Title of Invention

"A PROCESS FOR THE PRODUCTION OF PURIFIED BIO DIESEL FROM GLYCERIDES"

Abstract

A system and method for the conversion of free fatty acids to glycerides and the subsequent conversion of glycerides to glycerin and biodiesel includes the transesterification of a glyceride stream with an alcohol. The fatty acid alkyl esters are separated from the glycerin to produce a first liquid phase containing a fatty acid alkyl ester rich (concentrated) stream and a second liquid phase containing a glycerin rich (concentrated) stream. The fatty acid alkyl ester rich stream is then subjected to distillation, preferably reactive distillation, wherein the stream undergoes both physical separation and chemical reaction. The fatty acid alkyl ester rich stream is then purified to produce a purified biodiesel product and a glyceride rich residue stream. Biodiesel may be further recovered from the glyceride rich residue stream, by further separation of and/or processing of glycerides/free fatty acids contained therein. The glycerin rich second liquid phase stream may further be purified to produce a purified glycerin product and a (second) wet alcohol stream. Neutralization of the alkaline stream, formed during the alkali-catalyzed transesterification process, may proceed by the addition of a mineral or an organic acid.

Full Text

APPLICATION FOR PARENT INVENTORS. JOHN P. JACKAM; JOEL M. PIERCE, JEFFREY D. atClIARD II. TALLEV PRODUCTION OF DIQDIB6EL AND GLYCERIN PROM HIGH FREE. JATTY ACID FEEDSTOCKS SPECIFICATION Field of the Invention The present invention relates to improved processes and systems for biodiesel production. Background of the Invention There is continued and growing interest in the use of renewable resources as replacements for petroleum-derived chemicals. Fatty acid alkyl esters (FAAEs) produced from fats and oils have been investigated as replacements for such petroleum-derived materials, particularly diesel fuel. It has long been known that triglycerides from fats and oils can be used as fuels for diesel engines. However, such use typically results in engine failure. Remedies for such engine failure wherein conversion of fatty acids, found in lipids, into simple esters, such as methyl and ethyl esters, has been proposed. See, for instance, the process described in U. S. Patent No. 6,398,707. An increasing body of evidence indicates that these esters perform well in essentially unmodified diesel engines and that such esters may effectively reduce the output of particulate and hydrocarbon pollutants relative to petroleum-diesel fuel. The term "biodiesel" is applied to these esters. Processes for biodiesel production have been known for many years. For instance U. S. Patent No. 4,164,506 discloses a biodiesel synthesis wherein fatty acids are subjected to acid catalysis. The conversion of triglycerides with base catalysis is described in U. S. Patent Nos. 2,383,601 and 2,494,366. Conversion of both free fatty acids and triglycerides with enzyme catalysis is disclosed in U. S. Patent Nos. 4,956,286, 5,697,986 and 5,713,965. None of these processes, however, completely addresses the production of biodiesel from low value high free fatty acid feedstocks. An economic analysis of any process for the production of biodiesel indicates that feedstock cost is the largest portion of production cost for biodiesel. Whereas a 15 weight percent free fatty acid (FFA) feedstock is the highest content that any contemporary commercial process has proposed to handle, producers (in order to conserve costs) would prefer to use feedstocks having up to 100 weight percent FFA content. Further, most of the processes of the prior art are unattractive because they rely upon acid catalyzed esterification of fatty acids. Acid catalysis is not suitable for processing such feedstocks containing FFA concentrations for two principal reasons. First, an excessive amount of acid catalyst is required in order to fully convert feedstocks having high FFA content. Since the acid catalyst must be neutralized before processing the glycerides, the increased catalyst loading results in an excessive amount of generated salt. Further, such processes generate a large volume of waste water as disclosed in U. S. Patent Nos. 4,303,590, 5,399,731 and 6,399,800. While enzymatic catalysis has been reported in the literature for esterification of free fatty acids, it is disadvantageous because of reaction product inhibition from the presence of water which results when the tree fatty acids in the feedstock are esterified with enzymes. Another problem evidenced from enzymatic processing is the high cost of enzymatic catalysts. Further, enzymatic catalysts have a limited life. To avoid two-phase operation in packed bed and other reaction settings, some conventional processes for biodiesel production use volatile, toxic co-solvents. Such a process is disclosed in U. S. Patent No. 6,642,399 B2. The use of volatile, toxic co-solvents is environmentally unacceptable. Further, some prior art processes for producing biodiesel employ water to wash residual glycerin and salts from the FAAEs. This, however, generates a large volume of wastewater and increases the risk of forming FAAE emulsions, as disclosed in U. S. Patent No. 5,399,731. To gain market share in the fuels industry, biodiesel must be competitively priced with conventional hydrocarbon diesel. To be competitively priced, production of biodiesel must be economically profitable. Increased profitability requires that the biodiesel industry take advantage of lower cost feedstocks. In addition, overall yields of biodiesel from fats and oils be high. Increased yield is a very important criterion as feedstock costs approach two thirds of the total cost of production of biodiesel. Improvements in processes for biodiesel production therefore need to be developed which result in an increased yield of biodiesel from feedstocks and which minimize undesirable by-products. Alternative processes further need to be developed which do not require high pressures or acid catalysis. Such processes should not employ toxic co-solvents or water for the extraction of impurities. Such processes also need to produce high yield of biodiesel as well as employ inexpensive feedstocks. Further, such feedstocks need to have a high FFA content in order to be competitive with petrodiesel. Summary of the Invention A process is disclosed which combines several unit operations into an economical and unique process for the conversion of free fatty acids to glycerides and the subsequent conversion of glycerides to glycerin and FAAEs. The fatty acid alkyl esters of the invention produced in accordance with the invention are typically fatty acid methyl esters though other fatty acid alkyl esters may be produced. The invention relates to a process for converting low-value, high free fatty acid (FFA) feedstocks to biodiesel and high quality glycerin at a market price comparable to that of petroleum derived diesel fuels. The process of the invention therefore substantially departs from conventional concepts and designs of the background art. In so doing, the inventive process provides a process and apparatus primarily developed for the purpose of producing fatty acid alkyl esters and high quality glycerin from any low-value high free fatty acid feedstock. In a preferred aspect of the invention, streams enriched in fatty acid alkyl esters are subjected to successive treatment stages of distillation and/or non-evaporative separation in order lo maximize the yield of recovery of purified biodiesel. Another aspect of the invention relates to separation and purification of major byproducts of biodiesel production to render glycerin at a purity level greater than 95 or 99.7 percent, with non-detectable levels of alcohol and less than 0.5 percent weight/weight (w/w) salts. A further aspect of the invention relates to treatment of a by-product stream (from which biodiesel has been separated) in order to maximize the yield of recovery of purified biodiesel. The invention further relates to minimization of waste streams during normal operations, the use of lower operating conditions (such as pressures) than other commercial biodiesel processes, the non-use of toxic co-solvents and the production of a high quality glycerin byproduct. In a preferred embodiment, the process is a continuous process. The major steps of the process include the transesterification of a glyceride stream with an alcohol, preferably in the presence of base catalyst, to convert the glycerides to fatty acid alkyl esters and glycerin. The fatty acid alkyl esters are then separated from the glycerin to produce a first liquid phase containing a fatty acid alkyl ester rich stream and a second liquid phase containing a glycerin rich stream. The fatty acid alkyl ester rich stream is then subjected to a first distillation or to a non-evaporative separation process. Preferably, the tatty acid alkyl ester rich stream is subjected to reactive distillation, wherein the stream undergoes both separation and chemical reaction. By means of reactive distillation, the stream is separated into (i.) a bottoms fraction or biodiesel stream comprising a plurality of the fatty acid alkyl esters; and (ii.) an overhead fraction (principally composed of alcohol, a first wet alcohol stream), while simultaneously chemically reacting two or more stream components together in such a way as to remove unwanted impurities in one or more output stream(s). Such reactive distillation for example increases the yield amount of glycerides exiting the distillation column while increasing the purity of the biodiesel exiting the distillation column. The biodiesel exiting the distillation column may be separated into a purified biodiesel stream and a by-product stream. The biodiesel stream exiting the first distillation column may further be subjected to a second distillation or to a non-evaporative separation in order to render a purified second biodiesel stream along with a second by-product fuel stream. The preferred second distillation occurs in a wiped film evaporator or a falling film evaporator, or other such evaporative device. Non-evaporative separation typically is a physical separation technique, such as freeze crystallization, steam stripping or liquid-liquid separation. A free fatty acid stream and/or glyceride enriched stream may further be separated from the second by-product fuel stream and then re-introduced into the process for production of fatty acid alkyl esters. The glycerin rich stream of the second liquid phase may further be purified to produce a purified glycerin product and a (second) wet alcohol stream. A portion of the purified glycerin product may then be recycled into a glycerolysis reactor (in a glycerolysis process described in more detail below) for reaction with the free fatty acids. The wet alcohol streams may further be purified, preferably continuously, to produce a purified alcohol product. Further, at least a portion of the purified alcohol product may be recycled into the transesteriflcation reactor for reaction with the glycerides. Neutralization of the alkaline stream, formed during the alkali-catalyzed transesteriflcation process, may proceed by the addition of a mineral acid or more preferably an organic acid to the stream. Neutralization may occur by addition of the acid to the transesteriflcation effluent stream directly or to the fatty acid alkyl ester rich stream and/or glycerin rich stream after such streams have been separated from the transesteriflcation effluent stream. Brief Description of the Drawings. The features of the invention will be better understood by reference to the accompanying drawings which illustrate presently preferred embodiments of the invention. In the drawings: FIG. 1 is a schematic flow diagram of the process of the invention. FIG. 2 is a schematic block diagram of the biodiesel production system in accordance with the invention; FIG. 3 is a schematic block diagram showing the basic steps of the production of biodiesel in accordance with the process of the invention; FIG. 4 is a schematic flow diagram of the process of the invention wherein a mineral acid is used in the neutrali/ation of the alkali catalyst used during transesteriflcation; and FIG. 5 is a schematic flow diagram of the process of the invention wherein an organic acid is used in the neutrali/ation of the alkali catalyst used during transesteriflcation; FIG. 6 is a schematic block diagram which demonstrates reactive distillation of a fatty acid alkyl ester rich stream upon separation from the transesteriflcation effluent stream, as set forth in Example No. 6. FIG. 7 is a schematic block diagram which illustrates the recycling of a stream from a by- : stream for further recovery of fatty acid alkyl esters. FIG. 8 is a schematic block diagram of the process of the invention illustrating the use of a non-evaporative separator to generate streams enriched in fatty acid alkyl esters, glycerides and free fatty acids from which refined biodiesel may be recovered. FIG. 9 is a schematic diagram illustrating biodiesel refining wherein a biodiesel stream is treated in an evaporative device, such as a wiped film evaporator or falling film evaporator, for further recovery of fatty acid alkyl esters. FIG. 10 is a schematic diagram which demonstrates an embodiment of the invention wherein by-product (fuel) separated from a biodiesel stream is further recycled to an evaporative device, such as a wiped film evaporator or falling film evaporator, for further recovery of fatty acid alkyl esters. FIG. 11 shows another embodiment of the invention wherein the by-product (fuel) stream, separated from purified biodiesel, is further separated for additional recovery of fatty acid alkyl esters. FIG. 12 illustrates an embodiment of the invention wherein a biodiesel stream may be directed to a non-evaporative separator, separated into a fatty acid enriched stream and then redirected to a second evaporative device for purification. Detailed Description of the Preferred Embodiments. In the process of the invention, biodiesel is prepared by reacting glycerides with an alcohol in a transesterification reactor to produce fatty acid alkyl esters. This reaction