WO2009039151A1 - Procédé de fabrication de biodiesel à point de trouble réduit - Google Patents

Procédé de fabrication de biodiesel à point de trouble réduit Download PDF

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Publication number
WO2009039151A1
WO2009039151A1 PCT/US2008/076630 US2008076630W WO2009039151A1 WO 2009039151 A1 WO2009039151 A1 WO 2009039151A1 US 2008076630 W US2008076630 W US 2008076630W WO 2009039151 A1 WO2009039151 A1 WO 2009039151A1
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Prior art keywords
alkanol
biodiesel
catalyst
transesterification
glycerin
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PCT/US2008/076630
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English (en)
Inventor
Donald Leroy Bunning
Louis A. Kapicak
Thomas Arthur Maliszewski
David James Schreck
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Best Energies, Inc.
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Publication of WO2009039151A1 publication Critical patent/WO2009039151A1/fr

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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10L—FUELS NOT OTHERWISE PROVIDED FOR; NATURAL GAS; SYNTHETIC NATURAL GAS OBTAINED BY PROCESSES NOT COVERED BY SUBCLASSES C10G, C10K; LIQUEFIED PETROLEUM GAS; ADDING MATERIALS TO FUELS OR FIRES TO REDUCE SMOKE OR UNDESIRABLE DEPOSITS OR TO FACILITATE SOOT REMOVAL; FIRELIGHTERS
    • C10L1/00—Liquid carbonaceous fuels
    • C10L1/02—Liquid carbonaceous fuels essentially based on components consisting of carbon, hydrogen, and oxygen only
    • C10L1/026—Liquid carbonaceous fuels essentially based on components consisting of carbon, hydrogen, and oxygen only for compression ignition
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10L—FUELS NOT OTHERWISE PROVIDED FOR; NATURAL GAS; SYNTHETIC NATURAL GAS OBTAINED BY PROCESSES NOT COVERED BY SUBCLASSES C10G, C10K; LIQUEFIED PETROLEUM GAS; ADDING MATERIALS TO FUELS OR FIRES TO REDUCE SMOKE OR UNDESIRABLE DEPOSITS OR TO FACILITATE SOOT REMOVAL; FIRELIGHTERS
    • C10L1/00—Liquid carbonaceous fuels
    • C10L1/10—Liquid carbonaceous fuels containing additives
    • C10L1/14—Organic compounds
    • C10L1/18—Organic compounds containing oxygen
    • C10L1/19—Esters ester radical containing compounds; ester ethers; carbonic acid esters
    • C—CHEMISTRY; METALLURGY
    • C11—ANIMAL OR VEGETABLE OILS, FATS, FATTY SUBSTANCES OR WAXES; FATTY ACIDS THEREFROM; DETERGENTS; CANDLES
    • C11C—FATTY ACIDS FROM FATS, OILS OR WAXES; CANDLES; FATS, OILS OR FATTY ACIDS BY CHEMICAL MODIFICATION OF FATS, OILS, OR FATTY ACIDS OBTAINED THEREFROM
    • C11C3/00—Fats, oils, or fatty acids by chemical modification of fats, oils, or fatty acids obtained therefrom
    • C11C3/003—Fats, oils, or fatty acids by chemical modification of fats, oils, or fatty acids obtained therefrom by esterification of fatty acids with alcohols
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02E—REDUCTION OF GREENHOUSE GAS [GHG] EMISSIONS, RELATED TO ENERGY GENERATION, TRANSMISSION OR DISTRIBUTION
    • Y02E50/00—Technologies for the production of fuel of non-fossil origin
    • Y02E50/10—Biofuels, e.g. bio-diesel

Definitions

  • This invention pertains to processes for making biodiesel from fats and oils that has lower cloud point in an economically attractive and energy efficient manner.
  • the cloud point of biodiesel is the temperature at which crystals first appear in the biodiesel when it is cooled.
  • the higher the cloud point temperature the higher the temperature at which the biodiesel gels. For cold weather use, biodiesel needs to remain flowing and accordingly low cloud points are desired.
  • the cloud point for a biodiesel is determined by the oil or fat source from which the biodiesel is made. Those feed sources that provide saturated fatty acids or glycerides yield biodiesel having higher cloud points than feed sources having unsaturated fatty acids or glycerides. [0004]
  • Additives are commercially available to retard the formation of solid crystals in biodiesel. However, the use of additives increases the cost of the biodiesel.
  • Another approach is to provide a blend of a biodiesel having a higher cloud point with that having a lower cloud point. This approach requires that multiple biosources for oils and fats be available.
  • methanol have a higher reaction rate than higher alkanols, but also methanol is relatively easily recovered from biodiesel and recycled due to its boiling point being lower than that of water and due to the fact that no azeotrope is formed with water.
  • biodiesel from higher alkanol.
  • the processes of this invention convert a biodiesel comprising lower alkyl ester, preferably methyl ester, to a biodiesel comprising higher alkyl ester through an alkanol interchange.
  • the processes of this invention enable biodiesel to facilely be made from not only primary alkanol but also secondary alkanol which beneficially provides low cloud point temperatures.
  • the processes of this invention since they do not directly involve the esterification or transesterif ⁇ cation processes to make the lower alkyl ester biodiesel, especially methyl ester biodiesel, enable the efficiencies of such processes to be maintained and avoid additional complications such as adversely affecting reaction rates, equilibria points in the equilibrium-limited esterification and transesterification reactions, and refining. For instance, by using methanol, avoidance of the formation of azeotropes with water facilitates removal of methanol for recycle. Also, by the processes of this invention, alkanol that may be more readily available to the producer than methanol, e.g., ethanol in a farming region, can be used.
  • the volume of biodiesel is increased, yet the energy density of the alkanol-exchanged biodiesel is comparable to that of the biodiesel feed.
  • the increase in volume of the biodiesel may at least partially offset the additional cost.
  • any cloud point lowering may offset costs associated with additives for lowering cloud point.
  • the biodiesel provided by the processes of this invention can provide a biodiesel having a performance in a diesel engine at least comparable to petroleum diesel with the advantages of essentially no sulfur content and higher lubricity. In heating oil applications, the biodiesel of this invention provides a higher heating value than methyl biodiesel.
  • One broad aspect of the processes of this invention for treating biodiesel comprising lower alkyl ester of fatty acid to lower cloud point comprises: a. contacting the biodiesel with higher molecular weight alkanol under transesterification conditions comprising the presence of a catalytically effective amount of transesterification catalyst to provide lower alkanol and an alkanol-exchanged biodiesel containing ester of the higher molecular weight alkanol; and b. removing the lower alkanol from the alkanol-exchanged biodiesel.
  • the lower alkyl ester of the biodiesel feed is methyl or ethyl alkyl ester of fatty acid.
  • a preferred biodiesel comprising lower alkanol is biodiesel in which the lower alkyl is methyl.
  • the biodiesel can economically be made from glycerides by transesterification with methanol. Methanol can be removed from the alkanol replacement transesterification at lower temperatures.
  • steps (a) and (b) are conducted simultaneously as the transesterification is an equilibrium-limited reaction.
  • the alkanol exchange transesterification can be conducted in two or more reaction stages, preferably with lower alkanol being removed between stages. Where one reaction stage is used and the transesterification reaction has not achieved equilibrium, a portion of the effluent containing partially transesterified biodiesel can be recycled to the reaction stage
  • the alkanol replacement process may be continuous, semi-continuous or batch.
  • the higher alkanol is provided in a molar ratio to the lower alkyl ester in the feed biodiesel of between about 0.2:1 and 10:1, although in the broad aspects of the invention higher and lower molar ratios can be used.
  • a molar ratio to the lower alkyl ester in the feed biodiesel of between about 0.2:1 and 10:1, although in the broad aspects of the invention higher and lower molar ratios can be used.
  • the amount of alkanol exchange will depend upon, for example, the sought extent of cloud point reduction as well as process economics in light of availability and relative costs of the lower and higher alkanols.
  • the higher alkanol has at least one more carbon atom than the lower alkyl group of the biodiesel feed.
  • the higher alkanol comprises primary or secondary alkanol having 2 to 8 carbon atoms.
  • the preferred higher alkanol will depend upon the sought objectives of the producer.
  • ethanol may be a desired higher alkanol where ethanol is more cost effective than methanol for biodiesel.
  • higher alkanols are used with steric configurations suitable to achieve the sought reduction.
  • alkanols typically have at least three carbon atoms and are preferably branched.
  • the preferred higher alkanols include ethanol, propanol, isopropanol, butanol, iso-butanol, n-pentanol, isopentanol, and 2-ethyl hexanol.
  • the higher alkanol has a steric configuration sufficient and is reacted in an amount sufficient to lower the cloud point of the biodiesel by at least 2°C, and sometimes at least 4°C.
  • the source of the biodiesel feed is not critical to the broad aspects of this invention.
  • the biodiesel used for the feed may be derived from any suitable oil or fat.
