RU2795004C2 - Method for obtaining ethylene glycol - Google Patents

Abstract

FIELD: chemistry.

SUBSTANCE: liquid phase process for producing ethylene glycol. Invention comprises the following steps: obtaining a water stream and an aqueous ethylene oxide feed stream; separating the aqueous ethylene oxide feed stream into at least the first equal portion and the second equal portion; combining the first portion and the water stream to form the first feed stream; feeding the first feed stream to the inlet of the first adiabatic reactor, where the inlet of the first adiabatic reactor is at the inlet temperature; reacting ethylene oxide and water in the presence of the first ion exchange resin catalyst in the first adiabatic reactor to produce the first reactor effluent containing water, ethylene glycol and unreacted ethylene oxide; removing the first reactor effluent from the outlet of the first adiabatic reactor; cooling the first reactor effluent through a heat exchanger downstream of the first adiabatic reactor to a hydration temperature; and combining the second portion and the cooled product stream to form the second feed stream; transferring the second feed stream to the inlet of the second adiabatic reactor, where the inlet of the second adiabatic reactor is at the hydrolysis temperature and causing an interaction of ethylene oxide and water contained in the second feed stream in the presence of the second catalyst based on an ion exchange resin in the second adiabatic reactor to obtain an effluent of the second reactor containing water, ethylene glycol and unreacted ethylene oxide. In this case, each of the first adiabatic reactor and the second adiabatic reactor contains a catalyst based on ion-exchange resin in a fixed catalyst bed, and the temperature at the outlet of the adiabatic reactors is higher than the inlet temperature. The total molar ratio of water:ethylene oxide in the water stream and the added aqueous ethylene oxide feed stream together is from 5:1 to 15:1.

EFFECT: development of an improved method of adiabatic catalytic hydration, in which the swelling of the catalyst is reduced, the flexibility of the method is significantly improved.

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Description

FIELD OF TECHNOLOGY

This invention relates to a process for the oxidation of ethylene to ethylene oxide and the hydrogenation of ethylene oxide to ethylene glycol.

BACKGROUND OF THE INVENTION

The role of ethylene glycol in society has increased significantly since soon after its first synthesis, the properties of the molecule became the subject of a fierce dispute between two giants of early organic chemistry, Adolf Wurz and Hermann Kolbe. In particular, Würz and Kolbe disputed the functionality and chemical formula of ethylene glycol among the still emerging notions of alcohol homologues, which were the key to Kolbe's great theories on chemical structure. The laboratory soon became a proxy war of the growing industrial and technical power of the growing rivalry between Germany and France, with both countries devoting significant resources to scientific research, and Wurtz's magic in chemical synthesis giving France a significant advantage. The dispute was only ended by the use of arms when Bismarck's German Confederation annexed Wurtz's Alsatian homeland in the Franco-Prussian War and thus effectively turned the international dispute into an internal one.

Today, interest in ethylene glycol is more peaceful, but all the more competitive, because ethylene glycol is one of the most widely produced organic chemicals. Ever since large-scale industrial production of ethylene glycol began on the eve of World War I, the dramatic increase in the use of internal combustion engines for motor cars and other vehicles has stimulated demand for ethylene glycol as coolant and antifreeze. Growth in ethylene glycol production has only accelerated since then, with estimated global ethylene glycol production exceeding 25 billion tonnes by 2017.

Ethylene glycol is usually produced as one of the many derivatives of ethylene oxide, and although other production methods are available, the majority is produced from ethylene oxide by a liquid phase non-catalytic thermal hydration process. Since ethylene oxide reacts more easily with ethylene glycols than with water, it is inevitable that mixtures of monoethylene glycol (MEG) as well as higher glycol content by-products such as diethylene glycol (DEG), triethylene glycol (TEG) and even higher ethylene glycols will appear. Although these higher glycols have significant economic value, many manufacturers and plant operators would like to avoid their formation because the end user market for these products is not as well developed and it can be difficult to find industrial customers who need them and sell these higher glycols to them. .

To suppress the reaction between the glycol product and ethylene oxide, and thereby reduce the formation of these higher glycols, conventional non-catalytic hydration is carried out with an amount of water that is much higher than the stoichiometric amount of water for hydration of ethylene oxide to ethylene glycol, for example, 15-40 moles of water per mole of ethylene oxide. This addition of excess water is effective in balancing the kinetically favorable parallel reaction between the glycol product and ethylene oxide, which, as mentioned above, competes with the hydration of ethylene oxide to monoethylene glycol. However, while effective in suppressing the formation of higher ethylene glycols, the use of a large excess of water compared to ethylene oxide presents a problem for the plant operator in removing these large excesses of unreacted water, since such removal is energy intensive and requires evaporation/distillation facilities. Accordingly, there has been intensive research into alternatives to the standard thermal hydration of ethylene oxide for the production of ethylene glycol.

One early example of an alternative to traditional thermal hydration is the homogeneous catalytic hydration of ethylene oxide to monoethylene glycol. The earliest examples of this approach involved the homogeneous catalysis of sulfuric acid and its associated salts (see Othmer, D.F. and Thakar, M.S., Glycol Production - Hydration of Ethylene Oxide. Ind. Eng. Chem. 1958, 50, 1235). EP 0,123,700 describes the modification of these earlier generations of acid catalysts by treating them with, for example, ethylamines to partially neutralize them, in the hope of improving the selectivity of the hydration reaction to monoethylene glycol. Since then, other salts have been proposed for homogeneous systems, such as quaternary phosphonium salts as described in US Pat. No. 4,160,116, and metallate and bicarbonate salts as described in US Pat. No. 7,683,221. Increasingly creative combinations of organics such as EDTA and Salen compounds have also been proposed as homogeneous catalysts (see Hal, J. W., Ledford, J. S., and Zhang, X., Catalysis Today 123 (2007), 310-315).

Homogeneous catalyst systems are often used in a two-stage process for the production of ethylene glycol, see for example US Patent No. 4,519,875, in which ethylene oxide is first reacted with carbon dioxide to produce ethylene carbonate, which is then hydrolyzed to ethylene glycol, where the same catalyst is usually used in both stages. . Following this pioneering patent, continuous research has led to incremental improvements in the two step process. For example, in US Pat. No. 5,763,691, the carbonization reaction is catalyzed in ethylene oxide absorbate in the presence of a halogenated organic phosphonium salt carbonization catalyst. Additional research has greatly expanded the range of known catalysts; see, for example, macrocyclic chelating compounds ("crown ethers"), described in US patent No. 7,453,015.

Although homogeneous catalysts have improved the selectivity of the ethylene glycol hydration reaction over non-catalytic hydration, homogeneous catalysts have disadvantages that more than offset these improvements. In particular, the glycol product solution contains a dissolved or suspended homogeneous catalyst—essentially causing catalyst contamination of the product, which greatly degrades the quality of the product, and is also very aggressive, as is the case for example with acid catalysts. Thus, an additional step of separating the homogeneous catalyst from the glycol product solution is necessary, adding cost and complexity to the process. For example, although US Pat. No. 5,763,691, cited above, attempts to simplify the two-stage production of ethylene glycol through the reaction of ethylene oxide and carbon dioxide in the absorbate, the patent still requires a separate distillation stage to remove the carbonization catalyst.

This complexity of the method brings with it the additional disadvantage that it is not very versatile and flexible. For example, the process lacks the flexibility to operate partially catalytically, i.e. in mixed catalytic and non-catalytic modes, which would allow the production of more higher glycols if the operator so desires, as well as adapt the plant in case of some disruptions in logistics and operation. when it cannot operate in catalytic mode, such as when sufficient catalyst is not available. The two-stage hydration process also lacks the versatility to be used in both new plants and retrofits. The two-stage hydration method can only be used in new plants and cannot be used in retrofits, because in the event of an upgrade, all “internal interfaces” would have to be removed and replaced, making such an upgrade too costly. The disadvantage of the two-stage hydration method is the need to inventory concentrated ethylene oxide (concentration of more than 70%) for the production of intermediate ethylene carbonate in the first stage of the method. The carbonization reaction in the first step is carried out by bubbling carbon dioxide into such a concentrated ethylene oxide liquid solution at high temperature and pressure (for example, more than 1.6 MPa). Large volumes of concentrated ethylene oxide require significant additional measures (e.g., storage in a liquid refrigerant) to prevent reactions or events that could result in ignition, combustion, deflagration, detonation, or explosion of any gas stream, but especially those containing typical concentrations of ethylene oxide.

Not surprisingly, given the aforementioned disadvantages of both non-catalytic and homogeneous catalytic hydration systems, heterogeneous catalysis using solid or supported catalysts has also been explored. In search of the benefits of higher catalytic hydration selectivity without the aforementioned disadvantages, researchers have turned their attention to heterogeneous catalyst systems. Because heterogeneous catalysts do not adsorb into liquid phase reaction mixtures or product solutions, do not contaminate product solutions, do not require a separation step, and promote continuous operation of the hydration reactor, they are of considerable interest.

