U.S. patent application number 14/607321 was filed with the patent office on 2015-05-21 for steamcracking bio-naphtha produced from complex mixtures of natural occurring fats and oils.
The applicant listed for this patent is TOTAL RESEARCH & TECHNOLOGY FELUY. Invention is credited to Nicolas Van Gyseghem, Walter Vermeiren.
Application Number | 20150141715 14/607321 |
Document ID | / |
Family ID | 41510566 |
Filed Date | 2015-05-21 |
United States Patent
Application |
20150141715 |
Kind Code |
A1 |
Vermeiren; Walter ; et
al. |
May 21, 2015 |
STEAMCRACKING BIO-NAPHTHA PRODUCED FROM COMPLEX MIXTURES OF NATURAL
OCCURRING FATS AND OILS
Abstract
A process for making a bio-naphtha and optionally bio-propane
from a complex mixture of natural occurring fats & oils,
wherein said complex mixture is subjected to a refining treatment
for removing the major part of non-triglyceride and non-fatty acid
components, thereby obtaining refined fats & oils; said refined
fats & oils are transformed into linear or substantially linear
paraffin's as the bio-naphtha by an hydrodeoxygenation or from said
refined fats & oils are obtained fatty acids that are
transformed into linear or substantially linear paraffin's as the
bio-naphtha by hydrodeoxygenation or decarboxylation of the free
fatty acids or from said refined fats & oils are obtained fatty
acids soaps that are transformed into linear or substantially
linear paraffin's as the bio-naphtha by decarboxylation of the
soaps.
Inventors: |
Vermeiren; Walter;
(Houthalen, BE) ; Van Gyseghem; Nicolas;
(Brussels, BE) |
|
Applicant: |
Name |
City |
State |
Country |
Type |
TOTAL RESEARCH & TECHNOLOGY FELUY |
Seneffe (Feluy) |
|
BE |
|
|
Family ID: |
41510566 |
Appl. No.: |
14/607321 |
Filed: |
January 28, 2015 |
Related U.S. Patent Documents
|
|
|
|
|
|
Application
Number |
Filing Date |
Patent Number |
|
|
14095530 |
Dec 3, 2013 |
8975459 |
|
|
14607321 |
|
|
|
|
13382308 |
Mar 8, 2012 |
8648224 |
|
|
PCT/EP2010/060028 |
Jul 13, 2010 |
|
|
|
14095530 |
|
|
|
|
Current U.S.
Class: |
585/240 ;
585/310; 585/318; 585/319; 585/321; 585/324; 585/327 |
Current CPC
Class: |
C10G 3/50 20130101; C07C
4/04 20130101; C10G 2400/20 20130101; C11B 13/02 20130101; Y02P
30/20 20151101; C11C 3/12 20130101; C07C 1/22 20130101; C10G
2300/1007 20130101; C10G 2400/22 20130101; C07C 1/2078 20130101;
C10G 2400/02 20130101; C11B 3/12 20130101; C10G 9/00 20130101; C10G
2400/30 20130101; C10G 3/42 20130101; C10G 2300/807 20130101; C10G
2300/1014 20130101; C10G 2300/805 20130101; C10G 3/46 20130101;
C11B 3/14 20130101; C10G 2300/1018 20130101; C10G 3/45 20130101;
C10G 3/49 20130101 |
Class at
Publication: |
585/240 ;
585/310; 585/319; 585/324; 585/321; 585/318; 585/327 |
International
Class: |
C10G 3/00 20060101
C10G003/00; C07C 1/207 20060101 C07C001/207; C11B 13/02 20060101
C11B013/02; C11B 3/12 20060101 C11B003/12; C11B 3/14 20060101
C11B003/14; C07C 1/22 20060101 C07C001/22; C10G 9/00 20060101
C10G009/00 |
Foreign Application Data
Date |
Code |
Application Number |
Jul 27, 2009 |
EP |
09166485.4 |
Claims
1-16. (canceled)
17. A process comprising: (a) making a bio-naphtha and optionally
bio-propane from a complex mixture of natural occurring raw fats
and oils, wherein the raw fats and oils are subjected to a refining
pretreatment including a degumming and a bleaching for removing
non-triglyceride and non-fatty acid components, thereby obtaining
refined fats and oils; (b1) wherein the refined fats and oils are
physically refined by vacuum distillation or steam distillation to
recover mixed fatty acids as overhead product and triglycerides as
bottom product; and wherein either: (c1) the refined fats and oils
still containing free fatty acids or the physically refined
triglycerides are sent to a hydrodeoxygenation section and
converted into bio-naphtha and bio-propane; or (c2) the mixed fatty
acids are sent to a hydrodeoxygenation section and converted into
bio-naphtha or are sent to a decarboxylation section and converted
into bio-naphtha; or (b2) the refined fats and oils are
saponificated to recover soaps and glycerol or soaps are obtained
during a chemical refining step of the raw fats and oils by a
neutralisation step; or soaps are obtained via neutralisation of
fatty acids obtained by steam splitting/hydrolysis of the refined
fats and oils producing fatty acids and glycerol and the soaps are
sent to a decarboxylation section and converted into bio-naphtha;
and wherein the complex mixture of natural occurring raw fats and
oils comprises vegetable oils, animal fats, or mixtures
thereof.
18. The process of claim 17, wherein the complex mixture of natural
occurring raw fats and oils comprises inedible oils.
19. The process of claim 17, wherein the complex mixture of natural
occurring raw fats and oils comprises highly saturated oils.
20. The process of claim 17, wherein the complex mixture of natural
occurring raw fats and oils comprises waste food oils.
21. The process of claim 17, wherein the complex mixture of natural
occurring raw fats and oils comprises by-products of a refining of
vegetable oils.
22. The process of claim 17, wherein the refining pretreatment
comprises: a physical refining comprising degumming, bleaching, and
steam refining-deodorisation producing free fatty acids; or a
chemical refining comprising degumming, neutralisation producing
soaps that are split into free fatty acids, bleaching, and
deodorisation.
23. The process of claim 17, wherein the fatty acids are obtained
by acidulation of soaps.
24. The process of claim 17, wherein the refined fats and oils are
hydrolysed to produce mixed fatty acids and glycerol.
25. The process of claim 17, wherein a quality of the mixed fatty
acids is improved by hydrogenation of double bonds in the
acyl-moiety or before hydrolysis the fats and oils are hydrogenated
to remove remaining double bonds and subsequently sent to a
hydrolysis step.
26. The process of claim 17, wherein a quality of the soaps is
improved by hydrogenation of double bonds in the acyl-moiety or
before saponification or hydrolysis, the fats and oils are
hydrogenated to remove remaining double bonds and subsequently sent
to a saponification or hydrolysis step.
27. The process of claim 17, wherein in step (c1) the
hydrodeoxygenation is conducted in the presence of hydrogen and at
least one catalyst selected among Ni, Mo, Co, NiW, NiMo, CoMo,
NiCoW, NiCoMo, NiMoW, CoMoW oxides or sulphides, optionally
supported on high surface area carbon, alumina, silica, titania or
zirconia.
28. The process of claim 27, wherein the hydrodeoxygenation is
carried out at a temperature from 200 to 500.degree. C., under a
pressure from 1 MPa to 10 MPa and with a hydrogen to feed ratio
from 100 to 2000 Nl/l.
29. The process of claim 17, wherein in step (c2) the
hydrodeoxygenation or decarboxylation is conducted in the presence
of hydrogen and of at least one catalyst selected from: Ni, Mo, Co,
NiW, NiMo, CoMo, NiCoW, NiCoMo, NiMoW, and CoMoW oxides or
sulphides, optionally supported on high surface area carbon,
alumina, silica, titania, or zirconia; or group 10 metals, group 11
metals, or alloy mixtures thereof optionally supported on high
surface area carbon, magnesia, zinc-oxide, spinels, perovskites,
calciumsilicates, alumina, silica, silica-alumina, or mixtures
thereof
30. The process of claim 29, wherein in step (c2) the
hydrodeoxygenation is carried out at a temperature from 200 to
500.degree. C., under a pressure from 1 MPa 15 to 10 MPa and with a
hydrogen to feedstock ratio from 100 to 2000 Nl/l.
31. The process of claim 17, wherein the decarboxylation of the
fatty acids is carried out on basic oxides either as bulk material
or dispersed on basic or neutral carriers, dispersed on basic
zeolites.
32. The process of claim 31, wherein the decarboxylation is carried
out at a temperature from 100 to 550.degree. C., under a pressure
from 0.1 MPa to 10 MPa and with a hydrogen to feedstock ratio from
0 to 2000 Nl/l.
33. The process of claim 17, wherein the decarboxylation of the
soaps is carried out at from 100 to 550.degree. C. under pressure
from 0.1 MPa to 10 MPa and in presence of water.
34. The process of claim 17, wherein the decarboxylation of the
soaps is carried out with a water to feedstock ratio of at least 1
mole water per mole of soap.
35. The process of claim 17, further comprising using the
bio-naphtha as a direct feedstock of a steamcracker, optionally
combined with the bio-propane or blended with LPG, naphtha, or
gasoil, in order to obtain cracked products including bio-ethylene,
bio-propylene, bio-butadiene, bio-isoprene, bio-cyclopentadiene and
bio-piperylenes, bio-benzene, bio-toluene, bio-xylene, or
bio-gasoline, wherein in the steamcracking process the hydrocarbon
feedstock is mixed with steam in a ratio of 0.3 to 0.45 kg steam
per kg hydrocarbon feedstock.
