US20100242346A1

US20100242346A1 - Processes for the esterification of free fatty acids and the production of biodiesel

Info

Publication number: US20100242346A1
Application number: US12/679,112
Authority: US
Inventors: Donald LeRoy Bunning; Louis A. Kapicak; Thomas Arthur Maliszwski; David James Schreck
Original assignee: Best Energies Inc
Current assignee: Best Energies Inc
Priority date: 2007-09-19
Filing date: 2008-09-17
Publication date: 2010-09-30
Also published as: WO2009039144A1

Abstract

The effluent from an acid esterification of free fatty acid with alkanol to produce alkyl ester of fatty acid is contacted with glycerin to remove water and alkanol. The alkanol separated with the glycerin can be recycled to the acid esterification by contacting the glycerin with fatty acid-containing feed being passed to the acid esterification.

Description

FIELD OF THE INVENTION

This invention pertains to processes for making alkyl esters, especially biodiesel, from feeds containing free fatty acids and to the esterification by acid catalysis of free fatty acids with lower alkanol.

BACKGROUND TO THE INVENTION

Biodiesel is being used as an alternative or supplement to petroleum-derived diesel fuel. Biodiesel is a mixture of alkyl esters which can be made from various bio-generated oils and fats from vegetable and animal sources.
One process for making biodiesel involves the transesterification of triglycerides in the oils or fats with a lower alkanol in the presence of a base catalyst to produce alkyl ester and a glycerin co-product. Unfortunately, most oils and fats useful as triglyceride-containing feeds for transesterification also contain free fatty acids which are not converted under typical transesterification conditions to biodiesel. Moreover, biodiesel must meet demanding product specifications. See, for instance, ASTM D 6751, American Society for Testing and Materials. These specifications, among other requirements, limit the amount of free fatty acid that can be contained in biodiesel.
Free fatty acids, while not acceptable in biodiesel, can be converted to esters suitable for inclusion in biodiesel. Numerous processes have been proposed. See, for instance, U.S. Pat. No. 6,822,105; U.S. Patent Application Publication No. 2005/0204612; Canakci, et al., Transactions of ASAE, 42, 5, pp. 1203-10 (1999), King, “Esterification: Chemistry and Processing”, Biodiesel Short Course, Quebec City, Canada, May 12-13, 2007, and Van Gerpen, et al., “Biodiesel Production Technology, August 2002-January 2004, National Renewable Energy Laboratory NREL/SR-510-36244, July 2004.
The esterification of free fatty acids with alkanol results in water being generated as a co-product. As the acid esterification is equilibrium limited reaction, a large excess of alkanol is typically used to shift the equilibrium towards the production of the sought alkyl ester. The unreacted alkanol must be recovered from the esterification product and recycled to provide an economically attractive process.
Turck in U.S. Pat. No. 6,538,146 discloses a method for producing fatty acid esters of alkyl alcohols using oils that contain free fatty acids and phosphatides. He summarizes his process as treating the feed with a base mixture of glycerin and a catalyst to produce a two phase mixture with the neutralized free fatty acids passing into the glycerin phase. The oil phase containing the triglycerides is then subjected to transesterification. See column 2, lines 35 et seq. At column 4, lines 41 et seq., Turck poses that the free fatty acids can be separated per WO 95/02661 and then subjected to esterification with an alcohol. The esterified product can be added to the transesterification mixture.
Koncar, et al., in U.S. Pat. No. 6,696,583 disclose methods for preparing fatty acid alkyl esters in which fatty acids contained in a glycerin phase from a transesterification are separated and mixed with an esterification mixture containing triglycerides and is subjected to esterification to form fatty acid esters. The object of their process is to process the fatty acid phase in the untreated state, i.e., without purification and removal of sulfuric acid. The esterification product is then transesterified with alcohol. Koncar, et al, refer to EP-A-0 708 813 as disclosing the esterification of free fatty acids at column 2, lines 26 to 34.
Iyer in U.S. 2006/0293533 discloses a process for the esterification and transesterification of fats and oils using one or more heterogeneous catalysts. See also, Clements, US 2006/0224006.
Lin, et al., in U.S. Pat. No. 7,122,688 disclose the use of acidic mesoporous silicates as catalysts for esterifying fatty acids and transesterifying oils.
Various processes are commercially offered for making biodiesel by transesterification of triglycerides. Several of these processes provide options for the esterification of free fatty acids to the corresponding esters. Desmet Ballestera have a process in which feedstocks preferably containing more than 1 percent free fatty acid is subjected to vacuum-steam stripping to remove free fatty acids. The distillate can then be subjected to esterification conditions comprising elevated temperature, methanol and sulfuric acid catalyst to make a methyl ester. The esterification product is subjected to a flash and phase separation to recover methanol for recycle and separate water, glycerin and sulfuric acid from the methyl ester. Kemper, Desmet Ballestra Biodiesel Production Technology, Biodiesel Short Course, Quebec City, Canada, May 12-13, 2007.
Crown Iron Works Company also provides a biodiesel manufacturing process where acid esterification is used to convert free fatty acids to methyl esters for biodiesel. They caution that acid esterification should only be used if disposal of the fatty acids or soaps thereof is not economic or possible or the feed used generates a lot of fatty acids. They note that acid esterification increases capital and production costs, and that sulfuric acid creates sulfates which increase the removal cost from glycerin. Waranica, Crown Iron Works Biodiesel Production Technology, Biodiesel Short Course, Quebec Canada.

SUMMARY

This invention provides improved processes for the esterification of feeds containing free fatty acids especially glycerides feeds containing free fatty acids wherein the glycerides in the feed are suitable to be transesterified with alkanol to produce biodiesel. The esterification of the free fatty acids is conducted with a stoichiometric excess of alkanol to provide an esterification effluent containing alkyl esters, unreacted alkanol and water. In accordance with processes of this invention, at least a portion of the esterification effluent is contacted with glycerin to reduce the concentration of water and alkanol in the esterification effluent containing alkyl esters, and the glycerin is separated by phase separation.
In one broad aspect of the processes of this invention, a feed containing free fatty acid is subjected to acidic esterification conditions in the presence of a stoichiometric excess of alkanol to provide an esterification effluent containing alkyl ester, water and unreacted alkanol. The esterification effluent is contacted with glycerin to form a two phase mixture. Water and alkanol are extracted into the glycerin phase which can be phase separated from the oil phase comprising alkyl ester. The processes of this aspect of the invention are particularly useful in conjunction with facilities to produce biodiesel from glycerides by transesterification as glycerin is available as a co-product of the transesterification. Additionally, where the acid catalyst is contained in the esterification effluent, the glycerin phase can remove the acid catalyst from the esterification effluent.
In further detail, this broad aspect of the invention for esterifying feed containing free fatty acid, especially free fatty acid of from about 8 to 30, say 14 to 24, carbon atoms, comprises:

- a. contacting the feed with a stoichiometric excess of alkanol, preferably a molar ratio of alkanol to free fatty acid of at least about 1.1:1, say, about 2:1 to 20:1 or 30:1 or more, under acidic esterification conditions including the presence of acid catalyst at elevated temperature, e.g., up to 200° C. or more, preferably less than about 150° C., say, less than about 120° C., and more preferably between about 35° C. and 100° C., for a time sufficient to provide an esterification effluent comprising alkyl ester of the free fatty acid, water and unreacted alkanol;
- b. contacting at least a portion of the esterification effluent containing alkyl ester with sufficient glycerin to form a glycerin-containing phase and an oil phase comprising alkyl ester, said contacting being for a time and under conditions sufficient that said oil phase has a lower water and a lower alkanol concentration than the esterification effluent and said glycerin phase contains water, preferably at least about 50, and sometimes at least about 60, percent of the water contained in the esterification effluent, and alkanol, preferably at least about 20, preferably at least about 30, mass percent of the alkanol in contained in the esterification effluent; and
- c. phase separating the oil phase and the glycerin phase.

The glycerin used for the contacting may be derived from any suitable source. One suitable source is glycerin-containing co-product from the transesterification of glycerides. Typically, the glycerin used for the contacting with the esterification effluent comprises at least about 40, preferably at least about 50, mass percent glycerin. Often the mass ratio of glycerin to esterification effluent is at least about 0.01:1, more preferably from about 0.05:1 to 0.5:1.
All or a portion of the esterification effluent is contacted with glycerin. Where a portion of the esterification effluent is contacted, that portion, which may be an aliquot portion or a portion remaining after a separation unit operation, contains alkyl ester. For instance, the esterification effluent may be subjected to a stripping operation to remove some of the water and alkanol. Where higher molar ratios of alkanol to free fatty acid are used for the esterification, an alkanol-containing phase may form. This alkanol-containing phase may be separated by phase separation prior to the contacting with the glycerin. Preferably the contacting is under conditions that minimize reversion of alkyl ester to free fatty acid or conversion to glycerides. Generally these conditions are provided by removal or inactivation of at least a portion of the catalyst. Although temperature reduction may also suffice to reduce reversion or conversion, it is usually not necessary. Frequently the contacting is at a temperature of from about 35° C. to 150° C. for 0.01 to 10 hours.
In a preferred embodiment of the invention, after the contacting and phase separation, the glycerin phase containing alkanol is contacted with the feed to be subjected to esterification whereby the feed extracts from the glycerin phase a portion of the alkanol. In a further aspect of the invention, the acidic esterification is integrated with a transesterification process to make biodiesel from glycerides. Not only can the transesterification process provide glycerin for removal of water and alkanol from the esterification effluent but also integration can enhance the economics of the process by reducing energy consumption and capital expense.
Preferably the alkanol is lower alkanol of up to 6, preferably 1 to 3, carbon atoms, especially methanol. In another preferred embodiment of this aspect of the invention the feed comprises glycerides of fatty acid.
In another broad aspect of the processes of this invention, a portion of the alkanol for the acid esterification of free fatty acids is obtained from a glycerin-containing liquid that contains alkanol. In this aspect of the invention, the glycerin-containing liquid is contacted with feed for the esterification and a portion of the alkanol partitions to the feed. In processes for making biodiesel from glycerides, not only is glycerin a co-product, but also, glycerin contains unreacted alkanol used in the transesterification. Hence, the processes of this aspect of the invention provide a low energy and low capital means for recovering alkanol from the glycerin co-product. Alternatively, or in addition, the glycerin may be that used to remove alkanol from the esterification effluent.
In this broad aspect, the processes of this invention for the esterification of feed containing free fatty acid, especially free fatty acid of from about 8 to 30, say 14 to 24, carbon atoms, comprise:

- a. contacting the feed with a glycerin solution containing alkanol, preferably where the solution contains a mass ratio of alkanol to glycerin of at least about 0.05:1, say, 0.2:1 to 2:1, under conditions sufficient to provide a feed phase containing an increased concentration of alkanol and a spent glycerin phase having a decreased concentration of alkanol, preferably reduced by at least about 20, more preferably at least about 30, mass percent;
- b. phase separating the feed phase and the spent glycerin phase; and
- c. contacting the feed phase with a stoichiometric excess of alkanol under acidic esterification conditions including the presence of acid catalyst at elevated temperature, for a time sufficient to provide an esterification effluent comprising alkyl ester of the free fatty acid, water and unreacted alkanol.

If desired, alkyl ester can be added to enhance the solubility of alkanol in the feed, especially where the feed comprises glyceride. The amount of alkyl ester added can be relatively minor yet significant additional alkanol solubility can be obtained. Often the alkyl ester is provided in a mass ratio to feed of at least about 0.01:1, say, about 0.05:1 to 0.2:1. The solubility of alkanol in the oil phase will depend, among other things, upon the type of alkanol and the content of free fatty acid in the feed. The alkyl ester may be from any suitable source including, but not limited to, alkyl ester from the esterification process and biodiesel.
In yet another broad aspect of the invention, the acid esterification is conducted using a glycerin-soluble, acid catalyst. The esterification effluent, which contains alkyl ester of free fatty acid, water, acid catalyst and unreacted alkanol, is contacted with glycerin to remove acid catalyst from the esterification effluent. Simultaneously or sequentially the acidic glycerin stream can be contacted with soaps of free fatty acids to generate free fatty acids which can be fed to the esterification. For instance, the glycerin used for contacting the esterification effluent may already contain soaps. Alternatively, or in addition, the glycerin, after contact with the esterification effluent may contact a soap-containing stream.
By this aspect of the invention, catalyst that would otherwise be discarded is effectively used to provide free fatty acids. For instance, the acidic esterification may be conducted to effect only a partial conversion of free fatty acid in the feed, e.g., between about 50 and 95 or 97 mass percent of the free fatty acid is converted to alkyl esters. The unreacted free fatty acids can be saponified and removed from the oil phase, and then acidified for recycle. An advantage of this aspect of the invention is that the esterification need not achieve a high conversion per pass of free fatty acid. Thus additional reactor stages or high alkanol to free fatty acid ratios that have typically been used to achieve high conversion, can be avoided to save in capital and energy costs. Also, if the esterification is used in conjunction with a base transesterification of glyceride, soaps formed as a by-product can be recovered using the acidic glycerin as free fatty acids, and the free fatty acids can be subjected to acidic esterification. As often the glycerin co-product from a base catalyzed transesterification also contains base catalyst, the glycerin co-product may be used as at least a portion of the basic glycerin to saponify free fatty acids in the esterification effluent as well as provide soaps from the transesterification.
In this broad aspect, the processes of this invention for the esterification of feed containing free fatty acid, especially free fatty acid of from about 8 to 30, say 14 to 24, carbon atoms, comprise:

- a. contacting the feed with a stoichiometric excess of alkanol under acidic esterification conditions including the presence of glycerin-soluble, acid catalyst at elevated temperature, for a time sufficient to provide an esterification effluent comprising alkyl ester of the free fatty acid, acid catalyst, water and unreacted alkanol;
- b. contacting the esterification effluent with glycerin whereby a glycerin-containing phase and an oil phase comprising alkyl ester are formed, said contacting being for a time and under conditions sufficient that the oil phase has a lower acid catalyst concentration than the esterification effluent and a lower water concentration than the esterification effluent and said glycerin phase contains acid catalyst (which may be unneutralized, partially neutralized or essentially completely neutralized), water and alkanol;
- c. phase separating the oil phase and the glycerin phase;
- d. contacting the glycerin phase with soaps of free fatty acids to generate free fatty acids, preferably unneutralized or partially neutralized acid catalyst generates at least a portion of the free fatty acids;
- e. separating free fatty acids from said glycerin phase; and
- f. recycling the free fatty acids to step (a).

