FROM DIOLS

FIELD OF THE INVENTION

[0001] The present invention relates to the field of separating glycols, and more particularly to separating a target glycol from a mixture of one or more diols.

BACKGROUND OF THE INVENTION

[0002] Certain glycols, such as monoethylene glycol (MEG) and monopropylene glycol (MPG), are useful as heat transfer media, antifreeze and precursors to polymers, such as polyester and polyethylene terephthalate.

[0003] In a typical industrial process, MEG is prepared in a two-step process. In the first step, ethylene is converted to ethylene oxide by reaction with oxygen over a silver oxide catalyst. The ethylene oxide can then be converted into MEG. This may be carried out directly by catalytic or non-catalytic hydrolysis. Alternatively, in one well-known process, ethylene oxide is catalytically reacted with carbon dioxide to produce ethylene carbonate. The ethylene carbonate is subsequently hydrolyzed to provide ethylene glycol.

[0004] These routes rely for their starting material on ethylene, which is produced in the petrochemical industry by steam cracking of hydrocarbons derived from fossil fuels. By-products from this reaction include diethylene glycol (DEG), triethylene glycol (TEG), and tetraethylene glycol (TTEG).

[0005] In recent years, increased efforts have been focused on reducing the reliance on fossil fuels as a primary resource for the provision of fuels and commodity chemicals. Carbohydrates and related biomass are seen as key renewable resources in the efforts to provide new fuels and alternative routes to desirable chemicals.

[0006] In particular, certain carbohydrates can be reacted with hydrogen in the presence of a catalyst system to generate polyols and sugar alcohols. Current methods for the conversion of saccharides to glycols revolve around a hydrogenation/hydrogenolysis process. A process for the conversion of cellulose to products including MEG is described in Angew. Chem. Int. Ed. 2008, 47, 8510-8513. Continuous processes for generating at least one polyol from a sacchari decontaining feedstock are described in WO 2013/015955 and CN 103731258A. [0007] The products of these reactions are a mixture of materials comprising MEG, MPG, 1,2- butanediol (1,2-BDO), 1,2-hexanediol (1,2-HDO), and other by-products. Although the conversion of glucose to glycols can be carried out with high selectivity to MEG, a much lower selectivity and increased levels of MPG are obtained when using fructose (for example, in combination with glucose), and/or sucrose as a feedstock.

[0008] Whether the MEG and/or MPG is produced from a petroleum-derived feedstock or a renewal feedstock, the separation of the target glycols from reaction by-products by fractional distillation is complicated due to the similarity in boiling points. For example, MEG and 1,2-BDO have boiling points of 197.2 and 196.5°C at 101.300 kPa, respectively. The isolation of a pure MEG or MPG by fractional distillation is further complicated by the formation of a homogeneous minimum boiling azeotropes, for example between MEG and 1,2-BDO, at atmospheric pressure.

[0009] Perez Golf et al. (US2019/0062244A1) relates to a process for producing high purity MEG or MPG by subjecting a product stream to a first distillation column to separate a top stream comprising a mixture of two or more C2-C7diols. The top stream is then passed to a second distillation column, where an extractant, such as a C3-C6 sugar alcohol, is provided to separate the MEG or MPG and extractant in a bottoms stream. The bottoms stream is then passed to a third distillation column to separate high-purity MEG or MPG in a top stream.

[0010] A further disadvantage of conventional separation processes is the decomposition of the target glycol product and/or extractant, for example, by condensation and/or elimination reactions. One current method to mitigate this problem is to limit the bottom temperature of the distillation column to temperatures lower than 180°C and to limit the time to which the target glycol and other organic components are subjected to high temperatures. Zaboon et al. (“Recovery of mono-ethylene glycol by distillation and the impact of dissolved salts evaluated through simulation of field data” J Natural Gas Sci Eng 44:214-232; 2017) is an example of conventional approaches of limiting distillation operating temperatures to avoid thermal degradation of MEG. Zaboon et al. confirm that some conventional MEG recovery units operate at temperatures as low as 95°C, 140°C and up to temperatures of 160°C. Accordingly, Zaboon et al. limited operating temperatures to a range of 120°C to 160°C, while operation of the reboiler unit at temperatures above 160°C was not performed due to the possibility of thermal degradation effects of MEG.

