Improved ethanol dehydration process

Abstract

A process for dehydrating ethanol to ethylene can reduce steam consumption and utilize latent heat of condensation to produce steam.

Description

Technical Field

[0001] This invention relates to a method for dehydrating ethanol to ethylene and oxidizing ethylene to ethylene oxide. Background Technology

[0002] Ethylene oxide molecules exist in minute quantities within the vast clouds of cosmic dust, molecules that have long been a natural component of our universe. However, it wasn't until 1859 that this molecule was synthesized on Earth, a feat achieved by Charles-Adolphe Wurtz through a groundbreaking chlorohydrin process. The full industrial potential of this compound remained dormant until the rise of the automotive industry and the increased demand for ethylene glycol (an antifreeze extracted from ethylene oxide). This necessity catalyzed further development, leading Thèodore Lefort to discover a more economically viable production method in 1931—the direct catalytic oxidation of ethylene.

[0003] Since then, ethylene oxide production has skyrocketed, becoming one of the most produced chemicals, with global production projected to reach 34 billion tons by 2020. (Most of this ethylene oxide is further processed into derivatives such as ethylene glycol.) With the increase in production, research into ethylene oxide catalysis and processes has flourished. One area of ​​particular interest is developing more efficient and sustainable production processes, such as the production of ethylene oxide from ethanol. Ethanol can be produced from renewable resources, providing a sustainable alternative to conventionally petrochemical-derived ethylene. The conversion process involves the dehydration of ethanol to form ethylene, which is then oxidized to ethylene oxide.

[0004] While using ethanol as an alternative source to ethylene addresses sustainability concerns, conventional petrochemical ethylene sources and ethanol-based ethylene oxide production processes are energy-intensive. Ethanol dehydration is an endothermic reaction, meaning it requires an input of heat. This heat is typically provided as steam at a 3:1 dilute weight ratio of steam to ethanol. While dilution dehydration offers advantages such as near 100% conversion and selectivity, the demand for steam input at this dilution ratio is substantial. Therefore, despite the environmental advantages of using renewable feedstocks like ethanol, the energy requirements diminish its sustainability benefits.

[0005] Therefore, there is still a need for an ethanol dehydration process that combines its sustainability advantages with high conversion rates and selective operation, while also reducing its external energy requirements. Summary of the Invention

[0006] This invention relates to a method for dehydrating ethanol to ethylene, comprising the following steps: (a) mixing an ethanol vapor stream with a superheated steam stream to form a reactor feed mixture; (b) supplying the reactor feed mixture to an ethanol dehydration reactor; (c) dehydrating the ethanol in the reactor feed mixture in the ethanol dehydration reactor to form an ethylene product stream; (d) supplying the ethylene product stream to a steam generator at a pressure of about 0.2 MPa to about 0.3 MPa and a temperature of about 125°C to about 140°C; (e) cooling the ethylene product stream below its dew point to release heat of condensation; and (f) generating a second process steam stream using the heat of condensation. (g) The temperature of the second processing steam stream is about 95°C to about 120°C; (h) The processing steam stream is discharged from the steam generator at a pressure of about 0.125 MPa to about 0.2 MPa; and (h) The processing steam stream is compressed to a pressure of about 0.3 MPa to about 0.45 MPa. Attached Figure Description

[0007] The invention will be more readily understood by reading the above description and the following detailed description of preferred embodiments in conjunction with the accompanying drawings. For the purpose of illustrating the invention, the drawings show currently preferred embodiments. However, it should be understood that the invention is not limited to the precise arrangements and means shown. In the drawings: Figure 1 shows a schematic flow chart of a conventional prior art ethanol dehydration process; Figure 2 shows a schematic flow chart of the ethanol dehydration process according to the present invention; Figure 3 shows a schematic flow chart simulating the temperature and pressure values ​​of a conventional ethanol dehydration process. Figure 4 shows a schematic flowchart of the temperature and pressure values ​​of the simulated ethanol dehydration process according to the present invention. Detailed Implementation

[0008] Unless otherwise stated, all parts, percentages, and ratios used in this article are expressed in terms of volume. All references cited in this article are incorporated herein by reference.

[0009] "Water" refers to any type of water suitable for chemical and petrochemical processes, including deionized water, demineralized water, industrial water, drinking water, and distilled water.

[0010] "Steam" refers to a combination of inert gases, or non-condensable inert gases containing saturated water vapor. "Superheated steam" is defined as steam obtained by further heating saturated steam at its saturation temperature.

[0011] Unless otherwise stated, pressure is expressed in MPa (absolute pressure).