A variety of solid materials have been proposed as heterogeneous hydration catalysts. This list includes zeolites such as those described in US Pat. No. 3,028,434, as well as amorphous aluminosilicates, solid bases, and supported acid catalysts such as those described in US Pat. No. 4,165,440. Ion exchange resins have also been proposed, a sulfonic acid-based partially amine-neutralized styrene-divinylbenzene resin is described in EP 0 123 709; quaternary phosphonium ion exchange resin catalysts, see PCT publication WO 2000/35840; and strong base ion exchange resins, summarized in US Pat. No. 6,211,419, were among the first ion exchange resins proposed to catalyze the reaction of water and ethylene oxide. More recent publications have expanded the scope to a wide range of anion exchange resins, see for example Chinese Patent No. 1208135.

Although excellent results have been achieved using heterogeneous ion exchange resin catalysts, their use also has certain disadvantages. In particular, the ion exchange resin has limited resistance to heat, so that under typical reaction conditions, the resin swells for long periods of time, resulting in deactivation. Swelling, which results from internal polymerization in the resin layer and the overall accumulation of high molecular weight by-products, is a serious problem that is a significant limiting factor in the service life of an ion exchange resin. Resins typically undergo some reversible swelling during operation due to exposure to solvents; however, with careful handling, this swelling can be reduced and the life of the ion exchange resin extended. However, exposure of the ion exchange resin to ethylene oxide can cause more problematic and irreversible swelling of the resin, especially when such exposure occurs at higher hydration reactor temperatures, which, as mentioned above, may be outside the temperature range that the resin can satisfactorily tolerate.

Swelling is not only a problem in itself, but can also reduce efficiency and complicate the catalytic hydration process; for example, when excessive swelling requires changing the resin and thus shutting down the catalytic hydration unit in which it is contained. Various attempts have been made to address these problems, including attempts to find resins that are less likely to swell under normal operating conditions, and to develop process modifications in an attempt to improve performance in addition to reducing swelling. For example, patent EP1828086 describes the use of "protective beads" designed to remove impurities that can conceivably cause the resin to swell, but such attempts have been ineffective.

Instead of protecting resins with a protective layer in an attempt to minimize swelling, another more successful approach, described in PCT Publication WO2008/150338, does not aim to protect the resin, but to facilitate the removal of an expired ion exchange resin through previously unsuggested work. an upflow catalytic reactor which expands the catalyst bed to allow size separation of the catalyst particles and thereby selectively remove swollen particles. Another approach is shown in US Pat. No. 6,160,187 which states that swelling reduction can be achieved under adiabatic reaction conditions and using an adiabatic reactor. However, the examples contained in the "187" patent do not actually show any improvement in the swelling characteristics of the resins, nor has there been any subsequent development of the adiabatic regime, which has been neglected.

Accordingly, while there have been several promising developments in certain aspects of heterogeneous catalytic hydration, there remains a significant need for further improvements in catalytic hydration under adiabatic process conditions.

BRIEF DESCRIPTION OF THE INVENTION

The present invention relates to a liquid phase process for the production of ethylene glycol, comprising: obtaining a water stream and an aqueous ethylene oxide feed stream; separating the aqueous ethylene oxide feed stream into at least a first equal part and a second equal part; combining the first portion and the water stream to form a first feed stream; feeding the first feed stream to the inlet of the first adiabatic reactor, where the inlet of the first adiabatic reactor has an inlet temperature; reacting ethylene oxide and water in the presence of a first ion exchange resin catalyst in a first adiabatic reactor to produce a first reactor effluent containing water, ethylene glycol and unreacted ethylene oxide; removing the first reactor effluent from the outlet of the first adiabatic reactor; cooling the first reactor effluent through a heat exchanger downstream of the first adiabatic reactor to a hydration temperature; combining the second portion and the cooled product stream to form a second feed stream; transferring the second feed stream to the inlet of the second adiabatic reactor, where the inlet of the second adiabatic reactor has a hydrolysis temperature; reacting ethylene oxide and water contained in the second feed stream in the presence of a second ion exchange resin catalyst in a second adiabatic reactor to produce a second reactor effluent containing water, ethylene glycol, and unreacted ethylene oxide; where each of the first adiabatic reactor and the second adiabatic reactor contain an ion exchange resin catalyst in a fixed catalyst bed, and the outlet temperature is higher than the inlet temperature.

BRIEF DESCRIPTION OF MULTIPLE DRAWING TYPES

The foregoing summary, as well as the following detailed description of the preferred embodiments of the invention, will be better understood when read in conjunction with the accompanying drawings. To illustrate the invention, the drawings show the presently preferred embodiments. However, it should be understood that the invention is not limited to the precise devices and instruments shown. On the pictures:

Figure 1 shows a process flow diagram for obtaining a stream of concentrated ethylene oxide in accordance with this invention;

Figures 2-3 show process flow diagrams for ethylene glycol production options in accordance with the present invention; And

Figure 4 shows a flow diagram of a simulated exemplary process according to the present invention for producing ethylene glycol.

DETAILED DESCRIPTION OF THE INVENTION

All parts, percentages and ratios used herein are by volume unless otherwise indicated. All pressures are absolute. All documents cited here are incorporated by reference.

By "water" is meant any kind of water suitable for use in chemical and petrochemical processing, including deionized, demineralized, industrial, drinking and distilled water. Steam condensate and condensate obtained from the dehydration and evaporation part of the glycol production process can also be used.

The present invention provides an improved adiabatic catalytic hydration process in which catalyst swelling is reduced, process flexibility is greatly improved, and the process is configured to advantageously be used to upgrade a non-catalytic process to a catalytic one. A particular advantage of the present invention is that a number of adiabatic reactors are used with fresh aqueous ethylene oxide injected into each adiabatic reactor. Typically, the non-catalytic hydration reaction is carried out in a single adiabatic reactor, regardless of yield or maximum reactor temperature. Such operation can be particularly detrimental to the ion exchange resins of this invention, as temperatures outside the temperature range indicated for such resins can increase swelling and reduce catalyst life, as described above.

Instead of using such single-stage adiabatic reactors and their attendant disadvantages (such as high temperature in the presence of a catalyst), the present invention uses multi-stage adiabatic reactors arranged in series, with external cooling between stages and injection of fresh aqueous ethylene oxide at each stage. Since no heat is removed in the adiabatic reactor, the temperature of the effluent is higher than the temperature of the feed, and the heat of reaction must be removed from the effluent by a cooler at each stage. This is particularly important to prevent the aforementioned catalyst swelling and to prevent excessive non-catalytic reaction of glycol products with ethylene oxide to higher glycols. In addition, a conventional non-catalytic adiabatic reactor cannot advantageously be configured to allow a series of multiple reactor stages with fresh aqueous ethylene oxide injection. In such a configuration, introducing fresh ethylene oxide into the reactor effluent containing a significant concentration of glycol products will cause the glycol products to overreact to undesired higher glycols. In the present invention, these problems are solved by a catalytic hydration reactor unit with two or more adiabatic reactors in series, intercooling between the reactors, and injection of fresh aqueous ethylene oxide into each reactor. Intercooling removes the heat of reaction from the effluent, lowering its temperature so that it can be combined with fresh aqueous ethylene oxide for the next reactor stage. The present invention uses both the temperature of the cooled effluent from each intercooler and the proportion of fresh aqueous ethylene oxide injection into each reactor to control the maximum temperature in each reactor within the range desired for catalyst life and performance. Moreover, from a process point of view, cooling with external interstage heat exchangers is easier, more efficient and simpler than setting the appropriate mode for internal reactor cooling, and can be adjusted to absorb normal fluctuations in the process while maintaining the performance of catalytic reactors. Thus, the use of a range of adiabatic reactors in accordance with the present invention improves the versatility of the process and allows the conventional non-catalytic to catalytic hydration process to be upgraded or converted, as described in more detail below.

The present invention achieves these advantages while maintaining the high selectivity of prior catalytic hydration systems for monoethylene glycol. In each bed of an adiabatic reactor, ethylene oxide and water react in the presence of a heterogeneous catalyst (such as an ion exchange resin) to form monoethylene glycol, with at least 98% of the ethylene oxide converted to monoethylene glycol, with only a small proportion converted to higher glycols such as diethylene glycol or triethylene glycol.

Thus, the present invention provides an improved process scheme that includes a catalytic hydration reactor block containing adiabatic reactors in series, which maintains low resin swelling and excellent selectivity, and reduces process complexity and increases versatility.

This catalytic hydration reactor bank (also "reactor bank") obtained in accordance with the present invention as described above is shown as 25 in Figure 2 and is described in more detail below. The reactor block 25 is loaded with, among other feed streams, a water feed stream 36, a liquid cold feed stream through conduit 40 that is highly enriched in ethylene oxide and is supplied at a pressure high enough to keep the ethylene oxide in the liquid. Reactor bank 25 includes a final adiabatic reactor for non-catalyzed thermal hydrolysis of EO 26, which provides complete conversion of the remaining unconverted EO (present in the last adiabatic reactor effluent from the series of reactors 22) to ppm levels, significantly reducing the total volume of catalytic reactor of a number of reactors 22.

