U.S. patent application number 12/679112 was filed with the patent office on 2010-09-30 for processes for the esterification of free fatty acids and the production of biodiesel.
This patent application is currently assigned to Best Energies, Inc. Invention is credited to Donald Leroy Bunning, Louis A. Kapicak, Thomas Arthur Maliszwski, David James Schreck.
Application Number | 20100242346 12/679112 |
Document ID | / |
Family ID | 40468302 |
Filed Date | 2010-09-30 |
United States Patent
Application |
20100242346 |
Kind Code |
A1 |
Bunning; Donald Leroy ; et
al. |
September 30, 2010 |
PROCESSES FOR THE ESTERIFICATION OF FREE FATTY ACIDS AND THE
PRODUCTION OF BIODIESEL
Abstract
The effluent from an acid esterification of free fatty acid with
alkanol to produce alkyl ester of fatty acid is contacted with
glycerin to remove water and alkanol. The alkanol separated with
the glycerin can be recycled to the acid esterification by
contacting the glycerin with fatty acid-containing feed being
passed to the acid esterification.
Inventors: |
Bunning; Donald Leroy;
(South Charleston, WV) ; Kapicak; Louis A.; (Cross
Lanes, WV) ; Maliszwski; Thomas Arthur; (Charleston,
WV) ; Schreck; David James; (Lake City, MN) |
Correspondence
Address: |
PAULEY PETERSEN & ERICKSON
2800 WEST HIGGINS ROAD, SUITE 365
HOFFMAN ESTATES
IL
60169
US
|
Assignee: |
Best Energies, Inc
Madison
WI
|
Family ID: |
40468302 |
Appl. No.: |
12/679112 |
Filed: |
September 17, 2008 |
PCT Filed: |
September 17, 2008 |
PCT NO: |
PCT/US08/76613 |
371 Date: |
March 19, 2010 |
Related U.S. Patent Documents
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Application
Number |
Filing Date |
Patent Number |
|
|
60994455 |
Sep 19, 2007 |
|
|
|
Current U.S.
Class: |
44/388 |
Current CPC
Class: |
C11C 3/003 20130101;
Y02E 50/13 20130101; Y02E 50/10 20130101; C07C 67/08 20130101; C07C
67/08 20130101; C07C 69/52 20130101; C07C 67/08 20130101; C07C
69/24 20130101 |
Class at
Publication: |
44/388 |
International
Class: |
C10L 1/19 20060101
C10L001/19 |
Claims
1. A process for the esterification of feed containing free fatty
acid comprising: a. contacting the feed with a stoichiometric
excess of alkanol under acidic esterification conditions including
the presence of acid catalyst at elevated temperature, for a time
sufficient to provide an esterification effluent comprising alkyl
ester of the free fatty acid, water and unreacted alkanol; b.
contacting at least a portion of the esterification effluent with
sufficient glycerin to form a glycerin-containing phase and an oil
phase comprising alkyl ester, said contacting being for a time and
under conditions sufficient that said oil phase has a lower water
and a lower alkanol concentration than the esterification effluent
and said glycerin phase contains water and alkanol; and c. phase
separating the oil phase and the glycerin phase.
2. The process of claim 1 wherein the feed comprises glyceride
containing free fatty acid.
3. The process of claim 1 wherein the glycerin phase contains at
least about 60 percent of the water contained in the esterification
effluent.
4. The process of claim 3 wherein the glycerin phase contains at
least about 30 mass percent of the alkanol contained in the
esterification effluent.
5. The process of claim 1 wherein the alkanol comprises
methanol.
6. The process of claim 1 wherein the catalyst comprises sulfuric
acid.
7. The process of claim 1 wherein at least a portion of the
glycerin phase of step (c) is contacted with the feed prior to step
(a) under conditions sufficient to provide a feed phase containing
an increased concentration of alkanol and a spent glycerin phase
having a decreased concentration of alkanol; phase separating the
feed phase and the spent glycerin phase; and providing the feed
phase to step (a).
8. A process for the esterification of feed containing free fatty
acid comprising: a. contacting the feed with a glycerin solution
containing alkanol under conditions sufficient to provide a feed
phase containing an increased concentration of alkanol and a spent
glycerin phase having a decreased concentration of alkanol; b.
phase separating the feed phase and the spent glycerin phase; and
c. contacting the feed phase with a stoichiometric excess of
alkanol under acidic esterification conditions including the
presence of acid catalyst at elevated temperature, for a time
sufficient to provide an esterification effluent comprising alkyl
ester of the free fatty acid, water and unreacted alkanol.
9. The process of claim 8 wherein alkyl ester is added to the feed
prior to step (a) to increase the solubility of alkanol in the
feed.
10. The process of claim 8 wherein the alkanol comprises
methanol.
11. A process for the esterification of feed containing free fatty
acid comprising: a. contacting the feed with a stoichiometric
excess of alkanol under acidic esterification conditions including
the presence of glycerin-soluble, acid catalyst at elevated
temperature, for a time sufficient to provide an esterification
effluent comprising alkyl ester of the free fatty acid, water, acid
catalyst and unreacted alkanol; b. contacting at least a portion of
the esterification effluent with glycerin whereby a
glycerin-containing phase and an oil phase comprising alkyl ester
are formed, said contacting being for a time and under conditions
sufficient that the oil phase has a lower acid catalyst
concentration than the esterification effluent and a lower water
concentration than the esterification effluent and said glycerin
phase contains acid catalyst, water and alkanol; c. phase
separating the oil phase and the glycerin phase; d. contacting the
glycerin phase with soaps of free fatty acids to generate free
fatty acids; e. separating free fatty acids from said glycerin
phase; and f. recycling the free fatty acids to step (a).
12. A process for the esterification of feed containing free fatty
acid comprising: a. contacting the feed with a stoichiometric
excess of alkanol under acidic esterification conditions including
the presence of an oil-soluble, acid catalyst at elevated
temperature, for a time sufficient to provide an esterification
effluent comprising alkyl ester of the free fatty acid, acid
catalyst, water and unreacted alkanol; b. providing sufficient
alkanol in said esterification effluent to enable to form an
alkanol phase containing alkanol, water and oil-soluble acid
catalyst and an oil phase containing alkyl ester and unreacted free
fatty acid; c. phase separating the alkanol phase and the oil
phase; and d. passing at least a portion of the alkanol phase to
step (a).
13. The process of claim 12 wherein the oil-soluble, acid catalyst
comprises an organosulfonic acid.
14. The process of claim 12 wherein water is separated from at
least a portion of the alkanol phase of step (c) to provide a
higher boiling catalyst fraction.
15. The process of claim 14 wherein at least one of free fatty
acid, glyceride and glycerin is provided during distillation in an
amount sufficient to maintain the oil-soluble, acid catalyst in a
liquid phase.
16. A process for the esterification of feed containing at least
about 5 mass percent free fatty acid comprising: a. contacting the
feed with a stoichiometric excess of alkanol under acidic
esterification conditions including the presence of acid catalyst
at elevated temperature below about 120.degree. C., for a time
sufficient to convert between about 50 and 95 mass percent of the
free fatty acids to alkyl esters and provide an esterification
effluent comprising alkyl ester of the free fatty acid, free fatty
acid, water and unreacted alkanol; b. contacting at least a portion
of the esterification effluent with sufficient basic glycerin to
saponify free fatty acid to form soaps of said free fatty acid and
provide an oil phase comprising alkyl ester and a glycerin phase
comprising glycerin and soaps; and c. phase separating the oil
phase and glycerin phase.
17. The process of claim 16 wherein the acidic esterification
conditions comprise a temperature below about 120.degree. C.
18. A process for esterification of feed containing free fatty acid
comprising: a. contacting the feed with a stoichiometric excess of
alkanol, under acidic esterification conditions including the
presence of acid catalyst that is soluble in alkanol and
substantially insoluble in alkyl ester of fatty acid to provide an
esterification effluent comprising an alkanol and acid
catalyst-containing phase and a first oil phase containing alkyl
ester of the free fatty acid, water and unreacted alkanol; b. phase
separating the alkanol and acid catalyst-containing phase and the
first oil phase; c. contacting the first oil phase with glycerin to
provide a second oil phase having a reduced concentration of water
and alkanol and a glycerin-containing phase containing water and
alkanol; d. admixing at least a portion of the glycerin-containing
phase with the alkanol and catalyst-containing phase to provide an
admixture; and e. recovering alkanol from the admixture by vapor
fractionation.
19. The process of claim 18 wherein the alkanol comprises
methanol.
20. The process of claim 19 wherein the catalyst comprises sulfuric
acid.
Description
FIELD OF THE INVENTION
[0001] This invention pertains to processes for making alkyl
esters, especially biodiesel, from feeds containing free fatty
acids and to the esterification by acid catalysis of free fatty
acids with lower alkanol.
BACKGROUND TO THE INVENTION
[0002] Biodiesel is being used as an alternative or supplement to
petroleum-derived diesel fuel. Biodiesel is a mixture of alkyl
esters which can be made from various bio-generated oils and fats
from vegetable and animal sources.
[0003] One process for making biodiesel involves the
transesterification of triglycerides in the oils or fats with a
lower alkanol in the presence of a base catalyst to produce alkyl
ester and a glycerin co-product. Unfortunately, most oils and fats
useful as triglyceride-containing feeds for transesterification
also contain free fatty acids which are not converted under typical
transesterification conditions to biodiesel. Moreover, biodiesel
must meet demanding product specifications. See, for instance, ASTM
D 6751, American Society for Testing and Materials. These
specifications, among other requirements, limit the amount of free
fatty acid that can be contained in biodiesel.
[0004] Free fatty acids, while not acceptable in biodiesel, can be
converted to esters suitable for inclusion in biodiesel. Numerous
processes have been proposed. See, for instance, U.S. Pat. No.
6,822,105; U.S. Patent Application Publication No. 2005/0204612;
Canakci, et al., Transactions of ASAE, 42, 5, pp. 1203-10 (1999),
King, "Esterification: Chemistry and Processing", Biodiesel Short
Course, Quebec City, Canada, May 12-13, 2007, and Van Gerpen, et
al., "Biodiesel Production Technology, August 2002-January 2004,
National Renewable Energy Laboratory NREL/SR-510-36244, July
2004.
[0005] The esterification of free fatty acids with alkanol results
in water being generated as a co-product. As the acid
esterification is equilibrium limited reaction, a large excess of
alkanol is typically used to shift the equilibrium towards the
production of the sought alkyl ester. The unreacted alkanol must be
recovered from the esterification product and recycled to provide
an economically attractive process.
[0006] Turck in U.S. Pat. No. 6,538,146 discloses a method for
producing fatty acid esters of alkyl alcohols using oils that
contain free fatty acids and phosphatides. He summarizes his
process as treating the feed with a base mixture of glycerin and a
catalyst to produce a two phase mixture with the neutralized free
fatty acids passing into the glycerin phase. The oil phase
containing the triglycerides is then subjected to
transesterification. See column 2, lines 35 et seq. At column 4,
lines 41 et seq., Turck poses that the free fatty acids can be
separated per WO 95/02661 and then subjected to esterification with
an alcohol. The esterified product can be added to the
transesterification mixture.
[0007] Koncar, et al., in U.S. Pat. No. 6,696,583 disclose methods
for preparing fatty acid alkyl esters in which fatty acids
contained in a glycerin phase from a transesterification are
separated and mixed with an esterification mixture containing
triglycerides and is subjected to esterification to form fatty acid
esters. The object of their process is to process the fatty acid
phase in the untreated state, i.e., without purification and
removal of sulfuric acid. The esterification product is then
transesterified with alcohol. Koncar, et al, refer to EP-A-0 708
813 as disclosing the esterification of free fatty acids at column
2, lines 26 to 34.
[0008] Iyer in U.S. 2006/0293533 discloses a process for the
esterification and transesterification of fats and oils using one
or more heterogeneous catalysts. See also, Clements, US
2006/0224006.
[0009] Lin, et al., in U.S. Pat. No. 7,122,688 disclose the use of
acidic mesoporous silicates as catalysts for esterifying fatty
acids and transesterifying oils.
[0010] Various processes are commercially offered for making
biodiesel by transesterification of triglycerides. Several of these
processes provide options for the esterification of free fatty
acids to the corresponding esters. Desmet Ballestera have a process
in which feedstocks preferably containing more than 1 percent free
fatty acid is subjected to vacuum-steam stripping to remove free
fatty acids. The distillate can then be subjected to esterification
conditions comprising elevated temperature, methanol and sulfuric
acid catalyst to make a methyl ester. The esterification product is
subjected to a flash and phase separation to recover methanol for
recycle and separate water, glycerin and sulfuric acid from the
methyl ester. Kemper, Desmet Ballestra Biodiesel Production
Technology, Biodiesel Short Course, Quebec City, Canada, May 12-13,
2007.
[0011] Crown Iron Works Company also provides a biodiesel
manufacturing process where acid esterification is used to convert
free fatty acids to methyl esters for biodiesel. They caution that
acid esterification should only be used if disposal of the fatty
acids or soaps thereof is not economic or possible or the feed used
generates a lot of fatty acids. They note that acid esterification
increases capital and production costs, and that sulfuric acid
creates sulfates which increase the removal cost from glycerin.
Waranica, Crown Iron Works Biodiesel Production Technology,
Biodiesel Short Course, Quebec Canada.
SUMMARY
[0012] This invention provides improved processes for the
esterification of feeds containing free fatty acids especially
glycerides feeds containing free fatty acids wherein the glycerides
in the feed are suitable to be transesterified with alkanol to
produce biodiesel. The esterification of the free fatty acids is
conducted with a stoichiometric excess of alkanol to provide an
esterification effluent containing alkyl esters, unreacted alkanol
and water. In accordance with processes of this invention, at least
a portion of the esterification effluent is contacted with glycerin
to reduce the concentration of water and alkanol in the
esterification effluent containing alkyl esters, and the glycerin
is separated by phase separation.
[0013] In one broad aspect of the processes of this invention, a
feed containing free fatty acid is subjected to acidic
esterification conditions in the presence of a stoichiometric
excess of alkanol to provide an esterification effluent containing
alkyl ester, water and unreacted alkanol. The esterification
effluent is contacted with glycerin to form a two phase mixture.
Water and alkanol are extracted into the glycerin phase which can
be phase separated from the oil phase comprising alkyl ester. The
processes of this aspect of the invention are particularly useful
in conjunction with facilities to produce biodiesel from glycerides
by transesterification as glycerin is available as a co-product of
the transesterification. Additionally, where the acid catalyst is
contained in the esterification effluent, the glycerin phase can
remove the acid catalyst from the esterification effluent.
