Process for the regeneration of a titanium zeolite catalyst for propylene epoxidation

Abstract

The invention relates to process for the regeneration of a catalyst comprising a titanium containing zeolite as catalytically active material comprising a stage comprising introducing a feed stream comprising propene, hydrogen peroxide or a hydrogen peroxide source, and an organic solvent into a reactor containing a catalyst comprising the titanium containing zeolite, subjecting the feed stream in the reactor to epoxidation conditions in the presence of the catalyst, removing a product steam comprising propylene oxide and the organic solvent from the reactor, stopping introducing the feed stream, washing the catalyst with a liquid aqueous system and calcining the washed catalyst.

Claims

1. A process for regenerating a catalyst comprising a titanium comprising zeolite of MWW framework structure as catalytically active material, the process comprising: (i) a stage comprising (a) continuously preparing propylene oxide by a process comprising, in the following order: (a1) introducing a feed stream comprising propene, hydrogen peroxide or a hydrogen peroxide source, and an organic solvent into a reactor comprising a catalyst comprising the titanium comprising zeolite of MWW framework structure as catalytically active material; (a2) subjecting the feed stream according to (a1) in the reactor to epoxidation conditions in the presence of the catalyst, obtaining a reaction mixture comprising the propylene oxide and the organic solvent; and (a3) removing a product stream comprising the propylene oxide and the organic solvent from the reactor; (b) stopping introducing the feed stream into the reactor; (c) washing the catalyst with a liquid aqueous system comprising at least 50 weight-% water, based on the total weight of the liquid aqueous system; and then, (ii) a stage comprising calcining the catalyst obtained from (c), wherein the calcining is performed at a temperature of the catalyst in the range of from 300 C. to 600 C. using a gas stream comprising oxygen; wherein the sequence of steps (a) to (c) in stage (i) is performed n times, where n is an integer and at least 2, such that the washing of the catalyst with a liquid aqueous system according to step (c) is followed by the continuous preparation of propylene according to step (a) at least once without an intervening calcination stage (ii).

2. The process of claim 1, wherein the organic solvent is methanol or acetonitrile.

3. The process of claim 1, wherein at least 99 weight-% of the framework structure of the titanium comprising zeolite of MWW framework structure consists of silicon, titanium, and oxygen.

4. The process of claim 1, wherein the titanium comprising zeolite of MWW framework structure comprises Zn.

5. The process of claim 1, wherein the catalyst comprising a titanium comprising zeolite of MWW framework structure is present in the reactor as fixed-bed catalyst.

6. The process of claim 1, wherein the liquid aqueous system according to (c) comprises at least 95 weight-% water, based on the total weight of the liquid aqueous system.

7. The process of claim 1, wherein according to (c), the washing of the catalyst is carried out in the reactor comprising the catalyst in continuous mode.

8. The process of claim 7, wherein the reactor according to (a) is a tube reactor or a tube bundle reactor and the washing according to (c) is performed with the liquid aqueous system at a liquid hourly space velocity (LHSV) in the range of from 1 to 20 m/h.

9. The process of claim 1, wherein the washing according to (c) is performed at a temperature of the liquid aqueous system in the range of from 30 to 90 C.

10. The process of claim 1, wherein the washing according to (c) is performed until the total organic carbon concentration of the liquid aqueous system after having been contacted with the catalyst is at most 5% of the maximum total organic carbon concentration detected during the washing in (c).

11. The process of claim 1, wherein the sequence of steps (b) and (c) is carried out when the selectivity of the epoxidation reaction according to (a2) has decreased by 4% or less, relative to the average selectivity of the epoxidation reaction according to (a2) during the first 100 h of carrying out step (a), wherein the selectivity of the epoxidation reaction is defined as the molar amount of propylene oxide obtained in (a2) relative to the molar amount of hydrogen peroxide converted in (a2).

12. The process of claim 1, wherein n is in the range of from 2 to 7.

13. The process of claim 1, wherein the calcining according to stage (ii) is performed at a temperature of the catalyst in the range of from 350 to 550 C.

14. The process of claim 1, wherein the calcining according to stage (ii) is carried out in the reactor comprising the catalyst.

15. The process of claim 1, further comprising: repeating stages (i) and (ii) m times, wherein m is an integer and at least 1, and wherein in each repetition of stages (i) and (ii) the integer n is the same or different.

