EP1448749B1 - Method for converting synthetic gas in series-connected reactors - Google Patents

Description

PRIOR ART:

The production of Fischer-Tropsch liquid fuels offers significant prospects for the exploitation of gas deposits far from major markets. These developments are conditioned by the need to reduce costs and especially investment costs to improve the profitability of this sector.

One way to achieve this goal is to play on a scale factor to reduce investment costs per tonne of liquid product obtained.

The use of the catalyst used to promote the synthesis reaction in the form of suspension in the liquid phase ("slurry") makes it possible to produce very large units of unit size and to reach very high production levels, for example 10,000 barrels per day using a single triphasic reactor.

Such triphasic reactors comprising a catalyst in suspension in a solvent generally inert in the reaction. They are usually called slurry reactors. Among the various types of slurry reactor, there are notably well-agitated autoclave-type reactors, or even bubble-column type reactors which operate in variable hydrodynamic conditions ranging from the perfectly stirred reactor to the reactor operated in piston mode without dispersion, this for both the gas phase and the liquid phase.

Recently, such types of reactors have been considered for Fischer-Tropsch synthesis, rather than conventional fixed bed reactors which have the drawback of not so easily evacuating the heat generated by the reaction.

Thus patents US 5,961,933 and US 6,060,524 describe a method and apparatus for operating a bubble-column type slurry reactor for Fischer-Tropsch synthesis. In these patents, the slurry reactor comprises an internal or external recirculation system of the liquid, which makes it possible to achieve higher productivities for each Fischer-Tropsch reactor.

The patent application WO 01/00595 describes a process for synthesizing hydrocarbons from synthesis gas in a three-phase reactor, preferably of the bubble column type, and wherein the hydrodynamic conditions of the liquid phase are such that the number of Peclet of the liquid phase is greater than at 0 and lower than 10. Furthermore, the superficial gas velocity is preferably less than 35 cm.s-1.

The patent EP-B-450 860 discloses a method for optimally operating a three-phase reactor of the bubble column type. This patent seeks to optimize the operation of a single reactor of this type. it is stated that the performance depends essentially on the dispersion of the gas phase (number of Peclet for the gas phase) and the suspension maintenance of the catalyst in the liquid phase. In particular, the number of Peclet for the gas phase must imperatively be greater than 0.2. Thus, this patent recommends not to use a substantially perfectly stirred reactor with respect to the gas phase (number of gas Peclet close to 0), because this type of reactor leads to insufficient performance levels.

Thus, such a method encounters certain limitations, particularly related to axial mixing phenomena. In order to promote the gas-liquid and solid-liquid mass transfer and the heat transfer, it is advantageous to stir the liquid and gaseous phases in the presence, which increases the axial mixing. In addition, for large reactor diameters, for example from 8 to 11 m, significant internal recirculation movements occur which result in a very large mixing of the liquid phase. These phenomena are favorable in terms of gas-liquid mass transfer and / or liquid-solid and heat transfer, but also a very strong mixture may be unfavorable for the degree of progress of the reaction.

Requirement EP 0 823 470 A1 describes a Fischer-Tropsch process using a piston flow for both the suspension and the gas phase. The reactors are therefore not perfectly agitated, or substantially perfectly agitated.

The patent US6,156,809 describes a Fischer-Tropsch process with several reactors in series. It is not mentioned that these reactors are substantially perfectly agitated. A suspension is taken from each reactor, for separation of a portion of the liquid product, and a concentric suspension is reintroduced into each reactor.

The patent GB-615,381 describes a synthesis gas conversion process in a multi-stage reactor, supplemented with fresh catalyst at the top stage, but it is a two-phase gas / solid reactor.

None of these documents describe the use of highly mixed reactors with increased material transfer, operating with a suspension (three-phase regime).

The method according to the invention aims to overcome these problems by combining at least two triphasic reactors, preferably at least three triphasic reactors. It has indeed been observed that the use of highly mixed reactors in series makes it possible to obtain a correct progress of the reaction, while promoting the evacuation of calories. This sequence makes it possible to achieve high productivities in the desired products, that is to say essentially paraffins having essentially a carbon number greater than 5, preferably greater than 10, while limiting the formation of light products (C1-C4 hydrocarbons).

