Patent Application
20100094066
This invention relates to the gas phase processing of polyhydric compounds having from about four to about twelve carbon atoms into oxygen-containing products.
The processing of sugars having three or more carbon atoms has traditionally been performed in liquid phase. For example, U.S. Pat. Nos. 5,276,181, 5,214,219, 5,616,817, 4,642,394, 2,852,270, 3,030,429, 4,404,411, 5,600,028, 4,366,332, and 4,433,184 provide examples of known methods for converting glycerol (a three carbon sugar) and/or sugars having more than three carbon atoms into products having commercial value. Typically, these reactions are conducted in aqueous (liquid) phases with hydrogen present at pressures above 100 bars. Catalyst deactivation and recovery are generally problematic with liquid phase reactions, however.
Gas phase packed-bed reactor processes have operational advantages over slurry reactions primarily because solid catalysts are easier to remove from gases than from liquids. Additionally, gas phase reactions tend to have greater product recovery and less catalyst deactivation than liquid phase reactions. For these reasons, the processing of sugars in gas phase reactions would be beneficial. Sugars, however, are generally considered non-volatile. Suppes and Sutterlin (WO 2007/053705) recently disclosed that glycerol could be evaporated and maintained in gaseous phase, thereby providing a higher selectivity and an extended catalyst life. U.S. Pat. No. 5,387,720 also revealed a gas phase reaction process, but it was limited to temperatures above 250° C. due to inadequate approaches to overcome the low vapor pressure of glycerol at lower temperatures. Although glycerol can be reacted in gas phase under specific reaction conditions, it is unknown whether sugars having four or more carbon atoms could also be evaporated and reacted in gas phase. One challenge is that the boiling points of sugars having four or more carbon atoms are much higher than the boiling point of glycerol (which is about 290° C.). Thus, there is a need for processes for the gas phase conversion of a sugar having four or more carbon atoms into a commercially valuable product(s). Furthermore, there is a need for separation processes to efficiently recover the reaction products.
Among the various aspects of the present invention is one aspect that provides a process for converting a polyhydric feedstock into an oxygen-containing product. The polyhydric feedstock comprises at least about 10% of water by weight and at least about 2% by weight of at least one polyhydric compound having from about four to about twelve carbon atoms and more than three hydroxyl groups. The process comprises evaporating the feedstock at a temperature from about 150° C. to about 340° C. and at a pressure from about 0.01 bars to about 200 bars to form a substantially gas phase reaction mixture. The process further comprises contacting the gas phase reaction mixture with a heterogeneous catalyst to facilitate a reaction such that the oxygen-containing product is formed.
Another aspect of the invention encompasses a process for separating an oxygen-containing product from a reaction product mixture comprising a polyhydric compound and hydrogen. The process comprises introducing the reaction product mixture into a dividing wall column. The dividing wall column comprises a rectifying section comprised of more than four stages, a walled rectifying section comprised of more than four stages, a walled stripping section comprised of more than eight stages, and a stripping section comprised of more than eight stages. The process further comprises collecting the oxygen-containing product.
A further aspect of the invention provides a process for converting a three-carbon alcohol feedstock into acrolein. The process comprises evaporating the alcohol feedstock at a temperature from about 180° C. to about 250° C. and at a pressure from about 0.03 bars to about 5 bars to form a substantially gas phase reaction mixture. The process further comprises contacting the reaction mixture with a heterogeneous catalyst at a temperature from about 200° C. to about 250° C., wherein the catalyst has a Hammett acidity value from about −2 to about −10.
Still another aspect of the invention encompasses a process for evaporating a feedstock comprising at least one polyhydric compound and more than about 0.5% by weight of a salt. The process comprises heating the feedstock to a temperature from about 150° C. to about 300° C. and at a pressure from about 0.01 bars to about 20 bars, wherein at least about 25% by weight of the polyhydric compound in the feedstock evaporates to form a substantially gas phase while the remainder of the polyhydric compound forms a liquid residue stream. The process further comprises cooling at least part of the liquid residue stream to a temperature less than about 120° C. to form a salt precipitate stream. Lastly, the process comprises purging the salt precipitate stream from the process and recycling at least part of the liquid residue stream.
