Olefin oxides, such as propylene oxide, are used to alkoxylate alcohols to form polyether polyols, such as polypropylene polyether polyols, which find significant utility in the manufacture of polyurethanes and synthetic elastomers. Olefin oxides are also important intermediates in the manufacture of alkylene glycols, such as propylene glycol and dipropylene glycol, and alkanolamines, such as isopropanolamine, which are useful as solvents and surfactants.

Propylene oxide is produced commercially via the well-known chlorohydrin process wherein propylene is reacted with an aqueous solution of chlorine to produce a mixture of propylene chlorohydrins. The chlorohydrins are dehydrochlorinated with an excess of alkali to produce propylene oxide.

Another well-known route to olefin oxides relies on the transfer of an oxygen atom from an organic hydroperoxide or peroxycarboxylic acid to an olefin. In the first step of this oxidation route, a peroxide generator, such as isobutane or acetaldehyde, is autoxidized with oxygen to form a peroxy compound, such as t-butyl hydroperoxide or peracetic acid.

This compound is used to epoxidize the olefin, typically in the presence of a transition metal catalyst, including titanium, vanadium, molybdenum, and other heavy metal compounds or complexes.

Gas phase processes for the direct oxidation of olefins by molecular oxygen to the corresponding olefin have also been described in several publications.

The former East German patent DD 212 961 discloses a process for the oxidation of olefins containing 3 to 20 carbon atoms with oxygen-containing gases in liquid phase in the presence of transition metal complexes. The catalyst preferably is a mixture of a Cu (II) complex compound and a complex compound of molybdenum or wolfram. The oxidation is carried out at a temperature of from 20°C to 200°C at atmospheric pressure or at an elevated pressure of 2 to 10 MPa. Recommended solvents are benzene, chlorobenzene, o-dichlorobenzene, nitrobenzene or bromobenzene. The desired products are obtained by

fractionated distillation of the product mixture. Unfortunately, the patent does not teach how to separate the produced olefin oxide from unconverted raw material, by-products and solvent. However, it is well known in the art that such separation is very complex.

Particularly the example given for propylene oxide and using benzene as the process solvent would lead to an impractical separation scheme and overall uneconomic process due to the azeotropic behavior of the mixture.

U. S. Patent No. 3,238,229 relates to a process for preparing olefin oxides wherein an olefinically unsaturated hydrocarbon is oxidized with molecular oxygen at a temperature of from 50°C to 400°C and at a pressure of from 0.5 to 150 atmospheres in halogenated benzenes, such as o-dichlorobenzene, as a solvent. The desired products are recovered from the reactor effluent by conventional separation techniques, such as distillation.

U. S. Patent No. 4,420,625 relates to a process for preparing alkylene oxides with oxygen in the presence of a transition metal borate catalyst and a non-polar, aromatic organic solvent. The U. S. patent does not teach how to recover the alkylene oxide from the resulting product mixture.

U. S. Patent No. 3,350,418 discloses a liquid phase oxidation of propene with molecular oxygen in an ester, such as a fully esterified polyacyl ester, as a solvent with the following subsequent separation steps: a) the effluent stream of the reaction mixture is passed from a reaction zone through a combination let-down distillation zone which comprises a flashing zone followed by a stripping zone into which the bottoms from said flashing zone is passed; the flashing and stripping zone are maintained at the same temperature as the oxidation reaction, but each successive zone is maintained at pressures substantially lower than in the preceding zone and in the reaction zone to separate substantially all of the low and intermediate boiling products, including propylene oxide, as gas phase from the bulk of the solvent; in steps b) and c) the gas phase is condensed and further treated; in step d) a combined stream of condensed liquids from the condensing zones in step b) and c) is passed into an acid-separation distillation zone where organic acids are removed as bottoms and propylene oxide, propylene, propane, acetaldehyde and methyl formate are distille overhead;

in step e) the overhead from step d) is subjected to a distillation step where propylene and propane are distille overhead and propylene oxide, acetaldehyde and methyl formate are removed as bottoms; in step f) acetaldehyde is removed by distillation; and in step g) propylene oxide is separated from methyl formate by extractive distillation using a hydrocarbon solvent.

Unfortunately, this process requires large equipment sizes and high energy consumption, as will be explained in more detail with reference to Fig. 1 of U. S. Patent No.

3,350,418. All unconverted propene has to pass the whole separation scheme consisting of flasher and columns 26,28 and 30 which leads to uneconomical equipment sizes and energy consumption. All recycle propere is evaporated and taken as overhead stream in column 30, so conventional compressors will be needed to recycle the propene. No convenient pressure profile for the separation is applied, so that the overhead stream 27 of column 26 containing all propene and low boiling products such as propylene oxide has to be compressed from 10.5 bar to 21 bar (1.05 to 2.1 MPa) which will result in a large compression unit operation or using of a large amount of refrigeration for liquefying a portion of the stream. Column 29 is stated to operate at 21 bar (2.1 MPa) and 160°C bottom temperature. All propylene oxide and low boiling reaction products has to pass this hot zone where yield losses of propylene oxide will occur due to reaction with other by-products.

Despite this, vent gases, particularly CO2, will not be removed using this separation scheme.

Most of the CO2 will be removed together with propene in stream 25 of the flasher and will accumulate in stream 35 recycle propene, so that the CO2 cannot be removed in the absorber 20.

U. S. Patent No. 3,071,601 relates to the production of propylene oxide wherein propylene is oxidized with elemental oxygen in the liquid phase in a hydrocarbon solvent at elevated pressure and temperature in the presence of a catalyst. According to the Example, the reactor temperature is about 200°C and the pressure is 750 p. s. i. g. (51.5 bar or 5.15 MPa). The reaction mixture is fed into a flash tank 13 wherein the temperature and pressure are allowed to drop to 60°C and 65 p. s. i. g. (4.5 bar or 0.45 MPa). The non-reacted propylene and highly volatile products pass off and are recycled to the reactor. The liquid material within the flash tank is passed to a distillation column 14 from the top of which there is distille off propylene oxide and methyl formate. Propylene oxide and methyl formate are

fed to a further distillation column wherein methyl formate is taken off overhead by means of an azeotrope-former, such as n-pentane. The less volatile materials recovered as the still bottoms in the distillation column 14 comprise various acids, non-acidic polymer, benzene diluent and propylene glycol which are separated by a subsequent distillation in a distillation column 17 and an extraction step. In this distillation step, volatile acids and the benzene solvent are recovered overhead. However, in the flash tank 13 all propene is evaporated and has to be recycled, which requires an uneconomically large compression unit. The recycle of propene and highly volatile organics directly to the reactor includes the carry over of formic acid which is detrimental to product selectivity. Moreover, the suggested distillation in distillation column 17, wherein volatile acids and the benzene solvent are recovered overhead, requires uneconomical large equipment sizes and high energy consumption. The benzene also forms undesirable azeotropes, making the separation complex and leading to uneconomic solvent losses.

Considering that the production of olefin oxides, such as propylene oxide, from the corresponding olefin is carried out on a very large scale and in view of the disadvantages of the processes of the prior art, it would still be desirable to provide a liquid phase process for preparing an olefin oxide from an olefin which allows an efficient separation of the products in the product mixture.

Accordingly, the present invention relates to a liquid phase process for preparing an olefin oxide from an olefin comprising the steps of: a) contacting the olefin in a solvent having a boiling point above 130°C with oxygen or an oxygen-containing gas, b) passing the liquid product mixture comprising olefin oxide, non-converted olefin, solvent and by-products into a first separator wherein a lower pressure is maintained than in the reaction step a) and the product mixture is divided into b1) a gaseous stream (12) containing olefin, carbon dioxide and other volatile products; and b2) a liquid stream (13) containing olefin oxide, solvent, olefin, carbon dioxide and other by-products;

c) passing the liquid stream (13) into a second separator wherein a lower pressure is maintained than in the first separator in step b) and the stream is divided into c1) a gaseous stream (2 or 17) containing olefin, carbon dioxide and other products; and c2) a liquid stream (3 or 16) containing olefin oxide, solvent, olefin and oxygenated by-products.