[0011] In an effort to overcome the problem of thermal degradation effects, Huizenga et al. (US2020/0377438A1) provide a method for separating oxygenates in a distillation column operating at liquid phase temperatures higher than the thermal stability of the oxygenates, by providing a water feed stream to a bottom of the distillation column to generate a water partial pressure in a vapor flow in the distillation column to provide a thermodynamic driving force to avoid and/or limit thermal condensation and elimination reactions in the distillation column. In one embodiment, a heat exchanger may be used to cool the bottoms recovery stream.

[0012] In addition to using feedstocks that are not petroleum-derived, there is also a drive towards reducing the energy required in producing desired chemicals.

[0013] To that end, Bajaj et al. (US2020/0299215A1) describes a process for recovering ethylene glycol from an aqueous stream by subjecting the stream to a multiple-effect evaporator to obtain a concentrated ethylene glycol stream. The concentrated stream is then passed to a dehydration step to obtain a partially dehydrated ethylene glycol stream, which is then passed to a second dehydrator operating under vacuum to obtain a dehydrated ethylene glycol stream. In one embodiment, the aqueous stream is produced in an ethylene oxide section having an ethylene oxide stripper. The ethylene oxide stripper has a first and second reboiler wherein heat is supplied to the second reboiler by a low-pressure steam generated in a MEG purification column overhead condenser.

[0014] Multi-stage evaporators are also described in Cocuzza et al. (US3,875,019) process for separating ethylene glycol from dilute aqueous glycol solutions. Residual water is removed from the solution before the solution is passed to a rectification column. The overhead stream from the last stage of the multi-stage evaporators is compressed and used as a heating medium for the reboiler of the rectification column.

[0015] Unfortunately, several of these solutions disclosing heat integration also involve increased capital expenses.

[0016] There remains a need to reduce energy requirements, carbon dioxide footprint, operating expenses and/or capital expenses of product recovery schemes, while providing a good quality product.

SUMMARY OF THE INVENTION

[0017] According to one aspect of the present invention, there is provided a process for separating a target glycol from a mixture of the target glycol and one or more C2-C8 diols by distillation, the process comprising the steps of: providing a feed mixture comprising the target glycol and a C2- C8 diol; providing a water feed stream; directing the water feed stream to a bottom of a first distillation column operating at a temperature greater than or equal to 180°C to separate a first overhead stream comprising a first portion of the target glycol and a first bottoms stream comprising the C2-C8 diol; providing a second distillation column to separate a second overhead stream comprising a second portion of the target glycol and a second bottoms stream comprising the C2-C8 diol; directing the feed mixture to one of the first distillation column and the second distillation column; and passing the first overhead stream to a reboiler of the second distillation column to transfer heat from the first overhead stream to a portion of the second bottoms stream recycled to the second distillation column.

BRIEF DESCRIPTION OF THE DRAWINGS

[0018] The process of the present invention will be better understood by referring to the following detailed description of preferred embodiments and the drawings referenced therein, in which:

[0019] Fig. 1 is a flow diagram illustrating one embodiment of the present invention having two distillation columns, with heat integration between the first and second column;

[0020] Fig. 2 is a flow diagram illustrating the embodiment of Fig. 1, with three distillation columns;

[0021] Fig. 3 is a flow diagram illustrating the embodiment of Fig. 2, with heat integration between the first and second columns, and between the second and third columns; and

[0022] Fig. 4 is a flow diagram illustrating another embodiment of the present invention, directing the feed mixture to the second column.

