HK1120544A - Production of biodiesel and glycerin from high free fatty acid feedstocks - Google Patents

Description

Method for producing biodiesel and glycerin from high free fatty acid feedstock

Technical Field

The present invention relates to an improved process and system for producing biodiesel (biodiesel).

Background

There is currently a continuing interest in using renewable resources as a substitute for petroleum-derived chemicals. Fatty Acid Alkyl Esters (FAAE) produced from fats and oils have been investigated as alternatives to such petroleum-derived materials, especially diesel fuels.

Triglycerides from fats and oils have long been known as fuels for diesel engines. However, such use often results in engine failure. Conversion of fatty acids present in lipids to simple esters such as methyl and ethyl esters has been proposed as a remedy for such engine failures. See, for example, the method described in US patent No. 6,398,707. There is increasing evidence that these esters perform well in substantially unmodified diesel engines and that these esters are effective in reducing emissions of particulate and hydrocarbon pollutants relative to petroleum-diesel fuels. The word "biodiesel" refers to these esters.

Methods for producing biodiesel have been known for many years. For example, US patent No. 4,164,506 discloses a method for the synthesis of biodiesel, wherein fatty acids are acid catalyzed. The conversion of triglycerides by base catalysis is described in US patent nos. 2,383,601 and 2,494,366. The conversion of both free fatty acids and triglycerides with enzymatic catalysis is disclosed in US patent nos. 4,956,286, 5,697,986 and 5,713,965. However, none of these methods completely addresses the production of biodiesel from low value higher free fatty acid feedstocks.

Economic analysis of any process for producing biodiesel shows that the feedstock cost is the largest fraction of the biodiesel production cost. Although all contemporary industrial processes suggest the use of Free Fatty Acid (FFA) feedstocks up to 15 wt%, manufacturers (for conservation cost) prefer to use feedstocks with FFA contents up to 100 wt%.

Furthermore, most of the processes in the prior art are not attractive because they rely on acid catalyzed esterification of fatty acids. Acid catalysis is not suitable for treating such feedstocks containing FFA concentrations for two main reasons. First, an excess of acid catalyst is required to completely convert a feedstock with a high FFA content. Since the acid catalyst must be neutralized prior to processing the glycerides, an increased amount of catalyst addition will produce an excess of salt. Secondly, such a process will produce large amounts of waste water as disclosed in US patent nos. 4,303,590, 5,399,731 and 6,399,800.

Although enzymatic catalysis for esterifying free fatty acids has been reported in the literature, this is disadvantageous because the reaction product is inhibited by the water produced when the free fatty acids in the raw materials are esterified with the enzyme. Another problem with enzymatic treatment is the high cost of the enzymatic catalyst. In addition, enzymatic catalysts have a limited lifetime.

To avoid two-phase operation in packed beds and other reaction settings, some conventional methods of biodiesel production use volatile and toxic co-solvents. Such a method is disclosed in US patent No. 6,642,399B 2. The use of a readily volatile and toxic co-solvent is not acceptable to the environment.

In addition, some existing methods of producing biodiesel utilize water to wash residual glycerin and salts from FAAEs. However, this produces large amounts of waste water and increases the risk of forming FAAE emulsions, as disclosed in US patent No. 5,399,731.

To obtain market share in the fuel industry, biodiesel must have a price competitive with conventional hydrocarbon diesel. In order to be competitive in price, the production of biodiesel must be economically advantageous. The growing profitability requires the biodiesel industry to utilize lower cost feedstocks. In addition, the overall yield of biodiesel from fats and oils must be high. Since the cost of raw materials is close to two thirds of the total cost of producing biodiesel, increased yields are a very important criterion.

There is therefore a need to develop improvements to biodiesel production processes that will result in increased yields of biodiesel from the feedstock and minimize unwanted by-products. There is also a need to develop alternative processes that do not require high pressure or acid catalysts. These methods should not use toxic co-solvents or water in order to extract impurities. These processes also require the use of inexpensive feedstocks and high yields of biodiesel. In addition, these feedstocks need to have a high FFA content to compete with petroleum diesel.

Summary of The Invention

An economical and unique process consisting of several plant operations is disclosed for converting free fatty acids to glycerides and subsequently converting glycerides to glycerol and FAAE. The fatty acid alkyl esters of the invention produced according to the invention are typically fatty acid methyl esters, although other fatty acid alkyl esters can also be produced.

The present invention relates to a process for converting low value higher Free Fatty Acid (FFA) feedstocks to biodiesel and high quality glycerol at market prices comparable to petroleum derived diesel fuel. The method of the present invention therefore substantially departs from the conventional concepts and designs of the background art. To this end, the process of the present invention provides a process and apparatus primarily developed for the purpose of producing fatty acid alkyl esters and high quality glycerin from any low value higher free fatty acid feedstock.

In a preferred aspect of the invention, the stream rich in fatty acid alkyl esters is subjected to successive processing steps of distillation and/or non-evaporative separation to maximize the recovery yield of the purified biodiesel.

Another aspect of the invention relates to the isolation and purification of the major by-products of biodiesel production to a purity level of greater than 95% or 99.7% glycerol, with undetectable alcohol levels and less than 0.5% weight/weight (w/w) salt.

Another aspect of the invention relates to treating the byproduct stream from which the biodiesel has been separated to maximize the recovery yield of the purified biodiesel.

The present invention also relates to minimizing waste streams during normal operation, the use of lower operating conditions (such as pressure) than other industrial biodiesel processes, the lack of use of toxic co-solvents, and the production of high quality glycerol by-products.

In a preferred embodiment, the process is a continuous process.

The main step of the process involves a transesterification reaction of the glyceride stream with an alcohol, preferably in the presence of a basic catalyst to convert the glycerides to fatty acid alkyl esters and glycerol.

The fatty acid alkyl esters are then separated from the glycerol to produce a first liquid phase containing a fatty acid alkyl ester-rich stream and a second liquid phase containing a glycerol-rich stream.

The fatty acid alkyl ester rich stream is then subjected to a first distillation or non-evaporative separation process. Preferably, the stream rich in fatty acid alkyl esters is subjected to reactive distillation, wherein the stream is subjected to both separation and chemical reaction. Separating the stream by reactive distillation into (i.) a bottoms fraction or biodiesel stream comprising a plurality of fatty acid alkyl esters; and (ii.) an overhead fraction (consisting essentially of alcohol, i.e., the first wet alcohol stream) while chemically reacting two or more stream components together in a manner that removes undesirable impurities in one or more outlet streams. Such reactive distillation, for example, increases the yield of glycerides leaving the distillation column while increasing the purity of the biodiesel leaving the distillation column. The biodiesel exiting the distillation column can be separated into a purified biodiesel stream and a byproduct stream.

The biodiesel stream exiting the first distillation column can be further subjected to a second distillation or non-evaporative separation to obtain a purified second biodiesel stream and a second byproduct fuel stream. The second distillation is preferably carried out in a wiped-film evaporator or falling-film evaporator or other similar evaporation device. Non-evaporative separation is typically a physical separation technique such as freeze crystallization, stripping or liquid-liquid separation. A free fatty acid stream and/or a glyceride-rich stream may be further separated from the second byproduct fuel stream and then reintroduced into the fatty acid alkyl ester production process.

The glycerol-rich stream of the second liquid phase may also be further purified to produce a purified glycerol product and a (second) wet alcohol stream. A portion of the purified glycerol product is then recycled to the glycerolysis reactor (the glycerolysis process will be described in more detail below) for reaction with free fatty acids.

The wet alcohol stream may also be further purified, preferably continuously, to produce a purified alcohol product. In addition, at least a portion of the purified alcohol product may be recycled to the transesterification reactor for reaction with the glycerides.

The basic stream generated during the base-catalyzed transesterification reaction can be neutralized by adding an inorganic acid or more preferably an organic acid to the basic stream. Neutralization may also be carried out by adding acid directly to the transesterification reaction effluent stream or by adding acid to the fatty acid alkyl ester rich stream and/or the glycerol rich stream that has been separated from the transesterification reaction effluent stream.

Brief Description of Drawings

Features of the present invention may be more readily understood by reference to the accompanying drawings that illustrate preferred embodiments of the invention. In the drawings:

FIG. 1 is a schematic flow diagram of the process of the present invention.

FIG. 2 is a schematic block diagram of a biodiesel production system according to the present invention;

FIG. 3 is a schematic block diagram showing the basic steps of biodiesel production according to the method of the present invention;

FIG. 4 is a schematic flow diagram of the process of the present invention wherein the basic catalyst used during the transesterification reaction is neutralized with a mineral acid; and

FIG. 5 is a schematic flow diagram of the process of the present invention wherein the basic catalyst used during the transesterification reaction is neutralized with an organic acid;

FIG. 6 is a schematic block diagram showing reactive distillation to separate a stream rich in fatty acid alkyl esters from a transesterification reaction effluent stream as shown in example 6.

Fig. 7 is a schematic block diagram showing the recycle of a stream from a byproduct stream for further recovery of fatty acid alkyl esters.

FIG. 8 is a schematic block diagram of the process of the present invention illustrating the use of a non-evaporative separator that produces a stream rich in fatty acid alkyl esters, glycerides, and free fatty acids from which refined biodiesel can be recovered.

Fig. 9 is a schematic diagram showing the refining of biodiesel wherein the biodiesel stream is treated in an evaporation device such as a wiped film evaporator or a falling film evaporator to further recover fatty acid alkyl esters.

Fig. 10 is a schematic diagram illustrating an embodiment of the invention wherein the by-product (fuel) separated from the biodiesel stream is further recycled to an evaporation device such as a wiped film evaporator or a falling film evaporator for further recovery of fatty acid alkyl esters.

Fig. 11 shows another embodiment of the present invention, wherein the byproduct (fuel) stream separated from the purified biodiesel is further separated to additionally recover fatty acid alkyl esters.

Fig. 12 illustrates an embodiment of the invention in which a biodiesel stream may be directed to a non-evaporative separator, separated into a fatty acid-rich stream, and then again directed to a second evaporation unit for purification.

Description of The Preferred Embodiment

In the process of the present invention, biodiesel is produced by reacting glycerides with alcohols in a transesterification reactor to produce fatty acid alkyl esters. This reaction usually takes place in the presence of a basic catalyst. The alcohol is usually C₁-C₅An alcohol, preferably methanol.

The resulting transesterification effluent stream is then separated into a fatty acid alkyl ester rich stream and a glycerol rich stream. Each of these streams can then be purified or subjected to further separation processes to maximize the recovery efficiency of biodiesel, glycerol, and alcohols. The byproduct (fuel) stream separated from the purified biodiesel can be further processed to maximize the recovery efficiency of the biodiesel.