16. The process of claim 15, wherein m is an integer from 1 to 6.

17. A continuous process for preparing propylene oxide, the process comprising: (i) a stage comprising (a) continuously preparing propylene oxide in a first reactor by a process comprising, in the following order: (a1) introducing a feed stream comprising propene, hydrogen peroxide or a hydrogen peroxide source, and an organic solvent into the first reactor comprising a catalyst comprising a titanium comprising zeolite of MWW framework structure as catalytically active material; (a2) subjecting the feed stream according to (a1) in the first reactor to epoxidation conditions in the presence of the catalyst, obtaining a reaction mixture comprising the propylene oxide and the organic solvent; and (a3) removing a product stream comprising the propylene oxide and the organic solvent from the first reactor; (b) stopping introducing the feed stream into the first reactor; (c) washing the catalyst with a liquid aqueous system comprising at least 50 weight-% water, based on the total weight of the liquid aqueous system; and then (ii) a stage comprising calcining the catalyst obtained from (c), wherein the calcining is performed at a temperature of the catalyst in the range of from 300 C. to 600 C. using a gas stream comprising oxygen; wherein the sequence of steps (a) to (c) in stage (i) performed n times, where n is an integer and at least 2, such that the washing of the catalyst with a liquid aqueous system according to step (c) is followed by the continuous preparation of propylene according to step (a) at least once without an intervening calcination stage (ii); and the process for the preparation of propylene oxide further comprising (i) a stage comprising (a) continuously preparing propylene oxide in a second reactor by a process comprising, in the following order: (a1) introducing a feed stream comprising propene, hydrogen peroxide or a hydrogen peroxide source, and an organic solvent into the second reactor comprising a catalyst comprising a titanium comprising zeolite of MWW framework structure as catalytically active material; (a2) subjecting the feed stream according to (a1) in the second reactor to epoxidation conditions in the presence of the catalyst, obtaining a reaction mixture comprising the propylene oxide and the organic solvent; (a3) removing a product stream comprising the propylene oxide and the organic solvent from the second reactor; (b) stopping introducing the feed stream into the second reactor; and (c) washing the catalyst with a liquid aqueous system comprising at least 50 weight-% water, based on the total weight of the liquid aqueous system; and then (ii) a stage comprising calcining the catalyst obtained from (c), wherein the calcining is performed at a temperature of the catalyst in the range of from 300 C. to 600 C. using a gas stream comprising oxygen; wherein the sequence of steps (a) to (c) in stage (i) is performed n times, where n is an integer and at least 1, such that the washing of the catalyst with a liquid aqueous system according to step (c) is followed by the continuous preparation of propylene according to step (a) at least once without an intervening calcination stage (ii); and wherein during at least one sequence of steps (b) and (c) in the first reactor or during at least one sequence of steps (b), (c), and (ii) in the first reactor, propylene oxide is prepared according to (a) in the second reactor.

18. The process of claim 17, further comprising: repeating stages (i) and (ii) m times, wherein m is an integer and at least 1, and wherein in each repetition of stages (i) and (ii) the integer n is the same or different; and/or repeating stages (i) and (ii) m times, wherein m is an integer and at least 1, and wherein in each repetition of stages (i) and (ii) the integer n is the same or different.

Description

DESCRIPTION OF THE FIGURES

(1) FIG. 1 shows a block diagram of the process according to Reference Example 2. In FIG. 1, the letters and numbers have the following meanings: A epoxidation unit A B epoxidation unit B C distillation unit D distillation unit E distillation unit F part stream distillation unit G mixer-settler unit H acetonitrile recovery unit acetonitrile recycle unit (1)-(20) streams according to a specifically preferred process as described in the examples S0, S01, S02, S1, S2, S3, S4, S4b, S5, L1, L2, TL1, TL2, TL2, BL2 streams according to a preferred process as described in the general description and the examples

(2) FIG. 2 shows a block diagram the part stream distillation F of FIG. 1 unit in detail. In FIG. 2, the letters and numbers have the following meanings: F1 first fractionation unit of the part stream distillation unit F F2 second fractionation unit of the part stream distillation unit F (13), (13a), (14), (15), (15a), (15b), (15c), (16), (19), (20) streams according to a specifically preferred process as described in the examples S1, S2, S3, S4, S4a, S4b, S4c, S5, TL2 streams according to a preferred process as described in the general description and the examples

EXAMPLES

Reference Example 1: Preparation of a Molding Containing ZnTiMWW Spray Powder

(3) The catalyst in the form of a molding was prepared as described in WO 2013/117536 A2, in Example 5, in particular in Examples 5.1-5.6, on pages 83-99. The methods of characterizing said molding referred to in said Examples 5.1-5.6 are described in Reference Examples 2-10 on pages 66-71 of WO 2013/117536 A2. Characteristic properties of the catalyst are shown in FIGS. 20-27 of WO 2013/117536 A2 and the respective description of said Figures on page 104 of WO 2013/117536 A2.

Reference Example 2: Epoxidation Reaction Setup

(4) As to the abbreviations, reference is made to the schemes according to FIGS. 1 and 2, generally described in the section Description of the Figures. All pressures given are absolute pressures.

(5) 1.1 Preparation of Stream S0 (Step (a))

(6) a) Epoxidation in an Epoxidation Main Reactor (Epoxidation Unit A)

(7) The main reactor A was a vertically mounted tube-bundle reactor with 5 tubes (length of the tubes: 12 m, internal tube diameter: 38 mm), each tube being equipped with an axially placed multi-point thermocouple with 10 equally spaced measuring points encased in a suitable thermowell with a diameter of 18 mm. Each tube was charged with 17.5 kg of the ZnTiMWW catalyst moldings as prepared according to Reference Example 1 (post-treated moldings). Free space eventually remaining was filled with steatite spheres (diameter of 3 mm). The heat of reaction was removed by circulating a thermostatized heat transfer medium (water/glycol mixture) on the shell side in co-current to the feed. The flow rate of the heat transfer medium was adjusted so that the temperature difference between entrance and exit did not exceed 1 C. The reaction temperature referred to hereinbelow was defined as the temperature of the heat transfer medium entering the reactor shell. At the reactor exit, the pressure was controlled by a pressure regulator and kept constant at 20 bar.