DESCRIPTION OF THE INVENTION

The invention relates to a process for the synthesis of hydrocarbons preferably having at least 2 carbon atoms in their molecule and more preferably at least 5 carbon atoms in their molecule by contacting a gas containing essentially monoxide of carbon and hydrogen and in a reaction zone containing a suspension of solid particles in a liquid, which comprises solid particles of catalyst of the reaction. Said catalytic suspension is also called slurry. The process according to the invention is therefore carried out in a three-phase reactor. Preferably, the process according to the invention will be carried out in a three-phase reactor of the bubble column type.

The process according to the invention according to claim 1 is a process for converting a synthesis gas into liquid hydrocarbons used in at least two reactors in series, preferably at least three reactors in series containing at least one catalyst in series. suspension in a liquid phase, wherein said reactors are perfectly mixed, the last reactor is at least partly supplied with at least a portion of at least one of the gaseous fractions collected at the outlet of at least one of said reactors, and the product mixture in liquid phase and catalyst leaving at least one reactor is at least partly separated so as to obtain a substantially catalyst-free liquid product and a catalyst-enriched liquid fraction (catalyst-enriched catalyst suspension, or catalytic suspension concentrated), which is recycled to the last reactor. More precisely, the invention proposes a method according to claim 1. The method according to the invention can in particular be implemented in an installation according to the figure 8 . However, it is described below, and in Figures 1 to 7 technical elements useful for understanding the claimed invention, these elements may belong to the state of the art.

Each of the reactors used is a reactor of the bubble column type, with contacting of the gas with a liquid / solid mixture very divided ("slurry" reactor or "slurry bubble column" according to the English terminology)

The catalysts used can be of various kinds and usually contain at least one metal preferably selected from the metals of groups 5 to 11 of the new periodic table of elements.

The catalyst may contain at least one activating agent (also called promoter) preferably chosen from the elements of groups 1 to 7 of the new periodic classification. These promoters can be used alone or in combination.

The support is generally a porous material and often a porous inorganic refractory oxide. By way of example, this support may be chosen from the group formed by alumina, silica, titanium oxide, zirconia, rare earths or mixtures of at least two of these porous mineral oxides.

Typically the suspension may contain from 10 to 65% by weight of catalyst. The catalyst particles have a mean diameter most often between about 10 and about 100 microns. Finer particles may be produced by attrition, ie by fragmentation of the initial catalyst particles.

In the process according to the invention, each of the reactors is strongly mixed and approaches the conditions of perfect mixing. The reactors according to the invention are therefore defined as being substantially perfectly stirred and the number of Peclet can be advantageously used as a criterion for measuring the degree of agitation of said reactors.

Since the reaction takes place in the liquid phase, controlling the hydrodynamics of this phase is crucial. For each reactor, the piston-dispersion model can be applied to the liquid phase because it is well suited to continuous phases. The number of Peclet bound to this model is Pe liq = VI * H / D _ax where VI is the velocity of the liquid in the reactor, H is the expansion height of the catalytic bed and D _{ax is} the axial dispersion coefficient. It should preferably be less than 10, and more preferably less than 8. Such a model is less well suited to the representation of mixing phenomena in the gas phase. However, if it is nevertheless used to interpret a tracer experiment, by determining a Peclet number, for example from the variance of the output concentration profile, it appears that it is possible to reach preference values. less than 0.2, preferably less than 0.18, very preferably less than 0.15 and even more preferably less than 0.1 or even less than 0.05 and in some cases less than or equal to 0, 03.

These conditions are more easily met in the case of a reactor of very large diameter, for example greater than 6m. However it is also possible to reach such conditions in the case of a reactor of smaller diameter by adjusting the hydrodynamic conditions to promote stirring, and therefore mass transfers gas-liquid and liquid-solid. This agitation can be obtained by any means known to those skilled in the art, and in particular for example by generating recirculation movements of the liquid phase by means of internal structures to the reactors or external recirculation means such as loops recirculation.

The mixing effect in the gaseous phase will be increased if said gaseous phase is finely dispersed, in gas bubbles with a diameter not exceeding, for example, a few millimeters. Such a condition is also favorable to the kinetics of reaction.