The separation of free fatty acid from glycerin of step (e) may be conducted in any convenient manner. For instance, free fatty acid may form an oil phase and can be removed by phase separation. Preferably the feed is contacted with the glycerin phase of step (e) to effect separation of the free fatty acids.
In a fourth broad aspect of the processes of this invention, an oil-soluble, acid catalyst is used for the acidic esterification and is recovered from the esterification effluent by phase separation of an alkanol phase from the oil phase. The alkanol phase is formed by providing alkanol in an amount in excess of that which is miscible with the oil phase. The acid catalyst, due to the more polar nature of the alkanol phase, preferentially partitions to the alkanol phase. Thus not only can the acid catalyst be recovered, e.g., for recycle to the esterification, but also, the recovery involves low capital and energy costs. Water will also tend to preferentially partition to the alkanol phase. If desired, all or a portion of the alkanol phase can be subjected to a water removal unit operation. As the volume of the alkanol phase can be minor in comparison to the esterification effluent, the energy required for the water removal, e.g., by distillation, is not unduly high.
In this broad aspect of the processes of this invention, esterification of feed containing free fatty acid, especially free fatty acid of from about 8 to 30, say 14 to 24, carbon atoms, comprises:

- a. contacting the feed with a stoichiometric excess of alkanol under acidic esterification conditions including the presence of an oil-soluble, acid catalyst, preferably an organic acid having at least about 4 carbon atoms, preferably an organosulfonic acid, at elevated temperature, for a time sufficient to provide an esterification effluent comprising alkyl ester of the free fatty acid, water, acid catalyst and unreacted alkanol;
- b. providing sufficient alkanol in said esterification effluent to enable to form an alkanol phase containing alkanol, water and oil-soluble acid catalyst and an oil phase containing alkyl ester and unreacted free fatty acid;
- c. phase separating the alkanol phase and the oil phase; and
- d. passing at least a portion of the alkanol phase to step (a).

The alkanol required for forming the alkanol phase may be present in step (a) or may, preferably, be through the addition of alkanol subsequent to step (a) such that a single, homogenous phase exists in step (a). The amount of alkanol used to form the separate alkanol phase is preferably sufficiently small that the separate alkanol phase is only a minor portion of the esterification effluent, and sometimes is less than about 10, even less than about 5, volume percent of the total of the alkanol phase and the oil phase. Preferably water is removed from at least a portion of the alkanol phase, e.g., by distillation. Where the alkanol boils at a lower temperature than water or co-boils with water, it may be desirable to provide high boiling liquid such that the oil-soluble catalyst is maintained in a liquid phase to facilitate handling and avoid any decomposition of the oil-soluble catalyst.
In a fifth broad aspect of the invention, the esterification is conducted under esterification temperatures less than about 120° C. in the presence of acidic catalyst that is soluble in alkanol and substantially insoluble in alkyl ester of fatty acid. The alkanol is preferably provided in an amount sufficient to provide an esterification effluent having an alkanol and catalyst phase and an alkyl ester-containing phase. In this aspect of the invention, the esterification processes only convert a portion of the free fatty acid in the feed to alkyl ester, e.g., from about 50 to 95 or 97 mass percent. The alkyl-ester-containing phase is contacted with basic glycerin to convert unreacted free fatty acid to soaps and a glycerin phase containing the soaps is separated from the alkyl ester-containing phase. Accordingly, attractive energy costs can be obtained. Moreover, capital savings can be achieved in that adequate conversion of free fatty acids can often be achieved in a few reaction stages, and sometimes even a single reaction stage.
Most significantly in the processes of this aspect of the processes of this invention, mild reaction conditions including lower temperatures and, potentially lower concentrations of acid catalyst, attenuate the relative rate of the hydrolysis of esters to acids as compared to the rate of acid esterification. Often the concentration of acidic catalyst is below about 1, and sometimes less than about 0.5, mass percent. Hence, flexibility is provided by the processes of this aspect of the invention to balance rate and catalyst concentration while still attenuating the relative rate of hydrolysis.
Processes for esterification of this fifth aspect of the invention use a glycerides-containing feed containing at least about 5 mass percent free fatty acid and comprise:

- a. contacting the feed with a stoichiometric excess of alkanol, preferably with a molar ratio of alkanol to free fatty acid between about 10:1 to 30:1, under acidic esterification conditions including the presence of acid catalyst that is soluble in alkanol and substantially insoluble in alkyl ester of fatty acid, preferably sulfuric acid, at elevated temperature below about 120° C., more preferably below about 100° C., say 35° C. to 100° C., for a time sufficient to convert between about 50 and 95 or 97 mass percent of the free fatty acids to alkyl esters and provide an esterification effluent comprising alkyl ester of the free fatty acid, free fatty acid, water and unreacted alkanol;
- b. contacting the esterification effluent with sufficient basic glycerin to saponify free fatty acid to form soaps of said free fatty acid and provide an oil phase comprising alkyl ester and a glycerin phase comprising glycerin and soaps; and
- c. phase separating the oil phase and glycerin phase.

In a preferred embodiment using a glyceride-containing feed, the esterification effluent contains between about 0.5 and 3 mass percent free fatty acid. Most preferably, the free fatty acid in the esterification effluent is at least sufficient to neutralize at least about 80 mole percent of the base in the basic glycerin. Advantageously the basic glycerin removes water from the esterification effluent to provide an oil phase containing less than about 0.1 mass percent water.
In a sixth broad aspect of this invention, the esterification is conducted in the presence of acidic catalyst that is soluble in alkanol and substantially insoluble in alkyl ester of fatty acid and the alkanol is provided in an amount sufficient to provide an esterification effluent having an alkanol and catalyst phase and an alkyl ester-containing phase. The alkyl-ester-containing phase is contacted with glycerin to reduce the concentration of alkanol and water in the alkyl ester and form a glycerin-containing phase and an oil phase containing alkyl ester. The glycerin phase is separated and at least a portion is admixed with the alkanol and catalyst phase and alkanol is selectively recovered from the admixture by vapor fractionation.
Processes for esterification of this sixth aspect of the invention for the esterification of feed containing free fatty acid, especially free fatty acid of from about 8 to 30, say, 14 to 24, carbon atoms, comprises:

- a. contacting the feed with a stoichiometric excess of alkanol, preferably with a molar ratio of alkanol to free fatty acid between about 10:1 to 30:1, under acidic esterification conditions including the presence of acid catalyst that is soluble in alkanol and substantially insoluble in alkyl ester of fatty acid, preferably sulfuric acid, to provide an esterification effluent comprising an alkanol and acid catalyst-containing phase and a first oil phase containing alkyl ester of the free fatty acid, water and unreacted alkanol;
- b. phase separating the alkanol and acid catalyst-containing phase and the first oil phase;
- c. contacting the first oil phase with glycerin to provide a second oil phase having a reduced concentration of water and alkanol and a glycerin-containing phase containing water and alkanol;
- d. admixing at least a portion of the glycerin-containing phase with the alkanol and catalyst-containing phase to provide an admixture; and
- e. recovering alkanol from the admixture by vapor fractionation.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic representation of an integrated esterification and transesterification biodiesel facility using the processes of this invention.

FIG. 2 is a schematic representation of a two stage esterification reactor system useful in the processes of this invention.

FIG. 3 is a schematic representation of another integrated esterification and transesterification biodiesel facility using the processes of this invention.

FIG. 4 is a schematic representation of an acid esterification unit operation using an organic soluble catalyst having preferential solubility in an alkanol phase.

FIG. 5 is a schematic representation of an acid esterification unit operation in which glycerin which has been used to recover alkanol from alkyl ester is subjected to vapor fractionation for alkanol recovery.

DETAILED DISCUSSION

Esterification Conditions

The processes of this invention pertain to the acidic esterification of free fatty acids, particularly those containing about 8 to 30, say 14 to 24, carbon atoms, to form alkyl esters of the free fatty acids. The alkyl esters can find various utilities. With emphasis on developing renewable fuels, the demands for biodiesel through the base transesterification of glycerides from plant and animal oils and fats have created a need to convert free fatty acids in those oils and fats to alkyl esters for inclusion in the biodiesel product.
The feed for the esterification may be an oil phase comprising substantially all free fatty acid or a composition containing free fatty acid and other components including, but not limited to, triglycerides. In general, the oil phase may contain from about 1 or 2 to essentially 100, mass percent free fatty acid. Where the free fatty acid is derived from an oil or fat-containing component, the feed may be subjected to unit operations to selectively remove the free fatty acid which then is the feed to the esterification processes. Such removal may be effected in any suitable manner as is known in the art. For instance, the oil or fat may be contacted with base to saponify the free fatty acid which can then be removed by phase separation or extraction, e.g., with glycerin, ethylene glycol, or the like. Alternatively, the oil or fat may be used as the feed to the esterification. In the latter circumstances, the free fatty acid may comprise up to about 60 mass percent (dry basis) of the oil or fat depending upon the specific oil or fat. Feeds may also contain phospholipids which may be as much as about 2 to 5 mass percent (dry basis) of the feeds. The balance of the fats and oils is largely fatty acid triglycerides. The unsaturation of the free fatty acids and triglycerides may also vary over a wide range.
Examples of oils or fats derived from bio sources, especially vegetable oils and animal fats, include, but are not limited to rape seed oil, soybean oil, cotton seed oil, safflower seed oil, castor bean oil, olive oil, coconut oil, palm oil, corn oil, canola oil, jatropha oil, rice bran oil, tobacco seed oil, fats and oils from animals, including from rendering plants and fish oils. Mixtures of two or more oils and fats can be used.
Esterification conditions include the presence of alkanol, elevated temperature and the presence of acid catalyst. Broadly, esterification is conducted with alkanol, which may be a diol, but preferably is a monoalkanol, having a primary —OH, under esterification conditions. The preferred alkanols are lower alkanols, especially those having 1 to 3 carbon atoms, although butanol and isobutanol and higher alkanols are operable. Most preferably the alkanol is methanol which has the highest reactivity. Ethanol can be used but may pose separation difficulties if the esterification product is used to make methyl biodiesel.
The molar ratio of alkanol to free fatty acid can vary widely. As the acidic esterification is an equilibrium-limited reaction, a stoichiometric excess of alkanol is typically used. Where esterification is sought, the molar ratio of alkanol to free fatty acid is generally between about 0:5:1 to 30:1, and preferably between about 2:1 to 25:1, and most preferably between about 3:1 to 20:1. Preferably the alkanol is methanol and is present in an amount that exceeds the solubility of methanol in the oil phase and thus forms a separate phase in the reaction zone. A portion of water present in the reaction menstruum can partition to the methanol-containing phase. The alkanol also aids in increasing the minimum concentration of water co-product and glycerin from any transesterification that forms a separate water or glycerin-containing phase. However, advantageous operation includes the use of sufficient alkanol to form a separate phase containing both alkanol and alkanol-soluble catalyst.
Esterification conditions include the presence of acidic catalyst and elevated temperature, e.g., generally between about 30° C. and 200° C. High temperatures are often unnecessary to achieve high conversions and thus temperatures in the range of about 30° C. or 40° C. to 150° C., and sometimes, 60° C. to 100° C. or 120° C., provide sufficient conversions of fatty acids with relatively short residence times. Preferred esterification temperatures are below about 120° C., to attenuate the reaction rate of water with ester. The desired esterification temperatures will depend in part, upon the other acidic esterification conditions including the strength of the acid catalyst and its concentration. The reaction pressure can be any suitable pressure, e.g., from about 10 to 5000, preferably from about 90 to 1000, kPa absolute. Preferably an inerting gas such as nitrogen, hydrocarbon gas such as methane or carbon dioxide is used during the esterification.
The pressure for the acidic esterification is preferably sufficient to maintain a liquid phase. The esterification may be conducted under conditions such that the alkanol and water are removed by vaporization or may be under conditions such that the reaction occurs in the liquid phase.
The catalyst can be heterogeneous or homogeneous. Where heterogeneous, it may be a solid or a highly dispersed liquid phase. Any suitable acid catalyst (Bronsted acid or Lewis acid) for the esterification of free fatty acids can be used including homogeneous and heterogeneous catalysts. The preferred acid catalysts are mineral acids such as hydrochloric acid, sulfurous acid, sulfuric acid, phosphoric acid, and phosphorous acid. However other strong acids including organic and inorganic acids can be used. Examples of strong organic acids include alkyl sulfonic acids such as methylsulfonic acid; alkylbenzene sulfonic acids such as toluene sulfonic acid; naphthalenesulfonic acid; and trichloroacetic acid. Solid acid catalysts include NAFION® resins. Sulfuric acid and phosphoric acid are preferred due to non-volatility and low cost with sulfuric acid being most often used due to its availability and strong acidity. Sulfuric acid may be provided in any suitable grade including, but not limited to highly concentrated, e.g., 98 percent, sulfuric acid, or in concentrated aqueous solutions, e.g., at least 30 percent, sulfuric acid.
The amount of acid catalyst provided can vary over a wide range. Typically the catalyst is provided in a catalytically effective amount of at least about 0.1 mass percent based upon the feed. Where soaps are present, the amount of acid should be sufficient to convert such soaps to free fatty acids. Often the acid is present in an amount of at least about 0.2 to 5, say, 0.25 to 2, mass percent based upon the feed above that required to convert any soaps to free fatty acids. Solid heterogeneous catalysts are typically provided in greater amounts. Oil soluble catalysts tend to be more active which is believed to be due to the dispersion of the catalyst in the oil phase. Preferred oil soluble, acid catalysts are those having organic substituents of at least about 4 carbon atoms, e.g., from 6 to 24 carbon atoms, especially sulfonic acids such as toluene sulfonic acid and naphthalene sulfonic acid.
The residence time for the esterification will depend upon the amount of free fatty acid present, the conversion sought, the type and amount of catalyst used, the reactivity and amount of alkanol as well as the temperature of the process, and the type of reactor and extent of mixing. Residence times thus can range from less than 1 minute to over 1000 minutes. The residence times frequently are in the range of about 5 minutes to 120 minutes, preferably in the range of about 10 minutes to 90 minutes. Often, the reactivity of alkanol and the residence time is sufficient to convert at least about 30 mole percent, and preferably at least about 50 mole percent, and sometimes at least about 75 mole percent to essentially all, preferably between about 75 and 95 or 98 or even 99, mass percent of the free fatty acid to ester.
The esterification may be conducted in one or more stages. If desired, the effluent from one reaction stage may be subjected to a unit operation to remove water.