DETAILED DESCRIPTION OF THE INVENTION

[0023] The present invention provides a process for separating a target glycol from a mixture of the target glycol and one or more C2-C8 diols. Preferably, the target glycol is selected from the group consisting of monoethylene glycol (MEG) and monopropylene glycol (MPG). The C2-C8 diols in mixture with the target glycol will depend on whether the target glycol is produced from a petroleum-derived feedstock or a renewable feedstock. For example, if the feedstock is petroleum-derived, the C2-C8 diols in the mixture include DEG, TEG, TTEG, and combinations thereof. Where a renewable feedstock is used, the C2-C8 diols in the mixture include butanediols, pentanediols, and hexanediols. The process of the present invention is preferably directed to a mixture of a target glycol and a C2-C8 diol selected from the group consisting of butanediols, pentanediols, hexanediols, DEG, TEG, TTEG, and combinations thereof.

[0024] The target glycol is separated in at least two distillation columns. In the first distillation column, a water feed stream is supplied to the bottom of the first column to enable the column to be operated at a higher operating temperature and pressure exceeding the thermal stability of the target glycol to separate a first overhead stream and a first bottoms stream. Preferably, the first distillation column is operating at a temperature greater than or equal to 180°C. By providing the water feed stream to the bottom of the first column, the first overhead stream is a vapor phase stream obtained at a higher temperature than conventional separation processes, allowing for heat integration between the first and second distillation columns. By “bottom”, we mean at a liquid hold-up region in the distillation column, at a heating device for the distillation column, and combinations thereof.

[0025] Until now, it was not possible to operate at temperatures greater than or equal to 180°C without thermally degrading the target glycol. Accordingly, until now, the distillation temperature was limited to temperatures below, for example, 175 °C, to avoid significant thermal degradation of MEG, which renders MEG off-spec on the UV specification. The inventors have discovered that by using the thermal stabilizing effect of water to enable distillation at higher temperatures, the separated streams from distillation are at a higher temperature and offer heat integration of the top of one column with the bottom of another. This was not heretofore available for distillation temperatures limited to operating at a temperature and pressure below the thermal degradation of MEG.

[0026] In this way, the heat in the vapor phase stream can result in an energy savings of about 50%, in the case of a 2 column distillation, or even 67% in the case of a three column distillation. Furthermore, additional savings may be possible, by avoiding the need for a conventional dehydrator or by relaxing the water specifications of the dehydrator.

[0027] Referring now to Fig. 1, in the process of the present invention 10, a feed mixture 12 and a water feed stream 14 are directed to a first distillation column 20.

[0028] The process of the present invention 10 may be applied to any feed mixture 12 comprising a target glycol and a C2-C8 diol. Preferably, the feed mixture 12 is derived from the reaction product stream from a process for producing a target glycol. Preferably, the feed mixture 12 is a product stream from a petroleum-derived feedstock or a renewable feedstock. In a particularly preferred embodiment, the feed mixture 12 is derived from the reaction product stream from a process for the hydrogenolysis of a saccharide-containing feedstock.

[0029] Typically, the reaction product stream from a process for the hydrogenolysis of a saccharide-containing feedstock comprises, as glycols, at least MEG, MPG, 1 ,2-BDO and 1 ,2- HDO, while the reaction product from ethylene oxide hydrolysis comprises, as diols, at least MEG, DEG, TEG, and TTEG These glycols and diols are typically present at a concentration in the range of from 0.1 to 30 wt.% of the overall stream. Although formally DEG, TEG and TTEG are diols but not glycols, as they have no adjacent alcohol groups, they might be referred to as ‘glycols’ as well as ‘diols’ for ease of discussion in the remainder of the document.

[0030] In such a reaction product stream, MEG is suitably present as at least 10 wt.%, preferably as at least 30 wt.% of the organic fraction of the stream. MEG is suitably present as at most 99 wt.%, preferably as at most 95 wt.%, more preferably as at most 90 wt.%, most preferably as at most 80 wt.% of the organic fraction of the stream.

[0031] In the case of a renewable reaction product stream where the feedstock is glucose, MPG is suitably present in a range of from 2 wt.% to 10 wt.%, of the organic fraction of the stream. In the case of a glucose/fructose feedstock or a sucrose feedstock, the MPG is present in a range of from 20 wt.% to 25 wt.% of the organic fraction of the stream.

[0032] In a renewable reaction product stream, 1,2-BDO is present in a range of from 1 wt.% to 10 wt.% of the organic fraction of the stream.

