PROCESS FOR THE ACYLATION OF AN ALPHA, OMEGA-ALKANDIOL

Abstract

The invention relates to a safe and efficient process for the for the acylation of an α,ω-alkanediol, which can be used in the manufacture of ω-nitrooxy-C.sub.3-10alkane-1-ols. The process is safer to operators and allows to obtain advantageous yields on industrial scale.

Claims

1. A process for the acylation of an α,ω-alkanediol, preferably of 1,3-propanediol, with an acylation agent (acylation reaction), said process comprising the step of re-feeding recycled reaction components comprising α,ω-C.sub.3-10alkanediol, α,ω-C.sub.3-10alkanediol monoacylate and α,ω-C.sub.3-10alkanediol diacylate back into said acylation reaction and with the proviso that in said acylation reaction per mole recycled acylate groups 0.5 to 1.5 mole of water is added, and that the molar ratio of the (molar) sum of the acylating agent, α,ω-C.sub.3-10alkanediol monoacylate and 2 times of α,ω-C.sub.3-10alkanediol diacylate to the sum of the α,ω-C.sub.3-10alkanediol, α,ω-C.sub.3-10alkanediol monoacylate and α,ω-C.sub.3-10alkanediol diacylate is selected in the range from 0.5 to 1.1 mol per mol of α,ω-C.sub.3-10alkanediol.

2. The process for the acylation according to claim 1, wherein the process is a continuous process carried out in a vessel cascade set-up.

3. The process for the acylation according to claim 1, wherein the process comprises the step of separation of the α,ω-C.sub.3-10alkanediol monoacylate such that a) the amount of α,ω-C.sub.3-10alkanediol in the α,ω-C.sub.3-10alkanediol monoacylate is less than 0.5 wt.-% and/or b) the amount of the α,ω-C.sub.3-10alkanediol diacylate in the α,ω-C.sub.3-10alkanediol monoacylate is less than 5 wt. %.

4. The process for the acylation according to claim 1, wherein the acylation agent is selected from the group of carboxylic acids and/or derivatives thereof, more preferably of linear or branched C.sub.1-4 carboxylic acids, most preferably the acylation agent is acetic acid.

5. The process for the acylation according to claim 1, wherein the process is used for the preparation of an α,ω-C.sub.3-10alkanediol mononitrate, preferably 3-nitrooxypropane-1-ol, said process further comprising the consecutive steps (B) Continuous nitrate ester formation of the α,ω-C.sub.3-10alkanediol monoacylate by reacting a nitrating agent with a solution comprising the α,ω-C.sub.3-10alkanediol monoacylate and an inert solvent in a group of pieces of equipment comprising at least two reactors in series by simultaneously feeding said solution into the first and the second reactor to obtain the respective α,ω-C.sub.3-10alkanediol mononitrate monoacylate, (C) Two-phase hydrolysis of the α,ω-C.sub.3-10alkanediol mononitrate monoacylate by continuously feeding a base and a solution comprising the α,ω-C.sub.3-10alkanediol mononitrate monoacylate and an inert solvent into a stirred cascade reactor to obtain a solution comprising the inert solvent and α,ω-C.sub.3-10alkanediol mononitrate, and optionally (D) Removal and recovery of the inert solvent from the solution by distillation, said process comprising partial condensation and continuous back-feeding of liquid fractions comprising mixtures of inert solvent and α,ω-C.sub.3-10alkanediol mononitrate into said distillation.

6. The process according to claim 5, wherein in the continuous nitrate ester formation of step (B) the mass flow of the solution into the first reactor is selected in the range of 40 to 60% of the total mass flow of the solution, while the remaining solution is fed into the second reactor.

7. The process according to claim 5, wherein in the continuous nitrate ester formation of step (B) the concentration of the α,ω-C.sub.3-10alkanediol monoacylate in the inert solvent is selected in the range from 10 and 60 wt.-%, more preferably from 20 and 50 wt.-% and most preferably from 35 and 45 wt.-%.