The Drawings

Processes for making biodiesel will be further described in connection with FIG. 1 which schematically depicts biodiesel manufacturing facility 100. Facility 100 is provided with a transesterification component (generally designated by numerals in the 200 series) as well as pretreatment components (generally designated by numerals in the 100 series) and a refining component generally (designated by numerals in the 300 series).
Pretreatment by Esterification
As shown in FIG. 1, a glyceride feed containing free fatty acid can be provided to facility 100 via line 102 for pretreatment by acid. Line 104 is provided in the event that more than one feed is desired to be processed simultaneously in the esterification section. Catalyst, which for purposes of this discussion, is sulfuric acid, is provided via line 114.
The feed may be directly introduced into esterification reactor 106, or as shown, is subjected to a contact with an alkanol laden stream of glycerin to strip alkanol from the glycerin into the oil-containing feed phase. This contact will be described later.
The preferred conditions for the esterification will depend upon the nature of the feed and the apparatus type and configuration. Reactor 106 may comprise one or more stages or vessels and separation unit operations may be located between each stage or vessel. Where reactor 106 is staged, it is often desirable, but not essential, to remove water between stages to enhance conversion of free fatty acid to esters. Reactor 106 may be a vessel or a length of pipe. But preferably other types of vessels are used such as mechanical and sonically agitated reactors, and reactors with static mixing such as reactors containing contact structures such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, and other impingement structures. Suitable reactors include those providing high intensity mixing, including high shear.
The oil phase from the esterification section of facility 100 often contains at least about 0.5, say between about 0.5 and 2 or 3, mass percent free fatty acid. This free fatty acid serves to neutralize at least a portion of the base catalyst contained in a spent glycerin stream produced in the transesterification and base pretreatment sections of facility 100. Preferably, the molar ratio of free fatty acid in the oil phase from the esterification to mole of base in the glycerin phase introduced into base reactor 134 as discussed below will be at least about 0.3:1, often at least about 0.7:1 up to about 1:1. The use of ratios of free fatty acid to base catalyst of greater than 1:1 can adversely affect the performance of the base pretreatment. A number of advantages flow from this preferred embodiment. For instance, the equipment and conditions required for the esterification section need not be of the type required for essentially complete conversion of the free fatty acids, resulting in capital and operating cost savings. Since residual free fatty acid is converted to soap and removed in the base pretreatment section, the feed to the transesterification section can be substantially devoid of free fatty acid which adversely affects the base catalyst therein. Additionally, the neutralized spent glycerin stream from the base pretreatment section can be used effectively for enhancing phase separation and water and catalyst removal from the esterification product.
During the esterification in reactor 106 some conversion of glycerides to esters may occur. The esters, diglycerides and monoglycerides essentially remain in the oil phase. Some glycerin will be produced as a result of the transesterification of the glycerides in the feed. The extent of such conversion is not critical but results in reduced requirements of alkanol and catalyst in the transesterification section per unit of biodiesel produced as well as enabling increased performance such as rate of conversion and extent of conversion to be obtained. Generally up to about 20 mass percent, say, between about 0.1 to 15, and sometimes between 5 to 10, mass percent of the glyceride-containing feed is transesterified during acid esterification.
The esterification reaction product from reactor 106 is passed via line 108 to phase separator 110. Phase separator 110 is optional depending upon whether or not two phases exist. In some instances, an oil layer containing glycerides and fatty ester and a water-containing layer form. The water-containing layer can contain more polar components such as glycerin, water-soluble catalyst, and alkanol. As shown, a neutralized spent glycerin stream from the base pretreatment section is provided via line 170A and contacted with the esterification product. The spent glycerin aids in the extraction of water and water-soluble phosphorus compounds. Additionally, the glycerin assists in making the phase separation. In this embodiment, the amount of glycerin added can vary widely. As relatively small amounts of water are produced during the acid esterification of free fatty acids, beneficial results can be achieved with relatively little spent glycerin being added. Often the spent glycerin added is less than about 20, preferably between about 0.5 and 10, mass percent of the stream from esterification reactor 106. A separate phase may exist in reactor 106, e.g., from catalyst such as sulfuric acid, or water co-produced during the esterification or even alkanol above that miscible with the oil phase. Glycerin can aid in forming a defined phase containing, e.g., catalyst and water. As used herein, the formation of a glycerin phase or providing a glycerin phase contemplates that there may, or may not, be separate phases in the fluid contacted with glycerin. Spent glycerin that is in a separate phase may be separated and removed via line 112.
Phase separator 110 may be of any suitable design including a decanter, a phase separation facilitated decanter that contains coalescing sites, and a centrifuge. The lower, water-containing fraction exits separator 110 via line 112. This fraction contains some alkanol, water, water-soluble catalyst, glycerin and water-soluble phosphorus compounds.
The oil fraction of separator 110 contains virtually no sulfuric acid, often some alkanol, relatively little water, unreacted free fatty acids, if any, fatty ester and glycerides. The fraction is passed via line 118 from separator 110 to fractionation column 120 to provide an overhead fraction containing alkanol and a bottoms stream containing oil. The overhead from column 120 can be recycled to esterification reactor 106 via line 122. Make up alkanol is provided via line 124.
The fractionation column may be of any suitable design including a flash column, stripping column, falling film evaporator, or trayed or packed column. If desired, more than one fractionation column can be used with one effecting separation of water from alkanol. Similarly a side draw 116 may be taken from distillation column 120 for the removal of water, and fractionation column may be a divided wall column to enhance such separation. In an embodiment, a substantial portion of the water is removed by the phase separation in phase separator 110, and fractionation column does not separately recover water. Water will be contained in both the overhead and bottoms stream from column 120. However, the relatively small amount of water in the overhead can be recycled with alkanol via line 122 to reactor 106 without undue adverse effect. Water contained in the bottoms passes to the base pretreatment section and is removed from the oil phase therein.
In another embodiment, only a portion of the alkanol is removed by fractionation in column 120. The alkanol remaining in the oil phase is passed to the base pretreatment section. In the base pretreatment section alkanol can be reacted with glyceride to form esters and can be recovered in the spent glycerin phase for recycle to the esterification section. Thus, the capital and operating costs for fractionation column 120 can be reduced. Often the bottoms stream from fractionation column 120 contains between about 0.1 to 10, say, between about 0.5 and 5, e.g., 0.5 to 2, mass percent alkanol. In yet another embodiment, the oil-containing fraction from separator 110 can be passed directly to separator 128 or base reactor 134.
While shown as processing the oil phase from separator 110, fractionation column 120 may be positioned between esterification reactor 106 and separator 110 and serve to recover alkanol from the esterification product exiting reactor 106.
Pretreatment by Base
The base pretreatment uses glycerin produced in facility 100 to treat feed. The base pretreatment serves to recover alkanol contained in the glycerin phase from the transesterification section. Hence, the spent glycerin from the base pretreatment section may contain relatively little alkanol. Base pretreatment also serves to partially convert glycerides in the feed to fatty acid esters and mono- and di-glycerides. Thus, the amount of alkanol required to transesterify the pretreated feed will be less than had no base pretreatment occurred. Base pretreatment can also serve to remove phospholipids as glycerin-soluble components. Base pretreatment further removes free fatty acids from the glyceride-containing feed by saponification to glycerin-soluble soaps. Removal of the phospholipids and free fatty acids facilitates processing during transesterification and minimizes catalyst loss during transesterification cased by saponification of free fatty acids with base catalyst. Phospholipids, for instance, tend to make more difficult phase separations of oil and glycerin in the transesterification component. And biodiesel must meet stringent phosphorus specifications. See, for instance, ASTM D 6751, American Society for Testing and Materials.
As shown in the facility of FIG. 1, a glyceride-containing feed stream is provided by line 132 to base reactor 134. The feed stream may comprise a fresh glyceride-containing feed. Alternatively or in addition, the feed stream may comprise the oil phase from the esterification provided via lines 126 and 130. To base reactor is also provided a glycerin and base catalyst-containing stream via line 142 which will be further discussed below. Preferably a non-acidic inerting gas such as nitrogen or hydrocarbon gas such as methane is used during base pretreatment.
In base reactor 134, free fatty acids contained in the feed stream are reacted with base catalyst to form soaps. If the free fatty acid content of the feed stream requires more than the amount of base catalyst introduced via line 142 for the desired degree of saponification, additional base can be added via line 133. The additional base may be the same or different from that comprising the catalyst, and may be one or more of alkali metal hydroxides or alkoxides and alkaline earth metal hydroxides, oxides or alkoxides, including by way of examples and not in limitation, sodium hydroxide, sodium methoxide, potassium hydroxide, potassium methoxide, calcium hydroxide, calcium oxide and calcium methoxide.
To the extent that phospholipids are present in the feed stream to base reactor 134, at least a portion is chemically reacted, e.g., by a hydration or by a salt formation, to provide chemical compounds preferentially soluble in glycerin.
Base reactor 134 is maintained under base reaction conditions, which for free fatty acid-containing feed streams is that sufficient to react basic catalyst and free fatty acids to soaps and water, and for phospholipids-containing feed streams is that sufficient to react basic catalyst and phospholipids to chemical compounds preferentially soluble in a glycerin phase. Typical base reaction conditions include a temperature of at least about 10° C., say, 35° C. to 150° C., and most frequently between about 40° C. and 80° C. Pressure is not critical and subatmospheric, atmospheric and super atmospheric pressures may be used, e.g., between about 1 and 5000, preferably from about 90 to 1000, kPa absolute. The residence time is sufficient to provide the sought degree of saponification of fatty free acids and reaction of phospholipids. The residence time in base reactor 134 may range from about 1 minute to 10 hours.
Base reactor 134 may be of any suitable design. Reactor 134 may be a vessel or a length of pipe. But preferably other types of vessels are used such as mechanical and sonically agitated reactors, and reactors with static mixing such as reactors containing contact structures such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, and other impingement structures. Suitable reactors include those providing high intensity mixing, including high shear.
The base reaction product from reactor 134 contains glycerin, glycerides, soaps, water, and fatty acid ester and is passed via line 136 to separator 128. Separator 128 serves to separate the less dense oil layer from the more dense glycerin layer. The soaps and reacted phospholipids preferentially pass to the glycerin layer as does most of the water. The oil layer preferably contains less than about 0.5 mass percent soaps. Phase separator 128 may be of any suitable design including a decanter, a phase separation facilitated decanter that contains coalescing sites, and, if needed, a centrifuge.
The glycerin phase is withdrawn from separator 128 via line 137 and may be sent to glycerin recovery or another application. If the glycerin layer contains significant amounts of soaps, it may be desirable to recycle the soaps to esterification reactor 106 for conversion to fatty esters. As shown, a portion or all of the glycerin phase may be passed via line 170 to acidification reactor 172 where soaps are converted to free fatty acids. At least a portion of this glycerin phase is passed via line 170A to provide the glycerin to assist in the separation of water, water-soluble catalyst (or salts thereof) from the esterification product in phase separator 110. The glycerin-containing phase from separator 110 is passed via line 112 to line 170. Also as shown, a portion of the glycerin phase in line 172 is recycled to reactor 134 via line 170B. The recycle can serve several purposes. For instance, hydrated phospholipids are returned to reactor 134 where they may undergo transesterification to recover additional fatty acid ester. Also, any base contained in the recycled glycerin stream is available for saponification of free fatty acids.
Unless acid contained in the esterification effluent of line 108 is neutralized prior to being passed to separator 110, the glycerin-containing phase from separator 110 will contain water-soluble acid which can be used as acid for acidification reactor 172. Acid can also be provided via line 174. Acidification reactor 174 may be one or more vessels of any suitable design including a length of pipe and other types of vessels such as mechanical and sonically agitated reactors, and reactors with static mixing such as reactors containing contact structures such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, and other impingement structures. Suitable reactors include those providing high intensity mixing, including high shear. The acidification conditions usually encompass a temperature in the range of about 20° C. to 150° C., a pressure from about 1 to 5000, preferably 90 to 1000, kPa absolute, and a residence time of from about 1 second to 5 hours. Suitable acids include mineral acids and organic acids, but typically a readily available acid such as sulfuric or phosphoric acid is used. The amount of acid is usually sufficient to convert substantially all the soaps to free fatty acid. The use of excess acid is not deleterious to the formation of the free fatty acids, but can entail additional expense. Accordingly the molar ratio of acidifying acid function to soaps is in the range of about 1:1 to 1.5:1. Generally the pH of the glycerin stream is less than about 6, say, between about 1 and 5, e.g., 2 and 4. The acidity of the glycerin stream is determined by diluting the glycerin stream to 50 volume percent water and measuring the pH.
The glycerin stream from acidification reactor is passed via line 176 to contact vessel 178 into which glyceride-containing feed is provided via line 102. In contact vessel 178 the glycerin stream is contacted with fresh feed which serves to extract a portion of the alkanol from the glycerin phase. The contact with the glycerin also serves to remove water from the feed. Removal of water assists in the esterification of free fatty acids in esterification reactor 106 as the esterification is an equilibrium-limited reaction affected by water concentration.
Contact vessel 178 may be of any suitable design including a length of pipe and other types of vessels such as mechanical and sonically agitated reactors, and reactors with static mixing such as reactors containing contact structures such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, and other impingement structures. Suitable reactors include those providing high intensity mixing, including high shear. The contact conditions usually encompass a temperature in the range of about 20° C. to 150° C., a pressure from about 1 to 5000 kPa absolute, and a residence time of from about 1 second to 5 hours. Often at least about 50 mass percent of the alkanol in the glycerin stream passes to the oil phase as do essentially all of the free fatty acids. The amount of alkanol recovered from the glycerin will depend upon the alkanol content of the glycerin, the ratio of glycerin to fresh feed, and the contacting conditions. Frequently the mass ratio of glycerin to oil is in the range of between about 1:5 to 1:20, say 1:8 to 1:15, and at least about 30, and sometimes between about 50 and 99, mass percent of the alkanol in the glycerin phase passes to the oil phase.
The ability to recover alkanol from glycerin by extraction with fresh feed can effectively be used to use glycerin as a complementary means for recycling unreacted alkanol to reactor 106. FIG. 1 shows two glycerin loops for alkanol recovery and recycle to the esterification reactor. The first loop involves the glycerin layer from separator 110 and the second, the glycerin layer from separator 128.
The fluid mixture from contact vessel 178 is passed via line 180 to phase separator 182. In phase separator 182, a glyceride and free fatty acid oil layer is produced that is passed via line 184 to esterification reactor 106. A glycerin-containing layer is discharged via line 186 and contains water, acidification acid, and soluble phosphorus compound. Separator 182 may be of any suitable design including a decanter, a phase separation facilitated decanter that contains coalescing sites, and, if necessary, a centrifuge. Contact vessel 178 and decanter 182 may be a single vessel, including but not limited to, a countercurrent extraction column.
If the esterification product from esterification reactor 106 has a sufficiently low free fatty acid content and low phospholipids content, another option is to eliminate separator 110 and fractionation column 120 and provide the esterification product in line 108 directly to separator 128 or base reactor 134.
Returning to separator 128, the oil phase is withdrawn and passed via line 138 to second pretreatment reactor 139. Second pretreatment reactor 139 and third pretreatment reactor 148 are adapted to recover alkanol contained in the glycerin from the transesterification component of facility 100 through reaction, e.g., transesterification and extraction into the glyceride-containing phase. A base transesterification process is used in these pretreatment reactors. While two reactors are shown, the number of reactors will depend upon the sought consumption of the alkanol as well as the efficiency of the reactors. Hence one, two, or three or more pretreatment reactors may be used. Also, the pretreatment reactor can comprise a number of stages in a single vessel which could be a countercurrent contact vessel. Advantageously, the feed stream to the alkanol consumption pretreatment reactors is relatively free from free fatty acids so as to prevent undue consumption of the base catalyst. Typically the pretreatment reactors provide a glycerin stream from which most of the alkanol has been removed. Often, the alkanol content of the glycerin discharged from base reactor 134 is less than about 5, and preferably less than about 2, mass percent.
In an alternative mode of operation, a significant portion of the alkanol is contained in line 126 (or line 108 if separator 110 and distillation column 120 are not used) and passed to separator 128. The concentration of alkanol in the glycerin-containing stream in line 170 may be higher than 5 mass percent, and alkanol is recovered be partitioning to the glyceride-containing feed in contact vessel 178. The alkanol content of the glycerin may be sufficiently low that no distillation is required to recover alkanol yet the overall process to make biodiesel can still exhibit high efficiencies.
Second pretreatment reactor 139 also receives the glycerin phase from the third pretreatment reactor. This glycerin phase contains glycerin, base catalyst, and alkanol. Second pretreatment reactor 139 is maintained under base transesterification conditions including the presence of base catalyst provided by the glycerin phase feed and elevated temperatures, often between about 30° C. and 220° C., preferably between about 30° C. and 80° C. to provide a second pretreatment product. The pressure is typically in the range of between about 90 to 1000 kPa (absolute) although higher and lower pressures can be used. The reactor is typically batch, semi-batch, plug flow or continuous flow tank. Preferably other types of vessels are used such as mechanical and sonically agitated reactors, and reactors with static mixing such as reactors containing contact structures such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, and other impingement structures. Suitable reactors include those providing high intensity mixing, including high shear. However, depending upon the presence of soaps and phospholipids, care needs to be taken so as not to generate a product that cannot be readily separated by phase separation. The residence time will depend upon the desired degree of conversion of the contained alkanol, the ratio of alkanol to glyceride, reaction temperature, the degree of agitation and the like, and is often in the range of about 0.1 to 20, say, 0.5 to 10, hours.
The second pretreatment product contains glycerides, fatty esters, base catalyst and glycerin, and it has a reduced concentration of alkanol. The second pretreatment product is passed from second pretreatment reactor 139 via line 141 to separator 140. Separator 140 may be of any suitable design including a decanter, a phase separation facilitated decanter that contains coalescing sites, and, optionally, a centrifuge. The lower, glycerin-containing phase from separator 140 contains relatively little alkanol, preferably less than about 10 mass percent, and contains base catalyst, and is passed via line 142 to base reactor 134 where catalyst reacts with free fatty acids to form soaps which can then be removed from the glyceride-containing feed.
As depicted, line 142 is provided with holding tank 142A. Holding tank 142A can serve as a reservoir and enables the rate that glycerin, which contains base, is provided to base reactor 134, to be varied with changes in free fatty acid content of the esterification product. It also can permit additional reaction of glycerides with alkanol contained in the glycerin phase to occur prior to introduction into base reactor 134 where catalyst is consumed by conversion of free fatty acids to soaps.
The upper oil phase is removed from separator 140 via line 144 and is passed to line 146 which also receives the glycerin co-product from transesterification from line 248. The combined streams are passed to third pretreatment reactor 148. The stream is provided by line 146 and contains in addition to glycerin, alkanol, base catalyst, and usually some water and soaps. Table I sets forth typical compositions of the stream in line 248. The compositions, of course, will depend upon the operation of the transesterification component as well as which of the glycerin-containing streams from the transesterification component are used. The typical concentrations are based upon combining all glycerin-containing streams and operating under preferred parameters.