typically occurs in the presence of an alkali catalyst. The alcohol is typically a Ci-Cj alcohol, preferably methanol. The resulting transesterification effluent stream may then be separated into a fatty acid alkyl ester rich stream and a glycerin rich stream. Each of these streams may then be purified or subject to further separation processes in order to maximize the efficiency in recovery of biodiesel, glycerin and alcohol. By-product (fuel) streams, separated from purified biodiesel, further be subjected to further processing in order to maximize the efficiency of biodiesel recovery. The alkaline transesterification effluent stream formed during the alkali-catalyzed transesterification process may be directly treated with a neutralizing agent, such as a mineral acid or an organic acid. Alternatively, the neutralizing agent may be added to the fatty acid alkyl ester rich stream and/or the glycerin rich stream after the streams have been separated from the transesterification effluent stream. Fatty acid alkyl esters are recovered from this pH adjusted stream. Subsequent to neutralization, the neutralized stream may further be purified, such as by distillation or fractionation. The process of the invention may further consist of an esterification step wherein a free fatty acid feedstock is first converted to' glycerides. The resulting glycerides are then introduced into the transesterifieation reactor. The use of the acid as neutralizing agent converts soaps, formed in the transesterification reactor, to free fatty acids. The soap forms from the action of caustic with fatty acids in the transesterification reactor. The presence of the soap makes it very difficult to effectuate phase separation between the fatty acid alkyl esters and the solution of glycerin, water, alcohol and salt. As a result, the soap emulsifies and retains much of the fatty acid alkyl esters in the glycerin rich phase. Purification of the glycerin rich phase is therefore complicated by the presence of the soap and the yield of alkyl esters is decreased. An overview of the process of the invention may be presented in FIG. 3 wherein a feedstock 1 containing free fatty acids is introduced into a glycerolysis reactor 2 with glycerin wherein the free fatty acids are converted to glycerides. The glycerides are then introduced into transesterification reactor 4 with alcohol wherein the glycerides are transesterified to form fatty acid alkyl esters and glycerin. Alcohol/alkali stream 3 may be introduced into transesterification reactor 4 as a combined mixture of alkali catalyst and alcohol, or alternatively the alkali catalyst and alcohol may be introduced into the transesterification reactor as separate streams into transesterification reactor 4. The transesterification effluent stream 4a or a portion thereof is then neutralized during neutralization/phase separation step 5, either before or after the effluent stream 5a is separated into a fatty acid alkyl ester rich stream and a glycerin rich stream. Ultimately, alcohol, glycerin and biodiesel may be refined in alcohol refining step 6, glycerin step 7 and hiodiesel refining step 8, respectively. The alcohol typically exits the system as a small portion of waste stream 9a or is recycled via flow 11 back to the transesterification reactor. Refined glycerin is isolated in technical grade glycerin stream 13 and/or may be recycled back via flow 15 to glycerolysis reactor 2. Waste stream 9b may contain some unrefined glycerin. The alkyl esters may further be refined in biodiesel refining step 8 to produce purified biodiesel stream 18 and waste stream 19 which may be useful, for example, as a burner fuel. As illustrated in FIG. 7, at least a portion of the waste stream 19 may be reintroduced into prior processes, for example as stream 351, into biodiesel refining stage 8 to further recover fatty acid methyl esters, or into the transesterification reactor 4 to transesterify glycerides into fatty acid methyl esters, or into esterification reactor 2 to esterify fatty acids. Alternatively, as illustrated in FIG. 8, at least a portion 351 of waste stream 358 may be separated into (i.) laity acid alkyl ester enriched stream 371 and (ii.) glyceride enriched stream 376 and/or tree fatty acid enriched stream 374 in separator 370. Fatty acid alkyl ester enriched stream 371 may then be re-introduced into biodiesel refining stage 8. The glyceride 376 and/or free fatty acid 374 enriched streams may then be re-introduced into the transesterification reactor 4 and/or esterification reactor 2. The process of the invention may be a continuous process. For example, a continuous process, wherein one or more of the following steps are carried out in a continuous fashion, is apparent from the description provided herein: (1) the optional conditioning of a fatty acid containing feedstock by heating, mixing and filtering; (2) continuously reacting the free fatty acids in the feedstock with glycerin in a glycerolysis or esterification reactor to produce glycerides; (3) reacting the glycerides in a transesterification reactor with alcohol to render fatty acid alkyl esters and glycerin. This reaction preferably occurs in the presence of an alkali catalyst; (4) separating (e.g., by gravitational separation of two relatively immiscible phases), fatty acid alkyl esters and glycerin from the transesterification effluent stream to yield a fatty acid alkyl ester rich stream and a glycerin rich stream; (5) purifying the fatty acid alkyl ester rich stream by distillation and/or fractionation. In a preferred embodiment, the fatty acid alkyl ester rich stream is purified by reactive distillation wherein a reaction in the distillation or fractionation column assists in the reduction of unwanted impurities such as glycerin. The purified fatty acid alkyl ester is acceptable for use as biodiesel; (6) puri lying the glycerin rich stream, preferably by use of an organic acid, such as a weak organic acid like acetic acid, formic acid or propionic acid, and recovering alcohol from the stream. The purified glycerin may then be introduced into the glycerolysis reactor; (7) purifying the wet alcohol streams resulting from steps (5) and (6) above and removing water from the streams; and (8) recycling at least a portion of the purified alcohol to the transesterification reactor for reaction with the glyceride. The process may further consist of subjecting the biodiesel stream of step (5) to further separation by a second distillation or non-evaporative separation in order to render a more purified biodiesel stream (or second purified biodiesel stream) and a second by-product fuel stream. As another option, the biodiesel stream of step (5) may further be separated in a non-evaporative separator into (i) a fatty acid alkyl ester enriched stream and (ii) a glyceride and/or free fatty acid enriched stream. Preferred non-evaporative separators for use here include freeze crystallization processes and liquid-liquid separation processes. The fatty acid alkyl ester enriched stream, resulting from this separation, may then be combined with the biodiesel stream of step (5) and then subjected to the second distillation or non-evaporative separation. The glyceride and free fatty acid enriched stream may then be re-introduced to the transesterification or esterification reactors. The feedstock, from which the biodiesel may be produced, typically contains a plurality of free fatty acids. The feedstock typically contains between from about 3 to about 100 weight percent of free fatty acids and, optionally, a fat and/or oil. Typically, the feedstock is a lipid feedstock. The free fatty acid feedstock for use in the invention may be a low-grade lipid material derived from animal fats and vegetable oils, including recycled fats and oils. For instance, the feedstock for the production of biodiesel fuel may be a grease feedstock, such as a waste grease or a yellow grease. Such low-grade lipid are very complex and typically are difficult to economically process using current state of the art processes because of their high free fatty acid levels (ranging from a few percent to 50 percent, and higher). In addition, such materials contain unprocessable material and contaminants that must be removed prior to processing or during refinement of the products. The feedstock may be first introduced into a conditioning vessel or reactor that is operative to heat, mix and/or filter the feedstock to produce a conditioned feedstock. The feedstock may then be filtered, such as by using a traveling screen. Subsequent to filtration, the concentration of free fatty acids in the conditioned lipid feedstock may be measured. Optionally, the concentration of free fatty acids in the conditioned feedstock may be continuously measured throughout the process. Measurements may be made with an in-line free fatty acid measurement device, such as a titration device or near-infrared spectrophotometer, that is operative to quantify the concentration of the free fatty acid in the conditioned feedstock. During conditioning, the feedstock may be heated to a temperature in the range of about 35°C to about 65°C, preferably between from about 55°C to about 65°C, while mixed. A uniform mixture of glyccridcs, free fatty acids and unsaponifiable materials are typically present in the conditioned feedstock. During glycerolysis, glycerin is used as a reactant to convert the free fatty acids in the feedstock to glycerides (mono-, di-, and triglyceride). Reaction of the free fatty acids in the feedstock typically occurs in the absence of a catalyst. In the glycerolysis reactor, the free fatty acid in the feedstock is mixed and continuously reacted with glycerin at an appropriate temperature and pressure to render a glycerolysis reactor effluent stream that contains generally less than about 0.5 percent by weight of free fatty acids and a plurality of glycerides. (ilycerolysis preferably occurs in the absence of both catalyst and co-solvent. The glycerin, typically a purified glycerin product, is normally added to the glycerolysis reactor at a rate that is greater than the stoichiometric amount of glycerin required for the glycerolysis reaction. The amount of glycerin introduced to the glycerolysis reactor is generally in a stoichiometric proportion of about 35 percent to about 400 percent glycerin to free fatty acid in order to render the glyceride. In a preferred embodiment, the amount of glycerin added to the reactor is at a rate in the range of about 300 percent of the stoichiometric amount of free fatty acids in the feedstock. Preferably, glycerolysis is conducted at a temperature in the range of about 150°C to about 250°C, typically from about 180°C to about 250°C, more typically from about 180°C to 230°C. The reaction typically proceeds under agitation. The reaction is further typically conducted at a pressure of about 0.1 pounds per square inch absolute to about 15 pounds per square inch absolute, more typically about 2 pounds per square inch absolute. Reaction of the free fatty acids and glycerin typically occurs in the presence of a catalyst such as ZnCli, but in a preferred embodiment is performed in the absence of a catalyst. The glycerolysis reactor effluent stream may contain less than 0.5 percent by weight of free fatty acids and a plurality of glycerides. The glycerolysis is typically a continuous reaction. The continuous reaction of the free fatty acid in the feedstock with glycerin to produce the glyceride in the glycerolysis reactor may be conducted in response to a signal from the in-line fatty acid measurement device or spectrophotometer. During glycerolysis, water is removed; the produced glycerides being essentially water-free. Water is typically continuously removed from the glycerolysis reactor as a vapor through a tractionation column or a vent in the reactor headspace. Preferably, the vapor vented from the glycerolysis reactor is fractionated to yield three streams, the first fraction having a high concentration of unsaponifiables evaporated from the feedstock that are condensed as a liquid stream, the second fraction being a liquid fraction having a high concentration of glycerin, and a vapor fraction and a third liquid fraction having a high concentration of water. The liquid fraction containing the glycerin may then be returned to the glycerolysis reactor. The glycerolysis reactor may consist of two or more continuous stirred tank reactors operated in series. The residence time of such reactors is typically from about 30 to not more than about 500 minutes, and preferably not more than 200 minutes. A plurality of glycerides contained in the glycerolysis effluent stream is reacted with an alcohol in the transesterification reactor, such as a continuous stirred tank reactor. In this reaction, the glycerides in the glycerolysis reactor effluent stream are transesterified into fatty acid alkyl esters and glycerin. Transesterification proceeds at an appropriate temperature and pressure to produce the desired transesterification reactor effluent stream. Transeslerification, which preferably is a continuous process, occurs in the presence of a base catalyst. Suitable base catalysts include such alkali catalysts as potassium hydroxide and sodium hydroxide. The alkali catalyst may be added to the transesterification reactor at a rate sufficient to catalyze the reaction. Typically, the amount of alcohol added to the transesterification reactor is from about 1 mole to 5 moles alcohol to each mole of fatty acid portion of the glycerides present in the transesterification reactor inlet stream. More typically, the ratio is about 2 moles alcohol for each mole of fatty acid portion present in the glycerides introduced into the transesterification reactor. The catalyst, typically potassium hydroxide, is added at a ratio of about 0.5% to 3% by weight catalyst to weight glycerides, more typically about 1%. Alternatively, an alkoxide, such as potassium methylate, may be added to the transesterification reactor to facilitate the base catalysis. As such, the rapid conversion of glycerides to alkyl esters may occur in the presence of caustic alkoxide. such as caustic methoxidc catalysis The transesterification reaction typically occurs at a temperature