8. The process according to claim 5, wherein in the continuous nitrate ester formation of step (B) the nitrating agent is a mixture of H.sub.2SO.sub.4 and HNO.sub.3, wherein a) the mol ratio of HNO.sub.3 to the α,ω-C.sub.3-10alkanediol monoacylate is selected in the range from 1 to 1.5, preferably from 1.1 to 1.2 and b) the mole ratio of the H.sub.2SO.sub.4 to the HNO.sub.3 is selected in the range from 1.5 to 2.5, preferably from 1.7 to 2.3, most preferably from 1.9 to 2.0.

9. The process according to claim 5, wherein in the continuous nitrate ester formation of step (B) the reaction volume of the first reactor to the second reactor is selected in the range from 4:1 to 1:4, preferably from 3:1 to 1:3, most preferably from 2:1 to 1:2.

10. The process according to claim 5, wherein in the continuous nitrate ester formation of step (B) the nitrate ester formation is carried out for a mean residence time for the two reactors ranging from about 5 to about 30 seconds, preferably from about 10 to about 20 seconds, most preferably from about 15 to 19 seconds.

11. The process according to claim 5, wherein in the continuous nitrate ester formation of step (B) a) the outlet reaction temperature of the reactor 1 is equal to or below 40° C., preferably 30° C., more preferably 20° C., more preferably 10° C. and most preferably equal to or below 5° C., and b) the outlet reaction temperature of reactor 2 is equal to or below 25° C.

12. The process according to claim 5, wherein in the two-phase hydrolysis step (C) a base and a solution comprising an α,ω-C.sub.3-10 alkanediol mononitrate monoacylate and an inert solvent are continuously feed into a stirred cascade reactor.

13. The process according to claim 5, wherein in the two-phase hydrolysis step (C) the base is selected from the group of NaOH, KOH, Ca(OH).sub.2 or ammonia or an aqueous solution thereof, preferably the base is an aqueous solution of NaOH.

14. The process according to claim 5, wherein in the two-phase hydrolysis step (C) the concentration of the base in the aqueous solution is selected in the range from 1 and 50 wt.-%, more preferably from 5 and 30 wt.-%, most preferably from 7.5 and 15 wt.-%.

15. The process according to claim 5, wherein in the two-phase hydrolysis step (C) the the reaction temperature is selected in the range from 20 to 70° C., preferably from 30 to 60° C., and most preferably from 40 to 56° C.

Description

DESCRIPTION OF THE FIGURES

[0161] FIG. 1: In the embodiment of FIG. 1, an exemplary, but none limiting vessel cascade setup for the acylation process according to the present invention is shown: The acylating agent (AA), the α,ω-alkanediol (AD) and water are fed into the first vessel of a vessel cascade (V1) to form the reaction mixture (ARM). The reaction mixture (ARM) is then fed onto a first distillation column (V2) and the acylating agent and water are distilled off to form a mixture (AM-1). The mixture (AM-1) is subsequently fed onto a second distillation column (V3) and the ‘recycled reaction components (I)’ consisting essentially of α,ω-alkanediol mono- and diacylate (ADMA & ADDA) are distilled off to form a mixture (AM-Ia). The mixture (AM-Ia) is then fed onto a third distillation column (V4) and ADMA is distilled off while recovering ‘recycled reaction components (II)’ consisting essentially of AD. During the process the recycled reaction components (I) and (II) are continuously re-fed into the first reaction vessel (V1). A small fraction (below 5%) of the recycled reaction components (I) and (II) can be purged/removed to avoid accumulation of possible formed by-products if necessary. In addition, the acylating agent as well as (part of) the water is re-fed as deemed appropriate into the first reaction vessel (V1).

[0162] FIG. 2: In the embodiment of FIG. 2, an exemplary, but none limiting continuously operated flow-reactor setup for the nitrate ester formation according to the present invention is shown:

[0163] The nitrating agent as well as part of the solution (B-1) consisting of α,ω-C.sub.3-10alkanediol monoacylate and the inert solvent are fed into a first reactor (B1), followed by adding a second portion of the solution (B-1) into the second reactor (B2). The nitrate ester formation reaction mixture (NRM) obtained after reactor (B2) is quenched in reactor B3. The thus obtained reaction biphasic mixture (NBM) is split into two phases to obtain an organic (NOP) and an aqueous phase (NAP). The α,ω-C.sub.3-10alkanediol mononitrate monoacylate (ADMNMA) is in said organic phase and can be isolated thereof.