[0012] This invention introduces an ethanol dehydration process with improved energy efficiency. Even with a high (“diluted” steam-to-ethanol ratio, high energy and steam consumption are reduced, ensuring a highly efficient process where almost 100% of the ethanol is converted to ethylene. This ethanol dehydration process is particularly suitable for the production of ethylene oxide from ethanol.

[0013] The process of the present invention will now be described with reference to Figure 1, which shows a conventional prior art ethanol dehydration process 1.

[0014] In process 1, liquid ethanol is supplied from an external source / OSBL at an ambient temperature, preferably between about 10°C and 30°C. The liquid ethanol passes through heat exchanger 5, with ethylene product stream 60 located on the other side of the heat exchanger. Thus, the liquid ethanol is evaporated through indirect heat exchange with the ethylene product stream 60 in heat exchanger 5, generating an ethanol vapor stream 10. The ethanol vapor stream 10 from heat exchanger 5 passes through heat exchanger 15, with the ethylene product stream 35 located on the other side of heat exchanger 15. The ethanol vapor stream 10 is preheated in heat exchanger 15 to a temperature between about 250°C and 400°C, preferably between about 275°C and 350°C, through indirect heat exchange with the ethylene product stream 35. The reactor feed mixture 25 is formed by mixing the ethanol vapor stream 10 with the superheated vapor stream 30 at point AA at a weight ratio of 1:1 to 1:5, preferably 1:2.5 to 1:3.5. (A higher steam dilution ratio will reduce the conversion rate of ethanol in the dehydration reaction and should therefore be avoided.) The resulting reactor feed mixture 25 is at a temperature of about 425°C to 500°C and is fed or supplied to the ethanol dehydration reactor 32 at a pressure of about 0.15 MPa to about 0.4 MPa (absolute).

[0015] Ethanol dehydration reactor 32 is a single-stage adiabatic reactor with a packed bed of dehydration catalyst. The reactor feed mixture (as described above, containing at least ethanol and vapor) flows downwards under pressure through the packed bed of catalyst in reactor 32, where the ethanol in the reactor feed mixture is dehydrated to ethylene with very high selectivity and conversion rate – thus producing a high-purity ethylene source. The dehydration reaction is represented by the following equation: C2H5OH → C2H4 + H2O + X Here, "X" represents byproducts such as aldehydes, ethers, ethanol, and C3 / C4 compounds. Operating the process within the optimal temperature and pressure range will limit the formation of these byproducts while maximizing conversion and yield. Generally, the conversion of ethanol to ethylene increases with increasing temperature, but—at the same higher pressure—can sometimes decrease significantly. Therefore, the temperature and pressure ranges of the reactor feed mixture and the dehydration reactor described above are selected to maximize the selectivity and conversion of the dehydration reaction. It is worth noting that ethanol itself can decompose into acetaldehyde at temperatures above 480°C. This acetaldehyde can cause destructive catalyst coking. Therefore, if possible, the temperature of the reactor feed mixture should be maintained below 480°C.

[0016] As described above, the reactor feed mixture 25, containing steam and ethanol, flows downwards under pressure through the catalyst bed of the dehydration reactor 32, where ethanol is dehydrated to ethylene, producing an ethylene product stream 35. This ethylene product stream 35 is a gas and exits the reactor under pressure. As described above, the dehydration reaction operates with high selectivity and conversion, so almost all of the ethanol is converted to ethylene. Only trace amounts of unreacted ethanol and acetaldehyde, among other impurities, are found in the ethylene product stream 35. Therefore, in addition to ethylene, the ethylene product stream also contains water and trace amounts of unreacted ethanol and acetaldehyde, among other impurities. More specifically, the ethylene product stream 35 contains approximately 5 mol% to approximately 20 mol% ethylene, with the balance being water / steam and unreacted ethanol and impurities. Measured on a non-aqueous basis, the ethylene product stream 35 contains approximately 95 mol% to approximately 99.5 mol% ethylene.

[0017] Ethylene product stream 35 exits the reactor at a temperature of approximately 325°C to 425°C. On the other side of the exchanger, ethylene product stream 35 comes into thermal contact with ethanol vapor stream 10 in exchanger 15. Therefore, ethanol vapor stream 10 is preheated through indirect heat exchange with ethylene product stream 35, while ethylene product stream 35 is cooled by this heat exchange. After exiting exchanger 15, ethylene product stream 35 comes into thermal contact with process steam stream 2 at exchanger 40, which provides heat to preheat process steam stream 2 to a temperature of approximately 275°C to approximately 325°C. Due to the heat exchange in exchanger 40, the temperature of ethylene product stream 35 decreases to a range between approximately 175°C and 250°C.