Other aspects and features of the invention will be in part apparent and in part pointed out hereinafter.
It has been discovered that polyhydric compounds having from about four to about twelve carbon atoms may be evaporated and reacted in gas phase. Preferably, the reactions proceed over a packed bed of a heterogeneous catalyst, which permits ready recovery of the catalyst. Reactions include but are not limited to reactions in the presence of hydrogen, wherein a polyhydric compound having from about four to about twelve carbon atoms is converted to an oxygen-containing product such as acetol, acrolein, glycerol, sorbitol, propylene glycol, furfural, furfural derivatives, and the like.
One aspect of the invention provides a process for converting a polyhydric feedstock comprising at least one polyhydric compound into an oxygen-containing product. The process comprises 1) evaporating the polyhydric feedstock at a temperature between about 150° C. and about 340° C. and at a pressure between about 0.05 bars and about 200 bars to form a substantially gas phase reaction mixture, and 2) contacting the gas phase reaction mixture with a heterogeneous catalyst to facilitate a reaction.
In general, the polyhydric feedstock comprises at least one polyhydric compound and water. The polyhydric compound has from about four to about twelve carbon atoms and more than three hydroxyl groups. In one embodiment, the polyhydric compound may comprise four carbon atoms and more than three hydroxyl groups. In another embodiment, the polyhydric compound may comprise five carbon atoms and more than three hydroxyl groups. In yet another embodiment, the polyhydric compound may comprise six carbon atoms and more than three hydroxyl groups. In still another embodiment, the polyhydric compound may comprise seven carbon atoms and more than three hydroxyl groups. In an alternate embodiment, the polyhydric compound may comprise eight carbon atoms and more than three hydroxyl groups. In another alternate embodiment, the polyhydric compound may comprise nine carbon atoms and more than three hydroxyl groups. In yet another embodiment, the polyhydric compound may comprise ten carbon atoms and more than three hydroxyl groups. In a further embodiment, the polyhydric compound may comprise eleven carbon atoms and more than three hydroxyl groups. In still another one embodiment, the polyhydric compound may comprise twelve carbon atoms and more than three hydroxyl groups. Non-limiting examples of suitable polyhydric compounds include sugars (such as, for example, glucose, fructose, galactose, mannose, sucrose, and the like) and sugar alcohols (such as, for example, sorbitol, mannitol, erythritol, arabitol, maltitol, xylitol, and the like).
In general, the polyhydric compound will have a boiling point higher than about 290° C. The boiling point of the polyhydric compound may be about 295° C., about 300° C., about 305° C., about 310° C., about 315° C., about 320° C., about 325° C., about 330° C., about 335° C., about 340° C., about 345° C., about 350° C., about 355° C., about 360° C., about 365° C., about 370° C., about 375° C., about 380° C., about 385° C., or about 390° C. As an example, the boiling point of sorbitol has been estimated to be about 362° C. (see Example 1).
The polyhydric feedstock typically comprises at least about 2% by weight of the polyhydric compound. In general, the concentration of the polyhydric compound may range from about 2% to about 90% by weight of the feedstock. In one embodiment, the concentration of the polyhydric compound may range from about 2% to about 10% by weight of the feedstock. In another embodiment, the concentration of the polyhydric compound may range from about 10% to about 40% by weight of the feedstock. In a further embodiment, the concentration of the polyhydric compound may range from about 40% to about 90% by weight of the feedstock.
Typically, the polyhydric feedstock also comprises at least about 10% by weight of water. In general, the concentration of water may range from about 10% to about 95% by weight of the feedstock. In one embodiment, the concentration of water may range from about 10% to about 25% by weight of the feedstock. In another embodiment, the concentration of water may range from about 25% to about 50% by weight of the feedstock. In still another embodiment, the concentration of water may range from about 50% to about 95% by weight of the feedstock.