in the range of about 25°C to about 65°C, preferably from about 50°C to about 60°C, and at a pressure of about 14.5 psia to about 3,625 psia. The alcohol is normally added to the transesterification reactor at a rate that is greater than the stoichiometric amount of alcohol required for the alkali catalyzed transesterification reaction. For instance, the alcohol may be added to the transesterification reactor at a rate equal to about 200 percent of the stoichiometric amount of alcohol required for the catalyzed reaction. Preferably, multiple alcohol or catalyst additions are made to the transesterification reactor. The transesterification reactor typically contains at least two continuous stirred tank reactors that are operated in series. Each of the tank reactors typically has a residence time of about 5 minutes to about 90 minutes, typically about 60 minutes. The resulting transesterification reactor effluent stream contains a fatty acid alkyl ester and glycerin. Preferably, at least a portion of the glycerin is removed from the transesterification reactor before the plurality of glycerides is reacted with the alcohol. A plurality of the resulting fatty acid alkyl esters may then be separated from the glycerin in the transesteriflcation effluent stream. Separation into two distinct immiscible phases, i.e., a first liquid phase in which the plurality of fatty acid alkyl esters may be concentrated and a second liquid phase in which glycerin may be concentrated, is typically dependent upon the differences in densities in the two phases and employs gravitational force and/or centrifugal force. Typically, the two phases are separated at a temperature of about 25°C to about 65°C to produce the fatty acid alkyl ester rich stream and glycerin rich stream. This separation process may be a continuous operation and may be performed in a clarifier or by means of membrane filtration. In a preferred embodiment, the fatty acid alkyl ester rich stream is subjected to reactive distillation in biodiesel refining step 8 to separate the fatty acid alkyl ester rich stream into a bottoms traction, an overhead fraction (principally comprising excess alcohol) and a fatty acid alkyl ester product stream. Such separation utilizes the differences in the vapor pressures of the components of the fatty acid alkyl ester rich stream and the reactive loss of glycerin. The conditions in the distillation or fractionation column including temperature and pressure conditions, simultaneously with and in the same vessel wherein the said separation occurs, promote a chemical reaction to occur. Reactive distillation in the embodiment depicted in Fig. 6 decreases the concentration of glycerin and increases the amount of glycerides exiting the column. Thus, reactive distillation increases the efficiency of the production process. The end result of reactive distillation is that the amount of glycerin seen in the transesteriflcation effluent stream, or the first liquid phase, is greater than the total amount of glycerin which exits the distillation or fractionation column. This is attributable to the reaction of the glycerin 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 which comprises essentially the alcohol. Preferably the bottoms fraction comprises impurities having a high boiling point, unsaponifiable materials, monoglycerides, diglycerides, triglycerides and fatty acids. Preferably, the fatty acid alkyl ester distillation column or fractionation column is operated at a pressure below about 15 pounds per square inch absolute. More preferably, the cid alkyl ester distillation column or fractionation column is operated at a pressure in the range of about 0.1 pounds per square inch absolute to about 3 pounds per square inch absolute. Preferably, the fatty acid alkyl ester distillation column or fractionation column is operated at a temperature in the range of about 180°C to about 290°C, more preferably between from about 230°C to about 270°C. Preferably, the fatty acid alkyl ester distillation column or fractionation column contains a packing material. The glycerin rich second liquid phase stream may further be purified and alcohol recovered from it. The recovered alcohol is operative to produce a purified glycerin product and a (second) wet alcohol stream. In a preferred embodiment, this step employs one or more of glycerin fractionation (wherein the fractions within the glycerin rich stream are separated by distillation), phase separation (wherein the impurities that co-fractionate with glycerin are removed by immiscibility and differences in density) and glycerin polishing (wherein other impurities are removed from glycerin). The glycerin rich stream may further be subjected to phase separation wherein a fatty acid alkyl ester rich liquid phase and a glycerin rich liquid phase are separated and the two liquid phases may then be subject to purification as described in the paragraphs above. The glycerin rich stream may further be purified in a glycerin distillation or fractionation column to produce a bottoms material, a side stream and an overhead stream. Preferably, the bottoms material contains essentially waste materials; the side stream contains essentially glycerin and trace impurities; and the overhead stream contains essentially alcohol and water that is collected for further purification and recycled. Preferably, the glycerin distillation column is operated at an elevated temperature between about 180°C and about 280°C, more preferably between from about 180°C to about 230°C. The distillation column is typically operated at a reduced pressure, of below about 2 pounds per square inch absolute, typically the pressure is in the range of about 0.1 pounds per square inch absolute to about 2 pounds per square inch absolute. The glycerin rich stream may further be subjected to a decolorization column wherein colored impurities and odors are removed from the glycerin, i.e., "glycerin polishing". The decolorization column typically comprises a packed bed of activated carbon operated at a temperature in the range of about 35°C to about 200°C, preferably between from about 40°C to 100°C. The contact time is generally less than four hours. Activated carbon fines carried through the packed bed are removed by filtration. Water may further be removed from the wet alcohol streams to render purified alcohol by subjecting the wet alcohol stream to an alcohol distillation or fractionation column at a temperature in the range of about 60°C to about 110°C and at a pressure in the range of about 14 pounds per square inch absolute to about 20 pounds per square inch absolute. Preferably, this purification comprises adsorption onto molecular sieves that can then be dried and reused or distillation resulting in a bottoms product consisting mainly of water. At least a portion of the purified glycerin 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 being recycled into the transesterification reactor for reaction with glycerides. It is typically desired to neutralize the fatty acid alkyl ester and glycerin produced in the transesterification reactor. Neutralization is often required in light of the caustic conditions which character!/e transesterification. Such neutralization may occur by addition of an acid to the transesterification effluent stream or to either the fatty acid alkyl ester rich stream or glycerin rich stream after such streams are separated from the transesterification effluent stream. Suitable acid treatments include mineral or more preferably organic acid treatments. Suitable mineral acids include sulfuric acid and phosphoric acid. Reaction of the alkali catalyst with a mineral acid renders an insoluble salt that is removed from the glycerin rich stream in a solids separation operation. FIG. 4 is illustrative of the process wherein a mineral acid, such as phosphoric acid, is employed. In particular. FIG. 4 illustrates introduction of feedstock 1 containing free fatty acids into glycerolysis reactor 2 wherein the free fatty acids are converted to glycerides by eslerification. The glycerides are then introduced into transesterification reactor 4 with alcohol 3 and alkali catalyst 318 at 317 (illustrated in FIG. 7) wherein the glycerides are transesterified to form fatty acid alkyl esters and glycerin. The transesterification effluent stream 4a is first separated in l"sl phase separation 320, typically by gravitational separation techniques, into a fatty acid alkyl ester rich stream and a glycerin rich stream. Each of these streams may then be purified in 2nd phase separation 322 in accordance with the processes described herein. The neutralization acid, phosphoric acid, 324 is added either prior to 1st phase separation 320 or subsequent to lsl phase separation 320 of the transesterification effluent stream after the fatty acid alkyl ester rich stream and glycerin rich stream have been separated. Such alternative or combination ports of introduction of the acid into the process are represented by the dotted lines in FIG. 4. Unfortunately, use of phosphoric acid renders an insoluble precipitate. The formation of the insoluble precipitate mandates the use of a filter in filtration step 326 and/or a filter in filtration step 328. Suitable filters include rotary vacuum drum filters, plate and frame presses as well as belt presses. In addition to the use of a filtration unit, use of a mineral acid further requires the rinsing of the insoluble by-product salts in order to wash residual organic materials from them. Suitable solvents include Ci-Cs alcohols, such as methanol. Illustrated in FIG. 4 is the introduction of alcohol solvent 329 for use as alcohol rinse 330 which removes organic residue from the filter cake. Vacuum dry 332 is then used to remove alcohol from the filter cake and to dry the purified sail which then exits the process as waste stream 334. The solvent may then be recovered as stream 364 for reuse in the process. Preferably, the process comprises drying the insoluble salt in a dryer under conditions wherein the temperature of the dryer exceeds the boiling point of the solvent at the operating pressure of the dryer. The dryer may optionally be operated under a vacuum to improve the drying. The dryer may further include a condenser to recover the solvent for reuse. FIG. 4 further illustrates the refining of alcohol, 'glycerin and biodiesel in alcohol refinery vessel 6, glycerin refinery vessel 7 and biodiesel refinery vessel 8, respectively. The alcohol typically exits the system as byproduct stream 9a or is recycled via 11 back to transesterification reactor 4. Refined glycerin is isolated as purified glycerin 13. A portion of the glycerin stream may be recycled back as stream 15 to glycerolysis reactor 2. The alkyl esters may further be purified to produce purified biodiesel 18 or may exit the system as byproduct 19 in the form of, for example, burner fuel. It is more preferable to employ an organic acid versus a mineral acid, however. While there are inorganic acids that don't create precipitating salts upon neutralization with the transesterified stream, all suffer from serious disadvantages. For instance, hydrochloric and acid produce chlorides in the process streams which, in turn, cause undesirable corrosion of steel and stainless steel, especially at elevated temperatures. Sulfuric acid, sulfurous acid and hydrogen sulfide suffer serious disadvantages due to the presence of sulfur which increases the tendency of sulfur to exit with the final biodiesel product. This, in turn, causes potential failure of sulfur level limits and the formation of unwanted sulfur oxide in emissions from biodiesel-burning engines. Arsenic acid, chromic acid, hydrocyanic acid and hydrofluoric acid are undesirably hazardous to use and/or require unwanted additional treatment methods for the disposal of undesirable byproducts. Lastly, iodic acid does not produce undesirable precipitates, but it is economically not viable. When an organic acid is used, no insoluble salt is formed and thus it is unnecessary to subject the stream to any solids separation operation. Suitable organic acid include weak organic acids, such as formic acid, acetic acid and propionic acid. In such instances, the pH of the glycerin rich stream resulting from Iransesterification may first be adjusted below 8.0, preferably between from about 6.5 to about 7.0. FIG. 5 contrasts the inventive process wherein an organic acid 325 is used in the neutralization of the alkali catalyst versus a mineral acid. In one embodiment, organic acid is added to the transesterifi cation effluent prior to separation of the fatty acid alkyl ester rich stream form the glycerin rich stream, at a weight ratio of from about 0.1% to about 5 %, more typically about 0.9%. In another embodiment, organic acid is added to the glycerin rich stream at a weight ratio of from about 1 % to about 7%, more typically about 4%. The use of an organic acid renders the steps of filtration, rinsing of the filter cake and vacuum drying unnecessary and thus offers advantages over the use of the mineral acid. As illustrated in FIG. 7, a portion of by-product (fuel) stream 351 is shown as being directed back into biodiesel refining stage via 351A, into transesterification reactor via 351C, or into esterification reactor via 35 ID. The composition of stream 351 is not changed prior to being separated into streams 351 A, 3 5 1C and 35 ID. In contrast, in FIG. 8, a portion of by-product (fuel) stream 351 is separated in separator 370 into fatty acid alkyl ester enriched stream 371 and/or a second stream 374 enriched in free fatty acids and/or a third stream 376 enriched in glycerides. The portion of the second stream having the lower free fatty acid content is then introduced into transesterification reactor 4 and tJsgportion of stream having the higher free fatty acid content is introduced into esterification reactor 2. FIG. 9 illustrates an embodiment for biodiesel refining step 8 wherein an increased yield of biodiesel may result by the use of a second distillation reactor or non-evaporative separator. In a preferred embodiment, this second distillation reactor is one or more evaporative devices, such as wiped film evaporators or falling film evaporators known in the