[0018] The process steam stream 2 originates from steam generator 95, where it is generated by mixing with a medium-pressure stream from OSBL. The process steam stream exits the steam generator at a temperature of approximately 125°C to 145°C. Process steam stream 2 is heated in heat exchanger 40 as described above, and then further heated in furnace 50 to form superheated steam stream 30. The temperature of stream 30 is between approximately 500°C and 650°C, preferably between approximately 525°C and 600°C. Stream 30 is then combined with steam 10 to form the reactor feed mixture 25 described above.

[0019] Following exchanger 40, the ethylene-vapor product stream is split into two or more portions at point BB. Figure 1 illustrates an embodiment where the ethylene-vapor product stream is split into two portions at point BB: the first portion 60 is fed to exchanger 5, where heat is provided to evaporate the liquid ethanol stream 10. The second portion 65, in exchanger 70, is further cooled by providing output heat through a fluid (not shown) on the other side of the exchanger. This heated fluid can be used to preheat downstream processes and utility steams (not shown).

[0020] When the second portion 65 of the ethylene-steam product stream exits the exchanger 70 and combines with the first portion 60, the temperature of the recombined stream 73 has been significantly reduced: the temperature of the recombined stream is between approximately 90°C and approximately 125°C. This cooling is achieved by transferring heat to the exchanged streams in a continuous exchanger, as previously described. At this point, if the downstream unit no longer requires heat, the recombined ethylene-steam product stream is sent to the dehydration condenser 75. In the dehydration condenser, the recombined stream 73 is further cooled until condensation, thereby separating into dehydration condensate 85 and steam ethylene product 80.

[0021] The ethylene product 80 contains more than about 85 mol% ethylene, preferably more than about 90 mol%, with the balance being water, preferably at a saturated level. The stream is then preferably compressed to a pressure ranging from about 0.5 MPa to about 5 MPa (absolute), and optionally subjected to water washing and alkaline washing to remove residual aldehydes, carbon dioxide, sulfur, and other impurities, before being fed into downstream processes. For example, the ethylene product stream can be fed into an ethylene oxide reactor to produce ethylene oxide and ethylene oxide derivatives.

[0022] Liquid water condensate 85, containing some unreacted ethanol and other impurities (such as acetaldehyde), is fed to stripping tower 88, where ethanol, acetaldehyde, and other impurities are stripped at the top to form impurity stream 90, which is then fed to ethanol dehydration furnace 50 for incineration (the impurity vapor may be sent to a knock-out drum, not shown, before being sent to the furnace). This leaves stripping tower bottom stream 92, which contains water and trace amounts of unreacted ethanol. Stripping tower bottom stream 92 is fed to steam generator 95, where the required processing steam stream 2 for the ethanol dehydration reactor is generated using medium-pressure steam supplied from OSBL to steam generator 95. This processing steam stream 2 is then fed to exchanger 40 for superheating, as described above.

[0023] Figure 2 illustrates an improved, inventive ethanol dehydration process 150 constructed according to the present invention. While the conventional process 1 shown in Figure 1 combines excellent thermal integration and ethanol dehydration with high selectivity and conversion, process 1 requires a continuous steam input, which increases the process's utility costs. The inventive process 150 differs from the conventional process in many respects, but in particular, by utilizing the latent heat released by steam condensation, process 150 eliminates the need for a continuous input of medium-pressure steam during operation to generate the process steam used in the process. (It should be noted that while a continuous steam input is not required during operation, it may be necessary at other times, such as during startup, and in certain situations or during occasional operational deviations, the plant operator may occasionally choose a continuous steam input.) In addition to the brief description of process 150 of the present invention above, process 150 of the present invention will now be described in more detail with reference to conventional process 1. Many of the same flows, services, and equipment items present in process 1 also exist in process 150 with the same temperature range, pressure range, flow composition, and other ranges. Therefore, unless otherwise stated, the units and flows in process 150 have the same markings and values ​​as in 1.