(8) The reactor was fed from below with a liquid monophasic stream (1). Stream (1) was prepared by mixing four streams (2), (3), (3a) and (4). The temperature of stream (1) was not actively controlled, but was usually in the range from 20 to 40 C.: Stream (2) having a flow rate of 85 kg/h. At least 99.5 weight-% of stream (2) consisted of acetonitrile, propene and water. This stream (2) came from the bottoms of the acetonitrile recycle distillation unit W. Stream (3) having a flow rate of 15 kg/h was an aqueous hydrogen peroxide solution having a hydrogen peroxide concentration of 40 weight-% (crude/washed grade from Solvay with a TOC in the range of 100 to 400 mg/kg. The aqueous hydrogen peroxide solution was supplied from a storage tank, allowing for a continuous feeding, and fed using a suitable metering pump. Stream (3a) was an aqueous stream comprising dissolved potassium formate. The further stream was supplied from a storage tank, allowing for a continuous feeding, and was fed using a suitable metering pump. The concentration of the potassium formate was 2.5 weight-%, the feed rate of the stream (S3a) was 370 g/h. Stream (3a) was thoroughly mixed with stream (3) before the combined stream was mixed with the stream resulting from mixing stream (2) and (4). Stream (4) was a make-up stream of pure acetonitrile (chemical grade, from Ineos, purity about 99.9%, containing between 70-180 weight-ppm propionitrile, 5-20 weight-ppm acetamide and less than 100 weight-ppm water as impurities). Enough fresh acetonitrile was added to compensate for losses in the process. Under regular conditions, an average of from 100 to 150 g/h of make-up acetonitrile were added.

(9) The output stream (5) leaving the epoxidation unit A was sampled every 20 minutes in order to determine the hydrogen peroxide concentration using the titanyl sulfate method and to calculate the hydrogen peroxide conversion. The hydrogen peroxide conversion was defined as 100(1m.sub.out/m.sub.in) wherein m.sub.in is the molar flow rate of H.sub.2O.sub.2 in the reactor and m.sub.out is the molar flow rate of H.sub.2O.sub.2 in the reactor outlet. Based on the respectively obtained hydrogen peroxide conversion values, the inlet temperature of the heat transfer medium was adjusted in order to keep the hydrogen peroxide conversion essentially constant in the range of from 90 to 92%. The inlet temperature of the heat transfer medium was set at 30 C. at the start of a given run with a fresh batch of the epoxidation catalyst and was increased, if necessary, to maintain the hydrogen peroxide conversion in the mentioned range. The required temperature increase was usually less than 1 Kid.

(10) The output stream (5) leaving the epoxidation unit A was passed through a heat exchanging unit. The stream leaving the heat exchanging unit (stream (6), S0) was fed to Epoxidation Unit B.

(11) b) Epoxidation in a Finishing Reactor (Epoxidation Unit B)

(12) The finishing reactor B was a fixed bed reactor operated adiabatically. In this context, the term adiabatic refers to an operation mode according to which no active cooling is carried out and according to which the finishing reactor is suitably insulated in order to minimize heat losses). The finishing reactor B had a length of 4 m and a diameter of 100 mm. The reactor was filled with 9 kg of the same epoxidation catalyst which was used in the main epoxidation reactor A. Spare space was filled with steatite spheres (diameter of 3 mm). The operating pressure of the finishing reactor B was 10 bar which was kept constant by a suitable pressure regulator at the reactor exit. The output of the finishing reactor B was sampled every 20 min in order to determine the hydrogen peroxide concentration using the titanyl sulfate method.

(13) The effluent of the finishing reactor B, stream (6), was depressurized into a flash drum, and both the liquid and the gas from this drum were fed to a light boiler separation column (distillation unit C).

(14) Stream (6) (stream S0) had in average an acetonitrile content of from 69 to 70 weight-%, a propylene oxide content of from 9-11 weight-% such as 9.8 weight-%, a water content of 17 weight-%, a propene content of about 3 weight-%, a propane content of about 0.05 weight-%, a hydrogen peroxide content of about 250 weight-ppm, a propene glycol content of about 0.1 weight-% and an oxygen content of about 150 weight-ppm.

(15) 1.2 Separation of Propylene Oxide from Stream S0 to Obtain Stream S1 (Step (b))

(16) a) Separation of Light Boilers from Stream (6) (Stream S0) to Obtain a Stream (8) (Stream S01)

(17) Stream (6) was sent to a light boiler separation column (distillation unit C) operated at 1.1 bar. The distillation column had a length of 8.5 m, a diameter of 170 mm, and was equipped with 40 bubble trays, an evaporator at the bottom and a condenser at the top. The column was operated as a mixed washing/distillation tower. As a washing agent, part of the bottoms stream of distillation unit D (stream 11, about 20-30 kg/h) was taken off, cooled to 10 C. and introduced at the top of the column. Liquid and gaseous inlet streams were introduced the column at different points. The feed point of the liquid portion of stream (6) was above bubble tray 37; the gaseous portion of stream (6) was introduced into the column above bubble tray 28 (counted from the top). The gaseous stream (7) leaving the cooling means at the top of the column contained mainly propene, propane (which was contained as impurity in the polymer-grade propene used), oxygen formed as a by-product and small amounts of other light boilers (acetonitrile (1-2 volume-%), propionaldehyde (about 200 volume-ppm), acetone (about 100 volume-ppm, H.sub.2 (about 400 volume-ppm), CO.sub.2 (about 400 volume-ppm) and acetaldehyde (about 100 volume-ppm)), and was essentially free of propylene oxide (less than 300 volume-ppm). This top stream was sent to the flare for disposal. The bottom stream of the light boiler separation column (stream (8), that is stream S01) having a temperature of 70 C., had a propene content of from 100 to 200 weight-ppm.