[0014] In further detail, this broad aspect of the invention for
esterifying feed containing free fatty acid, especially free fatty
acid of from about 8 to 30, say 14 to 24, carbon atoms, comprises:
[0015] a. contacting the feed with a stoichiometric excess of
alkanol, preferably a molar ratio of alkanol to free fatty acid of
at least about 1.1:1, say, about 2:1 to 20:1 or 30:1 or more, under
acidic esterification conditions including the presence of acid
catalyst at elevated temperature, e.g., up to 200.degree. C. or
more, preferably less than about 150.degree. C., say, less than
about 120.degree. C., and more preferably between about 35.degree.
C. and 100.degree. C., for a time sufficient to provide an
esterification effluent comprising alkyl ester of the free fatty
acid, water and unreacted alkanol; [0016] b. contacting at least a
portion of the esterification effluent containing alkyl ester with
sufficient glycerin to form a glycerin-containing phase and an oil
phase comprising alkyl ester, said contacting being for a time and
under conditions sufficient that said oil phase has a lower water
and a lower alkanol concentration than the esterification effluent
and said glycerin phase contains water, preferably at least about
50, and sometimes at least about 60, percent of the water contained
in the esterification effluent, and alkanol, preferably at least
about 20, preferably at least about 30, mass percent of the alkanol
in contained in the esterification effluent; and [0017] c. phase
separating the oil phase and the glycerin phase.
[0018] The glycerin used for the contacting may be derived from any
suitable source. One suitable source is glycerin-containing
co-product from the transesterification of glycerides. Typically,
the glycerin used for the contacting with the esterification
effluent comprises at least about 40, preferably at least about 50,
mass percent glycerin. Often the mass ratio of glycerin to
esterification effluent is at least about 0.01:1, more preferably
from about 0.05:1 to 0.5:1.
[0019] All or a portion of the esterification effluent is contacted
with glycerin. Where a portion of the esterification effluent is
contacted, that portion, which may be an aliquot portion or a
portion remaining after a separation unit operation, contains alkyl
ester. For instance, the esterification effluent may be subjected
to a stripping operation to remove some of the water and alkanol.
Where higher molar ratios of alkanol to free fatty acid are used
for the esterification, an alkanol-containing phase may form. This
alkanol-containing phase may be separated by phase separation prior
to the contacting with the glycerin. Preferably the contacting is
under conditions that minimize reversion of alkyl ester to free
fatty acid or conversion to glycerides. Generally these conditions
are provided by removal or inactivation of at least a portion of
the catalyst. Although temperature reduction may also suffice to
reduce reversion or conversion, it is usually not necessary.
Frequently the contacting is at a temperature of from about
35.degree. C. to 150.degree. C. for 0.01 to 10 hours.
[0020] In a preferred embodiment of the invention, after the
contacting and phase separation, the glycerin phase containing
alkanol is contacted with the feed to be subjected to
esterification whereby the feed extracts from the glycerin phase a
portion of the alkanol. In a further aspect of the invention, the
acidic esterification is integrated with a transesterification
process to make biodiesel from glycerides. Not only can the
transesterification process provide glycerin for removal of water
and alkanol from the esterification effluent but also integration
can enhance the economics of the process by reducing energy
consumption and capital expense.
[0021] Preferably the alkanol is lower alkanol of up to 6,
preferably 1 to 3, carbon atoms, especially methanol. In another
preferred embodiment of this aspect of the invention the feed
comprises glycerides of fatty acid.
[0022] In another broad aspect of the processes of this invention,
a portion of the alkanol for the acid esterification of free fatty
acids is obtained from a glycerin-containing liquid that contains
alkanol. In this aspect of the invention, the glycerin-containing
liquid is contacted with feed for the esterification and a portion
of the alkanol partitions to the feed. In processes for making
biodiesel from glycerides, not only is glycerin a co-product, but
also, glycerin contains unreacted alkanol used in the
transesterification. Hence, the processes of this aspect of the
invention provide a low energy and low capital means for recovering
alkanol from the glycerin co-product. Alternatively, or in
addition, the glycerin may be that used to remove alkanol from the
esterification effluent.
[0023] In this broad aspect, the processes of this invention for
the esterification of feed containing free fatty acid, especially
free fatty acid of from about 8 to 30, say 14 to 24, carbon atoms,
comprise: [0024] a. contacting the feed with a glycerin solution
containing alkanol, preferably where the solution contains a mass
ratio of alkanol to glycerin of at least about 0.05:1, say, 0.2:1
to 2:1, under conditions sufficient to provide a feed phase
containing an increased concentration of alkanol and a spent
glycerin phase having a decreased concentration of alkanol,
preferably reduced by at least about 20, more preferably at least
about 30, mass percent; [0025] b. phase separating the feed phase
and the spent glycerin phase; and [0026] c. contacting the feed
phase with a stoichiometric excess of alkanol under acidic
esterification conditions including the presence of acid catalyst
at elevated temperature, for a time sufficient to provide an
esterification effluent comprising alkyl ester of the free fatty
acid, water and unreacted alkanol.
[0027] If desired, alkyl ester can be added to enhance the
solubility of alkanol in the feed, especially where the feed
comprises glyceride. The amount of alkyl ester added can be
relatively minor yet significant additional alkanol solubility can
be obtained. Often the alkyl ester is provided in a mass ratio to
feed of at least about 0.01:1, say, about 0.05:1 to 0.2:1. The
solubility of alkanol in the oil phase will depend, among other
things, upon the type of alkanol and the content of free fatty acid
in the feed. The alkyl ester may be from any suitable source
including, but not limited to, alkyl ester from the esterification
process and biodiesel.
[0028] In yet another broad aspect of the invention, the acid
esterification is conducted using a glycerin-soluble, acid
catalyst. The esterification effluent, which contains alkyl ester
of free fatty acid, water, acid catalyst and unreacted alkanol, is
contacted with glycerin to remove acid catalyst from the
esterification effluent. Simultaneously or sequentially the acidic
glycerin stream can be contacted with soaps of free fatty acids to
generate free fatty acids which can be fed to the esterification.
For instance, the glycerin used for contacting the esterification
effluent may already contain soaps. Alternatively, or in addition,
the glycerin, after contact with the esterification effluent may
contact a soap-containing stream.
[0029] By this aspect of the invention, catalyst that would
otherwise be discarded is effectively used to provide free fatty
acids. For instance, the acidic esterification may be conducted to
effect only a partial conversion of free fatty acid in the feed,
e.g., between about 50 and 95 or 97 mass percent of the free fatty
acid is converted to alkyl esters. The unreacted free fatty acids
can be saponified and removed from the oil phase, and then
acidified for recycle. An advantage of this aspect of the invention
is that the esterification need not achieve a high conversion per
pass of free fatty acid. Thus additional reactor stages or high
alkanol to free fatty acid ratios that have typically been used to
achieve high conversion, can be avoided to save in capital and
energy costs. Also, if the esterification is used in conjunction
with a base transesterification of glyceride, soaps formed as a
by-product can be recovered using the acidic glycerin as free fatty
acids, and the free fatty acids can be subjected to acidic
esterification. As often the glycerin co-product from a base
catalyzed transesterification also contains base catalyst, the
glycerin co-product may be used as at least a portion of the basic
glycerin to saponify free fatty acids in the esterification
effluent as well as provide soaps from the transesterification.
[0030] In this broad aspect, the processes of this invention for
the esterification of feed containing free fatty acid, especially
free fatty acid of from about 8 to 30, say 14 to 24, carbon atoms,
comprise: [0031] a. contacting the feed with a stoichiometric
excess of alkanol under acidic esterification conditions including
the presence of glycerin-soluble, acid catalyst at elevated
temperature, for a time sufficient to provide an esterification
effluent comprising alkyl ester of the free fatty acid, acid
catalyst, water and unreacted alkanol; [0032] b. contacting the
esterification effluent with glycerin whereby a glycerin-containing
phase and an oil phase comprising alkyl ester are formed, said
contacting being for a time and under conditions sufficient that
the oil phase has a lower acid catalyst concentration than the
esterification effluent and a lower water concentration than the
esterification effluent and said glycerin phase contains acid
catalyst (which may be unneutralized, partially neutralized or
essentially completely neutralized), water and alkanol; [0033] c.
phase separating the oil phase and the glycerin phase; [0034] d.
contacting the glycerin phase with soaps of free fatty acids to
generate free fatty acids, preferably unneutralized or partially
neutralized acid catalyst generates at least a portion of the free
fatty acids; [0035] e. separating free fatty acids from said
glycerin phase; and [0036] f. recycling the free fatty acids to
step (a).
[0037] The separation of free fatty acid from glycerin of step (e)
may be conducted in any convenient manner. For instance, free fatty
acid may form an oil phase and can be removed by phase separation.
Preferably the feed is contacted with the glycerin phase of step
(e) to effect separation of the free fatty acids.
[0038] In a fourth broad aspect of the processes of this invention,
an oil-soluble, acid catalyst is used for the acidic esterification
and is recovered from the esterification effluent by phase
separation of an alkanol phase from the oil phase. The alkanol
phase is formed by providing alkanol in an amount in excess of that
which is miscible with the oil phase. The acid catalyst, due to the
more polar nature of the alkanol phase, preferentially partitions
to the alkanol phase. Thus not only can the acid catalyst be
recovered, e.g., for recycle to the esterification, but also, the
recovery involves low capital and energy costs. Water will also
tend to preferentially partition to the alkanol phase. If desired,
all or a portion of the alkanol phase can be subjected to a water
removal unit operation. As the volume of the alkanol phase can be
minor in comparison to the esterification effluent, the energy
required for the water removal, e.g., by distillation, is not
unduly high.
[0039] In this broad aspect of the processes of this invention,
esterification of feed containing free fatty acid, especially free
fatty acid of from about 8 to 30, say 14 to 24, carbon atoms,
comprises: [0040] a. contacting the feed with a stoichiometric
excess of alkanol under acidic esterification conditions including
the presence of an oil-soluble, acid catalyst, preferably an
organic acid having at least about 4 carbon atoms, preferably an
organosulfonic acid, at elevated temperature, for a time sufficient
to provide an esterification effluent comprising alkyl ester of the
free fatty acid, water, acid catalyst and unreacted alkanol; [0041]
b. providing sufficient alkanol in said esterification effluent to
enable to form an alkanol phase containing alkanol, water and
oil-soluble acid catalyst and an oil phase containing alkyl ester
and unreacted free fatty acid; [0042] c. phase separating the
alkanol phase and the oil phase; and [0043] d. passing at least a
portion of the alkanol phase to step (a).
[0044] The alkanol required for forming the alkanol phase may be
present in step (a) or may, preferably, be through the addition of
alkanol subsequent to step (a) such that a single, homogenous phase
exists in step (a). The amount of alkanol used to form the separate
alkanol phase is preferably sufficiently small that the separate
alkanol phase is only a minor portion of the esterification
effluent, and sometimes is less than about 10, even less than about
5, volume percent of the total of the alkanol phase and the oil
phase. Preferably water is removed from at least a portion of the
alkanol phase, e.g., by distillation. Where the alkanol boils at a
lower temperature than water or co-boils with water, it may be
desirable to provide high boiling liquid such that the oil-soluble
catalyst is maintained in a liquid phase to facilitate handling and
avoid any decomposition of the oil-soluble catalyst.
[0045] In a fifth broad aspect of the invention, the esterification
is conducted under esterification temperatures less than about
120.degree. C. in the presence of acidic catalyst that is soluble
in alkanol and substantially insoluble in alkyl ester of fatty
acid. The alkanol is preferably provided in an amount sufficient to
provide an esterification effluent having an alkanol and catalyst
phase and an alkyl ester-containing phase. In this aspect of the
invention, the esterification processes only convert a portion of
the free fatty acid in the feed to alkyl ester, e.g., from about 50
to 95 or 97 mass percent. The alkyl-ester-containing phase is
contacted with basic glycerin to convert unreacted free fatty acid
to soaps and a glycerin phase containing the soaps is separated
from the alkyl ester-containing phase. Accordingly, attractive
energy costs can be obtained. Moreover, capital savings can be
achieved in that adequate conversion of free fatty acids can often
be achieved in a few reaction stages, and sometimes even a single
reaction stage.
[0046] Most significantly in the processes of this aspect of the
processes of this invention, mild reaction conditions including
lower temperatures and, potentially lower concentrations of acid
catalyst, attenuate the relative rate of the hydrolysis of esters
to acids as compared to the rate of acid esterification. Often the
concentration of acidic catalyst is below about 1, and sometimes
less than about 0.5, mass percent. Hence, flexibility is provided
by the processes of this aspect of the invention to balance rate
and catalyst concentration while still attenuating the relative
rate of hydrolysis.
[0047] Processes for esterification of this fifth aspect of the
invention use a glycerides-containing feed containing at least
about 5 mass percent free fatty acid and comprise: [0048] a.
contacting the feed with a stoichiometric excess of alkanol,
preferably with a molar ratio of alkanol to free fatty acid between
about 10:1 to 30:1, under acidic esterification conditions
including the presence of acid catalyst that is soluble in alkanol
and substantially insoluble in alkyl ester of fatty acid,
preferably sulfuric acid, at elevated temperature below about
120.degree. C., more preferably below about 100.degree. C., say
35.degree. C. to 100.degree. C., for a time sufficient to convert
between about 50 and 95 or 97 mass percent of the free fatty acids
to alkyl esters and provide an esterification effluent comprising
alkyl ester of the free fatty acid, free fatty acid, water and
unreacted alkanol; [0049] b. contacting the esterification effluent
with sufficient basic glycerin to saponify free fatty acid to form
soaps of said free fatty acid and provide an oil phase comprising
alkyl ester and a glycerin phase comprising glycerin and soaps; and
[0050] c. phase separating the oil phase and glycerin phase.
[0051] In a preferred embodiment using a glyceride-containing feed,
the esterification effluent contains between about 0.5 and 3 mass
percent free fatty acid. Most preferably, the free fatty acid in
the esterification effluent is at least sufficient to neutralize at
least about 80 mole percent of the base in the basic glycerin.
Advantageously the basic glycerin removes water from the
esterification effluent to provide an oil phase containing less
than about 0.1 mass percent water.
[0052] In a sixth broad aspect of this invention, the
esterification is conducted in the presence of acidic catalyst that
is soluble in alkanol and substantially insoluble in alkyl ester of
fatty acid and the alkanol is provided in an amount sufficient to
provide an esterification effluent having an alkanol and catalyst
phase and an alkyl ester-containing phase. The
alkyl-ester-containing phase is contacted with glycerin to reduce
the concentration of alkanol and water in the alkyl ester and form
a glycerin-containing phase and an oil phase containing alkyl
ester. The glycerin phase is separated and at least a portion is
admixed with the alkanol and catalyst phase and alkanol is
selectively recovered from the admixture by vapor
fractionation.
