METHOD FOR OLIGOMERISATION IN A REACTOR COMPRISING VARIABLE-DIAMETER ZONES, INCLUDING A STEP OF RECYCLING A PRE-COOLED SOLVENT

Abstract

The present invention relates to a process for oligomerization in a reactor with zones of variable diameters comprising a step of recycling a precooled solvent.

Claims

1. Process for the oligomerization of an olefinic feedstock, comprising: a) a step of oligomerization of the olefinic feedstock, carried out at a temperature of between 30? C. and 200? C. and a pressure of between 0.1 and 10 MPa, in the presence of a homogeneous catalytic oligomerization system and of a solvent, in a reaction section comprising: an oligomerization reactor with zones of variable diameter and comprising a liquid phase, and at least one recirculation loop allowing the cooling of at least a part of a liquid-phase fraction to a temperature T.sub.loop, b) a step of separating a reaction effluent resulting from the oligomerization step a), in a separation section so as to obtain a solvent fraction, c) a step of cooling the solvent fraction resulting from step b) to a temperature below the temperature T.sub.loop to which the liquid-phase fraction is cooled in the recirculation loop(s), d) a step of introducing, into the reaction section of the oligomerization step a), the cooled solvent fraction resulting from step c).

2. Process according to claim 1, in which the reactor with zones of variable diameter comprises n consecutive zones, n being a positive integer between 2 and 10, with: for each of the n zones having a diameter Dn which decreases in the direction of the bottom zone to the top zone of said reactor, a ratio (Dn/Dn?1) of the diameter of the upper zone, denoted Dn, to the diameter of the adjacent lower zone, denoted Dn?1, of less than or equal to 0.9, for a given zone, a ratio of the volume of said zone, denoted Vn, to the total volume of the reaction chamber, denoted Vtot, of between 0.2 and 0.8.

3. Process according to claim 2, in which the n consecutive zones of the reactor with zones of variable diameter are arranged in series along the vertical axis of the reactor so as to define zones in the reaction enclosure having diameters decreasing from the bottom to the top.

4. Process according to claim 1, in which the solvent fraction resulting from step b) is cooled in step c) to a temperature of between 0? ? C. and 150? C.

5. Process according to claim 1, in which the solvent fraction resulting from step b) is cooled in step c) to a temperature at least 40? C. lower relative to the temperature T.sub.loop of the liquid-phase fraction cooled in the recirculation loop(s).

6. Process according to claim 1, in which the cooling of the solvent fraction in step c) is carried out by one or more thermal exchangers, preferably chosen from one or more heat exchangers of process fluid/process fluid type, of air cooler type, of cooling water exchanger type.

7. Process according to claim 1, in which the separation section comprises at least two distillation columns, preferably at least three distillation columns, preferably at least four distillation columns.

8. Process according to claim 1, in which step d) of introducing the cooled solvent fraction is carried out in the reactor and/or in one or more of the recirculation loops.

9. Process according to claim 5, in which step d) of introducing at least a part of the cooled solvent fraction is carried out in a recirculation loop upstream or downstream of a thermal exchanger of the recirculation loop(s), preferably downstream of said thermal exchanger.

10. Process according to claim 1, in which the cooled solvent fraction has a flow rate, as a weight percentage relative to the flow rate of the liquid-phase fraction circulating in the recirculation loop(s), of between 0.05% and 15.0%, preferably between 0.1% and 10.0%.

11. Process according to claim 1, in which the olefinic feedstock comprises olefins having between 2 and 6 carbon atoms, preferably between 2 and 4 carbon atoms.

12. Process according to claim 1, in which the oligomerization step a) comprises at least one of the following substeps: step a1) of introducing the catalytic system, step a2) of bringing into contact with olefinic feedstock, substep a3) of withdrawing a liquid-phase fraction from the oligomerization reactor, substep a4) of cooling at least a part of the liquid-phase fraction withdrawn in step a3) in at least one recirculation loop, to a temperature T.sub.loop, step a5) of introducing the cooled liquid fraction into the reactor.

