Method for producing butadiene from ethanol with optimised in situ regeneration of the catalyst of the second reaction step

Abstract

The present invention relates to a process for producing butadiene from ethanol, in two reaction steps, comprising a step a) of converting ethanol into acetaldehyde and a step b) of conversion into butadiene, said step b) simultaneously implementing a reaction step and a regeneration step in (n+n/2) fixed-bed reactors, n being equal to 4 or a multiple thereof, comprising a catalyst, said regeneration step comprising four successive regeneration phases, said step b) also implementing three regeneration loops.

Claims

1. A process for producing butadiene from ethanol, comprising at least the following steps: a) a step of converting ethanol into an effluent comprising ethanol and acetaldehyde, comprising at least one reaction section (A) fed with a stream comprising ethanol and operated in the presence of a catalyst (Ca) at a temperature of between 200 and 500° C., and at a pressure of between 0.1 and 1.0 MPa; b) a butadiene conversion step comprising at least one reaction-regeneration section in which simultaneously implemented are a reaction step and a regeneration step in (n+n/2) fixed-bed reactors, n being an integer equal to 4 or a multiple thereof, said (n+n/2) fixed-bed reactors each comprising at least one fixed bed of a catalyst (Cb), said (n+n/2) fixed-bed reactors functioning in parallel and in sequence so that said reaction step starts in each of said reactors with a time shift equal to a quarter of the catalytic cycle time of said catalyst (Cb), said reaction-regeneration section comprising three regeneration loops, and wherein, at a given moment: b1) said reaction step is operated in n of said fixed-bed reactors, n being an integer equal to 4 or a multiple thereof, fed at least with a fraction of said effluent obtained from step a), at a temperature of 300 to 400° C., at a pressure of between 0.1 and 1.0 MPa, and for a time equal to the catalytic cycle time of said catalyst (Cb), to produce a reaction effluent, and b2) said regeneration step is operated in n/2 of said fixed-bed reactors for a total time equal to half of the catalytic cycle time of said catalyst (Cb), and comprises the following four successive phases: i. a stripping phase operated at a temperature of 300 to 400° C., under a stream of inert gas, said phase i) starting on conclusion of the reaction step b1); and then ii. a first combustion phase operated on conclusion of phase i) under a gas stream comprising said inert gas and oxygen in a content of less than or equal to 1 vol % relative to the total volume of said gas stream, at a temperature of 300 to 450° C.; and then iii. a second combustion phase operated on conclusion of the first combustion phase ii) under a gas stream comprising said inert gas and oxygen in a content of greater than or equal to 2 vol % relative to the total volume of said gas stream, at a temperature of 390 to 550° C.; and then iv. a final stripping phase operated at a temperature of 300° C. to 550° C., under a stream of said inert gas; said three regeneration loops of said reaction-regeneration section comprising a regeneration loop for the inert gas of the stripping phases i) and iv), a regeneration loop for the gas stream of the first combustion phase ii) and a regeneration loop for the gas stream of the second combustion phase iii).

2. The process as claimed in claim 1, wherein the reaction section of step a) is operated at a temperature of 250° C. to 300° C., and at a pressure of 0.1 to 0.5 MPa.

3. The process as claimed in claim 1, wherein said fixed-bed reactors in said reaction step b1) are further fed with a supply of ethanol and/or a supply of acetaldehyde, and wherein a mole ratio of ethanol to acetaldehyde in said reaction step b1) is 1 to 5.

4. The process as claimed in claim 1, wherein the integer n is equal to 4.

5. The process as claimed in claim 1, wherein said reaction step b1) is operated at a temperature of 300 to 360° C.

6. The process as claimed in claim 1, wherein said reaction step b1) is operated at a pressure of 0.2 to 0.4 MPa.

7. The process as claimed in claim 1, wherein the catalytic cycle time of said catalyst (Cb) for the butadiene conversion step b) is greater than or equal to 1 day, and less than or equal to 20 days.

8. The process as claimed in claim 1, wherein the inert gas of the regeneration step b2) is nitrogen, carbon dioxide (CO.sub.2) or a mixture thereof.