The basic transesterification effluent stream generated during the base-catalyzed transesterification process may be treated directly with a neutralizing agent such as an inorganic or organic acid. Alternatively, the neutralizing agent may be added to the fatty acid alkyl ester rich stream and/or the glycerol rich stream that has been separated from the transesterification reaction effluent stream. Fatty acid alkyl esters are recovered from this pH adjusted stream.

After neutralization, the neutralized stream can be further purified, such as by distillation or fractionation.

The process of the present invention may further comprise an esterification step wherein the free fatty acid feedstock is first converted to glycerides. The obtained glycerides are then introduced into a transesterification reactor.

The soap produced in the transesterification reactor is converted to free fatty acids using an acid as a neutralizing agent. Caustic reacts with the fatty acid in the transesterification reactor to form soap. The presence of soap makes phase separation between the fatty acid alkyl esters and the solution of glycerol, water, alcohol and salt difficult to achieve. As a result, in the glycerin-rich phase, the soap emulsifies and retains most of the fatty acid alkyl esters. Thus, purification of the glycerol-rich phase is complicated and the yield of alkyl ester is reduced due to the presence of soap.

An overview of the process of the invention is shown in fig. 3, wherein a feedstock 1 containing free fatty acids is introduced into a glycerolysis reactor 2 containing glycerol, where the free fatty acids are converted into glycerides. The glycerides are then introduced into a transesterification reactor 4 containing an alcohol, where the glycerides undergo transesterification to produce fatty acid alkyl esters and glycerol. The alcohol/base stream 3 may be introduced into the transesterification reactor 4 as a mixture of the basic catalyst and the alcohol, or the basic catalyst and the alcohol may be introduced into the transesterification reactor as separate streams introduced into the transesterification reactor 4. The transesterification reaction effluent stream 4a, or a portion thereof, is then neutralized during a neutralization/phase separation step 5, which neutralization may be carried out before or after the effluent stream 5a is separated into a fatty acid alkyl ester-rich stream and a glycerol-rich stream. Finally, the alcohol, glycerol and biodiesel can be refined in an alcohol refining step 6, a glycerol refining step 7 and a biodiesel refining step 8, respectively. The alcohol typically leaves the system as a small portion of waste stream 9a or is recycled back to the transesterification reactor via stream 11. The refined glycerol is separated into a technical grade glycerol stream 13 and/or recycled back to the glycerol hydrolysis reactor 2 via stream 15. Waste stream 9b may contain some unrefined glycerol. The alkyl esters may also be further refined in a biodiesel refining step 8 to produce a purified biodiesel stream 18 and a waste stream 19, which waste stream 19 may also be useful, for example, as boiler fuel.

As shown in fig. 7, at least a portion of the waste stream 19 may be reintroduced to the previous process, for example as stream 351 into the biodiesel refining stage 8 for further recovery of fatty acid methyl esters, or into the transesterification reactor 4 for transesterification of glycerides to fatty acid methyl esters, or into the esterification reactor 2 for esterification of fatty acids.

Alternatively, as shown in fig. 8, at least a portion 351 of the waste stream 358 may be separated in a separator 370 into (i.) a fatty acid alkyl ester rich stream 371 and (ii.) a glyceride rich stream 376 and/or a fatty acid rich stream 374. The fatty acid alkyl ester rich stream 371 may then be reintroduced to the biodiesel refining stage 8. The stream rich in glycerides 376 and/or the stream rich in free fatty acids 374 may then be reintroduced into the transesterification reactor 4 and/or the esterification reactor 2.

The process of the present invention may be a continuous process. For example, a continuous process in which one or more of the following steps may be carried out in a continuous manner, as will be apparent from the following description:

(1) optionally conditioning the fatty acid containing feedstock by heating, mixing and filtering;

(2) continuously reacting free fatty acids in the feedstock with glycerol in a glycerolysis or esterification reactor to produce glycerides;

(3) the glycerides are reacted with an alcohol in a transesterification reactor to produce fatty acid alkyl esters and glycerol. The reaction preferably takes place in the presence of a basic catalyst;

(4) separating the fatty acid alkyl esters and glycerol from the transesterification reaction effluent stream (e.g., by gravity separation of the two relatively immiscible phases) to produce a fatty acid alkyl ester-rich stream and a glycerol-rich stream;

(5) the stream enriched in fatty acid alkyl esters is purified by distillation and/or fractional distillation. In a preferred embodiment, the fatty acid alkyl ester rich stream is purified by reactive distillation, wherein the reaction in the distillation or fractionation column will help to reduce unwanted impurities such as glycerol. Purified fatty acid alkyl esters are acceptable for use as biodiesel;

(6) preferably, the glycerol-rich stream is purified using an organic acid, such as a weak organic acid, for example acetic acid, formic acid or propionic acid, and the alcohol is recovered from the stream. The purified glycerol may then be introduced into a glycerolysis reactor;

(7) purifying the wet alcohol stream obtained from steps (5) and (6) above and removing water from the stream; and

(8) at least a portion of the purified alcohol is recycled to the transesterification reactor for use in the transesterification reaction with the glyceride.

The process may also include further separating the biodiesel stream of step (5) by a second distillation or non-evaporative separation to obtain a purer biodiesel stream (or a second purified biodiesel stream) and a secondary byproduct fuel stream.

Alternatively, the biodiesel stream of step (5) may be further separated in a non-evaporative separator into (i) a stream rich in fatty acid alkyl esters and (ii) a stream rich in glycerides and/or free fatty acids. Preferred non-evaporative separators for use herein include freeze crystallization processes and liquid-liquid separation processes.

The stream rich in fatty acid alkyl esters obtained from this separation can then be combined with the biodiesel stream of step (5) and then subjected to a second distillation or non-evaporative separation. The stream rich in glycerides and free fatty acids may then be reintroduced into the transesterification reactor or esterification reactor.

Feedstocks useful for the production of biodiesel typically contain large amounts of free fatty acids. The feedstock typically contains from about 3 to about 100 wt% free fatty acids and optionally fats and/or oils.

Typically, the feedstock is a lipid feedstock. The free fatty acid feedstock used in the present invention may be a low grade lipid material derived from animal fats and vegetable oils, including recovered fats and oils. For example, the feedstock for producing biodiesel fuel may be a tallow feedstock, such as waste tallow or yellow grease. Such low-grade lipid materials are very complex and, due to their high content of free fatty acids (from a few percent to 50% or more), are difficult to process economically using prior art methods. In addition, this material contains substances and impurities which are unsuitable for processing and which have to be removed before processing or during the refining of the product.

The feedstock is first introduced into a conditioning vessel or reactor for heating, mixing, and/or filtering the feedstock to produce a conditioned feedstock. The material may then be filtered, such as with a moving screen.

After filtration, the concentration of free fatty acids in the conditioned lipid feedstock can be measured. Optionally, the concentration of free fatty acids in the conditioned feedstock is measured continuously throughout the process. The measurement is performed with an on-line free fatty acid measuring device, such as a titration device or a near infrared spectrophotometer, which is used to quantify the concentration of free fatty acids in the conditioned feedstock.

During conditioning, the feedstock may be heated to about 35 to about 65 ℃, preferably about 55 to about 65 ℃ while mixing. A homogeneous mixture of glycerides, free fatty acids and unsaponifiable matter is usually present in the conditioned feedstock.

During glycerolysis, glycerol is used as a reactant to convert free fatty acids in the feedstock into glycerides (mono-, di-and triglycerides). The reaction of the free fatty acids in the feedstock typically occurs in the absence of a catalyst. In the glycerolysis reactor, the free fatty acids in the feedstock are mixed and continuously reacted with glycerol at a suitable temperature and pressure to yield a glycerolysis reactor effluent stream, which typically contains less than about 0.5 wt% free fatty acids and a plurality of glycerides. Glycerolysis preferably occurs in the absence of a catalyst and a co-solvent.

Glycerol, typically a purified glycerol product, is typically added to the glycerol decomposition reactor in a proportion greater than the stoichiometric amount of glycerol required for the glycerolysis reaction. The amount of glycerol introduced into the glycerolysis reactor is typically from about 35 to about 400% of the stoichiometric ratio of glycerol to free fatty acids to produce glycerides. In a preferred embodiment, the amount of glycerol added to the glycerolysis reactor is about 300% of the stoichiometric amount of free fatty acids in the feedstock.

Preferably, glycerolysis is carried out at a temperature in the range of about 150 to about 250 ℃, typically between about 180 to about 250 ℃, and more typically between about 180 ℃ and 230 ℃. The reaction is generally carried out with stirring. The reaction is also typically carried out at a pressure of from about 0.1 pounds per square inch absolute to about 15 pounds per square inch absolute, more typically at a pressure of about 2 pounds per square inch absolute.

The reaction of free fatty acids and glycerol is usually carried out in the presence of a catalyst such as ZnCl₂But in the preferred embodiment in the absence of a catalyst. The glycerolysis reactor effluent stream may contain less than 0.5 wt% free fatty acids and a plurality of glycerides.

The glycerolysis reaction is generally a continuous reaction. To produce glycerides in the glycerolysis reactor, a continuous reaction of the free fatty acids in the feedstock with glycerol can be performed in response to the signal of an on-line fatty acid measurement device or spectrophotometer.

During the glycerolysis process, water is removed; the resulting glycerides are substantially free of water. Water is typically continuously removed from the glycerolysis reactor as a vapor through a fractionation column or vent in the headspace of the reactor. Preferably, the vapor exiting the glycerolysis reactor is fractionated to produce three streams, a first fraction having a high concentration of unsaponifiable matter vaporized from the feedstock that condenses into a liquid stream, a second fraction being a liquid fraction having a high concentration of glycerol and a vapor fraction, and a third liquid fraction having a high concentration of water. The liquid fraction containing glycerol may then be returned to the glycerolysis reactor.

The glycerol reactor may consist of two or more continuously stirred tank reactors operated in series. The residence time of such reactors is generally from about 30 to no more than about 500 minutes, and preferably no more than 200 minutes.

The glycerolysis effluent stream contains a plurality of glycerides that are reacted with an alcohol in a transesterification reactor, such as a continuously stirred tank reactor. In this reaction, the glycerides in the glycerolysis reactor effluent stream are transesterified into fatty acid alkyl esters and glycerol. The transesterification reaction is conducted at a suitable temperature and pressure to produce the desired transesterification reactor effluent stream.