In order to promote the progress of the reaction, in the process according to the invention, reactors in series, at least two, but preferably at least three, are used. This further allows and this is another object of the present invention, to stagger the injection of synthesis gas. In this way it is possible to optimize the configuration of the reactors in series. In particular, when high rail capacities are targeted for scale effect, the maximum diameter of a reactor is generally limited for reasons of road construction and transport. This diameter may be for example 11 m. In this case, to maximize the production capacity, it is advantageous to use reactors of the same diameter and this can be achieved by adjusting the amount of synthesis gas sent to each of the reactors.

Each of the reactors is operated at a temperature preferably between 180 ° C and 370 ° C, preferably between 180 ° C and 320 ° C, more preferably between 200 ° C and 250 ° C, and at a pressure preferably between 1 and 5 MPa (megapascal), preferably between 1 and 3 MPa.

In summary, the process according to the invention is a process for converting a synthesis gas into liquid hydrocarbons used in at least two reactors in series containing at least one catalyst in suspension in a liquid phase, in which said reactors are substantially fully mixed, the last reactor is at least partly supplied with at least a portion of at least one of the gaseous fractions collected at the outlet of at least one of said reactors, and the mixture of liquid phase product and catalyst leaving at least one reactor is at least partially separated so as to obtain a substantially catalyst-free liquid product and a catalyst-enriched liquid fraction, which is recycled to the last reactor. The process according to the invention preferably comprises at least 3 reactors in series.

In the process according to the invention, the number of liquid Peclet is preferably less than 8, and independently the number of Peclet gas is preferably less than 0.2 and more preferably less than 0.1.

According to a preferred mode of operation of the process according to the invention, at the outlet of each reactor, the gaseous phase is separated from the liquid phase containing the catalyst in suspension. More preferably, the gaseous fractions leaving the first reactors are combined, treated and sent to the inlet of the last reactor and very preferably, the gaseous fraction leaving the last reactor is recycled to the input of the production step. of synthesis gas.

According to a preferred mode of operation of the process according to the invention, the introduction of synthesis gas is distributed to the inlet of the reactors in series so that all the reactors are of identical size.

The catalyst of the process according to the invention is preferably formed of a porous mineral support and at least one metal deposited on this support. The catalyst is preferably suspended in the liquid phase in the form of particles with a diameter preferably of less than 200 microns.

Several possible embodiments of the invention are described below as well as other embodiments not covered by the claims, but useful for understanding the invention. In the figures presented, the references of the same flow or equipment are identical.

EXAMPLE 1

Several embodiments of features or technical variants are possible, one of these modes is presented on the figure 1 .

In this example of a process arrangement, 3 reactors in series are used. The synthesis gas arrives via line 100. It is sent to the first reactor R1, in which it is dispersed within the liquid phase formed by the reaction products which are recycled. At the outlet of this first reactor R1, the mixture of liquid product formed containing the catalyst in suspension (catalytic suspension) and the unreacted gas is discharged through line 101, in the form of a dispersed phase. Through line 102 a second supply of synthesis gas is introduced and the resulting mixture is sent through line 103 to the second reactor R2. At the outlet of this second reactor R2, the liquid product mixture containing the suspension catalyst and the unreacted gas in the form of a dispersed phase are discharged through line 104. Via the conduit 106 is introduced a third supply of synthesis gas and the resulting mixture is sent through the conduit 107 to the third reactor R3. At the outlet of this third reactor R3, the mixture of liquid product containing the suspended catalyst and the unreacted gas in the form of a dispersed phase is discharged through line 108. The gas phase is separated from the liquid phase in the SL separator. This gaseous phase is discharged through line 111, treated and recycled. The liquid phase containing the catalyst in suspension (catalytic suspension) is sent to the separation and filtration system SC. The liquid phase separated from the catalyst is discharged through line 110 while the concentrated liquid phase catalyst (concentrated catalytic suspension) is recycled via line 109 to the first reactor R1.

EXAMPLE 2

In synthesis gas conversion processes not in accordance with the claims, intermediate separations may optionally be carried out. In particular, it is possible to separate the residual gaseous fraction at the outlet of each reactor, as shown in the diagram of FIG. figure 2 .