9. The process as claimed in claim 1, wherein said stripping phase i) is operated at a temperature of 330 to 370° C.

10. The process as claimed in claim 1, wherein a flow rate of inert gas of said stripping phase i) is 0.5 to 1.5 Nm.sup.3/h/kg of catalyst.

11. The process as claimed in claim 1, wherein the oxygen content of said first combustion phase ii) is 0.1 to 1 vol %.

12. The process as claimed in claim 1, wherein said first combustion phase ii) is operated at a temperature of 330 to 430° C.

13. The process as claimed in claim 1, wherein said first combustion phase ii) is operated at a flow rate of gas stream of 1.7 to 2.5 Nm.sup.3/h/kg of catalyst.

14. The process as claimed in claim 1, wherein the oxygen content of said second combustion phase iii) is 2 to 20 vol %.

15. The process as claimed in claim 1, wherein said second combustion phase iii) is operated at a constant temperature of 390 to 430° C. followed by a temperature increase ramp of 10 to 30° C/h and then a phase at a constant temperature of 460 to 510° C.

16. The process as claimed in claim 1, wherein said second combustion phase iii) is operated at a flow rate of gas stream of 1.2 to 1.8 Nm.sup.3/h/kg of catalyst.

17. The process as claimed in claim 1, wherein said final stripping phase iv) is operated on a temperature decrease ramp of 50 to 150° C/h followed by a phase at a constant temperature of 300 to 400° C.

18. The process as claimed in claim 1, wherein said final stripping phase iv) is operated under a stream of said inert gas, at a flow rate of 0.5 to 1.5 Nm.sup.3/h/kg of catalyst.

19. The process as claimed in claim 1, wherein the inert gas of the regeneration step b2) is nitrogen.

20. The process as claimed in claim 1, wherein said first combustion phase ii) is operated at a constant temperature of 330 to 370° C. followed by a temperature increase ramp of 10 to 30° C/h and then a phase at a constant temperature of 390 to 430° C.

Description

BRIEF DESCRIPTION OF DRAWINGS

(1) FIG. 1 shows, schematically and in a nonlimiting manner, an operating diagram of fixed-bed reactors in a system containing six fixed-bed reactors, during step b) of conversion into butadiene, a line corresponding to the functioning of a reactor, the steps in dark gray corresponding to the reaction steps b1), the steps in light gray corresponding to the various phases of the regeneration step b2): the phases i) corresponding to the initial stripping phases i), the phases ii) corresponding to the first combustion phases ii), the phases iii) corresponding to the second combustion phases iii) and the phases iv) corresponding to the final stripping phases iv).

(2) FIG. 2 shows the scheme of the regeneration step b2) in two reactors (R1) and (R2) in non-operational mode (i.e. in a multi-reactor system comprising six fixed-bed reactors), comprising three regeneration loops, a regeneration loop (L1) for the inert gas, a regeneration loop (L2) for the gas stream with a low oxygen content and a regeneration loop (L3) for the gas stream with a high oxygen content. Each regeneration loop (L1), (L2) and (L3) comprises a purge, (P1), (P2) and (P3), respectively. The regeneration loop (L1) for the inert gas (or stripping loop) also comprises a supply (AA1) of inert gas. The regeneration loop (L2) for the gas stream with a low oxygen content (or low O.sub.2 loop) comprises a supply (AA2) of inert gas and a supply of dioxygen (AO2) (for example in the form of an air supply) and the regeneration loop (L3) for the gas stream with a high oxygen content (or high O.sub.2 loop) also comprises a supply (AA3) of inert gas and a supply of dioxygen (AO3) (for example in the form of an air supply). Before the reactors (R1) and (R2), in a regeneration phase, the gas streams can be heated in ovens (F1) and (F2). The lines (3) (i.e. (3-a), (3-b), (3-c), (3-d)) correspond to safety offtakes. The lines (1) (i.e. (1-a), (1-b)) represent feed lines for ethanol/acetaldehyde feedstock for the reactors (R1) and (R2) when the latter are operational. The lines (2) represent the lines at the outlet of reactors (R1) and (R2) via which the reaction effluent produced is recovered to feed a separation section, for example.

EXAMPLES

(3) The following examples are based on simulations of processes incorporating thermodynamic data set up on experimental points.