The polyhydric feedstock is generally heated to a temperature from about 150° C. to about 340° C. and a total pressure from about 0.01 bar to about 200 bars to evaporate the feedstock. The polyhydric feedstock may be heated to a temperature ranging from about 150° C. to about 200° C., from about 200° C. to about 250° C., from about 250° C. to about 300° C., or from about 300° C. to about 340° C. In a preferred embodiment, the polyhydric feedstock is heated to a temperature ranging from about 250° C. to about 300° C. The pressure of the evaporating step may range from about 0.01 bar to about 1 bar, from about 1 bar to about 5 bars, from about 5 bars to about 50 bars, or from about 50 bars to about 200 bars. In a preferred embodiment, the pressure may range from about 0.1 bar to about 2 bars. The temperature and the pressure of the evaporating step may be any combination of the above detailed ranges.
Evaporation of the polyhydric feedstock forms a substantially gas phase reaction mixture. In the context of the present invention, the phrase “substantially gas phase” means that there is no liquid in the gas phase reaction mixture. The substantially gas phase reaction mixture may comprise water vapor, however. In general, the concentration of water in the substantially gas phase reaction mixture depends upon the concentration of water in the starting polyhydric feedstock. The concentration of water in the gas phase reaction mixture may range from about 10% to about 95% by weight of the mixture. The concentration of water in the gas phase reaction mixture may range from about 10% to about 25%, from about 25% to about 50%, or from about 50% to about 95% by weight of the mixture.
The process further comprises contacting the substantially gas phase reaction mixture with a heterogenous catalyst to facilitate a reaction such that the oxygen-containing product is formed. Preferably, the heterogeneous catalyst is selected from the group consisting of 1) a catalyst containing at least one element from Groups I or VIII of the Periodic Table, 2) copper, 3) chromium, and 4) an acid catalyst having a Hammett acidity value of less than about +2. In a preferred embodiment, the acid catalyst has a Hammett acidity value from about −2 to about −10.
The reaction may be at least one reaction selected from the group consisting of 1) a dehydration reaction at a temperature between about 180° C. and about 300° C., 2) a hydrocracking reaction at a temperature between about 180° C. and about 300° C., 3) a hydrogenation reaction at a temperature less than about 320° C., and 4) hydrogenolysis reaction at a temperature less than about 320° C. Table 1 summarizes the preferred temperatures for the different reactions.
Hydrogenation and hydrogenolysis reactions are preferably performed with gas phase reaction mixtures further comprising water and hydrogen. The ratio of water to polyhydric feedstock may be about 1:0.33, about 1:0.5, about 1:1, about 1:2, or about 1:3. The ratio of hydrogen to polyhydric feedstock may be about 1:0.01, about 1:0.05, about 1:0.25, or about 1:0.5.
Dehydration reactions are also preferably performed with gas phase reaction mixtures further comprising water and hydrogen. The ratio of water to polyhydric feedstock may be about 1:0.33, about 1:0.5, about 1:1, about 1:2, or about 1:3. The ratio of hydrogen to polyhydric feedstock may be about 1:0.5, about 1:1, about 1:2.5, or about 1:5.
Typically, the pressure of the reaction ranges from about 0.01 bar and about 200 bars, and more preferably from about 0.1 bar to about 5 bars. Table 2 summarizes suitable reactions with the preferred pressures of operation. For dehydration reactions, the most preferred pressure is in the range from about 0.1 bar to about 2 bars. For hydrogenation and hydrogenolysis reactions, the most preferred pressure is in the range from about 0.1 bar to about 4 bars.
Non-limiting examples of oxygen-containing products formed by the process of the invention include acetol, acrolein, glycerol, sorbitol, propylene glycol, furfural, furfural derivatives, and the like.
One process of the invention comprises the hydrocracking of a polyhydric compound having from about four to about twelve carbon atoms into a three-carbon compound (i.e., Scheme 1). The general conditions of the process are summarized in Tables 1 and 2. Preferably, the reaction is in a gas phase packed reactor bed comprising the heterogeneous catalyst.