The residual gaseous fractions are separated at the outlet of each of the reactors by means of the separators SL1, SL2 and SL3.

This avoids sending the inert gases and the water contained in the residual gaseous fractions leaving a reactor to the next reactor. The separators SL1, SL2, SL3 operate for example by decantation, providing a residence time in the separation tank sufficient. The gaseous fractions thus collected by the conduits 111, 112 and 113 are combined, treated and recycled.

The gaseous fractions collected by the conduits 111, 112 and 113 contain water, carbon dioxide, light hydrocarbons and a mixture of carbon monoxide and hydrogen. It is advantageous to send the oxide mixture of carbon and hydrogen collected at the outlet of a reactor to the next reactor (not shown).

The other flows or equipment are identical to those of the figure 1 .

EXAMPLE 3

In the case of the example of an installation not in accordance with the claims, in which the catalytic catalyst-enriched suspension is not recycled to the last reactor (R ₃ ) shown in FIG. figure 3 the gaseous fractions collected by the conduits 112 and 113 at the outlet of the reactors R1 and R2 are combined and treated. The gaseous mixture is first cooled in the exchanger-condenser C1, so as to condense the water. A mixture of three phases is thus obtained, which are separated in the separator S4: an aqueous phase which is discharged through line 114, a liquid hydrocarbon phase which is discharged through line 115, and a gaseous phase which is discharged through the conduit. 116. The gaseous phase is sent to a treatment section T1, so as to separate at least part of the carbon dioxide that it contains. The gaseous fraction rich in carbon dioxide, which is thus separated, is discharged via line 117. The treatment section T1 can use the various known methods for separating carbon dioxide. For example, a washing method with a solvent, such as for example an amine, or a physical solvent such as refrigerated methanol, propylene carbonate or tetraethylene glycol dimethyl ether (DMETEG) can be used. Any other method based on, for example, adsorption separation or selective membrane separation may also be used. The gaseous mixture obtained, which is discharged from the treatment unit T1 through line 106, is enriched with carbon monoxide and hydrogen. It still contains light hydrocarbons including methane. It is sent to the entrance of the last reactor R3. It may be optionally mixed with an additional mixture of carbon monoxide and hydrogen from the synthesis gas production section (not shown). The light hydrocarbons that arrive via line 106 and which are not converted into reactor R3 are discharged through line 111 and can be recycled to the inlet of the synthesis gas production section.

EXAMPLE 4

On the figure 4 Another example of a non-conforming arrangement is presented in which the catalytic suspension from a reactor is not sent directly to another reactor.

The synthesis gas is sent to the first reactor R1 via line 100. At the outlet of reactor R1, the gaseous phase and the liquid phase are separated in separator SL1. The gas phase leaving the separator SL1 is cooled in the exchanger C1. This refrigeration leads to the condensation of an aqueous phase and the evacuation of this condensed phase through the conduit 210, moreover a condensed phase of light hydrocarbons is evacuated via the conduit 211. The resulting gas phase is discharged through the conduit. 113 and sent to the reactor R2, being mixed at the inlet of the reactor R2 with an addition of synthesis gas arriving via the conduit 102. At the outlet of the reactor R2, the gas phase and the liquid phase are separated in the separator SL2 . The gas phase leaving the separator SL2 is cooled in the exchanger C2. This refrigeration leads to the condensation of an aqueous phase and the evacuation of this condensed phase through the conduit 212 and also a condensed phase of light hydrocarbons which is discharged through the conduit 213. The resulting gas phase is evacuated by the conduit 112 and sent to the reactor R3, with a supplement of synthesis gas arriving via the conduit 106. At the outlet of the reactor R3, the gaseous phase and the liquid phase are separated in the separator SL3. The gas phase leaving the separator SL3 is cooled in the exchanger C3. This refrigeration leads to the condensation of an aqueous phase and the evacuation of this condensed phase through the conduit 213; in addition, a condensed phase of light hydrocarbons is evacuated via line 214.