[0024] In the process of this invention, the ethanol product stream 35 is maintained within the same temperature and pressure range as previously disclosed regarding the process of this invention. However, a difference can be observed in this invention when the ethanol product stream 35 enters the heat exchanger 140. In the backwards operation of this invention, the second processing steam stream 12, which has just been compressed in the dehydration steam compressor 101, enters the heat exchanger 40 at a much higher temperature than the corresponding stream 2 in the conventional process. Therefore, the ethanol product stream 35 is raised to even higher temperatures as it passes through the heat exchanger 140. As shown in FIG2, stream 35 passes through the heat exchanger 140, where it comes into thermal contact with the second processing steam stream 12 on the other side of the exchanger 140, and exits the exchanger as a high-temperature ethanol vapor stream 135 with a temperature between about 250°C and about 325°C. (The formation, temperature, and pressure of the second processing steam stream 12 will be described in more detail below.) A high-temperature ethanol vapor stream 135 flows from heat exchanger 140 across one side of heat exchanger 15, where it exchanges heat with liquid ethanol on the other side of the exchanger, causing the ethanol to evaporate and produce an ethanol vapor stream 10, as described with respect to the process shown in Figure 1. Here, the invention makes a significant operational change compared to conventional processes: after leaving exchanger 15, 100% of the stream flows to the second steam generator exchanger 195, instead of some being diverted for providing heat and utility supplies (see instead the diversion at point BB in Figure 1 and the accompanying description). This operational change in the invention can be seen in Figure 2, where, after heat exchange with the liquid ethanol stream, the temperature of stream 135 now decreases to approximately 125°C to 140°C; then, the entire high-temperature ethanol vapor stream 135 is sent to the second steam generator 195. It flows to the tube side of the steam generator heat exchanger 195 at a pressure of approximately 0.2 MPa to approximately 0.3 MPa (absolute), where it is sufficiently cooled within 195 to reach the dew point of the product ethylene stream, approximately 120°C. As a result, the steam in the product ethylene stream condenses, releasing the heat of condensation. This heat of condensation is collected to generate steam to form the second process steam stream 12. This saves a significant amount of energy because no more steam input is needed to provide heat for the process. This saving far outweighs the thermal integration loss caused by the product ethylene stream no longer being available to provide heat to downstream sources. This is because the heat generated when the latent heat of condensation in the steam generator 195 produces steam far exceeds the heat that could be transferred as sensible heat in further exchangers in process 1.

[0025] Then, due to the pressure gradient on the steam generator, the ethylene product stream flows out of the steam generator and is flash-separated at point DD into steam generator steam stream 105 and steam generator condensate 110. The steam generator condensate flows out of the steam generator 195, mixes with liquid water condensate at point CC, and then the combined stream is sent to the stripping tower 88, which operates as described above regarding process 1. The remaining steam generator steam stream 105 flows to the dehydration condenser 75, which operates as described above regarding process 1.

[0026] The second processing steam stream 12 of process 150 exits the steam generator at a temperature of approximately 95°C to approximately 120°C, comparable to the temperature of the processing steam stream 12 in process 1, but unlike process 1, it does not require the input of medium-pressure steam. However, without the addition of medium-pressure steam, the pressure on the shell-side steam generator is low, and the processing steam stream 2 is only pressurized to approximately 0.125 MPa to approximately 0.20 MPa (absolute) pressure upon exiting the steam generator. Therefore, in order to make the processing steam stream usable for the remainder of process 150, the processing steam stream is compressed by the dehydration steam compressor 101 to approximately 0.3 MPa to approximately 0.45 MPa (absolute) pressure. After passing through the compressor, the temperature correspondingly rises to between approximately 225°C and approximately 275°C.

[0027] As mentioned above, the key inputs to processes 1 and 150 are ethanol, which is supplied by an external source / OSBL. Particularly preferred ethanol is bioethanol, meaning it is produced from biomass material. Bioethanol itself is obtained through the fermentation of plant material: plant biomass and agricultural byproducts and waste; all of which are abundant and renewable. Fermenting biomass into ethanol produces a mixture containing approximately 95% water and 5% ethanol. The water can then be separated using a combination of azeotropic distillation or solvent extraction.

[0028] The dehydration reactor of this invention can use any suitable dehydration catalyst, including heterogeneous catalysts such as zeolites. - Alumina (Al2O3), etc. A particularly preferred option is the one sold by Scientific Design Company, Inc. under the trade name SynDol. - Alumina dehydration catalyst.

[0029] As described above, the gaseous ethylene in stream 80 can be further refined and then used downstream to produce ethylene oxide. Ethylene oxide is produced by continuously contacting an oxygen-containing gas with the ethylene produced as described above in a fixed-bed tubular reactor in the presence of a silver-based ethylene oxide (“epoxidation”) catalyst. (The silver-based epoxidation catalyst is described in more detail below.) Oxygen can be supplied to the reaction in substantially pure molecular form or in the form of a mixture such as air. For example, under operating conditions, a typical reactant feed mixture may contain about 0.5% to about 45%, preferably about 5% to about 30%, of ethylene and about 3% to about 15% of oxygen, and about 0.3% to about 10% of carbon dioxide, with the balance consisting of relatively inert materials, including substances such as water, inert gases, other hydrocarbons, and reaction modifiers described herein. 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 byproducts of the epoxidation process and are common contaminants in the feed gas. Both have adverse effects on the catalyst, so the concentrations of these components are usually kept to a minimum.