The hydrocracking process detailed above may produce products such as acetol, propylene glycol, and the like. Generally, the heterogeneous catalysts that are used are those known to be effective for hydrogenation. Non-limiting examples of suitable heterogeneous catalysts include as palladium, nickel, ruthenium, copper, copper zinc, copper chromium, and the like. The polyhydric feedstock generally comprises water from about 10% to about 90% by weight, and more preferably, from about 25% to about 85% by weight of water. The near optimal temperature of the hydrocracking process is about 260° C.
In a preferred embodiment, the process comprises a flash evaporation process and a polyhydric feedstock comprising equal masses of water and polyhydric compound having from about four to about twelve carbon atoms. The polyhydric feedstock along with at least some diluent (which may be water in the feedstock) may be heated to a temperature of about 300° C. in a preheater in the presence of a partial pressure of water of about 3 bars (e.g. a static pressure of 3 bars), after which the pressure may be reduced to about 0.2 bar as it passes through a throttling device such as a control valve to facilitate flash evaporation. After the throttling processing, the feedstock may be mixed with a gas or vapors to facilitate further evaporation. The mixture may be then heated in a heat exchanger to the desired reaction temperature. The gas phase mixture, free of liquids, optionally may be passed through a knockout pot and then introduced to the packed-bed reactor. The gas or vapors mixed with the feedstock after the throttling device are generally at a temperature at least as hot as the feedstock after throttling. The preferred temperature range for the evaporator temperature is from about 250° C. to about 300° C.
Water is effective in reducing by-product formation during gas phase reactions of this invention and generally increases product selectivity. Hydrogen also decreases byproduct formation, thus the preferred gas phase reaction mixtures comprise both hydrogen and water. Without being bound by a particular theory, it is believed that a synergy exists in using the combination of hydrogen and water in the feedstock. In particular, the presence of hydrogen may cause the water to evaporate more evenly over the temperature range from about 70° C. to about 150° C., which may allow for easier heat recovery through counter-current heat exchange between the reactor effluents and the entering reagents. The following preferred conditions with respect to reactor feeds generally apply:
For hydrogenolysis to a three-carbon compound:
In embodiments in which the polyhydric feedstock comprises a six-carbon polyhydric compound, the following reaction conditions and concentrations of the six-carbon polyhydric compound are preferred. At about 0.1 bar of pressure and about 250° C., the preferred concentration of the six-carbon polyhydric compound in water is about 20% by weight. At about 0.1 bar of pressure and about 270° C., the preferred concentration of the six-carbon polyhydric compound in water is about 33% by weight. At about 0.1 bar of pressure and about 290° C., the preferred concentration of the six-carbon polyhydric compound in water is about 50% by weight. In all of these embodiments, the hydrogen to six-carbon polyhydric compound flow rate is from about 8 moles to about 1 mole.
Exemplary embodiments of this invention encompass the formation of a gas phase mixture comprising hydrogen and the polyhydric compound having from about four to about twelve carbon atoms. Because water is known to reduce by-product formation, its presence is also preferred.
Furthermore, charring may be a problem for polyhydric compounds having four or more carbon atoms. Charring occurs rapidly in the absence of water and at temperatures greater than about 180° C. Thus, the process of the invention reduces charring by manipulating the temperature and pressure of the evaporation step. In one embodiment, a mixture of water and polyhydric compound may be introduced into a moderate pressure heater to attain a temperature from about 210° C. to about 300° C. During this process, heating may cause water to selectively evaporate leading to lower water and higher polyhydric compound concentrations in the liquid phase than is in equilibrium with the vapor phase during the heating. Higher pressures may keep more water in the liquid phase while accumulating higher sensible heats due to higher temperatures.
In a preferred process, the pressure of the polyhydric compound-water mixture may be reduced (e.g. by flow through a valve), and the mixture may be rapidly mixed with a heated vapor or gas (e.g., hydrogen) of similar temperature but lower pressure. The sensible heat and heat in the vapor/gas may allow for rapid evaporation of the polyhydric compound without contact with a heated surface (flash evaporation), wherein the lack of contact with a heated surface substantially eliminates charring.