The flow sheets in Figures 1,2 and 3 are graphical illustrations of preferred embodiments of the process of the present invention.

Ethylene can be employed in the process of this invention, however the olefin preferably contains three or more carbon atoms. Undiluted olefins or mixtures thereof are preferably used, however also olefin feedstock can be used which contains up to 50 weight percent of saturated compounds. Monoolefins are preferred, but compounds containing two or more olefins, such as dienes, can also be used. The olefins can be aliphatic or alicyclic.

The olefin can be a simple hydrocarbon containing only carbon and hydrogen atoms; or alternatively, the olefin can be substituted at any of the carbon atoms with an inert substituent. The term"inert", as used herein, requires the substituent to be non-reactive in the process of this invention. Suitable inert substituents include, but are not limited to, halides, ether, ester, alcool, or aromatic moieties, preferably chloro, Ct, 2-ether, ester, or alcohol moieties or C-aromatic moieties. Non-limiting examples of olefins which are suitable for the process of this invention include propylene, 1-butene, 2-butene, 2-methylpropene, 1-pentene, 2-pentene, 2-methyl-1-butene, 2-methyl-2-butene, 1-hexene, 2-hexene, 3-hexene, and analogously, the various isomers of methylpentene, ethylbutene, heptene, methylhexene, ethylpentene, propylbutene, the octenes, including preferably 1-octene, and other higher analogues of these; as well as butadiene, cyclopentadiene, dicyclopentadiene, styrene, a-methylstyrene, divinylbenzene, allyl chloride, allyl alcool, allyl ether, allyl ethyl ether, allyl butyrate, allyl acetate, allyl benzene, allyl phenyl ether, allyl propyl ether, and allyl anisole. Preferably, the olefin is an unsubstituted or substituted C3, 2-olefin, more preferably, an unsubstituted or substituted C38-olefin. Most preferably, the olefin is propylene. Propylene feedstock can be used which contains up to 50 weight percent propane, however the use of undiluted propylene is preferred. Accordingly, the subsequent detailed description of the present invention often relates to a process

wherein propylene is used as a starting material, although the process of the present invention is not limited thereto.

The quantity of olefin employed in the process can vary over a wide range provided that the corresponding olefin oxide is produced. Generally, the quantity of olefin depends upon the specific process features, including, for example, the design of the reactor, the specific olefin, and economic and safety considerations. Those skilled in the art will know how to determine a suitable range of olefin concentrations for the specific process features. Typically, on a molar basis an excess of olefin is used relative to the oxygen. This condition enhances the selectivity to olefin oxide and minimizes the selectivity to combustion products, such as carbon dioxide. The quantity of the olefin is typically greater than 1, preferably greater than 10, more preferably greater than 20, and most preferably greater than 25 mole percent, based on the total moles of olefin, oxygen and solvent. Typically, the quantity of the olefin is less than 99, preferably less than 85, more preferably, less than 70, and most preferably less than 65 mole percent, based on the total moles of olefin, oxygen and solvent.

The olefin is contacted with oxygen, such as essentially pure molecular oxygen, or an oxygen-containing gas, such as air or oxygen diluted with carbon dioxide. If the olefin is contacted with an oxygen-containing gas, the oxygen concentration in the gas preferably is from 15 to 60 volume percent, more preferably from 20 to 55 volume percent.

Other sources of oxygen may be suitable, including ozone and nitrogen oxides, such as nitrous oxide. Molecular oxygen or oxygen diluted with carbon dioxide is preferred. The quantity of oxygen employed can vary over a wide range provided that the quantity is sufficient for producing the desired olefin oxide. Ordinarily, the number of moles of oxygen per mole of olefin is less than 1. Generally, the quantity of oxygen is greater than 0.01, preferably greater than 1, and more preferably greater than 2 mole percent, based on the total moles of olefin, oxygen, and solvent. Generally, the quantity of oxygen is less than 30, preferably less than 25, more preferably less than 20, and most preferably less than 15 mole percent, based on the total moles of olefin, oxygen and solvent.

Step a) of the process of the present invention is carried out in a liquid phase in a solvent which has a boiling point above 130°C, at atmospheric pressure preferably above 150°C, more preferably above 170°C. Preferred solvents are halogenated benzenes, particularly monohalogenated benzenes and, more preferably, dihalogenated benzenes.

Exemplary thereof are monobromobenzene, chlorobenzene, o-, m-or p-dibromobenzene, o-

, m-or p-bromochlorobenzene, or, most preferably, o-, m-or p-dichlorobenzene. The most preferred solvent for the process of the present invention is o-dichlorobenzene. Other suitable solvents are polyethers, polyesters, polyalcohols or halogenated, preferably chlorinated, aliphatic alcohols, such as 2-chloro-1-propanol, 3-chloro-1-propanol, 1-bromo-2- propanol, dichloro-or dibromopropanols, provided that these solvents are liquid at the chosen reaction conditions and have a boiling point above 130°C. The amount of solvent is typically greater than 0.1, preferably, greater than 15, and more preferably greater than 25 mole percent, based on the total moles of olefin, oxygen and solvent. The amount of solvent is typically less than 90, preferably, less than 80, and more preferably, less than 75 mole percent, based on the total moles of olefin, oxygen and solvent.

Step a) of the process of the present invention can be carried out in the absence or in the presence of a catalyst. If a catalyst is used, homogeneous or heterogeneous catalysts are useful. Exemplary of heterogeneous catalysts are those disclosed in East German Patent Nos. DD-218,100; DD-213,436; DD-218,099; DD-212,959; DD-212,960; DD-212 902; U. S. Patent No. 3,957,690 and in published PCT application WO 96/20788. The disclosed heterogeneous catalysts contain active components, such as nickel, manganese, molybdenum and/or vanadium containing complexes or salts, on a carrier, such as an alumina, silica, aluminosilicate, titania, magnesia and/or carbon. As homogeneous oxidation catalysts common complexes or salts can be used, for examples those disclosed in U. S. Patent Nos. 3,505,359; 3,518,285; or 4,420,625 or in WO 96/37295.

Typical active components of these catalysts are elements of groups lb, llb, Illb, IVb, Vb, Vlb, Vllb, Illa, IVa, Va, Vla, VIII, and/or of the lanthanide group, such as Mo, Mn, Wo, Zn, Re, Au, Pd, Ag, V, Ru, La and/or Ti in any combination and ratio, but also Sc, Y, Ce, Zr, Nb, Ta, Cr, Fe, Os, Co, Rh, Ir, Ni, Pd, Pt, Cu, Ga, In, Ge, Sn, Se, Te, As, Sb and/or Bi in any combination and ratio with those mentioned above.

Furthermore, step a) of the process of the present invention can be carried out in the absence or in the presence of a promoter. A promoter is for example an aldehyde, like acetaldehyde, or an alkaline additive, like hydroxides of the group 1a and 1b, such as sodium hydroxide or magnesium hydroxide. The total quantity of promoter metal (s) generally is greater than 0.01, preferably, greater than 0.10, and more preferably, greater than 0.15 weight percent, based on the total weight of the catalyst. The total quantity of promoter metal (s) is generally less than 20, preferably, less than 15, and more preferably less than 10 weight percent, based on the total weight of the catalyst.

Step a) of the process of this invention can be conducted in one or more reactors of any conventional design suitable for liquid phase processes. These designs broadly include batch, fixed-bed, transport bed, fluidized bed, moving bed, shell and tube, and trickle bed reactors, as well as continuous and intermittent flow and swing reactor designs. Preferred reactor types are plug flow reactors, cascades of stirred tank reactors or bubble columns. Cascades of stirred tank reactors preferably contain from 2 to 10, more preferably from 2 to 5 stirred tank reactors.

