carried out for the tests
35 in a microreactor (10 grams of 1.5Pt/1.5%Fe/A1203 and
10 grams of 0.5Pt/0.5%Fe/Al203) and just one for the
other tests in the larger reactor (200 grams of
WO 2011 / 12482 4 PCT/` 'R2011 / 050692
- 32 -
5
0.5Pt/ 0.5%Fe /A1203).
Example 7: Tests in the fluidized microreactor starting
from the catalysts of example 6
Use was made of a microreactor consisting of a quartz
tube with an internal diameter of 7 mm in which 1 g of
catalyst was arranged on a sintered glass (20 μm)
placed in the middle of the tube acting as distributor.
10 The entire assembly is computer controlled and it is
possible to record all the relevant experimental
parameters (temperatures, gas flow rates).
The tests were carried out under 40 ml/min for 15 min.
15 This corresponds to a space velocity SV of
approximately 3000 h-1.
20
The composition of the inlet gas supplied to the
microreactor is as follows:
Gas Ar 02 C3H6 CO CO2 C3H9
vol. 5.530 7.61% 1.69% 5.19% 10.6% 66.79%
Between two tests, 40 ml/min of argon were :cent in
order to purge the reactor for 10 min. Argon is also
the inert material of choice during the heating
25 periods.
In order to have an idea of the exothermicity of the
oxidation reaction, monitoring of the temperature was
carried out on the temperature of the oven and not on
30 that of the catalyst bed. For each experiment carried
out, the operating conditions were kept constant during
the 15 min of the test.
The gases at the outlet of the reactor were analyzed by
35 an inline mass spectrometer.
WO 2011/124824 PCT/FR2011/050692
- 33 -
Figure 3 is presented as an example and gives the
composition of the outlet gas after reaction and also
the change in the temperature of the catalytic bed. All
5 the experiments carried out give the same type of
profile with stationary levels. It is from these data
that a mean conversion was calculated.
The tests carried out in the microreactor are
10 summarized in table 7 below.
Test Catalyst
Temperature of
the bed attire start
of the test (°C)
Maximum
6Tofthe
bed (°C)
Temperature of
the oven (°C)
(constant)
R co (°^a R >rr,>lrykne (9a) H pmpac (%)
m ..... 112
_
z .. N 100 4 0 II
110
II
1
-
1la..._ .-
-- _
4-7
? fx4 II
_
726 62 _,73
d 148 [30 58 - (3 6
U3Pe 161) 1b 14r,""" gO t1 6 ... 772 007th-
7 733 12. 161r 106 -Ltl
195 14 170 714) 0 Ir
16
1.1
206
-. 7io _. 228 _
13
., ( _
.g
166
.,. ^4tY ^..^.
Z64
GU
__. 1071_
100
7
'A
49
5
. ,....-, 2177 ..
17 777 ..... - R IPCr _.._ ci 1 ^^ @""'"^
1 1-5M 40 v .
11fI 62
......_._._
4
.....
iâ
74 5
1
140 d ;220 100
6
1
IS 152 IS 1.x70 100 13 Il
if, 16% _ 33 140 10(1
_
+1 I(t
17 _ 13 310 030 00 _ X13 16
7k 185 34 1100 100 49 79
7^7 196 33 170
_
1001 43 IU
20 207 i3 11400
^^
1011 4$ I(}
2l
--
219 32 190 . IOU - 13
^22 324 29 200 00 d5 16
23 13R n. a. .00 92 t3 a1
24
..
707
...
n. a. Elgi 3137! U p
146 75d n.a... 7F0 10t1... {}..
37 2Fd 161 ma. 140) IOU . 5. ... 1
70 n.a. 134}
_
(i I I U
1+37
J
197
'UR
N9
na
n.a
Ma.
170
0
1010
IUo
11
^ __ 235. n. a. 21113 100 20 4}
n.a. for not available
15 Table 7
The lines in bold characters are the experiments for
which the oxidation of CO was complete and selective.
20 These tests have shown that the selective oxidation of
WO 2011 / 12482 4 PCT/FR2011 / 050592
- 34 -
CO in the fluidized bed is highly efficient, with
conversions of 100% and the absence of oxidation of
propylene and propane, at well controlled temperatures,
since the increase in temperature in the bed was
5 minimal.
It is therefore entirely possible to envisage the
selective oxidation of CO in a fluidized bed on a
larger scale provided that mild oxidation conditions
10 and good heat transfer are provided. Starting from
mathematical principles using kinetic constants, it is
then entirely possible to determine the optimum
operating conditions to be employed on the industrial
scale.