TABLE I

Component	Broad, Mass %	Typical, Mass %

Glycerin	40 to 80	50 to 70
Alkanol (Methanol)	15 to 50	25 to 45
Catalyst (NaOCH₃)	0.2 to 5	0.5 to 5
Soaps	0.1 to 5	0.5 to 5
Water	0.01 to 0.5	0.05 to 0.3
Oil (glycerides and alkyl	0 to 5	0.5 to 2
esters)

Third pretreatment reactor 148 is maintained under base transesterification conditions including the presence of base catalyst provided by the glycerin—containing feed and elevated temperatures, often between about 30° C. and 220° C., preferably between about 30° C. and 80° C. to provide a first pretreatment product. Base catalyst in the transesterification component tends to partition to the glycerin phase and often adequate catalyst is provided for the base pretreatment section in the glycerin co-product from the transesterification section provided by line 248. In some instances, however, it may be desired to add additional base catalyst to third pretreatment reactor 148 or any preceding base pretreatment reactor. The pressure is typically in the range of between about 90 to 1000 kPa (absolute) although higher and lower pressures can be used. The reactor is typically batch, semi-batch, plug flow or continuous flow tank with some agitation or mixing. The preferred types of vessels are mechanical and sonically agitated reactors, and reactors with static mixing such as reactors containing contact structures such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, and other impingement structures. Suitable reactors include those providing high intensity mixing, including high shear. However, depending upon the presence of soaps and phospholipids, care needs to be taken so as not to generate a product that cannot be readily separated by phase separation. The residence time will depend upon the desired degree of conversion, the ratio of alkanol to glyceride, reaction temperature, the degree of agitation and the like, and is often in the range of about 0.1 to 20, say, 0.5 to 10, hours.
Typically the transesterification in third pretreatment reactor 148 recovers through transesterification and extraction to the glyceride-containing phase at least about 20, preferably at least about 30, and more preferably at least about 50, mass percent of the alkanol fed to the reactor. Any unreacted alkanol in the oil phase will be carried with the oil phase to the transesterification component of facility 100. Often the total amount of alkanol recovered from the glycerin-coproduct from transesterification using all pretreatment stages is at least about 50, and sometimes at least about 80, mass percent.
The third pretreatment product passes from third pretreatment reactor 148 through line 150 to separator 152. Separator 152 may be of any suitable design including a decanter, a phase separation facilitated decanter that contains coalescing sites, and, optionally, a centrifuge. Separator 152 serves to separate an oil phase containing glycerides, esters and alkanol and some catalyst, from a glycerin-containing phase containing glycerin, reduced concentration of alkanol, and catalyst. The glycerin-containing phase frequently contains less than about 15 mass percent alkanol. The glycerin-containing phase from separator 152 is passed via line 154 to second pretreatment reactor 139.
Facility 100 includes a chiller 158 to remove high molecular weight glycerides, waxes and esters that are insoluble at the chiller temperature. Some feeds, such as crude corn oil, contain high molecular weight glycerides and esters. The hydrocarbyl moieties in these high molecular weight components typically have between 30 and 40 carbon atoms. If they remain in the oil, the resultant biodiesel product tends to have unacceptably high cloud points and gel points. As shown, the oil phase from separator 152 passes through line 156 to chiller 158. Chiller 158 is maintained at a temperature sufficient to cause high molecular weight and other components that lead to and increase in gel point temperature to solidify. Typically this temperature is between about 0° C. and 20° C. In some instances, cooling will tend to remove monoglycerides and diglycerides. Cooling below the desired temperature and then warming to a temperature to liquefy the mono- and di-glycerides while still maintaining a solid wax, can minimize loss of components that can be converted to biodiesel. The chilled oil phase is then passed via line 160 to centrifuge 162 to remove higher density components including solids and any remaining glycerin phase. The higher density fraction is discharged via line 164. Rather than using a centrifuge, the solids can be filtered from the glyceride-containing stream. Filter aids can be used if desired. A producer composition is provided by centrifuge 162 and is provided to line 166.
Chiller 158 is optional, and chillers may also be used elsewhere in facility 100 to remove waxes. For instance, a chiller may be used to treat fresh feed in line 102 or can be used to treat biodiesel product from the refining component.
If desired all or a portion of the producer composition in line 166 may be withdrawn via line 168 as an intermediate product for storage or sale as a feedstock for transesterification. Line 168 also provides the feed for the transesterification component of facility 100 by introducing the producer composition into line 200.
Transesterification
Line 200 provides glyceride-containing feed to first transesterification reactor 202. Line 200 can also supply additional glyceride-containing feed. Preferably the additional feed is relatively free of free fatty acids and phospholipids such as refined oils sourced from rape seed, soybean, cotton seed, safflower seed, castor bean, olive, coconut, palm, corn, canola, fats and oils from animals, including from rendering plants and fish oils.
Alkanol for the transesterification is supplied to first transesterification reactor via line 206. The alkanol is preferably lower alkanol, preferably methanol, ethanol or isopropanol with methanol being the most preferred. The alkanol may be the same or different from the alkanol provided to esterification reactor 106 via line 124. Although line 206 is depicted as introducing alkanol into line 200, it is also contemplated that alkanol can be added directly to reactor 202 at one or more points. Generally the total alkanol (line 206 and from the producer composition of line 166) is in excess of that required to cause the sought degree of transesterification in reactor 202. Preferably, the amount of alkanol is from about 101 to 500, more preferably, from about 110 to 250, mass percent of that required for the sought degree of transesterification in reactor 202. In facility 100 three reactors are depicted as being used. One reactor may be used, but since the reaction is equilibrium limited, most often at least two and preferably three reactors are used. Often, where more than one reactor is used, at least about 60, preferably between about 70 and 96, percent of the glycerides in the feed are reacted in first transesterification reactor 202. It is possible to provide all the alkanol required for transesterification to first transesterification reactor 202, or a portion of the alkanol can be provided to each of the transesterification reactors.
The base catalyst is shown as being introduced via line 204 to first transesterification reactor 202. The amount of catalyst used is that which provides a desired reaction rate to achieve the sought degree of transesterification in first transesterification reactor 202. Preferably, catalyst is provided to each of the transesterification reactors since base catalyst preferentially partitions to the glycerin phase and is removed with phase separation of the glycerin after each transesterification reactor. The amount of catalyst used will be in excess of that required to react with the amount of free fatty acid contained in the feed oil, which due to the pretreatment, will be relatively little. The base catalyst may be an alkali or alkaline earth metal hydroxide or alkali or alkaline earth metal alkoxide, especially an alkoxide corresponding to the lower alkanol reactant. Preferred alkali metals are sodium and potassium. When the base is added as a hydroxide, it may react with the lower alkanol to form an alkoxide with the generation of water which in turn results in the formation of free fatty acid. Another type of catalyst is an alkali metal or alkaline earth metal glycerate. This catalyst converts to the corresponding alkoxide of the alkanol reactant in the reaction menstruum. Alternatively, the catalyst may be a heterogeneous base catalyst. Catalyst may need to be separately provided to the base pretreatment reactors if the base catalyst, e.g., a heterogeneous or oil soluble catalyst, is not carried with the co-product glycerin in the transesterification component to the base pretreatment reactors. However, homogeneous catalysts that have solubility in glycerin are preferred where the pretreatment component is used since the catalyst serves as at least a portion of the base used therein. The exact form of the catalyst is not critical to the understanding and practice of this invention. For the purposes of the following discussion, homogenous base catalyst is used. Preferably a non-acidic inerting gas such as nitrogen or hydrocarbon gas such as methane is used during base transesterification.
Often the transesterification is at a temperature between about 30° C. and 220° C., preferably between about 30° C. and 80° C. The pressure is preferably sufficient to maintain a liquid phase reaction menstruum and typically is in the range of between about 90 to 1000 kPa (absolute) although higher and lower pressures can be used. First transesterification reactor 202 is typically batch, semi-batch, plug flow or continuous flow tank with some agitation or mixing. Preferably the reactors are mechanical and sonically agitated reactors. Reactors with static mixing such as reactors containing contact structures such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, and other impingement structures can be used. Suitable reactors include those providing high intensity mixing, including high shear. As stated above, one of the advantages of the processes of this invention is that the producer compositions do not require an induction period for the transesterification reaction to initiate. Accordingly plug flow reactors have enhanced viability. The residence time will depend upon the desired degree of conversion, the ratio of alkanol to glyceride, reaction temperature, the base catalyst concentration, the degree of agitation and the like, and is often in the range of about 0.02 to 20, say, 0.1 to 10, hours.
The partially transesterified effluent from reactor 202 is passed via line 208 to phase separator 210. Phase separator 210 may be of any suitable design including a decanter, a phase separation facilitated decanter that contains coalescing sites, and, optionally, a centrifuge. A glycerin-containing bottoms phase is provided in the separator and is removed via line 212 and is passed to glycerin header 214. As depicted, this stream is used as a portion of the glycerin for the pretreatment component of facility 100. This glycerin phase also contains any soaps made in reactor 202 and a portion of the catalyst. The soaps can be recovered from this stream in acidifying reactor 172 as discussed above. The lighter phase contains alkyl esters and unreacted glycerides and is passed via line 216 to second transesterification reactor 218. A rag layer may form in separator 210. The rag layer may contain unreacted glycerides, alkyl esters, alkanol, soaps, catalyst and glycerin. An advantage of the process set forth in FIG. 1 is that withdrawing the rag layer with the glycerin phase does not result in a loss of glycerides, alkyl esters, alkanol, and catalyst since the glycerin phase is passed to the pretreatment component of facility 100.
Reactor 218 may be of any suitable design and may be similar to or different than reactor 202. As shown, additional alkanol is provided via line 206A, and additional catalyst is provided via line 204A. Preferably the transesterification conditions in reactor 218 are sufficient to react at least about 90, more preferably at least about 95, and sometimes at least about 97 to 99.9 or more, mass percent of the glycerides in the feed to the transesterification. The transesterification in reactor 218 is typically operated under conditions within the parameters set forth for reactor 202 although the conditions may be the same or different. The residence time will depend upon the desired degree of conversion, the ratio of alkanol to glyceride, reaction temperature, the degree of agitation and the like, and is often in the range of about 0.02 to 20, say, 0.1 to 10, hours.
The effluent from second transesterification reactor 218 is passed via line 220 to phase separator 222 which may be of any suitable design and may be the same as or different from the design of separator 210. A heavier, glycerin-containing phase is withdrawn via line 224 and passed to glycerin header 214. A lighter phase containing crude biodiesel is withdrawn from separator 222 via line 226.
As depicted, third transesterification reactor 228 is used and the crude biodiesel in line 226 is passed to this reactor. The transesterification conditions in reactor 228 are sufficient to provide essentially complete conversion, at least about 97 or 98 to 99.9, mass percent of the glycerides in the feed converted to alkyl ester. As shown, additional alkanol is provided via line 206B, and additional catalyst is provided via line 204B. The transesterification in reactor 228 is typically operated under conditions within the parameters set forth for reactor 202 although the conditions may be the same or different. The residence time will depend upon the desired degree of conversion. The reactor may be of the type described for reactor 202. The residence time will depend upon the desired degree of conversion, the ratio of alkanol to glyceride, reaction temperature, the degree of agitation and the like, and is often in the range of about 0.02 to 20, say, 0.1 to 10, hours. Advantageously, the transesterification product from third transesterification reactor 228 contains less than about 1, preferably less than about 0.8, and most preferably less than 0.5, mass percent soaps based upon the total mass of alkyl esters and soaps. The lighter phase also contains alkanol. In reactor 228 the reaction proceeds quickly to completion by the addition of additional alkanol and catalyst, and can be conveniently accomplished by a plug flow reactor.
The overall molar ratio of alkanol to glycerides in the feed to the reactors in the transesterification component, i.e., alkanol provided by lines 206, 206A and 206B, can vary over a wide range. Since transesterification is an equilibrium-limited reaction, the driving force toward the alkyl ester and the conversion of glycerides will be dependent upon the molar ratio of alkanol equivalents to glycerides. Alkanol equivalents are alkanol and alkyl group of the alkyl esters in the feed to the transesterification component. On the basis of transesterfiable substituents in the feed to the transesterification component, the mole ratio of alkanol equivalents to glyceride in the feed to the pretreatment component is frequently between about 3.05:1 to 15:1, say 4:1 to 9:1. Advantageously, the pretreatment processes of this invention permit the reuse of alkanol partitioned to the co-product glycerin without intermediate vaporization. Often the amount of total catalyst provided based on the mass of feed to the first transesterification reactor, i.e., the catalyst provided by lines 204, 204A and 204B, is between about 0.3 and 1 mass percent (calculated on the mass of sodium methoxide).
The effluent from third transesterification reactor 228 is passed via line 230 to phase separator 232 which may be of any suitable design and may be the same as or different from the design of separator 210. A heavier, glycerin-containing phase is withdrawn via line 234 and passed to glycerin header 214. A lighter phase containing crude biodiesel is withdrawn from separator 232 via line 236. Alternatively, separator 232 can be eliminated provided that in second transesterification reactor 218, the conversion of the glycerides in the feed is at least about 90, preferably 92 to 96 or 98, percent. In some instances, the effluent from reactor 228 may be a single phase containing relatively little glycerin. In some instances it may be advantageous to use a centrifuge to separate the glycerin phase from the oil phase following third transesterification reactor 228.
Facility 100 contains an optional alkanol replacement reactor 238. The alkanol replacement reactor serves to transesterify the alkyl ester with a different alkanol. For purposes of transesterification in reactors 202, 218 and 228, an alkanol such as methanol provides not only attractive reaction rates but also an effluent that is more easily separated than, say, a reaction effluent where ethanol is the alkanol. In some instances it may be desired to provide a biodiesel that contains fatty esters in which the alkyl group of the fatty ester is branched in order to reduce cloud and gel points. The transesterification between, say, a fatty acid methyl ester, and higher molecular weight alkanol results in methanol, rather than glycerin, being formed, and often is more readily accomplished than the transesterification of glyceride with that higher alkanol. The higher alkanols include those having 2 to 8 or more carbon atoms, and are preferably branched primary and secondary alkanols although tertiary alkanols may find application but generally are less reactive. Examples of higher alkanols include propanol, isopropanol, isobutanol, 2,2-dimethylbutan-1-ol, 2,3-dimethylbutan-1-ol, 2-pentanol, and the like. Other alkanols include benzyl alcohol and 2 ethylhexanol.
Where an alkanol replacement operation is desired, it may be located at various points in the process. For instance, the replacement alkanol may be provided via line 206B to reactor 228, or, as shown, it can follow reactor 228. In either case, alkanol replacement transesterification can take advantage of catalyst contained in the transesterification medium. Alternatively, alkanol replacement may be effected on a biodiesel product by adding catalyst. Thus, it can be located elsewhere in the refining component of facility 100 including, but not limited to, treating biodiesel in line 352.