The transesterification reaction, preferably a continuous process, takes place in the presence of a basic catalyst. Suitable basic catalysts include basic catalysts such as potassium hydroxide and sodium hydroxide. The basic catalyst may be added to the transesterification reactor in a proportion sufficient to catalyze the reaction. Typically, the amount of alcohol added to the transesterification reactor is from about 1 to 5 moles of alcohol per mole of fatty acid moieties of the glycerides present in the transesterification reactor inlet stream. More generally, the ratio is about 2 moles of alcohol per mole of fatty acid moieties present in the glycerides entering the transesterification reactor. The catalyst, typically potassium hydroxide, is added in a proportion of about 0.5 to 3% by weight of the catalyst relative to the weight of the glyceride, more typically about 1%.

Alternatively, an alkoxide, such as potassium methoxide, may be added to the transesterification reactor to promote base catalysis. In this way, glycerides can be rapidly converted to alkyl esters in the presence of a caustic alkoxide, such as a caustic methoxide catalyst.

The transesterification reaction is generally conducted at a temperature of from about 25 to about 65 deg.C, preferably from about 50to about 60 deg.C, and a pressure of from about 14.5 to about 3625 psia.

The alcohol is typically added to the transesterification reactor in a proportion greater than the stoichiometric amount of alcohol required for the base-catalyzed transesterification reaction. For example, the alcohol may be added to the transesterification reactor in a proportion equal to about 200% of the stoichiometric amount of alcohol required to catalyze the reaction.

Preferably, a plurality of alcohols or catalysts are added to the transesterification reactor.

The transesterification reactor generally comprises at least two continuously stirred tank reactors operating in series. The residence time of each tank reactor is typically from about 5 to about 90 minutes, typically about 60 minutes.

The resulting transesterification reactor effluent contains fatty acid alkyl esters and glycerol. Preferably, at least a portion of the glycerol is removed from the transesterification reactor prior to reacting the plurality of glycerides with the alcohol.

The various fatty acid alkyl esters obtained may then be separated from the glycerol in the transesterification effluent stream. Typically, the two phases are separated into distinct immiscible two phases, i.e., a first liquid phase enriched in the fatty acid alkyl esters and a second liquid phase enriched in the glycerol, based on the difference in density of the two phases and using gravity and/or centrifugal force.

Typically, the two phases are separated at about 25 to about 65 ℃ to produce a stream rich in fatty acid alkyl esters and a stream rich in glycerol. This separation process may be a continuous operation and may be carried out in a clarifier or by membrane filtration.

In a preferred embodiment, the fatty acid alkyl ester rich stream is subjected to reactive distillation in the biodiesel refining step 8 to separate the fatty acid alkyl ester rich stream into a bottoms fraction, an overheads fraction (comprising mainly excess alcohol) and a fatty acid alkyl ester product stream. This separation takes advantage of the difference in vapor pressure of the components of the fatty acid alkyl ester rich stream and the reaction loss of glycerol. Distillation column or fractionation column conditions, including temperature and pressure conditions, facilitate the chemical reaction to occur simultaneously with conditions in the same vessel in which the separation occurs. The reactive distillation in the embodiment shown in fig. 6 reduces the concentration of glycerol and increases the amount of glycerides leaving the column. Thus, reactive distillation increases the efficiency of the production process.

The end result of the reactive distillation is that the amount of glycerol visible in the transesterification effluent stream or first liquid phase is greater than the total amount of glycerol leaving the distillation or fractionation column. This is due to the reaction of glycerol with free fatty acids and/or fatty acid alkyl esters in the reactive distillation column to form glycerides.

Preferably, the overhead fraction produced by the fatty acid alkyl ester distillation column is a (first) alcohol stream comprising substantially alcohol. Preferably, the bottoms fraction contains high boiling impurities, unsaponifiable matter, monoglycerides, diglycerides, triglycerides and fatty acids.

Preferably, the fatty acid alkyl ester distillation or fractionation column is operated at a pressure of less than about 15 pounds per square inch absolute. More preferably, the fatty acid alkyl ester distillation or fractionation column is operated at a pressure of about 0.1 pounds per square inch absolute to about 3 pounds per square inch absolute. Preferably, the fatty acid alkyl ester distillation or fractionation column is operated at about 180 to about 290 ℃, more preferably at about 230 to about 270 ℃. Preferably, the fatty acid alkyl ester distillation column or fractionation column contains packing.

The glycerol-rich second liquid phase stream may also be further purified and the alcohol recovered therefrom. The recovered alcohol is used to produce a purified glycerol product and a (second) wet alcohol stream. In a preferred embodiment, this step utilizes one or more of glycerol fractionation (where fractions in the glycerol-rich stream are separated by distillation), phase separation (where impurities fractionated with glycerol are removed by immiscibility and density differences), and glycerol purification (where other impurities are removed from glycerol).

The glycerol-rich stream may be further subjected to a phase separation wherein a liquid phase rich in fatty acid alkyl esters and a liquid phase rich in glycerol are separated, which may then be subjected to the purification described in the preceding paragraph.

The glycerol-rich stream may be further purified in a glycerol distillation column or fractionation column to produce bottoms, a side draw, and an overhead stream. Preferably, the bottoms material contains mainly waste; the side cut mainly contains glycerol and trace impurities; the overhead stream, which contains primarily alcohol and water, is collected for further purification and recycle.

Preferably, the glycerol distillation column is operated at an elevated temperature between about 180 ℃ and 280 ℃, preferably about 180 ℃ and 230 ℃. The distillation column is typically operated at reduced pressure of less than about 2 pounds per square inch absolute, typically from about 0.1 to about 2 pounds per square inch absolute.

The glycerol-rich stream may also be passed through a decolorization column where colored impurities and odors are removed from the glycerol, i.e., "glycerol finishing". The decolorization column typically contains a packed bed of activated carbon operating at about 35 to about 200 ℃, preferably about 40 to about 100 ℃. The contact time is often less than 4 hours. The activated carbon fines entrained by the packed bed were removed by filtration.

The water may also be removed from the wet alcohol stream by passing the wet alcohol stream through an alcohol distillation column or fractionation column at from about 60 to about 110 ℃ and from about 14 to about 20 pounds per square inch absolute to obtain a purified alcohol. Preferably, this purification comprises adsorption on molecular sieves which can subsequently be dried and reused, or distillation to produce a bottom product consisting essentially of water.

At least a portion of the purified glycerol product may then be returned to the glycerolysis reactor for reaction with free fatty acids in the feedstock; at least a portion of the purified alcohol is recycled to the transesterification reactor for reaction with the glyceride.

It is generally desirable to neutralize the fatty acid alkyl esters and glycerin produced in the transesterification reactor. Neutralization is often required depending on the caustic conditions characterizing the transesterification reaction. Such neutralization may be carried out by adding an acid to the transesterification reaction effluent stream or by adding an acid to the fatty acid alkyl ester-rich stream or the glycerol-rich stream separated from the transesterification reaction effluent stream. Suitable acid treatments include inorganic acid treatments or more preferably organic acid treatments.

Suitable inorganic acids include sulfuric acid and phosphoric acid. The basic catalyst reacts with the mineral acid to form insoluble salts that can be removed from the glycerol-rich stream in a solids separation operation.

Fig. 4 is an example of a process in which a mineral acid such as phosphoric acid is employed. More specifically, fig. 4 shows that a feedstock 1 containing free fatty acids is introduced into a glycerolysis reactor 2, where the free fatty acids are converted to glycerides by an esterification reaction. The glycerides are then introduced into transesterification reactor 4, along with alcohol 3 and base catalyst 318 at 317 (shown in fig. 7), where the glycerides are transesterified to produce fatty acid alkyl esters and glycerol.

The transesterification reaction effluent stream 4a is first separated in a first phase separation 320, typically using gravity separation techniques, into a fatty acid alkyl ester-rich stream and a glycerol-rich stream. Each stream is then purified in second phase separation 322 according to the methods described herein.

After having been separated into a fatty acid alkyl ester rich stream and a glycerol rich stream, the neutralizing acid, phosphoric acid 324, is added either before the first phase separation 320 or after the first phase separation 320 of the transesterification reaction effluent stream. Such alternative or combined introduction of acid to the process is indicated by the dashed lines in fig. 4.

Unfortunately, the use of phosphoric acid will result in an insoluble precipitate. The formation of insoluble precipitates requires the use of a filter in filtration step 326 and/or the use of a filter in filtration step 328. Suitable filters include rotary drum vacuum filters, plate and frame filter presses and belt filter presses.

In addition to using a filtration device, the use of mineral acids also requires rinsing of the insoluble byproduct salts to wash residual organic material therefrom. Suitable solvents include C₁-C₅Alcohols, such as methanol. Figure 4 shows an alcohol wash 330 that introduces an alcohol solvent 329 as an organic residue on the filter cake. Vacuum drying 332 is then used to remove the alcohol from the filter cake and dry the purified salt, which will then exit the process as waste stream 334. The solvent may be recovered as stream 364 for reuse in the present process.

Preferably, the method comprises drying the insoluble salt in a dryer at the operating pressure of the dryer at a temperature in the dryer that exceeds the boiling point of the solvent. The dryer may optionally be operated under vacuum to improve drying. The dryer may also include a condenser to recover the solvent for reuse.

Fig. 4 also shows that alcohol, glycerin and biodiesel are refined in the alcohol refining vessel 6, the glycerin refining vessel 7 and the biodiesel refining vessel 8, respectively. The alcohol typically leaves the system as a by-product stream 9a or is recycled back to the transesterification reactor 4 via 11. The refined glycerol is separated as purified glycerol 13. A portion of the glycerol stream may be recycled back to the glycerolysis reactor 2 as stream 15. The alkyl esters may also be further purified to produce purified biodiesel 18 or exit the system as a by-product 19 in the form of, for example, boiler fuel.

However, it is more preferable to use an organic acid than an inorganic acid. Although there are mineral acids which do not produce precipitated salts upon neutralization with the transesterified stream, they all have serious disadvantages. For example, hydrochloric acid and perchloric acid will generate chlorides in the process stream, which will in turn cause undesirable corrosion of steel and stainless steel, especially when the temperature is elevated. Sulfuric acid, sulfurous acid and hydrogen sulfide also have serious drawbacks due to the presence of sulfur, which increases the tendency of sulfur to flow out with the finished biodiesel. This in turn will lead to potential failure of sulfur content limitations and the generation of unwanted sulfur oxides in the emissions of biodiesel burning engines. Arsenic acid, chromic acid, hydrocyanic acid and hydrofluoric acid have an undesirable risk of using and/or requiring undesirable additional treatment methods to dispose of undesirable byproducts. Finally, although iodic acid does not produce undesirable precipitates, it is not economically feasible.