36. The process of claim 35, wherein the hydrocarbon feedstock is
mixed with steam in a ratio of 0.3 to 0.4 kg steam per kg
hydrocarbon feedstock.
37. The process of claim 35, wherein the hydrocarbon feedstock is
heated up to a temperature of 750-950.degree. C. and is steamed
cracked for a residence time of 0.05 to 0.5 seconds.
38. The process of claim 35, wherein the hydrocarbon feedstock is
heated up to a temperature of 750-950.degree. C. is steamed cracked
for a residence time of 0.05 to 0.15 seconds.
39. The process of claim 35, wherein the obtained cracked products
exhibit an ethylene to methane weight ratio of at least 3.
Description
FIELD OF THE INVENTION
[0001] The present invention relates to the production of
bio-naphtha in an integrated bio-refinery from complex mixtures of
natural occurring fats & oils. The limited supply and
increasing cost of crude oil and the need to reduce emission of
fossil based carbon dioxides has prompted the search for
alternative processes for producing hydrocarbon products such as
bio-naphtha and bio-diesel. The bio-naphtha can be used as
feedstock of conventional steamcracking. Made up of organic matter
from living organisms, biomass is the world's leading renewable
energy source.
BACKGROUND OF THE INVENTION
[0002] The industrial chemistry of fats & oils is a mature
technology, with decades of experience and continuous improvements
over current practices. Natural fats & oils consist mainly of
triglycerides and to some extent of free fatty acids (FFA). Many
different types of triglycerides are produced in nature, either
from vegetable as from animal origin. Fatty acids in fats &
oils are found esterified to glycerol (triacylglycerol). The
acyl-group is a long-chain (C.sub.12-C.sub.22) hydrocarbon with a
carboxyl-group at the end that is generally esterified with
glycerol. Fats & oils are characterized by the chemical
composition and structure of its fatty acid moiety. The fatty acid
moiety can be saturated or contain one or more double bonds. Bulk
properties of fats & oils are often specified as
"saponification number", "Iodine Value", "unsaponification number".
The "saponification number", which is expressed as grams of fat
saponified by one mole of potassium hydroxide, is an indication of
the average molecular weight and hence chain length. The "Iodine
value", which is expressed as the weight percent of iodine consumed
by the fat in a reaction with iodine monochloride, is an index of
unsaturation.
[0003] Some typical sources of fats & oils and respective
composition in fatty acids are given by way of example in Table
1.
TABLE-US-00001 TABLE 1 Symbol Cotton- Coconut Corn Palm Peanut Palm
Linseed Rice Rape- Olive Saturated Caproic 6:0 0.4 0.2 Caprylic 8:0
7.3 3.3 Capric 10:0 6.6 3.5 Lauric 12:0 47.8 47.8 0.2 Myristic 14:0
0.9 18.1 16.3 0.1 1.1 0.4 0.02 Palmitic 16:0 24.7 8.9 10.9 8.5 11.6
44.1 6.0 19.8 3.9 10.5 Margaric 17:0 0.05 Stearic 18:0 2.3 2.7 1.8
2.4 3.1 4.4 2.5 1.9 1.9 2.6 Arachidic 20:0 0.1 0.1 1.5 0.2 0.5 0.9
0.6 0.4 Behenic 22:0 3.0 0.3 0.2 0.2 Lignoceric 24:0 1.0 0.2 0.1
TOTAL 28.0 91.9 22.7 82.0 20.3 50 9.0 23.3 6.8 13.87 Unsaturated
Myristoleic 14:1 w-5 Palmitoleic 16:1 w-7 0.7 0.5 0.1 0.2 0.6
Heptadecenoic 17:1 w-15 0.09 Oleic 18:1 w-9 17.6 6.4 24.2 15.4 38.0
37.5 19.0 42.3 64.1 76.9 Linoleic 18:2 w-6 53.3 1.6 58.0 2.4 41.0
1.0 24.1 31.9 18.7 7.5 Linolenic 18:3 w-3 0.3 0.7 47.4 1.2 9.2 0.6
Gadolenic 20:1 w-9 1.0 0.5 0.5 1.0 0.3 TOTAL 72.0 8.1 77.3 18.0
79.7 50 91 76.7 93.2 86.13 Polyunsaturated Ricinoleic 18 Rosin
acids -- % FFA 0.5-0.6 1.0-3.5 1.7 0.1 0.8 2-14 2 5-15 0.5-3.8
0.5-3.3 Soy- Sun- Linola Lard Butterfat Tallow Tall Castor Jatropha
Saturated Caproic 2 Caprylic 2 Capric 3 Lauric 0.5 0.5 3.5 Myristic
0.1 0.2 1.5 11 3 Palmitic 11.0 6.8 5.6 26 26 26 2 1.0 14.6 Margaric
0.5 0.5 Stearic 4.0 4.7 4.0 13.5 11 22.5 1 1.0 7.4 Arachidic 0.3
0.4 2 0.5 Behenic 0.1 Lignoceric TOTAL 15.5 12.6 9.6 42.0 60.5 52.0
3.5 2.0 22.0 Unsaturated Myristoleic 0.5 Palmitoleic 0.1 0.1 4 2
2.5 0.8 Heptadecenoic 0.5 3 0.5 Oleic 23.4 18.6 15.9 43 26 43 16
3.0 47.5 Linoleic 53.2 68.2 71.8 9 2.5 1.5 20 4.2 28.7 Linolenic
7.8 0.5 2.0 0.5 4 0.3 1.0 Gadolenic 1 0.5 TOTAL 84.5 87.4 90.4 58.0
37.5 48.0 54.5 7.5 78.0 Polyunsaturated 2 4 Ricinoleic 89.5 Rosin
acids 40 % FFA 0.3-1.6 0.1-1.5 0.3 0.5 5-20
[0004] There are other potential feedstock available at this time,
namely trap and sewage grease and other very high free fatty acid
greases who's FFA can exceed 50%.
[0005] The main sources of fats & oils are palm and palm
kernels, soybeans, rapeseed, sunflower, coconut, corn, animal fats,
milk fats.
[0006] Potentially new sources of triglycerides will become
available in the near future, namely those extracted from Jatropha
and those produced by microalgues. These microalgues can accumulate
more then 30 wt % of lipids on dry basis and they can either be
cultivated in open basin, using atmospheric CO.sub.2 or in closed
photobioreactors. In the latter case, the required CO.sub.2 can
originate from the use of fossil hydrocarbons that are captured and
injected into the photobioreactor. Main sources of fossil CO.sub.2
are power stations, boilers used in refineries and steamcrackers
furnaces used to bring hydrocarbon streams at high temperature or
to supply heat of reactions in hydrocarbon transformations in
refineries and steamcrackers. In particular steameracking furnaces
produce a lot of CO.sub.2. In order to enhance the CO.sub.2
concentration in flue gases of these furnaces, techniques like
oxycombustion, chemical looping or absorption of CO.sub.2 can be
employed. In oxycombustion, oxygen is extracted from air and this
pure oxygen is used to burn hydrocarbon fuels as to obtain a stream
only containing water and CO.sub.2, allowing concentrating easily
the CO.sub.2 for storage or re-utilisation. In chemical looping, a
solid material acts as oxygen-transfer agent from a re-oxidation
zone where the reduced solid is re-oxidised with air into oxidised
solid to a combustion zone, where the hydrocarbon fuel is burned
with the oxidised solid and hence the effluent resulting from the
combustion zone only contains water and CO.sub.2. Absorption of
CO.sub.2 can be done with the help of a lean solvent that has a
high preferential to absorb CO.sub.2 under pressure and typically
at low temperature and will release the CO.sub.2 when depressurised
and/or heated. Rectisol.RTM. and Selexol.RTM. are commercial
available technologies to remove and concentrate CO.sub.2. Other
sources of CO.sub.2 are the byproduct from carbohydrates
fermentation into ethanol or other alcohols and the removal of
excess CO.sub.2 from synthesis gas made from biomass or coal
gasification.
[0007] US 2007/0175795 reports the contacting of a hydrocarbon and
a triglyceride to form a mixture and contacting the mixture with a
hydrotreating catalyst in a fixed bed reactor under conditions
sufficient to produce a reaction product containing diesel boiling
range hydrocarbons. The example demonstrates that the
hydrotreatment of such mixture increases the cloud point and pour
point of the resulting hydrocarbon mixture.
[0008] US 2004/0230085 reports a process for producing a
hydrocarbon component of biological origin, characterized in that
the process comprises at least two steps, the first one of which is
a hydrodeoxygenation step and the second one is an isomerisation
step. The resulting products have low solidification points and
high cetane number and can be used as diesel or as solvent.
[0009] US 2007/0135669 reports the manufacture of branched
saturated hydrocarbons, characterized in that a feedstock
comprising unsaturated fatty acids or fatty acids esters with C1-C5
alcohols, or mixture thereof, is subjected to a skeletal
isomerisation step followed by a deoxygenation step. The results
demonstrate that very good cloud points can be obtained.
[0010] US 2007/0039240 reports on a process for cracking tallow
into diesel fuel comprising: thermally cracking the tallow in a
cracking vessel at a temperature of 260-371.degree. C., at ambient
pressure and in the absence of a catalyst to yield in part cracked
hydrocarbons.