The amount of higher alkanol provided via line 240 to alkanol replacement reactor 238 can vary over a wide range. Typically the molar ratio of higher alkanol to alkyl ester being fed to reactor 238 is less than 0.5:1, e.g., from about 1:100 to 1:5. Often the alkanol replacement transesterification is at a temperature between about 30° C. and 220° C., preferably between about 30° C. and 80° C. The pressure is preferably sufficient to maintain a liquid phase reaction menstruum and typically is in the range of between about 90 to 1000 kPa (absolute) although higher and lower pressures can be used. Alkanol replacement reactor 238 can be batch, semi-batch, plug flow or continuous flow tank with some agitation or mixing, e.g., mechanically stirred, ultrasonic, static mixer containing contact surfaces, e.g., trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, or other impingement structures. High intensity mixing reactors, including high shear reactors, may also be used. Preferred reactors are those in which the alkanol being replaced is continuously removed. For instance, a reactive distillation reactor can be used to continuously remove displaced methanol from a transesterification of methyl ester and isopropanol. As depicted, reactor 238 is a reactive distillation unit and lower alkanol is withdrawn via line 330A and passed to the transesterification reactors. Make-up alkanol is provided via line 332.
Where the alkanol replacement reactor is a batch reactor, driving the replacement reaction to either essentially complete conversion of the higher alkanol or essentially complete conversion of the methyl ester to the higher alkanol ester (depending upon whether the higher alkanol is provided below or at or above the stoichiometric amount required for complete conversion), since the vapor fractionation of methanol can continue until completion. With continuous reactors, having unreacted methanol and higher alkanol in the alkanol replacement product is likely. For purposes of this discussion, a continuous alkanol replacement reactor is used.
Where the base catalyst has been removed from the fatty acid ester of the lower alkanol, for instance, if the alkanol replacement were to be conducted on a refined or partially refined biodiesel, catalyst is provided. Suitable catalyst includes base catalyst such as is used for transesterification. Since a single liquid phase exists during the alkanol replacement unlike transesterification where a glycerin layer forms, heterogeneous catalysts and homogeneous catalysts having limited solubility in the reaction menstruum can be used. Solid catalysts are preferred to minimize or eliminate post treatment of the alkanol replacement product, but good contact with catalyst is desirable to timely achieve sought conversion. Homogeneous transesterification catalysts such as titanium tetra-isopropoxide are also advantageous as they are readily removed.
The residence time will depend upon the desired degree of conversion, the ratio of higher alkanol to alkyl ester, reaction temperature, the degree of agitation and the like, and is often in the range of about 0.02 to 20, say, 0.1 to 10, hours. Preferably at least about 80, and sometimes at least about 90, mass percent of the higher alkanol is reacted.
Refining
A crude biodiesel is withdrawn from reactor 238 via line 300 and is passed to the refining component of facility 100. The crude biodiesel may be contacted with acid to neutralize any catalyst therein and then refined to remove alkanol, soaps, water and glycerin.
In a preferred process, an acid, preferably an organic acid, is provided via line 302 in an amount sufficient to substantially neutralize residual base catalyst contained in the crude biodiesel. Inorganic acids such as sulfuric acid can be used as well as organic acids, particularly those less volatile than the alkanol, and acids that do not themselves or any potential reaction product formed in contact with the crude biodiesel, form azeotropes with the alkanol. Exemplary organic acids include acetic acid, citric acid, oxalic acid, glycolic, lactic, free fatty acid and the like. Generally the amount of catalyst contained in the crude biodiesel is quite small as base catalyst preferentially partitions to the glycerin phase. Accordingly, little acid is required to neutralize sufficient catalyst to enable refining without risk of reversion of alkyl ester. Often the amount of acid used is at least 0.95 times, sometimes between about 1 and 3 times, that required to neutralize the catalyst.
Crude biodiesel is passed via line 300 to an alkanol separation unit operation. As shown, a two stage separation unit is used. A single stage separator can be used if desired. The crude biodiesel in line 300 is passed to first alkanol separator stage 304. Separator 304 is of any convenient design including a stripper, wiped film evaporator, falling film evaporator, solid sorbent, and the like. Preferably the fractionation is by fast, vapor fractionation. Generally for a fast, vapor separation the residence time is less than about one minute, preferably less than about 30 seconds, and sometimes as little as 5 to 25 seconds. Preferably the vapor fractionation conditions comprise a maximum temperature of less than about 200° C., preferably less than about 150° C., and most preferably, when the lower alkanol is methanol, less than about 120° C. Depending upon the alkanol, the lower boiling fractionation may need to be conducted under subatmospheric pressure to maintain desired overhead and maximum temperatures. Where a falling film stripper is used, it may be a concurrent or countercurrent flow stripper. Concurrent strippers are preferred should there be a risk of undue vaporization of alkanol at the point of entry of the crude biodiesel. An inert gas such as nitrogen may be used to assist in removing the alkanol.
The fast fractionation may be effected by any suitable vapor fractionation technique including, but not limited to, distillation, stripping, wiped film evaporation, and falling film evaporation. Often the falling film evaporator has a tube length of at least about 1 meter, say, between about 1.5 and 5 meters, and an average tube diameter of between about 2 and 10 centimeters. Usually the vapor fractionation recovers at least about 70, preferably at least about 90, mass percent of the alkanol contained in the crude biodiesel. Any residual alkanol is substantially removed in any subsequent water washing of the crude biodiesel. Advantageously, the amount of alkanol contained in the spent water from the washing may be at a sufficiently low concentration that the water can be disposed without further treatment. However, from a process efficiency standpoint, alkanol can be recovered from the spent wash water for recycle to the transesterification reactors.
The lower boiling fraction containing the alkanol will contain a portion of any water contained in the crude biodiesel. Since the transesterification is conducted with little water being present, and a portion of the water is removed with the glycerin, the concentration of water in this fraction can be sufficiently low that it can be recycled to the transesterification reactors. This lower boiling fraction often contains less than about 1, and more preferably less than about 0.5, mass percent water. Alternatively, the lower boiling fraction may be passed to a methanol and water distillation column in the esterification section of facility 100.
Alkanol is exhausted from first alkanol separator stage via line 306 and may be exhausted from the facility as a by-product, e.g., for burning or other suitable use, or can be recycled. Where no alkanol replacement reaction is used, the alkanol will be the lower alkanol for the transesterification and is recycled to the transesterification section. The bottoms stream from first alkanol separation stage 304 is passed via line 308 to second alkanol separation stage 314 for additional alkanol recovery. The design of second alkanol separation stage 314 may be similar to or different than that of first alkanol separation stage 304 and may be operated under the same or different conditions. Alkanol exits via line 316 and is combined with alkanol from line 306 and is passed to condenser 318. In the process of facility 100, the condensed alkanol will contain both the lower alkanol and the higher alkanol. Condensed alkanol is recycled via line 330 to alkanol replacement reactor 238. Non-condensed gases exit condenser 318 via line 320. As shown, the alkanol separation operation is maintained under vacuum conditions and these gases are passed to liquid ring vacuum pump 322. The liquid for the liquid ring is provided via line 324 and exits via line 328. As the gases contain some alkanol, the liquid for the liquid ring vacuum pump will remove alkanol from the gases. The liquid may be water, in which case the water may need to be treated to remove alkanol. Alternative liquid streams can be used, including but not limited to glyceride-containing feed, biodiesel, and glycerin. Feed is preferred as the liquid for the liquid ring vacuum pump since it can be passed to a transesterification reactor and alkanol contained therein used for the transesterification. Gas is removed from liquid ring vacuum pump 322 via line 326.
The bottoms stream from the second alkanol separation stage exits via line 334 and is passed to separator 336 in which a glycerin-containing phase and a biodiesel-containing phase are separated. The presence of alkanol in the crude biodiesel enhances the solubility of glycerin therein. Upon removal of the alkanol, a separate glycerin-containing phase, which also contains soaps, tends to form during the alkanol separation operation. The glycerin fraction is removed from separator 336 via line 338 and can be combined with spent glycerin in line 186. The lighter, oil-containing phase is passed via line 340 to a water wash unit operation. If desired, techniques can be used to assist in the phase separation of glycerin in separator 336 such as adding an effective amount of water to assist in the separation. Other components useful in enhancing phase separations may also be used including water-soluble inorganic salts that are essentially insoluble in the biodiesel-containing phase. If desired, any water-containing phase can be passed to evaporator 374.
Line 340 serves as a reactor and mixer where strong acid is supplied. The amount of strong acid provided is sufficient to convert any soaps remaining to free fatty acids. Sufficient strong acid is used such that water used for washing the crude biodiesel is at a suitably low pH. The strong acid is supplied in admixture with a recycle stream in the wash operation as will be explained later. While line 340 serves as an in-line mixer, a separate vessel may be used for the acidification. Where a separate mixer is used, it may be of any convenient design, e.g., a mechanically or sonically agitated vessel, or static mixer containing static mixing devices such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, or other impingement structure. In any event, sufficient mixing and residence time should be provided such that essentially all of the soaps are converted to free fatty acids. Often the temperature during the mixing is in the range of about 30° C. to 220° C., preferably between about 60° C. to 180° C., and for a residence time of between about 0.01 to 4, preferably 0.02 and 1, hours.
For purposes of discussion only and not in limitation, the water wash operation uses a two stage water wash. Water wash operation may be of any suitable design. Typically, the water wash operates with a recycling water loop, often with the water recycle being at least about 20, say between about 30 and 500, mass percent of the crude biodiesel being fed to the column. Normally washing is operated at a temperature between about 20° C. and 120° C., preferably between about 35° C. and 90° C. The amount of water provided to each wash vessel is sufficient to effect a sought removal of glycerin, residual alkanol and any water-soluble contaminants from the crude biodiesel. Typically between about 20 and 200, preferably between about 30 and 100, mass parts of wash water are used per 100 mass parts of crude biodiesel. Usually the free fatty acid is present in an amount less than about 3000, most frequently less than about 2500, parts per million by mass in the biodiesel product, and thus no need exists to remove free fatty acid to provide a biodiesel product meeting current commercial specifications. Preferably between about 1000 and 2500 ppm-m free fatty acid is contained in the biodiesel product to aid in lubricity.
The vessels used for the water washing may be of any suitable design including a pipe reactor, mechanically or sonically agitated tank, a vessel containing static mixing devices such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, or other impingement structure. Each stage needs to effect a phase separation of the oil phase from the water phase. Such a separation may be inherent in, for instance, a wash column where the water and oil phases are moving countercurrently, or a separate phase separator may be provided. It is understood that other washing operations can be used such as a one vessel washing operation, an acid wash followed by a neutral wash, and the like. The washing may be effected in one or more stages and in one or more vessels. A single vessel, such as a wash column can contain a plurality of stages.
As shown, crude biodiesel is provided via line 340 to first wash stage 342. For purposes of discussion, wash stage 342 comprises an agitated vessel to provide desired contact between the oil and water phases and a decanter to effect separation. Typically, the agitated vessel provides a contact time of about 1 second and 10 minutes, say, 5 to 60 seconds. Crude biodiesel is contacted with acidic water from water loop 368. The washed biodiesel from first wash stage 342 is passed via line 344 to second wash stage 346 having a design similar to or different from that of stage 342. This biodiesel is contacted with water from water loop 364. In each stage the water, after contacting the biodiesel stream being processed, is returned to the respective loops. Acidic water is withdrawn from first wash stage 342 and recycled via line 368. Substantially neutral water is withdrawn from second wash stage 346 and recycled via line 364. Additional water is provided to line 364 via line 376 which will be described later.
As configured with separate water cycle loops, the pH of the water in second wash stage 346 may be neutral or less acidic than the water in first wash stage 342. Make-up water to line 368 is provided by line 366. A purge is taken from line 368 via line 372. The purge balances the amount of water in the wash loops and is at a suitable rate to maintain desirably low concentrations of impurities such as alkanol and glycerin in the water used for the washing. The purge is usually at a rate of between about 1 and 50, say 5 and 20, mass percent per unit time of the recycle rate in the loop.
Line 370 provides strong acid to the water recycled via line 368 for combining with crude biodiesel in line 340 or being passed to first wash stage 342. Adequate strong acid aqueous solution is provided that the water in line 368 has a pH sufficiently low to convert the soaps to free fatty acids. The acid may be any suitable acid to achieve the sought pH such as hydrochloric acid, sulfuric acid, sulfonic acid, phosphoric acid, perchloric acid and nitric acid. Sulfuric acid is preferred due to cost and availability and it is a non-oxidizing acid. The amount of strong acid aqueous solution provided is typically in a substantial excess of that required to convert the soaps to free fatty acid and to neutralize any remaining catalyst. The excess of acid is often at least about 5, preferably at least about 10, say between about 10 and 1000 times that required. Consequently the feed to first wash stage 342 provides a wash water in line 368 having a pH of up to about 4, preferably between about 0.1 and 4.
Returning to line 372, the purge water is passed to evaporator 374 which provides a lower boiling fraction and a higher boiling fraction. While an evaporator may be used, it is also possible to use a packed or trayed distillation column with or without reflux. Generally the bottoms temperature of evaporator 374 is less than about 150° C., preferably between about 120° C. and 150° C. The distillation may be at any suitable pressure. A membrane separation system may, alternatively or in combination, be used with evaporator 374 to effect the sought concentration of the spent water.
The lower boiling fraction contains water, potentially acid if not neutralized or salts, and some alkanol and is passed via line 376 to water wash loop 364. Fresh water is provided to line 376 by line 380. The higher boiling fraction contains glycerin, some alkanol and some water and potentially acid or salts thereof. The higher boiling fraction or a portion thereof is preferably passed via line 382 to line 170 or it can be combined with spent glycerin.
A washed biodiesel stream is withdrawn from second washing stage 346 via line 348 and is passed to drier 350 to remove water which exhausts via line 354. Preferably substantially all the alkanol has been removed from the crude biodiesel prior to drying to permit the water vapor to be exhausted without treatment to eliminate volatile organic components. Drier 350 may be of any suitable design such as stripper, wiped film evaporator, falling film evaporator, and solid sorbent. Generally the temperature of drying is between about 60° C. and 220° C., say, about 70° C. and 180° C. The pressure is generally in the range of about 5 to 200 kPa absolute. The dried biodiesel is withdrawn as product via line 352. The biodiesel product contains free fatty acid and preferably has a free fatty acid content of less than about 0.3 mass percent. An inert gas such as nitrogen may be used in facilitating drying.
The subatmospheric pressure is maintained in drier 350 by the use of liquid ring vacuum pump 356 which is in communication with line 354. Liquid ring vacuum pump 356 uses water as the sealing fluid which is provided by line 358 and water exits via line 362. The gases from liquid ring vacuum pump 356 exit via line 360.
Returning to glycerin header 214, the glycerin-containing streams are passed via line 242 to blending tank 246 such that a relatively uniform glycerin composition can be provided via line 248 to the pretreatment section of facility 100. Blending tank 246 may also provide sufficient residence time for any glycerides in the glycerin to transesterify with alkanol as well as permit any oil entrained in the glycerin phase to separate. As shown, an oil layer that forms in blending tank 246 can intermittently or continuously be withdrawn via line 247 for recycle to first transesterification reactor 202. Alternatively, the oil layer can be withdrawn with the glycerin and passed to the pretreatment section.