When organic acids are used, insoluble salts will not be formed and thus there is no need to subject the stream to any solids separation operation. Suitable organic acids include weak organic acids such as formic, acetic and propionic acids. In these examples, the pH of the glycerol-rich stream resulting from the transesterification reaction may first be adjusted to a pH of less than 8.0, preferably from about 6.5 to about 7.0.

Figure 5 compares the process of the present invention wherein an organic acid 325 is used in the neutralization reaction of the basic catalyst as compared to the inorganic acid. In one embodiment, the organic acid is added to the transesterification reaction effluent stream in a weight proportion of from about 0.1 to about 5%, more typically about 0.9%, prior to separation of the glycerol-rich stream from the fatty acid alkyl ester-rich stream. In another embodiment, the organic acid is added to the glycerol-rich stream in a weight proportion of from about 1 to about 7%, more typically about 4%. The use of organic acids would eliminate the need for filtration, cake washing and vacuum drying steps and would therefore be advantageous over the use of inorganic acids.

As shown in fig. 7, a portion of byproduct (fuel) stream 351 is directed back to the biodiesel refining step via 351A, to the transesterification reactor via 351C, or to the esterification reactor via 351D. The composition of stream 351 has not changed prior to separation into streams 351A, 351C and 351D.

In contrast, in fig. 8, a portion of byproduct (fuel) stream 351 is separated in separator 370 into a fatty acid alkyl ester-rich stream 371 and/or a free fatty acid-rich second stream 374 and/or a glyceride-rich third stream 376. A portion of the second stream having a lower free fatty acid content is then introduced into transesterification reactor 4, and a portion of the stream having a higher free fatty acid content is introduced into esterification reactor 2.

Figure 9 illustrates an embodiment of the biodiesel refining step 8 wherein increased yields of biodiesel can be obtained utilizing a second distillation reactor or non-evaporative separator. In a preferred embodiment, the second distillation reactor is one or more evaporation devices, such as a wiped film evaporator or a falling film evaporator as known in the art. Typically the second distillation reactor is located in a biodiesel refining unit. In addition, the separator apparatus can also be used to treat a byproduct (fuel) stream resulting from the purification of biodiesel.

A system for producing biodiesel from feedstocks such as lipid feedstocks having free fatty acids can be constructed in accordance with the teachings set forth herein above. The system may include:

(1) optionally a conditioning reactor for continuously converting the feedstock to a conditioned feedstock. The conditioning reactor is used to heat, mix and filter the feedstock to produce a conditioned feedstock;

(2) an optional system for continuously measuring the concentration of free fatty acids in the conditioned feedstock. Suitable systems include an on-line free fatty acid measurement device for quantifying the concentration of free fatty acids in the conditioned feedstock;

(3) a glycerolysis reactor in which free fatty acids in the feedstock are continuously reacted with glycerol to produce glycerides. This reaction may be responsive to the signal of an on-line free fatty acid measurement device;

(4) a transesterification reactor for continuously reacting glycerides with alcohols and for converting glycerides into fatty acid alkyl esters and glycerol, preferably by base-catalyzed reaction. This reaction can be performed in response to the signal of an on-line free fatty acid measurement device;

(5) a separator for continuously separating fatty acid alkyl esters from glycerol and for producing a fatty acid alkyl ester rich stream and a glycerol rich stream. Suitable separators include clarifiers or phase separation centrifuges for producing a (first) liquid phase enriched in fatty acid alkyl esters and a (second) liquid phase enriched in glycerol.

(6) A purifier for continuously purifying the fatty acid alkyl ester rich stream and recovering alcohol from the fatty acid alkyl ester rich stream; the purifier is used to produce a purified biodiesel product and a first wet alcohol stream. Suitable purifiers include fractionation columns and distillation columns. In a preferred embodiment, the stream rich in fatty acid alkyl esters is purified by reactive distillation to obtain biodiesel;

(7) an optional evaporative separator, such as a wiped or falling film evaporator, for further separation of the biodiesel into a fatty acid alkyl ester rich stream and a byproduct (fuel) rich stream;

(8) an optional non-evaporative separator for separating the byproduct (fuel) stream into a fatty acid alkyl ester rich stream and a free fatty acid/glycerol rich stream;

(9) a purifier for continuously purifying the glycerol-rich stream and recovering alcohol from the glycerol-rich stream; the purifier is used to produce a purified glycerol product and a second wet alcohol stream. Suitable purifiers include fractionation columns and distillation columns, including reactive distillation;

(10) a purifier for continuously purifying the wet alcohol stream to produce a purified alcohol product. Suitable purifiers include an alcohol fractionation column for treating an alcohol stream; and

(11) a channel for recycling at least a portion of the purified glycerol product to the glycerolysis reactor and recycling at least a portion of the purified alcohol to the transesterification reactor for continuous reaction with the glycerol ester.

Referring to fig. 1, a preferred embodiment of a biodiesel production process 10 for converting a high free fatty acid feedstock to biodiesel is shown.

In a feedstock introduction step 12, feedstock is introduced into the process 10. The introduced feedstock is preferably conditioned in a feedstock conditioning operation 14 in which the feedstock is heated and mixed in a conditioning reactor 16; the higher free fatty acid feedstock is heated and mixed to ensure a homogeneous mixture. The free fatty acids can be quantified, for example, with an in-line free fatty acid measuring device 18, wherein the concentration of free fatty acids in the feedstock is measured by spectroscopy, titration, or other suitable means. In the first separation, solid (insoluble) material is removed in filter 24.

The feedstock may include at least one free fatty acid at a concentration of about 3 to about 97 weight percent; moisture, impurities and unsaponifiable matter in concentrations of up to about 5 wt%; and a remainder comprising monoglycerides, diglycerides and/or triglycerides. The raw material can also comprise swill grease (trap grease).

Preferably, the conditioning step is carried out and a conditioned feedstock is produced, the conditioned feedstock having a temperature of from about 35 to about 250 ℃, and more preferably from about 45 to about 65 ℃. In a preferred embodiment, the feedstock is heated to a temperature of from about 55 to about 65 ℃. Preferably, the conditioned feedstock obtained is substantially free of insoluble solids.

The conditioned feedstock is introduced to a glycerolysis or esterification reaction at 26, preferably comprising a glycerol addition step 28, a heating step 32, a glycerolysis step 34 (wherein free fatty acids are converted to glycerides), and a glycerolysis effluent cooling step 38.

Preferably, the glycerolysis reaction step 26 further comprises conducting the glycerolysis reaction at a temperature of from about 150 to about 250 ℃; and removing water from the environment of the glycerolysis reaction. More preferably, the glycerolysis reaction step 26 also comprises the use of two or more continuous stirred tank reactors in series.

In a preferred embodiment, the free fatty acids and glycerol are reacted in the form of an esterification reaction, typically in the absence of a catalyst, continuously in a glycerolysis reactor at a temperature of about 220 ℃ and an absolute pressure of about 2 pounds per square inch, to produce an effluent stream containing less than 0.5 wt% free fatty acids and the plurality of glycerides. Preferably, the purified glycerol product is continuously fed to the glycerolysis reactor in a proportion of from about 35% to about 400% of the stoichiometric amount of free fatty acids, and water is continuously removed from the glycerolysis reactor as water vapor in a water removal step 35 by a fractionation column that returns condensed glycerol to the glycerolysis reactor.

Preferably, the reactor used in the glycerolysis step 34 comprises at least two continuous stirred tank reactors operating in series, with a total residence time in the reactors of no more than about 400 minutes for a feedstock having a free fatty acid concentration of 20 wt.%.

Water is preferably removed as water vapor by a fractionation or distillation column that returns condensed glycerol to the glycerolysis reactor.

The effluent from the glycerolysis reaction step 26 is introduced into a base-catalyzed transesterification reaction at 42, which preferably includes an alcohol metering step 44, a catalyst metering step 46, an alkoxide addition step 48, and a transesterification reaction step 50, wherein the glycerides are transesterified in the transesterification reactor.

In the transesterification step 50, the glyceride is contacted with an effective amount of an alcohol and an effective amount of a base catalyst under conditions wherein the glyceride, the alcohol, and the base catalyst are in sufficiently intimate contact. Preferably, the basic catalyst is selected from sodium hydroxide and potassium hydroxide.

The transesterification reaction step 42 is preferably conducted at a temperature of from about 20 ℃ to about 65 ℃ and at an absolute pressure of about 14.5 psia. More preferably, the transesterification step 42 includes conducting the transesterification reaction at a temperature of about 25 to about 65 ℃ and an absolute pressure near atmospheric pressure. In a preferred embodiment, the alcohol and the basic catalyst are mixed in a specified ratio before being added to the transesterification reaction operation.

In a preferred embodiment, the transesterification step 42 includes reacting a plurality of glycerides contained in the glycerolysis effluent stream with an alcohol in a transesterification reactor. In the transesterification reactor, a plurality of glycerides are mixed with an alcohol and an alkaline catalyst, preferably with a stirrer, and continuously reacted with the alcohol.

Preferably, the alcohol (most preferably methanol) is added to the transesterification reactor in a proportion equal to about 200% of the stoichiometric amount of alcohol required to catalyze the reaction, and the basic catalyst is added to the transesterification reactor in a proportion of about 0.5 to 2.0 wt% of the glycerides present in the glycerol decomposition effluent stream. More preferably, the basic catalyst is dissolved in the alcohol prior to introduction into the transesterification reactor.

Preferably, the transesterification reactor comprises at least two continuous stirred tank reactors operating in series, the total residence time of the reactors being no greater than about 90 minutes.

The transesterification reactor effluent stream contains a variety of fatty acid alkyl esters and glycerol. The effluent from the transesterification step 42 is preferably introduced into a second separation at 52 where the light phase (e.g., specific gravity 0.79-0.88) is separated from the heavy phase (e.g., specific gravity 0.90-1.20). In the biodiesel purification step (operation) 58 (indicated at 8 in fig. 3), excess methanol and high boiling impurities are preferably separated from the fatty acid alkyl esters in the light phase, and the alcohol is collected for reuse. Preferably, the separation of the fatty acid alkyl esters from the glycerol comprises separating the first light liquid phase and the second heavy liquid phase by taking advantage of the difference in density between them.

In the biodiesel purification step 56, the difference in vapor pressure between components is utilized to separate excess methanol and high boiling impurities from fatty acid alkyl esters in the light phase, and the alcohol is collected for reuse.