art. Typically, this second distillation reactor occurs in the biodiesel refining unit. Further, a separator unit may also be used to treat the by-product (fuel) stream which results from the purification of biodiesel. A system may be constructed in accordance with the teachings set forth herein for the production of biodiesel from a feedstock, such as a lipid feedstock having free fatty acids. The system may include: (1) an optional conditioning reactor which is operative to continuously convert the feedstock to a conditioned feedstock. The conditioning reactor is operative to heat, mix and filter the feedstock in order to produce a conditioned feedstock; (2) an optional system for continuously measuring the concentration of the free fatty acid in the conditioned feedstock. Suitable systems include an in-line free fatty acid measurement device which is operative to quantify the concentration of the free fatty acid in the conditioned feedstock; (3) a glycerolysis reactor wherein the free fatty acid in the feedstock is continuously reacted with glycerin to produce a glyceride. This reaction may be in response to a signal from the in-line free fatty acid measurement device; (4) a transesterification reactor for continuously reacting the glyceride with an alcohol and which is operative to convert the glyceride to a fatty acid alkyl ester and glycerin, preferably by an alkali catalyzed reaction. This reaction may proceed in response to the signal from the in line free fatty acid measurement device; (5) a separator for continuously separating the fatty acid alkyl ester from the glycerin and which is operative to produce a fatty acid alkyl ester rich stream and a glycerin rich stream. Suitable separators include a clarifier or a phase separation centrifuge which is operative to produce a (first) liquid phase in which the fatty acid alkyl ester is concentrated and a (second) liquid phase in which glycerin is concentrated. (6) a purifier for continuously purifying the fatty acid alkyl ester rich stream and recovering the alcohol from the fatty acid alkyl ester rich stream; the purifier being operative to produce a purified hiodiesel product and a first wet alcohol stream. Suitable purifiers include fractionation and distillation columns. In a preferred embodiment, the fatty acid alkyl ester rich stream is purified by reactive distillation to render biodiesel; (7) an optional evaporator separator, such as a wiped film evaporator or a falling film evaporator, for further separation of biodiesel into a fatty acid alkyl ester enriched stream and a by-product stream (fuel) stream; (8) an optional non-evaporative separator for separation of the by-product (fuel) stream into a fatty acid alkyl ester enriched stream and a free fatty acid/glyceride enriched stream; (9) a purifier for continuously purifying the glycerin rich stream and recovering alcohol from the glycerin rich stream; the purifier being operative to produce a purified glycerin product and a second wet alcohol stream. Suitable purifiers include fractionation and distillation columns, including reactive distillation; (10) a purifier for continuously purifying the wet alcohol streams that is operative to produce a purified alcohol product. Suitable purifiers include an alcohol fractionation column for treating the alcohol streams; and (11) pathways for recycling at least a portion of the purified glycerin product to the glycerolysis reactor and recycling at least a portion of the purified alcohol into the transesterification reactor for continuously reacting with the glyceride. Referring to FIG. 1, a preferred embodiment of a biodiesel production process 10 for the conversion of high free fatly acid feedstocks into biodiesel is presented. In feedstock introduction step 12, feedstock is introduced to process 10. The introduced feedstock is preferably conditioned in feedstock conditioning operation 14 wherein feedstock is heated and mixed in conditioning reactor 16; the high free fatty acid feedstock being heated and mixed to ensure a uniform mixture. The free fatty acid may be quantified, such as in an in-line free fatty acid measurement device 18, wherein the concentration of free fatty acids in the feedstock is determined by spectroscopy, titration or other suitable means. In a first separation, solid (insoluble) substances are removed in filter 24. The feedstock may include at least one free fatty acid at a concentration in the range of about 3 percent to about 97 percent by weight; moisture, impurities and unsaponifiable matter at a concentration up to about 5 percent by weight; and a remainder that includes monoglycerides, diglycerides and/or triglycerides. The feedstock may further include trap grease. Preferably, the conditioning step is carried out and produces a conditioned feedstock with a temperature in the range of about 35°C to about 250°C and more preferably in the range of about 45°C to about 65°C. In a preferred embodiment, the feedstock is heated to a temperature in the range of about 55°C to about 65°C. Preferably, the resulting conditioned feedstock is substantially free of insoluble solids. The conditioned feedstock is introduced to a glycerolysis or esterification reaction at 26 which preferably comprises glycerin addition step 28, heating step 32, glycerolysis step 34 in which free fatty acids are converted to glycerides and glycerolysis effluent cooling step 38. Preferably, glycerolysis reaction step 26 further comprises performing the glycerolysis reaction at a temperature in the range of about 150°C to about 250°C; and removing water from the environment of the glycerolysis reaction. More preferably, glycerolysis reaction step 26 further comprises using two or more continuous stirred tank reactors in series. In a preferred embodiment the free fatty acid and glycerin are continuously reacted, typically in the absence of a catalyst, in a glycerolysis reactor at a temperature of about 220°C and at a pressure of about 2 pounds per square inch absolute, in an esterification reaction to produce an effluent stream that contains less than 0.5 percent by weight of free fatty acids and a plurality of glycerides. Preferably, the purified glycerin product is continuously added to the glycerolysis reactor at a rate in the range of about 35 percent to about 400 percent of the sioichiometric amount of free fatty acids and water is continuously removed from the glycerolysis reactor as a vapor in water venting step 35 through a fractionation column that returns condensed glycerin to the glycerolysis reactor. Preferably, the reactor for glycerolysis step 34 comprises at least two continuous stirred tank reactors that are operated in series, the reactors having a combined residence time of not greater than about 400 minutes for feedstock with a 20 percent by weight free fatty acid concentration. Water is preferably removed as vapor through a fractionation column or a distillation column that returns condensed glycerin to the glycerolysis reactor. The effluent from glycerolysis reaction step 26 is introduced to alkali catalyzed transesterification reaction at 42 which preferably comprises alcohol metering step 44, catalyst metering step 46, alkoxide addition step 48 and transesterification step 50 wherein the glycerides undergo transesterification in the transesterification reactor. In transesterification step 50, glycerides are contacted with an effective amount of alcohol and an effective amount of alkali catalyst under conditions wherein the glycerides, alcohol and alkali catalyst come into substantially intimate contact. Preferably, the alkali catalyst is selected from the group consisting of sodium hydroxide and potassium hydroxide. The transesterification reaction step 42 is preferably conducted at a temperature in the range of about 20°C to about 65°C and at an absolute pressure in the range of about 14.5 psia. More preferably, transesterification reaction step 42 comprises conducting the transesterification at a temperature in the range of about 25°C to about 65°C and at an absolute pressure near atmospheric. In a preferred embodiment, the alcohol and alkali catalyst are mixed at prescribed rates prior to their addition to the transesterification reaction operation. In a preferred embodiment, transesterification reaction step 42 comprises reacting the plurality of glycerides contained in the glycerolysis effluent stream with an alcohol in the transesterification reactor. In the transesterification reactor, the plurality of glycerides are preferably mixed with the alcohol and alkali catalyst by an agitator and continuously reacted with the alcohol. Preferably, the alcohol, most preferably methanol, is added to the transesterification reactor at a rate equal to about 200 percent of the stoichiometric amount of alcohol required for the catalyzed reaction and the alkali catalyst is added to the transesterification reactor at a rate of about 0.5 percent by weight to 2.0 percent by weight of glycerides present in the glycerolysis effluent stream. More preferably, the alkali catalyst is dissolved in the alcohol prior to their introduction to the transesterification reactor. Preferably, the transesterification reactor comprises at least two continuous stirred tank reactors that are operated in series, said reactors having a combined residence time of not more than about 90 minutes. The transesterification reactor effluent stream contains a plurality of fatty acid alkyl esters and glycerin. The effluent from transesterification reaction step 42 is preferably introduced to second separation at 52 in which a light phase (for instance, specific gravity 0.79 -0.88) is separated from a heavy phase (for instance, specific gravity 0.90 - 1.20). In biodiesel purification step (operation) 58 (referenced as 8 in FIG. 3), excess methanol and high-boiling impurities are preferably separated from fatty acid alkyl esters in the light phase and the alcohol is collected for reuse. Preferably, separating the fatty acid alkyl esters from the glycerin involves using the density difference between the first light liquid phase and the second heavy liquid phase to separate them. In biodiesel purification step 56, differences in component vapor pressures are used to separate excess alcohol and high-boiling impurities from fatty acid alkyl esters in the light phase, and the alcohol is collected for reuse. In a preferred embodiment, second separation step 52 comprises separating the fatty acid alkyl esters from ihe glycerin in the transesterification effluent stream in a continuous clarifier in phase separation step 54. Preferably, in the continuous clarifier, a first light liquid phase in which the plurality of fatty acid alkyl esters are concentrated and a second heavy liquid phase in which glycerin is concentrated are continuously separated at a temperature of about 25°C to about 65°C to produce a fatty acid alkyl ester rich stream and a glycerin rich stream. Alternatively, the separation step may be a reactive distillation or fractionation column wherein the fatty acid alkyl ester and glycerin may be separated. The transesterification effluent stream entering the reactive column contains, in addition to fatty acid alkyl esters, a certain amount of glycerin, glycerides and unreacted or non-convertible lipid feedstock. In the reactive column, some of the glycerin reacts with unreacted fatty acids and/or fatty acid alkyl esters to form glycerides. In preferred embodiments, the light phase is separated in fatty acid alkyl esters purification step 56. In step 56, differences in component vapor pressures are used to separate excess alcohol and high-boiling impurities from fatty acid alkyl esters in the first liquid phase, and the alcohol is collected for reuse. Preferably, purifying the fatty acid alkyl ester rich stream step 58 further comprises using a distillation column to separate the fatty acid alkyl ester rich stream into a bottoms fraction, an fraction comprising primarily the alcohol, and a side stream fraction comprising a fatty acid alkyl ester product. Preferably, the bottoms fraction produced by the distillation column comprises impurities, unsaponifiable materials, monoglycerides, diglycerides, triglycerides and free fatty acids. Preferably, the fatty acid alkyl ester product produced by the distillation column meets ASTM specification D 6751. Preferably, the overhead fraction produced by the distillation column comprises essentially the alcohol. In preferred embodiments, the heavy phase from second separation step 52 is treated in catalyst separation step 62 comprising mineral acid addition step 64, catalyst precipitation step 66 in which the alkali catalyst is reacted with a mineral acid to produce a solid precipitate, catalyst precipitation reactor effluent filtration step 70 in which an alcohol washing step 68 occurs before the alkali salt precipitate is removed in salt recovery step 71, filtrate separation step 72 in which the precipitate-free filtrate is separated into two liquid phases, with the fatty acids and fatty acid ulkyl eslers floating to the top and the glycerin and most of the alcohol sinking to the bottom, pll neutralization step 74 in which the pH of the glycerin is increased, and free fatty acid recycling step 76. Crude glycerin may be treated in glycerin purification step 80 wherein glycerin is purified by differences in component vapor pressures. A preferred embodiment comprises distillation or fraclionation step 84 in which the alcohol and high boiling impurities are separated from the glycerin. Glycerin decolorization step 86 comprises using a packed bed of activated carbon to remove color and odor from the distilled glycerin. Preferably, in purifying the glycerin rich stream and recovering alcohol from it to produce the purified glycerin product and a wet alcohol stream, the alkali catalyst in the glycerin rich stream is reacted with a mineral acid, such as phosphoric acid or sulfuric acid, to produce an insoluble salt having fertilizer value that is removed from the glycerin rich stream in a solids separation operation and thereafter filtered and rinsed with the alcohol. The pH of the glycerin rich stream is adjusted to about neutral by adding a caustic alkali solution and then further purified in