Preferably, process step a) is conducted at a temperature of from 100°C to 210°C, more preferably from 130°C to 190°C, most preferably from 150°C to 180°C.

Preferably, the pressure ranges from 20 to 100 bar (2 to 10 MPa), more preferably from 40 to 80 bar (4 to 8 MPa), most preferably from 50 to 60 bar (5 to 6 MPa) absolute.

The liquid product mixture obtained in step a) of the present invention generally contains from 30 to 90 percent, typically from 40 to 70 percent of solvent, generally from 0.2 to 6.0 percent, typically from 0.5 to 4.5 percent of olefin oxide, generally from 10 to 60 percent, typically from 20 to 40 percent of non-converted olefin, based on the total weight of the liquid product mixture, the remaining amount being by-products, such as carbon dioxide, or oxygenated by-products, like water, aldehydes, such as acetaldehyde; ketones, such as acetone or 1-hydroxy-2-propanone; alcools such as 2- propanole or allyl alcohol; glycols, such as 1,2-propane diol; esters, such as methyl formate, methyl acetate or 1,2-propane glycol diacetate; acids, such as formic acid, acetic acid or propionic acid; acetals such as 2-ethyl-4-methyl-1,3-dioxolane, 2,4-dimethyl-1, 3-dioxolane or or ethers such as 1-methoxy-2- propanone.

Steps b) and c) of the process of the present invention and optional further separation steps are described in more detail with reference to Figures 1-3.

Now referring to Fig. 1, in step b) of the process of the present invention the liquid product mixture 1 which is obtained in step a) and which comprises olefin oxide, non- converted olefin, solvent and by-products is passed into a first separator F-1 wherein a lower pressure is maintained than in the reaction step a) and the product mixture is divided into b1) a gaseous stream 12 containing olefin, carbon dioxide and other volatile products; and b2) a liquid stream 13 containing olefin oxide, solvent, olefin, carbon dioxide and other by-products. Preferably, separation step b) is conducted at a temperature of from

100°C to 180°C, more preferably from 130°C to 180°C, most preferably from 130°C to 170°C. Preferably, the pressure in step b) ranges from 19 to 95 bar (1.9 to 9.5 MPa), more preferably from 39 to 78 bar (3.9 to 7.8 MPa), most preferably from 48 to 59 bar (4.8 to 5.9 MPa) absolute, provided that the pressure is lower than the pressure in process step a).

The separator F-1 used in step b) can be any known separator with or without inserts, such as a distillation column. The separator F-1 can also be a part of the reaction design and located inside the reactor of step a). However, most preferably a flash separator is used.

The weight ratio between the gaseous stream 12 and the liquid stream 13 preferably is from 0.03 to 1.0: 1, more preferably from 0.1 to 0.4: 1. The use of a flash separator F-1 in step b) is particularly useful if the olefin used in step a) is propylene and the solvent is 1,2-dichlorobenzene. An excellent separation between propylene and 1,2-dichlorobenzene is achieved in the flash separator. Depending on the pressure and other conditions in the flash evaporator, the separation between propylene and 1,2-dichlorobenzene is generally from 3 to 30 times better than the separation between propylene and benzene. Generally from 5 to 70 percent, preferably from 30 to 50 percent of propylene is comprised in the gaseous stream 12 after the separation step b), based on the total amount of non-reacted propylene in the product mixture obtained in step a). Because a large amount of non- reacted propylene can be separated from the solvent and the produced olefin oxide in the first separation step b), equipment required for the subsequent separation step c) and optional further separations steps can be minimized.

The gaseous stream 12 contains non-reacted olefin, carbon dioxide, non-converted oxygen and minor portions of other volatile products, such as propylene oxide and acetaldehyde. The gaseous stream preferably contains from 40 to 90, more preferably from 60 to 80 weight percent of non-reacted olefin and preferably from 10 to 60, more preferably from 20 to 40 weight percent of volatile products. The gaseous stream can be directly recycled to the reaction step a) to recycle non-reacted olefin and to make use of carbon dioxide to dilute oxygen used in the reaction step a). However, non-reacted olefin is preferably separated from the gaseous stream and recycled to the separation step a). A preferred separation step is a partial condensation to remove reaction products from the olefin recycle.

The liquid stream 13 contains olefin oxide, solvent, olefin, carbon dioxide and other by-products, generally oxygenated by-products such as those listed above. The liquid stream preferably contains from 0.2 to 6, more preferably from 0.6 to 4.5 weight percent of

olefin oxide, preferably from 50 to 85, more preferably from 60 to 80 weight percent of solvent, preferably from 10 to 35, more preferably from 15 to 22 weight percent of olefin and preferably from 0.2 to 8, more preferably from 1.5 to 5.5 weight percent of by-products, such as carbon dioxide.

The molar ratio between carbon dioxide in the gaseous stream 12 and in the liquid stream 13 generally is from 1 to 4: 1, typically from 1.4 to 2: 1. The molar ratio between olefin in the gaseous stream 12 and in the liquid stream 13 generally is from 0.4 to 1.6: 1, typically from 0.5 to 1: 1.

Although very efficient product separations are achieved in the separation steps of the present invention, it is to be understood that in none of the separation steps described herein a 100 percent separation is achieved. It is to be understood that the gaseous streams 12 and 2 and the liquid streams 13 and 3 can contain product components in addition to those specifically listed herein.

In step c) of the process of the present invention the liquid stream 13 is passed into a second separator F-2 wherein a lower pressure is maintained than in the first separator F-1 in step b) and the stream is divided into c1) a gaseous stream 2 containing olefin, carbon dioxide and other products and c2) a liquid stream 3 containing olefin oxide, solvent and oxygenated by-products. Preferably, separation step c) is conducted at a temperature of from 40°C to 180°C, more preferably from 50°C to 160°C, most preferably from 60°C to 150°C. Preferably, the pressure in step c) ranges from 4 to 70 bar (0.4 to 7 MPa), more preferably from 8 to 50 bar (0.8 to 5 MPa), most preferably from 10 to 45 bar (1 to 4.5 MPa) absolute, provided that the pressure is lower than the pressure in the separation step b). The separator used in step c) can be any known separator with or without inserts, such as a distillation column. However, most preferably a flash separator or evaporator is used.

The weight ratio between the gaseous stream 2 and the liquid stream 3 preferably is from 0.05 to 1: 1, more preferably from 0.05 to 0.5: 1, most preferably from 0.1 to 0.3: 1.

The gaseous stream 2 contains non-reacted olefin, carbon dioxide and other volatile products, such as propylene oxide, acetaldehyde, methyl formate, water, methanol or acrolein. The gaseous stream generally contains from 70 to 85 weight percent, typically from 75 to 80 weight percent, of non-reacted olefin and generally from 15 to 30 weight

percent, typically from 20 to 25 weight percent of volatile products. The gaseous stream can be directly recycled to the reaction step a) to recycle non-reacted olefin and to make use of carbon dioxide to dilute oxygen used in the reaction step a). According to a preferred embodiment of the process of the present invention the liquid stream 3 is subjected to a further separation step as described further below.

The liquid stream 3 contains olefin oxide, solvent, non-reacted olefin and oxygenated by-products, such as those listed further above, and optionally minor amounts of carbon dioxide. The liquid stream preferably contains from 0.1 to 6, more preferably from 0.6 to 4.5 weight percent of olefin oxide, preferably from 60 to 95, more preferably from 70 to 90 weight percent of solvent, preferably from 5 to 50, more preferably from 15 to 25 weight percent of non-reacted olefin, and preferably from 0.8 to 10, more preferably from 1.5 to 8 weight percent of oxygenated by-products.