[0030] As previously mentioned, one or more reaction regulators are also present in the reaction, non-limiting examples of which include organic halogenated compounds, such as C1 to C8 halogenated hydrocarbons; particularly preferred are chlorine-containing regulators, such as methyl chloride, ethyl chloride, dichloroethylene, vinyl chloride, or mixtures thereof. For rhenium-containing catalysts, controlling the chloride concentration level is especially important.

[0031] As described above, a common approach to ethylene epoxidation involves contacting ethylene with oxygen in a fixed-bed tubular reactor in the presence of an epoxidation catalyst. Conventional commercial fixed-bed ethylene oxide reactors typically consist of multiple parallel, elongated tubes (in a suitable shell) with an outer diameter (OD) of approximately 0.7 to 2.7 inches, an inner diameter (ID) of approximately 0.5 to 2.5 inches, and a length of 15–53 feet, each filled with catalyst. The reaction feed mixture (as described above) is introduced into these tubes, and the resulting reactor effluent gas contains ethylene oxide, unused reactants, and byproducts. Typically, in ethylene oxide processes, the reactor operating rate is around 130 kg / m³. 3 / h to 300kg / m 3 The working rate is between / h, while ΔEO is between 1.0% and 2.5%. The working rate is the production rate, expressed in kg / m² in this paper. 3 / h is the unit. ΔEO is defined as the number of moles of EO formed in the reactor per 100 moles of reactor feed. It essentially represents the concentration of ethylene oxide in the reactor effluent, as the concentration of ethylene oxide in the reactor feed must be maintained at a level very close to zero, typically only a few ppm. After startup and during normal operation, the feed composition at the reactor inlet typically includes (by volume%) 1-40% ethylene; 3-12% O2; 0.3% to 20%, preferably 0.3% to 5%, more preferably 0.3% to 1% CO2; 0-3% ethane; a certain amount of one or more chloride modifiers (described herein); the balance of the feed consists of argon, methane, nitrogen, or mixtures thereof.

[0032] Typical operating temperatures for ethylene epoxidation processes range from about 180°C to about 330°C, preferably from about 200°C to about 325°C, and more preferably from about 225°C to about 280°C. Operating pressures can vary from about atmospheres to about 30 atmospheres, depending on the desired mass velocity and productivity. Higher pressures can be used within the scope of this invention. Residence times in commercial-scale reactors are typically from about 2 seconds to about 20 seconds.

[0033] Silver-based epoxidation catalyst As described above, this ethylene oxide process utilizes a silver-based epoxidation catalyst. The silver-based epoxidation catalyst comprises a support, and at least an effective amount of silver or a silver-containing compound; optionally, a promoting amount of rhenium or a rhenium-containing compound; and optionally, a promoting amount of one or more alkali metals or alkali metal-containing compounds. The support used in this invention can be selected from a large number of solid refractory supports, which can be porous and can provide a preferred pore structure. Alumina is well known to be used as a catalyst support for olefin epoxidation and is a preferred support.

[0034] Regardless of the characteristics of the carrier used, it is typically manufactured in the form of granules, blocks, flakes, pellets, rings, spheres, wheels, cross-segmented hollow cylinders, etc., with dimensions suitable for use in fixed-bed epoxidation reactors. The equivalent diameter of the carrier particles is preferably in the range of about 3 mm to about 12 mm, more preferably in the range of about 5 mm to about 10 mm. (Equivalent diameter is the diameter of a sphere having the same outer surface (i.e., ignoring the surface within the pores of the particles) to volume ratio as the carrier particles used.) Suitable carriers are available from Saint-Gobain Norpro Co., Sud Chemie AG, Noritake Co., CeramTec AG, and Industrie Bitossi SpA. Further information on carrier compositions and methods of carrier preparation, without limitation on the specific compositions and formulations contained herein, can be found in U.S. Patent Publication No. 2007 / 0037991.