a glycerin distillation column that is operated at a temperature in the range of about 180°C to about 230°C and at a pressure below about 1 pound per square inch absolute and in a decolorization column comprising a packed bed of activated carbon operated at a temperature in the range of about 40°C to about 200°C. In a more preferred embodiment, the pH of the glycerin rich stream is adjusted to between about 6.5 and 8.0 by the addition of an acid. An organic acid, such as a weak organic acid, like acetic acid, propionic acid or formic acid, is then introduced to the glycerin rich stream. Salts present in the glycerin rich stream remain soluble. Thus, filtering and rinsing steps are unnecessary by use of the organic acid. Preferably, the wet alcohol is treated in alcohol purification step 88 in which water is removed from the wet alcohol. More preferably, the water is removed by vapor pressure differences or adsorption. In a preferred embodiment, the alcohol is purified by distillation or fractionation in alcohol distillation or fractionation step 90. In a preferred embodiment, purifying the wet alcohol stream comprises removing water from it to produce a purified alcohol product. Preferably, the wet alcohol stream is purified in an alcohol distillation column that is operated at a temperature in the range of about 60°C to about 110°C and at a pressure in the range of about 14 pounds per square inch absolute to about 20 pounds per square inch absolute. In glycerin recycling step 92, glycerin is preferably recycled to step 28 and in alcohol recycling step 94, alcohol is preferably recycled to step 44. Preferably, glycerin recycling step 92 involves recycling at least a portion of the purified glycerin product into the glycerolysis reactor for reaction with the plurality of free fatty acids in the feedstock. Preferably, the alcohol recycling step involves recycling at least a portion of the purified alcohol product into the transesterification reactor for reaction with the plurality of glycerides. The additional alcohol required for the transesterification reaction is supplied to the alkoxide tank. Biodiesel is delivered to its market in biodiesel delivery step 96 and glycerin is delivered to its market in glycerin delivery step 98. Referring to FIG. 2, a preferred embodiment of system 110 for the conversion of high free fatty acid feedstocks into biodiesel is presented. Biodiesel production system 110 preferably comprises the subsystems and reactors described below wherein the alcohol employed is methanol. In feedstock introduction subsystem 112, the feedstock is introduced to system 110. In a preferred embodiment, the feed material is composed of between 0 and 100 percent free fatty acid content, with the remainder comprising mono-, di- and triglycerides, moisture, impurities and unsaponifiables (MIU). The introduced feedstock may optionally be conditioned in feedstock conditioning subsystem 14 comprising feedstock heating and mixing vessel 16 in which the high free fatty acid feedstock is heated and mixed to ensure a uniform, homogeneous mixture with uniform viscosity. The concentration of free fatty acids in the feedstock may be measured by in-line measurement device 18. The concentration is measured continuously to allow continuous control of downstream process steps. Preferably, the feed material is heated in feedstock heating and mixing vessel 16 to ensure that all of the available lipids are liquid and that solids are suspended. Temperatures in the range of at least 35°C but not more than 200°C are adequate to melt the lipids, decrease their viscosity and allow thorough mixing of the feedstock. A jacketed stirred tank may be used to provide agitation and maintain the feedstock at increased temperature. The conditioned feedstock may then be introduced to glycerolysis reaction subsystem 26 which comprises glycerin addition apparatus 28, input heater 32, first glycerolysis reactor 1 34 and second glycerolysis reactor 136 and glycerolysis effluent cooler 38. The filtered product of step 24 is combined with glycerin and subjected to conditions that promote the glycerolysis reaction in glycerolysis reaction subsystem 126. In a preferred embodiment, these conditions include a reaction temperature between from about 150°C to about 250°C and a pressure between about 0.1 pounds per square inch, absolute (psia) and about 30 psia. A more preferred condition is a temperature of about 220°C and a pressure of about 2 psia. Glycerin is added to the filtered grease feedstock in excess of the free fatty acid molar quantity of the grease feedstock. This excess is in the range of 10 percent to 300 percent excess glycerin (from 110 percent to 400 percent of the stoichiometric amount). In this embodiment, the glycerolysis reactors used as elements 134 and 136 are configured as two heated, continuous stirred tank reactors in series. In these vessels, the mixture of glycerin and grease (containing free fatty acids) is agitated to keep the two immiscible fluids in intimate contact. In a preferred embodiment, mixing is provided by an agitator. Under these conditions, the free fatty acids are converted into glycerides (mono-, di-, or triglycerides) with the production of water. The water is vented as vapor and removed from the system together with any water that was initially present in the feedstock in water vapor vent 35. The free fatty acid content of the reactor effluent stream in this preferred embodiment of the invention can consistently be maintained at less than 0.5 percent w/w. Because of the corrosive nature of free fatty acids, the glycerolysis reactor is preferably constructed of materials resistant to organic acids. The effluent from glycerolysis reaction subsystem 126 contains mono-, di-, and triglycerides and residual fatty acids. The glycerolysis reaction effluent is introduced to alkali catalyzed transesterification subsystem 142 which preferably comprises methanol metering apparatus 144, potassium hydroxide metering apparatus 146, methoxide addition apparatus 148 and first transesterification reactor 150 and second transesterification reactor 151 in which the glycerides undergo transesterification. In transesterification reaction subsystem 142, the glycerides are transesterified with an alkali catalyst and a simple alcohol having 1 to 5 carbons. In a preferred embodiment, the alkali catalyst is potassium hydroxide and the alcohol is methanol. The residual free fatty acids are saponified consuming a molar quantity of alkali catalyst about equal to the number of moles of free fatty acid present. The transeslerification reaction is preferably catalyzed by potassium methoxide, which is formed from the addition of potassium hydroxide to methanol. The amount of potassium hydroxide added is preferably equivalent to 0.5 percent to 2.0 percent w/w of the glycerides present in the feed solution. The methanol and catalyst are combined and added to the solution of glycerides coming from the glycerolysis reactors by methoxide addition apparatus 148. A 200 percent stoichiometric excess of methanol based upon the number of moles of fatty acids available in the glycerides is added to the reaction mixture. Upon entering each transesterification reactor 150 and 151, the two-phase system undergoes vigorous mixing. Preferably, the reaction temperature is held between about 25°C and about 65°C. At this temperature, the miscibility of the phases is limited and mixing is required to achieve a high conversion rate. The residence time required is dependent on glyceride composition of the feed (between mono-, di- and triglycerides), temperature, catalyst concentration and mass transfer rate. Thus, agitation intensity is preferably considered in selecting a residence time. Typically, the residence time required for greater than (>) 99 percent conversion of glycerides to alkyl esters is 20 to 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 chosen for the reactors. The effluent from transesterification subsystem 142 may be introduced to phase separation subsystem 52 which comprise phase separation tank 54 in which a light phase (for instance, specific gravity 0.79 - 0.88) is separated from a heavy phase (for instance, specific gravity 0.90 - 1.2). The effluent streams from the phase separator are a light phase fatty acid alkyl esters comprised of methanol and alkyl esters (biodiesel), a fraction of the excess alcohol and some impurities, and a heavy phase (crude glycerin) containing glycerin, alcohol, FAAEs, soaps, alkali catalyst, a trace of water and some impurities. Phase separation unit 54 is preferably a conventional liquid/liquid separator, capable of separating of the heavy phase from the light phase. Suitable phase separation units include commercially available equipment, including continuous clarifier 54. In biodiesel purification subsystem 56, excess methanol and high-boiling impurities may be separated from the fatty acid methyl esters in the light phase in fractionation column 58 and methanol collected for reuse. Preferably, purifying the fatty acid methyl ester rich stream subsystem 56 further comprises a fatty acid alkyl ester distillation column 58 for separating the laity acid alkyl ester rich stream into a bottoms fraction, an overhead fraction comprising primarily methanol, and a side stream fraction comprising a fatty acid alkyl ester product. Preferably, the bottoms fraction produced by distillation column 58 comprises impurities, and unsaponifiable materials, monoglycerides, diglycerides, triglycerides and fatty acids. Preferably, the fatty acid methyl ester product produced by distillation column 58 in FIG. 2 meets ASTM specification D 6751. Preferably, the overhead fraction produced by distillation column 58 comprises essentially methanol. Preferably, distillation column 58 is operated under pressure below about 2 pounds per square inch absolute and at a temperature in the range of about 180°C to about 28()°C. More preferably, distillation column 58 is operated under pressure in the range of about 0.1 pounds per square inch absolute to about 2 pounds per square inch absolute and at a temperature in the range of about 180°C to about 230°C. Preferably, distillation column 58 contains high efficiency structured packing material. The heavy phase separated in phase separation tank 54 is preferably treated in catalyst separation subsystem 62 comprising a mineral acid (such as phosphoric acid) addition apparatus 64, catalyst precipitation reactor 66, catalyst precipitation reactor effluent filter 70 in which washing with methanol 68 occurs before the potassium phosphate precipitate 171 is removed from the filter, filtrate separation tank 72, pH neutralization tank and free fatty acid recycling apparatus 76. In catalyst separation subsystem 62, the crude glycerin phase is pumped to a catalyst precipitation reactor where a mineral acid 64 is added. Preferably, the amount of acid added is a molar quantity equal to the molar quantity of alkali catalyst used in the transesterification reaction. The product of the reaction is an insoluble salt that can be separated as a solid. In addition to forming an insoluble salt, the acid converts soaps formed in transesterification reaction subsystem 142 to free fatty acids. In a preferred embodiment, potassium hydroxide is used as the transesterification catalyst, and the precipitation reaction uses phosphoric acid to form monobasic potassium phosphate. This salt is not soluble in this system and can be removed by simple filtration. As the potassium phosphate salt is filtered in catalyst precipitation reactor effluent filter 70, methanol 68 is used to wash glycerin and other process chemicals off of the precipitate. The filtrate from catalyst precipitation reactor effluent filter 70 is sent to another phase separation operation where two liquid phases form and separate according to their relative specific gravities in filtrate separation tank 72. Glycerin, water, impurities and most of the methanol report to the bottom or heavy phase, while fatty acid alkyl ester, some alcohol and fatty acids report to the top, or light phase. The light phase is combined with the light phase from the previous phase separation subsystem (subsystem 52) and sent to the fractionation column 58. The heavy phase is sent to a reaction operation where any residual acid is neutralized in pH neutralization reactor 74 by adding a small amount of caustic. In a preferred embodiment, this is performed in a continuous stirred tank reactor. Following pll neutralization reactor 74, the crude glycerin phase is sent to the glycerin refining subsystem 80, where the methanol and water are separated and collected for further purification and the glycerin is separated from the high boiling impurities. In a preferred embodiment, glycerin separation is performed in glycerin distillation or fractionation column 84 a glycerin side draw. The distilled glycerin may further be treated in glycerin decolorization column 86 in which activated carbon is used to remove color and odor from the distilled glycerin. The methane I recovered from the distillation column contains trace amounts of water and is therefore considered a "wet" methanol stream that must be purified prior to reuse in the process in methanol purification subsystem 88. This "wet" methanol stream is collected and purified by distillation in methanol purification column 90 before being pumped back into the inventory storage tanks. The distilled glycerin stream is then subjected to decolorization and deodorization through activated carbon bed 86. The feed enters the column from the bottom and is allowed to flow upwards through the activated carbon bed resulting in a colorless, solventless and salt free glycerin that is >95 percent pure. Glycerin recycling pump 92 may be used to recycle glycerin to glycerin addition apparatus 28. Methanol recycling apparatus 94 is preferably used to recycle methanol to methanol metering apparatus 144. Biodiesel is then delivered to its market in biodiesel delivery vehicle 96 and glycerin is delivered to its market in glycerin delivery vehicle 98. The process may also consist of refinements to increase the yield of production of biodiesel. FIG. 7 illustrates the option of increasing the