(4) In each of the following examples, the process described is incorporated into a more global process such as the one described in French patent FR 3 026 100. The ethanol feedstock for the global process is obtained from a renewable source and comprises more than 93% by weight of ethanol. The flow rate of feedstock feeding the global process is adjusted so as to obtain an annual production of 150 kt/year of a butadiene having a purity of between 99.5% and 100% by weight (compatible with the current use of the product), with an annual duration of functioning of the process of 8000 hours.

(5) In the following examples, the term “variation(s) of composition” means the mean amplitude(s) of variation of the weight contents of the compounds of the reaction effluent, over the duration of functioning.

Example 1 (in Accordance)

(6) In this example, the second reaction unit comprises 6 (4+2) fixed-bed reactors and is operated at a weight hourly space velocity (WHSV) of 2 h.sup.−1. The regeneration section comprises 3 regeneration loops.

(7) The conversion of ethanol into acetaldehyde is carried out in a multitubular reactor comprising a catalyst based on copper oxide on a silica support, at 275° C. and 0.26 MPa. The ethanol/acetaldehyde effluent, which is separated from the hydrogen stream at the reactor outlet, is then sent to the second reaction unit.

(8) The conversion into butadiene is carried out in 6 radial fixed-bed reactors, each containing a fixed bed of a catalyst based on tantalum oxide on a silica-based matrix, at 380° C. and 0.2 MPa and at a weight hourly space velocity (WHSV) of 2 h.sup.−1. Supplies of ethanol and of acetaldehyde, obtained from purification and separation sections of the unconverted ethanol and acetaldehyde streams downstream of the reaction sections, are added to the ethanol/acetaldehyde effluent entering the second reaction unit, so that the total flow rate of ethanol/acetaldehyde mixture, at the inlet of the second reaction unit, is equal to 92.4 t/h. Under these conditions, the ethanol conversion is 34 mol % in reaction step 2 and the catalytic cycle of the catalyst based on tantalum oxide is 10 days. The six reactors function in parallel and in sequence as shown schematically in FIG. 1: each one starts the reaction phase with an offset of 5 days; at a given moment, 4 reactors are operational and 2 reactors are in the regeneration phase. The catalyst is regenerated in situ in each reactor according to the protocol presented in Table 1. The regeneration section comprises three regeneration loops: a nitrogen loop for providing the nitrogen stream required for stripping the catalyst, a loop for the gas stream with a low content of O.sub.2 making it possible to provide a gas stream comprising nitrogen and 0.5% by volume of O.sub.2 and a loop for the gas stream with a low content of O.sub.2 making it possible to provide a gas stream comprising nitrogen and 6.0% by volume of O.sub.2.

(9) TABLE-US-00001 TABLE 1 Regeneration protocol Initial T, Final T, Ramp, vol % Stage time, Period ° C. ° C. ° C./h O.sub.2 h I 350 350 — 0.0 12 II 350 350 — 0.5 30 III 350 410 20 0.5 3 IV 410 410 — 0.5 16.5 V 410 410 — 6.0 19 VI 410 480 20 6.0 3.5 VII 480 480 — 6.0 20 VIII 480 350 100  0.0 1.5 IX 350 350 — 0.0 12

(10) Table 2 indicates composition variations in the reaction effluent, over the duration of functioning.

(11) TABLE-US-00002 TABLE 2 Composition variations in the reaction effluent Component Variation, wt % Ethanol 0.8 Acetaldehyde 4.7 Butadiene 6.1 Diethyl ether 6.7

(12) It appears that the composition variations of the reaction effluent that are observed at the reaction unit outlet are low. Over the duration of functioning of the process, the butadiene production shows flow rate variations of only 6% by weight.

(13) Table 3 presents the consumptions of utilities for the regeneration of the tantalum oxide based catalyst, for all of the reactors, over the duration of functioning.

(14) TABLE-US-00003 TABLE 3 Consumptions of utilities for the regeneration Nitrogen, Nm.sup.3/h 48.79 Instrument Air, Nm.sup.3/h 317.19 Electricity, kW 2084.25 Boiler water, t/h 2.67 Gas fuel, kW 1923.95

(15) The consumption of utilities for the regeneration step is relatively low. In particular, the nitrogen consumption (48.79 Nm.sup.3/h) over the duration of functioning of the process remains very low in particular relative to the nitrogen consumption of a reactor system that is similar but with once-through regeneration (N.sub.2 consumed=42 400 Nm.sup.3/h).