15
Example 8: Tests in a fluidized bed reactor
A fluidized bed of greater size makes it possible to
carry out tests on amounts of catalyst which have
20 reached 300 g. The metal reactor is divided into two
sections: a reactive section (diameter 3.5 cm, height
60 cm) and a withdrawal region (diameter 4.5 cm, height
40 cm). The gases are supplied and regulated via flow
regulators. These make it possible to achieve a maximum
25 flow rate by volume in this plant of 4 1/min. Control
of the temperature is provided by an external sand bath
which makes it possible to provide good homogeneity of
the temperature in the reactor. Four measurement points
for the temperature were installed inside the reactor
30 (two in the catalytic bed and two in the withdrawal
section). A valve was also installed on the outlet line
in order to be able to raise the reactor up to the
desired pressure. A pump of HPLC type introduces an
unchanging flow of an aqueous acrolein solution
35 directly into the catalytic bed.
An inline mass spectrometer makes it possible to
WO 2011 / 124824 PCT/FR2011/050692
- 35 -
quantify the products. A line heated to 110°C connects
the mass spectrometer to the assembly and thus prevents
any condensation of water. Filters installed on the
outlet line prevent the particles from exiting from the
5 reactor and prevent the lines from becoming blocked.
Four different gas compositions were tested and are
summarized in the following table 8. The tests denoted
Tl were carried out under 2.2 bar absolute, while the
10 tests denoted T2 were carried out at atmospheric
pressure. Each Test, Tl or T2, was carried out "dry" or
"wet" (in the presence of water vapor).
In % Ar O C3H6 CO C02 C3H8 C3H4O H2O
T1 dry 3,8 11,3 1.7 4.4 23.9 54.9 0 0
Ti wet 3.4 10.1 1.5 4.0 21.4 49.2 2.7 7.7
T2 dry 3.2 4.5 1,8 5.5 14,0 71.0 0 0
T2 wet 2.9 4.0 1.6 4.9 12.6 63.6 3.0 7.3
Table 8
15
The, tests were carried out on 150 g of 0.5Pt catalyst
prepared according to example 6. The flow rate by
volume of the gases was maintained at approximately
600 ml/min for 15 to 40 minutes (depending on the test
20 and on the time for stabilization of the outlet
concentrations) for the tests and for the purges with
argen in order to provide good fluidization of the
particles. This represents an hourly space velocity of
approximately 180 h-1. It was not possible to operate
25 with higher gas flow rates, in order to achieve much
shorter contact times and a greater hourly space
velocity, without bringing about entrainment of the
catalyst as a result of the small diameter of the
reactor. Nevertheless, these results are representative
30 of the tests which a person skilled in the art is
capable of extrapolating for a unit of greater size,
with greater gas linear velocities and greater hourly
WO 2011 / 124024 PCT/FR2011./050692
- 36 -
space velocities.
A mean conversion was calculated over the whole of the
15-40 minutes of the test,
5
The results for test 1, carried out under a pressure of
2.2 bar, and test 2, carried out under a pressure of
1 bar, are summarized in tables 9 and 10 respectively.
Test 1 Tempera- Maximum Duration Con- Con- Con-
2.2 bar Lure of AT of of the version version version
the bed the experi- of the of the of the
at the bed ment CO (°s) propylene propane
start of (°C) (min ) (o) (o)
the test
(°C)
Dry 100 3 35 27 10 1
Dry 116 43 30 100 17 1
Wet 100 < 1 25 16 1 2
Wet 110 48 40 100 26 5
10 Table 9: Test 1 (2.2 bar)
Test Tempera- Maximum Duration Con- Con- Con- '.
2 ture of AT of of the version version version
1 bar the bed the experi- of the of the of the
at the bed ment CO (o) propylene propane
start of (°C) (min) ( o) (o)
the test
(°C)
Dry 100 2 15 39 0 0
Dry 120 4 20 93 0 0
Dry 140 4 20 100 0 0
Wet 110 1 15 27 1 1
Wet 120 1 15 12 1 1
Table 10: Test 2 (1 bar)
Wa 2011/124824 PCT/FR2O11/050692
. 37 -
The duration of the reaction was adjusted so as to
achieve a stationary state for the outlet
concentrations.
5 No conversion of propylene or of propane was observed
under 1 bar, even at temperatures above 140°C.
Furthermore, the increase in temperature in the bed was
very limited, indicating very good heat transfer in the
fluidized bed.

