DIRECT EPOXIDATION PROCESS

FIELD OF THE INVENTION

The present invention relates to a process for epoxidizing an olefin in a slurry bubble-column reactor. Epoxide is produced by reacting an olefin, oxygen, and hydrogen in a solvent in the presence of a catalyst.

BACKGROUND OF THE INVENTION

A slurry bubble-column reactor is operated by suspending catalytic particles in liquid and feeding gas reactants into the bottom of the reactor through a gas distributor, which produces small gas bubbles. As the gas bubbles rise through the reactor, the reactants are absorbed into the liquid and diffuse to the catalyst where, depending on the catalyst system, they are typically converted to gaseous and liquid products. The gaseous products are recovered from the top of the reactor. Liquid products are recovered from the suspending liquid by using different techniques like filtration, settling, hydrocyclones, magnetic techniques, etc. See YT. Shah, et al., AIChE Journal 28(3) (1982) 353-79. Epoxides such as propylene oxide are important industrial chemical intermediates. An epoxide can be produced by direct oxidation of an olefin with oxygen and hydrogen in the presence of a catalyst (U.S. Pat. Nos. 5,973,171 ; 7,138,535; 7,238,817; and 7,279,145).

Direct epoxidation of olefins with oxygen and hydrogen is highly exothermic. The reaction selectivity is highly sensitive to the reaction temperature. U.S. Pat. Appl. Pub. No. 20070260075 teaches a tower reactor having a plurality of separate reaction zones wherein the reactant gases are reacted in a slurry of catalyst particles in a solvent. Heat removal zones are provided between the reaction zones. The reactor is suitable for propylene oxide production by reacting propylene, oxygen and hydrogen reactant gases in a liquid medium comprising a slurry of catalyst particles in a solvent such as methanol or methanol and water.

SUMMARY OF THE INVENTION

The invention is a process for making epoxide from an olefin. It comprises reacting an olefin, oxygen, and hydrogen in a slurry comprising a catalyst and a solvent in a reactor to produce an epoxide. The gas bubbles in the slurry are in a churn-turbulent flow regime.

BRIEF DESCRIPTION OF DRAWINGS

Fig. 1 is a schematic presentation of one embodiment of the present invention.

DETAILED DESCRIPTION OF THE INVENTION

The process employs a catalyst. Suitable catalyst comprises a transition metal zeolite and a noble metal. Zeolites generally contain one or more of Si, Ge, Al ₁ B, P, or the like, in addition to oxygen. A transition metal zeolite (e.g., titanium zeolite, vanadium zeolite) is a crystalline material having a porous molecular sieve structure and containing a transition metal. A transition metal is a Group 3-12 element. The first row of these includes elements from Sc to Zn. Preferred transition metals are Ti, V, Mn ₁ Fe, Co, Cr, Zr, Nb, Mo, and W. Particularly preferred are Ti ₁ V ₁ Mo, and W. Most preferred is Ti. The type of transition metal zeolite employed depends upon a number of factors, including the size and shape of the olefin to be epoxidized. For example, it is especially advantageous to use titanium silicalite-1 (TS-1 , a titanium silicalite having an MFI topology analogous to that of the ZSM-5 aluminosilicate) for the epoxidation of propylene. For a bulky olefin such as cyclohexene, larger pore zeolites may be preferred.

Suitable titanium zeolites include titanium silicates (titanosilicates). Preferably, they contain no element other than titanium, silicon, and oxygen in the lattice framework (see R. Szostak, "Non-aluminosilicate Molecular Sieves," in Molecular Sieves: Principles of Synthesis and Identification (1989), Van Nostrand Reinhold, pp. 205-282). Small amounts of impurities, e.g., boron, iron, aluminum, phosphorous, copper, and mixtures thereof, may be present in the lattice. The amount of impurities is preferably less than 0.5 weight percent (wt%), more preferably less than 0.1 wt%. Preferred titanium silicates will generally have a composition corresponding to the following empirical formula: xTiO ₂ *(1 -X)SiO ₂ ,