The liquid products leaving the separators SL1, SL2 and SL3 via the ducts 200, 201 and 202, containing the catalyst in suspension (catalytic suspensions), are mixed in the separator SC, in which the liquid products discharged via the duct 110 are separated. a concentrated liquid phase catalyst (concentrated catalyst suspension), which is recycled to the reactors R1, R2 and R3.

On the diagram of the figure 4 the separators SL1, SL2 and SL3 are shown as distinct from the reactors R1, R2 and R3. The gas phase leaving each reactor could, alternatively, be separated from the liquid phase containing the catalyst in suspension in the reactor itself, the liquid phase containing the catalyst can then be removed under control level.

EXAMPLE 5

This example describes a usable mode of circulation of the catalyst between the various reactors. The figure 5 presents the corresponding diagram.

Each reactor being highly mixed, the catalyst introduced at the base of each reactor is distributed homogeneously throughout the liquid phase occupying the reactor. In the exemplary embodiment schematized on the figure 5 the unconverted gaseous fraction disengages at the top of each reactor and the liquid phase containing the catalyst in suspension (catalytic suspension) flows overflow and flows to the base of the next reactor by simple gravity. The transfer lines passing from one reactor to the next reactor must be designed to have the smoothest possible slope. The liquid phase collected at the outlet of the last reactor is at least partially separated from the catalyst it contains and filtered. It is then discharged through line 110. The catalyst which remains in suspension in a residual liquid phase (concentrated catalytic suspension) is recycled with this liquid phase to the first reactor by the line shown in dashed lines.

Such a circulation mode can also be implemented in cases where separating devices and in particular disengaging the gas phase are used at the outlet of each of the reactors as illustrated in Examples 2, 3 and 4. .

It is also possible to perform at the outlet of each of the reactors a separation between the liquid phase produced and a liquid phase concentrated catalyst which is returned to the reactor. Instead of a single separation device SC, there will be for example in such a case as many separation devices as reactors.

The Figures 6 and 7 present two circulation arrangement schemes for use in the process according to the invention or in other methods of the state of the art. These reactors comprise an internal exchanger, for example consisting of preferably tubular cooling beams. These arrangement schemes are usable according to the invention, but the invention is not related to this or these uses.

These reactors have a supply and an outlet, the water entering through line 1 and the generated steam exiting through line 2. A system for dispersing the charge 4 is also disposed inside the reactor. It can be a distributor plate of the gaseous feedstock (synthesis gas) fed by line 3. The supply of liquid comprising the catalyst in suspension can optionally be carried out by the same line, the gas / liquid / solid mixture being produced upstream, like this is the case on Figures 6 and 7 . It is also possible to use separate power supplies, only the gas feeding the dispersion system 4. In the figure 7 internal recirculation is favored by the design of the reactor.

EXAMPLE 6

The figure 8 more particularly represents a mode of arrangement of reactors according to the invention of claim 1 with particular circulation of the catalyst. As in Example 3, the plant comprises two (first) reactors R1, R2 operating in parallel with synthesis gas supplied by lines 100 and 102, and a reactor R3 operating in series with R1, R2, using the gas non-converted residual synthesis from reactors R1 and R2 by lines 101 and 104. This residual synthesis gas, or first stage is (advantageously) treated in unit S1, to substantially eliminate water, and possibly the carbon dioxide before supplying the reactor R3 via line 112. The section S1 can thus correspond to the equipment C1 and S4 of the figure 3 , possibly with the addition of the processing section T1 represented in this same figure. The particular layout of the installation of the figure 8 , compared to the installation of the figure 3 relates to the circulation of the catalyst, that is to say the catalytic suspension of at least one solid catalyst in a liquid phase typically composed of products of the reaction. This catalytic suspension circulates at least partly countercurrently between the different reactors, a catalytic suspension stream circulating from the last reactor R3 (last with respect to the circulation of synthesis gas) to a first reactor R2 via line 221. Another catalytic suspension stream flows from the reactor R2 to the reactor R1 via the line 222. A third catalytic suspension stream flows from the reactor R1 to the reactor R3, via the line 223, the separation section SC, then the line 109 in FIG. which circulates a catalytic suspension (relatively more) concentrated, a stream of purified liquid having been discharged through the line 110.