In a preferred embodiment, the second separation step 52 comprises separating the fatty acid alkyl esters and glycerol in the transesterification reaction effluent stream in a continuous clarifier in a phase separation step 54. Preferably, the first light liquid phase enriched in a substantial amount of fatty acid alkyl esters and the second heavy liquid phase enriched in glycerol are continuously separated in a continuous clarifier at a temperature of about 25 to about 65 ℃ to produce a fatty acid alkyl ester rich stream and a glycerol rich stream.

Alternatively, the separation step may be a reactive distillation column or a fractionation column that separates the fatty acid alkyl esters and glycerin. The transesterification reaction effluent stream entering the reaction column contains, in addition to fatty acid alkyl esters, a quantity of glycerol, glycerides and unreacted or non-convertible lipid feedstock. In the reaction column, some of the glycerol reacts with unreacted fatty acids and/or fatty acid alkyl esters to form glycerides.

In a preferred embodiment, the light phase is separated in a fatty acid alkyl ester purification step 56. In step 56, the difference in vapor pressure between the components is used to separate the fatty acid alkyl esters and excess alcohol from the high boiling impurities in the first light phase, and the alcohol is collected for reuse.

Preferably, the purification step 58 of the fatty acid alkyl ester rich stream also includes utilizing a distillation column to separate the fatty acid alkyl ester rich stream into a bottoms fraction, an overhead fraction containing primarily alcohol, and a side stream fraction containing fatty acid alkyl ester product. Preferably, the bottoms fraction produced by the distillation column contains impurities, unsaponifiable materials, monoglycerides, diglycerides, triglycerides and free fatty acids. Preferably, the distillation column produces a fatty acid alkyl ester product that meets ASTM standard D6751. Preferably, the distillation column produces an overhead fraction comprising predominantly alcohol.

In a preferred embodiment, the heavy phase from the second separation step 52 is treated in a catalyst separation step 62, step 62 comprising a mineral acid addition step 64; a catalyst precipitation step 66, wherein a basic catalyst is reacted with a mineral acid to produce a solid precipitate; a catalyst precipitation reactor effluent filtration step 70, wherein an alcohol wash step 68 is performed before removing alkali salt precipitates in a salt recovery step 71; a filtration separation step 72, in which the filtrate free of precipitates is separated into two liquid phases, fatty acids and fatty acid alkyl esters floating at the top and glycerol and most of the alcohol precipitating at the bottom; a pH neutralization step 74 to increase the pH of the glycerol; and a free fatty acid recycle step 76.

The crude glycerol is processed in a glycerol purification step 80, wherein the glycerol is purified by the difference in vapor pressure between the components. A preferred embodiment includes a distillation or fractionation step 84 wherein the alcohol and high boiling impurities are separated from the glycerol. The glycerol decolorization step 86 includes removing color and odor from the distilled glycerol using a packed bed of activated carbon.

Preferably, in the step of purifying the glycerol-rich stream and recovering the alcohol therefrom to produce a purified glycerol product and a wet alcohol stream, the alkaline catalyst in the glycerol-rich stream is reacted with a mineral acid, such as phosphoric acid or sulfuric acid, to produce fertilizer-valued insoluble salts which can be removed from the glycerol-rich stream in a solid separation operation, then filtered and washed with the alcohol.

The glycerol-rich stream is adjusted to a near neutral pH by the addition of a caustic solution and then further purified in a glycerol distillation column operating at a temperature of from about 180 to about 230 ℃ and an absolute pressure of less than 1 pound per square inch and a decolorization column containing a packed bed of activated carbon and operating at a temperature of from about 40 to about 200 ℃.

In a more preferred embodiment, the pH of the glycerol-rich stream is adjusted to about 6.5 to 8.0 by the addition of acid. Organic acids such as weak organic acids, for example acetic acid, propionic acid or formic acid, are then added to the glycerol-rich stream. The salts present in the glycerol-rich stream remain soluble. Thus, the use of organic acids would not require filtration and washing steps.

Preferably, the wet alcohol is treated in an alcohol purification step 88, wherein water is removed from the wet alcohol. More preferably, the water is removed by vapor pressure difference or adsorption. In a preferred embodiment, the alcohol is purified by distillation or fractionation in an alcohol distillation or fractionation step 90. In a preferred embodiment, purifying the wet alcohol stream comprises removing water therefrom to produce a purified alcohol product. Preferably, the wet alcohol stream is purified in an alcohol distillation column that is operated at a temperature of from about 60 to about 110 ℃ and a pressure of from about 14 to about 20 pounds per square inch absolute.

In the glycerol recycle step 92, the glycerol is preferably recycled to step 28, and in the alcohol recycle step 94, the alcohol is preferably recycled to step 44. Preferably, the glycerol recycling step 92 includes recycling at least a portion of the purified glycerol product to the glycerolysis reactor to react with the plurality of free fatty acids in the feedstock. Preferably, the alcohol recycling step comprises recycling at least a portion of the purified alcohol product to the transesterification reactor to react with the plurality of glycerides. Additional alcohol required for the transesterification reaction was supplied to the alkoxide tank. Biodiesel is marketed in a biodiesel transportation step 96 and glycerin is marketed in a glycerin transportation step 98.

Referring to fig. 2, a preferred embodiment of a system 110 for converting higher free fatty acid feedstocks to biodiesel is shown. Biodiesel production system 110 preferably includes subsystems and reactors as described below, wherein the alcohol used is methanol.

The feedstock is introduced into the system 110 in a feedstock introduction subsystem 112. In a preferred embodiment, the feedstock material consists of a free fatty acid content of 0-100% and a remainder comprising mono-, di-and triglycerides, moisture, impurities and unsaponifiables (MIU).

The incoming feedstock may optionally be conditioned in a feedstock conditioning subsystem 14, which subsystem 14 includes a feedstock heating and mixing vessel 16 in which the higher free fatty acid feedstock is heated and mixed to ensure a uniform, homogeneous mixture with a uniform viscosity is obtained. The concentration of free fatty acids in the feedstock can be measured with an online measurement device 18. The concentration is measured continuously to allow continuous control of downstream process steps.

Preferably, the feedstock material is heated in the feedstock heating and mixing vessel 16 to ensure that all available lipids are liquid and solids are suspended. Temperatures of at least 35 ℃ but not higher than 200 ℃ are sufficient to melt the lipids, reduce their viscosity and allow thorough mixing of the raw materials. A jacketed stirred tank may be used to provide agitation and to maintain the feedstock at an elevated temperature.

The conditioned feedstock may then be introduced into a glycerolysis reaction subsystem 26, which includes a glycerol addition device 28, an input heater 32, first and second glycerolysis reactors 134, 136, and a glycerolysis effluent cooler 38. The filtered product of step 24 is combined with glycerol and conditioned to promote glycerolysis reaction in glycerolysis reaction subsystem 126. In a preferred embodiment, these conditions include a reaction temperature of from about 150 to about 250 ℃ and a reaction pressure of from about 0.1 pounds per square inch absolute (psia) to about 30 psia. More preferred conditions are about 220 ℃ and about 2 psia.

The filtered lipid material is added with glycerol in an excess amount relative to the molar amount of free fatty acids of the lipid material. The excess is 10% -300% excess of glycerol (110% -400% of stoichiometric amount). In this embodiment, the glycerolysis reactors of elements 134 and 136 are configured as two heated, continuously stirred tank reactors in series. In these containers, the mixture of glycerol and lipid (containing free fatty acids) is agitated to maintain the two immiscible fluids in intimate contact.

In a preferred embodiment, mixing is performed with a stirrer. Under these conditions, free fatty acids are converted to glycerides (mono-, di-and triglycerides) and water is produced. Water is removed as steam and removed from the system at steam vent 35 along with all of the water originally present in the feed. In this preferred embodiment of the invention, the free fatty acid content in the reactor effluent stream may be consistently maintained at less than 0.5% w/w.

Because of the corrosive nature of the free fatty acids, the glycerolysis reactor is preferably constructed of materials that are resistant to organic acids.

The effluent of the glycerolysis reaction subsystem 126 contains mono-, di-, and tri-glycerides and residual fatty acids. The glycerolysis reaction effluent is introduced into a base catalyzed transesterification reaction subsystem 142 which preferably includes a methanol metering device 144, a potassium hydroxide metering device 146, a methoxide addition device 148, and a first transesterification reactor 150 and a second transesterification reactor 151 in which the glycerides undergo transesterification reactions.

In the transesterification reaction subsystem 142, glycerides are transesterified with simple alcohols having 1-5 carbons under basic catalyst. In a preferred embodiment, the basic catalyst is potassium hydroxide and the alcohol is methanol. Saponification of the remaining free fatty acids consumes about the same number of moles of alkaline catalyst as the number of moles of free fatty acids present.

The transesterification reaction is preferably catalysed by potassium methoxide, which is generated by adding potassium hydroxide to methanol. The amount of potassium hydroxide added is preferably equal to 0.5% -2.0% w/w of the glycerides present in the feed solution. Methanol and catalyst are combined and added to the glyceride solution from the glycerolysis reactor via methoxide addition means 148.

A 200% stoichiometric excess of methanol was added to the reaction mixture based on the moles of fatty acids available in the glycerides. Upon entering each of the transesterification reactors 150 and 151, the two-phase system is subjected to vigorous mixing.

Preferably, the reaction temperature is maintained at about 25 to about 65 ℃. At this temperature, miscibility of the phases is limited and mixing is required to obtain high conversion rates. The residence time required depends on the glyceride composition of the feed (composition between mono-, di-and triglycerides), the temperature, the catalyst concentration and the mass transfer rate.

Therefore, the stirring intensity is preferably considered in selecting the residence time. Generally, for conversions of glycerides to alkyl esters of greater than (>) 99%, the residence time required is 20-30 minutes.

In the transesterification reactor, the presence of potassium hydroxide, methanol and fatty acid esters can be corrosive. In a preferred embodiment, at least two continuous stirred-tank reactors in series are used. Suitable resistant materials are preferably selected for the reactor.

The effluent from the transesterification reaction subsystem 142 may be introduced into a phase separation subsystem 52, the phase separation subsystem 52 including a phase separation tank 54 in which a light phase (e.g., having a specific gravity of 0.79 to 0.88) and a heavy phase (e.g., having a specific gravity of 0.90 to 1.2) are separated. The effluent stream of the phase separator is a light phase fatty acid alkyl ester consisting of methanol and alkyl esters (biodiesel), a portion of excess alcohol and some impurities, and a heavy phase (crude glycerol) containing glycerol, alcohol, FAAE, soap, basic catalyst, traces of water and some impurities.