  • the biodiesel feed may be a refined biodiesel feed, e.g., one which is substantially devoid of catalyst and alkanol. If a solid catalyst is used for transesterification, stripping of the lower alkanol from the alkanol-exchanged biodiesel may provide a refined biodiesel. However, with soluble transesterification catalysts, a subsequent water washing, with neutralization if required, can remove the catalysts. Drying would thus be required to lower the moisture content of the biodiesel, usually to less than about 0.05 volume percent, to meet specifications.
  • the processes of this invention are also useful in treating crude biodiesel, that is, biodiesel containing alkanol, catalyst and potentially glycerin, soaps of fatty acids and water, and partially converted biodiesel, i.e., a transesterification product which contains at least about 2, say, up to about 10, mass percent glycerides.
  • the catalyst in the crude biodiesel may serve as a portion or all of the transesterification catalyst for the alkanol exchange.
  • the alkanol exchange can also serve to further convert glycerides to alkyl esters.
  • the alkanol exchanged biodiesel can then be refined as would the crude biodiesel of the lower alkyl ester. If, for example, the alkanol exchange is effected with a simultaneous removal of the co-produced lower alkanol, a subsequent alkanol stripping unit operation could be eliminated.
  • the biodiesel is a crude or partially converted biodiesel
  • glycerin may be present or generated and can lead to reversion. In such instances, maintaining a high ratio of alkanol to glycerin may be desirable until the catalyst is inactivated or decomposed.
  • Any suitable catalyst can be used for the alkanol exchange.
  • the catalyst may be homogeneous or heterogeneous, and may be solid or liquid.
  • Catalysts include metal hydroxides, basic metal oxides, and metal alkoxides.
  • catalysts include alkali or alkaline earth metal hydroxide, alkali metal or alkaline earth metal glycerate or alkali or alkaline earth metal alkoxide, especially an alkoxide corresponding to the lower alkanol reactant.
  • Preferred alkali metals are sodium and potassium.
  • Other suitable catalysts include alkoxides or glycerates of titanium, zirconium, tungsten, iron, cobalt, cobalt, copper, zinc or aluminum.
  • Insoluble catalysts or partially insoluble catalysts if not in a fixed bed, can be recovered from the alkanol exchange product by, for instance, filtration or centrifugation, and recycled if desired.
  • Transition metal alkoxides that are soluble in the biodiesel such as titanium tetraisopropoxide, and can be decomposed represent a useful class of catalysts.
  • Another broad aspect of this invention pertains to processes for treating biodiesel comprising lower alkyl ester of fatty acid to lower cloud point by transesterification using an oil soluble, decomposable transition metal catalyst that decomposes, upon contact with a deactivating agent to provide a substantially insoluble transition metal oxide.
  • a deactivating agent include water and acids.
  • this aspect of the processes of this invention comprises: a.
  • contacting the biodiesel with higher molecular weight alkanol under transesterification conditions comprising the presence of a catalytically effective amount of oil soluble, decomposable transition metal catalyst to provide lower alkanol and an alkanol-exchanged biodiesel containing ester of the higher molecular weight alkanol; b. removing unreacted alkanol from the alkanol-exchanged biodiesel; and c. contacting the catalyst before or after step (b) with a deactivating agent in an amount sufficient to deactivate the catalyst and provide a substantially oil-insoluble transition metal oxide.
  • the sequence of contacting the catalyst with a deactivating agent will depend upon the concentration of total glycerin (unbound glycerin and compounds that can yield glycerin such as mono-, di- and triglycerides).
  • the total glycerin concentration is maintained below about 0.24, and sometimes below about 0.18, mass percent.
  • the transition metal catalyst comprises transition metal tetraalkoxide, more preferably titanium or zirconium tetraalkoxide.
  • the alkoxide may contain from 1 to about 20 or 24, often 1 to 4, carbon atoms. Titanium tetraisopropoxide is a particularly convenient catalyst due to its commercial availability.
  • the amount of catalyst used can vary widely and is frequently in the range of from about 0.05 to 5, say, about 0.1 to 2, mass percent based upon the oil. Any water or free fatty acid contained in the biodiesel will decompose catalyst, and therefore the amount of catalyst used is above that which will be decomposed by contact with the biodiesel.
  • the decomposing agent is used in a substantially stoichiometric amount.
  • the products of the decomposition are alkanols if water is used as the deactivating agent, and if acid is used, metal carboxylates. Alkanols can be removed from the alkanol-exchanged biodiesel by vapor fractionation.
  • the metal carboxylates need to be removed will depend upon the acid used for the decomposition. While any Bronsted or Lewis acid can be used for the decomposition, the preferred acids are carboxylic acids. With higher molecular weight carboxylic acids, the ester may be able to be retained in the biodiesel. Water is the preferred deactivating agent. The amount of deactivating agent should be sufficient to decompose substantially all the catalyst but preferably is not in such an excess that the water or free acid content of the biodiesel do not meet specifications.
  • the oil-insoluble, metal oxide may be removed from the alkanol exchanged biodiesel in any convenient manner including filtration and centrifugation.
  • biodiesel compositions containing mixtures of lower and higher alkyl esters.
  • Preferred biodiesel compositions of this invention comprise about 1 to 50 mass percent methyl ester of fatty acid and about 50 to 99 mass percent isopropyl or isobutyl ester of fatty acid.
  • Figure 1 is a schematic depiction of a facility to make biodiesel from fats or oils which includes an alkanol exchange operation.
  • Figure 2 is a schematic depiction of a stand alone apparatus using oil-soluble, decomposable catalyst that is suitable for the practice of a process of this invention.
  • FIG. 1 which schematically depicts biodiesel manufacturing facility 100.
  • Facility 100 is provided with a transesterification component (generally designated by numerals in the 200 series) as well as pretreatment components (generally designated by numerals in the 100 series) and a refining component generally (designated by numerals in the 300 series).
  • a transesterification component generally designated by numerals in the 200 series
  • pretreatment components generally designated by numerals in the 100 series
  • refining component generally (designated by numerals in the 300 series).
  • Biodiesel manufacturing facility 100 uses a suitable raw material, glyceride-containing feed.
  • glyceride-containing feeds can be used, and the point of introduction of such feeds into facility 100 will be influenced by the nature of the feed, especially the impurities therein.
  • Glycerides are aliphatic glycerides where the aliphatic groups contain between about 8 and 30, often between about 14 and 24 carbon atoms. Triglycerides have three such aliphatic groups, diglycerides, two such groups, and monoglycerides, one such group.
  • the feed may be one or more suitable oils or fats derived from bio sources, especially vegetable oils and animal fats.
  • glyceride- containing feeds include, but are not limited to rape seed oil, soybean oil, cotton seed oil, safflower seed oil, castor bean oil, olive oil, coconut oil, palm oil, corn oil, canola oil, jatropha oil, rice bran oil, tobacco seed oil, fats and oils from animals, including from rendering plants and fish oils.
  • the free fatty acid in the raw material feed is less than about 60 mass percent (dry basis). Suitable feeds may also contain phospholipids which may be as much as about 2 to 5 mass percent (dry basis) of the feeds. The balance of the fats and oils is largely fatty acid triglycerides.
  • the unsaturation of the free fatty acids and triglycerides may also vary over a wide range. Typically, some degree of unsaturation is preferred to reduce the propensity of the biodiesel to gel at cold temperatures. Additionally, more than one glyceride-containing feed can be simultaneously used to provide a blended biodiesel product.
  • the blend can be designed to adjust the cloud point of the product to a suitable level, or to take advantage of multiple feeds available to the producer.
  • the facility is capable of handling less expensive glyceride-containing feedstocks, especially those containing free fatty acids and phospholipids.
  • Preferred feeds comprise unrefined or partially refined soy oil, crude corn oil removed from syrup or distillers dried grain from fermentation processes, e.g., to make ethanol, and animal fats, and mixtures thereof and mixtures with other glycerides.
  • Pretreatment by acid treatment [0028] As shown in Figure 1, feed containing at least one of phospholipid and free fatty acid can be provided to facility 100 via line 102 for pretreatment by acid.
  • Line 104 is provided in the event that more than one feed is desired to be processed in the acid treatment section.
  • a blend of crude soy oil, which contains low concentrations of free fatty acid but higher concentrations of phospholipids, and crude corn oil, which contains higher concentrations of free fatty acid and lower concentrations of phospholipids, can be acid treated.
  • the feed may be directly introduced into acid treatment reactor 106, or as shown, is subjected to a contact with an alkanol laden stream of glycerin to strip alkanol from the glycerin into the oil-containing feed phase. This contact will be described later.
  • the feed will contain both free fatty acids and phospholipids, and acid treatment reactor 106 serves both to convert free fatty acid to esters and to facilitate removal of phospholipids.
  • esterification is conducted with alkanol, which may be a diol, but preferably is a monoalkanol, having a primary -OH, under esterification conditions.
  • the preferred alkanols are lower alkanols, especially those having 1 to 3 carbon atoms, although butanol and isobutanol and higher alkanols are operable. Most preferably the alkanol is methanol.
  • Esterification conditions include the presence of acidic catalyst, elevated temperature, e.g., at least about 40°C and sometimes as high as 200°C or more, and especially where the reaction menstruum is to be in the liquid phase, such high temperatures may be accompanied with the use of superatmospheric pressures sufficient to maintain the liquid phase.