[0011] U.S. Pat. No. 4,554,397 reports on a process for
manufacturing olefins, comprising contacting a carboxylic acid or a
carboxylic ester with a catalyst at a temperature of
200-400.degree. C., wherein the catalyst simultaneously contains
nickel and at least one metal from the group consisting of tin,
germanium and lead.
[0012] It has been discovered a process to make bio-naphtha in an
integrated biorefinery from all kinds of natural triglycerides or
fatty acids. In said process crude fats & oils are refined,
either physically or chemically, to remove substantially all
non-triglyceride components and non-free fatty acids.
[0013] The use of a biofeed is a possible solution in the search of
alternative raw material for the naphthacracker. Nevertheless,
using this type of feed can lead to corrosion problems and
excessive fouling because of oxygenates forming from the oxygen
atoms in the biofeed. Also existing steamcrackers are not designed
to remove high amounts of carbonoxides that would result from the
steamcracking of these biofeedstock. According to the present
invention, such a problem can be solved by
hydrodeoxygenating/decarboxylating (or decarbonylating) this
biofeed before its injection into the steam cracker. Thanks to this
hydrodeoxygenation/decarboxylation (or decarbonylation), the
negative effect due to the production of CO and CO.sub.2 and traces
of low molecular weight oxygenates (aldehydes and acids) in the
steam cracker is reduced.
[0014] Another advantage is of course the production of
bio-monomers in the steam cracker.
BRIEF DESCRIPTION OF THE INVENTION
[0015] The subject-matter of the present invention is, in a first
embodiment, a process for making a bio-naphtha from a complex
mixture of natural occurring fats & oils, wherein said complex
mixture is subjected to a refining treatment for removing the major
part of the non-triglyceride and non-fatty acid components, thereby
obtaining refined oils; [0016] said refined fats & oils are
transformed into linear or substantially linear paraffin's as the
bio-naphtha by an hydrodeoxygenation [0017] or from said refined
fats & oils are obtained fatty acids that are transformed into
linear or substantially linear paraffin's as the bio-naphtha by
hydrodeoxygenation or decarboxylation of the fatty acids [0018] or
from said refined fats & oils are obtained fatty acids soaps
that are transformed into linear or substantially linear paraffin's
as the bio-naphtha by decarboxylation of the soaps.
[0019] In an embodiment 2 the invention is according to embodiment
I, wherein said complex mixture of natural occurring fats &
oils is selected among vegetable oils and animal fats,
preferentially inedible oils, highly saturated oils, waste food
oils, by-products of the refining of vegetable oils, and mixtures
thereof.
[0020] In an embodiment 3 the invention is according to embodiment
1 or 2, wherein said fatty acids are obtained by physical refining,
including a steam distillation or vacuum distillation of fats &
oils [0021] or said fatty acids are obtained by hydrolysis of
triglycerides of the fats & oils [0022] or said fatty acids are
obtained by acidulation of soaps.
[0023] In an embodiment 4 the invention is according to embodiment
1 or 2, wherein said fatty acid soaps are obtained by
saponification of fats & oils or by the chemical refining,
including neutralisation of free fatty acids, present in the fats
& oils [0024] or neutralisation of fatty acids, obtained from
hydrolysis of the fats & oils
[0025] In an embodiment 5 the invention is according to anyone of
embodiments 1 to 3, wherein said refined fats & oils are
transformed into linear or substantially linear paraffins as
bio-naphtha together with bio-propane by hydrodeoxygenation in the
presence of hydrogen and of at least one catalyst that can be
selected among Ni, Mo, Co or mixtures like NiW, NiMo, CoMo, NiCoW,
NiCoMo, NiMoW and CoMoW oxides or sulphides as catalytic phase,
preferably supported on high surface area carbon, alumina, silica,
titania or zirconia. .
[0026] In an embodiment 6 the invention is according to embodiment
5, wherein the hydrodeoxygenation is carried out at a temperature
from 200 to 500.degree. C., under a pressure from 1 MPa to 10 MPa
(10 to 100 bars) and with a hydrogen to feed ratio from 100 to 2000
Nl/l.
[0027] In an embodiment 7 the invention is according to anyone of
Embodiments 1 to 4, wherein said refined oils are transformed into
linear or substantially linear paraffins as bio-naphtha by
hydrolysis of the fats & oils into glycerol and fatty acids,
removal of the glycerol or by physical refining, including a steam
distillation or vacuum distillation of fats & oils or obtained
by acidulation of soaps and hydrodeoxygenation or decarboxylation
of the fatty acids, said hydrodeoxygenation or decarboxylation
being conducted in the presence of hydrogen and of at least one
catalyst that can be selected among Ni, Mo, Co or mixtures like
NiW, NiMo, CoMo, NiCoW, NiCoMo, NiMoW and CoMoW oxides or sulphides
as catalytic phase, preferably supported on high surface area
carbon, alumina, silica, titania or zirconia or group 10 (Ni, Pt
and Pd) and group 11 (Cu and Ag) metals or alloy mixtures supported
on high surface area carbon, magnesia, zinc-oxide, spinels
(Mg.sub.2Al.sub.2O.sub.4, ZnAl.sub.2O.sub.4), perovskites
(BaTiO.sub.3, ZnTiO.sub.3), calciumsilicates (like xonotlite),
alumina, silica or silica-alumina's or mixtures of the latter.
[0028] In an embodiment 8 the invention is according to anyone of
embodiments 1 to 4, wherein said refined oils are transformed into
linear or substantially linear paraffin's as bio-naphtha by
hydrolysis of the fats & oils into glycerol and fatty acids,
removal of the glycerol or by physical refining, including a steam
distillation or vacuum distillation of fats & oils or obtained
by acidulation of soaps and decarboxylation of the fatty acids is
carried out on basic oxides, like alkaline oxides, alkaline earth
oxides, lanthanide oxides, zinc-oxide, spinels
(Mg.sub.2Al.sub.2O.sub.4, ZnAl.sub.2O.sub.4), perovskites
(BaTiO.sub.3, ZnTiO.sub.3), calciumsilicates (like xonotlite),
either as bulk material or dispersed on neutral or basic carriers,
on basic zeolites (like alkali or alkaline earth low silica/alumina
zeolites obtained by exchange or impregnation).
[0029] In an embodiment 9 the invention is according to embodiments
7 and 8 , wherein the hydrodeoxygenation is carried out at a
temperature from 200 to 500.degree. C., under a pressure from 1 MPa
to 10 MPa (10 to 100 bars) and with a hydrogen to feedstock ratio
from 100 to 2000 Nl/l. or wherein the decarboxylation is carried
out at a temperature from 100 to 550.degree. C., under a pressure
from 0.1 MPa to 10 MPa (1 to 100 bars) and with a hydrogen to
feedstock ratio from 0 to 2000 Nl/l.
[0030] In an embodiment 10 the invention is according to embodiment
4, wherein the decarboxylation of the soaps is carried out at from
100 to 550.degree. C. under pressure from 0.1 Mpa to 10 Mpa and in
presence of water.
[0031] In an embodiment 11 the invention is according to
embodiments 4 and 10, wherein the decarboxylation of the soaps is
carried out with a water to feedstock ratio of at least 1 mole
water per mole of soap.
[0032] In an embodiment 12 the invention is the Use of the
bio-naphtha as obtained in the process of any one of embodiments 1
to 11, as a direct feedstock of a steamcracker, said bio-naphtha
being used as such, or together with the bio-propane when produced
by the process of embodiment 5, or as blended with at least a
conventional feedstock selected among LPG, naphtha and gasoil, in
order to obtain cracked products including bio-ethylene,
bio-propylene, bio-butadiene, bio-isoprene, bio-cyclopentadiene and
bio-piperylenes, bio-benzene, bio-toluene, bio-xylene and
bio-gasoline, [0033] wherein in the steamcracking process the
hydrocarbon feedstock is mixed with steam in a ratio of 0.3 to 0.45
kg steam per kg hydrocarbon feedstock.
[0034] In an embodiment 13 the invention is the process of
embodiment 12 wherein the hydrocarbon feedstock is mixed with steam
in a ratio of 0.3 to 0.4 kg steam per kg hydrocarbon feedstock.
[0035] In an embodiment 14 the invention is a process for steam
cracking a feedstock as defined in embodiment 12 or 13, wherein the
mixture is heated up to a temperature of 750-950.degree. C. at a
residence time of 0.05 to 0.5 seconds.
[0036] In an embodiment 15 the invention is a process for steam
cracking a feedstock as defined in any one of embodiments 12 to 14,
wherein the mixture is heated up to a temperature of
750-950.degree. C. at a residence time of 0.05 to 0.15 seconds.
[0037] In an embodiment 16 the invention is the Use of the
bio-naphtha as obtained in the process of any one of embodiments 1
to 11 for steamcracking such as to obtain a ethylene to methane
weight ratio, resulting from the cracking of bio-naphtha, of at
least 3.
[0038] By "bio-naphtha" we mean naphtha produced from renewable
sources by hydrotreatment of these renewable sources, It is a
hydrocarbon composition, consisting of mainly paraffin's and that
can be used for the steamcracking to produce light olefins, dienes
and aromatics. The molecular weight of this bio-naphtha ranges from
hydrocarbons having 8 to 24 carbons, preferably from 10 to 18
carbons.
[0039] By "substantially linear paraffins", we mean a composition
of paraffin's consisting of at least 90% by weight of linear
paraffin's.