The mostly vaporous mixture 106 exiting the lower pressure evaporator optionally flows through a demisting device to knock out any liquid 108 entrained in the gas/vapor. The gas 107 flows to the reactor. The reactor preferably operates at a pressure slightly less than the lower pressure evaporator. Reactor products 109 flow to a condenser that preferably operates near a temperature of about 40° C. The mixture 110 flows to a separator that separates the gas 111/112 from the liquid 113. The gas is mostly hydrogen and water vapor and may be recycled. The liquid 113 is preferably separated in a divided wall column that produces a distillate containing water 116, a bottoms product 114, and at least one sidestream product 115.
In general, the less volatile the polyhydric compound the greater the tendency for the catalyst to become coated and deactivated. A preferred method to reduce catalyst deactivation is to use guard reactors such as those illustrated by
An alternative approach to overcoming the deactivation associated with the low volatility of polyhydric compounds may be to operate at lower concentrations of the polyhydric compounds. However, lower concentrations may increase conversion costs.
Another aspect of the invention encompasses a separation process for the separation of the oxygen-containing products derived from the polyhydric compounds. Separation costs may be high for oxygen-containing products. These products include but are not limited to propylene glycol, lactic acid, epichlorohydrin, furfual, acetol, and the like. These oxygen-containing products are generated from polyhydric compounds including but not limited to glycerol, sorbitol, sucrose, or glucose. This invention provides an improved method for separating the oxygen-containing products from the reaction product mixture.
The reaction product mixture may be separated in a divided-wall distillation column with four sections as indicated in the diagram (i.e., the column contains rectifying, walled rectifying, walled stripping, and stripping sections). The reaction product mixture is generally introduced into the column at a divided-tray (tray at location in column with dividing wall). The process includes the removal of an oxygen-containing product at a divided tray at the opposite side of the wall relative to the side where the reaction product mixture is fed into the column. The reaction product is not necessarily removed from the column at the same tray as the oxygen-containing product is removed. An example oxygen-containing product is propylene glycol. The column also generally has a bottoms product removed at or below the bottom tray and an overhead product removed at or above the top tray. An additional product or products may be removed as sidestreams at divided trays or trays without dividing walls.
The dividing wall is generally a functionally continuous wall expanding the walled rectifying and walled stripping sections. The wall functionally prevents mixing and vapors in the walled rectifying and walled stripping sections. Each of the four sections (i.e., rectifying, walled rectifying, walled stripping, and stripping sections) of the dividing wall distillation column performs at least about one to about 40 stages of separation and more preferably from about 2 to about 30 stages of separation. Per
More preferably, the rectifying, walled rectifying, walled stripping, and stripping sections comprise about 4-10 stages, about 4-10 stages, about 8-20 stages, and about 8-20 stages, respectively. In an exemplary embodiment, the rectifying, walled rectifying, walled stripping, and stripping sections comprise about 4 stages, about 4 stages, about 8 stages, and about 8 stages, respectively. The most preferred stages are dependent upon the feedstock composition and the desired degree of separation. In general, the process comprises introducing the reaction product mixture to a dividing wall column, and collecting the oxygen-containing product. In one embodiment, the oxygen-containing product may comprise more than about 10% of water by weight.
Any vapor-liquid separation means may be used to achieve separations (including use of trays and use of packing). High-efficiency packing is generally preferred due to the number of stages of this column. For example, suitable methods include near-stage-located evaporation, near-stage-located condensation, and adjustment of vapor and liquid flow distribution.
Another aspect of the invention encompasses a process for the conversion of a three-carbon alcohol feedstock to acrolein. The three-carbon alcohol may be acetol or glycerol. Tables 1 and 2 summarize exemplary examples of reaction conditions for converting acetol (i.e., Scheme 2) or glycerol (i.e., Scheme 3) to acrolein in a gas phase packed bed reactor. The gas phase conversion of acetol or glycerol to acrolein is demonstrated in Example 7.
The process comprises evaporating the three-carbon alcohol feedstock at a temperature from about 180° C. to about 250° C. and at a pressure from about 0.03 bar to about 5 bars to from a substantially gas phase reaction mixture. The phrase “substantially gas phase” is as defined above. The temperature of the evaporating step may range from about 180° C. to about 200° C., from about 200° C. to about 220° C., or from about 220° C. to about 250° C. The pressure of the evaporating step may range from about 0.03 bar to about 0.2 bar, from about 0.2 bar to about 1 bar, or from about 1 bar to about 5 bars. The temperature and the pressure of the evaporating step may be any combination of the ranges listed above.