Alternatively, the reactor R1 is not fed by a catalytic suspension from R2, but by a catalytic suspension from R3, flowing in the beginning of the line 221 and in the dotted line 224, the flow of Catalytic suspension discharged from the reactor R2 is, in this alternative, sent to the section SC via the line 222, then the dashed line 225, then the line 223.

In these two configurations, a suspension stream circulates (directly, that is to say without crossing a separation section) of the (or a) last reactor R3, to a previous or first reactor R1 or R2 (relative to to the circulation of the synthesis gas), and a relatively concentrated suspension stream, from a separation section SC, feeds the last reactor or R3.

An advantage of these configurations results from the fact that the last reactor R3 operates with a concentration of the catalytic suspension greater than that of the preceding reactors or first reactor (s) R1, R2. Indeed, the average concentration (in catalyst) of the catalytic suspension in the reactor R3 is lower than that of the suspension feeding R3 via the line 109, because of the production of liquid products in R3. More generally, a catalytic suspension leaving a reactor is less concentrated than the catalytic suspension feeding the same reactor. The advantage of having a relatively more concentrated catalyst suspension in the last reactor is that this makes it possible to compensate for less favorable operating conditions. On the one hand, the reactor R3 being downstream from R1 and R2, operates under a lower pressure than that (s) of R1, R2. On the other hand, the synthesis gas is depleted in reagents (H2 / CO) in the reactors R1, R2, and enriched in inerts produced by the reaction, in particular methane. Therefore, because of these two phenomena, the partial pressure of reagents (H2 / CO) is significantly lower in the (or a) last reactor R3 than in a previous or first reactor R1, R2. The use of a relatively higher catalytic concentration in the (or a) last reactor makes it possible to compensate for the influence of this lower partial pressure and to be able to maintain a high conversion in the last step. The mass percentage of catalyst may for example be between 20 and 35% by weight, especially between 25 and 32% by weight in the first reactors R1, R2. In the reactor R3, the mass percentage of catalyst may be multiplied by a factor K of between 1.03 and 1.25, in particular between 1.06 and 1.20 and for example between 1.08 and 1.18 with respect to (x) percentage (s) of one (or more) first reactors R1, R2.

Often, in one or the other of the different configurations described in the preceding figures, or according to other configurations not described but obvious for the one skilled in the art, at least one reactor (R1, R2, or R3) is fed (typically directly, that is to say without intermediate fractionation of the type of a liquid separation / catalytic suspension) by a catalytic suspension stream from another reactor.

In general, an installation for implementing the method according to the invention (according to the configuration of the figure 8 ), at least one reactor is fed with a catalytic suspension stream directly from another reactor, and at least one catalytic suspension stream from a reactor is at least partly separated to obtain a substantially free liquid product of catalyst and catalytic suspension catalyst-enriched (concentrated), which is recycled. Each of the reactors is in communication with at least one other reactor, via a suspension stream sent directly to this other reactor or coming directly from this reactor.

According to the invention of claim 1, the catalyst-enriched catalyst suspension is recycled to the last reactor (for example R3), so as to enrich the catalytic suspension of the latter reactor relative to that of the other reactors, by example of one or more reactors (R1, R2).

The method may in particular comprise a first reaction stage carried out in several first reactors operating in parallel, in which the gaseous fractions leaving these first reactors are combined, treated and sent to the inlet of a last reactor. The conversion carried out in the first reactors can be determined so that all the reactors are of identical size.

The number of "first reactors" or "last reactor (s)" may be different, for example between 1 and 8. The number of reaction steps may be between 1 and 5. The reactors R1, R2 , R3 previously described may be replaced by reaction zones, possibly integrated in a smaller number of reactors etc ...

EXAMPLE 7

This example presents a material balance illustrating the possible use of several reactors substantially perfectly mixed in a synthesis gas conversion process, and shows the relevance of this technical option. Through line 100, a flow rate of 713t / h of synthesis gas is obtained, the molar composition of which is as follows: Water : 0.004 Hydrogen : 0.672 CO : 0.311 Methane : 0.013

The process used comprises 3 reactors R1, R2, R3 substantially perfectly mixed and having the Peclet numbers between 0.02 and 0.03.