[0053] Processes for esterification of this sixth aspect of the
invention for the esterification of feed containing free fatty
acid, especially free fatty acid of from about 8 to 30, say, 14 to
24, carbon atoms, comprises: [0054] a. contacting the feed with a
stoichiometric excess of alkanol, preferably with a molar ratio of
alkanol to free fatty acid between about 10:1 to 30:1, under acidic
esterification conditions including the presence of acid catalyst
that is soluble in alkanol and substantially insoluble in alkyl
ester of fatty acid, preferably sulfuric acid, to provide an
esterification effluent comprising an alkanol and acid
catalyst-containing phase and a first oil phase containing alkyl
ester of the free fatty acid, water and unreacted alkanol; [0055]
b. phase separating the alkanol and acid catalyst-containing phase
and the first oil phase; [0056] c. contacting the first oil phase
with glycerin to provide a second oil phase having a reduced
concentration of water and alkanol and a glycerin-containing phase
containing water and alkanol; [0057] d. admixing at least a portion
of the glycerin-containing phase with the alkanol and
catalyst-containing phase to provide an admixture; and [0058] e.
recovering alkanol from the admixture by vapor fractionation.
BRIEF DESCRIPTION OF THE DRAWINGS
[0059] FIG. 1 is a schematic representation of an integrated
esterification and transesterification biodiesel facility using the
processes of this invention.
[0060] FIG. 2 is a schematic representation of a two stage
esterification reactor system useful in the processes of this
invention.
[0061] FIG. 3 is a schematic representation of another integrated
esterification and transesterification biodiesel facility using the
processes of this invention.
[0062] FIG. 4 is a schematic representation of an acid
esterification unit operation using an organic soluble catalyst
having preferential solubility in an alkanol phase.
[0063] FIG. 5 is a schematic representation of an acid
esterification unit operation in which glycerin which has been used
to recover alkanol from alkyl ester is subjected to vapor
fractionation for alkanol recovery.
DETAILED DISCUSSION
Esterification Conditions
[0064] The processes of this invention pertain to the acidic
esterification of free fatty acids, particularly those containing
about 8 to 30, say 14 to 24, carbon atoms, to form alkyl esters of
the free fatty acids. The alkyl esters can find various utilities.
With emphasis on developing renewable fuels, the demands for
biodiesel through the base transesterification of glycerides from
plant and animal oils and fats have created a need to convert free
fatty acids in those oils and fats to alkyl esters for inclusion in
the biodiesel product.
[0065] The feed for the esterification may be an oil phase
comprising substantially all free fatty acid or a composition
containing free fatty acid and other components including, but not
limited to, triglycerides. In general, the oil phase may contain
from about 1 or 2 to essentially 100, mass percent free fatty acid.
Where the free fatty acid is derived from an oil or fat-containing
component, the feed may be subjected to unit operations to
selectively remove the free fatty acid which then is the feed to
the esterification processes. Such removal may be effected in any
suitable manner as is known in the art. For instance, the oil or
fat may be contacted with base to saponify the free fatty acid
which can then be removed by phase separation or extraction, e.g.,
with glycerin, ethylene glycol, or the like. Alternatively, the oil
or fat may be used as the feed to the esterification. In the latter
circumstances, the free fatty acid may comprise up to about 60 mass
percent (dry basis) of the oil or fat depending upon the specific
oil or fat. Feeds may also contain phospholipids which may be as
much as about 2 to 5 mass percent (dry basis) of the feeds. The
balance of the fats and oils is largely fatty acid triglycerides.
The unsaturation of the free fatty acids and triglycerides may also
vary over a wide range.
[0066] Examples of oils or fats derived from bio sources,
especially vegetable oils and animal fats, include, but are not
limited to rape seed oil, soybean oil, cotton seed oil, safflower
seed oil, castor bean oil, olive oil, coconut oil, palm oil, corn
oil, canola oil, jatropha oil, rice bran oil, tobacco seed oil,
fats and oils from animals, including from rendering plants and
fish oils. Mixtures of two or more oils and fats can be used.
[0067] Esterification conditions include the presence of alkanol,
elevated temperature and the presence of acid catalyst. Broadly,
esterification is conducted with alkanol, which may be a diol, but
preferably is a monoalkanol, having a primary --OH, under
esterification conditions. The preferred alkanols are lower
alkanols, especially those having 1 to 3 carbon atoms, although
butanol and isobutanol and higher alkanols are operable. Most
preferably the alkanol is methanol which has the highest
reactivity. Ethanol can be used but may pose separation
difficulties if the esterification product is used to make methyl
biodiesel.
[0068] The molar ratio of alkanol to free fatty acid can vary
widely. As the acidic esterification is an equilibrium-limited
reaction, a stoichiometric excess of alkanol is typically used.
Where esterification is sought, the molar ratio of alkanol to free
fatty acid is generally between about 0:5:1 to 30:1, and preferably
between about 2:1 to 25:1, and most preferably between about 3:1 to
20:1. Preferably the alkanol is methanol and is present in an
amount that exceeds the solubility of methanol in the oil phase and
thus forms a separate phase in the reaction zone. A portion of
water present in the reaction menstruum can partition to the
methanol-containing phase. The alkanol also aids in increasing the
minimum concentration of water co-product and glycerin from any
transesterification that forms a separate water or
glycerin-containing phase. However, advantageous operation includes
the use of sufficient alkanol to form a separate phase containing
both alkanol and alkanol-soluble catalyst.
[0069] Esterification conditions include the presence of acidic
catalyst and elevated temperature, e.g., generally between about
30.degree. C. and 200.degree. C. High temperatures are often
unnecessary to achieve high conversions and thus temperatures in
the range of about 30.degree. C. or 40.degree. C. to 150.degree.
C., and sometimes, 60.degree. C. to 100.degree. C. or 120.degree.
C., provide sufficient conversions of fatty acids with relatively
short residence times. Preferred esterification temperatures are
below about 120.degree. C., to attenuate the reaction rate of water
with ester. The desired esterification temperatures will depend in
part, upon the other acidic esterification conditions including the
strength of the acid catalyst and its concentration. The reaction
pressure can be any suitable pressure, e.g., from about 10 to 5000,
preferably from about 90 to 1000, kPa absolute. Preferably an
inerting gas such as nitrogen, hydrocarbon gas such as methane or
carbon dioxide is used during the esterification.
[0070] The pressure for the acidic esterification is preferably
sufficient to maintain a liquid phase. The esterification may be
conducted under conditions such that the alkanol and water are
removed by vaporization or may be under conditions such that the
reaction occurs in the liquid phase.
[0071] The catalyst can be heterogeneous or homogeneous. Where
heterogeneous, it may be a solid or a highly dispersed liquid
phase. Any suitable acid catalyst (Bronsted acid or Lewis acid) for
the esterification of free fatty acids can be used including
homogeneous and heterogeneous catalysts. The preferred acid
catalysts are mineral acids such as hydrochloric acid, sulfurous
acid, sulfuric acid, phosphoric acid, and phosphorous acid. However
other strong acids including organic and inorganic acids can be
used. Examples of strong organic acids include alkyl sulfonic acids
such as methylsulfonic acid; alkylbenzene sulfonic acids such as
toluene sulfonic acid; naphthalenesulfonic acid; and
trichloroacetic acid. Solid acid catalysts include NAFION.RTM.
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.
[0072] The amount of acid catalyst provided can vary over a wide
range. Typically the catalyst is provided in a catalytically
effective amount of at least about 0.1 mass percent based upon the
feed. Where soaps are present, the amount of acid should be
sufficient to convert such soaps to free fatty acids. Often the
acid is present in an amount of at least about 0.2 to 5, say, 0.25
to 2, mass percent based upon the feed above that required to
convert any soaps to free fatty acids. Solid heterogeneous
catalysts are typically provided in greater amounts. Oil soluble
catalysts tend to be more active which is believed to be due to the
dispersion of the catalyst in the oil phase. Preferred oil soluble,
acid catalysts are those having organic substituents of at least
about 4 carbon atoms, e.g., from 6 to 24 carbon atoms, especially
sulfonic acids such as toluene sulfonic acid and naphthalene
sulfonic acid.
[0073] The residence time for the esterification will depend upon
the amount of free fatty acid present, the conversion sought, the
type and amount of catalyst used, the reactivity and amount of
alkanol as well as the temperature of the process, and the type of
reactor and extent of mixing. Residence times thus can range from
less than 1 minute to over 1000 minutes. The residence times
frequently are in the range of about 5 minutes to 120 minutes,
preferably in the range of about 10 minutes to 90 minutes. Often,
the reactivity of alkanol and the residence time is sufficient to
convert at least about 30 mole percent, and preferably at least
about 50 mole percent, and sometimes at least about 75 mole percent
to essentially all, preferably between about 75 and 95 or 98 or
even 99, mass percent of the free fatty acid to ester.
[0074] The esterification may be conducted in one or more stages.
If desired, the effluent from one reaction stage may be subjected
to a unit operation to remove water.
The Drawings
[0075] Processes for making biodiesel will be further described in
connection with FIG. 1 which schematically depicts biodiesel
manufacturing facility 100. Facility 100 is provided with a
transesterification component (generally designated by numerals in
the 200 series) as well as pretreatment components (generally
designated by numerals in the 100 series) and a refining component
generally (designated by numerals in the 300 series).
[0076] Pretreatment by Esterification
[0077] As shown in FIG. 1, a glyceride feed containing free fatty
acid can be provided to facility 100 via line 102 for pretreatment
by acid. Line 104 is provided in the event that more than one feed
is desired to be processed simultaneously in the esterification
section. Catalyst, which for purposes of this discussion, is
sulfuric acid, is provided via line 114.
[0078] The feed may be directly introduced into esterification
reactor 106, or as shown, is subjected to a contact with an alkanol
laden stream of glycerin to strip alkanol from the glycerin into
the oil-containing feed phase. This contact will be described
later.
[0079] The preferred conditions for the esterification will depend
upon the nature of the feed and the apparatus type and
configuration. Reactor 106 may comprise one or more stages or
vessels and separation unit operations may be located between each
stage or vessel. Where reactor 106 is staged, it is often
desirable, but not essential, to remove water between stages to
enhance conversion of free fatty acid to esters. Reactor 106 may be
a vessel or a length of pipe. But preferably other types of vessels
are used such as mechanical and sonically agitated reactors, and
reactors with static mixing such as reactors containing contact
structures such as trays, packing, baffles, orifices, venturi
nozzles, tortuous flow path, and other impingement structures.
Suitable reactors include those providing high intensity mixing,
including high shear.
[0080] The oil phase from the esterification section of facility
100 often contains at least about 0.5, say between about 0.5 and 2
or 3, mass percent free fatty acid. This free fatty acid serves to
neutralize at least a portion of the base catalyst contained in a
spent glycerin stream produced in the transesterification and base
pretreatment sections of facility 100. Preferably, the molar ratio
of free fatty acid in the oil phase from the esterification to mole
of base in the glycerin phase introduced into base reactor 134 as
discussed below will be at least about 0.3:1, often at least about
0.7:1 up to about 1:1. The use of ratios of free fatty acid to base
catalyst of greater than 1:1 can adversely affect the performance
of the base pretreatment. A number of advantages flow from this
preferred embodiment. For instance, the equipment and conditions
required for the esterification section need not be of the type
required for essentially complete conversion of the free fatty
acids, resulting in capital and operating cost savings. Since
residual free fatty acid is converted to soap and removed in the
base pretreatment section, the feed to the transesterification
section can be substantially devoid of free fatty acid which
adversely affects the base catalyst therein. Additionally, the
neutralized spent glycerin stream from the base pretreatment
section can be used effectively for enhancing phase separation and
water and catalyst removal from the esterification product.
[0081] During the esterification in reactor 106 some conversion of
glycerides to esters may occur. The esters, diglycerides and
monoglycerides essentially remain in the oil phase. Some glycerin
will be produced as a result of the transesterification of the
glycerides in the feed. The extent of such conversion is not
critical but results in reduced requirements of alkanol and
catalyst in the transesterification section per unit of biodiesel
produced as well as enabling increased performance such as rate of
conversion and extent of conversion to be obtained. Generally up to
about 20 mass percent, say, between about 0.1 to 15, and sometimes
between 5 to 10, mass percent of the glyceride-containing feed is
transesterified during acid esterification.
[0082] The esterification reaction product from reactor 106 is
passed via line 108 to phase separator 110. Phase separator 110 is
optional depending upon whether or not two phases exist. In some
instances, an oil layer containing glycerides and fatty ester and a
water-containing layer form. The water-containing layer can contain
more polar components such as glycerin, water-soluble catalyst, and
alkanol. As shown, a neutralized spent glycerin stream from the
base pretreatment section is provided via line 170A and contacted
with the esterification product. The spent glycerin aids in the
extraction of water and water-soluble phosphorus compounds.
Additionally, the glycerin assists in making the phase separation.
In this embodiment, the amount of glycerin added can vary widely.
As relatively small amounts of water are produced during the acid
esterification of free fatty acids, beneficial results can be
achieved with relatively little spent glycerin being added. Often
the spent glycerin added is less than about 20, preferably between
about 0.5 and 10, mass percent of the stream from esterification
reactor 106. A separate phase may exist in reactor 106, e.g., from
catalyst such as sulfuric acid, or water co-produced during the
esterification or even alkanol above that miscible with the oil
phase. Glycerin can aid in forming a defined phase containing,
e.g., catalyst and water. As used herein, the formation of a
glycerin phase or providing a glycerin phase contemplates that
there may, or may not, be separate phases in the fluid contacted
with glycerin. Spent glycerin that is in a separate phase may be
separated and removed via line 112.
[0083] Phase separator 110 may be of any suitable design including
a decanter, a phase separation facilitated decanter that contains
coalescing sites, and a centrifuge. The lower, water-containing
fraction exits separator 110 via line 112. This fraction contains
some alkanol, water, water-soluble catalyst, glycerin and
water-soluble phosphorus compounds.
[0084] The oil fraction of separator 110 contains virtually no
sulfuric acid, often some alkanol, relatively little water,
unreacted free fatty acids, if any, fatty ester and glycerides. The
fraction is passed via line 118 from separator 110 to fractionation
column 120 to provide an overhead fraction containing alkanol and a
bottoms stream containing oil. The overhead from column 120 can be
recycled to esterification reactor 106 via line 122. Make up
alkanol is provided via line 124.
[0085] The fractionation column may be of any suitable design
including a flash column, stripping column, falling film
evaporator, or trayed or packed column. If desired, more than one
fractionation column can be used with one effecting separation of
water from alkanol. Similarly a side draw 116 may be taken from
distillation column 120 for the removal of water, and fractionation
column may be a divided wall column to enhance such separation. In
an embodiment, a substantial portion of the water is removed by the
phase separation in phase separator 110, and fractionation column
does not separately recover water. Water will be contained in both
the overhead and bottoms stream from column 120. However, the
relatively small amount of water in the overhead can be recycled
with alkanol via line 122 to reactor 106 without undue adverse
effect. Water contained in the bottoms passes to the base
pretreatment section and is removed from the oil phase therein.