  • an inerting gas such as nitrogen, hydrocarbon gas such as methane or carbon dioxide is used during the acid treatment.
  • Reactor 106 may comprise one or more stages or vessels and separation unit operations may be located between each stage or vessel. Where reactor 106 is staged, it is often desirable to remove water between stages to enhance conversion of free fatty acid to esters.
  • Reactor 106 may be a vessel or a length of pipe. But preferably other types of vessels are used such as mechanical and sonically agitated reactors, and reactors with static mixing such as reactors containing contact structures such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, and other impingement structures. Suitable reactors include those providing high intensity mixing, including high shear.
  • the molar ratio of alkanol to free fatty acid in the feed will vary depending upon the reactivity of the alkanol and the extent of esterification sought.
  • the esterification reaction is an equilibrium limited reaction and hence an excess of alkanol may be used to drive the reaction to the desired degree of completion.
  • any unreacted free fatty acid can be recovered elsewhere in the process and returned to acid treatment reactor 106.
  • the alkanol and the residence time in reactor 106 is sufficient to convert at least about 30 mole percent, and preferably at least about 50 mole percent, and sometimes at least about 75 mole percent to essentially all, the free fatty acid to ester.
  • the oil phase from the acid treatment section of facility 100 contain less than about 3, most preferably less than about 2, mass percent free fatty acid.
  • the use of alkanol is optional as it is not essential for phospholipid removal.
  • the molar ratio of alkanol to free fatty acid is generally between about 0:5:1 to 20:1, and preferably between about 0.9:1 to 10:1, and most preferably between about 3:1 to 9: 1.
  • the oil phase from the acid treatment section of facility 100 contains at least about 0.5, say between about 0.5 and 2 or 3, mass percent free fatty acid.
  • This free fatty acid serves to neutralize at least a portion of the base catalyst contained in a spent glycerin stream produced in the transesterification and base pretreatment sections of facility 100.
  • the molar ratio of free fatty acid in the oil phase from the acid treatment to mole of base in the glycerin phase introduced into base reactor 134 as discussed below will be at least about 0.3:1, often at least about 0.7: 1 up to about 1 :1.
  • ratios of free fatty acid to base catalyst of greater than 1 :1 can adversely affect the performance of the base pretreatment.
  • the equipment and conditions required for the esterification section need not be of the type required for essentially complete conversion of the free fatty acids, resulting in capital and operating cost savings. Since residual free fatty acid is converted to soap and removed in the base pretreatment section, the feed to the transesterification section can be substantially devoid of free fatty acid which adversely affects the base catalyst therein. Additionally, the neutralized spent glycerin stream from the base pretreatment section can be used effectively for enhancing phase separation and water and catalyst removal from the acid treatment product.
  • the water hydrolyzes or hydrates the phospholipids in the presence of acid to provide a water soluble phosphorus compound that can be removed from the oil phase by extraction into a water or glycerin phase.
  • Water for the hydrolysis or hydration may comprise that co-produced in the esterification of free fatty acids or provided by other streams within facility 100.
  • the catalyst can be heterogeneous or homogeneous. Where heterogeneous, it may be a solid or a highly dispersed liquid phase. As shown, liquid catalyst is provided via line 114 to acid treatment reactor 106. Any suitable acid catalyst (Bronsted acid or Lewis acid) for the esterification of free fatty acids can be used including homogeneous and heterogeneous catalysts.
  • the preferred acid catalysts are mineral acids such as hydrochloric acid, sulfurous acid, sulfuric acid, phosphoric acid, and phosphorous acid. However other strong acids including organic and inorganic acids can be used.
  • strong organic acids include alkyl sulfonic acids such as methylsulfonic acid; alkylbenzene sulfonic acids such as toluene sulfonic acid; naphthalenesulfonic acid; and trichloroacetic acid.
  • Solid acid catalysts include NAFION® resins.
  • Sulfuric acid and phosphoric acid are preferred due to non-volatility and low cost with sulfuric acid being most often used due to its availability and strong acidity.
  • Sulfuric acid may be provided in any suitable grade including, but not limited to highly concentrated, e.g., 98 percent, sulfuric acid, or in concentrated aqueous solutions, e.g., at least 30 percent, sulfuric acid. For the purposes of discussion in connection with facility 100, sulfuric acid is used as the acid.
  • the amount of acid catalyst provided can vary over a wide range. Typically the catalyst is provided in a catalytically effective amount of at least about 0.1 mass percent based upon the feed. Where soaps are present, the amount of acid should be sufficient to convert them to free fatty acids. Often the acid is present in an amount of at least about 0.2 to 5, say, 0.25 to 2, mass percent based upon the feed above that required to convert any soaps to free fatty acids.
  • the residence time for the acid treatment will depend upon the amounts of phospholipids and free fatty acid present, the conversion sought, the type and amount of catalyst used, the reactivity and amount of alkanol as well as the temperature of the process, and the type of reactor and extent of mixing. Residence times thus can range from less than 1 minute to over 1000 minutes. The residence times frequently are in the range of about 5 minutes to 120 minutes, preferably in the range of about 10 minutes to 90 minutes. [0038] Acid treatment temperatures are generally between about 3O 0 C and 200 0 C.
  • the reaction pressure can be any suitable pressure, e.g. from about 10 to 5000, preferably from about 90 to 1000, kPa absolute.
  • phase separator 110 is optional depending upon whether or not two phases exist. In some instances, an oil layer containing glycerides and fatty ester and a water- containing layer form.
  • the water-containing layer can contain more polar components such as glycerin, water-soluble catalyst, alkanol, and water-soluble phosphorus compounds.
  • a neutralized spent glycerin stream from the base pretreatment section is provided via line 170A and contacted with the acid treatment product.
  • the spent glycerin aids in the extraction of water and water-soluble phosphorus compounds. Additionally, the glycerin assists in making the phase separation.
  • the amount of glycerin added can vary widely. As relatively small amounts of water are produced during the acid esterification of free fatty acids, beneficial results can be achieved with relatively little spent glycerin being added.
  • Phase separator 110 may be of any suitable design including a decanter, a phase separation facilitated decanter that contains coalescing sites, and a centrifuge.
  • the lower, water-containing fraction exits separator 110 via line 112. This fraction contains some alkanol, water, water-soluble catalyst and water-soluble phosphorus compounds.
  • the oil fraction of separator 110 contains virtually no sulfuric acid, often some alkanol, relatively little water, unreacted free fatty acids, if any, fatty ester and glycerides.
  • the fraction is passed via line 118 from separator 110 to fractionation column 120 to provide an overhead fraction containing alkanol and a bottoms stream containing oil.
  • the overhead from column 120 can be recycled to acid treatment reactor 106 via line 122.
  • Make up alkanol is provided via line 124.
  • the fractionation column may be of any suitable design including a flash column, stripping column, falling film evaporator, or trayed or packed column. If desired, more than one fractionation column can be used with one effecting separation of water from alkanol.
  • a side draw 116 may be taken from distillation column 120 for the removal of water, and the fractionation column may be a divided wall column to enhance such separation.
  • a substantial portion of the water is removed by the phase separation in phase separator 110, and fractionation column does not separately recover water. Water will be contained in both the overhead and bottoms stream from column 120. However, the relatively small amount of water in the overhead can be recycled with alkanol via line 122 to reactor 106 without undue adverse effect. Water contained in the bottoms passes to the base pretreatment section and is removed from the oil phase therein. [0044] In another embodiment, only a portion of the alkanol is removed by fractionation in column 120. The alkanol remaining in the oil phase is passed to the base pretreatment section.
  • fractionation column 120 In the base pretreatment section alkanol can be reacted with glyceride to form esters and can be recovered in the spent glycerin phase for recycle to the acid treatment section. Thus, the capital and operating costs for fractionation column 120 can be reduced. Often the bottoms stream from fractionation column 120 contains between about 0.1 to 10, say, between about 0.5 and 5, e.g., 0.5 to 2, mass percent alkanol. [0045] While shown as processing the oil phase from separator 110, fractionation column 120 may be positioned between acid treatment reactor 106 and separator 110 and serve to recover alkanol from the acid treatment product exiting reactor 106.
  • the oil phase from the acid treatment contains little phospholipids and free fatty acids, preferably less than 0.1, more preferably less than 0.05, mass percent of each based on the oil phase, it can be directly passed to the transesterification component of the facility, i.e., to line 200. Alternatively, the oil phase can be passed to the base pretreatment component.
  • the base pretreatment uses glycerin produced in facility 100 to treat feed.
  • the base pretreatment serves to recover alkanol contained in the glycerin phase from the transesterification section. Hence, the spent glycerin from the base pretreatment section may contain relatively little alkanol.
  • Base pretreatment also serves to partially convert glycerides in the feed to fatty acid esters and mono- and di-glycerides. Thus, the amount of alkanol required to transesterify the pretreated feed will be less than had no base pretreatment occurred.
  • Base pretreatment can also serve to remove phospholipids as glycerin-soluble components.