The three-carbon alcohol feedstock generally comprises a three-carbon alcohol compound and water. In one embodiment, the three-carbon alcohol compound may be acetol, and the concentration of acetol in water may be about 20%, about 30%, about 40%, about 50%, about 60%, about 70%, about 80%, about 90%, or about 99% by weight. In another embodiment, the three-carbon alcohol compound may be glycerol, and the concentration of glycerol in water may be about 20%, about 30%, about 40%, about 50%, about 60%, about 70%, or about 75% by weight. The three-carbon alcohol feedstock may optionally be combined with other gases (such as, for example, hydrogen).
The process further comprises contacting the substantially gas phase reaction mixture with a heterogeneous catalyst at a temperature that ranges from about 200° C. to about 250° C. The temperature of the reacting step may be from about 200° C. to about 215° C., from about 215° C. to about 230° C., or from about 230° C. to about 250° C. The pressure of the reacting step may range from about 0.03 bar to about 0.2 bar, from about 0.2 bar to about 1 bar, or from about 1 bar to about 5 bars. The temperature and the pressure of the reacting step may be any combination of the ranges detailed above.
The heterogeneous catalyst may be an acidic catalyst having a Hammett acidity value between about −2 and about −10. The acid catalyst may be a solid acid catalyst or a solid, such as activated carbon, with phosphoric acid adsorbed on the catalyst. In a preferred embodiment, the catalyst may be phosphoric acid on an activated carbon. In one embodiment, the acetol feedstock may be a gaseous product of a dehydration reaction of glycerol. In another embodiment, the acetol feedstock may be a liquid comprising at least about 25% by weight of water, and the pressure of the reaction may range from about 0.1 bar to about 0.8 bar (such that lower densities of acrolein are formed at lower water contents). In a further embodiment, the three-carbon alcohol feedstock may be mixed with phosphoric acid prior to the process, wherein upon evaporation of the feedstock, the gas phase reaction mixture comprises less than about 0.5% by weight of phosphoric acid prior to contact with the heterogeneous catalyst. The concentration of phosphoric acid in the vapor phase reaction mixture may be less than about 0.5%, less than about 0.1%, less than about 0.05%, less than about 0.001%, or less than 0.0005% by weight.
Evaporation of Polyhydric Feedstock with Recycling
A further aspect of the invention provides a process for evaporating a polyhydric feedstock comprising at least about 0.5% by weight of a salt. The process comprises heating the feedstock to a temperature from about 150° C. to about 300° C. and at a pressure from about 0.01 bar to about 20 bars, wherein at least about 25% by weight of the polyhydric compound in the feedstock evaporates to form a substantially gas phase while the remainder of the polyhydric compound forms a liquid residue stream. The process further comprises cooling at least part of the liquid residue stream to a temperature less than about 120° C. to form a salt precipitate stream. Lastly, the process comprises separating and purging the salt precipitate stream from the process and recycling at least part of the liquid residue stream.
In general, the polyhydric feedstock may comprise from about 0.5% to about 10%, or more preferably from about 0.5% to about 5% of a salt. The concentration of salt in the polyhydric feedstock may range from about 0.5% to about 1%, from about 1% to about 2%, from about 2% to about 5%, or from about 5% to about 10% by weight. Non-limiting examples of suitable salts include KCl, NaCl, K2SO4, KHSO4, Na2SO4, NaHSO4, K3PO4, K2HPO4, KH2PO4, Na3PO4, Na2HPO4, NaH2PO4, and combinations thereof. The polyhydric feedstock may comprise at least one polyhydric compound having from about four to about twelve carbon atoms and more than three hydroxyl groups, or the polyhydric feedstock may comprise a three-carbon alcohol. The polyhydric feedstock may further comprise at least about 10% of water by weight.