Any suitable means of obtaining such a liquid feed stream that is highly enriched in ethylene oxide is applicable to this inventive process for the catalytic production of ethylene glycol. One suitable means for this is the straight two-stage evaporative absorber-stripper unit shown in Figure 1. Figure 2 shows the communications between the absorber-stripper unit 1 and the reactor bank 25, which are obtained in accordance with the present invention. The method for producing ethylene oxide with reference to this direct two-stage evaporator-absorber-desorber unit is described in more detail below.

Ethylene oxide is produced by continuously contacting an oxygen-containing gas with an olefin, preferably ethylene, in the presence of an ethylene oxide (“epoxidation”) catalyst (described in more detail below). The oxygen may be supplied to the reaction in substantially pure molecular form or in a mixture such as air. By way of example, conventional reactant feed mixtures under operating conditions may contain from about 0.5% to about 45%, preferably from about 5% to about 30% ethylene and from about 3% to about 15% oxygen, with the balance being relatively inert materials, including substances such as carbon dioxide, water, inert gases, other hydrocarbons and reaction moderators described here. Non-limiting examples of inert gases include nitrogen, argon, helium, and mixtures thereof. Non-limiting examples of other hydrocarbons include methane, ethane, propane, and mixtures thereof. Carbon dioxide and water are by-products of the epoxidation process, as well as common impurities in feed gases. Both have an adverse effect on the catalyst, so the concentrations of these components are usually kept at a minimum.

As noted previously, one or more reaction moderators are also present in the reaction, non-limiting examples of which include organic halogenated compounds such as C _{1 -C 8} halohydrocarbons; particularly preferred are chloride-containing moderators such as methyl chloride, ethyl chloride, ethylene dichloride, vinyl chloride or mixtures thereof. Controlling the level of chloride concentration is especially important for ruthenium-containing catalysts.

As stated above, a conventional process for epoxidizing ethylene involves the vapor phase oxidation of ethylene with molecular oxygen in the presence of an epoxidation catalyst, in a fixed bed tubular reactor. Typically, commercial fixed bed ethylene oxide reactors are in the form of a plurality of parallel tubes (suitably sheathed) approximately 0.7-2.7 inches OD and 0.5-2.5 inches ID and 15-53 feet long, each filled and packed with catalyst. The initial reaction mixture (described above) is introduced into these tubes and the resulting reactor effluent contains ethylene oxide, unused reactants and by-products.

For purposes of illustration only, the following are conditions that are often used in modern commercial ethylene oxide reactors: gas hourly space velocity (GOSV) 1500-10000 h ^-1 , reactor inlet pressure 150-400 psi. inch g, coolant temperature 180-315°C, oxygen conversion 10-60%, and EO production (output) 7-20 lb. EO/cfm catalyst/h Feedstock consumption at the reactor inlet after start-up is complete and during normal operation is typically (in % vol) 1-40% ethylene, 3-12% O ₂ ; 0.3%-20%, preferably 0.3-5%, more preferably 0.3-1% CO ₂ ; 0-3% ethane, some of one or more chloride moderators as described herein; and the remainder of the raw material is argon, methane, nitrogen, or mixtures thereof.

Typical operating conditions for the ethylene epoxidation process include temperatures ranging from about 180°C to about 330°C, and preferably from about 200°C to about 325°C, and more preferably from about 225°C to about 280°C. Operating pressure can vary from about atmospheric pressure to about 30 atmospheres, depending on the desired mass flow and performance. Higher pressures may be used within the scope of this invention. Residence times in commercial reactors are typically on the order of about 2 to about 20 seconds.

In figure 1, the liquid from the wash column (hereinafter referred to as the stream of "water from the enrichment cycle") flows through the pipeline 1 and is fed into the flash drum 5. The stream of water from the enrichment cycle is an aqueous stream containing from 1.5% mol to 4% mol ethylene oxide with some dissolved gases such as methane, ethylene and others. This water stream from the enrichment cycle is fed under pressure, preferably from about 0.4 MPa to about 0.6 MPa, into the flash drum 5. The absorber 9 is permanently and directly fixed to the top of the flash drum 5, and the flash drum has an opening for connection to the absorber 9. In this invention, the pressure in both the flash drum 5 and the absorber 9 is maintained at a level sufficient so that when water from the enrichment cycle enters the flash drum 5, the “light” gases that are dissolved in the water from the enrichment cycle , which have a volatility greater than that of ethylene oxide, were washed out of the water from the enriched cycle upon entering the evaporation drum 5, and the vapor of light gases is directed by the pressure difference between the evaporation drum 5 and the absorber 9 upwards through the hole into the absorber 9 with the formation of the upper layers of the absorber 9, and the dissolved ethylene oxide remains substantially solubilized in the upgrade water stream in the bottoms of the flash drum 5. These more volatile "light" gases include the aforementioned inert, unreacted or by-product gases dissolved in the upgrade water. In addition, traces of ethylene oxide can also evaporate with more volatile light gases in the water stream from the upgrade cycle. Small amounts of ethylene oxide that evaporate from the water from the upgrade cycle and rise to absorber 9 with dissolved light gas vapor can be recovered by a descending flushing stream (not shown) to absorb and recover the upward moving ethylene oxide.

As stated above, the absorber 9 is directly and permanently attached to the top surface of the flash drum 5, for example by welding. The absorber 9 is attached to the flash drum 5 in the same manner as in a flash column and reboiler configuration well known to those skilled in the art. More specifically, the absorber 9 is attached to the flash drum 5 along the longitudinal axis of the flash drum 5, but is preferably attached in the middle of the longitudinal axis. The cross section of the absorber 9 matches the shape of the opening in the flash drum 5 over which the absorber 9 is attached, and which allows communication between the flash drum 5 and the absorber 9. The cross section of the absorber and the shape of the hole in the flash drum 5 have the same matching shape, and this shape can be any shape between a circle and an ellipse, between eccentricity 0 and eccentricity 1. The opening in the flash drum 5 and the absorber 9 is positioned so that the center of the hole is collinear with the center of the absorber 9. The configuration of the aperture in the flash drum 5 and the shape of the absorber 9 and their respective positions provides proper fluid communication between the two elements. The absorber 9 is preferably flat welded to the inner diameter of the hole made in the flash drum 5 so that the outer diameter of the absorber 9 is flush with the diameter of the hole in the flash drum 5.

In order to achieve a balance between washing out the more volatile "light" solutes while maintaining high ethylene oxide solubility in water from the enrichment cycle, the pressure in the flash drum 5 is maintained at a level of about 0.1 MPa to 0.3 MPa, and the pressure in the absorber 9, with to which there is a direct connection, only slightly below the pressure in the flash drum 5. The temperature of the gases in the upper part of the absorber 9 is from about 30°C to about 45°C, and the temperature in the flash drum 5 is from about 50°C to about 60°C .

The top components of the absorber 9 are returned via conduit 12 to the pre-installation reactor block (not shown) as a possible recycle to the reactor. The water from the upgrade cycle flows as bottoms from flash drum 5 through pump 14 and through an optional heat exchanger 20 (which preheats the stream) and into stripper 15. Stripper 15 is pressurized by controlling the back pressure created by removing vapors from flash condenser 19. The temperature at the top of the stripper 15 should be from about 25°C to about 40°C, and at the bottom of the stripper 15 is from about 120°C to about 160°C. The water stream from the enrichment cycle may be preheated in the heat exchanger 20 to reduce the load required to operate the stripper 15. The pressure in the stripper 15 is maintained in the range from about 0.2 MPa to about 0.6 MPa. Stripper 15 may also employ a steam jet pump system well known to those skilled in the art, where steam is supplied from: (1) steam generated throughout the ethylene oxide production system, or (2) medium or high pressure steam supplied from external sources. /ZGU; or a combination of the two sources.

The pressurized stream from the top of the stripper flows through conduit 30 to reactor bank 25 as shown in Figure 2. The stream flows to condenser 16 and then to flash drum 17 which separates the scrubbed light gases from the concentrated aqueous ethylene oxide liquid stream which is fed to pump 18. The output from pump 18 is divided into two streams: one stream as a return stream back to stripper 15, and one liquid aqueous stream of pure product concentrated in ethylene oxide. Washed light gases from drum 17 flow to flash condenser 19 to further recover ethylene oxide from the washed steam stream and control the amount of dissolved light gases in concentrated ethylene oxide liquid solution 40. Compared to a highly dilute aqueous solution found at a comparable location in a conventional process ethylene oxide recovery unit , the pressurized overhead stream of the stripper is substantially enriched in ethylene oxide, containing at least about 40 mole %, preferably at least about 50 mole % ethylene oxide, with the remainder being primarily water. Thus, the pressurized stream from the top of the stripper produces both high concentrations of ethylene oxide and is under high pressure. This makes the stream extremely useful for further processing to ethylene glycols via catalytic hydrolysis on a very economical basis, since there is no need to further enrich the stream or pressurize the following processing steps.

Another preferred option of the process of the present invention is that the bottoms of stripper 15, which is a lean aqueous cycle solution containing some ethylene oxide (from about 1 molar ppm to about 50 molar ppm), and which is also free of light components as well as impurities, can preferably be used in a variety of other thermal energy management schemes throughout the plant, depending on needs and conditions.