[0086] In another embodiment, only a portion of the alkanol is
removed by fractionation in column 120. The alkanol remaining in
the oil phase is passed to the base pretreatment section. In the
base pretreatment section alkanol can be reacted with glyceride to
form esters and can be recovered in the spent glycerin phase for
recycle to the esterification section. Thus, the capital and
operating costs for fractionation column 120 can be reduced. Often
the bottoms stream from fractionation column 120 contains between
about 0.1 to 10, say, between about 0.5 and 5, e.g., 0.5 to 2, mass
percent alkanol. In yet another embodiment, the oil-containing
fraction from separator 110 can be passed directly to separator 128
or base reactor 134.
[0087] While shown as processing the oil phase from separator 110,
fractionation column 120 may be positioned between esterification
reactor 106 and separator 110 and serve to recover alkanol from the
esterification product exiting reactor 106.
[0088] Pretreatment by Base
[0089] 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.
[0090] As shown in the facility of FIG. 1, a glyceride-containing
feed stream is provided by line 132 to base reactor 134. The feed
stream may comprise a fresh glyceride-containing feed.
Alternatively or in addition, the feed stream may comprise the oil
phase from the esterification provided via lines 126 and 130. To
base reactor is also provided a glycerin and base
catalyst-containing stream via line 142 which will be further
discussed below. Preferably a non-acidic inerting gas such as
nitrogen or hydrocarbon gas such as methane is used during base
pretreatment.
[0091] In base reactor 134, free fatty acids contained in the feed
stream are reacted with base catalyst to form soaps. If the free
fatty acid content of the feed stream requires more than the amount
of base catalyst introduced via line 142 for the desired degree of
saponification, additional base can be added via line 133. The
additional base may be the same or different from that comprising
the catalyst, and may be one or more of alkali metal hydroxides or
alkoxides and alkaline earth metal hydroxides, oxides or alkoxides,
including by way of examples and not in limitation, sodium
hydroxide, sodium methoxide, potassium hydroxide, potassium
methoxide, calcium hydroxide, calcium oxide and calcium
methoxide.
[0092] To the extent that phospholipids are present in the feed
stream to base reactor 134, at least a portion is chemically
reacted, e.g., by a hydration or by a salt formation, to provide
chemical compounds preferentially soluble in glycerin.
[0093] 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.degree. C., say, 35.degree. C. to 150.degree. C., and most
frequently between about 40.degree. C. and 80.degree. 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.
[0094] 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.
[0095] The base reaction product from reactor 134 contains
glycerin, glycerides, soaps, water, and fatty acid ester and is
passed via line 136 to separator 128. Separator 128 serves to
separate the less dense oil layer from the more dense glycerin
layer. The soaps and reacted phospholipids preferentially pass to
the glycerin layer as does most of the water. The oil layer
preferably contains less than about 0.5 mass percent soaps. Phase
separator 128 may be of any suitable design including a decanter, a
phase separation facilitated decanter that contains coalescing
sites, and, if needed, a centrifuge.
[0096] The glycerin phase is withdrawn from separator 128 via line
137 and may be sent to glycerin recovery or another application. If
the glycerin layer contains significant amounts of soaps, it may be
desirable to recycle the soaps to esterification reactor 106 for
conversion to fatty esters. As shown, a portion or all of the
glycerin phase may be passed via line 170 to acidification reactor
172 where soaps are converted to free fatty acids. At least a
portion of this glycerin phase is passed via line 170A to provide
the glycerin to assist in the separation of water, water-soluble
catalyst (or salts thereof) from the esterification product in
phase separator 110. The glycerin-containing phase from separator
110 is passed via line 112 to line 170. Also as shown, a portion of
the glycerin phase in line 172 is recycled to reactor 134 via line
170B. The recycle can serve several purposes. For instance,
hydrated phospholipids are returned to reactor 134 where they may
undergo transesterification to recover additional fatty acid ester.
Also, any base contained in the recycled glycerin stream is
available for saponification of free fatty acids.
[0097] Unless acid contained in the esterification effluent of line
108 is neutralized prior to being passed to separator 110, the
glycerin-containing phase from separator 110 will contain
water-soluble acid which can be used as acid for acidification
reactor 172. Acid can also be provided via line 174. Acidification
reactor 174 may be one or more vessels of any suitable design
including a length of pipe and other types of vessels such as
mechanical and sonically agitated reactors, and reactors with
static mixing such as reactors containing contact structures such
as trays, packing, baffles, orifices, venturi nozzles, tortuous
flow path, and other impingement structures. Suitable reactors
include those providing high intensity mixing, including high
shear. The acidification conditions usually encompass a temperature
in the range of about 20.degree. C. to 150.degree. C., a pressure
from about 1 to 5000, preferably 90 to 1000, kPa absolute, and a
residence time of from about 1 second to 5 hours. Suitable acids
include mineral acids and organic acids, but typically a readily
available acid such as sulfuric or phosphoric acid is used. The
amount of acid is usually sufficient to convert substantially all
the soaps to free fatty acid. The use of excess acid is not
deleterious to the formation of the free fatty acids, but can
entail additional expense. Accordingly the molar ratio of
acidifying acid function to soaps is in the range of about 1:1 to
1.5:1. Generally the pH of the glycerin stream is less than about
6, say, between about 1 and 5, e.g., 2 and 4. The acidity of the
glycerin stream is determined by diluting the glycerin stream to 50
volume percent water and measuring the pH.
[0098] The glycerin stream from acidification reactor is passed via
line 176 to contact vessel 178 into which glyceride-containing feed
is provided via line 102. In contact vessel 178 the glycerin stream
is contacted with fresh feed which serves to extract a portion of
the alkanol from the glycerin phase. The contact with the glycerin
also serves to remove water from the feed. Removal of water assists
in the esterification of free fatty acids in esterification reactor
106 as the esterification is an equilibrium-limited reaction
affected by water concentration.
[0099] 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.degree. C.
to 150.degree. 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.
[0100] The ability to recover alkanol from glycerin by extraction
with fresh feed can effectively be used to use glycerin as a
complementary means for recycling unreacted alkanol to reactor 106.
FIG. 1 shows two glycerin loops for alkanol recovery and recycle to
the esterification reactor. The first loop involves the glycerin
layer from separator 110 and the second, the glycerin layer from
separator 128.
[0101] The fluid mixture from contact vessel 178 is passed via line
180 to phase separator 182. In phase separator 182, a glyceride and
free fatty acid oil layer is produced that is passed via line 184
to esterification reactor 106. A glycerin-containing layer is
discharged via line 186 and contains water, acidification acid, and
soluble phosphorus compound. Separator 182 may be of any suitable
design including a decanter, a phase separation facilitated
decanter that contains coalescing sites, and, if necessary, a
centrifuge. Contact vessel 178 and decanter 182 may be a single
vessel, including but not limited to, a countercurrent extraction
column.
[0102] If the esterification product from esterification reactor
106 has a sufficiently low free fatty acid content and low
phospholipids content, another option is to eliminate separator 110
and fractionation column 120 and provide the esterification product
in line 108 directly to separator 128 or base reactor 134.
[0103] Returning to separator 128, the oil phase is withdrawn and
passed via line 138 to second pretreatment reactor 139. Second
pretreatment reactor 139 and third pretreatment reactor 148 are
adapted to recover alkanol contained in the glycerin from the
transesterification component of facility 100 through reaction,
e.g., transesterification and extraction into the
glyceride-containing phase. A base transesterification process is
used in these pretreatment reactors. While two reactors are shown,
the number of reactors will depend upon the sought consumption of
the alkanol as well as the efficiency of the reactors. Hence one,
two, or three or more pretreatment reactors may be used. Also, the
pretreatment reactor can comprise a number of stages in a single
vessel which could be a countercurrent contact vessel.
Advantageously, the feed stream to the alkanol consumption
pretreatment reactors is relatively free from free fatty acids so
as to prevent undue consumption of the base catalyst. Typically the
pretreatment reactors provide a glycerin stream from which most of
the alkanol has been removed. Often, the alkanol content of the
glycerin discharged from base reactor 134 is less than about 5, and
preferably less than about 2, mass percent.
[0104] In an alternative mode of operation, a significant portion
of the alkanol is contained in line 126 (or line 108 if separator
110 and distillation column 120 are not used) and passed to
separator 128. The concentration of alkanol in the
glycerin-containing stream in line 170 may be higher than 5 mass
percent, and alkanol is recovered be partitioning to the
glyceride-containing feed in contact vessel 178. The alkanol
content of the glycerin may be sufficiently low that no
distillation is required to recover alkanol yet the overall process
to make biodiesel can still exhibit high efficiencies.
[0105] Second pretreatment reactor 139 also receives the glycerin
phase from the third pretreatment reactor. This glycerin phase
contains glycerin, base catalyst, and alkanol. Second pretreatment
reactor 139 is maintained under base transesterification conditions
including the presence of base catalyst provided by the glycerin
phase feed and elevated temperatures, often between about
30.degree. C. and 220.degree. C., preferably between about
30.degree. C. and 80.degree. C. to provide a second pretreatment
product. The pressure is typically in the range of between about 90
to 1000 kPa (absolute) although higher and lower pressures can be
used. The reactor is typically batch, semi-batch, plug flow or
continuous flow tank. Preferably other types of vessels are used
such as mechanical and sonically agitated reactors, and reactors
with static mixing such as reactors containing contact structures
such as trays, packing, baffles, orifices, venturi nozzles,
tortuous flow path, and other impingement structures. Suitable
reactors include those providing high intensity mixing, including
high shear. However, depending upon the presence of soaps and
phospholipids, care needs to be taken so as not to generate a
product that cannot be readily separated by phase separation. The
residence time will depend upon the desired degree of conversion of
the contained alkanol, the ratio of alkanol to glyceride, reaction
temperature, the degree of agitation and the like, and is often in
the range of about 0.1 to 20, say, 0.5 to 10, hours.
[0106] 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.
[0107] As depicted, line 142 is provided with holding tank 142A.
Holding tank 142A can serve as a reservoir and enables the rate
that glycerin, which contains base, is provided to base reactor
134, to be varied with changes in free fatty acid content of the
esterification product. It also can permit additional reaction of
glycerides with alkanol contained in the glycerin phase to occur
prior to introduction into base reactor 134 where catalyst is
consumed by conversion of free fatty acids to soaps.
[0108] The upper oil phase is removed from separator 140 via line
144 and is passed to line 146 which also receives the glycerin
co-product from transesterification from line 248. The combined
streams are passed to third pretreatment reactor 148. The stream is
provided by line 146 and contains in addition to glycerin, alkanol,
base catalyst, and usually some water and soaps. Table I sets forth
typical compositions of the stream in line 248. The compositions,
of course, will depend upon the operation of the
transesterification component as well as which of the
glycerin-containing streams from the transesterification component
are used. The typical concentrations are based upon combining all
glycerin-containing streams and operating under preferred
parameters.
TABLE-US-00001 TABLE I Component Broad, Mass % Typical, Mass %
Glycerin 40 to 80 50 to 70 Alkanol (Methanol) 15 to 50 25 to 45
Catalyst (NaOCH.sub.3) 0.2 to 5 0.5 to 5 Soaps 0.1 to 5 0.5 to 5
Water 0.01 to 0.5 0.05 to 0.3 Oil (glycerides and alkyl 0 to 5 0.5
to 2 esters)
[0109] Third pretreatment reactor 148 is maintained under base
transesterification conditions including the presence of base
catalyst provided by the glycerin--containing feed and elevated
temperatures, often between about 30.degree. C. and 220.degree. C.,
preferably between about 30.degree. C. and 80.degree. C. to provide
a first pretreatment product. Base catalyst in the
transesterification component tends to partition to the glycerin
phase and often adequate catalyst is provided for the base
pretreatment section in the glycerin co-product from the
transesterification section provided by line 248. In some
instances, however, it may be desired to add additional base
catalyst to third pretreatment reactor 148 or any preceding base
pretreatment reactor. The pressure is typically in the range of
between about 90 to 1000 kPa (absolute) although higher and lower
pressures can be used. The reactor is typically batch, semi-batch,
plug flow or continuous flow tank with some agitation or mixing.
The preferred types of vessels are mechanical and sonically
agitated reactors, and reactors with static mixing such as reactors
containing contact structures such as trays, packing, baffles,
orifices, venturi nozzles, tortuous flow path, and other
impingement structures. Suitable reactors include those providing
high intensity mixing, including high shear. However, depending
upon the presence of soaps and phospholipids, care needs to be
taken so as not to generate a product that cannot be readily
separated by phase separation. The residence time will depend upon
the desired degree of conversion, the ratio of alkanol to
glyceride, reaction temperature, the degree of agitation and the
like, and is often in the range of about 0.1 to 20, say, 0.5 to 10,
hours.
[0110] Typically the transesterification in third pretreatment
reactor 148 recovers through transesterification and extraction to
the glyceride-containing phase at least about 20, preferably at
least about 30, and more preferably at least about 50, mass percent
of the alkanol fed to the reactor. Any unreacted alkanol in the oil
phase will be carried with the oil phase to the transesterification
component of facility 100. Often the total amount of alkanol
recovered from the glycerin-coproduct from transesterification
using all pretreatment stages is at least about 50, and sometimes
at least about 80, mass percent.
[0111] 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.
[0112] Facility 100 includes a chiller 158 to remove high molecular
weight glycerides, waxes and esters that are insoluble at the
chiller temperature. Some feeds, such as crude corn oil, contain
high molecular weight glycerides and esters. The hydrocarbyl
moieties in these high molecular weight components typically have
between 30 and 40 carbon atoms. If they remain in the oil, the
resultant biodiesel product tends to have unacceptably high cloud
points and gel points. As shown, the oil phase from separator 152
passes through line 156 to chiller 158. Chiller 158 is maintained
at a temperature sufficient to cause high molecular weight and
other components that lead to and increase in gel point temperature
to solidify. Typically this temperature is between about 0.degree.
C. and 20.degree. C. In some instances, cooling will tend to remove
monoglycerides and diglycerides. Cooling below the desired
temperature and then warming to a temperature to liquefy the mono-
and di-glycerides while still maintaining a solid wax, can minimize
loss of components that can be converted to biodiesel. The chilled
oil phase is then passed via line 160 to centrifuge 162 to remove
higher density components including solids and any remaining
glycerin phase. The higher density fraction is discharged via line
164. Rather than using a centrifuge, the solids can be filtered
from the glyceride-containing stream. Filter aids can be used if
desired. A producer composition is provided by centrifuge 162 and
is provided to line 166.
[0113] Chiller 158 is optional, and chillers may also be used
elsewhere in facility 100 to remove waxes. For instance, a chiller
may be used to treat fresh feed in line 102 or can be used to treat
biodiesel product from the refining component.