[0035] To produce a catalyst for oxidizing olefins to olefin oxides, a catalytically effective amount of silver is then provided on the surface of a support having the above-described characteristics. In one embodiment, the catalytically effective amount of silver is from 10 wt% (by weight) to 45 wt%. The catalyst is prepared by impregnating the support with a silver compound, complex, or salt dissolved in a suitable solvent sufficient to allow the silver precursor compound to deposit on the support. An aqueous solution of silver is preferred.

[0036] A promoting amount of rhenium component (which may be a rhenium-containing compound or a rhenium-containing complex) may also be deposited on the support before, during, or after silver deposition. The amount of rhenium promoter present may be from about 0.001 wt% to about 1 wt%, preferably from about 0.005 wt% to about 0.5 wt%, more preferably from about 0.01 wt% to about 0.1 wt%, based on the weight of the total catalyst including the support (expressed in rhenium metal).

[0037] Other components that may be deposited on the support before, simultaneously with, or after the deposition of silver and rhenium include promoting amounts of alkali metals or mixtures of two or more alkali metals, and optionally promoting amounts of Group IIA alkaline earth metal components or mixtures of two or more Group IIA alkaline earth metal components, and / or transition metal components or mixtures 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 stepwise with various catalyst promoters. Specific combinations of the support, silver, alkali metal promoters, rhenium components, and optional additional promoters of the present invention will provide one or more improvements in catalytic performance superior to the same combination of silver and support and with or without only one promoter.

[0038] As used herein, the term "promoting amount" refers to the amount of a component that effectively improves the catalytic performance of a catalyst compared to a catalyst without that component. Of course, the exact concentration used will depend on a variety of factors, including the desired silver content, the nature of the support, the viscosity of the liquid, and the solubility of the specific compound used to deliver the promoter into the impregnation solution. Examples of catalytic performance include operability (resistance to runaway), selectivity, activity, conversion, stability, and yield. Those skilled in the art will understand that a "promoting amount" can enhance one or more individual catalytic properties, while other catalytic properties may or may not be enhanced, or may even be diminished.

[0039] Suitable alkali metal promoters can 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 should be a promoting amount. Preferably, this amount ranges from about 10 ppm to about 3000 ppm, more preferably from about 15 ppm to about 2000 ppm, even more preferably from about 20 ppm to about 1500 ppm, and particularly preferably from about 50 ppm to about 1000 ppm, based on the weight of the total catalyst and in terms of metal content.

[0040] Suitable alkaline earth metal promoters include elements from Group IIA of the periodic table, such as 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, or combinations thereof.

[0041] The amount of alkaline earth metal promoter and / or transition metal promoter deposited on the carrier is the promotion amount. The content of the transition metal promoter is typically from about 0.1 μmol / g to about 10 μmol / g, preferably from about 0.2 μmol / g to about 5 μmol / g.

[0042] The silver solution used for impregnating the carrier may also include optional solvents or complexing agents / solvents known in the art. A variety of solvents or complexing agents / solvents can be used to dissolve silver to the desired concentration in the impregnation medium. Useful complexing agents / solvents include amines, ammonia, oxalic acid, lactic acid, and combinations thereof. Amines include alkylene diamines having 1 to 5 carbon atoms. In a preferred embodiment, the solution comprises an aqueous solution of silver oxalate and ethylenediamine. The complexing agent / solvent may be present in the impregnation solution in an amount from about 0.1 mol to about 5.0 mol per mole of silver, preferably from about 0.2 mol to about 4.0 mol, more preferably from about 0.3 mol to about 3.0 mol.

[0043] When using a solvent, it can be an organic solvent or water, and can be polar or substantially or completely nonpolar. Generally, the solvent should have sufficient dissolving power to dissolve the solution components. At the same time, it is preferable to choose a solvent to avoid undue influence or interaction with the solvation accelerator. Organic-based solvents having 1 to about 8 carbon atoms per molecule are preferred. Mixtures of several organic solvents or mixtures of organic solvents with water can be used, provided that such mixed solvents function as required herein.

[0044] The silver concentration in the impregnation solution is typically in the range of about 0.1 wt% to the maximum solubility provided by the particular solvent / solvent combination used. Solutions containing about 0.5 wt% to about 45 wt% silver are generally well-suited, with a preferred silver concentration of 5 wt% to 35 wt%.

[0045] The impregnation of the selected carrier can be achieved using any conventional method; for example, overfill impregnation, initial wetting impregnation, spraying, etc. Typically, the carrier material is contacted with a silver-containing solution until the carrier absorbs a sufficient amount of solution. Preferably, the amount of silver-containing solution used to impregnate the porous carrier does not exceed the amount required to fill the carrier pores. A single impregnation or a series of impregnations can be used, with or without drying in between, depending in part on the concentration of the silver component in the solution. Impregnation steps are described, for example, in U.S. patents of patent numbers 4,761,394, 4,766,105, 4,908,343, 5,057,481, 5,187,140, ​​5,102,848, 5,011,807, 5,099,041, and 5,407,888. Known prior steps of pre-deposition, co-deposition, and post-deposition with various accelerators can be employed.