While all glycerin-containing streams from the transesterification and refining components of facility 100 have been shown to be directed to glycerin header 214, it is within the purview of the process to use fewer streams. As stated above, the bottoms from evaporator 374 may be passed via line 382 to line 170 or added to header 214 or removed from the facility as a by-product. Moreover, any of the glycerin-containing streams may be used elsewhere prior to being passed to blending tank 246, and the blended stream or a portion thereof in line 248 may be used elsewhere and either returned to glycerin header 214 or passed to pretreatment component of facility 100.
One such use may be to pretreat a feed provided by line 200 to dehydrate the feed. If the feed contains free fatty acids or phospholipids, its introduction into the pretreatment component rather than via line 200, may be preferred. In such a pretreatment, a portion of the alkanol contained in the glycerin phase as well as some of the base catalyst, will be partitioned to the oil phase.
FIG. 2 depicts one type of esterification reaction system 400 useful in the processes of this invention. The reaction system depicts two stages with glycerin treatment between stages and is adapted for use with an oil soluble esterification catalyst such as para-toluene sulfonic acid. It is apparent that the system can be used with other catalysts.
A fatty acid-containing feed is provided to apparatus 400 via line 402 and enters contact vessel 404. Contact vessel 404 is adapted to contact the feed with glycerin containing alkanol. Contact vessel 404 may be of any suitable design sufficient to promote contact between the oil and glycerin phases including static and mechanical mixing devises and may be an extraction column, in which case a subsequent phase separator may not be necessary. The contacting may be at any suitable pressure and temperature as set forth in connection with the description of FIG. 1. A mechanically agitated vessel is depicted.
The mixed stream from contact vessel 404 is passed via line 406 to phase separator 408. An oil phase containing free fatty acid and alkanol is passed via line 412 to first esterification reactor 414 and the glycerin phase is withdrawn via line 410. Alkanol is provided via line 416 to first esterification reactor 414. Additional catalyst, if required, can be provided via line 418. The esterification effluent from first esterification reactor 414 contains alkanol, catalyst, ester, water and free fatty acid. Usually at least about 40, preferably at least about 60, mass percent of the free fatty acid is converted to ester in first esterification reactor 414. This esterification effluent is passed via line 420 to contact vessel 422.
Contact vessel 422 is adapted to contact the feed with glycerin supplied by line 424. The glycerin may be from any suitable source, e.g., a glycerin containing stream from a transesterification process. Usually the mass ratio of glycerin to oil is in the range of about 0.05:1 to 1:1, preferably between about 0.1:1 to 0.5:1. Contact vessel 422 may be of any suitable design sufficient to promote contact between the oil and glycerin phases including static and mechanical mixing devises and may be an extraction column, in which case a subsequent phase separator may not be necessary. The contacting may be at any suitable pressure and temperature as set forth in connection with the description of FIG. 1. For the sake of convenience, the contacting is usually conducted at approximately the temperature and pressure conditions of the esterification in first esterification reactor 414. A mechanically agitated vessel is depicted.
The mixed stream from contact vessel 422 is passed via line 426 to phase separator 428. An oil phase containing free fatty acid and alkanol is passed via line 432 to second esterification reactor 434 and the glycerin phase is withdrawn via line 430. Alkanol is provided via line 436 to second esterification reactor 434. Typically all alkanol is provided to first esterification reactor 414. Additional catalyst, if required, can be provided via line 438. Usually, since the catalyst is oil soluble, no additional catalyst need be used. The esterification effluent from second esterification reactor 434 contains alkanol, catalyst, ester, water and free fatty acid. Often at least about 70, preferably at least about 90, mass percent of the free fatty acid in the feed is converted to ester in apparatus 400. Additional stages of esterification reactors can be used if desired.
The esterification effluent from second esterification reactor 434 is passed via line 440 to contact vessel 442. Contact vessel 442 is adapted to contact the feed with glycerin supplied by line 444. Usually the mass ratio of glycerin to oil is in the range of about 0.05:1 to 1:1, preferably between about 0.1:1 to 0.5:1. Contact vessel 442 may be of any suitable design sufficient to promote contact between the oil and glycerin phases including static and mechanical mixing devises and may be an extraction column, in which case a subsequent phase separator may not be necessary. The contacting may be at any suitable pressure and temperature as set forth in connection with the description of FIG. 1. For the sake of convenience, the contacting is usually conducted at approximately the temperature and pressure conditions of the esterification in second esterification reactor 434. A mechanically agitated vessel is depicted.
The mixed phase stream from phase separator 442 is passed via line 446 to phase separator 448. An oil phase is withdrawn from phase separator 448 via line 456 and a glycerin phase via line 450. As is described herein, two glycerin contact stages are used, the first to remove alkanol and water from the esterification effluent and the second to recover catalyst from the oil phase. The broad aspects of this invention contemplate that a single stage can be use.
Returning to FIG. 2, the oil phase in line 456 from separator 448 is passed via line 460 to neutralization reactor 462. If the neutralization and alkanol recovery occurs in a single stage, i.e., vessel 442 serves both functions, the esterification product can be withdrawn via line 458. To neutralization reactor 462 is fed a mixed stream of glycerin and base, e.g., sodium hydroxide or preferably potassium hydroxide, via line 464 in an amount sufficient to convert the catalyst to salt. Other bases can be used if desired. Usually the mass ratio of glycerin to oil is in the range of about 0.05:1 to 1:1, preferably between about 0.1:1 to 0.5:1. The glycerin source, if a transesterification waste stream, may already contain sufficient base that little, if any, additional base is required. Neutralization conditions can vary over a wide range in that the reaction between acid and base proceeds rapidly and does not require catalyst. Temperature and pressure are often with the range of about 10° C. to 150° C. and 90 to 1000 kPa absolute. A residence time of from about 0.1 to 100 minutes may be used. Under these conditions, any free fatty acid contained in the oil phase will also be saponified. Hence, the amount of base present should thus include an amount sufficient to effect the saponification as well as the neutralization of the catalyst. Neutralization reactor 462 may be of any suitable design. It is preferably a static or mechanically agitated reactor.
The effluent from neutralization reactor 462 is passed via line 466 to phase separator 468 with a neutral, esterification product being withdrawn via line 470. A glycerin phase is withdrawn via line 472 and passed to acidifier 476 to which mineral acid, e.g., sulfuric acid, is added via line 478. The acidification conditions in acidifier 476 are sufficient to provide free fatty acid and convert the catalyst from a salt to its acid form, e.g., toluene sulfonic acid. Typically a two phase mixture will result. Acidification conditions can vary over a wide range in that the reaction proceeds rapidly and does not require catalyst. Temperature and pressure are often with the range of about 10° C. to 150° C. and 90 to 1000 kPa absolute. A residence time of from about 0.1 to 100 minutes may be used.
The effluent from acidifier is depicted as being passed via line 480 to contact vessel 404 and phase separator 408. If desired a separate phase separator can be used with only the oil phase passing to contact vessel 404, or more preferably, directly to first esterification reactor 414. The glycerin phase from phase separator 408 will contain salts formed from the acidification in acidifier 476.
Returning to line 430 containing glycerin phase from phase separator 428, the glycerin phase contains alkanol and water. Line 430 directs the glycerin phase to contact vessel 404 for recovery of alkanol therefrom. With respect to line 450 containing glycerin phase from phase separator 448, that glycerin phase usually contains less water since less conversion to ester occurs in reactor 434 as compared to first esterification reactor 414. If the water content is sufficiently low, the glycerin phase, or a portion of the glycerin phase, can be passed to first esterification reactor 414 via line 452. Thus any catalyst contained therein as well as alkanol is available for use in first esterification reactor 414. Alternatively, the glycerin phase in line 450 may be passed via line 454 and 480 to contact vessel 404.
Reference is now made to FIG. 3. Biodiesel manufacturing facility 500 uses a suitable raw material feed provided via line 502 containing glycerides and free fatty acids. The raw material feed in line 502 is combined with a mixed stream of glycerin provided from the biodiesel process by line 614 as discussed below and base provided by line 504. Thus the base and the glycerin are simultaneously contacted with the raw material feed. Alternatively, the glycerin can be added to the raw material feed before or after the addition of the base to the raw material feed. The mixture is passed to contact vessel 506. As the saponification reaction occurs rapidly, the contact vessel may only be a length of pipe sufficient to provide distribution of the components. Other types of contact vessels can be used such as mechanical and sonically agitated reactors, and reactors with static mixing such as reactors containing static mixing devices such as a packed bed, baffles, orifices, venturi nozzles, tortuous flow path, or other impingement structure
The amount of glycerin used should be sufficient to not only provide a separate phase but also enable a substantial portion, preferably at least about 70, more preferably at least about 90, mass percent of the soaps formed from the free fatty acids to reside in the glycerin phase. The base used should be an alkali metal or alkaline earth metal base, preferably an alkali metal hydroxide or alkylate, and more preferably sodium or potassium hydroxide or methylate or ethylate. The amount of base provided should be at least sufficient to provide the sought degree of saponification of the free fatty acids. Typically the base is provided in an amount of between about 90 to 500, preferably between about 95 and 150, mole percent of that required to react on a stoichiometric basis with the free fatty acid. Often the glycerin contains between about 0.1 to 10 or 15 mass percent base.
The contact of glycerin with the raw material feed can also remove water from the feed. With some biomass, phospholipids may be present. Treatment by glycerin sorbent can reduce the phospholipids content of the glyceride phase.
In general, it is preferred to use only sufficient glycerin sorbent to effect the sought removal of fatty acids from the raw material feed and, if desired, to effect the sought degree of dehydration of the feed, although more can be used. The mass ratio of glycerin to raw material feed will vary depending upon the amount of free fatty acid in the raw material feed and, if desired to use glycerin for dehydration, the water content of the raw material feed and the glycerin sorbent. The glycerin sorbent is conveniently comprised of glycerin phase separated from transesterification reactor effluent, with or without intervening treatment. A number of advantages flow from using glycerin phase separated from the transesterification effluent. First, the glycerin phase contains some of the base catalyst. Second, soaps in the glycerin phase can be recovered as free fatty acids in a subsequent acidification.
The conditions of the contacting of the streams should be sufficient to convert the sought amount of free fatty acid to the corresponding soap. Since a glyceride phase will exist, the contacting should be under conditions such that good mixing of the components occurs. The temperature of the contacting may be within a wide range, say, from about 15° C. to 220° C., preferably within the range of about 20° C. to 120° C. The time of contacting is often in the range of from about 10 seconds to 6 hours, preferably between about 20 seconds and 1 hour.
The mixed phase system of glycerin and raw material feed is passed via line 508 to phase separator 510. Phase separator 510 may be of any convenient design including a decanter, a phase separation facilitated decanter that contains coalescing sites, and a centrifuge. A glycerin phase is withdrawn as the heavy phase via line 582 from phase separator 510 and will be discussed later. The lighter phase contains triglycerides and is passed via line 512 to reactor 520 for transesterification.
The transesterification is base catalyzed with a lower alkanol, preferably methanol, ethanol or isopropanol. For purposes of the discussion, methanol will be the alkanol.
As shown, methanol is supplied via line 516 from methanol header 514. Line 515 supplies fresh methanol to reactor 520. Although line 516 is depicted as introducing methanol into line 512, it is also contemplated that methanol can be added directly to reactor 520 at one or more points. Generally methanol is supplied in excess of that required to cause the sought degree of transesterification in reactor 520. More methanol can be supplied but it may be lost from the facility. Preferably, the amount of methanol is from about 101 to 500, more preferably, from about 110 to 250, mass percent of that required for the sought degree of transesterification in reactor 520. In the facility depicted, two reactors are used. One reactor may be used, but since the reaction is equilibrium limited, most often at least two reactors are used. Often, where more than one reactor is used, at least about 60, preferably between about 70 and 96, percent of the glycerides in the feed are reacted in the first reactor.
The base catalyst is shown as being introduced via line 518 to reactor 520. The amount of catalyst used is in excess of that amount of base that will react with free fatty acids to form soaps in the transesterification. The base catalyst may be an alkali or alkaline earth metal hydroxide or alkali or alkaline earth metal alkoxide, especially an alkoxide corresponding to the lower alkanol reactant. Preferred alkali metals are sodium and potassium. When the base is added as a hydroxide, it may react with the lower alkanol to form an alkoxide with the generation of water. Alternatively, the catalyst may be a heterogeneous base catalyst. The exact form of the catalyst is not critical to the understanding and practice of this invention. For purposes of discussion, potassium hydroxide is used as the catalyst and the catalysis is homogeneous.
Often the transesterification is at a temperature between about 30° C. and 220° C., preferably between about 30° C. and 80° C. The pressure is typically in the range of between about 90 to 1000 kPa (absolute) although higher and lower pressures can be used. The reactor is typically batch, semi-batch, plug flow or continuous flow tank with some agitation or mixing, e.g., mechanically stirred, ultrasonic, static mixer, e.g., a packed bed, baffles, orifices, venturi nozzles, tortuous flow path, or other impingement structure. The residence time will depend upon the desired degree of conversion, the ratio of methanol to glyceride, reaction temperature, the degree of agitation and the like, and is often in the range of about 0.1 to 20, say, 0.5 to 10, hours.
The partially transesterified effluent from reactor 520 is passed via line 521 to phase separator 522. Phase separator 522 may be of any suitable design including a decanter, a phase separation facilitated decanter that contains coalescing sites, and a centrifuge. A glycerin-containing bottoms phase is provided in the separator and is removed via line 524. This glycerin phase also contains any soaps made in reactor 520 and a portion of the catalyst. The soaps can be recovered from this stream as discussed later. The lighter phase contains alkyl esters and unreacted glycerides and is passed via line 526 to second transesterification reactor 528.
Reactor 528 may be of any suitable design and may be similar to or different than reactor 520. As shown, additional methanol is provided via line 530 from methanol header 514 and additional catalyst is provided via line 532. Preferably the transesterification conditions in reactor 528 are sufficient such that reactors 520 and 528 together react at least about 90, more preferably at least about 95, and sometimes at least about 97 to 99.9, mass percent of the glycerides in the feed to reactor 520. The transesterification in reactor 528 is typically operated under conditions within the parameters set forth for reactor 520 although the conditions may be the same or different. The residence time will depend upon the desired degree of conversion. The reactor is typically agitated, e.g., stirred or ultrasonically agitated, or the ingredients otherwise subjected to a mixing action such as by a static mixer, e.g., a packed bed, baffles, orifices, venturi nozzles, tortuous flow path, or other impingement structure. Plug flow reactors may be useful. The residence time will depend upon the desired degree of conversion, the ratio of methanol to glyceride, reaction temperature, the degree of agitation and the like, and is often in the range of about 0.1 to 20, say, 0.5 to 10, hours.
The effluent from reactor 528 is passed via line 534 to phase separator 536 which may be of any suitable design and may be the same as or different from the design of separator 522. A heavier, glycerin-containing phase is withdrawn via line 538. This stream contains soaps as well as some catalyst and methanol. As shown, this glycerin-containing layer is used as a portion of the glycerin sorbent for the raw material feed treatment. A lighter phase containing crude biodiesel is withdrawn from separator 536 via line 540. Advantageously, this stream contains less than about 1, preferably less than about 0.8, and most preferably less than 0.05, mass percent soaps based upon the total mass of alkyl esters and soaps. The lighter phase also contains methanol.
Alternatively, separator 536 can be eliminated provided that in reactor 520, the conversion of the glycerides in the feed is at least about 90, preferably 92 to 96 or 98, percent. Thus the lighter phase from phase separator 522 contains little glyceride. In reactor 528 the reaction proceeds quickly to completion by the addition of additional methanol and catalyst, and can be conveniently accomplished by a plug flow reactor. Especially with the higher conversions, the effluent from reactor 528 may be a single phase.