13. Process according to claim 12, in which the cooling substep a4) is carried out by circulating at least a part of the liquid-phase fraction withdrawn in step a3) through one or more thermal exchangers located in the recirculation loop(s).

14. Process according to claim 13, in which the thermal exchanger(s) used in substep a4) decreases the temperature of the liquid-phase fraction withdrawn in substep a3) by 1.0 to 30.0? C., preferably between 2.0 and 25.0? C.

15. Process according to claim 12, in which the reaction effluent is obtained by dividing the liquid fraction withdrawn in step a3) into two streams.

Description

DESCRIPTION OF THE FIGURES

[0190] FIG. 1 represents a schematic illustration of a plant carrying out an embodiment of the oligomerization process according to the invention. Said plant comprises a two-phase gas/liquid oligomerization reactor A, the zones of variable diameter of which are not represented, a recirculation loop comprising an exchanger B and a pump C, a separation section D, a pump E for circulating the solvent fraction, an exchanger for cooling the solvent fraction (F). In this plant, the stream 2 is a mixture of the stream 1 of fresh ethylene and of ethylene resulting from the separation section. The stream 2 is introduced into the reactor A. The stream 3 is the liquid-phase fraction withdrawn from the reactor and sent to the recirculation loop comprising an exchanger B and a pump C in order to obtain a cooled liquid fraction 4. The stream 6 is the solvent fraction separated in the separation section D which passes through a pump E and is cooled in an exchanger F to give a cooled solvent fraction 7. The fractions 7 and 4 are mixed to give a stream 8 before being introduced into the reactor A. The stream 5 corresponding to a part of the effluent withdrawn from the reactor A is sent to the separation section D. The separation section makes it possible to obtain the stream 6, a stream 9 corresponding to the light reaction products, a stream 10 corresponding to the heavy reaction products and a stream 11 corresponding to a heavy fraction comprising the spent catalyst.

[0191] FIG. 2 illustrates a gas/liquid reactor A with consecutive zones of decreasing diameter according to the invention comprising a lower part comprising a liquid phase, an upper part comprising a gas phase, and a means for introducing a gaseous olefinic feedstock 12 by means of a gaseous distributor 13 into the liquid phase. The upper part comprises a bleeding means 15. A fraction divided into two streams is withdrawn from the bottom of reactor A, a first main stream 3 sent to a heat exchanger B and a pump C in order to obtain a cooled fraction 4. The second stream 5 corresponds to the effluent sent to the separation section. The fractions 7 and 4 are mixed in a stream 8 before being introduced into the reactor A. The catalytic system 14 is introduced into the bottom of the reactor. The zone 1 located at the bottom of the reactor has a larger diameter than the zone located at the top of the reactor. The first bottom zone is characterized by its diameter denoted D1 and its height H1, these two parameters defining the volume, denoted V1, of said zone. Similarly, the second zone located at the top is characterized by its height denoted H2 and its diameter denoted D2, which is less than D1, defining the volume, V2, of the second zone. In this embodiment, the two zones making up the reactor A are formed by cylinders of decreasing diameter.

[0192] FIG. 3 illustrates another embodiment, which differs from that of FIG. 2 in that the second zone located at the top of the reactor A is delimited by an internal 11 placed inside the reactor A.

[0193] FIG. 4 illustrates another embodiment, which differs from that of FIG. 2 in that the reactor A comprises three consecutive zones of decreasing diameter.

[0194] FIGS. 1, 2, 3 and 4 provide schematic illustrations of the particular embodiments of the subject of the present invention, without limiting the scope of said invention.

EXAMPLES

[0195] The examples below illustrate the invention without limiting the scope thereof.