Under these conditions of streams with higher than normal concentrations of ethylene oxide, operators will continue to comply with the appropriate safety standards that are always followed when producing, processing or storing ethylene oxide. As always, precautions must be taken to prevent reactions or events that could result in ignition, combustion, deflagration, detonation, or explosion of any gas stream, especially those containing ethylene oxide in concentrations higher than typical. Accordingly, to prevent such events, safety devices may be used to relieve or reduce unwanted pressure buildup in the separation process, reaction or systems, or elsewhere in the ethylene oxide plant, both before and after what is shown. The separation in stripper 15 is so efficient that almost all of the ethylene oxide in the upgrade water that enters stripper 15 is successfully separated and recovered from the upgrade water and leaves stripper 15 as a clean liquid overhead stream, while only very a small portion of the EO leaves stripper 15 in other streams.

The pressurized overhead stripper stream obtained in stripper 15 flows at elevated pressure through conduit 30 from stripper unit 1 to reactor bank 25 as shown in Figure 2. In accordance with the present invention, a reactor bank comprises two or more reactors. Figure 2 shows a reactor block 25 with a series of four catalytic reactors arranged in series, each separated by a heat exchanger. More specifically, as shown in Figure 2, and in more detail in Figure 3, the reactor bank 25 shows a series of reactors 22 of four in-line adiabatic fixed bed downflow reactors 38, 40, 46 and 52 with three intercoolers, built-in heat exchangers 39, 43 and 49. The reactor is referred to as the first adiabatic reactor 38, the second adiabatic reactor 40, the third adiabatic reactor 46 and the fourth adiabatic reactor 52, but in this description, these adiabatic reactors 38, 40, 46 and 52 can also be referred to as "reactors".

As described above, this intercooling is necessary because the reaction of water with ethylene oxide is highly exothermic, so it is necessary to remove the heat of reaction from the product or effluents if the reaction takes place in an adiabatic reactor, as in this invention, to avoid excess temperature in the reactor bed, which can reduce the life of a temperature sensitive catalyst such as ion exchange resins. This is different from isothermal conditions where the heat of reaction is removed directly from the reactor. Keeping the ion exchange resin catalyst at temperatures that the resin can withstand is of course critical to reduce swelling, prevent degradation, and maximize catalyst life. Under proper operation, the rise in adiabatic temperature can be controlled at a level that maintains good efficiency and service life.

Finally, downstream of the series of reactors 22 is a conventional non-catalytic adiabatic tubular reactor 26 (this reactor, of course, does not contain a catalyst).

Since the overhead stripper stream flows under pressure through conduit 30 into the reactor block, it is divided equally between four reactors in a row; then, a first equal portion of the highly concentrated ethylene oxide stream is mixed with a source of water 36 (preheated in heat exchanger 35), in a mixing device, such as a static mixer, to form a dilute aqueous ethylene oxide feed stream 37. Stream 36 is heated in heat exchanger 35 to bring stream 37 to a temperature of 50°C to 90°C.

A water stream 36 is also present. The water stream 36 can be taken from any water source, including the PGI, but is preferably taken from another unit of the ethylene oxide/ethylene glycol plant. For example, the water may come from recovered condensed water from an ethylene glycol flash unit of a plant, more specifically, the water may be condensed from the overhead vapor of one or more evaporators in the flash unit.

When combined, the water stream 36 and the pressurized overhead stripper stream flowing through conduit 30 form a combined dilute feed stream to a series of reactors 22. This combined feed stream will have a stoichiometric excess of water to ethylene oxide, compared to that required to hydrate the ethylene oxide to ethylene glycol, however, the amount of this excess is significantly less than the excess used in the prior art. Therefore, in accordance with the present invention, the combined streams 40 and 36 have a molar ratio of water:ethylene oxide from about 5:1 to about 15:1, preferably from about 7:1 to about 12:1. And thus, since the combined total feed stream contains less water, the product stream from the last reactor of row 88 also has correspondingly reduced amounts of water to be removed from the product, as described below.

The aqueous ethylene oxide contained in the stripper liquid product stream 40 is divided into four approximately equal streams 57, 61,64 and 67, each of these streams containing approximately equal molar amounts of water and ethylene oxide. Of these, stream 57 is combined with water stream 36 (after being heated in heat exchanger 35) to form feed stream 37 to reactor 38 such that the molar ratio of water:ethylene oxide in feed stream 37 is from about 40:1 to about 10:1, preferably , from about 30:1 to about 20:1. This is a significant improvement over conventional catalytic and non-catalytic hydration. By co-feeding stream 57 with water stream 36 to first reactor 38 and then co-feeding other equal portions (61, 64, 67) of aqueous ethylene oxide feed stream 40 with reactor effluent (70, 73, 76), a high ratio of dilute water:ethylene oxide in the feed stream to each reactor while maintaining a low overall water:ethylene oxide ratio in the product stream 88 at the end of the reactor block, which is fed to the evaporators to remove water. Thus, the reaction of ethylene oxide and water is catalyzed in the presence of an ion exchange resin and with a significant excess of water to ethylene oxide. This contributes to the selectivity to the production of monoethylene glycol and to the suppression of the formation of the higher glycol homologue, thus both effects are advantageous. And yet this is retained by obtaining an end product that is much more concentrated in ethylene glycol and therefore requires much less energy in the evaporation unit. For example, in accordance with this invention, the end product of the reactor block, the combined product from the reactor 88, may have a water:ethylene glycol ratio of 10:1 or even lower, while on the contrary, in a conventional plant, the reactor effluent fed to the evaporation block, has a water:ethylene oxide molar ratio of about 25:1, thereby reducing the amount of heat load for evaporating excess water by a factor of at least 2.5. A significant advantage of this invention is that the water:ethylene oxide molar ratio can be adjusted and optimized as needed and desired for catalyst and energy costs, ethylene oxide availability and product yield requirements.

Having illustrated the general operation of the reactor block and the advantages of the present invention, the method according to the present invention will now be described in more detail with reference to FIG.

As noted above, streams 57 and 36 are combined to form feed stream 37 to the first reactor and thus the final temperature of the catalyst bed in reactor 38 is controlled by controlling the heating of water stream 36 in heat exchanger 35, which in turn controls the inlet temperature of the stream. 37. The final temperature of the catalyst bed in each of the series of reactors 40, 46 and 52 arranged in series is controlled by the use of heat exchangers 39, 43 and 49, respectively, used to cool the feed stream to each reactor.

Temperature and pressure control are important parts of the process according to the present invention. The hydration reaction in the present invention is carried out as a liquid phase process. In the present invention, by "liquid phase process" is meant that the feed streams to the reactor, reactor product or reactor effluents, and the reactants within the reactor are maintained in the liquid phase. Accordingly, the temperature and pressure in each reactor is such that as the streams are fed into and enter each reactor, they are maintained in the liquid phase. Thus, the temperature and pressure in each reactor is controlled to keep the contents of the reactor in the liquid phase. The choice of temperature is largely determined by the considerations of the above paragraphs. Thus, at the following temperature ranges in the present invention, when the contents of the reactor are in liquid form, the pressure will be from about 0.5 MPa to about 1.5 MPa, preferably from 0.8 MPa to 1.2 MPa.

When choosing a temperature, two competing requirements must be balanced. First, the inlet temperature according to the present invention must be high enough to drive the ethylene oxide hydration reaction so that the conversion of ethylene oxide to monoethylene glycol is complete or nearly complete. In particular, by "complete", "almost complete", or "completion" is meant that the percentage of conversion of ethylene oxide to monoethylene glycol should be at least 97%, preferably at least about 99%. Thus, the inlet temperature to reactors 38, 40, 46 and 52 should be in the range of about 50°C to about 90°C, preferably in the range of about 70°C to about 85°C. However, the temperature in the reactor (and therefore at the outlet of the reactor) should not be too high, because, as mentioned above, a significant increase in temperature inside the reactor causes swelling and destruction of the ion exchange resin. The high temperatures in the reactor are the result of an exothermic hydration reaction, so this heat of reaction must be removed to contain any temperature increase. In such applications, adiabatic reactors have significantly fewer process complications, but have the disadvantage that they do not remove the heat of reaction. Accordingly, it has been found in the present invention that the temperature increase can, however, be at least partially controlled by excess water in each reactor, which acts as a heat sink. Thus, although one of the objectives of the present invention is to reduce the amount of water used in the production of ethylene glycol, nevertheless, at least a minimum amount of water is necessary to contain the temperature increase in the reactors resulting from the exothermic reaction. As described above, both of these objectives are accomplished in the present invention by dividing the aqueous ethylene oxide feed stream 40 into four equal portions 57, 61, 64, and 67 and co-feeding (the term "gradual feed" may also be used) to produce feed streams in reactor 38, 40, 46, 52. With this inventive configuration, a high water:ethylene oxide ratio is obtained in each reactor by controlling the exotherm so that the outlet temperature of each reactor is from about 85° C. to about 120° C., preferably from about 100° C to about 110°C.