The evaporating step may comprise heating the polyhydric feedstock comprising the salt to a temperature ranging from about 150° C. to about 200° C., from about 200° C. to about 250° C., or from about 250° C. to about 300° C. The pressure of evaporating step may range from about 0.01 bar to about 0.1 bar, from about 0.1 bar to about 1 bar, from about 1 bar to about 10 bars, or from about 10 bars to about 20 bars. The evaporating step may comprise any combination of the above listed temperature and pressure ranges.
At least a part of the liquid residue steam may be cooled to a temperature less than about 120° C., less than about 100° C., less than about 80° C., less than about 60° C., or less than about 40° C. The salt precipitate stream may be removed from the process using any separation technique known to those of skill in the art. For example, the salt precipitate stream may be separated and purged from the process using a filtering technique. Alternatively, the solid-liquid mixture (i.e., the liquid residue and the salt precipitate) may be passed through a solid-liquid separator (preferably up-flow through a sand bed) that substantially removes the salt precipitate from the liquid residue stream.
The process may further comprise mixing the recycled residue stream with an entering feedstock stream to form a mixture stream. The mixing of the recycled residue stream with an entering feedstock stream essentially cools the recycled residue stream and heats the entering the feedstock stream such that a precipitate is formed. The process further comprises separating the precipitate from the mixture stream, and purging the precipitate from the process.
As various changes could be made in the above-described processes without departing from the scope of the invention, it is intended that all matter contained in the above description and the examples presented below, shall be interpreted as illustrative and not in a limiting sense.
The following examples illustrate various embodiments of the invention.
A thermogravimetric analysis (TGA) method was used to estimate the vapor pressure of sorbitol. Table 3 reports these estimated vapor pressures. An extrapolation of this data estimates the boiling point of sorbitol at 362° C.
An equal-mass mixture of sorbitol (MW=182) and water (MW=18) would have a sorbitol mole fraction of 9 mol %. Based on Raoult's Law, the dew point pressure of this 50:50 mixture at 250° C. and 275° C. are about 0.14 bars and 0.5 bars, respectively. Addition of hydrogen to the mixture would further increase the dew point.
The evaporation and subsequent condensation of sorbitol was evaluated in a continuous flow process with feed at 2.5 wt %. The data are summarized in Table 4.
1Recovery calculated from TGA analysis.
2Recovery calculated from HPLC analysis.
The condensed aqueous-sorbitol mixture was clear, without odor, and absent of degradation products. The data of Table 4 validate the hypothesis that sorbitol can be successfully evaporated to a gas phase and recovered as a condensate.
The evaporation and condensation of sorbitol 4 had recoveries less than would be expected based on the TGA estimates of vapor pressure. Higher water content, hydrogen dilution, and flash evaporation methods could be used promote sorbitol evaporation in preferred processes. Based on available data and extrapolating to account for improved flash evaporation means (as compared to the experimental system of Table 4), the more preferred operating conditions for a 50 wt % solution of sorbitol in water are about 0.01 to 0.04 bars at 250° C. and 0.04-0.1 bars at 275° C. The pressure range is under the dew point pressure, preferably with a margin of safety.
Evaporation of glucose at similar condition resulted is substantial degradation of the glucose. Enhanced flash evaporation methods are needed that would minimize the time at which glucose is on hot surfaces to promote evaporation.
The gas phase conversions of glycerol, sorbitol, and glucose were carried out at the temperature range from 240° C. to 290° C. and a pressure of 1 bar. The reaction system consisted of evaporator, trap, fix-bed reactor and ice-water condenser. For each run, 70 g of catalyst was loaded into copper fix-bed reactor (i.d.=1 inch). While the system reached the desired temperature and pressure, the aqueous sugar/sugar alcohol solution (mass concentration in solution was 2.5%) was pumped into reaction system by a micro-pump at the flow rate of approximately 200 gram per hour. Hydrogen was fed at 7-8 l/min. The samples were collected from ice-water condenser and analyzed by Hewlett-Packard gas chromatography (GC) and high performance liquid chromatography (HPLC) without further delay. Table 5 summarizes the conversions of the two prominent products for these reactions.