(18) b) Separation of Propylene Oxide from Stream (8) (Stream S01) to Obtain a Stream S02

(19) The stream S01 obtained according to section 1.2 a) above was introduced into a distillation column (distillation unit D) in order to separate propylene oxide from the stream S01. The column had a height of 50 m and a diameter of 220 mm and was equipped with a packing (Sulzer BX64) with a total packing length of 27.5 m divided into 8 beds with a length of 3060 mm each and two beds with a length of 1530 mm each. Between each bed intermediate flow distributors were installed. The column was operated at a top pressure of 750 mbar. The feed point of stream S01 was located below the fourth packing bed, counted from the top. The overhead stream of the column was condensed and partly returned to the column as reflux (reflux ratio approximately 5:1). The remainder (stream (9)), having a flow rate of 10.1 kg/h, was taken as overhead product and essentially consisted of propylene oxide having a purity of more than 99.9 weight-%. The bottoms evaporator was operated in such a way that the propylene oxide concentration in the bottoms stream was below 100 weight-ppm. The resulting temperature of the bottoms stream was about 69 C. The stream S02 was then divided in two. The major portion of it (stream (10), with a flow rate of ca. 85 kg/h) was sent to the next distillation column (distillation unit E). The remainder (stream (11), 20-30 kg/h) was cooled and recirculated to the top of the light boiler separation column (distillation unit C) as washing agent as described above in section 1.2 a). This stream S02 had an acetonitrile content of about 80 weight-%, a propylene oxide content of less than 100 wt.-ppm, a water content of about 20 weight-%, a propene glycol content of about 0.1 weight-% and a hydroxypropanol content of about 0.1 weight-%.

(20) c) Separation of Light Boiling Compounds from Stream (10) (Stream S02) to Obtain a Stream (13) (Stream S1)

(21) The stream S02 obtained according to section 1.2 b) above was introduced into a lights separation column (distillation unit F). This lights separation column had a height of 8 m and a nominal diameter of 150 mm and was equipped with 35 bubble trays. The column was operated at a top pressure of 2 bar, and the stream S02 was introduced above bubble tray number 7 (counted from the bottom). The overhead stream obtained (stream (12), flow rate about 1 kg/h) left the column with a temperature of from 40 to 45 C. and was not condensed as the column was operated with no internal reflux stream. Besides acetonitrile (6500 vol.-ppm), this overhead stream contained mainly nitrogen which was employed to keep the column operating pressure at a value of 2 bar) and small amounts of light boilers (acetaldehyde (900 vol.-ppm), oxygen (300 vol.-ppm), and propionaldehyde (320 vol.-ppm). This top stream was sent to the flare for disposal. The sump evaporator was operated by feeding it with a constant amount (5 kg/h) of saturated steam at a pressure of 16 bar. The bottom temperature of the column was 100 C. The bottoms stream, stream S1, mainly consisted of acetonitrile and water, the remainder being high boilers. This stream S1 had an acetonitrile content of about 80 weight-% and a water content of about 20 weight-%.

(22) 1.3 Dividing Stream S1 into Streams S2 and S3 (Step (C))

(23) Step (c), the stream S1, flow rate 86 kg/h, obtained according to section 1.2 c) above, was divided into two streams, streams S2 (stream (13a according to FIG. 1) and S3 (stream 14 according to FIG. 1). Stream S2 had a flow rate of 84 kg/h and stream S3 had a flow rate of 2 kg/h. Stream S3, 2.3% of stream S1, was subjected to part stream distillation unit F (part stream distillation columns).

(24) 1.4 Part-Stream Distillation of Stream S1 (Step (d))

(25) The first fractionation unit, i.e. the first distillation column, F1, had a height of 9.5 m and a diameter of 85 mm and was equipped with 6.5 meters of metal structured Rombopak 9M packing installed in three identical beds. Above the first bed of the structured packing counted from the top, the stream S3 ((stream 14)) was introduced in the first distillation column. The temperature of the stream S3 stream was 603 C. The first distillation column was operated at a top pressure of about 1.4 bar and a bottoms temperature of 925 C. No reflux was applied. The amount of steam fed to the bottoms evaporator of the first fractionation unit was controlled in such a way that the concentration of acetonitrile in the bottoms was in the range of from 10 to 25 weight-%. The bottoms stream S4b (stream (15b), about 3% of the stream S3) was removed. This stream consisted mainly of water (72-85 weight-%) and acetonitrile (10-24 weight-%). The sum of all the analyzed high-boiling components (27 components) varied in the range of 2-10 weight-%. The top stream, vapor fraction stream S4a (stream 15a), having a temperature of from 853 C., was not condensed and passed to the bottom of the second fractionation unit, i.e. the second distillation column, F2. S4a entered F2 below the last bed of the structured packing counted from the top. F2 had a height of 9.5 m and a diameter of 85 mm and was equipped with 6.5 m of metal structured Rombopak 9M packing installed in 3 identical beds. The second distillation column was operated at a top pressure of about 1.25 bar and a bottoms temperature of 855 C. The top stream, vapor fraction stream S4c (stream (15c), at most 1% of the stream S4a), was fully condensed by an external overhead condenser (not shown in FIG. 2) and applied essentially completely to use the condensed, liquid stream as reflux to the second distillation column. The liquid bottoms stream S4 (stream 15), was removed and passed to the next step (recycling of the stream 54). The stream S4 had an acetonitrile content of about 80 weight-% and a water content of about 20 weight-%.