Despite the success of the present invention in controlling the temperature increase in the reactor, it may nevertheless be necessary to further reduce the temperature of the reactor outlet stream to achieve the acceptable inlet temperature range mentioned above. To do this, streams 70, 73 and 76 are passed through heat exchangers 39, 43 and 49, respectively. This allows for heat integration because the stream (not shown) from outside the reactor bank can be heated in any of the heat exchangers 39, 43 and 49 by indirect heat exchange with the reactor effluent from one or more reactors in series.

As noted above, at higher temperatures almost complete conversion of ethylene oxide to monoethylene glycol occurs, at lower inlet temperatures not all of the ethylene glycol will react, and detectable amounts of ethylene oxide will appear and remain unreacted ("breakthrough") in the reactor effluent. In certain situations, it may be necessary to lower the inlet temperature in order to reduce ethylene oxide conversion to provide more ethylene oxide breakthrough, when, for example, it is desirable to extend the life of the ion exchange resin catalyst or in other situations, as described below.

In addition to the above parameters and operating conditions, other factors that may be considered during process operation include, but are not limited to, reactor and process configuration, ion exchange resin catalyst conversion and selectivity, and reactant flow rate (RFR). This unique reactor configuration provides the flexibility to select reactor operating parameters that best suit the plant's operating objectives in terms of catalyst and energy utilization, ethylene oxide availability, and product yield requirements. In addition, the production of the highly concentrated ethylene oxide 40 liquid feed stream that constitutes the present invention greatly reduces the number of steps required in an epoxidation process unit compared to conventional commercial plants, which require more steps due to the need for more dilute aqueous ethylene oxide and high temperatures for conventional non-catalytic thermal hydrolysis.

With the above considerations in mind, feed stream 37 at a suitable temperature is fed to reactor 38 where ethylene oxide is reacted with water in the presence of an ion exchange resin catalyst or catalyst bed to form monoethylene glycol. In adiabatic reactors in accordance with the present invention, for example, reactor 38, mixing ethylene oxide with hot dilution water to obtain a suitable temperature in the presence of a catalyst causes ethylene oxide hydration, which is highly selective for monoethylene glycol such that at least about 95%, preferably at at least about 98% of the ethylene oxide is converted to monoethylene glycol, and only a small proportion is converted to higher glycols such as diethylene glycol or triethylene glycol. Thus, for example, with respect to reactor 38, on a molar basis, the moles of monoethylene glycol in the effluent of reactor 70 will be approximately equal to the moles of ethylene oxide in feed stream 38. water for use in the next reactor), monoethylene glycol and trace amounts of other components, including ethylene oxide and higher glycol homologues.

Due to the pressure drop along the length of reactor 38, there is a significant pressure drop between reactor 38 and 40 for the first reactor effluent 70 to pass through heat exchanger 39 where it is cooled and then the first reactor effluent is combined with stream 61 to form a second feed stream. 81, which is fed into reactor 40 as in reactor 38, where ethylene oxide is reacted with water with high selectivity and conversion such that substantially all of the ethylene oxide is reacted with a stoichiometric amount of water to form monoethylene glycol at the above conversion percentages. The second reactor effluent, which, like the first reactor effluent, contains unreacted water, monoethylene glycol, and trace amounts of the aforementioned components, passes under pressure through heat exchanger 43. In fact, the only difference between the first reactor effluent 70 and the second reactor effluent 73 is that that the latter contains more monoethylene glycol because it contains the amount of monoethylene glycol produced in both reactors 38 and 40. Reactors 46 and 52 and their associated streams operate in the same manner as 38 and 40, where the second reactor 73 effluent is cooled into heat exchanger 43 and combined with an equal portion 64 of an aqueous concentrated ethylene oxide feed stream 40 to form a third feed stream 75 which is reacted in the presence of a third ion exchange resin catalyst bed in a third reactor 46 to form a third reactor effluent 76. Third reactor effluent 76 is cooled in a heat exchanger 49 and combined with an equal portion of 67 to form a fourth feed stream 77 which is reacted in the presence of a bed of a fourth ion exchange resin catalyst in a fourth reactor 52 to form a fourth reactor effluent 79 where each effluent is sequentially more enriched in monoethylene glycol as described higher.

In addition to the series of reactors 22 described above, including four downflow fixed bed reactors in series, the reactor bank 25 also includes a conventional tubular reactor 26 which can be fed with the effluent of the fourth reactor 79 after being heated in the heat exchanger 55. In the present invention, this tubular reactor 26 is configured and operated to impart previously unknown levels of versatility and controllability to the catalytic hydration process scheme, which is described in more detail below.

First, in addition to improving the conversion and selectivity of the reaction of ethylene oxide and water to MEG, the present invention can also be used and configured to increase the versatility and adaptability of the hydration process, allowing the process operator to produce di- and triethylene glycol. This can be done by physically bypassing one or more catalytic reactors 38, 40, 46, 52 (this bypass itself is not shown) and thus removing one or more catalytic reactors from the circulation of the feed stream. The ethylene oxide that "breaks through" and remains unreacted after bypassing one or more of the aforementioned catalytic reactors is sent to a non-catalytic tubular reactor 26 to thermally convert the ethylene oxide and water to ethylene glycol. Since tubular reactor 26 is non-catalytic and is entirely dependent on thermal conversion, the fourth reactor 79 effluent is preheated in heat exchanger 55 to fall within the effective temperature range for non-catalytic hydration of ethylene oxide to monoethylene glycol, preferably the inlet temperature of tubular reactor 26 is from about 130°C to about 160°C.

Of course, this non-catalytic thermal conversion is significantly less selective for MEG, producing significant amounts of higher grade glycols than MEG, such as DEG and TEG, so that by increasing the amount of ethylene oxide that is converted to ethylene glycol in the tubular reactor , the product of the combined reactor 88 contains less MEG and more DEG and TEG than would be produced if only catalytic reactors were used in the series of reactors 22. This gives the operator a versatile and flexible process for the production of ethylene glycol, since the production of mono-, di- and triethylene glycol can adjusted by the operator to meet demand. In the extreme case, all four consecutive reactors are bypassed so that only the last non-catalytic tubular reactor 26 remains, and all hydration will take place in this reactor, thereby making the catalytic process non-catalytic. Alternatively, catalyst bank 22 adds the flexibility to reduce the plant's energy consumption while maintaining a fixed monoethylene glycol production compared to the same production in a plant using non-catalytic EO hydrolysis. If about half of the ethylene glycol production is made in the catalytic reactors 38 and 40 of a row of reactors 22 (bypassing reactors 46 and 42), and the rest is made in the tubular thermal reactor 26, the resulting product distribution will be the same as in a conventional non-catalytic thermal hydrolysis process, but with the advantage of operating the unit at about 20% lower water:ethylene oxide molar ratio, resulting in significant savings in glycol evaporation.

Alternatively, instead of physically bypassing the reactors, a "thermal bypass" can be implemented by lowering the inlet temperature of one or more reactors in series to the required low level to reduce the reaction between ethylene oxide and water and reduce the amount of monoethylene glycol formed in the reactors, which means more "breakthrough" of unreacted ethylene oxide observed at the outlets of the catalytic reactors. This ethylene oxide is then converted to ethylene glycol in a tubular reactor as described above.

The tubular reactor may also play a "finishing" role. Even when a row of reactors 22 is operating at normal catalytic efficiency designed to maximize the conversion of ethylene oxide to ethylene glycol, however, a small amount of unreacted ethylene oxide will remain unreacted after all four reactors in row 22, and this ethylene oxide will be a "breakthrough" in the effluent of the fourth reactor 79 In the present invention, row 26 operates normally such that the concentration of ethylene oxide in the effluent of fourth reactor 79 will not exceed about 1 mol%. This remaining ethylene oxide can then be converted to monoethylene glycol in tubular reactor 26.

Finally, the tubular reactor 26 can also be used as a backup reactor in situations where one or more of the catalytic reactors fail or must otherwise be taken out of service due to depletion, plugging of the resin bed or nozzle, or other difficulties commonly seen in bed operation. ion exchange resin. When this happens, the failed catalytic reactor must be taken out of service. This means that there will be an excess of unreacted ethylene oxide in reactor row 22, which is undesirable. This excess ethylene oxide breakthrough can be converted to ethylene glycol in tubular reactor 26. Thus, the presence of tubular reactor 26 in the present invention provides considerable flexibility and operability for the process operator. In contrast to the prior art, where the inability to replace or repair the ion exchange resin catalyst or reactor is a serious disadvantage of solid/heterogeneous catalyst systems, in the present invention the reactors can be taken offline for such repairs without interrupting operation.

Although four reactors and three heat exchangers are shown in the reactor block of Figures 2 and 3, this is for purposes of illustration only. More or fewer reactors may be used depending on specific needs and requirements. Increasing the number of reactors has the advantage of reducing the exotherm in each reactor, thereby reducing the amount of irreversible temperature-induced swelling that the ion exchange resin catalyst experiences within each reactor. In addition, increasing the number of reactors also increases ethylene oxide dilution, and since ethylene oxide destroys most ion exchange resins, increased dilution is expected to prolong resin life. In addition, the total volume of the catalyst is reduced.