[0046] After impregnating the support with a silver-containing compound (i.e., a silver precursor, rhenium component, alkali metal component, and optionally other promoters), the impregnated support is calcined for a period of time sufficient to convert the silver-containing compound into active silver and remove volatile components from the impregnated support to produce a catalyst precursor. Calcination can be achieved by heating the impregnated support (preferably at a gradual rate) to approximately 200 °C. o C to approximately 600 o This is accomplished using temperatures ranging from approximately 0.5 bar to approximately 35 bar. Generally, the higher the temperature, the shorter the required heating time. A wide range of heating time ranges has been proposed in the art; for example, U.S. Patent No. 3,563,914 discloses heating times of less than 300 seconds, while U.S. Patent No. 3,702,259 discloses heating times of 100... o C to 375 oHeating at a temperature of C for 2 to 8 hours, typically lasting from about 0.5 to 8 hours. However, it is important that the heating time is temperature-dependent, ensuring that virtually all the silver contained is converted into active silver. Continuous or stepwise heating can be used for this purpose.

[0047] During calcination, the impregnated support may be exposed to a gaseous atmosphere containing an inert gas or a mixture of an inert gas and approximately 10 ppm to 21% by volume of an oxygen-containing oxidizing component. For the purposes of this invention, an inert gas is defined as a gas that does not substantially react with the catalyst or catalyst precursor under the selected calcination conditions. More information on catalyst manufacturing can be found in the aforementioned U.S. Patent Publication No. 2007 / 0037991.

[0048] Example The invention will now be described in more detail based on the following non-limiting simulated embodiments.

[0049] The conventional ethanol dehydration process 1 and the ethanol dehydration process prepared according to the present invention 150 are shown in Figures 1 and 2, respectively, and were simulated using PRO / II software. The temperature and pressure of each flow used in the simulation are shown in Figure 3 (simulation of process 1 shown in Figure 1) and Figure 4 (simulation of process 150 shown in Figure 2).

[0050] In addition to the temperature and pressure values ​​shown in Figures 3–4, the following reactor parameters are also part of the simulation: Table I Special attention is paid to the comparison of the temperature and pressure values ​​of the processed steam stream 2 (in process 1) and the second processed steam stream 12 (in process 150). As mentioned above, in conventional ethanol dehydration process 1, the ethylene-steam product stream is not directly sent to the steam generator, but is entirely sent to the dehydration condenser and then immediately leaves the process for further purification and / or for downstream processes. Therefore, in conventional processes, steam must be input to provide the heat and steam required by the process. Due to the steam input, the temperature and pressure of processed steam stream 2 in conventional process 1 are higher than those of the corresponding stream 12 in the process 150 of the present invention, which does not have a steam input, as shown in Table II: Table II The advantages of this invention are particularly evident in Table III, which compares the energy balance of a simulated conventional ethanol dehydration process 1 with the simulated energy balance of this invention 150. The energy requirement of the dehydration process of this invention is 14.45 MKal / hr, significantly lower than the simulated conventional process's 41.84 MKal / hr. This saving is primarily due to the elimination of the need for steam input. The specific differences in energy requirements are shown in Table III below.

[0051] Table III It is worth noting that, in addition to the significant savings shown in the table above, additional energy is required to power the steam compressor. However, this energy requirement of 4.5 Mkal / hr is small compared to the energy cost of the input steam. Furthermore, it is noteworthy that in the process of this invention, there is no heat available for output, as shown in Table IV below.

[0052] Table IV However, this output heat has almost no utilization value for downstream processes, so its loss will not reduce the significantly improved process economy of the present invention.

Claims

1. A method for dehydrating ethanol to ethylene, comprising the following steps: (a) A reactor feed mixture is formed by mixing an ethanol vapor stream with a superheated vapor stream; (b) Supplying the reactor feed mixture to the ethanol dehydration reactor; (c) Dehydrating the ethanol in the reactor feed mixture in an ethanol dehydration reactor to form an ethylene product stream; (d) The ethylene product stream is supplied to a steam generator at a pressure of about 0.2 MPa to about 0.3 MPa and a temperature of about 125°C to about 140°C; (e) Cool the ethylene product stream below the dew point of the ethylene product stream to release the heat of condensation; (f) A second processing steam stream is generated by the condensation heat, the temperature of the processing steam stream being about 95°C to about 120°C; (g) Allowing the second processed steam stream to exit from the steam generator at a pressure of about 0.125 MPa to about 0.2 MPa; and (h) Compress the second processing steam stream to a pressure of about 0.3 MPa to about 0.45 MPa.