where x is between 0.0001 and 0.5. More preferably, the value of x is from 0.01 to 0.125. The molar ratio of Si:Ti in the lattice framework of the zeolite is advantageously from 9.5:1 to 99:1 (most preferably from 9.5:1 to 60:1 ). The use of relatively titanium-rich zeolites may also be desirable. Particularly preferred titanium zeolites include the class of molecular sieves commonly known as titanium silicalites (see Catal. Rev.-Sci. Enq. 39(3) (1997) 209). Examples of these include TS-1 , TS-2 (having an MEL topology analogous to that of the ZSM-11 aluminosilicate), and TS-3 (as described in Belgian Pat. No. 1 ,001 ,038). Titanium zeolites having framework structures isomorphous to zeolite beta, mordenite, ZSM-12, MCM-22, MCM-41 , and MCM-48 are also suitable for use. Examples of MCM-22, MCM-41 , and MCM-48 zeolites are described in U.S. Pat. Nos. 4,954,325, 6,077,498, and 6,114,551 ; Maschmeyer, T., et al, Nature 378(9) (1995) 159; Tanev, P. T., et al., Nature 368 (1994) 321 ; Corma, A., J. Chem. Soc. Chem. Commun. (1998) 579; Wei D., et al., Catal. Today 5J. (1999) 501 ). The most preferred is TS-1. Suitable noble metals include, e.g., gold, silver, platinum, palladium, iridium, ruthenium, rhenium, rhodium, osmium, and mixtures thereof. Preferred noble metals are Pd, Pt, Au, Re, Ag, and mixtures thereof. A catalyst comprising palladium is particularly preferred. Typically, the amount of noble metal present in the epoxidation catalyst will be in the range of from 0.01 to 20 wt%, preferably 0.1 to 5 Wt%.

The noble metal and the transition metal zeolite may be on a single particle or on separate ones. For example, the noble metal may be supported on the transition metal zeolite. Alternatively, the catalyst may comprise a mixture of a transition metal zeolite and a noble metal supported on a carrier. Suitable carriers for the supported noble metal include carbons, titanias, zirconias, niobias, silicas, aluminas, silica- aluminas, titania-silicas, zirconia-silicas, niobia-silicas, ion-exchange resins, and the like, and mixtures thereof.

The manner in which the noble metal is incorporated into the catalyst is not critical. For example, the noble metal may be supported on the transition metal zeolite or other carriers by impregnation, ion exchange, adsorption, precipitation, or the like.

The weight ratio of the transition metal zeolite to noble metal is not particularly critical. However, a transition metal zeolite to noble metal ratio of from 10:1 to 10000:1 (grams of transition metal zeolite per gram of noble metal) is preferred.

The catalyst particles are generally very small in size and have a means mass diameter of 10-500 μm, preferably 20-100 μm. Catalyst particles of this size range can be produced by a number of means, one example of which is spray drying. Use of a catalyst of such size minimizes the mass transfer between the liquid and the catalyst due to large solid-liquid interfacial area.

The concentration of the catalyst in the slurry is typically in the range of from 1 to 40 wt%, preferably in the range of from 5 to 30 wt%. The present process has the advantage of being able to run the reactor with high catalyst loading (e.g., >10 wt%). In contrast, in a typical agitated slurry reactor, it is difficult to run the reaction when the catalyst loading is greater than about 10 wt% in the slurry without causing significant catalyst attrition. The process uses an olefin. Suitable olefins include any olefin having at least one carbon-carbon double bond and generally from 2 to 6 carbon atoms. Propylene is the most preferred. Any gas comprising propylene may be used. Typically, it comprises at least 90 wt% propylene. Preferably, it comprises greater than 95 wt% propylene. The process uses oxygen. Any gas comprising oxygen may be used.

Typically, it comprises at least 10 wt% oxygen. Preferably, it comprises greater than 90 wt% oxygen.

Similarly, The process uses hydrogen. Any gas comprising hydrogen may be used. Typically, it comprises at least 10 wt% hydrogen. Preferably, it comprises greater than 90 wt% hydrogen.

The molar ratio of hydrogen to oxygen can usually be varied in the range of H2:U2 = 1 :100 to 10:1 and is especially favorable at 1 :5 to 2:1. The molar ratio of oxygen to olefin is usually 1 :1 to 1 :20, and preferably 1 :1.5 to 1 :10.

In addition to olefin, oxygen, and hydrogen, an inert gas may be used. Suitable inert gases include nitrogen, helium, argon, and carbon dioxide. Saturated hydrocarbons with 1-4 carbon atoms, e.g., methane, ethane, propane, and n-butane, may be used. Mixtures of the listed inert gases can be used. The molar ratio of olefin to inert gas is usually in the range of 100:1 to 1 :20 and especially 20:1 to 1 :20.