CLAIMS
1. A process for the selective oxidation of the
5 carbon monoxide present in a gas mixture
comprising at least one hydrocarbon or one
hydrocarbon derivative, said process comprising
the stage consisting in bringing said gas mixture
into contact with a solid catalyst capable of
10 oxidizing the carbon monoxide to carbon dioxide at
a chosen temperature, characterized in that said
stage is carried out in a fluidized bed.
2. The process as claimed in claim 1, characterized
15 in that the hydrocarbon is chosen from saturated
or mono- or diunsaturated and linear or branched
hydrocarbons comprising from 2 to 6 carbon atoms,
or aromatic hydrocarbons, which can be
substituted, comprising from 6 to 12 carbon atoms.
20
3. The process as claimed in claim 1 or 2,
characterized in that the hydrocarbon derivative
is chosen from anhydrides, such as phthalic
anhydride or maleic anhydride, aldehydes, =;uch as
25 acrolein or methacrolein, unsaturated carboxylic
acids, such as acrylic acid or methacrylic acid,
unsaturated nitriles, such as acrylonitrile,
methacrylonitrile or atroponitrile, or their
mixtures, ethylene oxide, propylene oxide or 1,2-
30 dichloroethane.
4. The process as claimed in any one of claims 1 to
3, characterized in that the catalyst is in the
form of solid particles with a particle size
35 ranging from 20 to 1000 microns, preferably from
40 to 500 microns.
W9 2011 /124824 PCT3FR2011 /050692
- 39 -
5. The process as claimed in any one of the preceding
claims, characterized in that the catalyst is
chosen from catalysts based on noble metals, such
as platinum, palladium, rhodium or ruthenium,
5 supported on an inorganic support, such as silica,
titanium oxide, zirconium oxide, alumina, sodium
silicate or silicalite; or catalysts based on
copper, manganese, cobalt, nickel or iron,
optionally in the presence of at least one noble
10 metal, such as platinum, palladium, rhodium or
ruthenium, in the form of mixed oxides or of
alloys optionally supported on an inorganic
support, such as silica, titanium oxide, zirconium
oxide, alumina, sodium silicate or silicalite.
15
6. The process as claimed in any one of the preceding
claims, characterized in that the temperature of
the fluidized bed is lower than the temperature
corresponding to the start of the reaction for the
20 oxidation of the hydrocarbon and/or hydrocarbon
derivative.
7. The process as claimed in any one of the preceding
claims, characterized in that the space velocity
25 SV, expressed as hourly flow rate by volume of gas
mixture with respect to the volume of catalyst, is
between 1000 h- and 30 000 h 1.
8. A process for the production of a hydrocarbon
30 derivative comprising at least one stage of
selective oxidation of the carbon monoxide present
in a gas mixture comprising at least one
hydrocarbon or one hydrocarbon derivative using a
fluidized bed comprising a solid catalyst capable
35 of oxidizing carbon monoxide to carbon dioxide at
a chosen temperature.
WO 2011 / 124824 PCT /FR2011/050692
- 40 -
9. The process as claimed in claim 8, comprising at
least the following stages:
a) at least one hydrocarbon and oxygen or an
5 oxygen-comprising gas are brought into contact
with an appropriate catalyst, resulting in a gas
mixture comprising at least one hydrocarbon
derivative, unconverted hydrocarbon, oxygen and
carbon monoxide,
10
b) the hydrocarbon derivative is separated or
extracted from the reaction stream resulting from
stage a),
15 c) the carbon monoxide present in the gas stream
is then converted to carbon dioxide using a
fluidized bed comprising a solid catalyst capable
of oxidizing carbon monoxide to carbon dioxide at
a chosen temperature, producing a gas stream
20 depleted in carbon monoxide,
d) said stream depleted in carbon monoxide is
recycled to reaction stage a).
25 10. The process as claimed in claim 8 or 9,
characterized in that all or part of the stream
entering or exiting from stage c) is subjected to
a separation in a selective permeation unit,
resulting in the removal of at least a portion of
30 the carbon dioxide present in said stream.
11. The process as claimed in claim 9 or 10,
characterized in that stage a) is carried out in
the presence of a thermal ballast which is inert
35 under the conditions of the reaction carried out.
12. The process as claimed in any one of claims 8 to
O 2011/124824 PCT/F-112011/050692
- 41 --
11, characterized in that the hydrocarbon is
propylene and the hydrocarbon derivative is
acrylic acid,
5 13. The process as claimed in claim 11 or 12,
characterized in that propane is used as thermal
ballast.