  • Base pretreatment further removes free fatty acids from the glyceride-containing feed by saponification to glycerin-soluble soaps. Removal of the phospholipids and free fatty acids facilitates processing during transesterification and minimizes catalyst loss during transesterification cased by saponification of free fatty acids with base catalyst. Phospholipids, for instance, tend to make more difficult phase separations of oil and glycerin in the transesterification component. And biodiesel must meet stringent phosphorus specifications. See, for instance, ASTM D 6751, American Society for Testing and Materials.
  • a glyceride-containing feed stream is provided by line 132 to base reactor 134.
  • the feed stream may comprise a fresh glyceride-containing feed.
  • the feed stream may comprise the oil phase from the acid treatment provided via lines 126 and 130.
  • To base reactor is also provided a glycerin and base catalyst-containing stream via line 142 which will be further discussed below.
  • a non-acidic inerting gas such as nitrogen and hydrocarbon gas such as methane is used during base pretreatment.
  • base reactor 134 free fatty acids contained in the feed stream are reacted with base catalyst to form soaps.
  • additional base can be added via line 133.
  • the additional base may be the same or different from that comprising the catalyst, and may be one or more of alkali metal hydroxides or alkoxides and alkaline earth metal hydroxides, oxides or alkoxides, including by way of examples and not in limitation, sodium hydroxide, sodium methoxide, potassium hydroxide, potassium methoxide, calcium hydroxide, calcium oxide and calcium methoxide.
  • At least a portion is chemically reacted, e.g., by a hydration or by a salt formation, to provide chemical compounds preferentially soluble in glycerin.
  • Base reactor 134 is maintained under base reaction conditions, which for free fatty acid-containing feed streams is that sufficient to react basic catalyst and free fatty acids to soaps and water, and for phospholipids-containing feed streams is that sufficient to react basic catalyst and phospholipids to chemical compounds preferentially soluble in a glycerin phase.
  • Typical base reaction conditions include a temperature of at least about 10 0 C, say, 35°C to 150°C, and most frequently between about 40 0 C and 80 0 C.
  • Pressure is not critical and subatmospheric, atmospheric and super atmospheric pressures may be used, e.g., between about 1 and 5000, preferably from about 90 to 1000, kPa absolute.
  • the residence time is sufficient to provide the sought degree of saponification of fatty free acids and reaction of phospholipids.
  • the residence time in base reactor 134 may range from about 1 minute to 10 hours.
  • Base reactor 134 may be of any suitable design.
  • Reactor 134 may be a vessel or a length of pipe. But preferably other types of vessels are used such as mechanical and sonically agitated reactors, and reactors with static mixing such as reactors containing contact structures such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, and other impingement structures. Suitable reactors include those providing high intensity mixing, including high shear.
  • the base reaction product from reactor 134 contains glycerin, glycerides, soaps, water, reacted phospholipids and often fatty acid ester and is passed via line 136 to separator 128.
  • Separator 128 serves to separate the less dense oil layer from the more dense glycerin layer.
  • the soaps and reacted phospholipids preferentially pass to the glycerin layer as does most of the water.
  • the oil layer preferably contains less than about 0.5 mass percent soaps and less than about 500, preferably less than about 300, ppm-m phosphorus (calculated the elemental phosphorus).
  • Phase separator 128 may be of any suitable design including a decanter, a phase separation facilitated decanter that contains coalescing sites, and, if needed, a centrifuge.
  • the glycerin phase is withdrawn from separator 128 via line 137 and may be sent to glycerin recovery or another application. If the glycerin layer contains significant amounts of soaps, it may be desirable to recycle the soaps to acid treatment reactor 106 for conversion to fatty esters. As shown, a portion or all of the glycerin phase may be passed via line 170 to acidification reactor 172 where soaps are converted to free fatty acids. At least a portion of this glycerin phase is passed via line 17OA to provide the glycerin to assist in the separation of water, water-soluble catalyst (or salts thereof) and phosphorus-containing species from the acid treatment product in phase separator 110.
  • the glycerin-containing phase from separator 110 is passed via line 112 to line 170. Also as shown, a portion of the glycerin phase in line 170 is recycled to reactor 134 via line 170B.
  • the recycle can serve several purposes. For instance, hydrated phospholipids are returned to reactor 134 where they may undergo transesterii ⁇ cation to recover additional fatty acid ester. Also, any base contained in the recycled glycerin stream is available for saponification of free fatty acids.
  • Acidification reactor 172 may be one or more vessels of any suitable design including a length of pipe and other types of vessels such as mechanical and sonically agitated reactors, and reactors with static mixing such as reactors containing contact structures such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, and other impingement structures. Suitable reactors include those providing high intensity mixing, including high shear.
  • the acidification conditions usually encompass a temperature in the range of about 2O 0 C to 150°C, a pressure from about 1 to 5000, preferably 90 to 1000, kPa absolute, and a residence time of from about 1 second to 5 hours.
  • Suitable acids include mineral acids and organic acids, but typically a readily available acid such as sulfuric or phosphoric acid is used.
  • the amount of acid is usually sufficient to convert substantially all the soaps to free fatty acid. The use of excess acid is not deleterious to the formation of the free fatty acids, but can entail additional expense. Accordingly the molar ratio of acidifying acid function to soaps is in the range of about 1 :1 to 1.5:1.
  • the acidity of the glycerin stream is less than about 6, say, between about 1 and 5, e.g., 2 and 4.
  • the acidity of the glycerin stream is determined by diluting the glycerin stream to 50 volume percent water and measuring the pH.
  • the glycerin stream from acidification reactor is passed via line 176 to contact vessel 178 into which glyceride-containing feed is provided via line 102.
  • contact vessel 178 the glycerin stream is contacted with fresh feed which serves to extract a portion of the alkanol from the glycerin phase.
  • the contact with the glycerin also serves to remove water from the feed. Removal of water assists in the esterification of free fatty acids in acid treatment reactor 106 as the esterification is an equilibrium-limited reaction affected by water concentration.
  • Contact vessel 178 may be of any suitable design including a length of pipe and other types of vessels such as mechanical and sonically agitated reactors, and reactors with static mixing such as reactors containing contact structures such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, and other impingement structures. Suitable reactors include those providing high intensity mixing, including high shear.
  • the contact conditions usually encompass a temperature in the range of about 20°C to 150°C, a pressure from about 1 to 5000 kPa absolute, and a residence time of from about 1 second to 5 hours. Often at least about 50 mass percent of the alkanol in the glycerin stream passes to the oil phase as do essentially all of the free fatty acids.
  • the amount of alkanol recovered from the glycerin will depend upon the alkanol content of the glycerin, the ratio of glycerin to fresh feed, and the contacting conditions. Frequently the mass ratio of glycerin to oil is in the range of between about 1:5 to 1:20, say 1 :8 to 1 :15, and at least about 30, and sometimes between about 50 and 99, mass percent of the alkanol in the glycerin phase passes to the oil phase.
  • Figure 1 shows two glycerin loops for alkanol recovery and recycle to the acid treatment reactor. The first loop involves the glycerin layer from separator 110 and the second, the glycerin layer from separator 128.
  • phase separator 182 a glyceride and free fatty acid oil layer is produced that is passed via line 184 to acid treatment reactor 106.
  • a glycerin-containing layer is discharged via line 186 and contains water, acidification acid, and soluble phosphorus compound.
  • Separator 182 may be of any suitable design including a decanter, a phase separation facilitated decanter that contains coalescing sites, and, if necessary, a centrifuge.
  • Contact vessel 178 and phase separator 182 may be a single vessel, including but not limited to, a countercurrent extraction column.
  • the acid treatment product from acid treatment reactor 106 has a sufficiently low free fatty acid content and low phospholipids content
  • another option is to eliminate separator 110 and fractionation column 120 and provide the acid treatment product in line 108 directly to separator 128 or base reactor 134.
  • the oil phase is withdrawn and passed via line 138 to second pretreatment reactor 139.
  • Second pretreatment reactor 139 and third pretreatment reactor 148 are adapted to recover alkanol contained in the glycerin from the transesterification component of facility 100 through reaction, e.g., transesterification and extraction into the glyceride-containing phase. A base transesterification process is used in these pretreatment reactors.
  • the number of reactors will depend upon the sought consumption of the alkanol as well as the efficiency of the reactors. Hence one, two, or three or more pretreatment reactors may be used.
  • the pretreatment reactor can comprise a number of stages in a single vessel which could be a countercurrent contact vessel.
  • the feed stream to the alkanol consumption pretreatment reactors is relatively free from free fatty acids so as to prevent undue consumption of the base catalyst.
  • the pretreatment reactors provide a glycerin stream from which most of the alkanol has been removed.
  • the alkanol content of the glycerin discharged from base reactor 134 is less than about 5, and preferably less than about 2, mass percent.
  • a significant portion of the alkanol is contained in line 126 (or line 108 if separator 110 and distillation column 120 are not used) and passed to separator 128.
  • the concentration of alkanol in the glycerin-containing stream in line 170 may be higher than 5 mass percent, and alkanol is recovered be partitioning to the glyceride-containing feed in contact vessel 178.