[0196] The examples below describe a process for the oligomerization of ethylene carried out continuously in a two-phase gas/liquid reactor of bubble column type at a pressure of 6.1 MPa and a temperature of 135? C. The catalytic system is introduced into the reactor at a concentration of 1 ppm by weight of chromium, comprises the chromium precursor Cr(2-ethylhexanoate).sub.3, 2,5-dimethylpyrrole at a molar ratio relative to the chromium of 3, 11 molar equivalents of triethylaluminium and 8 molar equivalents of diethylaluminium chloride relative to the chromium, in the presence of o-xylene as additive at a molar ratio of 500 relative to the chromium, and cyclohexane as solvent.

[0197] The amount of cyclohexane, used as solvent 6, introduced into the reactor A is dependent on the amount of ethylene entering the same reactor A (stream 2); the amount of solvent is adjusted so as to have a solvent content of 58% in the reactor.

Example 1 (Comparative)

[0198] Example 1 illustrates an oligomerization process according to the prior art in which the solvent fraction (7) is separated in a downstream separation section and recycled to the reaction section without being cooled by a heat exchanger F and in which the oligomerization process uses a gas-liquid reactor of bubble column type.

[0199] The catalytic system is brought into contact with gaseous ethylene by introduction of said gaseous ethylene into the lower part of said reactor. The reaction effluent is subsequently recovered at the bottom of the reactor.

[0200] The production of hex-1-ene requires the conversion of 14 000 kg/h of ethylene. The flow rate of solvent under the operating conditions adopted is 19 500 kg/h. The temperature of the recycled solvent fraction is 101? C.

[0201] The residence time in the reaction section (reactors+recirculation loop(s)) is 40 minutes.

[0202] Since the oligomerization reaction is exothermic, the heat of the reaction is removed by heat exchangers placed on recirculation loops outside the reactor, having a total surface area of 1650 m.sup.2. The total reaction liquid volume of 41.4 m.sup.3 is distributed between the volume taken by the heat exchangers and their recirculation loop, and the reactor. The total reaction liquid volume is broken down in the following manner: 30.4 m.sup.3 for the heat exchange loops, and 11.0 m.sup.3 for the reactor. The height of liquid in the reactor is then 4.8 metres (denoted m) for a diameter of 1.7 m. The temperature of the mixture 8 of the solvent fraction 7 and of the recirculation fluid 4 is then 120? C. at the reactor inlet. The temperature of the stream 4, corresponding to the stream at the outlet of the exchanger B of the recirculation loop of the reaction section, is then 120.4? C.

[0203] The production of hex-1-ene is 9.32 tonnes/hour, the hex-1-ene selectivity is 93.2 wt %.

Example 2 (According to the Invention)

[0204] The oligomerization process according to the invention is illustrated in FIG. 1 and is performed under the same conditions as in Example 1. Example 2 comprises a step of cooling the solvent fraction resulting from the separation section before introduction into the reaction section, using an exchanger F, as illustrated in FIG. 1 and a reactor A with zones of variable diameters. Said solvent fraction is cooled to a temperature of 40? C.

[0205] The total surface area of the exchangers of the recirculation loops is 1440 m.sup.2. Thus, the process according to the invention makes it possible to reduce the exchange surface area of the heat exchangers by approximately 13% (=100?(1650?1440)/1650) compared to the exchange surface area of Example 1, which represents a saving in the operating cost of the unit.

[0206] In this example, the residence time is kept identical to that of Example 1, namely 40 minutes (min). The total reaction liquid volume is identical to that of Example 1. The step of cooling the solvent fraction therefore makes it possible to reduce the need for exchange in the recirculation loops, which results in a decrease of the surface area of the exchangers.

[0207] Owing to the decrease of the surface area required in the exchangers, the volume of the recirculation loops is reduced by 3% (volume of 29.5 m.sup.3). The liquid volume of the reactor may then be increased by 8% (working volume of 11.9 m.sup.3) which leads to a gain in the height of liquid of 8%.

[0208] Furthermore, the use of a reactor with zones of variable diameters makes it possible to obtain an additional gain of 35% in the height of liquid (total liquid height of 7.1 m). The height of liquid in the lower zone of the reactor is 2.1 m for a diameter of 1.7 m. The height of the upper part of the liquid reactor is then 4.9 m for a diameter of 1.35 m.