[0164] FIG. 3: In the embodiment of FIG. 3, an exemplary, but not limiting vertical, stirred cascade reactor setup for the hydrolysis process according to the present invention is shown.

[0165] The first (bottom) chamber (C1) is continuously loaded with a solution consisting essentially of an α,ω-C.sub.3-10alkanediol mononitrate monoacylate and an inert solvent (HS-1) (such e.g. with the NOP obtained as outlined in FIG. 2) and an aqueous solution of a base to form a reaction mixture (HRM). The reaction mixture (HRM) is transferred into a second vessel (C2) for phase separation resulting in an organic phase (HS-II) and an aqueous phase (HS-III). The organic phase is transferred to a evaporator setup (C4) for isolation of the α,ω-C.sub.3-10alkanediol mononitrate (ADMN) The aqueous phase (HS-III) is transferred into a third vessel (C3) for further extraction with the inert solvent to obtain an organic phase HS-IV) which is also transferred (combined with HS-II) to the evaporator setup (C4) to recover further ADMN.

[0166] FIG. 4: In the embodiment of FIG. 4, an exemplary, but none limiting evaporator setup for the removal and recovery of the inert solvent according to the present invention is shown: A solution comprising an α,ω-C.sub.3-10alkanediol mononitrate in an inert solvent (S-1) is fed onto a first evaporator (E1) and a first liquid fraction (LF-I) is removed from the distillate of evaporator (E1) through partial condensation in a first condenser (C1) while the remaining vapors pass onto a second condenser (C2) to liquefy the remaining inert solvent (S). The liquid phase (LP-I) from evaporator (E1) is fed onto a second evaporator (E2). A second liquid fraction (LF-II) is removed from the distillate of evaporator (E2) through partial condensation in a condenser (C3), while the remaining vapors (GP-IV) pass onto a fourth condenser (C4) to liquefy the remaining inert solvent. The liquid phase (LP-II) from evaporator (E2) is fed into a third evaporator (E-3) to remove the remaining inert solvent and to recover the pure α,ω-C.sub.3-10 alkanediol mononitrate.

Example

A) Acylation

[0167] The acetylation (equilibrium formation) was performed either batchwise without recycles or in a vessel cascade setup in a fully continuous process, by feeding the starting materials into a first vessel. The resulting reaction mixture from the last vessel is fed onto a first distillation column for separation (removal) of H.sub.2O/HAc from PDDA/PDMA/PD. This mixture of PDDA/PDMA/PD is fed to a second rectification column for removal of PDDA from PD/PDMA. This mixture of PD/PDMA is fed to a third rectification column for separation of PDMA from PD.

[0168] Pure PDMA was obtained by rectification. Recovered PDDA, PD and HAc were recycled and fed back together with the adjusted amount of water to the reaction vessel cascade, allowing for an overall yield of 90%.

[0169] Aa) Without Using Recycling Streams (Comparative)

[0170] 1,3-Propanediol (PD, 14.0 kg, 0.18 kmol, 99.7%) was mixed with Acetic Acid (HAc, 9.8 kg, 0.16 kmol, 100%). After inerting of the reactor by nitrogen flow, stirring was started (500 rpm) and the jacket temperature was increased from 20° C. to 135° C. within 70 minutes and kept at 135° C. at 4 hours and at reflux of reaction mixture. After 4 hours the jacket temperature was set to 100° C. and the pressure is slowly reduced to approx. 100 mbar abs. while removing 1.55 kg distillate. 22.0 kg residue were obtained comprising a mixture of acetic acid, water, unreacted PD (28 wt %), 3-acetylpropan-1-ol (PDMA, 44.1 wt %) and 1,3-propanedioldiacetate (PDDA, 11.3 wt %). Yield of PDMA was 44.4% and of PDDA was 8.5% based on PD.