The primary disadvantages of more reactors are cost and complexity, including additional capital costs for reactors, pumps, heat exchangers, and other equipment, as well as operating costs for large volumes of ion exchange resin catalyst, as well as additional piping, instrumentation, and operational complexity.

The alkylene glycol in the combined product of the reactor 88 can then be recovered using an evaporation battery or vacuum distillation, or any other suitable method known to those skilled in the art. Recovery methods can be combined for different stages of recovery. For example, an evaporator battery can be used to remove water from the reactor effluent, while vacuum distillation can be used to dry the glycol. Indeed, one of the advantages of the present invention is that complexity can be reduced in the plant's post-glycol stack of reactors because the amount of heat load required to boil off excess water has been reduced, as noted above and as detailed in the example below. Thus, although prior art plants often include evaporation systems having multiple effects or stages (see, for example, EP No. 2121646 B1), the present invention requires fewer stages.

As discussed above, this invention may include one or more ion exchange resin catalysts. Ion-exchange resins have a polymer matrix that contains ion exchange centers filled with ionic functional groups on the surface. Ion exchange resins are usually distinguished into cationic and anion exchange resins, although other types of ion exchange resins are available. For a more complete description of cation and anion exchange resins, see de Dardel, F. and Arden, TV Ion Exchangers in Ullman's Encyclopedia of Industrial Chemistry (2005).

Suitable polymeric matrices for the ion exchange resin catalyst include polystyrene matrix, polyacrylic matrix, polyalkylamine resin, and other polymeric materials. Preferably, the polymeric matrix is cross-linked with divenylbenzene sufficiently to increase operating power, but without increasing the density of the ion exchange material to the extent that the ion exchange material becomes too physically hard and too chemically resistant to chemical processing. Preferably, the matrix is a copolymer of styrene and divinylbenzene.

Attached to the regions of the polymer matrices described above are ionic stationary groups which determine whether the resin functions as a cationic or anionic ion exchange resin. In solutions, the positive or negative charge of stationary groups is compensated by ions with the opposite charge, which are called here functional groups.

Strongly acidic cationic ion exchange resins typically include sulfonic groups as stationary groups, in turn attached to the styrene-divenylbenzene polymer matrix. Examples of strongly acidic sulfonic acid cation exchange resins include Amberlite IR 120, Dowex HCR, Lewatit S 100 and Amberlyst 15, among others.

Cationic ion exchange resins may also include a polymeric material that is a copolymer obtained by addition polymerization of acrylic or methacrylic acid and divinylbenzene, such as described in US patent No. 3,957,698. Other suitable polymeric materials for cation exchange resins include a polymeric matrix obtained by crosslinking a polyvinylpyridine resin with divenylbenzene. Such materials are available under the brand name Reillex HPQ. Sulfonated phenolic polymer resins are also suitable cationic ion exchange resins.

Weak acid cationic ion exchange resins typically include carboxyl groups as stationary groups. Examples of weak acid cation exchange resins include, Amberlite IRC 86, Dowex Mac-3, Lewatit CNP, among others.

In the present invention, the strongly basic anion exchange resins contain quaternary ammonium stationary groups. They are further subdivided into type I, obtained by the reaction of trimethylamine with a styrene-divenylbenzene copolymer after chloromethylation, and type II, obtained by the reaction of a styrene-devenylbenzene copolymer with dimethylethanolamine. Suitable examples of such Type I resins include Lewatit MP 500 available from Lanxess and Amberlyst 26 and Amberlite IRA 402 and IRA 410 available from Dow. Suitable examples of strongly basic Type II resins include, for example, Purolite A510S (Purolite Corporation).

Weak base anion exchange resins typically include polyacrylic resins that have received stationary groups by reaction with a polyfunctional amine to form anion exchange resins such as the weak base tertiary ammonium resins Amberlite IRA 67 and Amberlyst 21 (available from Dow). Of particular note, this ion exchange resin can then be further treated with chloromethane or dimethyl sulfate to form a Type I strong base quaternary amine resin Amberlite IRA 458 (Dow). Weak base anion exchange resins may also include a free base group as a stationary group, such as Amberlite IRA-67 resin (Dow).

In a particular embodiment of the present invention, the strongly basic anion exchange resin contains a "bonding" group of 3-7 bonding atoms, in particular 3-5 carbon atoms between the quaternary ammonium stationary group and the benzene group of the polymer matrix/material. An example of such a link chain in a strong base resin is illustrated in US Pat. No. 5,945,568 and was produced under the name Diaion TSA1200 (Mitsubishi Chemical). As stated in the '568 patent, the linking group that links the stationary quaternary ammonium group to the benzene group of the polymer matrix is not particularly limited as long as it is long enough. Suitable examples of the linking group are an alkylene group or an alkyleneoxyalkylene group. A preferred example for use in this invention is an anion exchange resin which contains, as a substrate, a polymer of a vinylaromatic compound, and which has such a structure that the quaternary ammonium group is linked to the aromatic group through a linking group having a chain length of at least 3.

Without being limited by theory, it is believed that in this invention, this linking group prevents nucleophilic attack on the ammonium stationary group from the benzene ring, which would otherwise lead to accelerated deactivation and degradation of the ion exchange resin catalyst and reduce the stability and life of the catalyst.

A variety of ions are suitable for use as a functional group in the anionic resins of this invention and may be selected from the group consisting of bicarbonate, carbonate, hydroxide, bisulfite, carboxylate, citrate and metalate, and molybdate anions. These functional groups can be used with any of the stationary groups and polymer material mentioned above.

In the present invention, it is preferable that the ion exchange resin catalyst belongs to the group of strongly basic type I anion exchange resins, more preferably, strongly basic type I anion exchange resins with a bicarbonate or monocitrate functional group, and most preferably, strongly basic type I anion exchange resins with a bicarbonate or monocitrate functional group. with a link group.

In addition to the ion exchange materials mentioned above, any other suitable cation or anion exchange resin may be used, such as the cation and anion exchange resins mentioned above in de Dardel, F. and Arden, TV 2008, Ion Exchanger , in Ullman's Encyclopedia of Industrial Chemistry .

In addition to the above ion exchange material, a combination of a homogeneous and a heterogeneous catalyst, such as a solid catalyst on which a soluble metal catalyst is adsorbed, can also be used. In this case, a solid catalyst is first prepared with a metal catalyst adsorbed thereon. Then, during the liquid phase hydration reaction described above, the metal catalyst is desorbed from the solid catalyst and can effectively catalyze the reaction. Then the metal catalyst must be separated from the reaction products and can be reused for re-adsorption on the solid catalyst. The solid catalyst in this case may be an ion exchange resin.

The above paragraphs relate to the selection of a suitable ion exchange resin catalyst for use in the catalytic reactors in accordance with the present invention. With respect to the operation of the catalytic reactors containing the ion exchange resin, in Figures 2 and 3 the feed flows from the top of the reactor down through each of the reactors. This is called the "downflow" mode, but in the present invention, the reactors can be operated in "downflow" or "upflow" modes. Downflow processes have the advantage of increasing the density or volume of the catalyst within the reactor and thus reducing the size and cost of the reactors themselves. Downflow operation also minimizes the non-catalytic reaction of water and ethylene oxide to ethylene glycol. However, during downflow operation, the selectivity of the catalyst can be degraded, both due to local inhomogeneities and impurities that can form and be trapped in the catalyst bed, and due to "canalization" that develops in the catalyst bed.

As can be expected, in upflow operation, the feed stream moves from the bottom of the reactor to the top. PCT Publication No. WO2008/150338A1 reports an increase in the stability and service life of an ion exchange resin in upflow operating mode. In the present invention with liquid phase reactants and solid catalyst, upflow operation can produce a moderate degree of fluidization of the bed during upflow operation, which increases the proportion of voids between adjacent resin particles, allowing reduction of inhomogeneities and impurities that would otherwise be embedded and trapped in the catalyst bed. during downflow operation and degrade catalyst efficiency. US Pat. No. 6,160,187 has previously underestimated the performance of upflow catalytic reactors because the resulting bed fluidization supposedly causes more catalyst attrition and reduced monoethylene glycol selectivity due to longitudinal agitation. However, in the present invention, it has been found that in the liquid phase process of the present invention, these criticisms are unrealistic because they are more likely to occur in gas/solid catalyst systems where there are much higher velocities, large differences in densities (between the gaseous reactant and the solid catalyst, according to compared to liquid reactant and solid catalyst) and thus much greater bed expansion than in the liquid/solid catalyst system of the present invention. Thus, upflow operation remains a suitable mode of operation in accordance with the present invention.

Silver-based epoxidation catalyst.

The silver-based epoxidation catalyst includes a support and at least a catalytically effective amount of silver or a silver-containing compound; an activating amount of rhenium or a rhenium-containing compound is also optionally present; an activating amount of one or more alkali metals or alkali metal-containing compounds is also optionally present. The substrate used in accordance with this invention may be selected from a wide variety of solid refractory substrates, which may be porous and may provide a preferred porous structure. It is well known that alumina is useful as a catalyst support for olefin epoxidation and is the preferred support.