[0114] If desired all or a portion of the producer composition in
line 166 may be withdrawn via line 168 as an intermediate product
for storage or sale as a feedstock for transesterification. Line
168 also provides the feed for the transesterification component of
facility 100 by introducing the producer composition into line
200.
[0115] Transesterification
[0116] Line 200 provides glyceride-containing feed to first
transesterification reactor 202. Line 200 can also supply
additional glyceride-containing feed. Preferably the additional
feed is relatively free of free fatty acids and phospholipids such
as refined oils sourced from rape seed, soybean, cotton seed,
safflower seed, castor bean, olive, coconut, palm, corn, canola,
fats and oils from animals, including from rendering plants and
fish oils.
[0117] Alkanol for the transesterification is supplied to first
transesterification reactor via line 206. The alkanol is preferably
lower alkanol, preferably methanol, ethanol or isopropanol with
methanol being the most preferred. The alkanol may be the same or
different from the alkanol provided to esterification reactor 106
via line 124. Although line 206 is depicted as introducing alkanol
into line 200, it is also contemplated that alkanol can be added
directly to reactor 202 at one or more points. Generally the total
alkanol (line 206 and from the producer composition of line 166) is
in excess of that required to cause the sought degree of
transesterification in reactor 202. Preferably, the amount of
alkanol is from about 101 to 500, more preferably, from about 110
to 250, mass percent of that required for the sought degree of
transesterification in reactor 202. In facility 100 three reactors
are depicted as being used. One reactor may be used, but since the
reaction is equilibrium limited, most often at least two and
preferably three reactors are used. Often, where more than one
reactor is used, at least about 60, preferably between about 70 and
96, percent of the glycerides in the feed are reacted in first
transesterification reactor 202. It is possible to provide all the
alkanol required for transesterification to first
transesterification reactor 202, or a portion of the alkanol can be
provided to each of the transesterification reactors.
[0118] The base catalyst is shown as being introduced via line 204
to first transesterification reactor 202. The amount of catalyst
used is that which provides a desired reaction rate to achieve the
sought degree of transesterification in first transesterification
reactor 202. Preferably, catalyst is provided to each of the
transesterification reactors since base catalyst preferentially
partitions to the glycerin phase and is removed with phase
separation of the glycerin after each transesterification reactor.
The amount of catalyst used will be in excess of that required to
react with the amount of free fatty acid contained in the feed oil,
which due to the pretreatment, will be relatively little. The base
catalyst may be an alkali or alkaline earth metal hydroxide or
alkali or alkaline earth metal alkoxide, especially an alkoxide
corresponding to the lower alkanol reactant. Preferred alkali
metals are sodium and potassium. When the base is added as a
hydroxide, it may react with the lower alkanol to form an alkoxide
with the generation of water which in turn results in the formation
of free fatty acid. Another type of catalyst is an alkali metal or
alkaline earth metal glycerate. This catalyst converts to the
corresponding alkoxide of the alkanol reactant in the reaction
menstruum. Alternatively, the catalyst may be a heterogeneous base
catalyst. Catalyst may need to be separately provided to the base
pretreatment reactors if the base catalyst, e.g., a heterogeneous
or oil soluble catalyst, is not carried with the co-product
glycerin in the transesterification component to the base
pretreatment reactors. However, homogeneous catalysts that have
solubility in glycerin are preferred where the pretreatment
component is used since the catalyst serves as at least a portion
of the base used therein. The exact form of the catalyst is not
critical to the understanding and practice of this invention. For
the purposes of the following discussion, homogenous base catalyst
is used. Preferably a non-acidic inerting gas such as nitrogen or
hydrocarbon gas such as methane is used during base
transesterification.
[0119] Often the transesterification is at a temperature between
about 30.degree. C. and 220.degree. C., preferably between about
30.degree. C. and 80.degree. C. The pressure is preferably
sufficient to maintain a liquid phase reaction menstruum and
typically is in the range of between about 90 to 1000 kPa
(absolute) although higher and lower pressures can be used. First
transesterification reactor 202 is typically batch, semi-batch,
plug flow or continuous flow tank with some agitation or mixing.
Preferably the reactors are mechanical and sonically agitated
reactors. Reactors with static mixing such as reactors containing
contact structures such as trays, packing, baffles, orifices,
venturi nozzles, tortuous flow path, and other impingement
structures can be used. Suitable reactors include those providing
high intensity mixing, including high shear. As stated above, one
of the advantages of the processes of this invention is that the
producer compositions do not require an induction period for the
transesterification reaction to initiate. Accordingly plug flow
reactors have enhanced viability. The residence time will depend
upon the desired degree of conversion, the ratio of alkanol to
glyceride, reaction temperature, the base catalyst concentration,
the degree of agitation and the like, and is often in the range of
about 0.02 to 20, say, 0.1 to 10, hours.
[0120] The partially transesterified effluent from reactor 202 is
passed via line 208 to phase separator 210. Phase separator 210 may
be of any suitable design including a decanter, a phase separation
facilitated decanter that contains coalescing sites, and,
optionally, a centrifuge. A glycerin-containing bottoms phase is
provided in the separator and is removed via line 212 and is passed
to glycerin header 214. As depicted, this stream is used as a
portion of the glycerin for the pretreatment component of facility
100. This glycerin phase also contains any soaps made in reactor
202 and a portion of the catalyst. The soaps can be recovered from
this stream in acidifying reactor 172 as discussed above. The
lighter phase contains alkyl esters and unreacted glycerides and is
passed via line 216 to second transesterification reactor 218. A
rag layer may form in separator 210. The rag layer may contain
unreacted glycerides, alkyl esters, alkanol, soaps, catalyst and
glycerin. An advantage of the process set forth in FIG. 1 is that
withdrawing the rag layer with the glycerin phase does not result
in a loss of glycerides, alkyl esters, alkanol, and catalyst since
the glycerin phase is passed to the pretreatment component of
facility 100.
[0121] Reactor 218 may be of any suitable design and may be similar
to or different than reactor 202. As shown, additional alkanol is
provided via line 206A, and additional catalyst is provided via
line 204A. Preferably the transesterification conditions in reactor
218 are sufficient to react at least about 90, more preferably at
least about 95, and sometimes at least about 97 to 99.9 or more,
mass percent of the glycerides in the feed to the
transesterification. The transesterification in reactor 218 is
typically operated under conditions within the parameters set forth
for reactor 202 although the conditions may be the same or
different. The residence time will depend upon the desired degree
of conversion, the ratio of alkanol to glyceride, reaction
temperature, the degree of agitation and the like, and is often in
the range of about 0.02 to 20, say, 0.1 to 10, hours.
[0122] The effluent from second transesterification reactor 218 is
passed via line 220 to phase separator 222 which may be of any
suitable design and may be the same as or different from the design
of separator 210. A heavier, glycerin-containing phase is withdrawn
via line 224 and passed to glycerin header 214. A lighter phase
containing crude biodiesel is withdrawn from separator 222 via line
226.
[0123] As depicted, third transesterification reactor 228 is used
and the crude biodiesel in line 226 is passed to this reactor. The
transesterification conditions in reactor 228 are sufficient to
provide essentially complete conversion, at least about 97 or 98 to
99.9, mass percent of the glycerides in the feed converted to alkyl
ester. As shown, additional alkanol is provided via line 206B, and
additional catalyst is provided via line 204B. The
transesterification in reactor 228 is typically operated under
conditions within the parameters set forth for reactor 202 although
the conditions may be the same or different. The residence time
will depend upon the desired degree of conversion. The reactor may
be of the type described for reactor 202. The residence time will
depend upon the desired degree of conversion, the ratio of alkanol
to glyceride, reaction temperature, the degree of agitation and the
like, and is often in the range of about 0.02 to 20, say, 0.1 to
10, hours. Advantageously, the transesterification product from
third transesterification reactor 228 contains less than about 1,
preferably less than about 0.8, and most preferably less than 0.5,
mass percent soaps based upon the total mass of alkyl esters and
soaps. The lighter phase also contains alkanol. In reactor 228 the
reaction proceeds quickly to completion by the addition of
additional alkanol and catalyst, and can be conveniently
accomplished by a plug flow reactor.
[0124] The overall molar ratio of alkanol to glycerides in the feed
to the reactors in the transesterification component, i.e., alkanol
provided by lines 206, 206A and 206B, can vary over a wide range.
Since transesterification is an equilibrium-limited reaction, the
driving force toward the alkyl ester and the conversion of
glycerides will be dependent upon the molar ratio of alkanol
equivalents to glycerides. Alkanol equivalents are alkanol and
alkyl group of the alkyl esters in the feed to the
transesterification component. On the basis of transesterfiable
substituents in the feed to the transesterification component, the
mole ratio of alkanol equivalents to glyceride in the feed to the
pretreatment component is frequently between about 3.05:1 to 15:1,
say 4:1 to 9:1. Advantageously, the pretreatment processes of this
invention permit the reuse of alkanol partitioned to the co-product
glycerin without intermediate vaporization. Often the amount of
total catalyst provided based on the mass of feed to the first
transesterification reactor, i.e., the catalyst provided by lines
204, 204A and 204B, is between about 0.3 and 1 mass percent
(calculated on the mass of sodium methoxide).
[0125] The effluent from third transesterification reactor 228 is
passed via line 230 to phase separator 232 which may be of any
suitable design and may be the same as or different from the design
of separator 210. A heavier, glycerin-containing phase is withdrawn
via line 234 and passed to glycerin header 214. A lighter phase
containing crude biodiesel is withdrawn from separator 232 via line
236. Alternatively, separator 232 can be eliminated provided that
in second transesterification reactor 218, the conversion of the
glycerides in the feed is at least about 90, preferably 92 to 96 or
98, percent. In some instances, the effluent from reactor 228 may
be a single phase containing relatively little glycerin. In some
instances it may be advantageous to use a centrifuge to separate
the glycerin phase from the oil phase following third
transesterification reactor 228.
[0126] Facility 100 contains an optional alkanol replacement
reactor 238. The alkanol replacement reactor serves to
transesterify the alkyl ester with a different alkanol. For
purposes of transesterification in reactors 202, 218 and 228, an
alkanol such as methanol provides not only attractive reaction
rates but also an effluent that is more easily separated than, say,
a reaction effluent where ethanol is the alkanol. In some instances
it may be desired to provide a biodiesel that contains fatty esters
in which the alkyl group of the fatty ester is branched in order to
reduce cloud and gel points. The transesterification between, say,
a fatty acid methyl ester, and higher molecular weight alkanol
results in methanol, rather than glycerin, being formed, and often
is more readily accomplished than the transesterification of
glyceride with that higher alkanol. The higher alkanols include
those having 2 to 8 or more carbon atoms, and are preferably
branched primary and secondary alkanols although tertiary alkanols
may find application but generally are less reactive. Examples of
higher alkanols include propanol, isopropanol, isobutanol,
2,2-dimethylbutan-1-ol, 2,3-dimethylbutan-1-ol, 2-pentanol, and the
like. Other alkanols include benzyl alcohol and 2 ethylhexanol.
[0127] Where an alkanol replacement operation is desired, it may be
located at various points in the process. For instance, the
replacement alkanol may be provided via line 206B to reactor 228,
or, as shown, it can follow reactor 228. In either case, alkanol
replacement transesterification can take advantage of catalyst
contained in the transesterification medium. Alternatively, alkanol
replacement may be effected on a biodiesel product by adding
catalyst. Thus, it can be located elsewhere in the refining
component of facility 100 including, but not limited to, treating
biodiesel in line 352.
[0128] The amount of higher alkanol provided via line 240 to
alkanol replacement reactor 238 can vary over a wide range.
Typically the molar ratio of higher alkanol to alkyl ester being
fed to reactor 238 is less than 0.5:1, e.g., from about 1:100 to
1:5. Often the alkanol replacement transesterification is at a
temperature between about 30.degree. C. and 220.degree. C.,
preferably between about 30.degree. C. and 80.degree. C. The
pressure is preferably sufficient to maintain a liquid phase
reaction menstruum and typically is in the range of between about
90 to 1000 kPa (absolute) although higher and lower pressures can
be used. Alkanol replacement reactor 238 can be batch, semi-batch,
plug flow or continuous flow tank with some agitation or mixing,
e.g., mechanically stirred, ultrasonic, static mixer containing
contact surfaces, e.g., trays, packing, baffles, orifices, venturi
nozzles, tortuous flow path, or other impingement structures. High
intensity mixing reactors, including high shear reactors, may also
be used. Preferred reactors are those in which the alkanol being
replaced is continuously removed. For instance, a reactive
distillation reactor can be used to continuously remove displaced
methanol from a transesterification of methyl ester and
isopropanol. As depicted, reactor 238 is a reactive distillation
unit and lower alkanol is withdrawn via line 330A and passed to the
transesterification reactors. Make-up alkanol is provided via line
332.
[0129] Where the alkanol replacement reactor is a batch reactor,
driving the replacement reaction to either essentially complete
conversion of the higher alkanol or essentially complete conversion
of the methyl ester to the higher alkanol ester (depending upon
whether the higher alkanol is provided below or at or above the
stoichiometric amount required for complete conversion), since the
vapor fractionation of methanol can continue until completion. With
continuous reactors, having unreacted methanol and higher alkanol
in the alkanol replacement product is likely. For purposes of this
discussion, a continuous alkanol replacement reactor is used.
[0130] Where the base catalyst has been removed from the fatty acid
ester of the lower alkanol, for instance, if the alkanol
replacement were to be conducted on a refined or partially refined
biodiesel, catalyst is provided. Suitable catalyst includes base
catalyst such as is used for transesterification. Since a single
liquid phase exists during the alkanol replacement unlike
transesterification where a glycerin layer forms, heterogeneous
catalysts and homogeneous catalysts having limited solubility in
the reaction menstruum can be used. Solid catalysts are preferred
to minimize or eliminate post treatment of the alkanol replacement
product, but good contact with catalyst is desirable to timely
achieve sought conversion. Homogeneous transesterification
catalysts such as titanium tetra-isopropoxide are also advantageous
as they are readily removed.
[0131] 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.
[0132] Refining
[0133] 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.
[0134] In a preferred process, an acid, preferably an organic acid,
is provided via line 302 in an amount sufficient to substantially
neutralize residual base catalyst contained in the crude biodiesel.
Inorganic acids such as sulfuric acid can be used as well as
organic acids, particularly those less volatile than the alkanol,
and acids that do not themselves or any potential reaction product
formed in contact with the crude biodiesel, form azeotropes with
the alkanol. Exemplary organic acids include acetic acid, citric
acid, oxalic acid, glycolic, lactic, free fatty acid and the like.
Generally the amount of catalyst contained in the crude biodiesel
is quite small as base catalyst preferentially partitions to the
glycerin phase. Accordingly, little acid is required to neutralize
sufficient catalyst to enable refining without risk of reversion of
alkyl ester. Often the amount of acid used is at least 0.95 times,
sometimes between about 1 and 3 times, that required to neutralize
the catalyst.