[0209] The temperature target of the mixture 8 of the solvent fraction 7 and of the liquid fraction 4 circulating in the recirculation loops was kept identical to Example 1, namely 120? C. The temperature of the cooled liquid fraction 4, corresponding to the outlet of the exchanger B, is then 121.9?C.

[0210] In addition, the 3% reduction in the volume of the recirculation loops and the use of a gas/liquid reactor with zones of variable diameters also makes it possible to increase the volume of liquid in the reactor by 9% and the liquid height by 47%, thereby making it possible to maximize the ethylene saturation in the liquid contained in the reactor.

[0211] The production of hex-1-ene is 9.32 tonnes/hour, the hex-1-ene selectivity is 93.2 wt %.

Example 3 (According to the Invention)

[0212] The oligomerization process according to the invention is performed under the same conditions as in Example 2.

[0213] The step of cooling the solvent fraction makes it possible to reduce the need for exchange in the recirculation loops, which results in a decrease of the surface area of the exchangers. The solvent fraction is cooled to a temperature of 40? C. The total surface area of the exchangers of the exchange loop is then limited to 1440 m.sup.2. Thus, the process according to the invention makes it possible to reduce the exchange surface area of the heat exchangers by approximately 13% (=100?(1650?1440)/1650) compared to the exchange surface area of Example 1, which represents a saving in the operating cost of the unit.

[0214] The liquid height in the reactor is kept identical to Example 1. The reactor then comprises two zones, a lower zone 1.6 m high for a diameter of 1.7 m and an upper zone 3.1 m high for a diameter of 1.35 m. Owing to the decrease of the surface area required in the exchangers, the volume of the recirculation loops is reduced by 3% (volume of 29.5 m.sup.3). The liquid volume of the reactor with zones of variable diameters is 8.4 m.sup.3. The total reaction liquid volume of the reaction section is therefore then 37.8 m.sup.3, namely an 8% saving compared to Example 1.

[0215] The temperature target of the mixture 8 of the solvent 7 and of the recirculation fluid 4 was kept identical to Example 1, namely 120? C. The temperature of the stream 4, corresponding to the outlet of the exchanger B, is then 121.9? C.

[0216] In this example, the residence time is 36.6 minutes. The process according to the invention allows a reduction in the residence time, which leads to a gain in selectivity of 0.1%. This gain enables, for a constant hex-1-ene production, a decrease in the ethylene consumption of 0.1% and therefore a saving in the operating cost of the unit. The decrease in the residence time also enables a saving in the chromium concentration needed for achieving this performance of 8% (corresponding to 4 ppm of chromium), namely a saving in the catalyst consumption and therefore a saving in the operating cost of the unit.

[0217] The production of hex-1-ene is 9.32 tonnes/hour, the hex-1-ene selectivity is 93.3 wt %.

[0218] The table below summarizes the results obtained for Examples 1 to 3.

TABLE-US-00001 Example 1 Example 2 Example 3 Temperature of the recycled ? C. 101 40 40 solvent fraction Temperature of the cooled ? C. 120.4 121.9 121.9 liquid fraction Temperature of solvent ? C. 120 120 120 fraction + cooled liquid fraction Exchange surface area of the m.sup.2 1650 1440 1440 recirculation loops Volume of the recirculation m.sup.3 30.4 29.5 29.5 loops Reactor liquid volume m.sup.3 11.0 11.9 8.4 Reaction section volume m.sup.3 41.3 41.3 37.8 Diameter of the lower zone of m 1.7 1.7 1.7 the Reactor Liquid height in the lower m 4.8 2.1 1.6 zone of the Reactor Diameter of the upper zone m 1.35 1.35 of the Reactor Liquid height in the upper 4.9 3.1 zone of the Reactor Total liquid height 4.8 7.1 4.8 Residence time min 40 40 36.6 Selectivity % 93.2 93.2 93.3