Regardless of the nature of the support used, it is usually in the form of particles, large pieces, small pieces, granules, rings, spheres, wagon wheels, baffled hollow cylinders and the like, the size of which is suitable for use in a fixed bed epoxidation reactor. The substrate particles will preferably have equivalent diameters in the range of about 3 mm to about 12 mm, and more preferably in the range of about 5 mm to about 10 mm. The equivalent diameter is the diameter of a sphere having the same ratio of outer surface (i.e., excluding surface in the pores of the particle) to volume as the particles of the substrate used. Suitable substrates are available from Saint-Gobain Norpro Co., Sud Chemie AG, Noritake Co., CeramTec AG and Industrie Bitossi S.p.A. Without being limited to the specific compositions and formulations contained herein, additional information on substrate compositions and methods for preparing substrates can be found in US Patent Publication No. 2007/0037991.

To prepare a catalyst for oxidizing an olefin to an olefin oxide, a catalytically effective amount of silver is then applied to the surface of a support having the above characteristics. In one embodiment, the catalytically effective amount of silver is from 10% by weight to 45% by weight. The catalyst is prepared by impregnating the support with a silver compound, complex or salt dissolved in a suitable solvent sufficient to cause precipitation of the silver precursor compound onto the support. Preferably, an aqueous solution of silver is used.

An activating amount of the rhenium component, which may be a rhenium-containing compound or a rhenium-containing complex, may also be applied to the substrate before, simultaneously with, or after the silver is deposited. The rhenium promoter may be present in an amount of from about 0.001 wt%. up to about 1% wt., preferably from about 0.005% wt. up to about 0.5% of the mass. and more preferably, about 0.01% of the mass. up to about 0.1% of the mass. calculated on the weight of the entire catalyst, including the substrate, in terms of metallic rhenium.

Other components that may also be applied to the substrate before, simultaneously with or after the deposition of silver and rhenium are activation amounts of an alkali metal or a mixture of two or more alkali metals, as well as optional activation amounts of a Group IIA alkaline earth metal component or a mixture of two or more components a Group IIA alkaline earth metal, and/or a transition metal component, or a mixture of two or more transition metal components, all of which may be in the form of metal ions, metal compounds, metal complexes, and/or metal salts, dissolved in a suitable solvent. The support may be impregnated simultaneously or in separate stages with different catalyst promoters. A particular combination of support, silver, alkali metal promoter(s), rhenium component and optional additional promoter(s) in accordance with this invention provides an improvement in one or more catalytic properties over the same combination of silver and support in the absence or use of only one of promoters.

As used herein, the "activating amount" of a particular catalyst component refers to the amount of that component that works effectively to improve the catalytic performance of a catalyst compared to a catalyst that does not contain that component. The exact concentrations used will, of course, depend on the desired silver content, the nature of the support, the viscosity of the liquid, and the solubility of the particular compound used to deliver the promoter to the impregnation solution, among other factors. Examples of catalytic properties include, among others, functionality (resistance to degradation), selectivity, activity, conversion, stability, and yield. One skilled in the art will recognize that one or more individual catalytic properties may be improved by an "activating amount" while other catalytic properties may or may not be improved, or may even be reduced.

Suitable alkali metal promoters may be selected from lithium, sodium, potassium, rubidium, cesium, or combinations thereof, with cesium being preferred, and combinations of cesium with other alkali metals being particularly preferred. The amount of alkali metal deposited or present on the support must be an activating amount. Preferably, the amount varies from about 10 ppm. up to about 3000 ppm, more preferably from about 15 ppm. up to about 2000 ppm and even more preferably from about 20 ppm. up to about 1500 ppm and particularly preferably from about 50 ppm. up to about 1000 ppm mass from the entire catalyst, in terms of metal.

Suitable alkaline earth metal promoters include elements from Group IIA of the Periodic Table of the Elements, which may be beryllium, magnesium, calcium, strontium and barium, or combinations thereof. Suitable transition metal promoters may include elements from Groups IVA, VA, VIA, VIIA and VIIIA of the Periodic Table of the Elements and combinations thereof.

The amount of alkaline earth metal promoter(s) and/or transition metal promoter(s) deposited on the support is the activating amount. The transition metal promoter may typically be present in an amount of from about 0.1 micromoles per gram to about 10 micromoles per gram, preferably from about 0.2 micromoles per gram to about 5 micromoles per gram.

The silver solution used to impregnate the support may also contain an optional solvent or complexing/solubilizing agent known in the art. A variety of solvents or complexing/solubilizing agents can be used to solubilize the silver to the desired concentration in the impregnation medium. Suitable complexing/solubilizing agents include amines, ammonia, oxalic acid, lactic acid, and combinations thereof. Amines include alkylenediamine containing from 1 to 5 carbon atoms. In one preferred embodiment, the solution contains an aqueous solution of silver oxalate and ethylenediamine. The complexing/solubilizing agent may be present in the impregnating solution in an amount of from about 0.1 to about 5.0 moles per mole of silver, preferably from about 0.2 to about 4.0 moles, and more preferably from about 0.3 to about 3 .0 mole for every mole of silver.

When a solvent is used, it may be an organic solvent or water and may be polar or substantially or completely non-polar. In general, the solvent must have sufficient solvating power to dissolve the components of the solution. At the same time, it is preferred that the solvent be chosen so as to avoid undue influence on or interaction with the solvated promoters. Organic based solvents containing from 1 to about 8 carbon atoms per molecule are preferred. Mixtures of several organic solvents or mixtures of organic solvents with water may be used, provided such mixed solvents function as desired herein.

The concentration of silver in the impregnating solution is typically from about 0.1% by weight up to the maximum solubility provided by the specific solvent/solubilizing agent combination used. It is generally very convenient to use solutions containing from 0.5% to about 45% by weight silver, with concentrations of 5 to 35% by weight silver being preferred.

The impregnation of the selected substrate is carried out by any of the conventional methods; e.g. excess solution impregnation, moisture capacity impregnation, spray application, etc. Typically, the substrate material is placed in contact with the silver containing solution until a sufficient amount of the solution is absorbed by the substrate. Preferably, the amount of silver-containing solution used to impregnate the porous substrate is not more than is necessary to fill the pores of the substrate. A single impregnation or a series of impregnations with or without intermediate drying may be used, depending in part on the concentration of the silver component in the solution. Impregnation methods are described, for example, in US Pat. Techniques known in the art for pre-precipitation, co-precipitation and post-precipitation of different promoters can be used.

After impregnation of the substrate with a silver-containing compound, i.e. silver precursor, rhenium component, alkali metal component, and optional other promoters, the impregnated support is fired for a time sufficient to convert the silver-containing compound to active silver species and to remove volatiles from the impregnated support to form a catalyst precursor. The firing may be performed by heating the impregnated substrate, preferably gradually, to a temperature in the range of about 200° C. to about 600° C. under a pressure in the range of 0.05 to about 3.5 MPa. In general, the higher the temperature, the shorter the required heating period. A wide range of heating periods has been proposed in the art; for example, US Pat. No. 3,563,914 discloses heating for less than 300 seconds and US Pat. about 8 hours. However, it is only important that the heating time correlates with the temperature so that all of the silver contained is converted to active silver species. For this purpose, continuous or stepped heating can be used.

During firing, the impregnated substrate may be exposed to a gas atmosphere containing an inert gas or an inert gas mixture with from about 10 ppm. up to 21% by volume of the oxygen-containing oxidizing component. For the purposes of this invention, an inert gas is defined as a gas that does not substantially react with a catalyst or catalyst precursor under the conditions chosen for firing. Further information on the manufacture of the catalyst can be found in the aforementioned US Patent Publication No. 2007/0037991.

For illustrative purposes only, the following conditions are often used in modern commercial ethylene oxide reactor plants: gas hourly space velocity (GSP) 1500-10000 h ^-1 , pressure at the reactor inlet from 1 MPa to 3 MPa, coolant temperature 180 -315°C, oxygen conversion rate 10-60% and EO production rate (operating rate) 100-350 kg EO/m ³ catalyst/h and change in ethylene oxide concentration, ΔEO, from about 1.5% to about 4.5% . The feedstock composition at the reactor inlet after startup is complete and during normal operation typically contains (in % by volume) 1-40% ethylene, 3-12% O ₂ ; 0.2%-10%, preferably 0.2%-6%, more preferably 0.2%-5% CO ₂ ; 0-5% ethane, some of one or more chloride moderators as described herein; and the rest of the raw material is argon, methane, nitrogen, or mixtures thereof.

EXAMPLES

The invention is further described in more detail with the help of the following non-limiting examples.

Example 1 (Prior Art)

The prior art ethylene glycol production process is simulated using a single non-catalytic tubular/tube reactor. In the process, a feed stream is fed into a reactor from which a reactor effluent is obtained and the table below gives the values in kilograms per hour:

Raw material flow Reactor Outgoing stream Temperature 150 -- 190 Ethylene oxide 1000 -- 0 Water 10191 -- 9798 MEG 45 -- 1315 Heavy glycols 0 -- 124 EO Conversion (%) -- 100% -- SWE (mol) 25 -- --

Above, "WEO" means the molar ratio of water to ethylene oxide. This prior art process produces ethylene glycol with good conversions, but uses a large excess of water, which must then be removed from the ethylene glycol product, making the process significantly more energy intensive than the process of the present invention, as described below.