According to a preferred embodiment of the process of the present invention the gaseous stream 2 and the overhead stream 6 of separator C-2, as described further below, are passed to a further separator C-1, wherein the combined stream is separated into an overhead stream 4 and a bottom stream 5. Preferably, the separation step in separator C-1 is conducted at a temperature of from 80°C to 150°C, more preferably from 90°C to 130°C at bottom stage and from-10°C to +70°C, more preferably from 0°C to 60°C at the top stage and at a pressure of preferably from 5 to 20 bar (0.5 to 2 MPa), more preferably from 8 to 14 bar (0.8 to 1.4 MPa) absolute. The separator can be any known separator with or without inserts, preferably a distillation column. The weight ratio between the overhead stream 4 and the bottom stream 5 resulting from this separation preferably is from 3 to 7: 1.

The overhead effluent stream 4 from the separator C-1 contains non-reacted olefin, carbon dioxide and may contain minor amounts of olefin oxide. The overhead stream 4 preferably contains from 80 to 98 weight percent, more preferably from 85 to 95 weight percent of non-reacted olefin and preferably from 2 to 20, more preferably from 5 to 15 weight percent of volatile products, such as carbon dioxide and olefin oxide. The overhead stream 4 can be directly recycled to the reaction step a) to recycle non-reacted olefin and to make use of carbon dioxide to dilute oxygen used in the reaction step a). Alternatively, non- reacted olefin can be separated from the overhead stream 4 and is preferably recycled to the reaction step a).

The bottom effluent stream 5 from the separator C-1 mainly contains crude olefin oxide. The bottom stream 5 preferably contains from 30 to 90, more preferably from 50 to 80 weight percent of olefin oxide, preferably from 1 to 9, more preferably from 5 to 7 weight percent of methyl formate, preferably from 1 to 9, more preferably from 4 to 6 weight percent of water, preferably from 1 to 9, more preferably from 4 to 6 weight percent of acetaldehyde, preferably from 0.5 to 10, more preferably from 2 to 5 weight percent of olefin and preferably from 1 to 25, more preferably from 5 to 15 weight percent of other oxygenated products like acetone, acroleine, or methanol. From the bottom stream 5 propylene oxide can be isolated from the other oxygenated by-products.

Preferably, the liquid effluent stream 3 from the separator F-2 is passed to a further separator C-2, wherein the stream 3 is separated into a overhead stream 6 and a bottom stream 7. Preferably, the separation step in separator C-2 is conducted at a temperature of from 0 to 80°C, more preferably from 0 to 60°C at the top stage and from 120°C to 220°C, more preferably from 140°C to 200°C at the bottom stage and at a pressure of from 0.1 to 10 bar (0.01 to 1 MPa), more preferably from 0.9 to 4 bar (0.09 to 0.4 MPa), most preferably from 1 to 2 bar (0.1 to 0.2 MPa) absolute. The separator can be any known separator with or without inserts, preferably a distillation column. The weight ratio between the overhead stream 6 and the bottom stream 7 resulting from this separation preferably is from 0.01 to 1: 1, more preferably from 0.03 to 0.2: 1.

The overhead stream 6 from the separator C-2 preferably contains from 5 to 65, more preferably from 15 percent to 30 weight percent of olefin oxide, preferably from 15 to 90, more preferably from 50 to 80 weight percent of non-reacted olefin, preferably from 0.1 to 20, more preferably from 2 to 10 weight percent of carbon dioxide and preferably from 0.1 to 25, more preferably from 2 to 10 weight percent of volatile oxygenated products like acetaldehyde or methyl formate. The overhead stream 6 is preferably directed to separator C-1, where crude olefin oxide is separated from olefin. The bottom effluent stream 7 from the separator C-2 contains solvent and oxygenated by-products. The bottom effluent stream 7 preferably contains from 70 to 98, more preferably from 85 to 95 weight percent of solvent and preferably from 2 to 30, more preferably from 5 to 15 weight percent of oxygenated by- products. The separation in separator C-2 is particularly effective if the olefin used in step a) is propylene and the solvent is 1,2-dichlorobenzene. An excellent separation between propylene oxide and 1,2-dichlorobenzene is achieved. Depending on the pressure and other conditions in the separator C-2, the separation between propylene oxide and

1,2-dichlorobenzene is generally from 1.7 to 11 times better than the separation between propylene oxide and benzene.

The bottom effluent stream 7 is preferably passed to a further separator C-3, wherein the stream is separated into an overhead stream 8 and a bottom stream 9.

Preferably the separation step in separator C-3 is conducted at a temperature of from 100°C to 180°C, more preferably from 100°C to 160°C at the top stage and from 150°C to 220°C, more preferably from 160°C to 210°C at the bottom stage and at a pressure of from 0.1 to 4 bar (0.01 to 0.4 MPa), more preferably from 0.8 to 2.5 bar (0.08 to 0.25 MPa) absolute. The separator can be any known separator with or without inserts, preferably a distillation column. The weight ratio between the overhead stream 8 and the bottom stream 9 resulting from this separation preferably is from 0.005 to 0.5: 1, more preferably from 0.01 to 0.1: 1.

The overhead stream 8 from the separator C-3 mainly contains oxygenated by-products which have a lower boiling point than the solvent, such as 2-propanol, allyl alcohol, 1-methoxy-2-propanone, 2,4-dimethyl-1, 3-dioxolane, water, formic acid or acetic acid. The bottom effluent stream 9 from the separator C-3 mainly contains solvent and a minor amount of high-boiling oxygenated by-products, such as propylene glycols or propylene glycol esters. The bottom effluent stream 9 generally contains 90 to 99 weight percent of solvent and from 1 to 10 weight percent of high-boiling oxygenated by-products.

It is preferred to recycle the major amount of the bottom effluent stream 9 to the reaction step a) as stream 10. In order to avoid accumulation of high-boiling oxygenated by-products in the reaction, it is preferred to separate a small side-stream 11 from the bottom effluent stream 9, preferably from 0.1 to 50 volume percent, more preferably from 5 to 25 volume percent. The side stream 11 can be passed to a further separator C-4, preferably a distillation column, wherein the solvent is separated from the high-boiling oxygenated by-products. Preferably, the separation step in separator C-4 is conducted at a temperature of from 140°C to 230°C, more preferably from 160°C to 210°C, and at a pressure of from 0.1 to 4 bar (0.01 to 0.4 MPa), preferably from 0.5 to 1.5 bar (0.05 to 0.15 MPa) absolute. The overhead effluent stream 20 from the separator C-4 contains the solvent and may be recycled to the reaction step a). The bottom effluent stream 21 containing high-boiling oxygenated by-products can be reused for other purposes.

We refer now to Fig. 2, which illustrates a more preferred embodiment of the present invention. Fig. 2 illustrates a similar separation process as the one illustrated in Fig.

1, except that the liquid stream 3 of separator F-2 is passed into an additional separator F-3 wherein a lower pressure is maintained than in separator F-2. In separator F-2 the liquid stream 3 is divided into a gaseous stream 14 containing olefin, olefin oxide, usually carbon dioxide and other volatile by-products like acetaldehyde or methyl formate; and a liquid stream 15 containing olefin oxide, solvent, oxygenated by-products, such as those listed further above and optionally minor amounts of olefin. Preferably, separation step F-3 is conducted at a temperature of from 0°C to 180°C, more preferably from 20°C to 160°C, most preferably from 40°C to 150°C. Preferably, the pressure in step c) ranges from 0.8 to 45 bar (0.08 to 4.5 MPa), more preferably from 1 to 12 bar (0.1 to 1.2 MPa), most preferably from 2 to 7 bar (0.2 to 0.7 MPa) absolute, provided that the pressure is lower than the pressure in the separation step c). The separator F-3 can be any known separator with or without inserts, such as a distillation column. However, most preferably a flash separator or evaporator is used. The weight ratio between the gaseous stream 14 and the liquid stream 15 preferably is from 0.005 to 0.6: 1, more preferably from 0.02 to 0.2: 1.