[0171] Removal of acetic acid/water was performed at 50 mbar abs top pressure in a rectification column DN50 with 3.5 m BX packing equipped with condenser, liquid separator for reflux adjustment and falling film evaporator with a feed rate of 6.7 kg/h and reflux ratio of 0.4-0.5 resulting in a top take off of 1.1 kg/h containing acetic acid and water and s sump stream of 5.6 kg/h (34 wt % PD, 52 wt % PDMA, 13 wt % PDDA).

[0172] Removal of PDDA was performed at 20 mbar abs top pressure in a rectification column DN50 with 3.5 m BX packing equipped with condenser, liquid separator for reflux adjustment and falling film evaporator with a feed rate of 1.6 kg/h and reflux ratio of 7-8 resulting in a top take off of 0.4 kg/h containing 1 wt % PD, 40 wt % PDMA, and 54 wt % PDDA. The corresponding sump stream (1.2 kg/h) consisted of 44 wt % PD, 55 wt % PDMA, 0.3 wt % PDDA.

[0173] Separation of PDMA from PD was performed at 20 mbar abs top pressure in a rectification column DN50 with 3.5 m BX packing equipped with condenser, liquid separator for reflux adjustment and falling film evaporator with a feed rate of 1.2 kg/h and reflux ratio of 3-4 resulting in a top take off of 0.6 kg/h containing 0.5 wt % PD, 97-98 wt % PDMA, and 1 wt % PDDA. The corresponding sump stream (0.6 kg/h) consisted of 91-92 wt % PD, 8-9% PDMA. Overall Yield of PDMA during the three rectification steps was 71-73%.

[0174] Overall yield of PDMA (reaction and rectification steps) based on PD was 31-33%.

[0175] Ab) Using Recycling Streams in Fully Continuous Mode (Inventive)

[0176] 1,3-Propanediol (PD, 76 kg/h, 0.99 kmol/h, 99.7%) was mixed with fresh Acetic Acid (HAc, 57 kg/h, 100%), 89 kg/h distillate from the first rectification column (56 wt % acetic acid, 44 wt % water), 90 kg/h distillate of the 2.sup.nd rectification column (2 wt % PD, 36.5 wt % PDMA, 61 wt % PDDA) and 110 kg/h sump stream from the third rectification column (97 wt % PD, 3% PDMA). The reaction was performed in a continuous stirred tank reactor at reflux temperature (atmospheric pressure) with a mean residence time of 5-6 hours to deliver 400 kg/h reaction mixture (mixture of acetic acid, water, unreacted PD (29 wt %), 3-acetylpropan-1-ol (PDMA, 35 wt %) and 1,3-propandioldiacetate (PDDA, 14.5 wt %)).

[0177] Removal of acetic acid/water was performed at 50 mbar abs top pressure in a rectification column DN500 with 3.7 m BX packing equipped with condenser, liquid separator for reflux adjustment and falling film evaporator with a feed rate of 400 kg/h and reflux ratio of 0.5-1 resulting in a top take off of 85 kg/h containing acetic acid and water and sump stream of 315 kg/h (36 wt % PD, 45 wt % PDMA, 19 wt % PDDA).

[0178] Removal of PDDA was performed at 20 mbar abs top pressure in a rectification column DN1000 with 10.8 m BX packing equipped with condenser, liquid separator for reflux adjustment and falling film evaporator with a feed rate of 315 kg/h and reflux ratio of 10-15 resulting in a top take off of 92 kg/h containing 2 wt % PD, 36.5 wt % PDMA, and 61 wt % PDDA. The corresponding sump stream (223 kg/h) consisted of 50 wt % PD, 48-49 wt % PDMA, 1-2 wt % PDDA.

[0179] Separation of PDMA from PD was performed at 10 mbar abs top pressure in a rectification column DN1000 with 7.5 m BX packing equipped with condenser, liquid separator for reflux adjustment and falling film evaporator with a feed rate of 223 kg/h and reflux ratio of 5-10 resulting in a top take off of 108 kg/h containing 0.1 wt % PD, 98-99 wt % PDMA, and 1 wt % PDDA. The corresponding sump stream (115 kg/h) consisted of 98-99 wt % PD, 1-2% PDMA.