[0040] Said complex mixture of natural occurring fats & oils
can be selected among vegetable oils and animal fats,
preferentially inedible highly saturated oils, waste food oils,
by-products of the refining of vegetable oils, and mixtures
thereof. Specific examples of these fats & oils have been
previously mentioned in the present specification.
[0041] Said refined oils, eventually still containing some free
fatty acids can be transformed into linear or substantially linear
paraffin's as bio-naphtha together with bio-propane by
hydrodeoxygenation in the presence of hydrogen and of at least one
hydrodeoxygenation catalyst. The hydrodeoxygenation catalyst can be
selected among Ni, Mo, Co or mixtures like NiW, NiMo, CoMo, NiCoW,
NiCoMo, NiMoW and CoMoW oxides or sulphides as catalytic active
phase, preferably supported on high surface area carbon, alumina,
silica, titania or zirconia or group 10 (Ni, Pt or Pd) or group 11
(Cu or Ag) metals or alloy mixtures supported on high surface area
carbon, magnesia, zinc-oxide, spinels (Mg.sub.2Al.sub.2O.sub.4,
ZnAl.sub.2O.sub.4), perovskites (BaTiO.sub.3, ZnTiO.sub.3),
calciumsilicates (like xonotlite), alumina, silica or mixtures of
the latter. It is preferred that the support for the catalytic
active phase exhibit low acidity, preferable neutral or basic in
order to avoid hydro-isomerisation reactions that would result in
branched paraffin's and cracking. The hydrodeoxygenation of the
fats & oils can be carried out at a temperature from 200 to
500.degree. C., preferably from 280 to 400.degree. C., under a
pressure from 1 MPa to 10 MPa (10 to 100 bars), for example of 6
MPa, and with a hydrogen to refined oils ratio from 100 to 2000,
but preferably from 350 to 1500 for example of 600 Nl H2/l oil.
[0042] Said refined oils can also be transformed into linear or
substantially linear paraffin's as bio-naphtha by producing fatty
acids by (i) hydrolysis of the fats & oils into glycerol and
fatty acids, removal of the glycerol,by (ii) physical refining
(steam/vacuum distillation) of fats & oils or by acidulation of
fatty acid soaps and subsequently hydrodeoxygenation or
decarboxylation (or decarbonylation) of the fatty acids, said
hydrodeoxygenation being conducted in the presence of hydrogen and
of at least one hydrodeoxygenation or decarboxylation catalyst. The
hydrodeoxygenation or decarboxylation catalyst can be selected
among Ni, Mo, Co or mixtures like NiW, NiMo, CoMa, NiCoW, NiCoMo,
NiMoW and CoMoW oxides or sulphides as catalytic phase, preferably
supported on high surface area carbon, alumina, silica, titania or
zirconia or group 10 (Ni, Pt or Pd) or group 11 (Cu or Ag) metals
or alloy mixtures supported on high surface area carbon, magnesia,
zinc-oxide, spinels (Mg.sub.2Al.sub.2O.sub.4, ZnAl.sub.2O.sub.4),
perovskites (BaTiO.sub.3, ZnTiO.sub.3), calciumsilicates (like
xonotlite), alumina, silica or silica-alumina's or mixtures of the
latter. It is preferred that the support for the catalytic active
phase exhibit low acidity, preferable neutral or basic in order to
avoid hydro-isomerisation reactions that would result in branched
paraffin's and cracking. The hydrolysis (splitting) can be carried
out in presence of steam thermally at 15 to 75 bars and at 50 -
300.degree. C. or catalytically, for example with basic catalysts,
like MgO, CaO, ZnO, spinels (Mg2Al2O4, ZnAl2O4), perovskites
(BaTiO3, ZnTiO3), calciumsilicates (like xonotlite) or basic
alumina or with acidic catalysts, like sulphuric acid. Detailed
information about fat & oil splitting has been published by
Sonntag (Sonntag, N., J. Am. Oil. Chem. Soc., 56, p. 729, 1979 and
Bailey's Industrial Oil and Fat Products, ed. F. Shahidi, 2005,
John Wiley & Sons). In the Colgate--Emery process, heated
liquid lipid is introduced at the bottom of a vertical tubular
reactor. Heated water enters at the top. As the fats & oils
rises through the descending water under pressure, a continuous
zone of high water solubility in oil establishes, wherein
hydrolysis occurs. Effluent from the column is recovered, fatty
acids from one outlet and an aqueous glycerol stream from the
other. The presence of small amounts of mineral acids, such as
sulfuric acid or sulfonic acids or certain metal oxides, such as
zinc or magnesium oxide, accelerates the splitting reaction. These
metal oxides are true catalysts and they assist also in the
formation of emulsions.
[0043] The hydrodeoxygenation of the fatty acids can be carried out
at a temperature from 200 to 500.degree. C., preferably from 280 to
400.degree. C., under a pressure from 1 MPa to 10 MPa (10 to 100
bars), for example of 6 MPa, and with a hydrogen to refined oils
ratio from 100 to2000Nl/l, for example of 600 Nl H2/l oil. The
decarboxylation of the fatty acids can be carried out at 100 to
550.degree. C. in absence or presence of hydrogen at pressures
ranging from 0.01 up to 10 MPa.
[0044] Said refined oils can also be transformed into linear or
substantially linear paraffin's as bio-naphtha by thermal
decarboxylation of fatty acid soaps. These soaps are obtained
during chemical refining by neutralisation to convert free fatty
acids into soaps, by neutralisation of fatty acids obtained by
hydrolysis of fats & oils or by complete saponification of
triglycerides into glycerol and soap. A soap is a metal salt of the
corresponding fatty acid.
[0045] The present invention also relates to the use of the
bio-naphtha as obtained in the above mentioned process, as a direct
feedstock of a steamcracker, said bio-naphtha being used as such,
or together with the bio-propane when produced by the
above-mentioned process, or as blended with at least a conventional
feedstock selected among LPG, naphtha and gasoil, in order to
obtain cracked products including bio-ethylene, bio-propylene,
bio-butadiene, bio-isoprene, bio-(di)cyclopentadiene,
bio-piperylenes, bio-benzene, bio-toluene, bio-xylene and
bio-gasoline.
[0046] Moreover, the present invention relates to a process for
steam cracking a feedstock as defined above, wherein said feedstock
is mixed with steam, having a steam/feedstock ratio of at least 0.2
kg per kg of feedstock. This mixture is sent through the heated
coils, having a coil outlet temperature of at least 700 .degree. C.
and a coil outlet pressure of at least 1.2 bara.
DETAILED DESCRIPTION OF THE INVENTION
[0047] All crude fats & oils obtained after rendering, crushing
or solvent extraction inevitably contain variable amounts of
non-triglyceride components such as free fatty acids, mono and
diglycerides, phosphatides, sterols, tocopherols, tocotrienols
hydrocarbons, pigments (gossypol, chlorophyll), vitamins
(carotenoids), sterols glucosides, glycolipids, protein fragments,
traces of pesticides and traces metals, as well as resinous and
mucilaginous materials. The quantities of the non-glycerides vary
with the oil source, extraction process, season and geographical
source. Removal of the non-triglyceride components, which interfere
with further processing and cause the oil to darken, foam, smoke,
precipitate and develop off-flavours, is the objective of the
refining process.
Refining Pretreatment
Choice of the Refining Method
[0048] FIG. 1 illustrates the refining pretreatment in which crude
oils are processed through various routes, physical or chemical, to
Refined Bleached Deodorized (RBD) oils. Physical refining and
alkali/chemical refining differ principally in the way free fatty
acids are removed.
[0049] In chemical refining, FFA, most of the phosphatides, and
other impurities are removed during neutralization with an alkaline
solution, usually NaOH.
[0050] In physical refining, the FFA is removed by distillation
during deodorization and the phosphatides and other impurities must
be removed prior to steam distillation of fats & oils
[0051] Currently, the refining method of choice is determined by
the characteristics of the individual crude fats & oils:
[0052] (1) fats and oils that are normally physically refined;
[0053] (2) fats and oils that can be physically or chemically
refined; and
[0054] (3) fats and oils that can only be chemically refined.
[0055] Table 2 below summarizes advantages and disadvantages of
each treatment:
TABLE-US-00002 TABLE 2 Refining type Advantages Disadvantages
Chemical Functional process Production of by-products refining Not
restricted by the oil type Expensive process Successful reduction
of FFA High loss of oil Physical Cheaper Not suitable for all types
refining of oils Less by-products Requires high temperature and
vacuum Less energy consumed Can form undesired side reaction
products
Physical Refining
[0056] The physical refining can remove the FFA, as well as the
unsaponifiables and other impurities by steam stripping, thus
eliminating the production of soapstock and keeping neutral oil
loss to a minimum. However, degumming pretreatments of the crude
fats & oils are still required to remove those impurities that
darken or otherwise cause a poor-quality product when heated to the
temperature required for steam distillation. A degumming process is
crucial for physical refining but optional for chemical refining.
It consists of the treatment of crude oils, with water, salt
solutions, enzymes, caustic soda, or diluted acids such as
phosphoric, citric or maleic to remove phosphatides, waxes,
pro-oxidants and other impurities. The degumming processes convert
the phosphatides to hydrated gums, which are insoluble in oil and
readily separated as a sludge by settling, filtering or centrifugal
action. After degumming, phosphorous must be less than 30 ppm. So
that bleaching or dry degumming can further reduce this level to
less than 5 ppm and remove all traces of iron and copper. Acid or
enzymatic degumming processes are normally employed to achieve
these results.