The results of Table 5 validate the gas phase reactions of C4-C12 sugar embodiments of this invention. The trap prior to the reactor assured that only gases made it through the reactor. The changing conversions with different catalysts indicated that the conversions were catalyst-dependent (not homogeneous). The product yields were calculated based on a normalization method. A GCMS library was used to identify the products (other than acetol and propylene glycol) with said products being the most probable but with less than 100% certainty in the identification.
Preliminary data were obtained for the hydrogenation of sorbitol in the gas phase over a solid catalyst. Prior to operating the reactor, the evaporation was evaluated at 80% water and an excess of hydrogen at temperatures of 260° C. and 290° C. A trap between the evaporator and condenser was estimated to knock out about a fifth of the sorbitol, indicating that about four-fifths had evaporated—this evaporation was performed for about 3 hours. Next, the plug flow reactor was placed between the evaporator and condenser.
The GC registered >90% selectivity to the mixture of acetol and propylene glycol (known to be in equilibrium). Variations in the reactor temperature produced different ratios of acetol and propylene glycol that are completely consistent with trends observed with the hydrogenolysis of glycerol.
Ethylene glycol is formed from the direct hydrogenolysis of glycerol. It appears that the essential absence of glycerol in the sorbitol reaction (it immediately reacts in gas phase) leads to essentially zero ethylene glycol production—actually an improvement over the glycerol reaction.
These preliminary runs were operated at about 1.2 bars of pressure. Pressure is an important degree of freedom—lower pressures can result in the need for lower water contents (e.g. 25% versus 80%) and evaporation at lower temperatures (totally avoiding forming of oligomers). It is anticipated that fructose and glucose will exhibit similar volatility and reactivity as sorbitol.
The hypothesized mechanism includes the following steps: 1) adsorption of C6 sugar onto catalyst; 2) hydrocracking of C6 sugar to glycerol and another C3 compound; 3) essentially immediate dehydration of glycerol to acetol, and 4) rapidly established chemical equilibrium between acetol and propylene glycol.
The gas-phase reaction of glycerol was performed in the reactor systems illustrated by
The isothermal data illustrates that good conversions and selectivities were attained with the gas phase reaction at the specified conditions. The impact of water in going from 60% to 90% water was minimal. When the heat of reaction was allowed to increase the reaction temperature, less propylene glycol was formed and more of the product remained as acetol (equilibrium limitations). Acetol tends to be more reactive and leads to more undesirable by-products.
The gas-phase reaction of sorbitol was performed in the reactor systems illustrated by
The gas-phase reaction of sorbitol was performed in the reactor systems illustrated by
Catalyst was prepared by taking 30 grams of activated carbon and soaking the carbon in about 60 grams of a 40% by weight phosphoric acid solution. After an hour at 25° C., free liquid was drained and the solids were dried in an oven at about 100° C. The catalyst was loaded in a 0.5″ ID tube to form a packed be reactor.
Feed mixtures prepared for separate reactor runs included 20% glycerol in water and 50% acetol in water. Feed mixtures were fed through a heat exchanger at about 200 ml/hr to a temperature of 290° C. to form a vapor phase that was sent through a knockout drum to remove entrained liquids. The liquid-free mixture was then fed to the packed-bed reactor maintained at 250° C. The reactor effluents were condensed to a temperature of about 40° C. and the liquid product was evaluated using gas chromatography. Table 9 summarizes the compositions for the two feed compositions.
Preliminary data in Table 9 confirms that acetol does undergo this reaction with higher conversions than when glycerol is the feed. A process option would be to feed gas phase products from a glycerol-to-acetol conversion over a copper-chromite catalyst directly into a gas phase reactor with solid acid catalyst to produce an overall acrolein yield in excess of 85%. Higher yields would be expected at lower temperatures.
| Filing Document | Filing Date | Country | Kind | 371c Date |
|---|---|---|---|---|
| PCT/US08/51257 | 1/17/2008 | WO | 00 | 12/29/2009 |
| Number | Date | Country | |
|---|---|---|---|
| 60880870 | Jan 2007 | US | |
| 60962630 | Jul 2007 | US |