[0135] Crude biodiesel is passed via line 300 to an alkanol
separation unit operation. As shown, a two stage separation unit is
used. A single stage separator can be used if desired. The crude
biodiesel in line 300 is passed to first alkanol separator stage
304. Separator 304 is of any convenient design including a
stripper, wiped film evaporator, falling film evaporator, solid
sorbent, and the like. Preferably the fractionation is by fast,
vapor fractionation. Generally for a fast, vapor separation the
residence time is less than about one minute, preferably less than
about 30 seconds, and sometimes as little as 5 to 25 seconds.
Preferably the vapor fractionation conditions comprise a maximum
temperature of less than about 200.degree. C., preferably less than
about 150.degree. C., and most preferably, when the lower alkanol
is methanol, less than about 120.degree. C. Depending upon the
alkanol, the lower boiling fractionation may need to be conducted
under subatmospheric pressure to maintain desired overhead and
maximum temperatures. Where a falling film stripper is used, it may
be a concurrent or countercurrent flow stripper. Concurrent
strippers are preferred should there be a risk of undue
vaporization of alkanol at the point of entry of the crude
biodiesel. An inert gas such as nitrogen may be used to assist in
removing the alkanol.
[0136] The fast fractionation may be effected by any suitable vapor
fractionation technique including, but not limited to,
distillation, stripping, wiped film evaporation, and falling film
evaporation. Often the falling film evaporator has a tube length of
at least about 1 meter, say, between about 1.5 and 5 meters, and an
average tube diameter of between about 2 and 10 centimeters.
Usually the vapor fractionation recovers at least about 70,
preferably at least about 90, mass percent of the alkanol contained
in the crude biodiesel. Any residual alkanol is substantially
removed in any subsequent water washing of the crude biodiesel.
Advantageously, the amount of alkanol contained in the spent water
from the washing may be at a sufficiently low concentration that
the water can be disposed without further treatment. However, from
a process efficiency standpoint, alkanol can be recovered from the
spent wash water for recycle to the transesterification
reactors.
[0137] The lower boiling fraction containing the alkanol will
contain a portion of any water contained in the crude biodiesel.
Since the transesterification is conducted with little water being
present, and a portion of the water is removed with the glycerin,
the concentration of water in this fraction can be sufficiently low
that it can be recycled to the transesterification reactors. This
lower boiling fraction often contains less than about 1, and more
preferably less than about 0.5, mass percent water. Alternatively,
the lower boiling fraction may be passed to a methanol and water
distillation column in the esterification section of facility
100.
[0138] Alkanol is exhausted from first alkanol separator stage via
line 306 and may be exhausted from the facility as a by-product,
e.g., for burning or other suitable use, or can be recycled. Where
no alkanol replacement reaction is used, the alkanol will be the
lower alkanol for the transesterification and is recycled to the
transesterification section. The bottoms stream from first alkanol
separation stage 304 is passed via line 308 to second alkanol
separation stage 314 for additional alkanol recovery. The design of
second alkanol separation stage 314 may be similar to or different
than that of first alkanol separation stage 304 and may be operated
under the same or different conditions. Alkanol exits via line 316
and is combined with alkanol from line 306 and is passed to
condenser 318. In the process of facility 100, the condensed
alkanol will contain both the lower alkanol and the higher alkanol.
Condensed alkanol is recycled via line 330 to alkanol replacement
reactor 238. Non-condensed gases exit condenser 318 via line 320.
As shown, the alkanol separation operation is maintained under
vacuum conditions and these gases are passed to liquid ring vacuum
pump 322. The liquid for the liquid ring is provided via line 324
and exits via line 328. As the gases contain some alkanol, the
liquid for the liquid ring vacuum pump will remove alkanol from the
gases. The liquid may be water, in which case the water may need to
be treated to remove alkanol. Alternative liquid streams can be
used, including but not limited to glyceride-containing feed,
biodiesel, and glycerin. Feed is preferred as the liquid for the
liquid ring vacuum pump since it can be passed to a
transesterification reactor and alkanol contained therein used for
the transesterification. Gas is removed from liquid ring vacuum
pump 322 via line 326.
[0139] The bottoms stream from the second alkanol separation stage
exits via line 334 and is passed to separator 336 in which a
glycerin-containing phase and a biodiesel-containing phase are
separated. The presence of alkanol in the crude biodiesel enhances
the solubility of glycerin therein. Upon removal of the alkanol, a
separate glycerin-containing phase, which also contains soaps,
tends to form during the alkanol separation operation. The glycerin
fraction is removed from separator 336 via line 338 and can be
combined with spent glycerin in line 186. The lighter,
oil-containing phase is passed via line 340 to a water wash unit
operation. If desired, techniques can be used to assist in the
phase separation of glycerin in separator 336 such as adding an
effective amount of water to assist in the separation. Other
components useful in enhancing phase separations may also be used
including water-soluble inorganic salts that are essentially
insoluble in the biodiesel-containing phase. If desired, any
water-containing phase can be passed to evaporator 374.
[0140] Line 340 serves as a reactor and mixer where strong acid is
supplied. The amount of strong acid provided is sufficient to
convert any soaps remaining to free fatty acids. Sufficient strong
acid is used such that water used for washing the crude biodiesel
is at a suitably low pH. The strong acid is supplied in admixture
with a recycle stream in the wash operation as will be explained
later. While line 340 serves as an in-line mixer, a separate vessel
may be used for the acidification. Where a separate mixer is used,
it may be of any convenient design, e.g., a mechanically or
sonically agitated vessel, or static mixer containing static mixing
devices such as trays, packing, baffles, orifices, venturi nozzles,
tortuous flow path, or other impingement structure. In any event,
sufficient mixing and residence time should be provided such that
essentially all of the soaps are converted to free fatty acids.
Often the temperature during the mixing is in the range of about
30.degree. C. to 220.degree. C., preferably between about
60.degree. C. to 180.degree. C., and for a residence time of
between about 0.01 to 4, preferably 0.02 and 1, hours.
[0141] For purposes of discussion only and not in limitation, the
water wash operation uses a two stage water wash. Water wash
operation may be of any suitable design. Typically, the water wash
operates with a recycling water loop, often with the water recycle
being at least about 20, say between about 30 and 500, mass percent
of the crude biodiesel being fed to the column. Normally washing is
operated at a temperature between about 20.degree. C. and
120.degree. C., preferably between about 35.degree. C. and
90.degree. C. The amount of water provided to each wash vessel is
sufficient to effect a sought removal of glycerin, residual alkanol
and any water-soluble contaminants from the crude biodiesel.
Typically between about 20 and 200, preferably between about 30 and
100, mass parts of wash water are used per 100 mass parts of crude
biodiesel. Usually the free fatty acid is present in an amount less
than about 3000, most frequently less than about 2500, parts per
million by mass in the biodiesel product, and thus no need exists
to remove free fatty acid to provide a biodiesel product meeting
current commercial specifications. Preferably between about 1000
and 2500 ppm-m free fatty acid is contained in the biodiesel
product to aid in lubricity.
[0142] 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.
[0143] As shown, crude biodiesel is provided via line 340 to first
wash stage 342. For purposes of discussion, wash stage 342
comprises an agitated vessel to provide desired contact between the
oil and water phases and a decanter to effect separation.
Typically, the agitated vessel provides a contact time of about 1
second and 10 minutes, say, 5 to 60 seconds. Crude biodiesel is
contacted with acidic water from water loop 368. The washed
biodiesel from first wash stage 342 is passed via line 344 to
second wash stage 346 having a design similar to or different from
that of stage 342. This biodiesel is contacted with water from
water loop 364. In each stage the water, after contacting the
biodiesel stream being processed, is returned to the respective
loops. Acidic water is withdrawn from first wash stage 342 and
recycled via line 368. Substantially neutral water is withdrawn
from second wash stage 346 and recycled via line 364. Additional
water is provided to line 364 via line 376 which will be described
later.
[0144] As configured with separate water cycle loops, the pH of the
water in second wash stage 346 may be neutral or less acidic than
the water in first wash stage 342. Make-up water to line 368 is
provided by line 366. A purge is taken from line 368 via line 372.
The purge balances the amount of water in the wash loops and is at
a suitable rate to maintain desirably low concentrations of
impurities such as alkanol and glycerin in the water used for the
washing. The purge is usually at a rate of between about 1 and 50,
say 5 and 20, mass percent per unit time of the recycle rate in the
loop.
[0145] Line 370 provides strong acid to the water recycled via line
368 for combining with crude biodiesel in line 340 or being passed
to first wash stage 342. Adequate strong acid aqueous solution is
provided that the water in line 368 has a pH sufficiently low to
convert the soaps to free fatty acids. The acid may be any suitable
acid to achieve the sought pH such as hydrochloric acid, sulfuric
acid, sulfonic acid, phosphoric acid, perchloric acid and nitric
acid. Sulfuric acid is preferred due to cost and availability and
it is a non-oxidizing acid. The amount of strong acid aqueous
solution provided is typically in a substantial excess of that
required to convert the soaps to free fatty acid and to neutralize
any remaining catalyst. The excess of acid is often at least about
5, preferably at least about 10, say between about 10 and 1000
times that required. Consequently the feed to first wash stage 342
provides a wash water in line 368 having a pH of up to about 4,
preferably between about 0.1 and 4.
[0146] Returning to line 372, the purge water is passed to
evaporator 374 which provides a lower boiling fraction and a higher
boiling fraction. While an evaporator may be used, it is also
possible to use a packed or trayed distillation column with or
without reflux. Generally the bottoms temperature of evaporator 374
is less than about 150.degree. C., preferably between about
120.degree. C. and 150.degree. 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.
[0147] 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.
[0148] A washed biodiesel stream is withdrawn from second washing
stage 346 via line 348 and is passed to drier 350 to remove water
which exhausts via line 354. Preferably substantially all the
alkanol has been removed from the crude biodiesel prior to drying
to permit the water vapor to be exhausted without treatment to
eliminate volatile organic components. Drier 350 may be of any
suitable design such as stripper, wiped film evaporator, falling
film evaporator, and solid sorbent. Generally the temperature of
drying is between about 60.degree. C. and 220.degree. C., say,
about 70.degree. C. and 180.degree. 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.
[0149] The subatmospheric pressure is maintained in drier 350 by
the use of liquid ring vacuum pump 356 which is in communication
with line 354. Liquid ring vacuum pump 356 uses water as the
sealing fluid which is provided by line 358 and water exits via
line 362. The gases from liquid ring vacuum pump 356 exit via line
360.
[0150] Returning to glycerin header 214, the glycerin-containing
streams are passed via line 242 to blending tank 246 such that a
relatively uniform glycerin composition can be provided via line
248 to the pretreatment section of facility 100. Blending tank 246
may also provide sufficient residence time for any glycerides in
the glycerin to transesterify with alkanol as well as permit any
oil entrained in the glycerin phase to separate. As shown, an oil
layer that forms in blending tank 246 can intermittently or
continuously be withdrawn via line 247 for recycle to first
transesterification reactor 202. Alternatively, the oil layer can
be withdrawn with the glycerin and passed to the pretreatment
section.
[0151] While all glycerin-containing streams from the
transesterification and refining components of facility 100 have
been shown to be directed to glycerin header 214, it is within the
purview of the process to use fewer streams. As stated above, the
bottoms from evaporator 374 may be passed via line 382 to line 170
or added to header 214 or removed from the facility as a
by-product. Moreover, any of the glycerin-containing streams may be
used elsewhere prior to being passed to blending tank 246, and the
blended stream or a portion thereof in line 248 may be used
elsewhere and either returned to glycerin header 214 or passed to
pretreatment component of facility 100.
[0152] One such use may be to pretreat a feed provided by line 200
to dehydrate the feed. If the feed contains free fatty acids or
phospholipids, its introduction into the pretreatment component
rather than via line 200, may be preferred. In such a pretreatment,
a portion of the alkanol contained in the glycerin phase as well as
some of the base catalyst, will be partitioned to the oil
phase.
[0153] FIG. 2 depicts one type of esterification reaction system
400 useful in the processes of this invention. The reaction system
depicts two stages with glycerin treatment between stages and is
adapted for use with an oil soluble esterification catalyst such as
para-toluene sulfonic acid. It is apparent that the system can be
used with other catalysts.
[0154] A fatty acid-containing feed is provided to apparatus 400
via line 402 and enters contact vessel 404. Contact vessel 404 is
adapted to contact the feed with glycerin containing alkanol.
Contact vessel 404 may be of any suitable design sufficient to
promote contact between the oil and glycerin phases including
static and mechanical mixing devises and may be an extraction
column, in which case a subsequent phase separator may not be
necessary. The contacting may be at any suitable pressure and
temperature as set forth in connection with the description of FIG.
1. A mechanically agitated vessel is depicted.
[0155] The mixed stream from contact vessel 404 is passed via line
406 to phase separator 408. An oil phase containing free fatty acid
and alkanol is passed via line 412 to first esterification reactor
414 and the glycerin phase is withdrawn via line 410. Alkanol is
provided via line 416 to first esterification reactor 414.
Additional catalyst, if required, can be provided via line 418. The
esterification effluent from first esterification reactor 414
contains alkanol, catalyst, ester, water and free fatty acid.
Usually at least about 40, preferably at least about 60, mass
percent of the free fatty acid is converted to ester in first
esterification reactor 414. This esterification effluent is passed
via line 420 to contact vessel 422.
[0156] Contact vessel 422 is adapted to contact the feed with
glycerin supplied by line 424. The glycerin may be from any
suitable source, e.g., a glycerin containing stream from a
transesterification process. Usually the mass ratio of glycerin to
oil is in the range of about 0.05:1 to 1:1, preferably between
about 0.1:1 to 0.5:1. Contact vessel 422 may be of any suitable
design sufficient to promote contact between the oil and glycerin
phases including static and mechanical mixing devises and may be an
extraction column, in which case a subsequent phase separator may
not be necessary. The contacting may be at any suitable pressure
and temperature as set forth in connection with the description of
FIG. 1. For the sake of convenience, the contacting is usually
conducted at approximately the temperature and pressure conditions
of the esterification in first esterification reactor 414. A
mechanically agitated vessel is depicted.
[0157] The mixed stream from contact vessel 422 is passed via line
426 to phase separator 428. An oil phase containing free fatty acid
and alkanol is passed via line 432 to second esterification reactor
434 and the glycerin phase is withdrawn via line 430. Alkanol is
provided via line 436 to second esterification reactor 434.
Typically all alkanol is provided to first esterification reactor
414. Additional catalyst, if required, can be provided via line
438. Usually, since the catalyst is oil soluble, no additional
catalyst need be used. The esterification effluent from second
esterification reactor 434 contains alkanol, catalyst, ester, water
and free fatty acid. Often at least about 70, preferably at least
about 90, mass percent of the free fatty acid in the feed is
converted to ester in apparatus 400. Additional stages of
esterification reactors can be used if desired.