The crude biodiesel may be contacted with acid to neutralize any catalyst therein and then refined to remove methanol, soaps, water and glycerin. Crude biodiesel is then passed to methanol separator 542. Methanol separator 542 is of any convenient design including a stripper, wiped film evaporator, falling film evaporator, solid sorbent, and the like. Preferably the fractionation is by fast, vapor fractionation.
The lower boiling fraction containing the lower alkanol will contain a portion of any water contained in the crude biodiesel. Since the transesterification is conducted with little water being present, and a portion of the water is removed with the glycerin, the concentration of water in this fraction can be sufficiently low that it can be recycled to the transesterification reactors. This lower boiling fraction often contains less than about 0.1, and more preferably less than about 0.05, mass percent water.
Vaporized methanol is exhausted via line 544 and may be exhausted from the facility as a waste stream, e.g., for burning or other suitable disposal, or can be added to the methanol header 514. The bottoms stream from methanol separator 542 is passed via line 546 to mixer 548. In another embodiment, line 546 can pass the bottoms stream to an intermediate mixer for contact with water, and then the oil phase passed to mixer 548. Into mixer 548 is passed a strong acid aqueous solution via line 552. Mixer 548 may be an in-line mixer or a separate vessel. Mixer 548 should provide sufficient mixing and residence time that essentially all of the soaps are converted to free fatty acids. Mixer 548 may be of any convenient design, e.g., a length of pipe, a mechanically or sonically agitated vessel, or static mixer containing static mixing devices such as a packed bed, baffles, orifices, venturi nozzles, tortuous flow path, or other impingement structure. Often the temperature during the mixing is in the range of about 30° C. to 220° C., preferably between about 60° C. to 180° C., and for a residence time of between about 0.01 to 4, preferably 0.02 and 1, hours. A convenient mode of practice is to pass the bottoms stream from methanol separator 542 to mixer 548 without intervening cooling.
A strong acid aqueous solution introduced via line 552 has a pH sufficiently low to convert the soaps to free fatty acids. Often the pH is less than about 4, and more preferably less than about 3, say, between about 0.1 and 2.5. The acid may be any suitable acid to achieve the sought pH such as hydrochloric acid, sulfuric acid, sulfonic acid, phosphoric acid, perchloric acid and nitric acid. Sulfuric acid is preferred due to cost and availability. The amount of strong acid aqueous solution provided is typically in a substantial excess of that required to convert the soaps to free fatty acid and to neutralize any remaining catalyst. The excess of acid is often at least about 5, preferably at least about 10, say between about 10 and 1000 times that required. Consequently the effluent from mixer 148 is at a pH of up to about 4, preferably between about 0.1 and 3.
The effluent from mixer 548 is passed via line 560 to phase separator 562. Phase separator 562 may be of any suitable design including a decanter, a phase separation facilitated decanter that contains coalescing sites, and a centrifuge. A lower aqueous phase is withdrawn via line 564. A portion of this aqueous phase is purged and the remaining portion is recycled via line 552 to mixer 548. Make-up acid is provided via line 550 to line 552.
The lighter phase which contains crude biodiesel and free fatty acid is withdrawn via line 566 and is passed to water wash vessel 568. Usually the free fatty acid is present in an amount less than about 500, most frequently less than about 300, parts per million by mass, and thus no need exists to remove free fatty acid to provide a biodiesel product meeting current commercial specifications. Fresh water enters vessel 568 via line 570 and serves to remove residual methanol and salts from the crude biodiesel. Water wash vessel 568 may be of any suitable design including a mechanically or sonically agitated tank, a vessel containing static mixing devices such as a packed bed, baffles, orifices, venturi nozzles, tortuous flow path, or other impingement structure, or a wash column. Normally washing is operated at a temperature between about 20° C. and 120° C., preferably between about 35° C. and 90° C. The amount of water provided is sufficient to effect a sought removal of glycerin and residual methanol from the crude biodiesel. Typically between about 2 and 50, preferably between about 5 and 20, mass parts of wash water are used per 100 mass parts of crude biodiesel.
In a preferred embodiment, the spent water from wash vessel 568 is passed via line 572 to mixer 548 or combined with the aqueous solution in line 552. Most preferably, the water provided via line 570 is in an amount to replace the volume of purge from line 564 to maintain steady state conditions. Often the purge from line 564 is less than 20, preferably between about 5 and 15, volume percent of the heavier, aqueous phase withdrawn from separator 562.
A washed biodiesel stream is withdrawn from washing column 568 via line 574 and is passed to drier 576 to remove water which exhausts via line 578. Preferably substantially all the methanol has been removed from the crude biodiesel prior to drying to permit the water vapor to be exhausted without further treatment to eliminate volatile organic components. Drier 576 may be of any suitable design such as stripper, wiped film evaporator, falling film evaporator, and solid sorbent. Generally the temperature of drying is between about 60° C. and 220° C., say, about 70° C. and 180° C. The pressure is generally in the range of about 5 to 200 kPa absolute. The dried biodiesel is withdrawn as product via line 580. The biodiesel product contains free fatty acid and preferably has a free fatty acid content of less than about 0.8, and more preferably less than about 0.5, mass percent.
Returning to separator 510, the heavier glycerin-containing phase is withdrawn via line 582 and passed to mixer 584. Mixer 584 may be an in-line mixer or a separate vessel including a mechanically or sonically agitated tank, a vessel containing static mixing devices such as a packed bed, baffles, orifices, venturi nozzles, tortuous flow path, or other impingement structure, or a wash column. Mixer 584 should provide sufficient mixing and residence time that essentially all of the soaps are converted to free fatty acids. Often the temperature during the mixing is in the range of about 30° C. to 220° C., preferably between about 40° C. and 180° C., and for a residence time of between about 0.01 to 4, preferably 0.02 and 1, hours.
An aqueous solution of strong acid is passed via line 586 to mixer 584 where the soaps contained in the glycerin are converted to free fatty acids which form a separate, lighter phase. Often the pH of the aqueous solution is less than about 4, and more preferably less than about 3, say, between about 0.1 and 2.5. The acid may be any suitable acid to achieve the sought pH such as hydrochloric acid, sulfuric acid, sulfonic acid, phosphoric acid, perchloric acid, and nitric acid. Sulfuric acid is preferred due to cost and availability. The amount of strong acid aqueous solution provided is typically in a substantial excess of that required to convert the soaps to free fatty acid and to neutralize any remaining catalyst. The excess of acid is often at least about 5, preferably at least about 10, say between about 10 and 1000 times that required.
The mixed phase glycerin and free fatty acid is passed from mixer 584 via line 588 to phase separator 590. Phase separator 590 may be of any suitable design to provide a heavy glycerin-containing phase which is removed via line 594. The glycerin phase may be treated in any suitable manner. For instance, the glycerin layer may be neutralized and subjected to distillation to remove water and methanol, if present. It can also be used as a fuel. This glycerin-containing stream may be used as the source of glycerin for treating the feed and thus at least a portion would be provided to line 504.
The free fatty acid from separator 590 passed via line 602 to acid catalyzed esterification reactor system 600 for conversion to the methyl ester. Methanol is provided to reactor system 600 via line 604 and any additional catalyst required via line 606. Any suitable acid catalyzed process for the esterification of free fatty acids to make biodiesel can be used including homogeneous and heterogeneous catalysis processes.
The effluent from esterification reactor system 600 is passed via line 608 to counter current extraction column 610 where it is contacted with glycerin from line 524. The glycerin serves to remove methanol and water from the esterification effluent. The oil phase is withdrawn via line 612. Depending upon the volume of the stream and its methanol content, the stream may be introduced at various points in the refining section of the transesterification unit. As shown, the effluent in line 612 is directed to the water wash unit operation. The water washing will serve to remove methanol still remaining in the effluent.
If the volume of the stream is larger and thus the absolute amount of methanol is greater, it can be directed to methanol separator 542.
The glycerin phase from column 610 is passed via line 614 to contact vessel 506.
The apparatus 700 depicted in FIG. 4 is adapted to acid esterify glyceride-containing feed with an oil-soluble, acid catalyst with recovery and recycle of the catalyst. As shown, glyceride-containing feed which also contains free fatty acid is provided by line 702 to esterification reactor 704. Alkanol and oil-soluble catalyst are also supplied as will be discussed below.
Since the catalyst is organic miscible, reactor 704 may be any of a widely divergent type of reactor. Suitable designs include a pipe reactor, mechanically or sonically agitated tank, a vessel containing static mixing devices such as a packed bed, baffles, orifices, venturi nozzles, tortuous flow path, or other impingement structure. Reactor 702 may be one or more vessels, each defining a reaction stage. Preferably the reactor provides static or mechanical missing of the liquid.
Reactor 704 is maintained under esterification conditions to provide an esterification effluent containing alkyl ester, catalyst, water from the esterification and a reduced concentration of alkanol and free fatty acid. This effluent is passed via line 706 to separator 708. The alkanol concentration in the reaction menstruum in reactor 704 may contain more alkanol than is miscible in the oil phase and/or additional alkanol can be added to the esterification effluent via line 736 to provide an alkanol phase and an oil phase in separator 708. Due to the more polar nature of the alkanol, acid catalyst preferentially partitions to the alkanol phase in separator 708. An oil phase having a reduced concentration of acid catalyst is removed via line 710.
The alkanol phase from separator 708 is withdrawn via line 712. All or a portion of the alkanol phase in line 712 can be withdrawn via line 714 and passed to lights column 716. The alkanol phase withdrawn via line 714 will contain some water. Although the esterification of free fatty acid is an equilibrium-limited reaction, water is tolerable to some extent and hence a portion of the alkanol phase can be passed to reactor 704. Preferably, no separate water phase is formed in reactor 704 as the catalyst tends to preferentially partition to a water phase.
In lights column, water is removed by vaporization. As methanol, if it is the alkanol, has a lower boiling temperature than water, and other alkanols such as ethanol and isopropanol form azeotropes with water, at least a portion of the alkanol will be vaporized with water. Where a substantial portion of the alkanol will be removed by vaporization, a higher boiling liquid in which the catalyst is soluble can be supplied to lights column 716. For instance, glycerin or biodiesel can be provided by line 718 to lights column 716 in an amount sufficient to maintain the oil-soluble, acid catalyst in a liquid medium. Especially where glycerin is used, the amount of glycerin is preferably less than that which would enable a glycerin phase to form in reactor 704. Alternatively or in addition, glyceride-containing feed may be used and provided via line 720.
Lights column 716 may be of any suitable design including a flash distillation column, a trayed or packed distillation column, or the like.
The higher boiling fraction containing the acid catalyst is passed via line 722 to line 712 which is in fluid communication with reactor 704. A lower boiling fraction from lights column 716 passes via line 724 to condenser 726 with a water phase being withdrawn via line 728 and an alkanol phase, after condensation, being withdrawn via line 730. If the alkanol and water form an azeotrope, unit operation 726 may be a suitable unit operation for selectively removing water, e.g., a selective extraction. A portion of the alkanol may be passed via line 731 as reflux to lights column 716. All, or the balance, of the alkanol can be passed via line 730 to line 712 for recycle to reactor 704. Make-up alkanol can be provided via line 732. As shown, make-up catalyst is provided to the make-up alkanol stream via line 734. All or a portion of the alkanol provided by line 736, if needed, can be supplied from line 730 or another source of alkanol.
Where the alkanol is methanol and the acid catalyst is toluene sulfonic acid, the apparatus of FIG. 4 is capable of recovering over 98 mass percent of the acid per pass in the methanol phase in separator 708. Moreover, as the methanol phase can be relatively small in comparison to the oil phase yet still recover a high percentage of the catalyst, energy requirements for water removal by lights column 716 can be economically viable.
FIG. 5 depicts an additional embodiment of the invention. To an esterification unit designated generally by numeral 800 is fed a glycerides-containing feedstock that also contains free fatty acid via line 802. The feedstock is fed to first esterification reactor 804. First esterification reactor 804 may be a vessel or a length of pipe. But preferably other types of vessels are used such as mechanical and sonically agitated reactors, and reactors with static mixing such as reactors containing contact structures such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, and other impingement structures. Suitable reactors include those providing high intensity mixing, including high shear. First esterification reactor 804 is maintained under esterification conditions including the presence of catalyst. For purposes of this discussion, the catalyst will be sulfuric acid which is provided by line 806. As the alkanol, methanol is used and is supplied to first esterification reactor 804 via line 852.
Esterification reactor 804 is operated to provide a partial conversion of free fatty acid to methyl ester, e.g., between about 30 to 90, say, 40 to 80, percent of the free fatty acid is converted. A partially converted effluent is withdrawn from first esterification reactor 804 and is passed via line 808 to second esterification reactor 810. As shown, no additional methanol or catalyst is added nor is any removal of water effected. Second esterification reactor 810 may be the same or different from first esterification reactor. Second esterification reactor 810 is maintained under esterification conditions and provides an esterification effluent having an increased conversion of the free fatty acid to methyl ester. Often at least about 75 mole percent to essentially all, preferably between about 75 and 95 or 98, mass percent of the free fatty acid is converted to ester.
The esterification effluent from second esterification reactor 810 is passed via line 812 to decanter 814 wherein a sulfuric acid and methanol-containing phase is formed and removed via line 830 and an oil phase containing methanol is formed and removed via line 816.
Oil phase in line 816 is contacted in vessel 818 with basic glycerin supplied via line 820 wherein free fatty acid contained in the oil phase is converted to soap and a treated product is provided. The treated product is passed via line 822 to decanter 824. In decanter 824 a glycerin-containing phase is formed which also contains salt of free fatty acid and methanol. This glycerin-containing phase is passed via line 828 to contact vessel 832 where it is contacted with the sulfuric acid and methanol-containing phase supplied by line 830. An oil phase is also formed in decanter 824 and is withdrawn via line 826 for further processing, e.g., to make biodiesel.
Returning to contact vessel 832, a mixture of glycerin, sulfuric acid and methanol is generated therein. Acid and water are partitioned to the glycerin and a methanol-containing phase is formed. The methanol-containing phase may also contain some free fatty acid. The mixture is passed via line 833 to decanter 834. A methanol-containing phase is obtained in decanter 834 and is passed via line 836 to first esterification reactor 804. A glycerin-containing phase which also contains methanol and water is withdrawn from decanter 834 via line 838 and passed to stripper 842. As shown, base such as sodium or potassium hydroxide or alkoxide is added via line 840 to the glycerin-containing phase being passed through line 838. If desired a contact vessel may be used rather than a length of pipe. The base is provided in an amount sufficient to neutralize the sulfuric acid. Alternatively base can be added to contact vessel 832 or to line 833 with the elimination of decanter 834 and line 836.
Stripper 842 is operated under conditions of temperature and pressure to provide an overhead containing methanol and water and a bottoms fraction containing glycerin and sulfate salt and a reduced concentration of methanol, preferably less than about 10, more preferably less than about 5, mass percent methanol. The bottoms fraction is withdrawn from stripper 842 via line 844. The overhead from stripper 842 is passed via line 846 to methanol column 848 which serves to provide an overhead containing virtually no water, e.g., less than about 0.1, more preferably less than about 0.01, volume percent water. This methanol stream is passed to first esterification reactor 804 via line 852. Make-up methanol is supplied via line 854 to maintain the sought methanol to free fatty acid mole ratios in the esterification reactors. A bottoms fraction provided by methanol column 848 preferably contains less than 0.1 mass percent methanol and is withdrawn via line 850.