  • the alkanol content of the glycerin may be sufficiently low that no distillation is required to recover alkanol yet the overall process to make biodiesel can still exhibit high efficiencies.
  • Second pretreatment reactor 139 also receives the glycerin phase from the third pretreatment reactor.
  • This glycerin phase contains glycerin, base catalyst, and alkanol.
  • Second pretreatment reactor 139 is maintained under base transesterification conditions including the presence of base catalyst provided by the glycerin phase feed and elevated temperatures, often between about 3O 0 C and 22O 0 C, preferably between about 30 0 C and 8O 0 C to provide a second pretreatment product.
  • the pressure is typically in the range of between about 90 to 1000 kPa (absolute) although higher and lower pressures can be used.
  • the reactor is typically batch, semi-batch, plug flow or continuous flow tank.
  • reactors such as mechanical and sonically agitated reactors, and reactors with static mixing such as reactors containing contact structures such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, and other impingement structures.
  • Suitable reactors include those providing high intensity mixing, including high shear.
  • the residence time will depend upon the desired degree of conversion of the contained alkanol, the ratio of alkanol to glyceride, reaction temperature, the degree of agitation and the like, and is often in the range of about 0.1 to 20, say, 0.5 to 10, hours.
  • the second pretreatment product contains glycerides, fatty esters, base catalyst and glycerin, and it has a reduced concentration of alkanol.
  • the second pretreatment product is passed from second pretreatment reactor 139 via line 141 to separator 140.
  • Separator 140 may be of any suitable design including a decanter, a phase separation facilitated decanter that contains coalescing sites, and, optionally, a centrifuge.
  • the lower, glycerin-containing phase from separator 140 contains relatively little alkanol, preferably less than about 10 mass percent, and contains base catalyst, and is passed via line 142 to base reactor 134 where catalyst reacts with free fatty acids to form soaps which can then be removed from the glyceride-containing feed.
  • line 142 is provided with holding tank 142A.
  • Holding tank 142A can serve as a reservoir and enables the rate that glycerin, which contains base, is provided to base reactor 134, to be varied with changes in free fatty acid content of the acid treatment product. It also can permit additional reaction of glycerides with alkanol contained in the glycerin phase to occur prior to introduction into base reactor 134 where catalyst is consumed by conversion of free fatty acids to soaps.
  • the upper oil phase is removed from separator 140 via line 144 and is passed to line 146 which also receives the glycerin co-product from transesterification from line 248.
  • the combined streams are passed to third pretreatment reactor 148 by line 146 and contain in addition to glycerin, alkanol, base catalyst, and usually some water and soaps.
  • Table I sets forth typical compositions of the stream in line 248. The compositions, of course, will depend upon the operation of the transesterification component as well as which of the glycerin-containing streams from the transesterification component are used. The typical concentrations are based upon combining all glycerin-containing streams and operating under preferred parameters.
  • Third pretreatment reactor 148 is maintained under base transesterification conditions including the presence of base catalyst provided by the glycerin-containing feed and elevated temperatures, often between about 3O 0 C and 220 0 C, preferably between about 3O 0 C and 8O 0 C to provide a first pretreatment product.
  • Base catalyst in the transesterification component tends to partition to the glycerin phase and often adequate catalyst is provided for the base pretreatment section. In some instances, however, it may be desired to add additional base catalyst to third pretreatment reactor 148 or any preceding base pretreatment reactor.
  • the pressure is typically in the range of between about 90 to 1000 kPa (absolute) although higher and lower pressures can be used.
  • the reactor is typically batch, semi-batch, plug flow or continuous flow tank with some agitation or mixing.
  • the preferred types of vessels are mechanical and sonically agitated reactors, and reactors with static mixing such as reactors containing contact structures such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, and other impingement structures.
  • Suitable reactors include those providing high intensity mixing, including high shear. However, depending upon the presence of soaps and phospholipids, care needs to be taken so as not to generate a product that cannot be readily separated by phase separation.
  • the residence time will depend upon the desired degree of conversion, the ratio of alkanol to glyceride, reaction temperature, the degree of agitation and the like, and is often in the range of about 0.1 to 20, say, 0.5 to 10, hours.
  • the transesterification in third pretreatment reactor 148 recovers through transesterification and extraction to the glyceride-containing phase at least about 20, preferably at least about 30, and more preferably at least about 50, mass percent of the alkanol fed to the reactor. Any unreacted alkanol in the oil phase will be carried with the oil phase to the transesterification component of facility 100.
  • the total amount of alkanol recovered from the glycerin-coproduct from transesterification using all pretreatment stages is at least about 50, and sometimes at least about 80, mass percent.
  • the third pretreatment product passes from third pretreatment reactor 148 through line 150 to separator 152.
  • Separator 152 may be of any suitable design including a decanter, a phase separation facilitated decanter that contains coalescing sites, and, optionally, a centrifuge. Separator 152 serves to separate an oil phase containing glycerides, esters and alkanol and some catalyst, from a glycerin-containing phase containing glycerin, reduced concentration of alkanol, and catalyst. The glycerin-containing phase frequently contains less than about 15 mass percent alkanol.
  • the glycerin-containing phase from separator 152 is passed via line 154 to second pretreatment reactor 139.
  • Facility 100 includes a chiller 158 to remove high molecular weight glycerides, waxes and esters that are insoluble at the chiller temperature.
  • Some feeds such as crude corn oil, contain high molecular weight glycerides and esters.
  • the hydrocarbyl moieties in these high molecular weight components typically have between 30 and 40 carbon atoms. If they remain in the oil, the resultant biodiesel product tends to have unacceptably high cloud points and gel points.
  • the oil phase from separator 152 passes through line 156 to chiller 158.
  • Chiller 158 is maintained at a temperature sufficient to cause high molecular weight and other components that lead to and increase in gel point temperature to solidify.
  • this temperature is between about 0 0 C and 20°C.
  • cooling will tend to remove monoglycerides and diglycerides. Cooling below the desired temperature and then warming to a temperature to liquefy the mono- and di-glycerides while still maintaining a solid wax, can minimize loss of components that can be converted to biodiesel.
  • the chilled oil phase is then passed via line 160 to centrifuge 162 to remove higher density components including solids and any remaining glycerin phase.
  • the higher density fraction is discharged via line 164. Rather than using a centrifuge, the solids can be filtered from the glyceride-containing stream. Filter aids can be used if desired.
  • Chiller 158 is optional, and a chiller may also be used elsewhere in facility
  • the pretreated glyceride product in line 156 is a unique producer composition that is a highly desirable feed for transesterification processes to make biodiesel.
  • all or a portion of the producer composition in line 166 may be withdrawn via line 168 as an intermediate product for storage or sale as a feedstock for transesterification.
  • Line 168 also provides the feed for the transesterification component of facility 100 by introducing the producer composition into line 200.
  • Line 200 provides glyceride-containing feed to first transesterification reactor 202.
  • Line 200 can also supply additional glyceride-containing feed.
  • the additional feed is relative free of free fatty acids and phospholipids such as refined oils sourced from rape seed, soybean, cotton seed, safflower seed, castor bean, olive, coconut, palm, corn, canola, fats and oils from animals, including from rendering plants and fish oils.
  • Alkanol for the transesterification is supplied to first transesterification reactor via line 206.
  • the alkanol is preferably lower alkanol, preferably methanol, ethanol or isopropanol with methanol being the most preferred.
  • the alkanol may be the same or different from the alkanol provided to acid pretreatment reactor 106 via line 124.
  • line 206 is depicted as introducing alkanol into line 200, it is also contemplated that alkanol can be added directly to reactor 202 at one or more points.
  • the total alkanol (line 206 and from the producer composition of line 166) is in excess of that required to cause the sought degree of transesterification in reactor 202.
  • the amount of alkanol is from about 101 to 500, more preferably, from about 110 to 250, mass percent of that required for the sought degree of transesterification in reactor 202. In facility 100 three reactors are depicted as being used.
  • One reactor may be used, but since the reaction is equilibrium limited, most often at least two and preferably three reactors are used. Often, where more than one reactor is used, at least about 60, preferably between about 70 and 96, percent of the glycerides in the feed are reacted in first transesterification reactor 202. It is possible to provide all the alkanol required for transesterification to first transesterification reactor 202, or a portion of the alkanol can be provided to each of the transesterification reactors. [0076] The base catalyst is shown as being introduced via line 204 to first transesterification reactor 202. The amount of catalyst used is that which provides a desired reaction rate to achieve the sought degree of transesterification in first transesterification reactor 202.
  • catalyst is provided to each of the transesterification reactors since base catalyst preferentially partitions to the glycerin phase and is removed with phase separation of the glycerin after each transesterification reactor.
  • the amount of catalyst used will be in excess of that required to react with the amount of free fatty acid contained in the feed oil, which due to the pretreatment, will be relatively little.
  • the base catalyst may be an alkali or alkaline earth metal hydroxide or alkali or alkaline earth metal alkoxide, especially an alkoxide corresponding to the lower alkanol reactant.
  • Preferred alkali metals are sodium and potassium.
  • the base When the base is added as a hydroxide, it may react with the lower alkanol to form an alkoxide with the generation of water which in turn results in the formation of free fatty acid.