(26) 1.5 Recycling of the Stream S4 (Step (4))

(27) a) Preparing a Liquid Stream S5

(28) The stream S4, (stream 15 according to FIG. 1 and FIG. 2) was admixed with stream S2 (stream (13a) according to FIG. 1 and FIG. 2). Thus, the stream S4 was pumped back into the bulk process acetonitrile solvent stream. Mixing took place at a point downstream of where stream S3 was diverted from stream S1. This combined stream having a flow rate of 86 kg/h was mixed with a liquid stream P (referred to as stream (20) in FIG. 1 and FIG. 2) to obtain a stream S5. Stream P was fresh propene stream containing propane (polymer grade, purity >96 weight-%, liquefied under pressure, feed rate: 10.9 kg/h). In order to obtain the stream S5, the combined stream of S2 and S4 was further mixed with two other streams: the first one of these streams is stream (16) according to FIG. 1, said stream being obtained from the top of the distillation unit H. The second one of these streams is stream (19) according to FIG. 1, said stream being obtained from the acetonitrile recovery unit I. Both streams (16) and (19) are described in detail hereinunder.

(29) b) Adjusting the temperature of stream S5 and separating liquid phases L1 and L2

(30) The stream S5 having a flow rate of 130 kg/h10 kg/h was then fed to a mixer-settler unit operated at 18 bar and a temperature in the range of 155 C. The settler tank had a volume of 5.3 liters. Two liquid phases L1 and L2 were obtained, an aqueous phase L2 and an organic phase L1. The upper organic phase L1 was removed from the settler tank as stream (17), the lower aqueous phase L2 was removed from the settler tank as stream (18). The stream (17) had a flow rate in the range of 110 kg/h11 kg/h. The stream (17) then was passed to the acetonitrile recycle unit I, the stream (18) was passed to the acetonitrile recovery unit H from which the stream (16) mentioned above was obtained. The stream (17) thus obtained had an acetonitrile content of about 45-51 weight-%, a propene content of about 49-55 weight-% and a water content of about 2 to 5 weight-%. The stream (18) thus obtained had an acetonitrile content of about 19-21 weight-%, a water content of about 79-81 weight-% and a propene content of less than 0.5 weight-%.

(31) c) Acetonitrile Recovery (Acetonitrile Recovery Unit H)

(32) In order to recycle as much solvent as possible, and in order to minimize acetonitrile losses, the stream (18) was introduced into a distillation column from which the stream (16), also referred to as stream TL2, was obtained as top stream which in turn was recycled into the solvent stream as described above. For this purpose, a distillation column with a height of 9.5 m and a diameter of 100 mm, equipped with 50 bubble trays was used. The column was operated at a top pressure of 1.5 bar with a reflux ratio of 1:4. Stream (18) was fed to the column above bubble tray 26 (counted from the bottom). The bottoms temperature was about 113 C., and the bottoms product consists mainly of water containing high boiling by-products. A typical composition of the bottoms stream was as follows (weight-% given in parenthesis): water (>99.0), propene glycol (0.5), acetonitrile (at most 0.001), dipropylene glycol (0.06), acetamide (0.01), acetic acid (0.03), TOC (2.4)). After optional metering and analyzing, this stream was discarded. The overhead product (stream (16)=stream TL2) had the following typical composition ranges (weight-% given in parenthesis): acetonitrile (75-80), water (15-20), low boilers (e.g. propene, 1). As described above stream (16) is recycled to the feed stream which is passed to the mixer-settler unit.

(33) d) Acetonitrile Recycling (Acetonitrile Recycling Unit I)

(34) For acetonitrile recycle, the stream (17) obtained from the mixer-settler unit G was introduced into a distillation column with a height of 10 m and a nominal diameter of 200 mm, equipped with 40 bubble trays. The column was operated at a top pressure of 18 bar with a reflux ratio of 1:4. Stream (17) was fed to the column above bubble tray 26 (counted from the top). The top product (stream (19)), also referred to as stream TL1, containing mainly propene (ca. 97 vol.-%) with small amounts of propane (ca. 1-3 vol.-%) was returned to the feed of the mixer-settler unit G as described above. Thus, excess propene was removed from steam (17) and recycled. The bottoms stream (stream (2), also referred to as stream BL1), had a temperature in the range of from 106 to 110 C. The precise operation parameters of the column, like energy input in the sump, are adjusted in such a way that the amount of propene returned to the reactor with stream (2) is in a range such that the molar ratio of propene to hydrogen peroxide in stream (1) was about 1:1.43. For the above mentioned feed rate of 15 kg/h of aqueous hydrogen peroxide, this means that the conditions needed to be adjusted such that the flow rate of propene in stream (2) was about 9.7 kg/h. Prior to feeding stream (2) to the main epoxidation reactor A, acetonitrile (stream (4), chemical grade, from Ineos, purity about 99.9%, containing between 70-180 weight-ppm propionitrile, 5-20 weight-ppm acetamide and <100 weight-ppm water as impurities) was optionally added to compensate for possible solvent losses.

(35) The exact amount of additionally added acetonitrile required depended on the losses in exit streams and in by-products but also on the number of samples taken for analytics. A typical amount of additionally added acetonitrile for the above-described process design may be in the range of from 100 to 150 g/h.

Example 1: One-Time ZnTiMWW Catalyst Regeneration

(36) The partial ZnTiMWW catalyst regeneration was performed based on the epoxidation reaction setup as described in Reference Example 2 above. The reactor was loaded with the ZnTiMWW catalyst according to Reference Example 1. The epoxidation reaction was continuously operated until the selectivity regarding the formation of propylene oxide in (a2) based on hydrogen peroxide decreased by 2% relative to the average selectivity regarding the formation of propylene oxide in (a2) based on hydrogen peroxide during the first 1000 h of carrying out step (a). Then the reaction was stopped and the reactor flushed free of H.sub.2O.sub.2, propylene and propylene oxide. After decompression and emptying, the reactor was flushed at 70 C. with demineralized water from top to bottom with a LHSV of about 7 m/h (the LHSV being based on the cross section of the empty tubes) for at least 3 hours. After this time, the liquid aqueous system leaving the reactor showed a total organic carbon concentration of less than 0.1% of the maximum value detected and the total volume of the liquid aqueous system obtained from the reactor in washing process was larger than the volume of the reactor. Then, the reactor was emptied and restarted. Continuous operation according to Reference Example 2 was resumed and carried out for approximately 1500 h until the reaction temperature reached again the value it had before the regeneration.