The phase separation apparatus 54 is preferably a conventional liquid/liquid separator capable of separating the heavy and light phases. Suitable phase separation equipment includes commercially available equipment, including a continuous clarifier 54.

In the biodiesel purification subsystem 56, excess methanol and high boiling impurities can be separated from fatty acid methyl esters in the light phase in a fractionation column 58 and the methanol collected for reuse. Preferably, the subsystem 56 for purifying the fatty acid methyl ester rich stream also includes a fatty acid alkyl ester distillation column 58 to separate the fatty acid alkyl ester rich stream into a bottoms fraction, an overhead fraction containing primarily methanol, and a side stream fraction containing fatty acid alkyl ester product.

Preferably, the bottoms fraction produced from distillation column 58 contains impurities and unsaponifiable matter, monoglycerides, diglycerides, triglycerides, and fatty acids. Preferably, the fatty acid methyl ester product produced by distillation column 58 of fig. 2 meets ASTM standard D6751.

Preferably, the overhead fraction produced by distillation column 58 comprises primarily methanol. Preferably, distillation column 58 is operated at a pressure of less than about 2 pounds per square inch absolute and a temperature of from about 180 to about 280 ℃. More preferably, distillation column 58 operates at an absolute pressure of from about 0.1 to about 2 pounds per square inch and a temperature of from about 180 to about 230 ℃. Preferably, distillation column 58 contains high efficiency structured packing.

The heavy phase separated by phase separation tank 54 is preferably treated in a catalyst separation subsystem 62, catalyst separation subsystem 62 comprising a mineral acid (such as phosphoric acid) addition apparatus 64, a catalyst precipitation reactor 66, a catalyst precipitation reactor effluent filter 70, a filtrate separation tank 72, a pH neutralization tank, and a free fatty acid recycle apparatus 76, in filter 70, washed with methanol 68 prior to removal of potassium phosphate precipitate 171 from the filter.

In the catalyst separation subsystem 62, the crude glycerol phase is pumped to a catalyst precipitation reactor to which mineral acid 64 is added. Preferably, the number of moles of acid added is equal to the number of moles of basic catalyst used in the transesterification reaction. The reaction product is an insoluble salt that can be isolated as a solid. In addition to forming insoluble salts, the acids convert soaps formed in transesterification reaction subsystem 142 to free fatty acids.

In a preferred embodiment, potassium hydroxide is used as a catalyst for the transesterification reaction, and the precipitation reaction uses phosphoric acid to produce potassium dihydrogen phosphate. The salt is insoluble in this system and can be removed by simple filtration. As the potassium phosphate salt is filtered in the catalyst precipitation reactor effluent filter 70, the methanol 68 is used to wash off the glycerin and other process chemicals from the precipitate.

The filtrate from the catalyst precipitation reactor effluent filter 70 is sent to another phase separation operation where two liquid phases are formed and separated in a filtrate separation tank 72 depending on their relative specific gravities. Glycerol, water, impurities and most of the methanol go to the bottom or heavy phase, while the fatty acid alkyl esters, some alcohols and fatty acids go to the top or light phase. This light phase is combined with the light phase of the previous phase separation subsystem (subsystem 52) and then sent to fractionation column 58. The heavy phase is sent to a reaction operation where a small amount of caustic is added to the pH neutralization reactor 74 to neutralize any residual acid. In a preferred embodiment, this is carried out in a continuously stirred tank reactor.

After the pH neutralization reactor 74, the crude glycerol phase is sent to a glycerol refining subsystem 80 where methanol and water are separated and collected for further purification, and glycerol is separated from high boiling impurities. In a preferred embodiment, the glycerol separation is performed in a glycerol distillation or fractionation column 84 having a glycerol side draw. The distilled glycerol may also be treated in a glycerol decolorization column 86, wherein activated carbon is used to decolorize and deodorize the distilled glycerol.

The methanol recovered from the distillation column contains trace amounts of water and is therefore considered a "wet" methanol stream that must be purified in methanol purification subsystem 88 before being reused in the process. This "wet" methanol stream is collected and purified by distillation in methanol purification column 90 before being pumped back to the inventory storage tank.

The distilled glycerol stream is then passed through activated carbon bed 86 for decolorization and deodorization. The feed entered the column from the bottom and flowed upward through the activated carbon bed to obtain > 95% pure colorless, solvent-free and salt-free glycerol.

A glycerol circulation pump 92 may be used to recycle glycerol to the glycerol addition apparatus 28. Methanol recycle device 94 is preferably used to recycle methanol to methanol metering device 144.

Biodiesel is then marketed by biodiesel delivery vehicle 96 and glycerin is marketed by glycerin delivery vehicle 98.

The process may also include a refining step to increase biodiesel production. Fig. 7 shows that the option of increasing the yield of biodiesel by further processing byproduct stream 358 depends largely on the relative concentrations of fatty acid alkyl esters, glycerides, and free fatty acids in byproduct stream 358. As shown, a portion of the byproduct stream 358 may be treated in the biodiesel refining step 8. As shown in fig. 7, the fatty acid alkyl ester rich stream 351A of the byproduct fuel stream 351 is reintroduced into the biodiesel refining step 8 for further recovery of fatty acid alkyl esters. Stream 358 may also be introduced into transesterification reactor 4 or esterification reactor 2 when a significant portion of the glycerides are present. As shown, a portion of byproduct stream 358 is introduced into transesterification reactor 4 as stream 351C. Alternatively, stream 351D is preferably introduced into esterification reactor 2 when it contains a higher content of free fatty acids.

In fig. 8, a portion of byproduct stream 358, represented as stream 351, may first be separated, preferably in non-evaporative separator 370, into a fatty acid alkyl ester-rich stream 371 and/or a glyceride-rich stream 376 and/or a free fatty acid-rich stream 374. The fraction containing a low level of free fatty acids can then be introduced into transesterification reactor 4 as stream 376, and stream 374 containing a higher level of free fatty acids can be introduced into esterification reactor 2. Suitable non-evaporative separation techniques which may be used are freeze crystallization, steam stripping or liquid-liquid separation.

A second distillation reactor or non-evaporative separator may also be used in the biodiesel refining step 8 to obtain increased yields of biodiesel. As shown in fig. 9, a fatty acid rich stream, such as fatty acid rich stream 323 separated from the transesterification effluent stream in first phase separation 320, is introduced into heat exchanger 405 and into flash drum 410 via pump 406. The flash tank 410 typically operates at a temperature of from about 60 to about 205 c, more typically about 140 c, and typically operates at a pressure of from about 1 to about 15 pounds per square inch absolute, preferably about 5 pounds per square inch absolute. Vapor 412 is removed and liquid stream 411 is then pumped into distillation column 420 via pump 415. In a preferred embodiment, as discussed above, distillation column 420 is a reactive distillation column. Overhead fraction 422 enters heat exchanger 440 and exits the system as stream 442, primarily excess alcohol, in vapor form. Condensate 441A leaving heat exchanger 440 leaves the system and liquid stream 441B reenters the distillation column. The bottoms fraction 421 of distillation column 420 is primarily a stream rich in fatty acid alkyl esters, which can then be introduced into reboiler 430 where it is either further separated in distillation column 420 as vapor stream 432 or exits as biodiesel stream 431A. Biodiesel stream 431A consists primarily of fatty acid alkyl esters, glycerides, and trace amounts of glycerin, as well as some fatty acids depending on the sour updraft. This stream may also be subjected to a second distillation in distillation column 450 via storage tank 440 to obtain a purified biodiesel stream 350C and a byproduct (fuel) stream 350A. In a preferred embodiment, distillation column 450 is one or more wiped or falling film evaporators as are known in the art. The temperature in the second distillation column 450 is approximately the same as the temperature in the distillation column 420. In an alternative embodiment, as shown in fig. 10, a portion of the byproduct (fuel) stream 350A may be reintroduced into the second distillation column 450 via storage tank 440.

The second distillation procedure may be carried out in one or more distillation columns. For example, a single wiped or falling film evaporator may be used. Furthermore, it is also possible to use a plurality of wiped-film evaporators or falling-film evaporators in parallel or in series. The residence time of the biodiesel in the wiped-film and falling-film evaporators is generally very short.

The wiped film evaporator consists of an internal rotating distribution plate which serves to distribute the biodiesel evenly over the heating plate of the evaporator onto the inner surface of the heated cylindrical shell. The spatula then spreads, stirs and moves the biodiesel down the heating shell in a very fast time, while the fatty acid alkyl ester evaporates quickly and condenses again on the cooled surface, usually in the center of the evaporator. Due to this particular configuration, the purified biodiesel stream then exits the bottom in the center of the evaporator, while the byproduct (fuel) stream exits the bottom at the periphery of the evaporator.

Falling film evaporators consist of a shell filled with steam or other heating medium and vertical parallel tubes through which the biodiesel falls. The flow of biodiesel is controlled such that the biodiesel creates a downwardly advancing membrane along the inner tube wall while the biodiesel selectively evaporates from the liquid. The separation of the biodiesel vapour from the residual liquid, which is usually composed of a mixture of glycerides, fatty acids and some unevaporated fatty acid alkyl esters, is carried out in a pipe. The biodiesel vapor is liquefied and recovered in a cold condenser.

Like distillation column 420, these second distillation columns 450 are typically operated at less than about 250torr absolute pressure and at a temperature of from about 150 ℃ to about 320 ℃. More preferably, distillation column 450 is operated at an absolute pressure of from about 0.1 to about 2torr and a temperature of from about 180 to about 230 ℃.

Fig. 11 illustrates another embodiment of the invention, wherein a byproduct (fuel) stream 350A is introduced into separator 370. Separator 370 is preferably a non-evaporative separator. In separator 370, a stream 371 rich in fatty acid alkyl esters can be separated from a stream 372 rich in glycerides and/or free fatty acids. The fatty acid alkyl ester rich stream 371 may then be reintroduced into the second distillation column 450 via storage tank 440 for further separation into purified biodiesel. Stream 372, which is rich in glycerides and/or free fatty acids, can then be reintroduced into transesterification reactor 4 and esterification reactor 2.

Fig. 12 shows another embodiment, wherein a stream 371 rich in fatty acid alkyl esters can be branched and combined as stream 371A with a purified biodiesel stream 350C. Additionally, a portion of stream 371 can be reintroduced into second distillation column 450. Further, fig. 12 illustrates the option of directing a portion or all of biodiesel stream 431A as 452 from first distillation column 420 to separator 370 for separation into a stream rich in fatty acid alkyl esters and a stream 372 rich in glycerides and/or free fatty acids. This stream 372 rich in glycerides and/or free fatty acids can then be reintroduced into the transesterification reactor 4 and/or esterification reactor 2.