[0033] In a petroleum-derived reaction product stream, the yield of DEG, TEG and TTEG depends on the initial concentration of ethylene oxide. For example, at an initial concentration of 9 wt.% ethylene oxide, with a full conversion, the product yield may be 90 wt.% MEG, 9 wt.% DEG, 0.9 wt.% TEG and 0.1 wt.% TTEG, based on the organic fraction of the product stream. Typically, DEG is present in a range of from 5 wt.% to 20 wt.% of the organic fraction of the stream.

[0034] As well as the glycols, the reaction product stream from hydrogenolysis reactions of saccharides may comprise solvent, oxygenates, hydrocarbons, catalyst, degradation products, and gases. The solvent in the reaction product stream is often water. The variety of compounds and their concentration depend on the saccharide-containing feedstock and the various hydrogenation and hydrogenolysis conversion conditions, including catalysts, reaction conditions such as temperature, pressure, and saccharide concentration. However, suitably the hydrogenolysis reactions have gone to completion and the aqueous stream contains less than 5 wt.%, preferably less than 2 wt.%, more preferably less than 1 wt.%, even more preferably less than 0.5 wt.%, most preferably substantially no saccharides when considered as a weight percentage of the overall stream. Typically, the aqueous stream also contains less than 5 wt.%, preferably less than 2 wt.%, more preferably less than 1 wt.%, even more preferably less than 0.5 wt.%, most preferably substantially no glycerol, when considered as a weight percentage of the overall stream.

[0035] Where the feed mixture 12 is derived from a reaction product stream from a petroleum- derived feedstock or a renewable feedstock, one or more treatment, separation and/or purification steps may be applied to the reaction product stream before the process of the present invention. Such steps may include one or more of: removal of at least a portion of the solvent present, for example by distillation; removal of light ends; fractional distillation to produce a glycols stream and removal of heavy organics and any inorganics present, such as catalyst material; and initial separation steps to achieve preliminary separation of glycols, e.g. separation of MPG by fractional distillation or other distillation process that results in a stream in which essentially all of the glycols remaining are MEG and 1,2-BDO, with trace amounts of higher glycols, such as 2,3-BDO, 1,2- pentanediol and 1,2-HDO. In another embodiment, the remaining glycols comprise 1,2-HDO. In the case where MEG is produced from ethylene oxide hydrolysis, the remaining glycols include DEG, TEG and TTEG

[0036] In a conventional process, the feed mixture 12 is the resulting stream from a concentrator and a dehydrator. For example, the concentrator may be a series of three concentrators, the first operating at high temperature and high pressure, while the second and third concentrators operate at progressively lower temperatures and pressures. The product of the concentrator stream typically has a water content of about 10 wt.%. Conventionally, the product is then routed to a high temperature dehydrator to further remove water before distillation. Advantageously, the process of the present invention can either eliminate the conventional dehydrator or allow the dehydrator to operate at a less costly conditions resulting in a higher water content in the feed mixture 12.

[0037] In the embodiment of Fig. 1, the feed mixture 12 is provided as a feed to a first distillation column 20. The first distillation column 20 may be any suitable distillation column known to those skilled in the art and may be equipped with trays, structured packing, and/or unstructured packing. The column preferably has a number of theoretical trays in a range of from 3 to 140. The actual number of theoretical trays is readily determined by a skilled person.

[0038] The process of the present invention 10 has at least two distillation columns 20, 40. The distillation columns may be, for example, without limitation, selected from simple distillation, fractional distillation, extractive distillation, and azeotropic distillation. The two distillation columns 20 and 40 may be the same or different.

[0039] In accordance with the present invention 10, the water feed stream 14 may be pure water or may comprise one or more contaminants. The water feed stream 14 may be provided as liquid water, water vapor or steam.

[0040] The ratio of the water feed stream 14 to the feed mixture 12 provided to the first distillation column 20 is in a weight ratio range of from 1 : 10,000 to 1: 10, preferably from 1 : 1000 to 1 :20, more preferably from 1:500 to 1:50, most preferably from 1: 100 to 1:10.