Comparative Example 1: Continuous Operation According to Reference Example 2 without Regeneration

(37) As a comparative Example, the process according to Reference Example 2 was carried out for 2,500 h without regenerating the catalyst.

Results from Example 1 and Comparative Example 1

(38) The results from example 1 and comparative example 1 are discussed hereinafter and shown in Table 1. The inlet temperature of the heat transfer medium of epoxidation unit A both in Example 1 and Comparative Example 1 was adjusted during the continuous operation according to Reference Example 2 in order to keep the hydrogen peroxide conversion essentially constant in the range of from 90 to 92% (cf. Reference Example 2, section 1.1 a). The resulting temperature of the heat transfer medium (cooling water) of epoxidation unit A after a total operation time of 2,500 h is shown in the table 1 both for Comparative Example 1 and for Example 1. Further, the deactivation rate is given as the ratio of the difference of the temperature of the heat transfer medium at the inlet at the end of the experiment and the beginning of the experiment divided by the total operation time of 2,500 h, expressed in C./day. Yet further, the selectivity S of the epoxidation reaction based on hydrogen peroxide after the total operation time of 2,500 h is provided in Table 1. The selectivity S/%=(n(PO)/n(H.sub.2O.sub.2))100 wherein n(PO) is the molar amount of propylene oxide detected directly downstream to unit B and n(H.sub.2O.sub.2) is the molar amount of hydrogen peroxide converted in the epoxidation reaction.

(39) TABLE-US-00001 TABLE 1 Comparative Example 1 Example 1 Temperature of the cooling water of epoxidation 55 49 unit A after 2,500 h time on stream/ C. Deactivation rate (delta(T)/delta(t))/ C./day 0.240 0.182 Selectivity S/% 96.5 97.2

(40) As clearly shown in Table 1, the process according to the present invention allows to run the epoxidation unit A after a total operation time of 2,500 h at a significantly lower temperature of the cooling water, i.e. the catalyst exhibits a significantly better performance than in comparative example 1. Further, also the deactivation rate (delta(T)/delta(t)) differs for the two experiments and is significantly lower with regard to the inventive example. Yet further, the process according to the present invention provides, after a total operation time of 2,500 h, a significantly higher selectivity than the comparative example 1.

(41) The regenerated catalyst of example 1 exhibited a differential conversion temperature of 1 K, wherein the differential conversion temperature is defined as the absolute difference between (A1) the temperature at which the conversion of the hydrogen peroxide 90-92% was achieved when the regenerated catalyst is used as catalyst, and (B1) the temperature at which said determined conversion of the hydrogen peroxide was achieved when the respective fresh catalyst was used under otherwise identical epoxidation reaction conditions. Further, the regenerated catalyst of example 1 exhibited a differential selectivity of 0.3% points wherein the differential selectivity is defined as the absolute difference in % points between (A2) the selectivity based on the hydrogen peroxide in the epoxidation process in which the regenerated catalyst is used as catalyst, and (B2) the selectivity based on the hydrogen peroxide in said epoxidation process in which the respective fresh catalyst is used as catalyst under otherwise identical epoxidation reaction conditions.

Example 2.1: ZnTiMWW Catalyst Regeneration According to the Invention Including Three Washing Steps

(42) The partial ZnTiMWW catalyst regeneration was performed based on the epoxidation reaction setup as described in Reference Example 2 above.

(43) Epoxidation

(44) The partial ZnTiMWW catalyst regeneration was performed based on the epoxidation reaction setup as described in Reference Example 2 above. The main reactor was loaded with the ZnTiMWW catalyst according to Reference Example 1. The epoxidation reaction was continuously operated for approximately 1300 hours until the epoxidation reaction temperature, defined as the temperature of the heat transfer medium at the entrance of the jacket of the main reactor whose staffing value at the beginning of the epoxidation reaction was 30 C., had reached a value of 50 C. At that point in time, the epoxidation reaction was stopped by stopping the flow of hydrogen peroxide; the acetonitrile and propene flows were continued until the epoxidation in the tubes was completed. Then, the reactor was flushed free of H.sub.2O.sub.2, propylene and propylene oxide by passing a mixture of acetonitrile and water through the reactor (80 weight-% acetonitrile, 20 weight-% water) wherein the temperature of this mixture was 50 C., followed by draining the major portion of the mixture of acetonitrile and water from the reactor.

(45) First Washing Step

(46) Subsequently, the reaction tubes were filled from below with water having a temperature of 70 C., and water having a temperature of 70 C. was then passed from the top downward through the tubes for about 6 h. The water wash was stopped once the conductivity of the outflowing water was below 200 microSiemens. The water was then drained from the reactor.

(47) Epoxidation

(48) Continuous operation according to Reference Example 2 was resumed and carried out for approximately 1400 h until the reaction temperature reached again the value of 50 C. At that point in time, the epoxidation reaction was stopped by stopping the flow of hydrogen peroxide; the acetonitrile and propene flows were continued until the epoxidation in the tubes was completed. Then, the reactor was flushed free of H.sub.2O.sub.2, propylene and propylene oxide by passing a mixture of acetonitrile and water through the reactor (80 weight-% acetonitrile, 20 weight-% water) wherein the temperature of this mixture was 50 C., followed by draining the major portion of the mixture of acetonitrile and water from the reactor.