The reactor R1 operates at a temperature of 236 ° C. At the outlet of the reactor R1, after separation, 200.66 t / h of liquid products are collected via the conduit, comprising 87% of mole fraction of constituents, the molecule of which comprises at least 10 carbon atoms. After cooling of the gaseous phase, 234 t / h of water (conduit 210), 67 t / h of condensed hydrocarbons (conduit 211) and 347 t / h of synthesis gas are recovered at a pressure of 2.8 MPa , which is sent to the reactor R2 via the conduit 113 being mixed with 327 t / h of synthesis gas arriving via the conduit 102.

At the outlet of the reactor R2, after separation, is collected through the conduit 101, 63 t / h of liquid products. After cooling of the gaseous phase, 212, 224 t / h of water are recovered via line 212, via line 213 76 t / h of condensate and via line 112, 311 t / h of synthesis gas, which is sent to the reactor. R3 being mixed with 293t / h of syngas arriving via line 106.

At the outlet of the reactor R3, 202.58 t / h of liquid products are collected via line 202. After cooling the gas phase, 205 t / h of water, 75 t / h of condensate and 266 t / h of synthesis gas are recovered.

The overall conversion yield is 91%.

It is possible to realize this example with reactors of different size. It is also possible to use identical size reactors, by adapting the temperatures and conversions to liquid products used for reactors R1, R2, R3, associated with the distribution of synthesis gas. Adaptation of conditions for to increase the relative size of a given reactor making it possible to obtain these conditions, can be achieved by increasing the relative flow rate of synthesis gas at the inlet of this reactor, and / or by increasing the conversion in this reactor, and / or by reducing the temperature of this reactor. Preferably one plays only on the first two parameters, the temperature of the three reactors remaining substantially identical. In the preceding example, the conditions mentioned can be obtained with identical size reactors, operating at similar pressures (differing only in the pressure losses), and maintained at the same temperature of 236 ° C.

Claims (10)

1. Process for converting a synthesis gas into liquid hydrocarbons used in at least two reactors that are arranged in series and that contain a catalytic suspension of at least one solid catalyst in suspension in a liquid phase, in which said reactors are essentially perfectly mixed, the last reactor is at least in part fed by at least a portion of at least one of the gaseous fractions that are collected at the outlet of at least one of the other reactors, at least one reactor is fed by a flow of catalytic suspension that is obtained directly from another reactor, at least one flow of catalytic suspension that is obtained from a reactor is at least in part separated so as to obtain a liquid product that is essentially free of catalyst and a catalytic suspension that is enriched in catalyst, which is recycled, in which each of the reactors is linked with at least one other reactor via a suspension flow that is sent directly to this other reactor or that is obtained directly from this reactor, and in which said catalytic suspension that is enriched in catalyst is recycled in last reactor (R3), so as to enrich the catalytic suspension of this last reactor relative to that (those) of other reactors, for example of one or more reactors (R1, R2).
2. Process according to claim 1, comprising a first reaction stage that is carried out in several first reactors that operate in parallel, in which the gaseous fractions that exit from these first reactors are combined, treated and sent to the inlet of a last reactor.
3. Process according to claim 2, in which the conversion that is carried out in the first reactors is determined such that all of the reactors are identical in size.
4. Process according to one of claims 1 or 2, in which the liquid Péclet number is less than 8.
5. Process according to one of claims 1 to 4, in which the gas Péclet number is less than 0.2.
6. Process according to one of claims 1 to 3, in which the gas Péclet number is less than 0.1.
7. Process according to one of claims 1 to 5, in which at the outlet of each reactor, the gaseous phase is separated from the liquid phase that contains the catalyst in suspension.
8. Process according to one of claims 1 to 7, in which the catalyst is formed by a porous mineral substrate and at least one metal that is deposited on this substrate, whereby the catalyst is suspended in the liquid phase in the form of particles with a diameter that is less than 200 microns.
9. Process according to one of claims 1 to 8, in which the distribution of the introduction of synthesis gas at the inlet of the reactors that are arranged in series is determined so as to allow the use of reactors of identical size.
10. Process according to one of claims 1 to 9, in which the gaseous fraction that exits from the last reactor is recycled in a stage for production of synthesis gas for feeding said reactors.

Images