[0041] Adding the water feed stream 14 provides a partial pressure in the first distillation column 20 that reduces unwanted thermal condensation reactions and/or elimination reactions, such as condensation of MEG to DEG or dehydration of MEG to form acetaldehyde, that decompose the target glycol. More specifically, by directing the water feed stream 14 into the bottom of the first distillation column 20, thermal degradation of molecules comprising OH groups is reduced. This, in turn, reduces the formation of ethylene oxide, polyethylene glycols, 1,4- di oxane, acetaldehyde, hydroxyacetone, 3-hydroxypropanal, acrolein, polyglycerols, formaldehyde, and various glycerol condensation dehydration products that occur during conventional distillation of mixtures containing MEG, glycols, alcohols, polyhydric alcohols, glycerol and sugar alcohols.

[0042] Although the feed mixture 12 may include water, it is the addition of the water feed stream 14 in the bottom portion of the distillation column, below the inlet of the feed mixture 12, that provides the advantages of the present invention 10.

[0043] Without being bound by theory, the water feed stream 14 may act as a stripping agent to strip impurities from the feed mixture 12 and/or it may affect the chemical equilibrium in the distillation column 20 to provide a higher degree of separation of the target glycol.

[0044] In one embodiment, the water feed stream 14 may be directed to the bottom of the first distillation column 20 by providing the water feed stream 14a to a reboiler 22. In a preferred embodiment, the water feed stream 14b may be directly provided to a liquid hold-up region at the bottom of the distillation column 20.

[0045] The first distillation column 20 is operated at a liquid bottom temperature in a range of from 180 to 220°C, and a column bottom pressure in a range of from 60 to 200 kPa. Temperature and pressure cannot be varied independently. As an example, the relationship between temperature and pressure for high-purity MEG is provided in Table 1.

TABLE 1

MEG vapor pressures

[0046] In the first distillation column 20, a first overhead stream 24 is separated from a first bottoms stream 26. The first overhead stream 24 comprises a first portion of the target glycol from the feed mixture 12. The first overhead stream 24 also contains water that was vaporized from the water feed stream 14. The flow rate of the target glycol in the first overhead stream 24 is in a range of from 10 to 90 wt.% based on the flow rate of the target glycol present in the feed mixture 12.

[0047] The first bottoms stream 26 contains the remainder of the target glycol and the C2-C8 diol from the feed mixture 12. The flow rate of the target glycol in the first bottoms stream 26 is in a range of from 10 to 90 wt.% based on the flow rate of the target glycol present in the feed mixture 12.

[0048] The first bottoms stream 26 is directed to the second distillation columne 40 for further recovery of the target glycol. Preferably, a portion of the first bottoms stream 26 is passed through reboiler 22. The reboiler 22 is preferably a forced circulation reboiler using a pump (not shown) to circulate a portion of the first bottoms stream 26. A portion of the first bottoms stream 26 is vaporized in the reboiler 22 using a heating medium. The vaporized first bottoms stream 28 is returned to the first distillation column 20 to drive fractionation. Alternatively, or in addition, the water feed stream 14 may be provided as stream 14a, which is vaporized in the reboiler 22 to drive fractionation. The heat medium may be selected from steam, hot oil and/or a reactor effluent. Reboiler 22 heating might also be performed by electrical heating.

[0049] The second distillation column 40 may be the same or different as the first distillation column.

[0050] The second distillation column 40 is operated at a liquid phase bottom temperature in a range of from 120 to 179°C, and a vapor phase bottom pressure in a range of from 5 to 55 kPa. Preferably, the bottom of second distillation column 40 is operated at a liquid phase temperature at least 5 Celsius degrees, preferably at least 10 Celsius degrees, below the operating temperature of the top of the first distillation column 20. A water feed stream (not shown) may optionally be provided to the bottom of the second distillation column 40.