The gaseous stream 14 from the separator F-3 preferably contains from 40 to 95, more preferably from 60 to 80 weight percent of non-reacted olefin, preferably from 2 to 40, more preferably from 8 to 25 weight percent of olefin oxide and preferably from 1 to 35, more preferably from 5 to 20 weight percent of other volatile products like carbon dioxide, acetaldehyde or methyl formate.

The liquid stream 15 from separator F-3 preferably contains from 0.1 to 10 weight percent, more preferably from 0.5 to 4.5 weight percent of olefin oxide, preferably from 50 to 99 weight percent, more preferably from 75 to 95 weight percent of solvent, preferably from 0.1 to 10 weight percent, more preferably from 0.1 to 5 weight percent of non-reacted olefin and preferably from 0.1 to 10, more preferably from 1.0 to 8 weight percent of oxygenated by-products.

According to a preferred embodiment of the present invention the gaseous stream 14 is passed to a further separator C-1, wherein the combined stream of the gaseous stream 14 from separator F-3, the gaseous stream 2 of separator F-2 and the overhead stream 6 from separator C-2 is separated into an overhead stream 4 and bottom stream 5, as described with reference to Fig. 1.

According to a preferred embodiment of the process of the present invention the liquid stream 15 from separator F-3 is passed to a further separator C-2. The liquid

stream 15 from separator F-3 is separated in separator C-2 in an overhead stream 6 and a bottom stream 7. The separator used and the conditions of temperature and pressure for separator C-2 are the same as described in Fig 1. Due to the fact that some of the olefin and olefin oxide have been removed via the gaseous stream 14 before stream 15 enters the separator C-2, the weight ratio between the overhead stream 6 and the bottom stream 7 and the composition of the overhead stream 6 are different from those stated with reference to Fig. 1. The weight ratio between the overhead stream 6 and the bottom stream 7 resulting from this separation preferably is from 0.01 to 0.5, more preferably from 0.03 to 0.15: 1. The overhead stream 6 from the separator C-2 preferably contains from 5 to 70, more preferably from 15 to 40 weight percent of olefin oxide, preferably from 15 to 85, more preferably from 40 to 65 weight percent of non-reacted olefin, preferably from 0.1 to 30, more preferably from 5 to 15 weight percent of carbon dioxide and preferably from 1 to 30 more preferably from 5 to 15 weight percent of volatile oxygenated products like acetaldehyde or methyl formate. The overhead stream 6 is preferably directed to separator C-1, where the crude olefin oxide is separated from the olefin.

By making use of the separator F-3, olefin and olefin oxide is removed from the liquid stream 3 after separator F-2. Due to this additional separation step, the weight of the overhead stream 6 from separator C-2 can generally be reduced to a level of from 30 to 45 percent, often from 35 to 40 percent, as compared to the weight of the overhead stream 6 in Fig. 1, where no separator F-3 is used. The significant reduction saves equipment and energy costs of separator C-2.

We refer now to Fig. 3, which illustrates the most preferred embodiment of the present invention. Fig. 3 illustrates a similar separation process as the one illustrated in Fig.

2, except that an additional separator F-4 and an additional separator F-5 are added.

The liquid stream 13 of separator F-1 is passed into an additional separator F-4 wherein a lower pressure is maintained than in separator F-1 and the stream 13 is divided into a gaseous stream 17 containing non-reacted olefin, carbon dioxide and other products, such as propylene oxide, acetaldehyde, methyl formate or water, and into a liquid stream 16 containing olefin oxide, solvent, non-reacted olefin and oxygenated by-products.

Preferably, separation step F-4 is conducted at a temperature of from 40°C to 180°C, more preferably from 50°C to 160°C, most preferably from 60°C to 150°C. Preferably, the pressure in step c) ranges from 4 to 70 bar (0.4 to 7 MPa), more preferably from 8 to 50 bar (0.8 to 5 MPa), most preferably from 10 to 45 bar (1 to 4.5 MPa) absolute, provided that the

pressure is lower than the pressure in the separation step b). The separator F-4 can be any known separator with or without inserts, such as a distillation column. However, most preferably a flash separator or evaporator is used. The weight ratio between the gaseous stream 17 and the liquid stream 16 preferably is from 0.005 to 0.5: 1, more preferably from 0.02 to 0.15: 1.

The liquid stream 16 from separator F-4 preferably contains from 0.1 to 6, more preferably from 0.6 to 4.5 weight percent of olefin oxide, preferably from 50 to 95, more preferably from 65 to 95 weight percent of solvent, preferably from 5 to 50, more preferably from 10 to 20 weight percent of non-reacted olefin, and preferably from 0.1 to 10, more preferably from 1.5 to 8 weight percent of oxygenated by-products, such as those listed further above.

The gaseous stream 17 from the separator F-4 preferably contains from 30 to 95, more preferably from 60 to 80 weight percent of non-reacted olefin, preferably from 1 to 35, more preferably from 10 to 25 weight percent of carbon dioxide, preferably from 0.05 to 10, more preferably from 1 to 7 weight percent of olefin oxide and preferably from 0.05 to 12, more preferably from 1 to 8 weight percent of volatile by-products like acetaldehyde, methyl formate or water.

According to a most preferred embodiment of the process of the present invention the gaseous stream 17 from separator F-4 is passed together with gaseous stream 12 of separator F-1 to an additional separator F-5. The combined streams 12 and 17 are separated into a gaseous stream 19 and liquid stream 18. Preferably, the separation step F-5 is conducted at a temperature of from 20°C to 180°C, more preferably from 40°C to 140°C, most preferably from 50°C to 100°C. Preferably, the pressure in step c) ranges from 4 to 60 bar (0.4 to 6 MPa), more preferably from 12 to 60 bar (1.2 to 6 MPa), most preferably from 45 to 55 bar (4.5 to 5.5 MPa) absolute, provided that the pressure is lower than the pressure in the separation step b). The separator F-5 can be any known separator with or without inserts, such as a distillation column. However, most preferably a combination of heat exchanger and flash separator is used to condense a portion of the combined gaseous streams 12 and 17. The weight ratio between gaseous stream 19 and liquid stream 18 preferably is from 0.01 to 2: 1, more preferably from 0.05 to 1: 1, most preferably from 0.1 to 0.5: 1.

The gaseous stream 19 from the separator F-5 contains non-reacted olefin, carbon dioxide, olefin oxide and minor amounts of other volatile products like acetaldehyde or methyl formate. The gaseous stream preferably contains from 30 to 95, more preferably from 60 to 80 weight percent of non-reacted olefin, preferably from 2 to 40, more preferably from 20 to 30 weight percent of carbon dioxide, preferably from 0.02 to 3, more preferably from 0.1 to 1.5 weight percent of olefin oxide and preferably less than 3, more preferably less than 1 weight percent of other volatile products like acetaldehyde or methyl formate.

The gaseous stream 19 can be directly recycled to make use of carbon dioxide as diluent or a portion of stream 19 can be used to remove some carbon dioxide from the process.

The preferred use of the separator F-5 has the advantage to remove olefin oxide and other reaction products from the recycle olefin stream which increases the yield and the selectivity for the olefin oxide.

According to a preferred embodiment of the process of the present invention the liquid stream 16 from separator F-4 and the liquid stream 18 from separator F-5 are passed to a further separator F-2 where they are separated into a gaseous stream 2 and a liquid stream 3. The separator used and the conditions of temperature and pressure for F-2 are the same described in Fig. 1.