[0158] The esterification effluent from second esterification
reactor 434 is passed via line 440 to contact vessel 442. Contact
vessel 442 is adapted to contact the feed with glycerin supplied by
line 444. Usually the mass ratio of glycerin to oil is in the range
of about 0.05:1 to 1:1, preferably between about 0.1:1 to 0.5:1.
Contact vessel 442 may be of any suitable design sufficient to
promote contact between the oil and glycerin phases including
static and mechanical mixing devises and may be an extraction
column, in which case a subsequent phase separator may not be
necessary. The contacting may be at any suitable pressure and
temperature as set forth in connection with the description of FIG.
1. For the sake of convenience, the contacting is usually conducted
at approximately the temperature and pressure conditions of the
esterification in second esterification reactor 434. A mechanically
agitated vessel is depicted.
[0159] The mixed phase stream from phase separator 442 is passed
via line 446 to phase separator 448. An oil phase is withdrawn from
phase separator 448 via line 456 and a glycerin phase via line 450.
As is described herein, two glycerin contact stages are used, the
first to remove alkanol and water from the esterification effluent
and the second to recover catalyst from the oil phase. The broad
aspects of this invention contemplate that a single stage can be
use.
[0160] Returning to FIG. 2, the oil phase in line 456 from
separator 448 is passed via line 460 to neutralization reactor 462.
If the neutralization and alkanol recovery occurs in a single
stage, i.e., vessel 442 serves both functions, the esterification
product can be withdrawn via line 458. To neutralization reactor
462 is fed a mixed stream of glycerin and base, e.g., sodium
hydroxide or preferably potassium hydroxide, via line 464 in an
amount sufficient to convert the catalyst to salt. Other bases can
be used if desired. Usually the mass ratio of glycerin to oil is in
the range of about 0.05:1 to 1:1, preferably between about 0.1:1 to
0.5:1. The glycerin source, if a transesterification waste stream,
may already contain sufficient base that little, if any, additional
base is required. Neutralization conditions can vary over a wide
range in that the reaction between acid and base proceeds rapidly
and does not require catalyst. Temperature and pressure are often
with the range of about 10.degree. C. to 150.degree. C. and 90 to
1000 kPa absolute. A residence time of from about 0.1 to 100
minutes may be used. Under these conditions, any free fatty acid
contained in the oil phase will also be saponified. Hence, the
amount of base present should thus include an amount sufficient to
effect the saponification as well as the neutralization of the
catalyst. Neutralization reactor 462 may be of any suitable design.
It is preferably a static or mechanically agitated reactor.
[0161] The effluent from neutralization reactor 462 is passed via
line 466 to phase separator 468 with a neutral, esterification
product being withdrawn via line 470. A glycerin phase is withdrawn
via line 472 and passed to acidifier 476 to which mineral acid,
e.g., sulfuric acid, is added via line 478. The acidification
conditions in acidifier 476 are sufficient to provide free fatty
acid and convert the catalyst from a salt to its acid form, e.g.,
toluene sulfonic acid. Typically a two phase mixture will result.
Acidification conditions can vary over a wide range in that the
reaction proceeds rapidly and does not require catalyst.
Temperature and pressure are often with the range of about
10.degree. C. to 150.degree. C. and 90 to 1000 kPa absolute. A
residence time of from about 0.1 to 100 minutes may be used.
[0162] The effluent from acidifier is depicted as being passed via
line 480 to contact vessel 404 and phase separator 408. If desired
a separate phase separator can be used with only the oil phase
passing to contact vessel 404, or more preferably, directly to
first esterification reactor 414. The glycerin phase from phase
separator 408 will contain salts formed from the acidification in
acidifier 476.
[0163] Returning to line 430 containing glycerin phase from phase
separator 428, the glycerin phase contains alkanol and water. Line
430 directs the glycerin phase to contact vessel 404 for recovery
of alkanol therefrom. With respect to line 450 containing glycerin
phase from phase separator 448, that glycerin phase usually
contains less water since less conversion to ester occurs in
reactor 434 as compared to first esterification reactor 414. If the
water content is sufficiently low, the glycerin phase, or a portion
of the glycerin phase, can be passed to first esterification
reactor 414 via line 452. Thus any catalyst contained therein as
well as alkanol is available for use in first esterification
reactor 414. Alternatively, the glycerin phase in line 450 may be
passed via line 454 and 480 to contact vessel 404.
[0164] Reference is now made to FIG. 3. Biodiesel manufacturing
facility 500 uses a suitable raw material feed provided via line
502 containing glycerides and free fatty acids. The raw material
feed in line 502 is combined with a mixed stream of glycerin
provided from the biodiesel process by line 614 as discussed below
and base provided by line 504. Thus the base and the glycerin are
simultaneously contacted with the raw material feed. Alternatively,
the glycerin can be added to the raw material feed before or after
the addition of the base to the raw material feed. The mixture is
passed to contact vessel 506. As the saponification reaction occurs
rapidly, the contact vessel may only be a length of pipe sufficient
to provide distribution of the components. Other types of contact
vessels can be used such as mechanical and sonically agitated
reactors, and reactors with static mixing such as reactors
containing static mixing devices such as a packed bed, baffles,
orifices, venturi nozzles, tortuous flow path, or other impingement
structure
[0165] The amount of glycerin used should be sufficient to not only
provide a separate phase but also enable a substantial portion,
preferably at least about 70, more preferably at least about 90,
mass percent of the soaps formed from the free fatty acids to
reside in the glycerin phase. The base used should be an alkali
metal or alkaline earth metal base, preferably an alkali metal
hydroxide or alkylate, and more preferably sodium or potassium
hydroxide or methylate or ethylate. The amount of base provided
should be at least sufficient to provide the sought degree of
saponification of the free fatty acids. Typically the base is
provided in an amount of between about 90 to 500, preferably
between about 95 and 150, mole percent of that required to react on
a stoichiometric basis with the free fatty acid. Often the glycerin
contains between about 0.1 to 10 or 15 mass percent base.
[0166] The contact of glycerin with the raw material feed can also
remove water from the feed. With some biomass, phospholipids may be
present. Treatment by glycerin sorbent can reduce the phospholipids
content of the glyceride phase.
[0167] In general, it is preferred to use only sufficient glycerin
sorbent to effect the sought removal of fatty acids from the raw
material feed and, if desired, to effect the sought degree of
dehydration of the feed, although more can be used. The mass ratio
of glycerin to raw material feed will vary depending upon the
amount of free fatty acid in the raw material feed and, if desired
to use glycerin for dehydration, the water content of the raw
material feed and the glycerin sorbent. The glycerin sorbent is
conveniently comprised of glycerin phase separated from
transesterification reactor effluent, with or without intervening
treatment. A number of advantages flow from using glycerin phase
separated from the transesterification effluent. First, the
glycerin phase contains some of the base catalyst. Second, soaps in
the glycerin phase can be recovered as free fatty acids in a
subsequent acidification.
[0168] The conditions of the contacting of the streams should be
sufficient to convert the sought amount of free fatty acid to the
corresponding soap. Since a glyceride phase will exist, the
contacting should be under conditions such that good mixing of the
components occurs. The temperature of the contacting may be within
a wide range, say, from about 15.degree. C. to 220.degree. C.,
preferably within the range of about 20.degree. C. to 120.degree.
C. The time of contacting is often in the range of from about 10
seconds to 6 hours, preferably between about 20 seconds and 1
hour.
[0169] The mixed phase system of glycerin and raw material feed is
passed via line 508 to phase separator 510. Phase separator 510 may
be of any convenient design including a decanter, a phase
separation facilitated decanter that contains coalescing sites, and
a centrifuge. A glycerin phase is withdrawn as the heavy phase via
line 582 from phase separator 510 and will be discussed later. The
lighter phase contains triglycerides and is passed via line 512 to
reactor 520 for transesterification.
[0170] The transesterification is base catalyzed with a lower
alkanol, preferably methanol, ethanol or isopropanol. For purposes
of the discussion, methanol will be the alkanol.
[0171] As shown, methanol is supplied via line 516 from methanol
header 514. Line 515 supplies fresh methanol to reactor 520.
Although line 516 is depicted as introducing methanol into line
512, it is also contemplated that methanol can be added directly to
reactor 520 at one or more points. Generally methanol is supplied
in excess of that required to cause the sought degree of
transesterification in reactor 520. More methanol can be supplied
but it may be lost from the facility. Preferably, the amount of
methanol is from about 101 to 500, more preferably, from about 110
to 250, mass percent of that required for the sought degree of
transesterification in reactor 520. In the facility depicted, two
reactors are used. One reactor may be used, but since the reaction
is equilibrium limited, most often at least two reactors are used.
Often, where more than one reactor is used, at least about 60,
preferably between about 70 and 96, percent of the glycerides in
the feed are reacted in the first reactor.
[0172] The base catalyst is shown as being introduced via line 518
to reactor 520. The amount of catalyst used is in excess of that
amount of base that will react with free fatty acids to form soaps
in the transesterification. The base catalyst may be an alkali or
alkaline earth metal hydroxide or alkali or alkaline earth metal
alkoxide, especially an alkoxide corresponding to the lower alkanol
reactant. Preferred alkali metals are sodium and potassium. When
the base is added as a hydroxide, it may react with the lower
alkanol to form an alkoxide with the generation of water.
Alternatively, the catalyst may be a heterogeneous base catalyst.
The exact form of the catalyst is not critical to the understanding
and practice of this invention. For purposes of discussion,
potassium hydroxide is used as the catalyst and the catalysis is
homogeneous.
[0173] Often the transesterification is at a temperature between
about 30.degree. C. and 220.degree. C., preferably between about
30.degree. C. and 80.degree. C. The pressure is typically in the
range of between about 90 to 1000 kPa (absolute) although higher
and lower pressures can be used. The reactor is typically batch,
semi-batch, plug flow or continuous flow tank with some agitation
or mixing, e.g., mechanically stirred, ultrasonic, static mixer,
e.g., a packed bed, baffles, orifices, venturi nozzles, tortuous
flow path, or other impingement structure. The residence time will
depend upon the desired degree of conversion, the ratio of methanol
to glyceride, reaction temperature, the degree of agitation and the
like, and is often in the range of about 0.1 to 20, say, 0.5 to 10,
hours.
[0174] The partially transesterified effluent from reactor 520 is
passed via line 521 to phase separator 522. Phase separator 522 may
be of any suitable design including a decanter, a phase separation
facilitated decanter that contains coalescing sites, and a
centrifuge. A glycerin-containing bottoms phase is provided in the
separator and is removed via line 524. This glycerin phase also
contains any soaps made in reactor 520 and a portion of the
catalyst. The soaps can be recovered from this stream as discussed
later. The lighter phase contains alkyl esters and unreacted
glycerides and is passed via line 526 to second transesterification
reactor 528.
[0175] Reactor 528 may be of any suitable design and may be similar
to or different than reactor 520. As shown, additional methanol is
provided via line 530 from methanol header 514 and additional
catalyst is provided via line 532. Preferably the
transesterification conditions in reactor 528 are sufficient such
that reactors 520 and 528 together react at least about 90, more
preferably at least about 95, and sometimes at least about 97 to
99.9, mass percent of the glycerides in the feed to reactor 520.
The transesterification in reactor 528 is typically operated under
conditions within the parameters set forth for reactor 520 although
the conditions may be the same or different. The residence time
will depend upon the desired degree of conversion. The reactor is
typically agitated, e.g., stirred or ultrasonically agitated, or
the ingredients otherwise subjected to a mixing action such as by a
static mixer, e.g., a packed bed, baffles, orifices, venturi
nozzles, tortuous flow path, or other impingement structure. Plug
flow reactors may be useful. The residence time will depend upon
the desired degree of conversion, the ratio of methanol to
glyceride, reaction temperature, the degree of agitation and the
like, and is often in the range of about 0.1 to 20, say, 0.5 to 10,
hours.
[0176] The effluent from reactor 528 is passed via line 534 to
phase separator 536 which may be of any suitable design and may be
the same as or different from the design of separator 522. A
heavier, glycerin-containing phase is withdrawn via line 538. This
stream contains soaps as well as some catalyst and methanol. As
shown, this glycerin-containing layer is used as a portion of the
glycerin sorbent for the raw material feed treatment. A lighter
phase containing crude biodiesel is withdrawn from separator 536
via line 540. Advantageously, this stream contains less than about
1, preferably less than about 0.8, and most preferably less than
0.05, mass percent soaps based upon the total mass of alkyl esters
and soaps. The lighter phase also contains methanol.
[0177] Alternatively, separator 536 can be eliminated provided that
in reactor 520, the conversion of the glycerides in the feed is at
least about 90, preferably 92 to 96 or 98, percent. Thus the
lighter phase from phase separator 522 contains little glyceride.
In reactor 528 the reaction proceeds quickly to completion by the
addition of additional methanol and catalyst, and can be
conveniently accomplished by a plug flow reactor. Especially with
the higher conversions, the effluent from reactor 528 may be a
single phase.
[0178] The crude biodiesel may be contacted with acid to neutralize
any catalyst therein and then refined to remove methanol, soaps,
water and glycerin. Crude biodiesel is then passed to methanol
separator 542. Methanol separator 542 is of any convenient design
including a stripper, wiped film evaporator, falling film
evaporator, solid sorbent, and the like. Preferably the
fractionation is by fast, vapor fractionation.
[0179] The lower boiling fraction containing the lower alkanol will
contain a portion of any water contained in the crude biodiesel.
Since the transesterification is conducted with little water being
present, and a portion of the water is removed with the glycerin,
the concentration of water in this fraction can be sufficiently low
that it can be recycled to the transesterification reactors. This
lower boiling fraction often contains less than about 0.1, and more
preferably less than about 0.05, mass percent water.
[0180] Vaporized methanol is exhausted via line 544 and may be
exhausted from the facility as a waste stream, e.g., for burning or
other suitable disposal, or can be added to the methanol header
514. The bottoms stream from methanol separator 542 is passed via
line 546 to mixer 548. In another embodiment, line 546 can pass the
bottoms stream to an intermediate mixer for contact with water, and
then the oil phase passed to mixer 548. Into mixer 548 is passed a
strong acid aqueous solution via line 552. Mixer 548 may be an
in-line mixer or a separate vessel. Mixer 548 should provide
sufficient mixing and residence time that essentially all of the
soaps are converted to free fatty acids. Mixer 548 may be of any
convenient design, e.g., a length of pipe, a mechanically or
sonically agitated vessel, or static mixer containing static mixing
devices such as a packed bed, baffles, orifices, venturi nozzles,
tortuous flow path, or other impingement structure. Often the
temperature during the mixing is in the range of about 30.degree.