2. The method according to claim 1, further comprising the following steps: The ethylene product stream is directed out of the steam generator; as well as The ethylene product stream is separated into steam generator condensate and steam generator vapor stream by flash evaporation.

3. The method according to claim 1, wherein the ethanol dehydration reactor is a single-stage adiabatic reactor having a packing bed of dehydration catalyst.

4. The method according to claim 1, wherein the ethanol dehydration reactor has - A single-stage adiabatic reactor with a packed bed of alumina dehydration catalyst.

5. The method of claim 1, wherein the temperature of the reactor feed mixture is from about 425°C to about 500°C, and the ethylene product stream exits the reactor at a temperature of about 325°C to 425°C.

6. The method according to claim 1, wherein in step (a), the reactor feed mixture is formed by mixing the ethanol vapor stream with the superheated vapor stream at a weight ratio of 1:1 to 1:

5.

7. The method of claim 1, wherein the ethylene product stream contains about 95 mol% to about 99.5 mol% ethylene on a non-aqueous basis.

8. The method of claim 1, wherein the method further comprises preheating the ethanol vapor stream in a heat exchanger by heat exchange with the ethylene product stream.

9. The method of claim 1, wherein after the compression step, the temperature of the processed steam stream is about 225°C to about 275°C.

10. The method of claim 1, wherein the temperature of the ethylene product stream is about 325°C to 425°C.

11. The method of claim 1, wherein the ethylene product stream contains, on a non-aqueous basis, from about 95 mol% to about 99.5 mol% ethylene.

12. The method of claim 1, wherein the temperature of the reactor feed mixture is below about 480°.

13. The method of claim 1, wherein the reactor feed mixture is supplied to the ethanol dehydration reactor at a pressure of about 0.15 MPa to about 0.4 MPa (absolute).

14. The method according to claim 1, wherein the ethylene is subjected to alkaline washing to remove impurities.

15. The method of claim 1, wherein bioethanol is the source of the ethanol vapor stream.

16. The method of claim 1, wherein the ethylene is contacted with oxygen in a fixed-bed tubular reactor in the presence of a silver-based epoxidation catalyst to produce ethylene oxide.

17. The method of claim 1, wherein the ethylene is contacted with oxygen in a fixed-bed tubular reactor in the presence of a silver-based epoxidation catalyst to produce ethylene oxide, wherein the reactor operates at approximately 130 kg / m³. 3 / h to approximately 300 kg / m 3 Running at a rate of / h.

18. The method of claim 1, wherein the silver-based epoxidation catalyst contains about 15 wt% to about 40 wt% silver.

19. The method of claim 1, wherein the ethylene is contacted with oxygen in a fixed-bed tubular reactor in the presence of a silver-based epoxidation catalyst to produce ethylene oxide, wherein the reactor is operated at a temperature of about 240°C to about 280°C.

20. A method for dehydrating ethanol to ethylene, comprising the following steps: (a) A reactor feed mixture is formed by mixing an ethanol vapor stream with a superheated vapor stream at a weight ratio of 1:1 to 1:5, the temperature of the reactor feed mixture being about 425°C to about 500°C; (b) The reactor feed mixture is supplied to the ethanol dehydration reactor, wherein the ethanol dehydration reactor is a reactor having - A single-stage adiabatic reactor with a packed bed of alumina dehydration catalyst; (c) Dehydrating the ethanol in the reactor feed mixture in an ethanol dehydration reactor to form an ethylene product stream, the ethylene product stream leaving the reactor at a temperature of about 325°C to about 425°C; (d) The ethylene product stream is supplied to a steam generator at a pressure of about 0.2 MPa to about 0.3 MPa and a temperature of about 125°C to about 140°C; (e) Cool the ethylene product stream below the dew point of the ethylene product stream to release the heat of condensation; (f) A second processing steam stream is generated by the condensation heat, the temperature of the processing steam stream being about 95°C to about 120°C; (g) Allowing the second processed steam stream to exit from the steam generator at a pressure of about 0.125 MPa to about 0.2 MPa; and (h) Compress the second processing steam stream to a pressure of about 0.3 MPa to about 0.45 MPa.

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