Claims

1. A process for the esterification of feed containing free fatty acid comprising:

a. contacting the feed with a stoichiometric excess of alkanol under acidic esterification conditions including the presence of acid catalyst at elevated temperature, for a time sufficient to provide an esterification effluent comprising alkyl ester of the free fatty acid, water and unreacted alkanol;

b. contacting at least a portion of the esterification effluent with sufficient glycerin to form a glycerin-containing phase and an oil phase comprising alkyl ester, said contacting being for a time and under conditions sufficient that said oil phase has a lower water and a lower alkanol concentration than the esterification effluent and said glycerin phase contains water and alkanol; and

c. phase separating the oil phase and the glycerin phase.

2. The process of claim 1 wherein the feed comprises glyceride containing free fatty acid.

3. The process of claim 1 wherein the glycerin phase contains at least about 60 percent of the water contained in the esterification effluent.

4. The process of claim 3 wherein the glycerin phase contains at least about 30 mass percent of the alkanol contained in the esterification effluent.

5. The process of claim 1 wherein the alkanol comprises methanol.

6. The process of claim 1 wherein the catalyst comprises sulfuric acid.

7. The process of claim 1 wherein at least a portion of the glycerin phase of step (c) is contacted with the feed prior to step (a) under conditions sufficient to provide a feed phase containing an increased concentration of alkanol and a spent glycerin phase having a decreased concentration of alkanol; phase separating the feed phase and the spent glycerin phase; and providing the feed phase to step (a).

8. A process for the esterification of feed containing free fatty acid comprising:

a. contacting the feed with a glycerin solution containing alkanol under conditions sufficient to provide a feed phase containing an increased concentration of alkanol and a spent glycerin phase having a decreased concentration of alkanol;

b. phase separating the feed phase and the spent glycerin phase; and

c. contacting the feed phase with a stoichiometric excess of alkanol under acidic esterification conditions including the presence of acid catalyst at elevated temperature, for a time sufficient to provide an esterification effluent comprising alkyl ester of the free fatty acid, water and unreacted alkanol.

9. The process of claim 8 wherein alkyl ester is added to the feed prior to step (a) to increase the solubility of alkanol in the feed.

10. The process of claim 8 wherein the alkanol comprises methanol.

11. A process for the esterification of feed containing free fatty acid comprising:

a. contacting the feed with a stoichiometric excess of alkanol under acidic esterification conditions including the presence of glycerin-soluble, acid catalyst at elevated temperature, for a time sufficient to provide an esterification effluent comprising alkyl ester of the free fatty acid, water, acid catalyst and unreacted alkanol;

b. contacting at least a portion of the esterification effluent with glycerin whereby a glycerin-containing phase and an oil phase comprising alkyl ester are formed, said contacting being for a time and under conditions sufficient that the oil phase has a lower acid catalyst concentration than the esterification effluent and a lower water concentration than the esterification effluent and said glycerin phase contains acid catalyst, water and alkanol;

c. phase separating the oil phase and the glycerin phase;

d. contacting the glycerin phase with soaps of free fatty acids to generate free fatty acids;

e. separating free fatty acids from said glycerin phase; and

f. recycling the free fatty acids to step (a).

12. A process for the esterification of feed containing free fatty acid comprising:

a. contacting the feed with a stoichiometric excess of alkanol under acidic esterification conditions including the presence of an oil-soluble, acid catalyst at elevated temperature, for a time sufficient to provide an esterification effluent comprising alkyl ester of the free fatty acid, acid catalyst, water and unreacted alkanol;

b. providing sufficient alkanol in said esterification effluent to enable to form an alkanol phase containing alkanol, water and oil-soluble acid catalyst and an oil phase containing alkyl ester and unreacted free fatty acid;

c. phase separating the alkanol phase and the oil phase; and

d. passing at least a portion of the alkanol phase to step (a).

13. The process of claim 12 wherein the oil-soluble, acid catalyst comprises an organosulfonic acid.

14. The process of claim 12 wherein water is separated from at least a portion of the alkanol phase of step (c) to provide a higher boiling catalyst fraction.

15. The process of claim 14 wherein at least one of free fatty acid, glyceride and glycerin is provided during distillation in an amount sufficient to maintain the oil-soluble, acid catalyst in a liquid phase.

16. A process for the esterification of feed containing at least about 5 mass percent free fatty acid comprising:

a. contacting the feed with a stoichiometric excess of alkanol under acidic esterification conditions including the presence of acid catalyst at elevated temperature below about 120° C., for a time sufficient to convert between about 50 and 95 mass percent of the free fatty acids to alkyl esters and provide an esterification effluent comprising alkyl ester of the free fatty acid, free fatty acid, water and unreacted alkanol;

b. contacting at least a portion of the esterification effluent with sufficient basic glycerin to saponify free fatty acid to form soaps of said free fatty acid and provide an oil phase comprising alkyl ester and a glycerin phase comprising glycerin and soaps; and

c. phase separating the oil phase and glycerin phase.

17. The process of claim 16 wherein the acidic esterification conditions comprise a temperature below about 120° C.

18. A process for esterification of feed containing free fatty acid comprising:

a. contacting the feed with a stoichiometric excess of alkanol, under acidic esterification conditions including the presence of acid catalyst that is soluble in alkanol and substantially insoluble in alkyl ester of fatty acid to provide an esterification effluent comprising an alkanol and acid catalyst-containing phase and a first oil phase containing alkyl ester of the free fatty acid, water and unreacted alkanol;

b. phase separating the alkanol and acid catalyst-containing phase and the first oil phase;

c. contacting the first oil phase with glycerin to provide a second oil phase having a reduced concentration of water and alkanol and a glycerin-containing phase containing water and alkanol;

d. admixing at least a portion of the glycerin-containing phase with the alkanol and catalyst-containing phase to provide an admixture; and

e. recovering alkanol from the admixture by vapor fractionation.

19. The process of claim 18 wherein the alkanol comprises methanol.

20. The process of claim 19 wherein the catalyst comprises sulfuric acid.