The process uses a solvent. Suitable solvents are liquid under the reaction conditions. They include, for example, oxygen-containing hydrocarbons such as alcohols, aromatic and aliphatic solvents such as toluene and hexane, nitriles such as acetonitrile, and water. Suitable oxygenated solvents include alcohols, ethers, esters, ketones, water, and the like, and mixtures thereof. Preferred oxygenated solvents include water, methanol, ethanol, isopropanol, tert-butanol, and mixtures thereof.

The process may use a buffer. The buffer is employed in the reaction to inhibit the formation of glycols or glycol ethers during the epoxidation, and it can improve the reaction rate and selectivities. The buffer is typically added to the solvent to form a buffer solution, or the solvent and the buffer are added separately. Useful buffers include any suitable salts of oxyacids, the nature and proportions of which in the mixture are such that the pH of their solutions preferably ranges from 3 to 12, more preferably from 4 to 10, and most preferably from 5 to 9. Suitable salts of oxyacids contain an anion and a cation. The anion may include phosphate, carbonate, bicarbonate, sulfate, carboxylates (e.g., acetate), borate, hydroxide, silicate, aluminosilicate, or the like. The cation may include ammonium, alkylammonium (e.g., tetraalkylammoniums, pyridiniums), alkylphosphonium, alkali metal, and alkaline earth metal ions, or the like. Examples include NH ₄ , NBu ₄ , NMe ₄ , Li, Na, K, Cs ₁ Mg, and Ca cations. The preferred buffer comprises an anion selected from the group consisting of phosphate, carbonate, bicarbonate, sulfate, hydroxide, and acetate; and a cation selected from the group consisting of ammonium, alkylammonium, alkylphosphonium, alkali metal, and alkaline earth metal ions. Buffers may preferably contain a combination of more than one suitable salt. Typically, the concentration of the buffer in the solvent is from 0.0001 M to 1 M, preferably from 0.0005 M to 0.3 M. The buffer may include ammonium hydroxide which can be formed by adding ammonia gas to the reaction system. For instance, one may use a pH = 12-14 solution of ammonium hydroxide to balance the pH of the reaction system. More preferred buffers include alkali metal phosphates, ammonium phosphates, and ammonium hydroxide. The ammonium phosphates buffer is particularly preferred.

The process comprises reacting the olefin, oxygen, and hydrogen in a slurry comprising the catalyst and the solvent in a reactor wherein the reactor comprises

gas bubbles in a churn-turbulent flow ^' regime. Typically three flow regimes have been identified in a slurry bubble-column reactor: homogeneous, heterogeneous (or churn-turbulent), and slug flow regime. See YT. Shah, et al., AIChE Journal 28(3) (1982) 353-79. The superficial gas velocity and the column diameter are the main variables that determine the flow regime of the slurry bubble-column reactors. The superficial gas velocity is the gas flow rate divided by the internal cross-section area of the bubble column at the reaction conditions. The homogeneous flow regime occurs when the superficial gas velocity is relatively low (e.g., <0.05 meter per second, m/s) and is marked by a narrow bubble-size distribution, and bubbles are distributed relatively uniformly over the cross section of a reactor, a column, or other vessels. In this regime the mixing intensity and axial dispersion are low due to weak bubble-bubble interactions. As the superficial gas velocity increases the uniform distribution of gas bubbles vanishes, and a highly turbulent flow structure appears. This is called a heterogeneous flow regime, where large bubbles or agglomerates of bubbles form and travel upward at high velocity, mainly in the axis of the vessel. The heterogeneous regime is also dependent on the column diameter and occurs at a certain column diameters typically above 0.10 m depending on the operating conditions and physical properties of the three-phase system. At lower diameter (i.e., <0.10 m) and high superficial gas velocities (e.g., >0.05 m/s) slug flow regime prevails where large bubbles are stabilized by the column wall and take on the characteristic slug shape. In this regime the mixing and mass transfer are very poor. Consequently, for a three-phase operation in a slurry bubble-column reactor, the churn-turbulent flow regime is preferred if good mixing is beneficial to the operation. Due to high gas flow rate and mixing intensity, there is a significant bubble-bubble interaction that results in a high level of bubble coalescence and break-up. Similarly a bimodal bubble size distribution (large and small gas bubbles) is observed. This is referred to as the two-bubble class mode. The large gas bubbles rise fast in the center of the column in a plug-flow manner and are responsible for the mixing, dispersion and solid suspension, whereas the small gas bubbles with small rise velocity are entrained in the liquid-backmixing and contribute mostly toward the enhancement of the gas holdup and mass transfer. Furthermore, due to high liquid mixing in the churn-turbulent flow regime, most slurry bubble-column reactors in this regime can operate at isothermal conditions.