Example 2 (present invention)

The method for producing ethylene glycol in accordance with this invention is simulated according to figure 4. This option is configured in the same way as the block of reactors 25 in figure 2 described above, where the conventional tubular thermal reactor 26 is equivalent to a tubular reactor "DD".

As shown in Figure 4, the simulation shows four downflow fixed bed catalytic reactors arranged in series with three intermediate inline heat exchangers as heating means. In this simulation, the reactors are filled with one or more suitable ion exchange resins as described above. The method operates in accordance with the present invention such that the temperatures and compositions of the streams, entry and exit points, and reactors are as indicated in Table I below, where the contents of the streams in the table are given in units of kilograms per hour.

Flow "Heated" flow Flow Flow Entrance Reactor Exit Flow Flow A B C D E F G H I Temperature (°C) 45 84 45 78 78 -- 110 84 45 Ethylene oxide 1000 0 250 250 250 -- 22 22 750 Water 1000 3000 250 3250 3250 -- 3157 3157 750 MEG 0 0 0 0 0 -- 318 318 0 Heavy glycols 0 0 0 0 0 -- 2 2 0 EO Conversion (%) -- -- -- -- -- 91% -- -- -- SWE (mol) -- -- -- -- -- 32:1 -- -- -- Flow Flow Flow Entrance Reactor Exit Flow Flow Flow J K L M N O P Q R Temperature (°C) 45 45 79 79 -- 110 86 45 45 Ethylene oxide 250 500 272 272 -- 25 25 250 250 Water 250 500 3407 3407 -- 3307 3307 250 250 MEG 0 0 318 318 -- 659 659 0 0 Heavy glycols 0 0 2 2 -- 10 10 0 0 EO Conversion (%) -- -- -- -- 91% -- -- -- -- SWE (mol) -- -- -- -- 31:1 -- -- -- -- Flow Entrance Reactor Exit Flow Flow Entrance Reactor Exit S T U V W Y Z AA BB Temperature (°C) 81 81 -- 110 88 84 84 -- 110 Ethylene oxide 275 275 -- 24 24 274 274 -- 25 Water 3557 3557 -- 3457 3457 3706 3706 -- 3607 MEG 659 659 -- 998 998 998 998 -- 1331 Heavy glycols 10 10 -- 21 21 21 21 -- 37 EO Conversion (%) -- -- 91% -- -- -- -- 91% -- SWE (mol) -- -- 32:1 -- -- -- -- 33:1 -- Flow Reactor Exit CC DD EE Temperature (°C) 160 -- 163 Ethylene oxide 25 -- 0 Water 3607 -- 3600 MEG 1331 -- 1351 Heavy glycols 37 -- 50 EO Conversion (%) -- 100% -- SWE (mol) -- 360:1 --

As can be seen from the table above, water and ethylene oxide are fed into the process in a 10:1 molar ratio (a combination of streams A and B), which is much less than in a prior art process such as that shown in Example 1 above. However, by dividing the input streams into multiple streams fed to multiple reactors as described in this invention, much higher ratios of water to ethylene oxide are obtained in each reactor, as shown above, the molar ratio of water to ethylene oxide in the streams entering the reactor is approximately 32:1 in each case. This high HEO not only improves the selectivity and conversion of ethylene oxide to MEG (as shown above, the conversion to MEG in this simulation is 99%), but also helps to reduce temperature rise in each reactor, reducing the amount of coolant needed in the intermediate heat exchangers and lowering the high temperatures that could potentially damage the ion exchange resin catalyst.

Moreover, as noted above, the use of a lower ROEC in the process according to the present invention allows significant energy savings compared to the prior art. In this simulation of the process according to the invention (Example 2) and the simulation of the prior art comparative process (Example 1), the final evaporative load is:

General SVEO Evaporative load (t/h) The present invention (example 2) 10:1 196 Prior Art (Example 1) 25:1 523

Thus, the prior art method has a evaporation load 2.5 times higher than the method according to the present invention.

Those skilled in the art will appreciate that variations may be made to variations of the liquid phase process described above without departing from the broad concept of the invention. Thus, it is to be understood that this invention is not limited to the specific embodiments, configurations, or compositions disclosed, but is intended to cover modifications within the spirit and scope of this invention as defined in the appended claims.

Claims (40)

1. Liquid-phase method for producing ethylene glycol, including: obtaining a water stream and an aqueous ethylene oxide feed stream; separating the aqueous ethylene oxide feed stream into at least a first equal part and a second equal part; combining the first portion and the water stream to form a first feed stream; feeding the first feed stream to the inlet of the first adiabatic reactor, where the inlet of the first adiabatic reactor has an inlet temperature; reacting ethylene oxide and water in the presence of a first ion exchange resin catalyst in a first adiabatic reactor to produce a first reactor effluent containing water, ethylene glycol and unreacted ethylene oxide; removing the first reactor effluent from the outlet of the first adiabatic reactor; cooling the first reactor effluent through a heat exchanger downstream of the first adiabatic reactor to a hydration temperature; And combining the second portion and the cooled product stream to form a second feed stream; transferring the second feed stream to the inlet of the second adiabatic reactor, where the inlet of the second adiabatic reactor has a hydrolysis temperature; reacting ethylene oxide and water contained in the second feed stream in the presence of a second ion exchange resin catalyst in a second adiabatic reactor to produce a second reactor effluent containing water, ethylene glycol and unreacted ethylene oxide; where each of the first adiabatic reactor and the second adiabatic reactor contain an ion exchange resin catalyst in a fixed catalyst bed, and the outlet temperature of the adiabatic reactors is higher than the inlet temperature, and where the total water:ethylene oxide molar ratio in the water stream and the added aqueous ethylene oxide feed stream together is from 5:1 to 15:1, preferably from 7:1 to 12:1. 2. Liquid phase method according to claim 1, wherein the inlet temperature is between 50 and 90°C. 3. Liquid-phase method according to paragraphs. 1, 2, where the outlet temperature is from 85 to 120°C. 4. Liquid-phase method according to paragraphs. 1-3, where the molar ratio of water:ethylene oxide in the first feed stream is from 40:1 to 10:1, preferably from 30:1 to 20:1. 5. Liquid-phase method according to paragraphs. 1-4 held continuously. 6. Liquid-phase method according to paragraphs. 1-5, where the aqueous ethylene oxide feed stream is further divided into a third equal part and a fourth equal part, and which includes the following stages: removing the second reactor effluent from the outlet of the second adiabatic reactor; cooling the second reactor effluent through a heat exchanger downstream of the second adiabatic reactor; combining the third portion and the cooled effluent from the second reactor to form a third feed stream; transferring the third feed stream to the inlet of the third adiabatic reactor; reacting ethylene oxide and water contained in the third feed stream in the presence of a third bed of ion exchange resin catalyst in a third adiabatic reactor to produce a third reactor effluent containing water, ethylene glycol and unreacted ethylene oxide; extracting the effluent of the third reactor from the outlet of the third adiabatic reactor; cooling the outgoing stream of the third reactor through a heat exchanger located after the third adiabatic reactor; combining the fourth portion and the cooled effluent from the third reactor to form a fourth feed stream; transferring the third feed stream to the inlet of the fourth adiabatic reactor; And reacting ethylene oxide and water contained in the fourth feed stream in the presence of a fourth bed of ion exchange resin catalyst in a fourth adiabatic reactor to form a fourth reactor effluent containing water, ethylene glycol and unreacted ethylene oxide. 7. The liquid phase process of claim 6, further comprising feeding the fourth reactor effluent to a non-catalytic reactor. 8. Liquid-phase method according to paragraphs. 1-7, where the aqueous ethylene oxide feed stream is obtained according to the following steps: obtaining a stream of water from the enrichment cycle containing ethylene oxide, methane, ethylene and other dissolved light gases; separating, in the flash drum, the dissolved light gas vapor from the water stream from the enrichment cycle; directing upward vapor of the dissolved light gas through an opening in the flash drum allowing fluid communication with an absorber mounted on the top surface of the flash drum to form an absorber vapor overhead stream; heating water from the enrichment cycle, obtained from the bottoms of the flash drum, and pumping it into the desorber; And separating into: (1) an enriched stripper overhead liquid stream containing at least 40 mol%, preferably at least 50 mol% ethylene oxide; and (2) a lean aqueous stripper bottoms solution containing 5 to 50 molar ppm ethylene oxide. 9. Liquid-phase method according to paragraphs. 1-8, wherein the ion exchange resin catalyst is a type I strongly basic anion exchange resin. 10. Liquid-phase method according to paragraphs. 1-9, wherein the ion exchange resin catalyst has a bicarbonate or monocitrate functional group. 11. Liquid-phase method according to paragraphs. 1-10 wherein the ion exchange resin catalyst includes a linking group. 12. Liquid-phase method according to paragraphs. 1-11 wherein the first and second feed streams are substantially free of carbon dioxide. 13. The liquid phase process of claim 9, wherein the type I strongly basic anion exchange resin includes a stationary quaternary ammonium group.

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