The advantage of using the separator F-4 is that non-reacted olefin and carbon dioxide is removed into the gaseous stream 17 of separator F-4 at a high pressure before entering the separator F-2. As a result, the following process streams 2,3,4,6, and 15 contain smaller amounts of olefin and carbon dioxide saving equipment and energy costs for the separators F-2, F-3, C-1 and C-2.

The invention is illustrated by the following examples which should not be construed to limit the scope of the present invention. Unless stated otherwise all parts and percentages are given by mol.

Example 1 The example refers to Fig. 1.10.1 mol/hour of propylene was oxidized with 1.35 mol/hour oxygen and 4.5 mol/hour nitrogen (resulting in an oxygen concentration of 22.9 volume percent) at a temperature of 165°C, a pressure of 54 bar (5.4 MPa) gauge and at a residence time of 13 minutes in the presence of 5.65 mol/hour o-dichlorobenzene (boiling point: 180.4°C). The oxidation was carried out in a continuously stirred tank reactor.

The reactor outlet 1, consisting of propylene oxide, not converted propene, other reaction products and o-dichlorobenzene was fed continuously to the flash separator F-1 at a rate of 20.8 mol/hour. The composition of this mixture was: 2.4 percent propylene oxide, 27.1 percent o-dichlorobenzene, 42.9 percent propylene, 21.6 percent nitrogen, 1.5 percent carbon dioxide, in total 1.3 percent of acids such as, formic acid, acetic acid and propionic acid, 0.1 percent oxygen, 1.4 percent water and in total 1.7 percent of oxygenated by-products and impurities, including aldehydes, ketones, glycols, acetals or esters, such as acetaldehyde, acetone, 1,2-propane diol, methyl formate, 1-methoxy-2-propanol, 1-methoxy-2-propanone, 1-hydroxy-2-propanone or methyl acetate.

The flash separator F-1 was operated at a pressure of 53 bar (5.3 MPa) gauge and at a temperature of 151 °C. The vapor stream 12 from flash separator F-1 contained 55.1 percent propene, 37.3 percent nitrogen, 2.3 percent carbon dioxide, 1.7 percent of propylene oxide and small amounts of other low boiling components and oxygen at a rate of 11.5 mol/hour. It can either be directly recycled to the reactor or be used to remove small amounts of carbon dioxide and other inert components such as nitrogen.

The liquid stream 13 from flash separator F-1 was fed to the flash separator F-2 at a rate of 9.3 mol/hour. The flash separator F-2 was operated at a pressure of 12 bar (1.2 MPa) absolute and at a temperature of 67°C. The composition of stream 13 was 58.6 percent o-dichlorobenzene, 27.9 percent propylene, 3.2 percent propylene oxide, 2.1 percent nitrogen and the rest were other reaction products.

The gaseous stream 2 containing 70.4 percent propylene, 3.0 percent CO2, 24.4 percent nitrogen, 1.0 percent propylene oxide and other volatile reaction products was directed to column C-1 at a rate of 0.7 mol/hour, where propene and carbon dioxide were separated.

The liquid stream 3 of the flash separator F-2 with a composition of 63.5 percent o-dichlorobenzene, 24.3 percent propylene, 3.4 percent propylene oxide, 0.3 percent nitrogen, 0.4 percent carbon dioxide and the rest, other reaction products, was fed to column C-2 at a rate of 8.6 mol/hour, where propylene and crude propylene oxide was separated as overhead stream 6 at a rate of 2.6 mol/hour, containing 81.8 percent propylene, 11.5 percent propylene oxide, 0.8 percent nitrogen, 1.2 percent carbon dioxide and the rest, other reaction products. Column C-2 was operated at 1.6 bar (0.16 MPa)

absolute at the bottom stage, the distillate temperature was 1°C, the bottom temperature was 154°C. The ratio of distillate molar flow rate 6 to bottom molar flow rate 7 was 0.42: 1.

The gaseous stream 2 of separator F-2 and the overhead stream 6 of column C-2 were separated in column C-1 into the distillate stream 4 and the bottom stream 5. The column C-1 was operated at 11.6 bar (1.16 MPa) absolute at the top stage, the distillate temperature was 21 °C and the bottom temperature was 126°C. The ratio of distillate molar flow rate 4 to bottom molar flow rate 5 was 6.6: 1. The overhead stream 4, containing 91.2 percent propylene, 6.9 percent nitrogen and 1.8 percent carbon dioxide can be directly recycled to the reactor at a rate of 2.9 mol/hour.

The crude propylene oxide bottom stream 5 can be used at a rate of 0.4 mol/hour to isolate the propylene oxide using known distillation techniques such as extractive distillation with a composition of 69.9 percent propylene oxide, 0.9 percent water, in total 28.3 percent other volatile products such as acetaldehyde, acetone, acrolein, methyl formate, methanol, 2-propanol, and minor amounts of other reaction products such as methyl acetate, formic acid, acetic acid, 1-methoxy-2-propanol, 1-methoxy-2-propanone or 1-hydroxy-2-propanone.

The bottom stream 7 from column C-2, containing 90.4 percent 1,2-dichloro- benzene and 8.0 percent reaction products with a normal boiling point within a range of 55°C to 150°C was fed to column C-3 at a rate of 6.0 mol/hour to remove those reaction products in overhead stream 8 with a lower boiling point than the 1,2-dichlorobenzene. The column C-3 was operated at a pressure of 1.5 bar (0.15 MPa) absolute at the bottom stage, a distillate temperature of 107°C and a bottom temperature of 196°C. The ratio of distillate molar flow rate 8 to bottom molar flow rate 9 was 0.084: 1 The composition of stream 8 was 36 percent water, in total 42.9 percent acids such as formic acid, acetic acid and propionic acid, 5.9 percent 1,2-dichlorobenzene and totally 15.2 percent of other oxygenated by-products and impurities at a rate of 0.5 mol/hour.

Most of the reaction products were removed from the solvent in line 9, such that substantially all 1,2-dichlorobenzene from the bottom stream 9 of column C-3 could be directly recycled to the oxidation reactor as stream 10 at a rate of 4.8 mol/hour containing 98.3 percent 1,2-dichlorobenzene and 1.7 percent other high boiling oxygenated products such as, glycols or glycol esters such as 1,2-propylene glycol or 1,2-propylene glycol diacetate.

Only a small portion (14 percent) of stream 9 was fed in line 11 to column C-4 to remove high boiling products from the solvent to avoid accumulation. The column C-4 was operated at 0.95 bar (0.095 MPa) absolute at the top stage, the distillate temperature was 175°C, the bottom temperature was 184°C. 1,2-dichlorobenzene going overhead in line 20 was also recycled to the reactor at a rate of 0.8 mol/hour. The bottom effluent stream containing high boiling by-products and 1,2-dichlorobenzene can be re-used for other purposes at a rate of 0.01 mol/hour.

As a result of these separation steps, 0.4 mol/hour of crude propylene oxide were obtained with a composition of 69.9 percent propylene oxide and in total 31.1 percent of low boiling oxygenated by-products and impurities, such as acetaldehyde, methyl formate or methanol. 11.5 mol/hour of propene/nitrogen/carbon dioxide were separated at a high pressure of 53 bar (5.3 MPa) gauge with a composition of 55.0 percent propene, 37.3 percent nitrogen, 2.3 percent carbon dioxide, 1.7 percent of propylene oxide and small amounts of other low boiling components and oxygen. 2.9 mol/hour of propene with minor amounts of carbon dioxide were separated at a medium pressure of 11 bar (1.1 MPa) gauge. 1,2-dichlorobenzene was recycled with a flow rate of totally 5.5 mol/hour and a composition of 98.3 percent o-dichlorobenzene and 1.7 percent of other compounds.