yield in the production of biodiesel by further treatment of byproduct stream 358, depending to a large extent on its relative concentration of tatty acid alkyl esters, glycerides, and free fatty acids in by-product stream 358. As illustrated, a portion of by-product stream 358 may be treated in biodiesel refining step 8. As shown in FIG. 7. fatty acid alkyl ester enriched stream 351A of by-product fuel stream 351 is redirected to biodiesel refining stage 8 for further recovery of fatty acid alkyl esters. Stream 358, when containing significant portions of glycerides, may further be introduced into transesterification reactor 4 or esterification reactor 2. As illustrated, a fraction of by-product stream 358 is introduced as stream 351C into transesterification reactor 4. Alternatively, stream 3511), when containing higher free fatty acid content is preferably introduced into esterification reactor 2. In FIG. 8, a portion of by-product stream 358, represented as stream 351, may first be separated, preferably in non-evaporative separator 370, as fatty acid alkyl ester rich stream 371 glycerides enriched rich stream 376 and/or free fatty acids enriched stream 374. The fraction containing low free fatty acid content may then be introduced as stream 376 into transesterification reactor 4 and stream 374 containing higher free fatty acid content may be introduced into esterification reactor 2. Suitable non-evaporative separation techniques that may be used are freeze crystallization, steam stripping or liquid-liquid separation. Increased yield of biodiesel may further result by the use of a second distillation reactor or non-evaporative separator in biodiesel refining stage 8. As shown in FIG. 9, a fatty acid rich stream, such as enriched stream 323 separated from the transesterification effluent stream in 1st phase separation 320, is introduced to heat exchanger 405 and introduced via pump 406 into Hash drum 410. Typical operating temperature range for flash drum 410 is from about 60°C to about 205°C, more typically about 140°C, and typical operating pressure is from about 1 pound per square inch absolute to about 15 pounds per square inch absolute, more typically about 5 pounds per square inch absolute. Vapor 412 is removed and the liquid stream 411 is then pumped through pump 415 into distillation column 420. 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 in vapor form, principally as excess alcohol, as stream 442. Condensate 441A exiting heat exchanger 440 exits the system and liquid stream 441B re-enters the distillation column. The bottoms fraction 421 from distillation column 420 is principally the fatty acid alkyl ester rich stream and may then be introduced into reboiler 430 where it is either further separated as vapor stream 432 in distillation column 420 or exits as biodiesel stream 431 A. Biodiesel stream 431A consists principally of fatty acid alkyl esters, glycerides and a trace amount of glycerin and, depending on the acidity upstream, some fatty acids. This stream may further be subjected to a second distillation in distillation column 450, via holding tank 440, to render purified biodiesel stream 350C and by-product (fuel) stream 350A. In a preferred embodiment, distillation column 450 is either one or more wiped film evaporators or falling film evaporators known in the art. The temperature in the second distillation column 450 is approximately the same as the temperature in distillation column 420. In an alternative embodiment, shown in FIG. 10, a portion of by-product (fuel) stream 350A may be re-introduced to second distillation column 450 via holding tank 440. The second distillation procedure may occur in one or more distillation columns. For instance, a single wiped film evaporator or falling film evaporator may be used. Further, multiple wiped film evaporators or falling film evaporators in parallel or series may be used. Residence time of the biodiesel stream in the wiped film evaporator and failing film evaporator is generally short. The wiped film evaporator consists of internal rotating distributor plates which serve to evenly disperse the biodiesel at the top of the heated plate of the evaporator to the interior surfaces of a heated cylindrical shell. Wiper blades then spread, agitate and move the biodiesel downwards along the heated shell in rapid time while fatty acid alkyl esters are quickly evaporated and re-condensed on a cooled surface, typically at the center of the evaporator. With this particular configuration, the purified biodiesel stream then exits the bottom of the center of the evaporator, and the byproduct (fuel) stream exits the bottom of the outer perimeter of the evaporator. The falling film evaporator consists of an outer shell filled with steam or other heating media and vertical, parallel tubes through which the biodiesel falls. The flow of biodiesel is controlled such thai the biodiesel creates a film along the inner tube walls, which progresses downwards while the biodiesel is selectively evaporated from the liquid. Separation between biodiesel vapors and the residual liquid typically consisting of a mixture of glycerides, fatty acids and some unevaporated fatty acid alkyl esters occurs in the tubes. The biodiesel vapor is liquefied in a cooled condenser and recovered. As in distillation column 420, these second distillation columns 450, are typically operated at a pressure below about 250 torr absolute and at a temperature in the range of about 150°C to about 320°C. More preferably, distillation column 450 is operated at a pressure in the range of about 0.1 torr absolute to about 2 torr absolute and at a temperature in the range of about 180°C to about 230°C. FIG. 11 presents a further embodiment of the invention wherein the by-product (fuel) stream 350A is introduced to separator 370. Separator 370 is preferably a non-evaporative separator. A fatty acid alkyl ester enriched stream 371 may be separated from a stream 372 enriched in glycerides and/or tree fatty acids in separator 370. The fatty acid alkyl ester enriched stream 371 may then be re-introduced to second distillation column 450 via holding tank 440 for separation into purified biodiesel. Stream 372 enriched in glycerides and/or free fatty acids may then be re-introduced into transesterification reactor 4 and esterification reactor 2. FIG. 12 presents another embodiment wherein the fatty acid alkyl ester rich stream 371 may be branched and introduced as stream 371A with the purified biodiesel stream 350C. In addition, a portion of stream 371 may be re-introduced to the second distillation column 450. Further, FIG. 12 illustrates the option of introducing either a portion or all of biodiesel stream 431 A, as 452, from first distillation column 420 into separator 370 for separation into a fatty acid alkyl ester rich stream and a glyceride and/or free fatty acid rich stream 372. The glyceride and/or free i'atty acid enriched stream 372 may then be re-introduced into the transesterification reactor 4 and/or esterification reactor 2. With respect to the above description then, it is to be realized that the optimum dimensional 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 all equivalent relationships to those illustrated in the drawings and described in the specification are intended to be encompassed by the present invention. Therefore, the foregoing 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 shown and described, and accordingly, all suitable modifications and equivalents may be resorted to, falling within the scope of the invention. Examples Example No. 1 Rendered yellow grease with a free fatty acid concentration of 20 percent by weight and 2 percent moisture, impurities and unsaponifiables (MIU) was fed to continuous stirred tank glycerolysis reactors at 100 pounds per minute (Ibs/min). The grease was filtered and titrated internmittently as it was fed to the glycerolysis reactor. Glycerin was added at a rate of 13 Ibs/min. The temperature of the grease and glycerin mixture was raised to 210°C as it was fed into the first of the glycerolysis continuous stirred tank reactors. In the reactor, the pressure was reduced to 2 psia and the temperature was maintained at 210°C. The vessel was fitted with a high intensity agitator to keep the immiscible liquids in contact. Water vapor produced by the reaction was removed through vents in the reactor headspace. The residence time in each of the glycerolysis reactors was 2.5 hours. The conversion of fatty acids to glycerides in the first vessel was 85 percent. The fatty acid concentration leaving the second reactor was maintained at 0.5 percent w/w. The product from the glycerolysis reactors was cooled to 50°C and fed continuously to the transesterification reactors in which a solution of potassium hydroxide in methanol was added. The potassium hydroxide was added at a rate of 1.1 Ibs/min and mixed with 22 Ibs/min of methanol. The transesterification took place in two continuous stirred tank reactors in series, each with a two-hour residence time. The transesterified product was then fed to a phase separation tank where the majority of the fatty acid methyl esters, a small amount of unreacted glycerides and a small concentration of the unreacted methanol floated to the top. The glycerin, the majority of the unreacted methanol. some fatty acid methyl esters, potassium hydroxide and soaps sank to the bottom. The bottom, or heavy phase was sent to an acidification reactor where the potassium hydroxide catalyst added in the transesterification step was reacted with 1.96 Ibs/min phosphoric acid. The soaps converted to free fatty acids and the potassium hydroxide was neutralized. The product of this acidification was monobasic potassium phosphate, which was not soluble in this system. The monobasic potassium phosphate precipitate was filtered out and the filtrate was fed to a second phase separation tank where the fatty acid methyl esters and free fatty acids present in the filtrate floated to the top and the glycerin and methanol sank to the bottom. The top, or light, phase was mixed with the light phase from the first phase separation tank and fed to the 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 glycerin fractionation column. The glycerin fractionation column recovered 10 Ibs/min of methanol and 18 Ibs/min of glycerin. The glycerin produced was more than 95 percent pure with non-detectable CQUKcentrations of salts and methanol. This glycerin stream was split into two streams: 13 Ibs/min was recycled back to the glycerin feed tank for the glycerolysis reaction and 5 Ibs/min was pumped through the decolorization column and collected for market. The two light phase streams were fed to the fatty acid methyl ester fractionation column where 2 Ibs/rnin of methanol was recovered and 92 Ibs/min of fatty acid methyl esters meeting ASTM D 6751-02 (Standard Specification for Biodiesel Fuel (B100) Blend Stock for Distillate Fuels) were produced. Example No. 2 Fancy bleachable inedible tallow with a free fatty acid concentration of 4 percent by weight and 0.5 percent MIL) (moisture, impurities and unsaponifiables) was fed to a continuous stirred tank reactor at 100 Ibs/min. The grease was filtered and titrated continuously as it was fed to the glycerolysis reactors. Glycerin was added at a rate of 2.6 Ibs/min. The temperature of the grease and glycerin mixture was raised to 210°C as it was fed into the first of the glycerolysis continuous stirred tank reactors. In the reactor the pressure was reduced to 2 psia and the temperature was maintained. The vessel was fitted with an agitator to keep the immiscible liquids in contact. Water vapor produced by the reaction was removed through vents in the reactor headspace. The residence time in each of the glycerolysis reactors was 2.5 hours. The conversion of fatty acids to glycerides in the first vessel was 92 percent. The fatty acid concentration leaving the second reactor was maintained at 0.5 percent by weight. The product from the glycerolysis reactors was cooled to SOT and fed to the transesterification reactors in which a solution of potassium hydroxide in methanol was added. The potassium hydroxide was added at a rate of 1.0 Ibs/min and mixed with 22 Ibs/min of methanol. The transesterification took place in two continuous stirred tank reactors in series, each with a two-hour residence time. The transesterified product was then fed to a phase separation tank where the majority of the fatty acid methyl esters and a small concentration of the unreacted methanol floated to the top. The glycerin, the majority of the unreacted methanol, some fatty acid methyl esters, potassium hydroxide and soaps sank to the bottom. The bottom, or heavy phase was sent to an acidification reactor where the potassium hydroxide catalyst added in the transesterification operation was reacted with 1.79 Ibs/min phosphoric acid. The soaps converted back to free fatty acids and the potassium hydroxide was neutralized. The product of this acidification was monobasic potassium phosphate, which was not soluble in this system. The monobasic potassium phosphate precipitate was filtered out 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 the glycerin and methanol sank to the bottom. The top, or light, phase was mixed with the light phase from the first phase separation tank and fed to the fatty acid methyl ester fractionation column. The pH of the heavy phase was adjusted to 7.8 with 0.1 Ibs/min potassium hydroxide and fed to the glycerin fractionation column. The glycerin fractionation column recovered 10 Ibs/min of methanol and 10.2 Ibs/min of glycerin. The glycerin produced was more than 95 percent pure with non-detectable concentrations of salts and methanol. The glycerin stream was split into two streams: 2.6 Ibs/min was recycled back to the glycerin feed tank for the glycerolysis reaction and 7.6 Ibs/min was collected for market. The two light phase streams were fed to the fatty acid methyl ester fractionation column in which 2.1 Ibs/min of methanol was recovered and 93 Ibs/min of fatty acid methyl esters meeting ASTM 1) 6751-02 (Standard Specification for Biodiesel Fuel (B100) Blend Stock for Distillate Fuels) was produced. Example No. 3 Degummed, food-grade soybean