[0180] Overall yield of PDMA (reaction and rectification steps) based on (fresh) PD was 90%.

B) Nitrate ester formation

[0181] A 40% w/w solution of PDMA in Dichloromethane (DCM) was reacted at 5° C. in a flow-reactor with Nitrosulfonic acid (1.1 eq HNO.sub.3, 2.2 eq H.sub.2SO.sub.4, less than 3 wt % water).

[0182] The nitrate ester formation reaction was performed in a continuously operated flow-reactor, by mixing PDMA in DCM (60 wt % DCM/40 wt % PDMA) with Nitrosulfuric acid in a constant ratio and a steady flow of both components. To control the reaction temperature below 40° C., the reaction was partitioned by massflow between two serial flow-reactors by feeding PDMA in 2 portions (reactor 1/reactor 2=40%: 60%). The overall residence time in both reactors was kept at 15-19 seconds.

[0183] Directly after the 2 sequential reactors, the reaction was diluted/quenched with water at 10° C., followed by a phase separation. The organic phase, containing the intermediate 3-acyl-propan-1-nitrate (MAMN) was washed once with water, stabilizing the mixture for intermediate storage in a buffer tank. The organic phase, containing MAMN can be subjected as—is to the next step or optionally washed with water prior to the next step, with an overall yield of 99%

[0184] The aqueous phase consisting mainly of diluted H.sub.2SO.sub.4 was concentrated to 65 or 96% H.sub.2SO.sub.4 for use in other applications.

C) Hydrolysis

[0185] PDMNMA (ca 50% in DCM) was reacted at 40-56° C. with 1,3 eq. NaOH (10-11% solution in water).

[0186] The hydrolysis of PDMNMA was performed in a vertical, stirred cascade reactor, by continuously feeding PDMNMA (ca 50% in DCM) together with 10-11% NaOH solution (in a ratio 1/1.3 eq.) from the bottom. Residence time was 4 hours, at a reaction temperature of 40-56° C. After complete conversion (>99.9%), the phases were cooled to appr. 20° C., split, and the aq. phase was washed/extracted in continuous mode with DCM (back-extraction of PDMN) at room temperature. The combined organic phases were subjected to solvent removal (see D) Workup).

[0187] The desired product is obtained in 97% yield, after removal of DCM from the combined organic phases.

D) Solvent removal and recovery (partial condensation)

[0188] After hydrolysis, the combined organic phases PDMN/DCM (77% DCM) were subjected to solvent removal in a 3-stage evaporator setup, by feeding the organic phases into a first evaporator where PDMN solution (containing 7-8% DCM) was produced at 500 mbar. The distillates (vapour stream) were directed to a partial condenser, where a liquid fraction (PDMN/DCM, ca. 55-60% PDMN) was recovered at 30° C. and fed back to the first evaporator. The remaining vapors passed to a (total) condenser operated at 0° C. to recover DCM in high purity (<0.03% PDMN).

[0189] The PDMN solution from the first evaporator (containing 7-8% DCM) is fed to a second evaporator operated at 100 mbar to produce a liquid solution containing ca. 1 wt % DCM. The distillates (vapour stream) were directed to a partial condenser, where a liquid fraction (PDMN/DCM, ca. 70-75% PDMN) was recovered at 15° C. and fed back to the first evaporator. The remaining vapors passed to a (total) condenser operated at 0° C. to recover DCM (ca. 0.1% PDMN).

[0190] The PDMN solution from the second evaporator (containing ca. 1% DCM) is fed to a third evaporator operated at 10 mbar to produce a liquid solution containing less than 0.1 wt % DCM. The distillates (vapour stream) were directed to a partial condenser, where a liquid fraction (PDMN/DCM, ca. 90% PDMN) was recovered at 0° C. and fed back to the first evaporator. The remaining vapors were discarded.