Example 2 The example refers to Fig. 1.10.2 mol/hour of propylene was oxidized with 1.35 mol/hour oxygen at a temperature of 165°C, a pressure of 54 bar (5.4 MPa) gauge and at a residence time of 13 minutes in the presence of 5.7 mol/hour o-dichlorobenzene (boiling point: 180.4°C) and at a carbon dioxide recycle rate of 1.3 mol/hour. The oxidation was carried out in a continuously stirred tank reactor.

The reactor outlet 1, consisting of propylene oxide, not converted propene, other reaction products and o-dichlorobenzene, was fed continuously to the flash separator F-1 at a rate 17.6 mol/hour. The composition of this mixture was: 2.8 percent propylene oxide, 32.4 percent o-dichlorobenzene, 51.3 percent propylene, 8.0 percent carbon dioxide, in total 1.5 percent of acids such as formic acid, acetic acid and propionic acid, 0.1 percent oxygen, 1.7 percent water and in total 2.2 percent of oxygenated by-products and impurities, including aldehydes, ketones, glycols, acetals or esters, such as acetaldehyde, acetone, 1,2-propane diol, methyl formate, 1-methoxy-2-propanol, 1-methoxy-2-propanone, 1-hydroxy-2-propanone or methyl acetate.

The flash separator F-1 was operated at a pressure of 53 bar (5.3 MPa) gauge and a at temperature of 151 °C. The vapor stream 12 from flash separator F-1 contained 77.0 percent propene, 16.6 percent carbon dioxide, 1.9 percent of propylene oxide and small amounts of other low boiling components and oxygen at a rate of 4.8 mol/hour. It can either be directly recycled to the reactor or be used to remove small amounts of carbon dioxide.

The liquid stream 13 from flash separator F-1 was fed to the flash separator F-2 at a rate of 12.8 mol/hour. The flash separator F-2 was operated at a pressure of 12 bar (1.2 MPa) absolute and at a temperature of 67°C. The composition of stream 13 was 43.9 percent o-dichlorobenzene, 41.6 percent propylene, 3.2 percent propylene oxide, 4.8 percent carbon dioxide and the rest were other reaction products.

The gaseous stream 2 containing 82.8 percent propylene, 14.7 percent CO 21 1.2 percent propylene oxide and other volatile reaction products was directed to column C-1 at a rate of 3 mol/hour, where propene and carbon dioxide were separated.

The liquid stream 3 of the flash separator F-2 with a composition of 57.1 percent o-dichlorobenzene, 29.8 percent propylene, 3.8 percent propylene oxide, 1.8 percent carbon dioxide and the rest other reaction products, was fed to column C-2 at a rate of 9.8 mol/hour, where propylene and crude propylene oxide was separated as overhead stream 6 at a rate of 3.6 mol/hour, containing 80.4 percent propylene, 10.5 percent propylene oxide, 4.8 percent carbon dioxide and the rest were other reaction products.

Column C-2 was operated at 1.6 bar (0.16 MPa) absolute at the bottom stage, the distillate temperature was 1°C, the bottom temperature was 152°C. The ratio of distillate mol flow rate 6 to bottom mol flow rate 7 was 0.57: 1.

The gaseous stream 2 of separator F-2 and the overhead stream 6 of column C-2 were separated in column C-1 into the distillate stream 4 and the bottom stream 5. The column C-1 was operated at 11.6 bar (1.16 MPa) absolute at the top stage, the distillate temperature was 21 OC and the bottom temperature was 127°C. The ratio of distillate mol flow rate 4 to bottom mol flow rate 5 was 9.9: 1. The overhead stream 4, containing 89.7 percent propylene and 10.2 percent carbon dioxide can be directly recycled to the reactor at a rate of 5.9 mol/hour.

The crude propylene oxide bottom stream 5 can be used at a rate of 0.6 mol/hour to isolate the propylene oxide using known distillation techniques such as extractive

distillation with a composition of 67.7 percent propylene oxide, 2.8 percent water, in total 29.4 percent other volatile products such as acetaldehyde, acetone, acrolein, methyl formate, methanol, 2-propanol, and minor amounts of other reaction products such as methyl acetate, formic acid, acetic acid, 1-methoxy-2-propanol, 1-methoxy-2-propanone or 1-hydroxy-2-propanone.

The bottom stream 7 from column C-2, containing 89.6 percent 1,2-dichlorobenzene and 8.7 percent reaction products with a normal boiling point within a range of 55°C to 150°C was fed to column C-3 at a rate of 6.3 mol/hour to remove those reaction products in overhead stream 8 with a lower boiling point than the 1,2-dichlorobenzene. The column C-3 was operated at a pressure of 1.5 bar (0.15 MPa) absolute at the bottom stage, a distillate temperature of 107°C and a bottom temperature of 195°C. The ratio of distillate molar flow rate 8 to bottom molar flow rate 9 is 0.1: 1 The composition of stream 8 was 40.1 percent water, in total 42.7 percent acids such as formic acid, acetic acid and propionic acid, 3.8 percent 1,2-dichlorobenzene and totally 13.3 percent of other oxygenated by-products and impurities at a rate of 0.5 mol/hour. Most of the reaction products were removed from the solvent in line 9, such that substantially all 1,2-dichlorobenzene from the bottom stream 9 of column C-3 could be directly recycled to the oxidation reactor as stream 10 at a rate of 4.9 mol/hour containing 98.3 percent 1,2-dichlorobenzene and 1.7 percent other high boiling oxygenated products such as glycols or glycol esters such as 1,2-propylene glycol or 1,2-propylene glycol diacetate.

Only a small portion (14 percent) of stream 9 was fed in line 11 to column C-4 to remove high boiling products from the solvent to avoid accumulation. The column C-4 was operated at 0.95 bar (0.095 MPa) absolute at the top stage, the distillate temperature was 175°C, the bottom temperature was 184°C. 1,2-dichlorobenzene going overhead in line 20 was also recycled to the reactor at a rate of 0.8 mol/hour. The bottom effluent stream containing high boiling by-products and 1,2-dichlorobenzene can be re-used for other purposes at a rate of 0.01 mol/hour.

As a result of these separation steps, 0.6 mol/hour of crude propylene oxide were obtained with a composition of 67.7 percent propylene oxide and in total 32.3 percent of low boiling oxygenated by-products and impurities, such as acetaldehyde, methyl formate or methanol. 4.8 mol/hour of propene/carbon dioxide were separated at a high pressure of

53 bar (5.3 MPa) gauge with a composition of 77.0 percent propene, 16.6 percent carbon dioxide, 1.9 percent of propylene oxide and small amounts of other low boiling components and oxygen. 5.9 mol/hour of propene with minor amounts of carbon dioxide were separated at a medium pressure of 11 bar (1.1 MPa) gauge. 1,2-dichlorobenzene was recycled with a flow rate of totally 5.6 mol/hour and a composition of 98.3 percent o-dichlorobenzene and 1.7 percent of other compounds.

Example 3 The Example refers to Fig. 3. A product mixture was obtained by the oxidation of 15.4 mol/hour propylene with 2.0 mol/hour oxygen in 8.5 mol/hour o-dichlorobenzene at a temperature of 165°C, a pressure of 54 bar (5.4 MPa) gauge and at a residence time of 9 minutes in a continuously stirred tank reactor. The product mixture consisted of 14.5 mol/hour propylene, 0.5 mol/h propylene oxide, 0.1 mol/hour formic acid, 8.5 mol/hour dichlorobenzene and minor amounts of by-products.

The product mixture was fed into a second continuously stirred tank reactor and contacted with 2.0 mol/hour oxygen, 1.3 mol/hour propylene and 2.2 mol/hour carbon dioxide at a temperature of 165°C, a pressure of 54 bar (5.4 MPa) gauge and at a residence time of 9 minutes.