(49) Second Washing Step

(50) Subsequently, the reaction tubes were filled from below with water having a temperature of 70 C., and water having a temperature of 70 C. was then passed from the top downward through the tubes for about 6 h. The water wash was stopped once the conductivity of the outflowing water was below 200 microSiemens. The water was then drained from the reactor.

(51) Epoxidation

(52) Continuous operation according to Reference Example 2 was resumed and carried out for approximately 1200 h until the reaction temperature reached again the value of 50 C. At that point in time, the epoxidation reaction was stopped by stopping the flow of hydrogen peroxide; the acetonitrile and propene flows were continued until the epoxidation in the tubes was completed. Then, the reactor was flushed free of H.sub.2O.sub.2, propylene and propylene oxide by passing a mixture of acetonitrile and water through the reactor (80 weight-% acetonitrile, 20 weight-% water) wherein the temperature of this mixture was 50 C., followed by draining the major portion of the mixture of acetonitrile and water from the reactor.

(53) Residual amounts of acetonitrile and water were removed from the reactor by passing nitrogen through the catalyst bed at a temperature of the nitrogen of 70 C.

(54) Third Washing Step

(55) Subsequently, the reaction tubes were filled from below with water having a temperature of 70 C., and water having a temperature of 70 C. was then passed from the top downward through the tubes for about 6 h. The water wash was stopped once the conductivity of the outflowing water was below 200 microSiemens. The water was then drained from the reactor.

(56) Calcining Step

(57) Then, nitrogen was passed through the catalyst bed for approximately 60 hours at an initial temperature of the nitrogen of 70 C. wherein this temperature was ramped to a value of 100 C. at 10 K/h. Subsequently, the catalyst was calcined by passing a gas mixture of air and nitrogen through the catalyst bed at a flow rate of 0.58 kg/s. The temperature of said mixture was further continuously ramped at 10 K/h from its initial value of 100 C. to a final value of 450 C. and then kept essentially constant at a value of 450 C. for 6 hours. At the beginning of the calcination, the gas mixture passed into the reactor had an oxygen content of 4 volume-%. Once having reached 450 C., said oxygen content was adjusted to a value of 21 volume-%. After the calcination at 450 C., the reactor was cooled to a temperature of approximately 30 C.

(58) Epoxidation

(59) Continuous operation according to Reference Example 2 was resumed and carried out for approximately 1000 h.

Example 2.2: ZnTiMWW Catalyst Regeneration According to the Invention Including Two Washing Steps

(60) The partial ZnTiMWW catalyst regeneration was performed based on the epoxidation reaction setup as described in Reference Example 2 above.

(61) Epoxidation

(62) The main reactor was loaded with the ZnTiMWW catalyst according to Reference Example 1. The epoxidation reaction was continuously operated for approximately 1200 hours until the epoxidation reaction temperature, defined as the temperature of the heat transfer medium at the entrance of the jacket of the main reactor whose starting value at the beginning of the epoxidation reaction was 30 C., had reached a value of 50 C. At that point in time, the epoxidation reaction was stopped by stopping the flow of hydrogen peroxide; the acetonitrile and propene flows were continued until the epoxidation in the tubes was completed. Then, the reactor was flushed free of H.sub.2O.sub.2, propylene and propylene oxide by passing a mixture of acetonitrile and water through the reactor (80 weight-% acetonitrile, 20 weight-% water) wherein the temperature of this mixture was 50 C., followed by draining the major portion of the mixture of acetonitrile and water from the reactor.

(63) First Washing Step

(64) Subsequently, the reaction tubes were filled from below with water having a temperature of 70 C., and water having a temperature of 70 C. was then passed from the top downward through the tubes for about 6 h. The water wash was stopped once the conductivity of the outflowing water was below 200 microSiemens. The water was then drained from the reactor.

(65) Epoxidation

(66) Continuous operation according to Reference Example 2 was resumed and carried out for approximately 1300 h until the reaction temperature reached again the value of 50 C. At that point in time, the epoxidation reaction was stopped by stopping the flow of hydrogen peroxide; the acetonitrile and propene flows were continued until the epoxidation in the tubes was completed. Then, the reactor was flushed free of 1-1202, propylene and propylene oxide by passing a mixture of acetonitrile and water through the reactor (80 weight-% acetonitrile, 20 weight-% water) wherein the temperature of this mixture was 50 C., followed by draining the major portion of the mixture of acetonitrile and water from the reactor.

(67) Residual amounts of acetonitrile and water were removed from the reactor by passing nitrogen through the catalyst bed at a temperature of the nitrogen of 70 C.

(68) Second Washing Step

(69) Subsequently, the reaction tubes were filled from below with water having a temperature of 70 C., and water having a temperature of 70 C. was then passed from the top downward through the tubes for about 6 h. The water wash was stopped once the conductivity of the outflowing water was below 200 microSiemens. The water was then drained from the reactor.

(70) Calcining Step

(71) Then, nitrogen was passed through the catalyst bed for approximately 60 hours at an initial temperature of the nitrogen of 70 C. wherein this temperature was ramped to a value of 100 C. at 10 K/h. Subsequently, the catalyst was calcined by passing a gas mixture of air and nitrogen through the catalyst bed at a flow rate of 0.58 kg/s. The temperature of said mixture was further continuously ramped at 10 K/h from its initial value of 100 C. to a final value of 450 C. and then kept essentially constant at a value of 450 C. for 6 hours. At the beginning of the calcination, the gas mixture passed into the reactor had an oxygen content of 4 volume-%. Once having reached 450 C., said oxygen content was adjusted to a value of 21 volume-%. After the calcination at 450 C., the reactor was cooled to a temperature of approximately 30 C.