In a slurry bubble-column reactor where gas bubbles are in churn-turbulent flow regime, sufficient mixing is achieved without the use of a mechanical agitator. Elimination of a mechanical agitator reduces time and cost associated with the shutdown for the repair and maintenance of the agitator and necessary seals. It allows higher loading of the catalyst in the reactor without causing significant attrition. It also frees up substantially more space within the reactor for heat transfer tubes, thus improving overall heat removal capability. Thus, the process improves the product yield and gives improved operational and maintenance advantages.

Preferably the olefin, oxygen and hydrogen are fed to the reactor at or near the bottom of the reactor. More preferably they are fed to the reactor via a gas sparger. The sparging of the gas bubbles is the main mechanism for mixing, catalyst suspension, and dispersion.

The reactor is preferably cylindrical in shape and has a diameter of greater than 0.10 meter (m) and has a height-to-diameter ratio of greater than 3:1. The reactor diameter has a significant effect on the flow regime of a slurry bubble-column reactor.

The superficial gas velocity in the reactor is preferably in the range of from 0.05 to 0.60 m/s. More preferably, it is in the range of from 0.08 to 0.20 m/s.

The process is performed at a temperature in the range of from 30 to 100 ⁰ C, preferably in the range of from 50 to 7O ⁰ C, and at a pressure of from 100 to 800 psig, preferably at 150 to 400 psig. A cooling coil is preferably used to remove the heat of the reaction.

A vapor zone and a slurry zone are present in the reactor. A gaseous product stream is present in the vapor zone. The gaseous product stream in the reactor comprises the epoxide, olefin, hydrogen, oxygen, solvent vapor, and inert gases if used. The gaseous product stream exits the reactor from the vapor zone. The epoxide may be isolated from the gaseous product stream by standard techniques. The remaining gases are preferably recycled to the reactor after the epoxide is separated. The catalyst and a liquid product stream that comprises solvent, epoxide, and dissolved gases are present in the slurry zone. The liquid product stream exits the reactor after the catalyst is separated from the slurry. After the epoxide is removed from the liquid product stream, the remaining liquid stream is further processed for solvent recycle and/or byproduct removal.

In one embodiment, a process for the epoxidation of propylene is illustrated in Figure 1. It comprises a cylindrical reactor 1 , gas inlets 2 and 3 near the bottom of the reactor, and a gas outlet 5 at the top of the reactor. A heat-exchanging tube (not shown) is installed in the reactor to remove the reaction heat. A liquid feed is added to the reactor via line 8. The liquid feed may comprise a solvent, a buffer, and other recycled liquid streams (e.g., via line 24). These components may be mixed outside the reactor and enter the reactor as a single feed stream, or they may be fed to the reactor through individual lines.

Gases are fed to the bottom of the reactor. A gas distribution apparatus 4 is used to distribute the gases to the cross section of the reactor to help the uniform mixing of the gases with the slurry in the reactor. The gases are dispersed to create small bubbles and distributed uniformly over the cross section of the reactor to maximize the intensity of mass transfer. Perforated plates and perforated ring spargers are preferably used for gas distribution. Fresh propylene and hydrogen, and optionally recycled gases may be combined into one feed stream and fed to the reactor via line 3. Oxygen enters the reactor through a separate line 2 and sparger 4.

The gaseous product stream exits the reactor from line 5 at the top of the reactor. The liquid product stream passes through an internal filter 6 and exits the reactor via line 7. Preferably at least 10%, more preferably at least 20%, most preferably at least 30% of the total propylene oxide produced exits the reactor in the gaseous product stream. Propylene oxide present in the reactor liquid product stream may react with the solvents (such as water or an alcohol) to form propylene glycol or propylene glycol ethers more readily at high propylene oxide concentration due to the presence of catalyst. Such reactions in the gaseous product stream are negligible. Stripping propylene oxide from the slurry zone to the gas zone in the reactor in the present process reduces the byproduct formation.