  • Another type of catalyst is an alkali metal or alkaline earth metal glycerate. This catalyst converts to the corresponding alkoxide of the alkanol reactant in the reaction menstruum.
  • the catalyst may be a heterogeneous base catalyst. Catalyst may need to be separately provided to the base pretreatment reactors if the base catalyst, e.g., a heterogeneous or oil soluble catalyst, is not carried with the co-product glycerin in the transesterification component to the base pretreatment reactors.
  • homogeneous catalysts that have solubility in glycerin are preferred where the pretreatment component is used since the catalyst serves as at least a portion of the base used therein.
  • the exact form of the catalyst is not critical to the understanding and practice of this invention.
  • homogenous base catalyst is used.
  • a non-acidic inerting gas such as nitrogen and hydrocarbon gas such as methane is used during base transesterification.
  • the transesterification is at a temperature between about 3O 0 C and
  • First transesterification reactor 202 is typically batch, semi-batch, plug flow or continuous flow tank with some agitation or mixing. Preferably the reactors are mechanical and sonically agitated reactors. Reactors with static mixing such as reactors containing contact structures such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, and other impingement structures can be used. Suitable reactors include those providing high intensity mixing, including high shear.
  • phase separator 210 may be of any suitable design including a decanter, a phase separation facilitated decanter that contains coalescing sites, and, optionally, a centrifuge.
  • a glycerin-containing bottoms phase is provided in the separator and is removed via line 212 and is passed to glycerin header 214. As depicted, this stream is used as a portion of the glycerin for the pretreatment component of facility 100.
  • This glycerin phase also contains any soaps made in reactor 202 and a portion of the catalyst. The soaps can be recovered from this stream in acidifying reactor 172 as discussed above.
  • the lighter phase contains alkyl esters and unreacted glycerides and is passed via line 216 to second transesterificati on reactor 218.
  • a rag layer may form in separator 210.
  • the rag layer may contain unreacted glycerides, alkyl esters, alkanol, soaps, catalyst and glycerin.
  • An advantage of the process set forth in Figure 1 is that withdrawing the rag layer with the glycerin phase does not result in a loss of glycerides, alkyl esters, alkanol, and catalyst since the glycerin phase is passed to the pretreatment component of facility 100.
  • Reactor 218 may be of any suitable design and may be similar to or different than reactor 202. As shown, additional alkanol is provided via line 206A, and additional catalyst is provided via line 204 A.
  • the transesterification conditions in reactor 218 are sufficient to react at least about 90, more preferably at least about 95, and sometimes at least about 97 to 99.9, mass percent of the glycerides in the feed to the transesterification.
  • the transesterification in reactor 218 is typically operated under conditions within the parameters set forth for reactor 202 although the conditions may be the same or different.
  • the residence time will depend upon the desired degree of conversion, the ratio of alkanol to glyceride, reaction temperature, the degree of agitation and the like, and is often in the range of about 0.02 to 20, say, 0.1 to 10, hours.
  • phase separator 222 which may be of any suitable design and may be the same as or different from the design of separator 210.
  • a heavier, glycerin-containing phase is withdrawn via line 224 and passed to glycerin header 214.
  • a lighter phase containing crude biodiesel is withdrawn from separator 222 via line 226.
  • third transesterification reactor 228 is used and the crude biodiesel in line 226 is passed to this reactor.
  • the transesterification conditions in reactor 228 are sufficient to provide essentially complete conversion, at least about 97 or 98 to 99.9, mass percent of the glycerides in the feed converted to alkyl ester. As shown, additional alkanol is provided via line 206B, and additional catalyst is provided via line 204B.
  • the transesterification in reactor 228 is typically operated under conditions within the parameters set forth for reactor 202 although the conditions may be the same or different. The residence time will depend upon the desired degree of conversion.
  • the reactor may be of the type described for reactor 202.
  • the residence time will depend upon the desired degree of conversion, the ratio of alkanol to glyceride, reaction temperature, the degree of agitation and the like, and is often in the range of about 0.02 to 20, say, 0.1 to 10, hours.
  • the transesterification product from third transesterification reactor 228 contains less than about 1, preferably less than about 0.8, and most preferably less than 0.5, mass percent soaps based upon the total mass of alkyl esters and soaps.
  • the lighter phase also contains alkanol.
  • the reaction proceeds quickly to completion by the addition of additional alkanol and catalyst, and can be conveniently accomplished by a plug flow reactor.
  • the overall molar ratio of alkanol to glycerides in the feed to the reactors in the transesterification component can vary over a wide range. Since transesterification is an equilibrium-limited reaction, the driving force toward the alkyl ester and the conversion of glycerides will be dependent upon the molar ratio of alkanol equivalents to glycerides.
  • Alkanol equivalents are alkanol and alkyl group of the alkyl esters in the feed to the transesterification component.
  • the mole ratio of alkanol equivalents to glyceride in the feed to the pretreatment component is frequently between about 3.05:1 to 15:1, say 4:1 to 9:1.
  • the pretreatment processes of this invention permit the reuse of alkanol partitioned to the co-product glycerin without intermediate vaporization.
  • the amount of total catalyst provided based upon the mass of feed to the first transesterification reactor, i.e., the catalyst provided by lines 204, 204A and 204B, is between about 0.3 and 1 mass percent (calculated on the mass of sodium methoxide).
  • phase separator 232 which may be of any suitable design and may be the same as or different from the design of separator 210.
  • a heavier, glycerin-containing phase is withdrawn via line 234 and passed to glycerin header 214.
  • a lighter phase containing crude biodiesel is withdrawn from separator 232 via line 236.
  • separator 232 can be eliminated provided that in second transesterification reactor 218, the conversion of the glycerides in the feed is at least about 90, preferably 92 to 96 or 98, percent.
  • the effluent from reactor 228 may be a single phase containing relatively little glycerin.
  • Facility 100 contains an optional alkanol replacement reactor 238.
  • the alkanol replacement reactor serves to transesterify the alkyl ester with a different alkanol.
  • an alkanol such as methanol provides not only attractive reaction rates but also an effluent that is more easily separated than, say, a reaction effluent where ethanol is the alkanol.
  • the transesterification between, say, a fatty acid methyl ester, and higher molecular weight alkanol results in methanol, rather than glycerin, being formed, and often is more readily accomplished than the transesterification of glyceride with that higher alkanol.
  • the higher alkanols include those having 2 to 8 or more carbon atoms, and are preferably branched primary and secondary alkanols although tertiary alkanols may find application but generally are less reactive.
  • higher alkanols examples include propanol, isopropanol, isobutanol, 2,2-dimethylbutan-l-ol, 2,3-dimethylbutan-l-ol, 2-pentanol, and the like.
  • Other alkanols include benzyl alcohol and 2 ethylhexanol.
  • An advantageous higher alkanol feed is fusel oil from the production of ethanol which contains a range of alkanols.
  • An alkanol replacement operation may be located at various points in the process.
  • the replacement alkanol may be provided via line 206B to reactor 228, or, as shown, it can follow reactor 228.
  • alkanol replacement transesterification can take advantage of catalyst contained in the transesterification medium.
  • alkanol replacement may be effected on a biodiesel product by adding catalyst.
  • it can be located elsewhere in the refining component of facility 100 including, but not limited to, treating biodiesel in line 352.
  • the biodiesel fed to the alkanol replacement operation has a relatively low total glycerin concentration (glycerin and compounds that can yield glycerin), most preferably less than about 0.24, and even more preferably less than about 0.18, mass percent.
  • glycerin and compounds that can yield glycerin glycerin and compounds that can yield glycerin
  • Figure 1 depicts the alkanol replacement unit operation between the transesterification section and the refining section of facility 100 to take advantage of the refining section.
  • the amount of higher alkanol provided via line 240 to alkanol replacement reactor 238 can vary over a wide range. Typically the molar ratio of higher alkanol to alkyl ester being fed to reactor 238 is less than 0.5:1, e.g., from about 1:100 to 1 :5. Often the alkanol replacement transesterification is at a temperature between about 3O 0 C and 22O 0 C, preferably between about 30 0 C and 8O 0 C.
  • the pressure is preferably sufficient to maintain a liquid phase reaction menstruum and typically is in the range of between about 90 to 1000 kPa (absolute) although higher and lower pressures can be used.
  • the temperature and pressure conditions are such that the displaced alkanol is able to be removed by vapor fractionation from reactor 238 during the reaction.
  • Alkanol replacement reactor 238 can be batch, semi-batch, plug flow or continuous flow tank with some agitation or mixing, e.g., mechanically stirred, ultrasonic, static mixer containing contact surfaces, e.g., trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, or other impingement structures.
  • High intensity mixing reactors, including high shear reactors may also be used.
  • Preferred reactors are those in which the alkanol being replaced is continuously removed.
  • a reactive distillation reactor can be used to continuously remove displaced methanol from a transesterification of methyl ester and isopropanol.
  • reactor 238 is a reactive distillation unit and lower alkanol is withdrawn via line 330A and passed to the transesterification reactors. Make-up alkanol is provided via line 332.