[0057] The various industrial degumming processes have different
aims. Fats & oils to be degummed vary widely in gum content and
gum properties and finally, the means of gum disposal available,
what equipment is needed and/or available, and the cost of
auxiliaries also influence the choice of the most appropriated
degumming process. The lipid handbook (The lipid handbook, edited
by Frank D. Gunstone, John L. Harwood, Albert J. Dijkstra. 3rd ed.,
chapter 3.4) deals with these aspects in details. Next is briefly
described the four major degumming process applied on the
market.
[0058] The main purposes of the water degumming process are to
produce oil that does not deposit a residue during transportation
and storage, and to control the phosphorus content of crude oils
just below 200 ppm. This process involves the addition of live
steam to raw oil for a short period. The proper amount of water is
normally about 75% of the phosphatides content of the oil. Too
little water produces dark viscous gums and hazy oil, while too
much water causes excess oil losses through hydrolysis.
Water-degummed oil still contains phosphatides (between 80 and 200
ppm); only hydratable phosphatides are removed with this process.
The nonhydratable phosphatides, which are calcium and magnesium
salts of phosphatic acid and phosphatidyl ethanolamine, remain in
the oil after water degumming.
[0059] Acid degumming process leads to a lower residual phosphorus
content than water degumming and is therefore a good alternative if
dry degumming and physical refining are to be the next refining
steps. The acid degumming process might be considered as a variant
of the water degumming process in that it uses a combination of
water and acid. The non-hydratable phosphatides can be conditioned
into hydratable forms with acid degumming. Phosphoric and citric
acids are used because they are food grade, sufficiently strong and
they bind divalent metal ions. Several acid degumming processes
have been developed to attain a phosphorus value lower than 5 ppm
that is required for good quality physically refined oils.
[0060] An acid refining differs from the acid degumming by the
neutralisation of the liberated phosphatides (the action of the
degumming acid does not lead to full hydration of the phosphatides)
to make them hydratable by the addition of a base.
[0061] In dry degumming process, the oil is treated with an acid
(principle is that strong acids displace weaker acids from their
salts) to decompose the metal ion/phosphatides complex and is then
mixed with bleaching earth. The earth containing the degumming
acid, phosphatides, pigments and other impurities is then removed
by filtration. Seed oils that have been water or acid-degummed may
also be dry degummed to ensure a low phosphorus oil to steam
distillation. An increase in FFA of less than 0.2% should be
expected but the final phosphorus content must be reduced to less
than 5 ppm. This process constitutes the main treatment for palm
oil, laurie oils, canola oil and low phosphatides animal fats, such
as tallow or lard. The dry degumming process allows crude oil to be
fully refined in only two steps: dry degumming and physical
refining.
[0062] In enzymatic degumming process, Phospholipase Al, the
lastest developed degumming enzyme, changes the phospholipids into
lysophospholipids and free fatty acids. This process has three
important steps:
[0063] (1) adjustement of the pH with a buffer;
[0064] (2) enzymatic reaction in the holding tanks; and
[0065] (3) separation of the sludge from the oil.
[0066] Oil to be degummed enzymatically by this way can be crude or
water degummed.
[0067] The lipid handbook (The lipid handbook, edited by Frank D.
Gunstone, John L. Harwood, Albert J. Dijkstra. 3rd ed.) describes
many variants and details of the degumming processes.
[0068] The purpose of bleaching is to provide a decoloured oil but
also to purify it in preparation for further processing. All fully
refined oils have been subjected to one or the other bleaching
process. Refined oil contains traces of a number of undesirable
impurities either in solution or as colloidal suspensions. The
bleaching process does more than just increasing the transmission
of light through the oil and is often called "adsorptive cleaning".
The bleaching process is often the first filtration encountered by
the oil, so it ensures the removal of soaps, residual phosphatides,
trace metals, and some oxidation products, and it catalyses the
decomposition of carotene and the adsorbent also catalyses the
decomposition of peroxides. These non-pigment materials, such as
soap, gums, and pro-oxidants metals, which hinder filtration,
poison hydrogenation catalyst, darken the oils, and affect finished
oil flavour. Another function is the removal of the peroxides and
secondary oxidation products. The key parameters for the bleaching
process are procedure, adsorbent type and dosage, temperature,
time, moisture and filtration, as shown in the Lipid Handbook (The
lipid handbook, edited by Frank D. Gunstone, John L. Harwood,
Albert J. Dijkstra. 3rd ed., chapter 3.7). The three most common
types of contact bleaching methods used for edible fats and oils
are batch atmospheric, batch vacuum and continuous vacuum. Chemical
agents have been used or proposed for use but practically all
edible oil decolouration and purification is accomplished with
adsorptive clays, synthetic amorphous silica and activated
carbons.
[0069] Before the last major processing step, bleached oil can be
hydrogenated, for two reasons. One reason is to change naturally
occurring fats & oils into physical forms with the consistency
and handling characteristics required for functionality. The second
reason for hydrogenation is to increase the oxidation and thermal
stability. Instead of purification in other described processes,
this step consists in fats & oils molecular modification.
[0070] Hydrogen is added directly to react with unsaturated oil in
the presence of catalysts, mostly nickel. This process greatly
influences the desired stability and properties of many edible oil
products. The hydrogenation process is easily controlled and can be
stopped at any point. A gradual increase in the melting point of
fats and oils is one of the advantages. If the double bonds are
eliminated entirely with hydrogenation, the product is a hard
brittle solid at room temperature. Shortening and margarine are
typical examples. A wide range of fats and oils products can be
produced with the hydrogenation process depending upon the
conditions used, the starting oils, and the degree of saturation or
isomerization.
[0071] To obtain good-quality fats and oils with physical refining,
it is advantageous to have a phosphorous content lower than 5 ppm
before steam stripping.
[0072] The degummed-bleached oils are vacuum stripped. This process
encompasses the deodorization process, applied after the alkali
routes, as well as physical refining. Deodorization, the last major
processing step during which the FFA can be removed, is a
vacuum-steam distillation process (1-2 mbar of residual pressure)
at elevated temperature (180-240.degree. C.) during which FFAs and
minute levels of odoriferous materials, mostly arising from
oxidation, are removed to obtain a bland and odourless oil. In
order to volatilise the undesired high-boiling components, a deep
vacuum and dilution with steam is applied so that the boiling
temperature can be minimised. The deodorization utilizes the
differences in volatility between off-flavour and off-odor
substances and the triglycerides.
[0073] The odoriferous substances, FFAs, aldehydes, ketones,
peroxides, alcohols, and others organic compounds are concentrated
in a deodorizer distillate. Efficient removal of these substances
depends upon their vapour pressure, for a given constituent is a
function of the temperature and increases with the temperature.
[0074] As usually the last stage in the refining process,
deodorization has an important effect an overall refined oil
quality and distillate composition, Its main purposes are giving a
bland taste and smell, low FFA content, high oxidative stability
and light and stable colour. Because of the need of a rather high
temperature to remove the undesired components, unwanted side
effects are, isomerisation of double bond, polymerisation,
intra-esterification and degradation of vitamins and anti-oxidants.
New dry condensing (steam is condensed into ice) vacuum systems
capable of reaching a very low operating pressure in the deodorizer
were introduced (close to 0.1 kPa). This progress allows a
reduction of the deodorization temperature without affecting the
stripping efficiency in a negative way. In order to minimise the
time that the oil is at high temperature, deodorizers can operate
at dual temperatures to reach the best compromise between required
residence time for deodorizing (at moderate temperature) and heat
bleaching and final stripping at high temperature.
[0075] Deodorizer distillate is the material collected from the
steam distillation of edible oils. The distillate from physically
refined oils consists mainly of FFAs with low levels of
unsaponifiable components. The concentration of FFA can be improved
from typical 80% up to 98% by applying double condensing system
that produces an enriched FFA cut. The distillate can be used as a
source of industrial fatty acids or mixed with the fuel oil used to
fire the steam boilers.
[0076] A physical refining will be preferred due to higher
remaining FFA content in refined oils before steam stripping.
Chemical Refining
[0077] As applied to crude oils, it includes degumming (removal of
phospholipids), neutralization (removal of free fatty acids),
bleaching (decolourisation) and deodorization (FIG. 1).
[0078] Degumming involves for instance the addition of water to
hydrate any gums present, followed by centrifugal separation.
Non-hydratable gums are removed by converting them first to a
hydratable form using phosphoric or citric acid, followed by the
addition of water and centrifugation. Acid degumming can also be
used (see the description above).
[0079] The following step is neutralisation in which an aqueous
alkali, typically caustic soda or sodium carbonate, is sprayed into
the oil which has been preheated to around 75-95.degree. C. The
alkali reacts with free fatty acids in the oil to form soaps, which
are separated by settling or centrifugation. Selection of the
aqueous alkali strength, mixing time, mixing energy, temperature,
and the quantity of excess caustic all have an important impact on
making the chemical refining process operate efficiently and
effectively. A drying step may be incorporated after neutralisation
to ensure the complete removal of the added water. The soap can be
used as such or can be hydrolysed (acidulation) with sulphuric acid
into the corresponding FFA.