[0051] In the second distillation column 40, a second overhead stream 44 is separated from a second bottoms stream 46. The second overhead stream 44 comprises a second portion of the target glycol. The flow rate of the target glycol in the second overhead stream 44 is in a range of from 10 to 90 wt.% of the flow rate of the target glycol present in the feed mixture 12, depending on the yield of target glycol in the first overhead stream 24, such that the concentration of the target glycol of the combined first overhead stream 24 and the first bottoms stream 26 is within the range of 90 wt.% - 100 wt.%, more preferably in the range of 95 wt.% - 100 wt.%. The second bottoms stream 46 comprises the C2-C8 diol. Preferably, the second bottoms stream 46 may comprise a portion of the target glycol, for example, 10 wt.%, more preferably 5 wt.% of the amount of target glycol present in the feed mixture 12, to provide a higher purity of the target glycol in the second overhead stream 44, to decrease the temperature delta between top and bottom and to improve the ease of operation. In one embodiment of Fig. 1, 50 wt.% of the target glycol in the feed mixture 12 is recovered in the first overhead stream 24, while from 40 - 50 wt.% of the target glycol is recovered in the second overhead stream 44, and 0 to 10 wt.% of the target glycol is present in the second bottoms stream 46.

[0052] The second bottoms stream 46 is directed to a third distillation column 60 (as illustrated in Fig. 2) for further recovery of the residual target glycol or disposed of. Recovered residual target glycol may be recycled to the feed mixture 12 and/or the first distillation column 20. Alternatively, the second bottoms stream 46 may be burned as a fuel source for the process. A portion of the second bottoms stream 46 is passed through reboiler 42, which may be the same type as reboiler 22.

[0053] The first overhead stream 24 is directed to the reboiler 42 to heat and vaporize a portion of the second bottoms stream 46. The vaporized portion of the bottoms stream 48 is returned to the second distillation column 40 to drive fractionation.

[0054] The cooled first overhead stream 32 is then sent for final processing. Preferably, a portion of the cooled first overhead stream 32 is returned as a reflux stream 34 to the first distillation column 20. The flow rate of the reflux stream 34 is used for controlling temperature in the first distillation column 20.

[0055] Fig. 2 illustrates the embodiment of Fig. 1 with a third distillation column 60. In the embodiment of Fig. 2, the second bottoms stream 46 is directed to the third distillation column 60 to separate a third overhead stream 64 from a third bottoms stream 66. The third overhead stream 64 comprises a third portion of the target glycol. The third bottoms stream 66 comprises the C2- C8 diol. In one embodiment of Fig. 2, one-third of the target glycol in the feed mixture 12 is recovered in the first overhead stream 24, one-third of the target glycol is recovered in the second overhead stream 44, up to one-third of the target glycol is recovered in the third overhead stream 64, and 0 to 10 wt.% of the target glycol is present in the third bottoms stream 66.

[0056] The third bottoms stream 66 may be processed to recover one or more of the compounds in the stream, recycled, disposed of, or used as a fuel source for generating heat in another part of the process. Recovered target glycol may be recycled to the feed mixture 12 and/or the first distillation column 20. A portion of the third bottoms stream 66 is passed through reboiler 62, which may be the same type as reboiler 22. The vaporized bottoms stream 68 is returned to the third distillation column 60 to drive fractionation.

[0057] The embodiment of Fig. 2 is particularly advantageous when excess heat is available from other sources, when excess heat can be obtained from the first distillation column 20, or to provide operational flexibility when potential heat sources should be optimized due to variations in temperature, duty and availability. An additional advantage is the possibility to reduce the temperature delta between top and bottom in column 40, which also aids in further improving the purity of the target glycol. [0058] Fig. 3 illustrates the embodiment of Fig. 2 with heat integration between the first 20 and second 40 distillation columns and between the second 40 and third 60 distillation columns. In this embodiment, the second overhead stream 44 is directed to the reboiler 62 to heat and vaporize a portion of the third bottoms stream 66. The vaporized bottoms stream 68 is returned to the third distillation column 60 to drive fractionation. Preferably, the water feed stream 14 is also directed to the second distillation column 40 to allow for increased operating temperature in the second column 40. In this embodiment, the operating conditions for the second distillation column 40 are in the same range as for the first distillation column 20. The third distillation column 60 is operated at a temperature in a range of from 120 to 179°C, and a pressure in a range of from 5 to 55 kPa. Preferably, the second distillation column 40 is operated at a bottom temperature at least 5 Celsius degrees, preferably at least 10 Celsius degrees, below the operating temperature of the top of the first distillation column 20. Preferably, the third distillation column 60 is operated at a bottom temperature at least 5 Celsius degrees, preferably at least 10 Celsius degrees, below the operating temperature of the top of the second distillation column 40. In another embodiment, the overhead vapor phase stream of column 20 can be heat integrated with the bottom of column 40, as well as with the bottom of column 60. Alternative heat integration options can be designed by a skilled person, within the scope of the current invention.