oil with a free fatty acid concentration of 0.5 percent by weight and 0.5 percent M1U (moisture, impurities and unsapom'fiables) was fed to a conditioning chamber at 100 Ibs/min. The grease was filtered and titrated continuously as it was transferred from the feedstock conditioner. Due to the low concentration of free fatty acids, the glycerolysis section of the process was bypassed when using this feedstock. The fatty acid concentration entering the transesterification reactors was 0.5 percent by weight. The potassium hydroxide was added at a rate of 1.0 Ibs/min and mixed with 22 Ibs/min of methanol. The transesterification took place in two continuous stirred tank reactors in series, each with a two-hour residence time. The transesterified product was then fed to a phase separation tank where the majority of the fatty acid methyl esters and a small concentration of the unreacted methanol floated to the top. The glycerin, the majority of the unreacted methanol, some fatty acid methyl esters, potassium hydroxide and soaps sank to the bottom. The bottom, or heavy, phase was sent to an acidification reactor where the potassium hydroxide catalyst added in the transesterification operation was reacted with 1.76 Ibs/min phosphoric acid. The pH of the solution was decreased, and the product of this acidification was monobasic potassium phosphate, which was not soluble in this system. The precipitate was filtered out at 2.2 Ibs/min and the filtrate was fed to a phase separation tank in which the fatty acid methyl esters and free fatty acids floated to the top and the glycerin and methanol sank to the bottom. The top, or light, phase was mixed with the light phase from the first phase separation tank and fed to the fatty acid methyl ester fractionation column. The heavy phase was transferred to another tank and the pH was adjusted to 7.4 with 0.1 Ibs/min potassium hydroxide. Then, the glycerin/methanol mixture was fed to the glycerin fractionation column. The glycerin fractionation column recovered 10 Ibs/min of methanol and 8.5 Ibs/min of glycerin. The glycerin produced had a purity greater than 95 percent with non-detectable concentrations of salts and methanol. The glycerin was collected for market. The two light phase streams were fed to the fatty acid methyl ester fractionation column where 2.1 Ibs/min of methanol was recovered and 93 Ibs/min of fatty acid methyl esters meeting ASTM D 6751-02 (Standard Specification for Biodiesel Fuel (B100) Blend Stock for Distillate Fuels) were produced. Example No. 4 Rendered trap grease with a free fatty acid concentration of 68 percent by weight and 5% M1U (moisture, impurities and unsaponifiables) was fed to the invention at 100 Ibs/min. The grease was filtered and titrated continuously as it was fed to the glycerolysis reactors. Glycerin was added at a rate of 44 Ibs/min. The temperature of the grease and glycerin mixture was raised to 210°C as it was fed into the first of the glycerolysis continuous stirred tank reactors. In the reactor, the pressure was reduced to 2 psia and the temperature was maintained. Water vapor by the reaction was removed through vents in the reactor headspace. The residence time in each of the glycerolysis reactors was 3.5 hours. The conversion of fatty acids to glycerides in the first vessel was 87 percent. The fatty acid concentration leaving the second reactor was maintained at 0.5 percent by weight. The product from the glycerolysis reactors was cooled to 50°C and fed to the transesterification reactors where a solution of potassium hydroxide in methanol was added. The potassium hydroxide was added at a rate of 1.4 Ibs/min and mixed with 21 Ibs/min of methanol. The transesterification took place in two continuous stirred tank reactors in series, each with a two-hour residence time. The transesterified product was then fed to a phase separation tank where the majority of the fatty acid methyl esters and 10 percent of the unreacted methanol floated to the top and the glycerin, the majority of the unreacted methanol, some fatty acid methyl esters, potassium hydroxide and soaps sank to the bottom. The bottom, or heavy, phase was sent to an acidification reactor where the potassium hydroxide catalyst added in the transesterification operation was reacted with 2.45 Ibs/min phosphoric acid. The soaps converted back to free fatty acids and the potassium hydroxide was neutralized. The product of this acidification was monobasic potassium phosphate, which was not soluble in this system. The monobasic potassium phosphate precipitate was filtered out at 3.1 Ibs/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 the glycerin and methanol sank to the bottom. The top, or light, phase was mixed with the light phase from the first phase separation tank and fed to the fatty acid methyl esters fractionation column. The pH of the heavy phase was adjusted back to 7.3 with 0.14 Ibs/min potassium hydroxide and fed to the glycerin fractionation column. The glycerin fractionation column recovered 10 Ibs/min of methanol and 40 Ibs/min of glycerin. The glycerin produced had a purity greater than 95 percent with non-detectable concentrations of salts and methanol. This glycerin stream was recycled back to the glycerin feed tank for the glycerolysis reaction and an additional 4 Ibs/min of fresh glycerin was added to the glycerin feed tank to provide enough glycerin feed for the glycerolysis reaction. The two light phase streams were fed to the fatty acid methyl ester fractionation column where 2.1 Ibs/min of methanol was recovered and 91 Ibs/min of fatty acid methyl esters meeting A$TM D 6751-02 (Standard Specification for Biodiesel Fuel (B100) Blend Stock for Distillate Fuels) were produced. Eixample No. 5 Rendered brown grease with a free fatty acid concentration of 37 percent by weight and 5 percent MIU (moisture, impurities and unsaponifiables) was fed to the invention at 100 Ibs/min. The grease was filtered and titrated continuously as it was fed to the glycerolysis reactors. Glycerin was added at a rate of 24 Ibs/min. The temperature of the grease and glycerin mixture was raised to 21()"C as it was fed into the first of the glycerolysis continuous stirred tank reactors. In the reactor, the pressure was reduced to 2 psia and the temperature was maintained. The vessel was fitted with an agitator to keep the immiscible liquids in contact. Water vapor produced by the reaction was removed through vents in the reactor headspace. The residence time in each of the glycerolysis reactors was 3.0 hours. The conversion of fatty acids to glycerides in the first vessel was 90 percent. The fatty acid concentration leaving the second reactor was maintained at 0.5 percent by weight. The product from the glycerolysis reactors was cooled to 50°C and fed to the transesterification reactors where a solution of potassium hydroxide in methanol was added. The potassium hydroxide was added at a rate of 1.2 Ibs/min and mixed with 21 Ibs/min of methanol. The transesterification took place in two continuous stirred tank reactors in series, each with a two-hour residence time. The transesterified product was then fed to a phase separation tank where the majority of the fatty acid methyl esters and 10 percent of the unreacted methanol floated to the top. The glycerin, the majority of the unreacted methanol, some fatty acid methyl esters, potassium hydroxide and soaps sank to the bottom. The bottom, or heavy, phase was sent to an acidification reactor where the potassium hydroxide catalyst added in the transesterification was reacted with 2.13 Ibs/min phosphoric acid. The soaps converted back to free fatty acids and the potassium hydroxide was neutralized. The product of this acidification was monobasic potassium phosphate, which is not soluble in this system. The monobasic potassium phosphate precipitate was filtered out at 2.7 Ibs/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 the glycerin and methanol sank to the bottom. The top, or light, phase was mixed with the light phase from the first phase separation tank and fed to the fatty acid methyl ester fractionation column. The pH of the heavy phase was adjusted to 7.5 with 0.12 Ibs/min potassium hydroxide and fed to the glycerin fractionation column. The glycerin fractionation column recovered 10 Ibs/min of methanol and 25.2 Ibs/min of glycerin. The glycerin produced had a purity greater than 95 percent with non-detectable concentrations of salts and methanol. This glycerin stream was split into two streams: 24 Ibs/min was recycled back to the glycerin feed tank for the glycerolysis reaction, and 1.2 Ibs/min was collected for market. The two light phase streams were fed to the fatty acid methyl ester fractionation column where 2.0 Ibs/min of methanol was recovered, and 89.8 Ibs/min of fatty acid methyl esters meeting AS I'M I) 6751-02 (Standard Specification for Biodiesel Fuel (B100) Blend Stock for Distillate Fuels) were produced. i-xample No. 6 A feedstock containing about 0.3 weight percent of free fatty acids and about 99.3 weight percent of glycerides (the remainder being water and insoluble and unsaponifiable solids), at a How rate of about 40.9 pounds per hour, was heated to SOT and added to a solution of potassium hydroxide (1 percent of the feedstock flow on a weight basis) in methanol (stoichiometric ratio of 2:1 methanol:bound fatty acids in glycerides). The transesterification took place in a single continuous stirred tank reactor with a ten-hour residence time. The transesterification effluent stream flow rate was approximately 50.3 pounds per hour and consisted of approximately 79 weight percent of fatty acid methyl esters, 8 weight percent glycerin, 9 weight percent methanol, 1.6 weight percent glycerides, with the remainder being water, insoluble and unsaponifiable solids, and soaps. This stream was separated in a flow-through separator into a light phase stream and a heavy phase stream, the light phase stream having a flow of 41.5 pounds per hour and a composition of approximately 94.26 weight percent fatty acid methyl esters, 5.6 weight percent methanol, 0.09 weight percent glycerides and 0.05 weight percent free glycerin. Free glycerin concentrations in this and the other samples in this example were determined using an enzyme assay solution provided by Sigma-Aldrich, Inc. of St. Louis, MO in a kit with product code BQP-02. With this kit, free glycerin was measured by coupled, enzymatic reactions that ultimately produce a quinoneimine dye that shows an absorbance maximum at 540 nm. The absorbance peak was measured using a Bausch & Lomb Spectronic 20 spectrophotometer. The light phase stream was analyzed for glycerin and found to contain approximately 490 ppm glycerin by weight. The light phase stream was introduced into a reactive distillation column maintained at 260°C at a pressure of 150 mmHg. The overhead vapor stream from the column was condensed, producing a liquid stream with a flow rate of about 2.1 pounds per hour, consisting primarily of methanol with a glycerin content of 135 ppm. The bottoms liquid stream, having a flow rate of approximately 39.3 pounds per hour, consisted of approximately 98.5 weight percent fatly acid methyl esters, 1.5 weight percent glycerides, and only 3 ppm glycerin. The reactive distillation referenced in this paragraph is schematically displayed as FIG. 6. The gravimetric flow rates calculated using these analyses of free glycerin in the feed to the column versus in the overhead and bottoms streams indicated that about 98 percent of the glycerin was reacted into other moieties in the distillation column rather than simply flowing either to the overhead or bottoms streams. This bottoms liquid stream was further refined to produce a biodiesel stream of fatty acid methyl esters meeting ASTM D 6751-06 SI5 (Standard Specification for Biodiesel Fuel (B100) Blend Stock for Distillate Fuels). 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 concepts of the invention. We Claim:- 1. A process for the production of purified bio diesel from glycerides comprising: (A) reacting glycerides with at least one alcohol in the presence of an alkali catalyst to produce a fatty acid alkyl ester stream; (B) purifying the fatty acid alkyl ester stream by reactive distillation; and (C) producing a bio diesel stream therefrom. 2. The process as claimed in Claim 1, wherein the bio diesel stream produced in step (C) is further subjected to a second distillation or non-evaporative separation to render a purified second bio diesel stream and a by-product fuel stream. 3. The process as claimed in Claim 2, wherein the separation of the biodiesel stream of step (C) into a purified second biodiesel stream and a by-product fuel stream occurs in at least one wiped film evaporator or at least one falling film evaporator. 4. The process as claimed in any of Claims 1 to 3, wherein the glycerides of step (A) are produced by converting fatty acids in a feedstock. 5. The process as claimed in Claim 4, wherein the feedstock is conditioned to remove solids prior to converting the feedstock to glycerides. 6. The process as claimed in Claim 4 or 5, wherein the glycerides of step (A) are prepared by mixing the feedstock and glycerin at a temperature from 35°C to 65°C. 7. The process as claimed in any of Claims 1 to 5, wherein step (A) comprises reacting the glycerides with the at least one alcohol in the presence of an alkali catalyst to produce a fatty acid alkyl ester stream containing glycerin and fatty acid alkyl esters. 8. The process as claimed in Claims 1 to 7, wherein the fatty acid alkyl ester stream in step (B) is separated by reactive distillation into a fatty acid alkyl ester rich stream and a glyceride rich stream. 9. The process as claimed in any of Claims 1 to 8, wherein the product of step (A) is separated into a fatty acid ester rich stream and a glycerin rich stream and further wherein an organic acid is introduced to the glycerin rich stream to produce a purified glycerin rich stream therefrom. 10. The process as claimed in Claim 9, wherein the organic acid is a weak organic acid selected from the group consisting of acetic acid, formic acid and propionic acid. 11. The process as claimed in Claim 10, wherein the weak organic acid is acetic acid.