  • the alkanol replacement reactor is a batch reactor, driving the replacement reaction to either essentially complete conversion of the higher alkanol or complete conversion of the methyl ester to the higher alkanol ester (depending upon whether the higher alkanol is provided below or at or above the stoichiometric amount required for complete conversion), since the vapor fractionation of methanol can continue until completion.
  • continuous reactors having unreacted methanol and higher alkanol in the alkanol replacement product is likely. For purposes of this discussion, a continuous alkanol replacement reactor is used.
  • Suitable catalyst includes base catalyst such as is used for transesterification. Since a single liquid phase exists during the alkanol replacement unlike transesterification where a glycerin layer forms, heterogeneous catalysts and homogeneous catalysts having limited solubility in the reaction menstruum can be used. Solid catalysts are preferred to minimize or eliminate post treatment of the alkanol replacement product, but good contact with catalyst is desirable to timely achieve sought conversion.
  • Homogeneous transesterification catalysts such as titanium tetra-isopropoxide are also advantageous as they are readily removed, e.g., by decomposition with the decomposition products being removed.
  • the residence time will depend upon the desired degree of conversion, the ratio of higher alkanol to alkyl ester, reaction temperature, the degree of agitation and the like, and is often in the range of about 0.02 to 20, say, 0.1 to 10, hours. Preferably at least about 80, and sometimes at least about 90, mass percent of the higher alkanol is reacted. Refining
  • a crude biodiesel is withdrawn from reactor 238 via line 300 and is passed to the refining component of facility 100.
  • the crude biodiesel may be contacted with acid to neutralize any catalyst therein and then refined to remove alkanol, soaps, water and glycerin.
  • an acid preferably an organic acid, is provided via line
  • inorganic acids can be used, organic acids, particularly those less volatile than the alkanol, and acids that do not themselves or any potential reaction product formed in contact with the crude biodiesel, form azeotropes with the alkanol, facilitate processing and minimize the potential of contamination of recovered alkanol.
  • Exemplary acids include acetic acid, citric acid, oxalic acid, glycolic, lactic, free fatty acid and the like.
  • the amount of catalyst contained in the crude biodiesel is quite small as base catalyst preferentially partitions to the glycerin phase. Accordingly, little acid is required to neutralize sufficient catalyst to enable refining without risk of reversion of alkyl ester. Often the amount of acid used is at least 0.95 times, sometimes between about 1 and 3 times, that required to neutralize the catalyst.
  • Crude biodiesel is passed via line 300 to an alkanol separation unit operation.
  • a two stage separation unit is used.
  • a single stage separator can be used if desired.
  • the crude biodiesel in line 300 is passed to first alkanol separator stage 304.
  • Separator 304 is of any convenient design including a stripper, wiped film evaporator, falling film evaporator, solid sorbent, and the like.
  • the fractionation is by fast, vapor fractionation.
  • the residence time is less than about one minute, preferably less than about 30 seconds, and sometimes as little as 5 to 25 seconds.
  • the vapor fractionation conditions comprise a maximum temperature of less than about 200 0 C, preferably less than about 150°C, and most preferably, when the lower alkanol is methanol, less than about 120°C.
  • the lower boiling fractionation may need to be conducted under subatmospheric pressure to maintain desired overhead and maximum temperatures.
  • a falling film stripper it may be a concurrent or countercurrent flow stripper. Concurrent strippers are preferred should there be a risk of undue vaporization of alkanol at the point of entry of the crude biodiesel.
  • An inert gas such as nitrogen may be used to assist in removing the alkanol.
  • the fast fractionation may be effected by any suitable vapor fractionation technique including, but not limited to, distillation, stripping, wiped film evaporation, and falling film evaporation.
  • the falling film evaporator has a tube length of at least about 1 meter, say, between about 1.5 and 5 meters, and an average tube diameter of between about
  • the vapor fractionation recovers at least about 70, preferably at least about 90, mass percent of the alkanol contained in the crude biodiesel. Any residual alkanol is substantially removed in any subsequent water washing of the crude biodiesel.
  • the amount of alkanol contained in the spent water from the washing may be at a sufficiently low concentration that the water can be disposed without further treatment.
  • alkanol can be recovered from the spent wash water for recycle to the transesterification reactors.
  • the lower boiling fraction containing the alkanol will contain a portion of any water contained in the crude biodiesel.
  • Alkanol is exhausted from first alkanol separator stage via line 306 and may be exhausted from the facility as a by-product, e.g., for burning or other suitable use, or can be recycled. Where no alkanol replacement reaction is used, the alkanol will be the lower alkanol for the transesterification and is recycled to the transesterification section.
  • the bottoms stream from first alkanol separation stage 304 is passed via line 308 to second alkanol separation stage 314 for additional alkanol recovery.
  • the design of second alkanol separation stage 314 may be similar to or different than that of first alkanol separation stage 304 and may be operated under the same or different conditions.
  • Alkanol exits via line 316 and is combined with alkanol from line 306 and is passed to condenser 318. In the process of facility 100, the condensed alkanol will contain both the lower alkanol and the higher alkanol. Condensed alkanol is recycled via line 330 to alkanol replacement reactor 238. Non-condensed gases exit condenser 318 via line 320.
  • the alkanol separation operation is maintained under vacuum conditions and these gases are passed to liquid ring vacuum pump 322.
  • the liquid for the liquid ring is provided via line 324 and exits via line 328.
  • the gases contain some alkanol
  • the liquid for the liquid ring vacuum pump will remove alkanol from the gases.
  • the liquid may be water, in which case the water may need to be treated to remove alkanol.
  • Alternative liquid streams can be used, including but not limited to glyceride-containing feed, biodiesel, and glycerin. Feed is preferred as the liquid for the liquid ring vacuum pump since it can be passed to a transesterification reactor and alkanol contained therein used for the transesterification.
  • Gas is removed from liquid ring vacuum pump 322 via line 326.
  • the bottoms stream from the second alkanol separation stage exits via line 334 and is passed to separator 336 in which a glycerin-containing phase and a biodiesel-containing phase are separated.
  • Agents to facilitate the separation such as water and ionic salts can also be provided to separator 336.
  • the presence of alkanol in the crude biodiesel enhances the solubility of glycerin therein.
  • a separate glycerin-containing phase which also contains soaps, tends to form during the alkanol separation operation.
  • the glycerin fraction is removed from separator 336 via line 338 and can be combined with spent glycerin in line 186.
  • the lighter, oil-containing phase is passed via line 340 to a water wash unit operation.
  • Line 340 serves as a reactor and mixer where strong acid is supplied.
  • the amount of strong acid provided is sufficient to convert any soaps remaining to free fatty acids.
  • Sufficient strong acid is used such that water used for washing the crude biodiesel is at a suitably low pH.
  • the strong acid is supplied in admixture with a recycle stream in the wash operation as will be explained later. While line 340 serves as an in-line mixer, a separate vessel may be used for the acidification.
  • a separate mixer may be of any convenient design, e.g., a mechanically or sonically agitated vessel, or static mixer containing static mixing devices such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, or other impingement structure.
  • static mixer containing static mixing devices such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, or other impingement structure.
  • sufficient mixing and residence time should be provided such that essentially all of the soaps are converted to free fatty acids.
  • the temperature during the mixing is in the range of about 3O 0 C to 220 0 C, preferably between about 60 0 C to 180°C, and for a residence time of between about 0.01 to 4, preferably 0.02 and 1, hours.
  • the water wash operation uses a two stage water wash.
  • Water wash operation may be of any suitable design.
  • the water wash operates with a recycling water loop, often with the water recycle being at least about 20, say between about 30 and 500, mass percent of the crude biodiesel being fed to the column.
  • Normally washing is operated at a temperature between about 20 0 C and 12O 0 C, preferably between about 35°C and 90 0 C.
  • the amount of water provided to each wash vessel is sufficient to effect a sought removal of glycerin, residual alkanol and any water-soluble contaminants from the crude biodiesel.
  • mass parts of wash water are used per 100 mass parts of crude biodiesel.
  • the free fatty acid is present in an amount less than about 3000, most frequently less than about 2500, parts per million by mass in the biodiesel product, and thus no need exists to remove free fatty acid to provide a biodiesel product meeting current commercial specifications.
  • the vessels used for the water washing may be of any suitable design including a pipe reactor, mechanically or sonically agitated tank, a vessel containing static mixing devices such as trays, packing, baffles, orifices, venturi nozzles, tortuous flow path, or other impingement structure. Each stage needs to effect a phase separation of the oil phase from the water phase.
  • Such a separation may be inherent in, for instance, a wash column where the water and oil phases are moving countercurrently, or a separate phase separator may be provided. It is understood that other washing operations can be used such as a one vessel washing operation, an acid wash followed by a neutral wash, and the like. The washing may be effected in one or more stages and in one or more vessels. A single vessel, such as a wash column can contain a plurality of stages. [00100] As shown, crude biodiesel is provided via line 340 to first wash stage 342.
  • wash stage 342 comprises an agitated vessel to provide desired contact between the oil and water phases and a decanter to effect separation.