[0080] The neutralised oil is bleached to remove colouring matter
(such as carotenoids) and other minor constituents, such as
oxidative degradation products or traces of metals. Bleaching uses
activated fuller's earth with treatments typically in the
90-130.degree. C. range for 10-60 minutes. The earth is sucked into
the oil under vacuum and is removed by filtration.
[0081] The bleached oil is steam distilled at low pressure to
remove volatile impurities including undesirable odours and
flavours. This process, known as deodorisation, takes place in the
temperature range of 180-270.degree. C. and may last 15 minutes to
five hours depending upon the nature of the oil, the quantity, and
the type of equipment used.
Obtention of Bio-Naphtha from Refined Oils
[0082] Three options exist to convert fats & oils into LPG and
naphtha-like hydrocarbons that can be used for the steamcracking in
order to produce light olefins, dienes and aromatics.
[0083] These are summarised in table 3.
TABLE-US-00003 TABLE 3 Catalyst/intermediate Feedstock Process
compounds Triglycerides, Catalytic Supported Ni, Mo, Co, NiW,
eventually Hydrodeoxygenation NiMo, CoMo, NiCoW, containing
Catalytic NiCoMo, NiMoW and fatty acids Decarboxylation CoMoW
oxides or sulphides Fatty acids Supported group 10 (Ni, Pt, Pd) or
group 11 (Cu, Ag) metals or alloys Basic oxides or mixed basic
oxides Fatty acids Thermal Decarboxylation Soaps of alkali,
alkaline Soaps earth, lanthanides or group 12 or 13
[0084] The first option consists in hydrodeoxygenation, which
removes the oxygen atoms from the fats & oils. This can be done
on the triglycerides as such, the triglycerides containing FFA's or
on only FFA's. Hydrodeoxygenation of fats & oils has been
reported in 1989 (W. H. Craig and D. W. Soveran, "Production of
hydrocarbons with relatively high cetane rating", U.S. Pat. No.
4,992,605 and Gusmao J, Brodzki D, Djega-Mariadassou G, Frety R.,
"Utilization of vegetable oils as an alternative source for
diesel-type fuel: Hydrocracking on reduced Ni/SiO2 and sulphided
Ni-Mo/.gamma.-Al2O3", Cat. Today 1989 (5) 533) in which
conventional CoMo or NiMo sulphided catalysts are used. These
catalysts are well known in hydrodesulphurization and are known to
catalyze also hydrodeoxygenation (E. Furimsky, Applied Catalysis A,
General, 199, pages 147-190, 2000).
[0085] Hydrodeoxygenation of fats & oils is preferentially done
in continuous fixed bed reactors, continuous stirred tank reactors
or slurry type reactors containing solid catalyst that can be
selected among Ni, Mo, Co or mixtures like NiW, NiMo, CoMo, NiCoW,
NiCoMo, NiMoW and CoMoW oxides or sulphides as catalytic phase,
preferably supported on high surface area carbon, alumina, silica,
titania or zirconia. It is preferred that the support for the
catalytic active phase exhibit low acidity, preferable neutral or
basic in order to avoid hydro-isomerisation reactions that would
result in branched paraffin's and cracking at elevated temperature
and pressure in the presence of hydrogen. Temperature ranges from
200 to 500.degree. C., pressure from 1 MPa to 10 MPa (10 to 100
bars) and hydrogen to oil feed ratio from 100 to 2000
Nm.sup.3/m.sup.3 of liquid. For optimum performance and stable
continuous operation, it is preferred that the active metal
component of the catalyst is in the form of sulfides. Thereto, it
is preferred that traces amounts of decomposable sulphur compounds
are present or added on purpose to the feedstock in order to keep
the metal sulphide in its sulphide state. By way of example, these
sulphur compounds can be H.sub.2S, COS, CS.sub.2, mercaptans (e.g.
methylsulfide), thio-ethers (e.g. DiMethylSulfide), disulfides
(e.g. DiMethyldiSulfide), thiophenic and tetrahydrothiophenic
compounds.
[0086] Under hydrodeoxygenation conditions several reactions occur.
The easiest is the hydrogenation of the double bonds in the
alkyl-chain. The more difficult reaction is the removal of oxygen
atoms from the C--O bonds. Both the carboxyl-group of the fatty
acid as the hydroxyl-group of the glycerol-moiety are
hydrodeoxygenated. This results in the production of linear
paraffin, resulting from the fatty acid and in propane, resulting
from glycerol. Depending on the conditions (catalyst, temperature,
hydrogen etc), the carboxyl-group can also be decomposed into
CO/CO.sub.2 (decarboxylation) and which on their turn can be even
further hydrogenated into methane. These hydrodeoxygenation
reactions consume a lot of hydrogen.
[0087] As way of example is given the equation for triolein
hydrodeoxygenation:
[C.sub.18H.sub.33O].sub.3C.sub.3H.sub.5O.sub.3+15H.sub.2.fwdarw.3C.sub.1-
8H.sub.38+C.sub.3H.sub.8+6H.sub.2O
[0088] Hydrodeoxygenation of fatty acids:
R--CH.sub.2--CH.sub.2--COOH+3H.sub.2--R--CH.sub.2--CH.sub.2--CH.sub.3+2H-
.sub.2O
[0089] Further hydrogenation of the intermediate CO/CO.sub.2 can
occur depending on the amount of available hydrogen, the catalyst
and the operating conditions:
CO+3H.sub.2.fwdarw.CH.sub.4+H.sub.2O
CO.sub.2+4H.sub.2.fwdarw.CH.sub.4+2H.sub.2O
[0090] The second option consists in decarboxylation or
decarbonylation of fatty acids. These fatty acids can be obtained
from fats & oils by physical refining (including steam/vacuum
distillation), by (steam) splitting of triglycerides or by
splitting of soaps (acidulation) using acids. Decarboxylation of
carboxylic acids has been reported in 1982 (W. F. Maier, Chemische
Berichte, 115, pages 808-812, 1982) over Pd/SiO.sub.2 and
Ni/Al.sub.2O.sub.3 catalysts in the gas phase. A highly selective
decarboxylation has been reported in 2005 (I. Kubickova, Catalysis
Today, 106, pages 197-200, 2005 and M. Snare, Industrial
Engineering, Chemistry Research, 45, p. 5708-5715, 2006) using
transition metal catalysts. Palladium based catalysts exhibit the
highest selectivity towards decarboxylation. Carboxylic acids can
also be decarboxylated under catalytic conditions using basic
catalyst, like MgO, ZnO and mixed basic oxides (A. Zhang *, Q. Ma,
K. Wang, X. Liu, P. Shuler, Y. Tang, "Naphthenic acid removal from
crude oil through catalytic decarboxylation on magnesium oxide",
Applied Catalysis A: General 303, p. 103, 2006; A. More, John R.
Schlup, and Keith L. Hohn "Preliminary Investigations of the
Catalytic Deoxygenation of Fatty Acids", AIChe, The 2006 annual
meeting, San Francisco and B. Kitiyanan, C. Ung-jinda, V. Meeyoo,
"Catalytic deoxygenation of oleic acid over ceria-zirconia
catalysts", AIChe The 2008 annual meeting).
[0091] The following reactions can occur:
[0092] Decarboxylation:
R--CH.sub.2--CH.sub.2--COOH--R--CH.sub.2--CH.sub.3+CO.sub.2
[0093] Decarbonylation:
R--CH.sub.2--CH.sub.2--COOH.fwdarw.R--CH.dbd.CH.sub.2+CO+H.sub.2O
[0094] Decarboxylation is preferentially done in presence of solid
catalyst in batch type tank reactors, continuous fixed bed type
reactors, continuous stirred tank reactors or slurry type reactors.
The catalyst can be selected among Ni, Mo, Co or mixtures like NiW,
NiMo, CoMo, NiCoW, NiCoMo, NiMoW and CoMoW oxides or sulphides as
catalytic phase, preferably supported on high surface area carbon,
alumina, silica, titania or zirconia or group 10 (Ni, Pt and Pd)
and group 11 (Cu and Ag) metals or alloy mixtures supported on high
surface area carbon, magnesia, zinc-oxide, spinels
(Mg.sub.2Al.sub.2O.sub.4, ZnAl.sub.2O.sub.4), perovskites
(BaTiO.sub.3, ZnTiO.sub.3), calciumsilicates (like xonotlite),
alumina, silica or silica-alumina's or mixtures of the latter. It
is preferred that the support for the catalytic active phase
exhibit low acidity, preferable neutral or basic in order to avoid
hydro-isomerisation reactions that would result in branched
paraffin's and cracking. Decarboxylation can also be carried out on
basic oxides, like alkaline oxides, alkaline earth oxides,
lanthanide oxides, zinc-oxide, spinels (Mg.sub.2Al.sub.2O.sub.4,
ZnAl.sub.2O.sub.4), perovskites (BaTiO.sub.3, ZnTiO.sub.3),
calciumsilicates (like xonotlite), either as bulk material or
dispersed on neutral or basic carriers, on basic zeolites (like
alkali or alkaline earth low silica/alumina zeolites obtained by
exchange or impregnation).