The reactor outlet 1, consisting of propylene oxide, not converted propene, other reaction products and o-dichlorobenzene was fed continuously to the flash separator F-1 at a rate of 28 mol/hour. The composition of this mixture was: 3.7 percent propylene oxide, 30.1 percent o-dichlorobenzene, 48.1 percent propylene, 10.5 percent carbon dioxide, in total 2.2 percent of acids such as formic acid, acetic acid and propionic acid, 0.3 percent oxygen, 2.3 percent water and in total 2.8 percent of oxygenated by-products and impurities, including aldehydes, ketones, glycols, acetals or esters, such as acetaldehyde, acetone, 1,2- propane diol, methyl formate, 1-methoxy-2-propanol, 1-methoxy-2-propanone, 1-hydroxy-2- propanone or methyl acetate.

The flash separator F-1 is operated at a pressure of 53 bar (5.3 MPa) gauge and at a temperature of 151 °C. The vapor stream 12 from flash separator F-1 contained 70.8 percent propene, 20.9 percent carbon dioxide, 2.5 percent of propylene oxide and small amounts of other low boiling components and oxygen at a rate of 8.7 mol/hour and was directed to flash separator F-5.

The liquid stream 13 from flash separator F-1 was fed to the flash separator F-4 at a rate of 19.2 mol/hour. The composition of stream 13 was 43 percent o-dichlorobenzene, 37.8 percent propylene, 4.3 percent propylene oxide and the rest, other reaction products. The flash separator F-4 was operated at a pressure of 41 bar (4.1 MPa) absolute and a temperature of 144°C. The gaseous stream 17 leaving the flash separator F-4 with a composition of 73.5 percent propylene, 18.5 percent carbon dioxide, 2.7 percent propylene oxide and the rest other reaction products, was fed at a rate of 2.4 mol/hour to flash separator F-5.

The flash separator F-5 was operated at a pressure of 50.5 bar (5.05 MPa) absolute and at a temperature of 78°C. The inlet streams to the separator F-5 were the gaseous streams 12 and 17. The combined stream was partly condensed and separated into the liquid stream 18 containing 73.6 percent propene, 16 percent carbon dioxide, 3 percent propylene oxide and the rest, other reaction products and solvents at a rate of 8.1 mol/hour, and into the gaseous stream 19 at a rate of 3 mol/hour and a composition of 65.6 percent propylene, 30.3 percent carbon dioxide and minor amounts of propylene oxide, oxygen and other volatile oxygenated products such as methyl formate and acetaldehyde.

The stream 19 could either be directly recycled to the reactor or be used to remove small amounts of carbon dioxide.

The liquid stream 16 of separator F-4 containing 48.9 percent o-dichlorobenzene, 32.6 percent propylene, 4.5 percent propylene oxide and the rest other reaction products, was fed at a rate of 16.8 mol/hour together with the liquid stream 18 of separator F-5 to the flash separator F-2. The flash separator F-2 was operated at a pressure of 12 bar (1.2 MPa) absolute and at a temperature of 68°C.

The gaseous stream 2, containing 78 percent propylene, 18.1 percent Cl2, 1.7 percent propylene oxide and other volatile reaction products is directed to column C-1 at a rate of 9.4 mol/hour, where propene and carbon dioxide were separated.

The liquid stream 3 of the flash separator F-2 was fed to the flash separator F-3 at a rate of 15.6 mol/hour, which was operated at a pressure of 3.9 bar (0.39 MPa) absolute and a temperature of 56°C. The gaseous stream 14 of F-3 was directed to column C-1 with a composition of 84.5 percent propylene, 3.5 percent propylene oxide, 8.7 percent carbon dioxide and the rest other volatile reaction products at a rate of 3.2 mol/hour.

The liquid stream 15 of the flash separator F-3 was fed to column C-2 at a rate of 12.3 mol/hour, where propylene and crude propylene oxide was separated as overhead stream 6 at a rate of 2.5 mol/hour, containing 57.3 percent propylene, 29.6 percent propylene oxide and the rest, other reaction products. Column C-2 was operated at 1.6 bar (0.16 MPa) absolute at the bottom stage, the distillate temperature was 19°C, the bottom temperature was 144°C. The ratio of distillate molar flow rate 6 to bottom molar flow rate 7 was 0.26: 1.

The gaseous stream 2 of separator F-2, the gaseous stream 14 of separator F-3 and the overhead stream 6 of column C-2 were separated in column C-1 into the distillate stream 4 and the bottom stream 5. The column C-1 was operated at 11.9 bar (1.19 MPa) absolute at the top stage, the distillate temperature was 19°C and the bottom temperature was 127°C. The ratio of distillate mol flow rate 4 to bottom mol flow rate 5 was 8.4: 1. The overhead stream 4, containing propylene and carbon dioxide can be directly recycled to the reactor at a rate of 13.5 mol/hour.

The crude propylene oxide bottom stream 5 leaving column C-1 at a rate of 1.6 mol/hour contained 64.4 percent propylene oxide, 8.4 percent water, in total 27.1 percent other volatile products such as acetaldehyde, acetone, acrolein, methyl formate, methanol, 2-propanol, and minor amounts of other reaction products, such as methyl acetate, formic acid, acetic acid, 1-methoxy-2-propanol, 1-methoxy-2-propanone or 1-hydroxy-2-propanone.

Propylene oxide can be isolated from this mixture using known distillation techniques such as extractive distillation.

The bottom stream 7 from column C-2, containing 85.2 percent 1,2- dichlorobenzene and 12.5 percent reaction products with a normal boiling point within a range of 55°C to 150°C was fed to column C-3 at a rate of 9.8 mol/hour to remove those reaction products in overhead stream 8 with a lower boiling point than 1,2-dichlorobenzene. The column C-3 was operated at a pressure of 1.5 bar (0.15 MPa) absolute at the bottom stage, a distillate temperature of 102°C and a bottom temperature of 194°C. The ratio of distillate molar flow rate 8 to bottom molar flow rate 9 was 0.15: 1 The composition of stream 8 was 38.5 percent water, in total 41.3 percent acids such as formic acid, acetic acid and propionic acid and totally 20.2 percent of other oxygenated by-products and impurities at a rate of 1.3 mol/hour. Most of the reaction products were removed from the solvent in line 9, such that substantially all

1,2-dichlorobenzene from the bottom stream 9 of column C-3 could be directly recycled to the oxidation reactor as stream 10 at a rate of 7.35 mol/hour containing 97.8 percent 1,2-dichlorobenzene and 2.2 percent other high boiling oxygenated products such as glycols or glycol esters, such as 1,2-propyiene glycol or 1,2-propylene glycol diacetate.

Only a small portion (14 percent) of stream 9 was fed in line 11 to column C-4 to remove high boiling products from the solvent to avoid accumulation. The column C-4 was operated at 0.9 bar (0.09 MPa) absolute at the top stage, the distillate temperature was 174°C, the bottom temperature was 182°C. 1,2-dichlorobenzene going overhead in line 20 was also recycled to the reactor at a rate of 1.15 mol/hour. The bottom effluent stream containing high boiling by-products and 1,2-dichlorobenzene can be re-used for other purposes at a rate of 0.02 mol/hour.

As a result of these separation steps, 1.6 mol/hour of crude propylene oxide are obtained with a composition of 64.4 percent propylene oxide and in total 35.6 percent of low boiling oxygenated by-products and impurities, such as acetaldehyde, methyl formate or methanol. 3 mol/hour of propene/carbon dioxide are separated at a high pressure of 50.5 bar (5.05 MPa) absolute with a composition of 65.6 percent propene, 30.3 percent carbon dioxide, 1.1 percent of propylene oxide and small amounts of other low boiling components and oxygen. 13.5 mol/hour of propene/carbon dioxide were separated at a medium pressure of 11 bar (1.1 MPa) gauge. o-dichlorobenzene was recycled with a flow rate of totally 8.5 mol/hour and a composition of 97.8 percent o-dichlorobenzene and 2.2 percent of other compounds.