  • the agitated vessel provides a contact time of about 1 second and 10 minutes, say, 5 to 60 seconds.
  • Crude biodiesel is contacted with acidic water from water loop 368.
  • the washed biodiesel from first wash stage 342 is passed via line 344 to second wash stage 346 having a design similar to or different from that of stage 342.
  • This biodiesel is contacted with water from water loop 364.
  • Acidic water is withdrawn from first wash stage 342 and recycled via line 368.
  • Substantially neutral water is withdrawn from second wash stage 346 and recycled via line 364. Additional water is provided to line 364 via line 376 which will be described later.
  • the pH of the water in second wash stage 346 may be neutral or less acidic than the water in first wash stage 342.
  • Make-up water to line 368 is provided by line 366.
  • a purge is taken from line 368 via line 372.
  • the purge balances the amount of water in the wash loops and is at a suitable rate to maintain desirably low concentrations of impurities such as alkanol and glycerin in the water used for the washing.
  • the purge is usually at a rate of between about 1 and 50, say 5 and 20, mass percent per unit time of the recycle rate in the loop.
  • Line 370 provides strong acid to the water recycled via line 368 for combining with crude biodiesel in line 340 or being passed to first wash stage 342.
  • Adequate strong acid aqueous solution is provided that the water in line 368 has a pH sufficiently low to convert the soaps to free fatty acids.
  • the acid may be any suitable acid to achieve the sought pH such as hydrochloric acid, sulfuric acid, sulfonic acid, phosphoric acid, perchloric acid and nitric acid. Sulfuric acid is preferred due to cost and availability and it is a non-oxidizing acid.
  • the amount of strong acid aqueous solution provided is typically in a substantial excess of that required to convert the soaps to free fatty acid and to neutralize any remaining catalyst.
  • the feed to first wash stage 342 provides a wash water in line 368 having a pH of up to about 4, preferably between about 0.1 and 4.
  • the purge water is passed to evaporator 374 which provides a lower boiling fraction and a higher boiling fraction. While an evaporator may be used, it is also possible to use a packed or trayed distillation column with or without reflux. Generally the bottoms temperature of evaporator 374 is less than about 150°C, preferably between about 120°C and 150°C. The distillation may be at any suitable pressure.
  • a membrane separation system may, alternatively or in combination, be used with evaporator 374 to effect the sought concentration of the spent water.
  • the lower boiling fraction contains water, potentially acid if not neutralized or salts, and some alkanol and is passed via line 376 to water wash loop 364. Fresh water is provided to line 376 by line 380.
  • the higher boiling fraction contains glycerin, some alkanol and some water and potentially acid or salts thereof. The higher boiling fraction or a portion thereof is preferably passed via line 382 to line 170 or it can be combined with spent glycerin.
  • a washed biodiesel stream is withdrawn from second washing stage 346 via line 348 and is passed to drier 350 to remove water which exhausts via line 354.
  • Drier 350 may be of any suitable design such as stripper, wiped film evaporator, falling film evaporator, and solid sorbent. Generally the temperature of drying is between about 6O 0 C and 22O 0 C, say, about 70 0 C and 180 0 C. The pressure is generally in the range of about 5 to 200 kPa absolute.
  • the dried biodiesel is withdrawn as product via line 352.
  • the biodiesel product contains free fatty acid and preferably has a free fatty acid content of less than about 0.3 mass percent.
  • An inert gas such as nitrogen may be used in facilitating drying.
  • the subatmospheric pressure is maintained in drier 350 by the use of liquid ring vacuum pump 356 which is in communication with line 354.
  • Liquid ring vacuum pump 356 uses water as the sealing fluid which provided by line 358 and water exits via line 362.
  • the gases from liquid ring vacuum pump 356 exit via line 360.
  • the glycerin-containing streams are passed via line 242 to blending tank 246 such that a relatively uniform glycerin composition can be provided via line 248 to the pretreatment section of facility 100.
  • Blending tank 246 may also provide sufficient residence time for any glycerides in the glycerin to transesterify with alkanol as well as permit any oil entrained in the glycerin phase to separate. As shown, an oil layer that forms in blending tank 246 can intermittently or continuously be withdrawn via line 247 for recycle to first transesterification reactor 202. Alternatively, the oil layer can be withdrawn with the glycerin and passed to the pretreatment section.
  • glycerin-containing streams from the transesterification and refining components of facility 100 have been shown to be directed to glycerin header 214, it is with in the purview of the process to use fewer streams. Moreover, any of the glycerin-containing streams may be used elsewhere prior to being passed to blending tank 246, and the blended stream or a portion thereof in line 248 may be used elsewhere and either returned to glycerin header 214 or passed to pretreatment component of facility 100. [00109] One such use may be to pretreat a feed provided by line 200 to dehydrate the feed. If the feed contains free fatty acids or phospholipids, its introduction into the pretreatment component rather than via line 200, may be preferred.
  • apparatus 400 is adapted to effect alkanol exchange on a refined biodiesel stream, which for purposes of discussion is a methyl biodiesel, provided by line 402 to reactor 406.
  • oil soluble catalyst is provided by line 404 and admixed with the biodiesel prior to passage to reactor 406.
  • catalyst may be separately added.
  • the catalyst for purposes of this illustration, is titanium tetraisopropoxide.
  • Higher alkanol is supplied by line 408 to reactor 406. Again, for purposes of this discussion, the higher alkanol is isopropanol.
  • Reactor 404 may be of any suitable design including mechanically or sonically agitated vessel, or static mixer containing static mixing devices such as a packed bed, baffles, orifices, venturi nozzles, tortuous flow path, or other impingement structure.
  • the reactor is maintained under transesterification conditions including the use of reflux temperature for the lower alkanol co-product, methanol.
  • the transesterification conditions may be at, above or below ambient pressure sufficient to provide a temperature sufficient to effect the transesterification, e.g., between about 35°C and 15O 0 C, and effect vaporization of methanol. Isopropyl biodiesel is formed.
  • the catalyst may equilibrate in the reaction menstruum to provide both titanium tetramethoxide and titanium tetraisopropoxide.
  • distillation column 410 The vapors from reactor 406 pass into distillation column 410. Distillation column 410 may be trayed or packed. The overhead is passed via line 414 to condenser 416. Methanol is withdrawn via line 418. A reflux is provided via line 420. Typically the distillation including reflux ratio is sufficient to maintain a concentration of higher alkanol, isopropanol, in the lower alkanol in line 418 of less than about 5, preferably less than about 2, mass percent.
  • a liquid phase containing alkanol exchanged biodiesel is withdrawn via line
  • Water is provided via line 428 to the alkanol exchanged biodiesel in line 422 and the mixture is passed to distillation column 424.
  • the amount of water used is sufficient to react on a stoichiometric basis the catalyst to titanium carboxylate and alkanol.
  • Distillation column 424 may be a flash column or may be trayed or packed. For purposes of this discussion, the decomposition of the catalyst occurs in the base of column 424. A separate reactor could be used, if desired, for the catalyst destruction.
  • the overhead from column 424 comprises methanol and isopropanol and is passed via line 426 to condenser 416.
  • the higher boiling fraction which contains alkanol-exchanged biodiesel is passed via line 430 to filter 440. Biodiesel meeting specifications is withdrawn via line 442 and solids, titanium dioxide, is removed via line 450.

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  • Engineering & Computer Science (AREA)
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Abstract

Le point de trouble du biodiesel d'ester alkyle inférieur d'acide gras est réduit par transestérification du biodiesel avec un alcanol supérieur.
PCT/US2008/076630 2007-09-19 2008-09-17 Procédé de fabrication de biodiesel à point de trouble réduit WO2009039151A1 (fr)

Applications Claiming Priority (2)

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US60/994,454 2007-09-19

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Cited By (1)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US20110239529A1 (en) * 2010-03-31 2011-10-06 Texaco Inc. Biodiesels useful for improving cloud point

Citations (3)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US20030139599A1 (en) * 2000-05-29 2003-07-24 Gerhard Nestler Method for the production of esters of unsaturated carboxylic acids
US20070039239A1 (en) * 2003-09-15 2007-02-22 Forester David R Low temperature operable fatty acid ester fuel composition and method thereof
US20070066838A1 (en) * 2005-09-16 2007-03-22 Gerard Hillion Method of manufacturing fatty acid ethyl esters from triglycerides and alcohols

Patent Citations (3)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US20030139599A1 (en) * 2000-05-29 2003-07-24 Gerhard Nestler Method for the production of esters of unsaturated carboxylic acids
US20070039239A1 (en) * 2003-09-15 2007-02-22 Forester David R Low temperature operable fatty acid ester fuel composition and method thereof
US20070066838A1 (en) * 2005-09-16 2007-03-22 Gerard Hillion Method of manufacturing fatty acid ethyl esters from triglycerides and alcohols

Cited By (2)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US20110239529A1 (en) * 2010-03-31 2011-10-06 Texaco Inc. Biodiesels useful for improving cloud point
US8709107B2 (en) * 2010-03-31 2014-04-29 Chevron U.S.A. Inc. Biodiesels useful for improving cloud point

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