Example 2 (in Accordance)

(16) In this example, the second reaction unit comprises 6 (4+2) fixed-bed reactors and is operated at a weight hourly space velocity (WHSV) of 1.2 h.sup.−1. The regeneration section comprises three regeneration loops.

(17) The conversion of ethanol into acetaldehyde is carried out in a multitubular reactor comprising a catalyst based on copper oxide on a silica support, at 275° C. and 0.26 MPa. The ethanol/acetaldehyde effluent, which is separated from the hydrogen stream at the reactor outlet, is then sent to the second reaction unit.

(18) The conversion into butadiene is carried out in six radial fixed-bed reactors, each containing a fixed bed of a catalyst based on tantalum oxide on a silica-based matrix, at 380° C. and 0.2 MPa and at a weight hourly space velocity (WHSV) of 1.2 h.sup.−1. Supplies of ethanol and of acetaldehyde, obtained from the sections for purification and separation of the unconverted ethanol and acetaldehyde streams downstream of the reaction sections, are added to the ethanol/acetaldehyde effluent entering the second reaction unit, so that the total flow rate of ethanol/acetaldehyde mixture entering the second reaction unit is equal to 129.7 t/h. Under these conditions, the ethanol conversion is 23 mol % in the reaction step 2 and the catalytic cycle of the catalyst based on tantalum oxide is 10 days. The six reactors function in parallel and sequentially as represented schematically in FIG. 1: each starts the reaction phase with an offset of 5 days; at a given moment, four reactors are operational and two reactors are in the regeneration phase. The catalyst is regenerated in situ in each reactor according to the same protocol as that of Example 1 and presented in Table 1. The regeneration section comprises three regeneration loops: a nitrogen loop for providing the nitrogen stream necessary for stripping the catalyst, a loop for a gas stream with a low O.sub.2 content for providing a gas stream comprising nitrogen and 0.5 vol % and a loop for a gas stream with a low O.sub.2 content for providing a gas stream comprising nitrogen and 6.0 vol % of O.sub.2.

(19) Table 4 indicates variations in composition in the reaction effluent.

(20) TABLE-US-00004 TABLE 4 Variations of compositions in the reaction effluent Component Variation, weight % Ethanol 0.8 Acetaldehyde 4.7 Butadiene 6.1 Diethyl ether 6.7

(21) It appears that the variations in composition of the reaction effluent observed leaving a reaction unit are low, throughout the functioning. Over the duration of functioning of the process, the butadiene production has flow rate variations of only 6% by weight.

(22) Table 5 shows the consumptions of utilities, for the regeneration of the catalyst based on tantalum oxide, of all of the reactors, over the duration of functioning.

(23) TABLE-US-00005 TABLE 5 Consumptions of the utilities for the regeneration Nitrogen, Nm3/h 59.69 Instrument Air, Nm3/h 787.89 Electricity, kW 4658.61 Boiler water t/h 3.80 Gas fuel, kW 4229.62

(24) The consumption of utilities for the regeneration step is relatively low. In particular, the nitrogen consumption (59.69 Nm.sup.3/h) over the duration of functioning of the process remains very low, in particular relative to the nitrogen consumption of a similar system of reactors but with a once-through regeneration (N.sub.2 consumed=47 619 Nm.sup.3/h).

Example 3 (not in Accordance)

(25) In this example, the multi-reactor system comprises two reactors: one reactor is in operational mode while the second is in regeneration (non-operational) mode.

(26) All the other reaction and regeneration parameters are similar to those of Example 1.

(27) Table 6 indicates variations of compositions in the reaction effluent.

(28) TABLE-US-00006 TABLE 6 Variations of compositions in the reaction effluent Component Variation, weight % Ethanol  4.6 Acetaldehyde 22.6 Butadiene 29.3 Diethyl ether 27.4

(29) It appears that the variations in composition of the reaction effluent are greater at the outlet of the unit with two reactors such as that in Example 3 (not in accordance) than those observed in the systems with six reactors described in Examples 1 and 2 (in accordance). The butadiene production, in the process described in Example 3, has variations of 30% by weight.