[0059] Preferably, a portion of the cooled second overhead stream 52 is returned as a reflux stream 54 to the second distillation column 40. The flow rate of the reflux stream 54 is used for controlling temperature in the second distillation column 40.

[0060] The embodiment of Fig. 3 is particularly advantageous when further reduction in energy requirements, operational costs or carbon footprint reduction is desired, in balance with higher capital investments for a 3 -column device relative to a 2-column system.

[0061] Fig. 4 illustrates another embodiment of the process of the present invention 10, wherein the feed mixture 12 is directed to the second distillation column 40, instead of the first distillation column 20. The second bottoms stream 46 is directed to the first distillation column 20.

[0062] The embodiment of Fig. 4 is particularly advantageous when volatile components are to be removed in a pasteurization section or via partial condensation, or when reactive volatile components require a low distillation temperature to avoid undesired reactions. [0063] For each of the embodiments of Figs. 1 - 4, preferably, MEG separated from C2-C8 diols by the process of the present invention meet or exceed the specifications for fiber-grade MEG, for example as provide in Table 2.

TABLE 2

[0064] The specifications provided in Table 2 provide a maximum of 0.05 wt.% DEG in the purified MEG. This would apply to a feed mixture 12 derived from a petroleum feedstock. For a feed mixture 12 derived from a renewable feedstock, the specifications would likely be adjusted to recite a maximum of 0.05 wt.% 1,2-BDO and/or 1.2-HDO.

[0065] The first overhead stream 24 may be passed through a partial condenser (not shown) to separate the water vapor from liquid-phase target glycol.

EXAMPLES

[0066] The following non-limiting examples of embodiments of the method of the present invention as claimed herein are provided for illustrative purposes only. [0067] Aspen+, Version 11 , was applied to model fractional distillation, applying a proprietary data package, which has been validated with available open-source Vapor Liquid equilibria data of the relevant components. Results are summarized in Table 3.

Comparative Example 1

[0068] A feed mixture of MEG, DEG, TEG and TTEG was supplied to a 10-tray distillation column at a feed rate of the separate components of 1000 kta (kiloton per annum), 100 kta, 10 kta and 1 kta, respectively, at tray number 7. Trays were numbered from top to bottom.

[0069] The pressure was set at 0.1 bara (10 kPa) at the condenser. A heat duty of 79.3 MW was required at a bottom temperature of 167.7°C, while a heat duty of -75.75 MW was obtained at a top temperature of 133.3°C. A temperature delta of 34.4 Celsius degrees between the bottom and top temperatures was derived. Further details are provided in Table 3, indicated as Comp. Ex. 1.

Comparative Example 2

[0070] Comparative Example 1 was repeated at 0.2 bara (20 kPa) pressure. A heat duty of 96.9 MW was required at a bottom temperature of 187.1°C, while a heat duty of -91.6 MW was obtained at a top temperature of 150.2°C. The bottom temperature was higher than the critical temperature where glycol thermal degradation occurs, and additional water feeding was required for stabilization (not included in the model calculation). Further details are provided in Table 3, indicated as Comp. Ex. 2.