[0095] Although, the decarboxylation reaction does not require
hydrogen, it is preferred that the decarboxylation is done in
presence of hydrogen that will stabilise the catalytic activity by
removing strongly adsorbed unsaturated species (for instance when
decarbonylation is the prevalent reaction pathway) from the
catalyst surface by hydrogen-addition reactions. The presence of
hydrogen can also hydrogenate the double bonds present in the
acyl-moiety of the fatty acid in order to obtain paraffinic
reaction products from the decarboxylation process. The
decarboxylation of the fatty acids can be carried out at 100 to
550.degree. C. in absence or presence of hydrogen at pressures
ranging from 0.01 up to 10 MPa. The hydrogen to feedstock ratio is
from 0 to 2000 Nl/l
[0096] Other reactions that can occur under the decarboxylation
conditions are:
R--CH.dbd.CH.sub.2+H.sub.2.fwdarw.R--CH.sub.2--CH.sub.3
[0097] Hydrodeoxygenation of fatty acids:
R--CH.sub.2--CH.sub.2--COOH+3H.sub.2.fwdarw.R--CH.sub.2--CH.sub.2--CH.su-
b.3+2H.sub.2O
[0098] Further hydrogenation of the intermediate CO/CO.sub.2 can
occur depending on the amount of available hydrogen, the catalyst
and the operating conditions:
CO+3H.sub.2.fwdarw.CH.sub.4+H.sub.2O
CO.sub.2+4H.sub.2.fwdarw.CH.sub.4+2H.sub.2O
[0099] A third option to obtain bio-naphtha from fats & oils is
through the thermal decarboxylation of soaps of fatty acids. The
soaps can be obtained from the chemical refining of fats & oils
by neutralisation, producing refined triglycerides and soaps, by
neutralisation of fatty acids obtained after (steam) splitting of
fats & oils or by direct saponification of fats & oils
using basic oxides or basic hydroxides, producing a soap and
glycerol.
[0100] Decarboxylation has been carried out by decomposition of
fatty acids in hot compressed water with the aid of
alkali-hydroxides, resulting in the production of alkanes and
CO.sub.2 (M. Watanabe, Energy Conversion and Management, 47, p.
3344, 2006). Calcium-soaps of Tung oil have been reported to
decompose by distillation as early as 1947 (C.C, Chang, S. W, Wan,
"China's Motor Fuels from Tung Oil", Ind. Eng. Chem, 39 (12), p.
1543, 1947; Hsu, H. L., Osburn, J. O., Grove, C. S., "Pyrolysis of
the calcium salts of fatty acids", Ind. Eng. Chem. 42 (10), p.
2141, 1950; Craveiro, A. A.; Matos, F. J. A.; Alencar, J. W.;
Silveira E. R. Energia: Fontes Altemativas 3, p. 44, 1981; A.
Demirbas, "Diesel fuel from vegetable oil via transesterification
and soap pyrolysis", Energy Sources 24 9, p. 835, 2002).
[0101] The preferred soaps are those made of alkaline, alkaline
earth, lanthanide, zinc or aluminium cations. The thermal
decarboxylation of soap can be carried out by heating until the
molten soap starts to decompose into the corresponding paraffin's
or olefins and the corresponding metal-carbonate or
metal-oxide/hydroxide and CO.sub.2. Without willing to be bound to
any theory, it is believed that the following overall reactions
occur:
[R--CH.sub.2--CH.sub.2--COO.sup.-].sub.xM.sup.x++xH.sub.2O.fwdarw.xR--CH-
.sub.2--CH.sub.3+M[HCO.sub.3].sub.x
M[HCO.sub.3].sub.x.revreaction.M[OH].sub.x+CO.sub.2
[0102] It is preferred that the thermal decomposition of the soaps
is carried out in the presence of liquid, supercritical or vaporous
water.
Steamcracking
[0103] Steamcrackers are complex industrial facilities that can be
divided into three main zones, each of which has several types of
equipment with very specific functions: (i) the hot zone including:
pyrolysis or cracking furnaces, quench exchanger and quench ring,
the columns of the hot separation train (ii) the compression zone
including: a cracked gas compressor, purification and separation
columns, dryers and (iii) the cold zone including: the cold box,
de-methaniser, fractionating columns of the cold separation train,
the C.sub.2 and C.sub.3 converters, the gasoline hydrostabilization
reactor Hydrocarbon cracking is carried out in tubular reactors in
direct-fired heaters (furnaces). Various tube sizes and
configurations can be used, such as coiled tube, U-tube, or
straight tube layouts. Tube diameters range from 1 to 4 inches.
Each furnace consists of a convection zone in which the waste heat
is recovered and a radiant zone in which pyrolysis takes place. The
feedstock-steam mixture is preheated in the convection zone to
about 530-650.degree. C. or the feedstock is preheated in the
convection section and subsequently mixed with dilution steam
before it flows over to the radiant zone, where pyrolysis takes
place at temperatures varying from 750 to 950.degree. C. and
residence times from 0.05 to 0.5 second, depending on the feedstock
type and the cracking severity desired. In an advantageous
embodiment the residence time is from 0.05 to 0.15 second. The
steam/feedstock (the steam/[hydrocarbon feedstock]) weight ratio is
between 0.2 and 1.0 kg/kg, preferentially between 0.3 and 0.5
kg/kg. In an advantageous embodiment the steam/feedstock weight
ratio is between 0.2 and 0.45 and preferably between 0.3 and 0.4.
For steamcracking furnaces, the severity can be modulated by:
temperature, residence time, total pressure and partial pressure of
hydrocarbons. In general the ethylene yield increases with the
temperature while the yield of propylene decreases. At high
temperatures, propylene is cracked and hence contributes to more
ethylene yield. The increase in severity thus obtained leads to a
moderate decrease in selectivity and a substantial decrease of the
ratio C.sub.3'/C.sub.2.dbd.. So high severity operation favors
ethylene, while low severity operation favors propylene production.
The residence time of the feed in the coil and the temperature are
to be considered together. Rate of coke formation will determine
maximum acceptable severity. A lower operating pressure results in
easier light olefins formation and reduced coke formation. The
lowest pressure possible is accomplished by (i) maintaining the
output pressure of the coils as close as possible to atmospheric
pressure at the suction of the cracked gas compressor (ii) reducing
the pressure of the hydrocarbons by dilution with steam (which has
a substantial influence on slowing down coke formation). The
steam/feed ratio must be maintained at a level sufficient to limit
coke formation.
[0104] Effluent from the pyrolysis furnaces contains unreacted
feedstock, desired olefins (mainly ethylene and propylene),
hydrogen, methane, a mixture of C.sub.4's (primarily isobutylene
and butadiene), pyrolysis gasoline (aromatics in the C.sub.6 to
C.sub.8 range), ethane, propane, di-olefins (acetylene, methyl
acetylene, propadiene), and heavier hydrocarbons that boil in the
temperature range of fuel oil. This cracked gas is rapidly quenched
to 338-510.degree. C. to stop the pyrolysis reactions, minimize
consecutive reactions and to recover the sensible heat in the gas
by generating high-pressure steam in parallel transfer-line heat
exchangers (TLE's). In gaseous feedstock based plants, the
TLE-quenched gas stream flows forward to a direct water quench
tower, where the gas is cooled further with recirculating cold
water. In liquid feedstock based plants, a prefractionator precedes
the water quench tower to condense and separate the fuel oil
fraction from the cracked gas. In both types of plants, the major
portions of the dilution steam and heavy gasoline in the cracked
gas are condensed in the water quench tower at 35-40.degree. C. The
water-quench gas is subsequently compressed to about 25-35 Bars in
4 or 5 stages. Between compression stages, the condensed water and
light gasoline are removed, and the cracked gas is washed with a
caustic solution or with a regenerative amine solution, followed by
a caustic solution, to remove acid gases (CO.sub.2, H.sub.2S and
SO.sub.2). The compressed cracked gas is dried with a desiccant and
cooled with propylene and ethylene refrigerants to cryogenic
temperatures for the subsequent product fractionation: Front-end
demethanization, Front-end depropanization or Front-end
deethanization.
[0105] In a front-end demethanization configuration, tail gases
(CO, H.sub.2, and CH.sub.4) are separated from the C.sub.2+
components first by de-methanization column at about 30 bars. The
bottom product flows to the de-ethanization, of which the overhead
product is treated in the acetylene hydrogenation unit and further
fractionated in the C.sub.2 splitting column. The bottom product of
the de-ethanization goes to the de-propanization, of which the
overhead product is treated in the methyl acetylene/propadiene
hydrogenation unit and further fractionated in the C.sub.3
splitting column. The bottom product of the de-propaniser goes to
the de-butanization where the C.sub.4's are separated from the
pyrolysis gasoline fraction, In this separation sequence, the
H.sub.2 required for hydrogenation is externally added to C.sub.2
and C.sub.3 streams. The required H.sub.2 is typically recovered
from the tail gas by methanation of the residual CO and eventually
further concentrated in a pressure swing adsorption unit.