Following examples merely illustrate the invention. Those skilled in the art will recognize many variations that are within the spirit of the invention and scope of the claims.

EXAMPLE 1. PROPYLENE EPOXIDATION

The following is one proposed method of practicing the process of the invention.

A catalyst is prepared by following the procedure taught by Example 5 of the copending Application Ser. No. 11/891 ,460, filed on August 10, 2007. The catalyst (3 kg) prepared by using the above procedure is charged to a cylindrical reactor 1 (Fig. 1 ) with a diameter of 0.20 m and a height of 5.2 m, equipped with two gas spargers (represented by 4) at the bottom, a gas outlet 5 at the top of the reactor, and a liquid inlet 8. An internal filter 6 is installed on the liquid outlet line 7. A cooling coil (not shown) is used to remove the heat of the reaction. Fresh propylene (0.92 kg/h) and hydrogen (660 standard liter per hour, SLPH) are fed to reactor via line 3. Oxygen (600 SLPH) is fed to reactor via line 2. The overall superficial gas velocity (including recycled gases from lines 16 and 21 ) is 0.15 m/s. A methanol/water mixture (85/15 by weight, flow rate 15 kg/h), recycled from a solvent- recovery unit 23, containing 5 mM ammonium phosphate is fed to the reactor via line 24. The reaction is operated at 6O ⁰ C and under pressure of 300 psig. A gaseous product stream exits the reactor via line 5 at a flow rate of 18 kg/h. The gaseous product stream comprises 0.38 mole percent (mol%) propylene oxide (PO), 1.58 mol% propane, 7.9 mol% oxygen, and 1.8 mol% hydrogen. PO is removed from the gaseous product stream through the use of the absorber 9. The absorber 9 has about 20 theoretical stages. A methanol-containing recycled solvent is chilled to 15°C and fed to the top of the absorber 9 at a flow rate of 6 kg/h via line 10. The gaseous product stream is fed to the bottom of the absorber via line 5. The absorber gas effluent in line 11 after a small purge via line 22 is recompressed (not shown) and recycled to the reactor via line 16. The absorber liquid effluent exits the absorber via line 12. The reactor liquid product stream exits the reactor via line 7 at a flow rate of 16 kg/h. The liquid product stream comprises 3.3 wt% PO, 0.87 wt% propylene, and 74 wt% methanol. The liquid product stream 7 from the reactor and the absorber liquid effluent 12 are combined into a liquid stream 13, which is fed to a depropanizer 14. The depropanizer 14 has 20 theoretical stages. Its overhead temperature is 3 ⁰ C and pressure is 10 psig. Its bottoms temperature is 82 ⁰ C and pressure is 10.2 psig. A C3 stream, comprising mostly propylene and propane exiting the depropanizer 14 via line 15, is sent to a C3 splitter (a distillation column,

not shown) for propylene enrichment before it is recycled to the reactor 1 via line 21. The depropanizer liquid stream 17 is fed to a crude PO column 18. The crude PO column 18 has 44 theoretical stages. Its overhead temperature is 36 ⁰ C and pressure is 1 psig. Its bottoms temperature is 74 ⁰ C and pressure is 1.4 psig. Crude PO stream is obtained as overhead 19. The remaining liquid stream exits column 18 via line 20. The remaining liquid stream, comprising 72 wt% methanol, 20 wt% water, and derivatives of propylene oxide such as propylene glycol, propylene glycol monomethyl ethers, dipropylene glycol, and dipropylene glycol methyl ethers, is sent to the solvent-recovery column 23 via line 20 for water and heavies removal. Column 23 has 16 theoretical stages. Its overhead temperature is 72 ⁰ C and pressure is about 0 psig. Its bottoms temperature is 100.5 ⁰ C and pressure is 0.2 psig. A methanol-containing solvent is recovered as overhead and is recycled to the reactor 1 via line 24. Water and heavies are removed via line 25.

The overall PO/POE selectivity is expected to be about 89%. POE represents the total amount of PO and PO equivalents (propylene glycol, dipropylene glycol, and propylene glycol ethers) formed in the process. POE (mole) = moles of PO + moles of PO units in the PO derivatives. PO/POE mole selectivity = (moles of PO)/(moles of POE) x 100.