Example 3

[0071] Comparative Example 1 was repeated in a first distillation at 0.8 bara (80 kPa) pressure at the condenser. 750.1 kta MEG was recovered, representing 75% of the MEG introduced in the first column. A heat duty of 47.5 MW was required at a bottom temperature of 196.3°C, while a heat duty of -38.2 MW was obtained at a top temperature of 190.2°C. The bottom temperature was higher than the critical temperature where glycol thermal degradation occurs, and additional water feeding was required for stabilization (not included in the model calculation). A temperature delta of 6.1 Celsius degrees between the bottom and top temperatures was derived. This low temperature delta is advantageous for heat integration.

[0072] The bottom stream of the first distillation column was introduced in a second 10-tray distillation column, at tray number 7, operated at 0.1 bara (10 kPa) condenser pressure. An additional amount of 248 kta MEG was recovered, representing 24.8% of the MEG introduced into the first column. A heat duty of 37.0 MW was required at a bottom temperature of 170.8°C, while a heat duty of -38.7 MW was obtained at a top temperature of 133.3°C.

[0073] The heat duty of -38.2 MW at 190.2°C, obtained from the top of the first column, was integrated in the bottom of the second distillation column, requiring a heat duty of 37.0 MW at 170.8°C. After integration, the total heat duty of the two columns was 47.5 MW, which is lower than the heat duties of 79.3 MW and 96.9 MW required for Comparative Examples 1 and 2, respectively. Further details are provided in Table 3, indicated as Example 3.

TABLE 3

Units: 1 kta (kiloton per annum) equals 1.0e+6 kg/8000 hours operational time; MW (megawatt). Example 4

[0074] The distillation rates of the two columns in Example 3 were 75% MEG and 24.8% MEG, respectively. Example 3 was repeated with 900 kta (90%) MEG and 99.2 kta (9.92%) MEG distillation rates for the first and second column, respectively. A heat duty of -52.4 MW was obtained at 190.2°C in the top of the first column, which was integrated with the heat duty of 71.7 MW at 172.3 °C in the bottom of the second column. The total heat duty required was 61.7 MW for the bottom of the first column combined with the duty generated in the top of the first column, heat integrated with the bottom of the second column (71.7 MW - 52.4 MW) = 19.3 MW, which results in a total heat duty of 81.0 MW, which is lower than the heat duty of 96.9 MW required in Comparative Example 2, and closely matches (within experimental error, in the absence of process optimisation) the duty of 79.3 MW required in Comparative Example 1. Preferably, the distillation rate ratio of the first column versus the second column should be equal to or lower than 90% glycol/10% glycol. Further details are provided in Table 4.

Example 5

[0075] Example 4 was repeated with 100 kta (10%) MEG and 893 kta (89.3%) MEG distillation rates for the first and second column, respectively. A heat duty of -4.7 MW was obtained at 190.2°C in the top of the first column, which was integrated with the heat duty of 61.8 MW at 165.7°C in the bottom of the second column. The total heat duty required was 14 MW for the bottom of the first column, combined with the delta between the top of the first column and the bottom of the second column (61.8 MW - 4.7 MW) = 57.1 MW, which results in a total heat duty of 71.1 MW, which is lower than the heat duties of 79.3 MW and 96.9 MW required for Comparative Examples 1 and 2, respectively. Preferably, the distillation rate ratio of the first column versus the second column should be equal to or higher than 10% glycol/90% glycol. Further details are provided in Table 4.

[0076] These examples have been provided for separation of MEG from DEG, TEG and TTEG, which are commercially derived from ethylene oxide. Comparable advantageous results could be obtained for separation of MPG from mixtures of MPG, DPG and TPG In other embodiments, mixtures of MEG and 1,2-HDO could be separated with high efficiency. Target glycols can be advantageously separated from a mixture of the target glycol and one or more C2- C8 diols, with reduced energy requirements, lower carbon footprint and lower variable costs, while realizing high target glycol purity, as demonstrated by the examples provided.

TABLE 4

Units: 1 kta (kiloton per annum) equals 1.0e+6 kg/8000 hours operational time; MW (megawatt).

[0077] While the embodiments are described with reference to various implementations and exploitations, it will be understood that these embodiments are illustrative and that the scope of the inventive subject matter is not limited to them. Many variations, modifications, additions and improvements are possible.