The optimum spatial relationships for the parts of the invention, to include variations in size, materials, shape, form, function and manner of operation, assembly and use, are deemed readily apparent and obvious to one skilled in the art, and are intended to include equivalents to those illustrated in the drawings and described in the specification.

Accordingly, the foregoing description is considered as illustrative only of the principles of the invention. Further, since numerous modifications and changes will readily occur to those skilled in the art, it is not desired to limit the invention to the exact construction and operation illustrated and described, and accordingly, all suitable modifications and equivalents may be resorted to, falling within the scope of the invention.

Examples

Example 1

A refined (rendered) yellow grease having a concentration of 20 wt% free fatty acids and 2% moisture, impurities and unsaponifiables (MIU) was fed into a continuous stirred tank glycerolysis reactor at 100 pounds per minute (lbs/min). Filtration and titration were performed immediately as the lipid entered the glycerol decomposer. Glycerol was added at a rate of 13 lbs/min. The temperature of the lipid and glycerol mixture was raised to 210 ℃ as it entered the first glycerolysis continuous stirred tank reactor. In the reactor, the pressure was reduced to 2psia and the temperature was maintained at 210 ℃. The vessel is equipped with a high intensity stirrer to keep the immiscible liquids in contact. The water vapor produced by the reaction was removed through a vent in the reactor headspace. The residence time in each glycerolysis reactor was 2.5 hours. The conversion of fatty acids to glycerides in the first vessel was 85%. The fatty acid concentration leaving the second reactor was maintained at 0.5% w/w.

The product of the glycerolysis reactor was cooled to 50 ℃ and continuously fed to the transesterification reactor to which a solution of potassium hydroxide in methanol was added. Potassium hydroxide was added at a rate of 1.1lbs/min and mixed with 22lbs/min of methanol. The transesterification reaction took place in two continuous stirred tank reactors in series, each with a 2 hour residence time.

The transesterified product is then sent to a phase separation tank where most of the fatty acid methyl esters, a small amount of unreacted glycerides and a small concentration of unreacted methanol float to the top. Glycerol, most of the unreacted methanol, some fatty acid methyl esters, potassium hydroxide and soap sink to the bottom.

The bottom product or heavy phase was sent to an acidification reactor where the potassium hydroxide catalyst added in the transesterification step was reacted with 1.96lbs/min phosphoric acid. The soap is converted to free fatty acid and the potassium hydroxide is neutralized. The product of the acidification reaction is potassium dihydrogen phosphate, which is insoluble in the system.

The potassium dihydrogen phosphate precipitate was filtered off and the filtrate was sent to a second phase separation tank where the fatty acid methyl esters and free fatty acids in the filtrate floated to the top and glycerol and methanol settled to the bottom. The top or light phase is mixed with the light phase from the first phase separation tank and sent to a fatty acid methyl ester fractionation column. The pH of the heavy phase was adjusted back to 7.5 with potassium hydroxide and fed to the glycerol fractionation column.

The glycerol fractionation column recovered 10lbs/min of methanol and 18lbs/min of glycerol. The produced glycerol has a purity of more than 95% and no detectable concentrations of salts and methanol. The glycerol stream is divided into two streams: 13lbs/min were recycled back to the glycerol feed tank for glycerolysis reaction and 5lbs/min were pumped through the decolorizer, collected and placed on the market.

The two light phase streams were fed to a fatty acid methyl ester fractionation column where 2lbs/min of methanol was recovered and 92lbs/min of fatty acid methyl esters were produced, which met ASTM D6751-02 (Standard specification for biodiesel Fuel (B100) blendstocks used as distillate fuels).

Example 2

Purpose-made (funcy) color-off inedible tallow, having a free fatty acid concentration of 4 wt% and an MIU (moisture, impurities and unsaponifiable) concentration of 0.5%, was fed to the continuous stirred tank reactor at 100 lbs/min. The lipid is continuously filtered and titrated as it is fed to the glycerolysis reactor. Glycerol was added at a rate of 2.6 lbs/min. The temperature of the lipid and glycerol mixture was raised to 210 ℃ as it entered the first glycerolysis continuous stirred tank reactor. In the reactor, the pressure was reduced to 2psia and the temperature was maintained. The vessel is equipped with a stirrer to keep the immiscible liquids in contact. The water vapor produced by the reaction was removed through a vent in the reactor headspace. The residence time in each glycerolysis reactor was 2.5 hours. The conversion of fatty acids to glycerides in the first vessel was 92%. The fatty acid concentration exiting the second reactor was maintained at 0.5 wt%.

The product of the glycerolysis reactor was cooled to 50 ℃ and sent to the transesterification reactor where a solution of potassium hydroxide in methanol was added. Potassium hydroxide was added at a rate of 1.0lbs/min and mixed with 22lbs/min methanol. The transesterification reaction takes place in two continuously stirred tank reactors connected in series, each with a residence time of 2 hours.

The product of the transesterification reaction is then fed to a phase separation tank where most of the fatty acid methyl esters and a small concentration of unreacted methanol float to the top. Glycerol, most of the unreacted methanol, some fatty acid methyl esters, potassium hydroxide and soap sink to the bottom.

The bottom product or heavy phase was sent to an acidification reactor where the potassium hydroxide catalyst added during the transesterification operation was reacted with 1.79lbs/min phosphoric acid. The soap is converted back to free fatty acids and the potassium hydroxide is neutralized. The product of this acidification reaction is potassium dihydrogen phosphate, which is insoluble in this system.

The potassium dihydrogen phosphate precipitate was filtered off and the filtrate was fed to a second phase separation tank where the fatty acid methyl esters and free fatty acids floated to the top and glycerol and methanol settled to the bottom. The top phase or light phase is mixed with the light phase from the first phase separation tank and fed to a fatty acid methyl ester fractionation column. The pH of the heavy phase was adjusted to 7.8 with 0.1lbs/min of potassium hydroxide and the heavy phase was fed to a glycerol fractionation column.

The glycerol fractionation column recovered 10lbs/min of methanol and 10.2lbs/min of glycerol. The produced glycerol has a purity of more than 95% and no detectable concentrations of salts and methanol. The glycerol stream is divided into two streams: 2.6lbs/min were recycled back to the glycerol feed tank for glycerolysis reaction and collected for 7.6lbs/min for marketing.

The two light phase streams were fed to a fatty acid methyl ester fractionation column where 2.1lbs/min of methanol was recovered and 93lbs/min of fatty acid methyl ester was produced that met the requirements of ASTM D6751-02 (biodiesel fuel (B100) blend stock specification for distillate fuel).

Example 3

Degummed food grade soybean oil having a free fatty acid concentration of 0.5 wt% and MIU (moisture, impurities and unsaponifiables) was fed into the conditioning chamber at 100 lbs/min. The lipids were continuously filtered and titrated as they were removed from the feed conditioner. Due to the lower concentration of free fatty acids, the glycerolysis portion of the process can be skipped when using this feedstock.

The fatty acid concentration entering the transesterification reactor was 0.5 wt%. Potassium hydroxide was added at 1.0lbs/min and mixed with 22lbs/min of methanol. The transesterification reaction takes place in two continuously stirred tank reactors in series, the residence time of each tank being 2 hours.

The product of the transesterification reaction was then fed to a phase separation tank where most of the fatty acid methyl esters and a small concentration of unreacted methanol floated to the top. Glycerin, most of the unreacted methanol, some fatty acid methyl esters, potassium hydroxide and soap settle to the bottom.

The bottom or heavy phase was fed to an acidification reactor where the potassium hydroxide catalyst added for the transesterification operation was reacted with 1.76lbs/min phosphoric acid. The pH of the solution is lowered and the product of the acidification reaction is potassium dihydrogen phosphate, which is insoluble in this system.

The precipitate was filtered off at 2.2lbs/min and the filtrate was fed to a phase separation tank where the fatty acid methyl esters and free fatty acids floated to the top and glycerol and methanol were precipitated at the bottom. The overhead or light phase is mixed with the light phase of the first phase separation tank and fed to a fatty acid methyl ester fractionation column. The heavy phase was transferred to another tank and the pH was adjusted to 7.4 with 0.1lbs/min potassium hydroxide. Subsequently, the glycerol/methanol mixture is fed to a glycerol fractionation column.

The glycerol fractionation column recovered 10lbs/min of methanol and 8.5lbs/min of glycerol. The produced glycerol had greater than 95% purity and no detectable concentrations of salts and methanol. Collecting glycerol and putting the glycerol into the market.

The two light phase streams were fed to a fatty acid methyl ester fractionation column where 2.1lbs/min of methanol was recovered and 93lbs/min of fatty acid methyl ester was produced that met the requirements of ASTM D6751-02 (biodiesel fuel (B100) blend stock specification for distillate fuel).

Example 4

Refined swill grease (trap grease) with a concentration of 68 wt% free fatty acids and 5% MIU (moisture, impurities and unsaponifiables) was fed to the present invention at 100 lbs/min. The lipid was continuously filtered and titrated as it was fed to the glycerolysis reactor. Glycerol was added at 44 lbs/min. The temperature of the lipid and glycerol mixture was raised to 210 ℃ as it entered the first glycerolysis continuous stirred tank reactor. The pressure in the reactor was reduced to 2psia and the temperature maintained. The water vapor produced by the reaction was removed through a vent in the reactor headspace. The residence time in each glycerolysis reactor was 3.5 hours. The conversion of fatty acids to glycerides in the first vessel was 87%. The fatty acid concentration exiting the second reactor was maintained at 0.5 wt%.

The product of the glycerolysis reactor was cooled to 50 ℃ and fed to a transesterification reactor, where a solution of potassium hydroxide in methanol was added. Potassium hydroxide was added at 1.4lbs/min and mixed with 21lbs/min of methanol. The transesterification reaction takes place in two continuously stirred tank reactors in series, each with a residence time of 2 hours.

The product of the transesterification reaction was then fed to a phase separation tank where most of the fatty acid methyl esters and 10% of the unreacted methanol floated to the top, while the glycerol, most of the unreacted methanol, some of the fatty acid methyl esters, potassium hydroxide and soap settled to the bottom.

The bottom or heavy phase was fed to an acidification reactor where the potassium hydroxide catalyst added during the transesterification operation was reacted with 2.45lbs/min phosphoric acid. The soap converts back to free fatty acids and neutralizes the potassium hydroxide. The product of this acidification reaction is potassium dihydrogen phosphate, which is insoluble in this system.