(72) Epoxidation

(73) Continuous operation according to Reference Example 2 was resumed and carried out for approximately 1000 h.

Comparative Example 2: ZnTiMWW Catalyst Regeneration Including One Washing Step

(74) Epoxidation

(75) The ZnTiMWW catalyst regeneration was performed based on the epoxidation reaction setup as described in Reference Example 2 above. The main reactor was loaded with the ZnTiMWW catalyst according to Reference Example 1. The epoxidation reaction was continuously operated for approximately 2200 hours until the epoxidation reaction temperature, defined as the temperature of the heat transfer medium at the entrance of the jacket of the main reactor whose starting value at the beginning of the epoxidation reaction was 30 C., had reached a value of 50 C. At that point in time, the epoxidation reaction was stopped by stopping the flow of hydrogen peroxide; the acetonitrile and propene flows were continued until the epoxidation in the tubes was completed. Then, the reactor was flushed free of H.sub.2O.sub.2, propylene and propylene oxide by passing a mixture of acetonitrile and water through the reactor (80 weight-% acetonitrile, 20 weight-% water) wherein the temperature of this mixture was 50 C., followed by draining the major portion of the mixture of acetonitrile and water from the reactor.

(76) Residual amounts of acetonitrile and water were removed from the reactor by passing nitrogen through the catalyst bed at a temperature of the nitrogen of 70 C.

(77) Washing Step

(78) Subsequently, the reaction tubes were filled from below with water having a temperature of 70 C., and water having a temperature of 70 C. was then passed from the top downward through the tubes for about 6 h. The water wash was stopped once the conductivity of the outflowing water was below 200 microSiemens. The water was then drained from the reactor.

(79) Calcining Step

(80) Then, nitrogen was passed through the catalyst bed for approximately 60 hours at an initial temperature of the nitrogen of 70 C. wherein this temperature was ramped to a value of 100 C. at 10 K/h. Subsequently, the catalyst was calcined by passing a gas mixture of air and nitrogen through the catalyst bed at a flow rate of 0.58 kg/s. The temperature of said mixture was further continuously ramped at 10 K/h from its initial value of 100 C. to a final value of 450 C. and then kept essentially constant at a value of 450 C. for 6 hours. At the beginning of the calcination, the gas mixture passed into the reactor had an oxygen content of 4 volume-%. Once having reached 450 C., said oxygen content was adjusted to a value of 21 volume-%. After the calcination at 450 C., the reactor was cooled to a temperature of approximately 30 C.

(81) Epoxidation

(82) Continuous operation according to Reference Example 2 was resumed and carried out for approximately 1000 h.

Results from Examples 2.1 and 2.2 and Comparative Example 2

(83) The results from examples 2.1 and 2.2 and comparative example 2 are discussed hereinafter and shown in Table 2 below. The inlet temperature of the heat transfer medium of epoxidation unit A both in Examples 2.1 and 2.2 and Comparative Example 2 was adjusted during the continuous operation according to Reference Example 2 in order to keep the hydrogen peroxide conversion essentially constant at approximately 96%. The selectivity S of the epoxidation reaction based on hydrogen peroxide as provided in Table 2 is defined as (n(PO)/n(H.sub.2O.sub.2))100 wherein n(PO) is the molar amount of propylene oxide detected directly downstream to unit B and n(H.sub.2O.sub.2) is the molar amount of hydrogen peroxide converted in the epoxidation reaction.

(84) TABLE-US-00002 TABLE 2 Example Example Comparative 2.1 2.2 Example 2 Epoxidation time before first 1300 1200 2200 washing step/h Epoxidation time after the first 1400 1300 until the second washing step/h Epoxidation time after the second 1200 until the third washing step/h Total epoxidation time until 3900 2500 2200 calcination/h Selectivity immediately before 97.2 97.1 96.8 calcination/% Selectivity 1000 hours after 97.3 97.4 97.2 calcination/%

(85) As clearly shown in Table 2, a process according to the invention wherein one washing step (example 2.2) and two washing steps (example 2.1) is/are carried out without a subsequent calcination step lead to a slight, but nevertheless existing improve in selectivity (1000 hours after calcination) compared to a common process (comparative example 2) according to which a calcination is carried out directly after the (only) washing step. This remarkable result is achieved although the total epoxidation time until the calcination step is higher (plus 300 hours with regard to example 2.2) or even significantly higher (plus 1700 hours with regard to example 2.1). Thus, it is convincingly shown that the inventive concept of a partial regeneration according to stage (i) which avoids a calcination step after each washing step and, thus, represents a significantly less complex process, unexpectedly leads to even improved results with regard to the most important parameter ofin particularan industrial-scale process, i.e. the selectivity with regard to the valuable product.

CITED LITERATURE

(86) WO 98/55229 A1 EP 1 221 442 A1 WO 2005/000827 A1 WO 2007/013739 A1 Ullmann's Encyclopedia of Industrial Chemistry, 5th edition, volume 13, pp. 447-456 Atlas of Zeolite Framework Types, 5th edition, Elsevier, London, England (2001) US 2007043226 A1 U.S. Pat. No. 6,114,551 Wu et al., Hydrothermal Synthesis of a novel Titanosilicate with MWW Topology, Chemistry Letters (2000), pp. 774-775 WO 2013/117536 A2