Documents:

1744-DEL-2007-Abstract-(13-04-2012).pdf

1744-del-2007-abstract.pdf

1744-DEL-2007-Assignment-(07-07-2010).pdf

1744-del-2007-Assignment-(14-05-2012).pdf

1744-del-2007-Assignment-(15-06-2012).pdf

1744-DEL-2007-Claims-(13-04-2012).pdf

1744-del-2007-claims.pdf

1744-DEL-2007-Correspondence Others-(13-04-2012).pdf

1744-del-2007-Correspondence Others-(14-05-2012).pdf

1744-del-2007-Correspondence Others-(15-06-2012).pdf

1744-DEL-2007-Correspondence-Others-(07-07-2010).pdf

1744-del-2007-correspondence-others-1.pdf

1744-del-2007-correspondence-others.pdf

1744-del-2007-description (complete).pdf

1744-DEL-2007-Drawings-(13-04-2012).pdf

1744-del-2007-drawings.pdf

1744-DEL-2007-Form-1-(07-07-2010).pdf

1744-DEL-2007-Form-1-(13-04-2012).pdf

1744-del-2007-form-1.pdf

1744-del-2007-form-18.pdf

1744-DEL-2007-Form-2-(07-07-2010).pdf

1744-del-2007-form-2.pdf

1744-DEL-2007-Form-3-(13-04-2012).pdf

1744-del-2007-form-3.pdf

1744-DEL-2007-Form-5-(13-04-2012).pdf

1744-del-2007-form-5.pdf

1744-DEL-2007-GPA-(07-07-2010).pdf

1744-del-2007-Petition-137-(14-05-2012).pdf