C. to 220.degree. C., preferably between about 60.degree. C. to
180.degree. C., and for a residence time of between about 0.01 to
4, preferably 0.02 and 1, hours. A convenient mode of practice is
to pass the bottoms stream from methanol separator 542 to mixer 548
without intervening cooling.
[0181] A strong acid aqueous solution introduced via line 552 has a
pH sufficiently low to convert the soaps to free fatty acids. Often
the pH is less than about 4, and more preferably less than about 3,
say, between about 0.1 and 2.5. The acid may be any suitable acid
to achieve the sought pH such as hydrochloric acid, sulfuric acid,
sulfonic acid, phosphoric acid, perchloric acid and nitric acid.
Sulfuric acid is preferred due to cost and availability. The amount
of strong acid aqueous solution provided is typically in a
substantial excess of that required to convert the soaps to free
fatty acid and to neutralize any remaining catalyst. The excess of
acid is often at least about 5, preferably at least about 10, say
between about 10 and 1000 times that required. Consequently the
effluent from mixer 148 is at a pH of up to about 4, preferably
between about 0.1 and 3.
[0182] The effluent from mixer 548 is passed via line 560 to phase
separator 562. Phase separator 562 may be of any suitable design
including a decanter, a phase separation facilitated decanter that
contains coalescing sites, and a centrifuge. A lower aqueous phase
is withdrawn via line 564. A portion of this aqueous phase is
purged and the remaining portion is recycled via line 552 to mixer
548. Make-up acid is provided via line 550 to line 552.
[0183] The lighter phase which contains crude biodiesel and free
fatty acid is withdrawn via line 566 and is passed to water wash
vessel 568. Usually the free fatty acid is present in an amount
less than about 500, most frequently less than about 300, parts per
million by mass, and thus no need exists to remove free fatty acid
to provide a biodiesel product meeting current commercial
specifications. Fresh water enters vessel 568 via line 570 and
serves to remove residual methanol and salts from the crude
biodiesel. Water wash vessel 568 may be of any suitable design
including a mechanically or sonically agitated tank, a vessel
containing static mixing devices such as a packed bed, baffles,
orifices, venturi nozzles, tortuous flow path, or other impingement
structure, or a wash column. Normally washing is operated at a
temperature between about 20.degree. C. and 120.degree. C.,
preferably between about 35.degree. C. and 90.degree. C. The amount
of water provided is sufficient to effect a sought removal of
glycerin and residual methanol from the crude biodiesel. Typically
between about 2 and 50, preferably between about 5 and 20, mass
parts of wash water are used per 100 mass parts of crude
biodiesel.
[0184] In a preferred embodiment, the spent water from wash vessel
568 is passed via line 572 to mixer 548 or combined with the
aqueous solution in line 552. Most preferably, the water provided
via line 570 is in an amount to replace the volume of purge from
line 564 to maintain steady state conditions. Often the purge from
line 564 is less than 20, preferably between about 5 and 15, volume
percent of the heavier, aqueous phase withdrawn from separator
562.
[0185] A washed biodiesel stream is withdrawn from washing column
568 via line 574 and is passed to drier 576 to remove water which
exhausts via line 578. Preferably substantially all the methanol
has been removed from the crude biodiesel prior to drying to permit
the water vapor to be exhausted without further treatment to
eliminate volatile organic components. Drier 576 may be of any
suitable design such as stripper, wiped film evaporator, falling
film evaporator, and solid sorbent. Generally the temperature of
drying is between about 60.degree. C. and 220.degree. C., say,
about 70.degree. C. and 180.degree. C. The pressure is generally in
the range of about 5 to 200 kPa absolute. The dried biodiesel is
withdrawn as product via line 580. The biodiesel product contains
free fatty acid and preferably has a free fatty acid content of
less than about 0.8, and more preferably less than about 0.5, mass
percent.
[0186] Returning to separator 510, the heavier glycerin-containing
phase is withdrawn via line 582 and passed to mixer 584. Mixer 584
may be an in-line mixer or a separate vessel including a
mechanically or sonically agitated tank, a vessel containing static
mixing devices such as a packed bed, baffles, orifices, venturi
nozzles, tortuous flow path, or other impingement structure, or a
wash column. Mixer 584 should provide sufficient mixing and
residence time that essentially all of the soaps are converted to
free fatty acids. Often the temperature during the mixing is in the
range of about 30.degree. C. to 220.degree. C., preferably between
about 40.degree. C. and 180.degree. C., and for a residence time of
between about 0.01 to 4, preferably 0.02 and 1, hours.
[0187] An aqueous solution of strong acid is passed via line 586 to
mixer 584 where the soaps contained in the glycerin are converted
to free fatty acids which form a separate, lighter phase. Often the
pH of the aqueous solution is less than about 4, and more
preferably less than about 3, say, between about 0.1 and 2.5. The
acid may be any suitable acid to achieve the sought pH such as
hydrochloric acid, sulfuric acid, sulfonic acid, phosphoric acid,
perchloric acid, and nitric acid. Sulfuric acid is preferred due to
cost and availability. The amount of strong acid aqueous solution
provided is typically in a substantial excess of that required to
convert the soaps to free fatty acid and to neutralize any
remaining catalyst. The excess of acid is often at least about 5,
preferably at least about 10, say between about 10 and 1000 times
that required.
[0188] The mixed phase glycerin and free fatty acid is passed from
mixer 584 via line 588 to phase separator 590. Phase separator 590
may be of any suitable design to provide a heavy
glycerin-containing phase which is removed via line 594. The
glycerin phase may be treated in any suitable manner. For instance,
the glycerin layer may be neutralized and subjected to distillation
to remove water and methanol, if present. It can also be used as a
fuel. This glycerin-containing stream may be used as the source of
glycerin for treating the feed and thus at least a portion would be
provided to line 504.
[0189] The free fatty acid from separator 590 passed via line 602
to acid catalyzed esterification reactor system 600 for conversion
to the methyl ester. Methanol is provided to reactor system 600 via
line 604 and any additional catalyst required via line 606. Any
suitable acid catalyzed process for the esterification of free
fatty acids to make biodiesel can be used including homogeneous and
heterogeneous catalysis processes.
[0190] The effluent from esterification reactor system 600 is
passed via line 608 to counter current extraction column 610 where
it is contacted with glycerin from line 524. The glycerin serves to
remove methanol and water from the esterification effluent. The oil
phase is withdrawn via line 612. Depending upon the volume of the
stream and its methanol content, the stream may be introduced at
various points in the refining section of the transesterification
unit. As shown, the effluent in line 612 is directed to the water
wash unit operation. The water washing will serve to remove
methanol still remaining in the effluent.
[0191] If the volume of the stream is larger and thus the absolute
amount of methanol is greater, it can be directed to methanol
separator 542.
[0192] The glycerin phase from column 610 is passed via line 614 to
contact vessel 506.
[0193] The apparatus 700 depicted in FIG. 4 is adapted to acid
esterify glyceride-containing feed with an oil-soluble, acid
catalyst with recovery and recycle of the catalyst. As shown,
glyceride-containing feed which also contains free fatty acid is
provided by line 702 to esterification reactor 704. Alkanol and
oil-soluble catalyst are also supplied as will be discussed
below.
[0194] Since the catalyst is organic miscible, reactor 704 may be
any of a widely divergent type of reactor. Suitable designs include
a pipe reactor, mechanically or sonically agitated tank, a vessel
containing static mixing devices such as a packed bed, baffles,
orifices, venturi nozzles, tortuous flow path, or other impingement
structure. Reactor 702 may be one or more vessels, each defining a
reaction stage. Preferably the reactor provides static or
mechanical missing of the liquid.
[0195] Reactor 704 is maintained under esterification conditions to
provide an esterification effluent containing alkyl ester,
catalyst, water from the esterification and a reduced concentration
of alkanol and free fatty acid. This effluent is passed via line
706 to separator 708. The alkanol concentration in the reaction
menstruum in reactor 704 may contain more alkanol than is miscible
in the oil phase and/or additional alkanol can be added to the
esterification effluent via line 736 to provide an alkanol phase
and an oil phase in separator 708. Due to the more polar nature of
the alkanol, acid catalyst preferentially partitions to the alkanol
phase in separator 708. An oil phase having a reduced concentration
of acid catalyst is removed via line 710.
[0196] The alkanol phase from separator 708 is withdrawn via line
712. All or a portion of the alkanol phase in line 712 can be
withdrawn via line 714 and passed to lights column 716. The alkanol
phase withdrawn via line 714 will contain some water. Although the
esterification of free fatty acid is an equilibrium-limited
reaction, water is tolerable to some extent and hence a portion of
the alkanol phase can be passed to reactor 704. Preferably, no
separate water phase is formed in reactor 704 as the catalyst tends
to preferentially partition to a water phase.
[0197] In lights column, water is removed by vaporization. As
methanol, if it is the alkanol, has a lower boiling temperature
than water, and other alkanols such as ethanol and isopropanol form
azeotropes with water, at least a portion of the alkanol will be
vaporized with water. Where a substantial portion of the alkanol
will be removed by vaporization, a higher boiling liquid in which
the catalyst is soluble can be supplied to lights column 716. For
instance, glycerin or biodiesel can be provided by line 718 to
lights column 716 in an amount sufficient to maintain the
oil-soluble, acid catalyst in a liquid medium. Especially where
glycerin is used, the amount of glycerin is preferably less than
that which would enable a glycerin phase to form in reactor 704.
Alternatively or in addition, glyceride-containing feed may be used
and provided via line 720.
[0198] Lights column 716 may be of any suitable design including a
flash distillation column, a trayed or packed distillation column,
or the like.
[0199] The higher boiling fraction containing the acid catalyst is
passed via line 722 to line 712 which is in fluid communication
with reactor 704. A lower boiling fraction from lights column 716
passes via line 724 to condenser 726 with a water phase being
withdrawn via line 728 and an alkanol phase, after condensation,
being withdrawn via line 730. If the alkanol and water form an
azeotrope, unit operation 726 may be a suitable unit operation for
selectively removing water, e.g., a selective extraction. A portion
of the alkanol may be passed via line 731 as reflux to lights
column 716. All, or the balance, of the alkanol can be passed via
line 730 to line 712 for recycle to reactor 704. Make-up alkanol
can be provided via line 732. As shown, make-up catalyst is
provided to the make-up alkanol stream via line 734. All or a
portion of the alkanol provided by line 736, if needed, can be
supplied from line 730 or another source of alkanol.
[0200] Where the alkanol is methanol and the acid catalyst is
toluene sulfonic acid, the apparatus of FIG. 4 is capable of
recovering over 98 mass percent of the acid per pass in the
methanol phase in separator 708. Moreover, as the methanol phase
can be relatively small in comparison to the oil phase yet still
recover a high percentage of the catalyst, energy requirements for
water removal by lights column 716 can be economically viable.
[0201] FIG. 5 depicts an additional embodiment of the invention. To
an esterification unit designated generally by numeral 800 is fed a
glycerides-containing feedstock that also contains free fatty acid
via line 802. The feedstock is fed to first esterification reactor
804. First esterification reactor 804 may be a vessel or a length
of pipe. But preferably other types of vessels are used such as
mechanical and sonically agitated reactors, and reactors with
static mixing such as reactors containing contact structures such
as trays, packing, baffles, orifices, venturi nozzles, tortuous
flow path, and other impingement structures. Suitable reactors
include those providing high intensity mixing, including high
shear. First esterification reactor 804 is maintained under
esterification conditions including the presence of catalyst. For
purposes of this discussion, the catalyst will be sulfuric acid
which is provided by line 806. As the alkanol, methanol is used and
is supplied to first esterification reactor 804 via line 852.
[0202] Esterification reactor 804 is operated to provide a partial
conversion of free fatty acid to methyl ester, e.g., between about
30 to 90, say, 40 to 80, percent of the free fatty acid is
converted. A partially converted effluent is withdrawn from first
esterification reactor 804 and is passed via line 808 to second
esterification reactor 810. As shown, no additional methanol or
catalyst is added nor is any removal of water effected. Second
esterification reactor 810 may be the same or different from first
esterification reactor. Second esterification reactor 810 is
maintained under esterification conditions and provides an
esterification effluent having an increased conversion of the free
fatty acid to methyl ester. Often at least about 75 mole percent to
essentially all, preferably between about 75 and 95 or 98, mass
percent of the free fatty acid is converted to ester.
[0203] The esterification effluent from second esterification
reactor 810 is passed via line 812 to decanter 814 wherein a
sulfuric acid and methanol-containing phase is formed and removed
via line 830 and an oil phase containing methanol is formed and
removed via line 816.
[0204] Oil phase in line 816 is contacted in vessel 818 with basic
glycerin supplied via line 820 wherein free fatty acid contained in
the oil phase is converted to soap and a treated product is
provided. The treated product is passed via line 822 to decanter
824. In decanter 824 a glycerin-containing phase is formed which
also contains salt of free fatty acid and methanol. This
glycerin-containing phase is passed via line 828 to contact vessel
832 where it is contacted with the sulfuric acid and
methanol-containing phase supplied by line 830. An oil phase is
also formed in decanter 824 and is withdrawn via line 826 for
further processing, e.g., to make biodiesel.
[0205] Returning to contact vessel 832, a mixture of glycerin,
sulfuric acid and methanol is generated therein. Acid and water are
partitioned to the glycerin and a methanol-containing phase is
formed. The methanol-containing phase may also contain some free
fatty acid. The mixture is passed via line 833 to decanter 834. A
methanol-containing phase is obtained in decanter 834 and is passed
via line 836 to first esterification reactor 804. A
glycerin-containing phase which also contains methanol and water is
withdrawn from decanter 834 via line 838 and passed to stripper
842. As shown, base such as sodium or potassium hydroxide or
alkoxide is added via line 840 to the glycerin-containing phase
being passed through line 838. If desired a contact vessel may be
used rather than a length of pipe. The base is provided in an
amount sufficient to neutralize the sulfuric acid. Alternatively
base can be added to contact vessel 832 or to line 833 with the
elimination of decanter 834 and line 836.
[0206] Stripper 842 is operated under conditions of temperature and
pressure to provide an overhead containing methanol and water and a
bottoms fraction containing glycerin and sulfate salt and a reduced
concentration of methanol, preferably less than about 10, more
preferably less than about 5, mass percent methanol. The bottoms
fraction is withdrawn from stripper 842 via line 844. The overhead
from stripper 842 is passed via line 846 to methanol column 848
which serves to provide an overhead containing virtually no water,
e.g., less than about 0.1, more preferably less than about 0.01,
volume percent water. This methanol stream is passed to first
esterification reactor 804 via line 852. Make-up methanol is
supplied via line 854 to maintain the sought methanol to free fatty
acid mole ratios in the esterification reactors. A bottoms fraction
provided by methanol column 848 preferably contains less than 0.1
mass percent methanol and is withdrawn via line 850.
* * * * *