The monopotassium phosphate precipitate was filtered off at 3.1lbs/min and the filtrate was fed to a second phase separation tank where the fatty acid methyl esters and free fatty acids floated to the top and glycerol and methanol were precipitated at the bottom. The overhead or light phase is mixed with the light phase of the first phase separation tank and fed to a fatty acid methyl ester fractionation column. The pH of the heavy phase was adjusted to 7.3 with 0.14lbs/min potassium hydroxide and fed to the glycerol fractionation column.

The glycerol fractionation column recovered 10lbs/min of methanol and 40lbs/min of glycerol. The produced glycerol had greater than 95% purity and no detectable concentrations of salts and methanol. The glycerol liquid was recycled back to the glycerol feed tank for glycerolysis reaction and an additional 4lbs/min of fresh glycerol was added to the glycerol feed tank to provide sufficient glycerol feed for glycerolysis reaction.

The two light phase streams were fed to a fatty acid methyl ester fractionation column where 2.1lbs/min of methanol was recovered and 91lbs/min of fatty acid methyl ester was produced that met the requirements of ASTM D6751-02 (biodiesel fuel (B100) blend stock specification for distillate fuel).

Example 5

A cooked brown grease having a concentration of free fatty acids of 37 wt% and 5% MIU (moisture, impurities and unsaponifiables) was fed to the present invention at 100 lbs/min. The lipid was continuously filtered and titrated as it was fed to the glycerolysis reactor. Glycerol was added at a rate of 24 lbs/min. The temperature of the lipid and glycerol mixture was raised to 210 ℃ as it entered the first glycerolysis continuous stirred tank reactor. The pressure in the reactor was reduced to 2psia and the temperature maintained. The vessel is equipped with an agitator to bring immiscible liquids into intimate contact. The water vapor produced by the reaction is removed through a vent in the reactor headspace. The residence time in each glycerolysis reactor was 3.0 hours. The conversion of fatty acids to glycerides in the first vessel was 90%. The fatty acid concentration exiting the second reactor was maintained at 0.5 wt%.

The product of the glycerolysis reactor was cooled to 50 ℃ and fed to a transesterification reactor, where a solution of potassium hydroxide in methanol was added. Potassium hydroxide was added at 1.2lbs/min and mixed with 21lbs/min of methanol. The transesterification reaction takes place in two continuously stirred tank reactors in series, each with a residence time of 2 hours.

The product of the transesterification reaction was then fed to a phase separation tank where most of the fatty acid methyl esters and 10% of the unreacted methanol floated to the top, while the glycerol, most of the unreacted methanol, some of the fatty acid methyl esters, potassium hydroxide and soap settled to the bottom.

The bottom or heavy phase was fed to an acidification reactor where the potassium hydroxide catalyst added during the transesterification operation was reacted with 2.13lbs/min phosphoric acid. The soap converts back to free fatty acids and neutralizes the potassium hydroxide. The product of this acidification reaction is potassium dihydrogen phosphate, which is insoluble in this system.

The monopotassium phosphate precipitate was filtered off at 2.7lbs/min and the filtrate was fed to a second phase separation tank where the fatty acid methyl esters and free fatty acids floated to the top and glycerol and methanol were precipitated at the bottom. The overhead or light phase is mixed with the light phase of the first phase separation tank and fed to a fatty acid methyl ester fractionation column. The pH of the heavy phase was adjusted to 7.5 with 0.12lbs/min potassium hydroxide and fed to a glycerol fractionation column.

The glycerol fractionation column recovered 10lbs/min of methanol and 25.2lbs/min of glycerol. The produced glycerol had greater than 95% purity and no detectable concentrations of salts and methanol. The glycerol liquid is divided into two streams: 24lbs/min were recycled back to the glycerol feed tank for glycerolysis reaction, collected at 1.2lbs/min and marketed.

The two light phase streams were fed to a fatty acid methyl ester fractionation column where 2.0lbs/min of methanol was recovered and 89.8lbs/min of fatty acid methyl ester was produced meeting the requirements of ASTM D6751-02 (biodiesel fuel (B100) blend stock specification for distillate fuel).

Example 6

A feedstock containing about 0.3 wt% free fatty acids and about 99.3 wt% glycerides (with the remainder being water and insoluble and unsaponifiable solids) was heated to 50c at a flow rate of about 40.9 pounds per hour and added to a solution of potassium hydroxide (1% by weight of the feed stream) in methanol (with the stoichiometric ratio of 2: 1 bound fatty acids in the methanolic glycerides). The transesterification reaction takes place in a single continuous stirred tank reactor with a residence time of 10 hours.

The flow rate of the transesterification reaction effluent stream was about 50.3 pounds per hour and consisted of about 79 wt% fatty acid methyl esters, 8 wt% glycerol, 9 wt% methanol, 1.6 wt% glycerides, and water, insoluble and unsaponifiable solids and soap as the remainder.

The stream was separated in a non-countercurrent separator into a light phase stream and a heavy phase stream, the light phase stream having a flow rate of 41.5 pounds per hour and a composition of about 94.26 wt% fatty acid methyl esters, 5.6 wt% methanol, 0.09 wt% glycerides, and 0.05 wt% free glycerol.

The free glycerol concentration of this sample and the other samples in this example were measured using the enzyme detection solution in the kit supplied by Sigma-Aldrich, inc., st.louis, MO, product code BQP-02. Using this kit, free glycerol can be measured by a coupled enzymatic reaction that ultimately produces a quinoneimine dye with an absorption maximum at 540 nm. Absorption peaks were measured using a Baush & Lomb Spectronic 20 spectrophotometer.

The light phase stream was analyzed for glycerol and found to contain about 490ppm by weight glycerol. The light phase stream is introduced into a reactive distillation column maintained at 260 ℃ and 150mmHg pressure. The overhead vapor stream of this column was condensed to produce a stream having a flow rate of about 2.1 pounds per hour consisting essentially of methanol having a glycerol content of 135 ppm. The bottom stream has a flow rate of about 39.3 pounds per hour and is comprised of about 98.5 wt% fatty acid methyl esters, 1.5 wt% glycerides, and only 3ppm glycerol. The reactive distillation referred to in this paragraph is schematically illustrated in FIG. 6.

The flow rates calculated using these gravimetric measurements in the feed to the column relative to the analysis of the free glycerol in the top and bottom streams show: about 98% of the glycerol reacts in the distillation column into other molecules rather than simply flowing as an overhead or bottoms stream.

The bottoms stream is further refined to produce a biodiesel oil that meets the requirements of ASTM D6751-06S 15 (Standard Specification for blend stocks of biodiesel Fuel (B100) for distillate fuels) for fatty acid methyl esters.

From the foregoing it will be observed that numerous variations and modifications may be effected without departing from the true spirit and scope of the novel concept of the invention.

Claims (11)

1. A method of producing biodiesel from glycerides, comprising:

(A) reacting the glycerides with at least one alcohol to produce a transesterification reaction effluent stream containing fatty acid alkyl esters and glycerol;

(B) separating the transesterification reaction effluent stream into a fatty acid alkyl ester-rich stream and a glycerol-rich stream;

(C) recovery of biodiesel

Wherein at least one of the following conditions is used:

(1) purifying the stream enriched in fatty acid alkyl esters by reactive distillation;

(2) introducing an organic acid into at least one of:

(a) transesterification effluent stream:

(b) a stream rich in fatty acid esters; and

(c) a glycerol-rich stream;

(3) purifying the fatty acid alkyl ester rich stream by distillation or fractionation in a column operating at less than about 3 pounds per square inch absolute and/or a temperature in a range of from about 180 ℃ to about 290 ℃;

(4) producing the fatty acid alkyl esters and glycerol of step (a) by reacting the glycerides with at least one alcohol in a transesterification reactor, wherein the at least one alcohol is also added to the transesterification reactor at a rate greater than the stoichiometric amount of alcohol required for the transesterification reaction;

(5) further subjecting at least a portion of the first biodiesel stream to a second distillation or non-evaporative separation to obtain a purified second biodiesel stream and a byproduct fuel stream, wherein the biodiesel of step (C) is recovered from the first biodiesel stream, wherein the byproduct fuel stream further:

(a) separating into a stream enriched in (i.) fatty acid alkyl esters and/or (ii.) free fatty acids and/or glycerides; or

(b) Reintroduced into the second distillation or non-evaporative separation.

2. The process of claim 1, wherein any of conditions (1), (2), (3), or (4) are used, and wherein at least a portion of the first biodiesel stream is subjected to a second distillation or non-evaporative separation to obtain a purified second biodiesel stream and a byproduct fuel stream, wherein the biodiesel of step (C) is recovered from the first biodiesel stream.

3. The process of claim 1 or 2, wherein the second distillation or non-evaporative separation is carried out in at least one wiped film evaporator or at least one falling film evaporator.

4. The process of any of claims 1-3, wherein the organic acid is a weak organic acid selected from the group consisting of acetic acid, formic acid, and propionic acid.

5. The process of any of claims 1-4, wherein the wet alcohol stream is recovered during purification of the glycerol-rich stream.

6. The process of any of claims 1-5, further comprising separating a free fatty acid and/or glyceride rich stream and/or a fatty acid ester rich stream from at least a portion of the first biodiesel stream.

7. The process of claim 6, wherein free fatty acids in the free fatty acid and/or glyceride rich stream separated from the portion of the first biodiesel stream are converted to glycerides that react to form the transesterification reaction effluent stream of step (A).

8. The process of claim 6, wherein the fatty acid ester rich stream separated from the second portion of the first biodiesel stream is further subjected to a second distillation or non-evaporative separation step.

9. The process of any one of claims 1-8, wherein the process is continuous.

10. The process of any one of claims 1-9, wherein the alcohol is methanol.

11. A method of producing purified biodiesel from glycerides, comprising:

(A) reacting the glycerides with at least one alcohol to produce a fatty acid alkyl ester stream;

(B) separating a biodiesel stream and a byproduct fuel stream from the fatty acid alkyl ester stream by distillation or non-evaporative separation; and

(C) thereby producing biodiesel;

wherein at least one of the following conditions is used:

(1) separating a second biodiesel stream and a fatty acid and/or glyceride rich stream from the byproduct fuel stream;

(2) separating the glycerides from the byproduct fuel stream and then reacting in step (a);

(3) separating a stream enriched in free fatty acids and/or glycerides from at least a portion of the first biodiesel stream, wherein the glycerides obtained are also introduced into the transesterification reactor.