[0106] Front-end de-propanization configuration is used typically
in steamerackers based on gaseous feedstock, In this configuration,
after removing the acid gases at the end of the third compression
stage, the C.sub.3 and lighter components are separated from the
C.sub.4+ by de-propanization. The de-propanizer C.sub.3- overhead
is compressed by a fourth stage to about 30-35 bars. The acetylenes
and/or dienes in the C.sub.3- cut are catalytically hydrogenated
with H.sub.2 still present in the stream. Following hydrogenation,
the light gas stream is de-methanized, de-ethanized and C.sub.2
split. The bottom product of the de-ethanization can eventually be
C.sub.3 split. In an alternative configuration, the C.sub.3-
overhead is first de-ethanised and the C.sub.2- treated as
described above while the C.sub.3's are treated in the C.sub.3
acetylene/diene hydrogenation unit and C.sub.3 split. The C.sub.4+
de-propanizer bottom is de-butanized to separate Co from pyrolysis
gasoline. There are two versions of the front-end de-ethanization
separation sequence. The product separation sequence is identical
to the front-end de-methanization and front-end depropanization
separation sequence to the third compression stage. The gas is
de-ethanized first at about 27 bars to separate C.sub.2- components
from C.sub.3+ components. The overhead C.sub.2- stream flows to a
catalytic hydrogenation unit, where acetylene in the stream is
selectively hydrogenated. The hydrogenated stream is chilled to
cryogenic temperatures and de-methanized at low pressure of about
9-10 bars to strip off tail gases. The C.sub.2 bottom stream is
split to produce an overhead ethylene product and an ethane bottom
stream for recycle. In parallel, the C.sub.3+ bottom stream from
the front-end de-ethanizer undergoes further product separation in
a de-propaniser, of which the overhead product is treated in the
methyl acetylene/propadiene hydrogenation unit and further
fractionated in the C.sub.3 splitting column. The bottom product of
the de-propaniser goes to the de-butanization where the C.sub.4's
are separated from the pyrolysis gasoline fraction. In the more
recent version of the front-end de-ethanization separation
configuration, the cracked gas is caustic washed after three
compression stages, pre-chilled and is then de-ethanized at about
16-18 bars top pressure. The net overhead stream (C.sub.2-) is
compressed further in the next stage to about 35-37 bars before it
passes to a catalytic converter to hydrogenate acetylene, with
hydrogen still contained in the stream. Following hydrogenation,
the stream is chilled and de-methanized to strip off the tail gases
from the C.sub.2 bottom stream. The C.sub.2's are split in a low
pressure column operating at 9-10 bars pressure, instead of 19-24
bars customarily employed in high pressure C.sub.2 splitters that
use a propylene refrigerant to condense reflux for the column. For
the low-pressure C.sub.2 splitter separation scheme, the overhead
cooling and compression system is integrated into a heat-pump,
open-cycle ethylene refrigeration circuit. The ethylene product
becomes a purged stream of the ethylene refrigeration recirculation
system.
[0107] The ethane bottom product of the C.sub.2 splitter is
recycled back to steam cracking. Propane may also be re-cracked,
depending on its market value. Recycle steam cracking is
accomplished in two or more dedicated pyrolysis furnaces to assure
that the plant continues operating while one of the recycle
furnaces is being decoked.
[0108] Many other variations exist of the above-described
configurations, in particular in the way the undesired
acetylene/dienes are removed from the ethylene and propylene
cuts.
[0109] Various embodiments are represented in FIGS. 2 to 4.
[0110] In a first embodiment (FIG. 2), Fats & Oils are
physically refined by vacuum distillation or steam distillation
(10) to recover the mixed fatty acids (12) as overhead product and
the triglycerides (11) as bottom product. Either the fats &
oils, eventually still containing free fatty acids (21) or the
physically refined triglycerides (20) acids can be sent to a
hydrodeoxygenation section where they are converted into
bio-naphtha (31) and bio-propane (30). This bio-naphtha (41) and
bio-propane (43) are sent to the to steamcracking (50) or blended
with fossil LPG, naphtha or gasoil (40) and hence the blend is
streamcracked (50). The products of the steamcracking are cooled,
compressed, fractionated and purified (51). This results in light
olefins (ethylene, propylene and butenes), dienes (butadiene,
isoprene, (di)cyclopentadiene and piperylenes), aromatics (benzene,
toluene and mixed xylenes) and gasoline as main components.
[0111] In a second embodiment (FIG. 3), Fats & Oils are
physically refined by vacuum distillation or steam distillation
(10) to recover the mixed fatty acids (12) as overhead product and
the triglycerides (11) as bottom product. Optionally fats &
Oils can be hydrolysed (21) to produce mixed fatty acids (22) and
glycerol (23). The quality of the mixed fatty acids can be further
improved by hydrogenation of double bonds in the acyl-moiety or
before hydrolysis, the fats & oils can be hydrogenated to
remove the remaining double bonds and subsequently sent (21) to the
hydrolysis step. The mixed fatty acids can be sent (30) to a
hydrodeoxygenation section where they are converted into
bio-naphtha (36) or alternatively they can be sent to the
decarboxylation section (31) where they are converted into
bio-naphtha (35). This bio-naphtha (41) is sent to the to
steamcracking (50) or blended with fossil LPG, naphtha or gasoil
(40) and hence the blend is streamcracked (50). The products of the
steamcracking are cooled, compressed, fractionated and purified
(51). This results in light olefins (ethylene, propylene and
butenes), dienes (butadiene, isoprene, (di)cyclopentadiene and
piperylenes), aromatics (benzene, toluene and mixed xylenes) and
gasoline as main components.
[0112] In a third embodiment (FIG. 4), fats & Oils are
saponificated (21) to recover the soap (22) and glycerol (23).
Optionally fats & Oils can hydrolysed (21) to produce mixed
fatty acids (22) and glycerol (23). Alternatively, soap (25) can be
obtained during a chemical refining step of raw fats & oils
(24) by the neutralisation step. Still another source of soap (30)
is via neutralisation (29) of fatty acids, obtained by (steam)
splitting (26) of fats & oils, producing fatty acids (28) and
glycerol (27). The quality of the soaps can be further improved by
hydrogenation of double bonds in the acyl-moiety or before
saponification or hydrolysis, the fats & oils can be
hydrogenated to remove the remaining double bonds and subsequently
sent to the saponification (21) or hydrolysis (26) step. The soaps
can be sent (31) to the decarboxylation section where they are
converted into bio-naphtha (35) and metal-carbonates or CO.sub.2
(36). This bio-naphtha (41) is sent to the to steamcracking (50) or
blended with fossil LPG, naphtha or gasoil (40) and hence the blend
is streamcracked (50). The products of the steamcracking are
cooled, compressed, fractionated and purified (51). This results in
light olefins (ethylene, propylene and butenes), dienes (butadiene,
isoprene, (di)cyclopentadiene and piperylenes), aromatics (benzene,
toluene and mixed xylenes) and gasoline as main components.
EXAMPLES
Example 1
[0113] Hydrodeoxygenation of a triglyceride feed has been evaluated
under the following conditions:
[0114] In an isothermal reactor, 10 ml of a hydrotreating catalyst
composed of Molybdenum and Nickel supported on alumina (KF848
obtained from Albemarle) was loaded, the catalyst dried and
pre-sulfurised under standard conditions with straightrun gasoil
doped with DMDS.
[0115] The hydrodeoxygenation of rapeseed is done at:
[0116] LHSV=1 h.sup.-1
[0117] Inlet Temperature=320.degree. C.
[0118] Outlet pressure=60 bars
[0119] H2/oil ratio=630 Nl/l
[0120] Feedstock=rapeseed doped with 1 wt % DMDS
[0121] Table 4 shows a typical composition of the rapeseed oil.
[0122] The gas and liquid effluent are separated by means of a
separator (gas/liquid) at atmospheric pressure. Gases are sent to a
.mu.-GC analyser and liquids are sent to a sampler. The mass
balance is around 99% and all product weights are calculated for
100 g of treated feed.
TABLE-US-00004 TABLE 4 Typical composition of rapeseed oil
Components wt % tetradecanoate 0.1 hexadecenoate 0.2 hexadecanoate
4.8 heptadecanoate 0.1 octadecadienoate 20.6 octadecenoate 61.3
octadecatrienoate 8.6 octadecanoate 1.8 eisosenoate 1.2 eicosanoate
0.7 docosenoate 0.3 docosanoate 0.3 100
[0123] The total liquid effluent is biphasic and need a separation
step. The organic phase was analyzed via GC-MS. A complete analysis
is reported in table 5.
[0124] The liquid effluent is composed of 94.4 wt % of n-paraffins
but it is composed of 99.94 wt % of interesting components, which
could be sent to the naphtha-cracker.
TABLE-US-00005 TABLE 5 Material balance and complete GC analysis of
hydrocarbon phase Feed Products 5.96 gr hydrogen 6.48 gr CO2 100 gr
rapeseed 0.55 gr CO 3.52 H2 5.98 gr propane 0.18 gr methane 2.77 gr
water phase 85 gr hydrocarbon phase Hydrocarbon phase composition
Wt % C3 0.005 n-paraffin's with C.sub.5 to C.sub.14 0.268 other
paraffin's with C.sub.5 to C.sub.14 0.238 other C15 0.061 n-C15
2.353 other C16 0.100 n-C16 2.754 other C17 1.633 n-C17 41.077
other C18 2.108 n-C18 44.344 dodecyl-cyclohexane 0.168
tridecyl-cyclopentane 0.110 n-paraffin's with C.sub.19 to C.sub.35
3.599 other paraffin's with C.sub.19 to C.sub.35 1.1 >n-C35
0.013 2-butanone 0.034 Other oxygenates 0.025 Total 100.00
[0125] 94.4 wt % of the hydrocarbon phase are comprised of
n-paraffin's that is high quality bio-naphtha feedstock for a
steameracker. About 0.059 wt % of remaining oxygenates are found in
the hydrocarbon phase. That corresponds to 112 wppm O-atoms.
Considering the O content in the triglyceride feed, that represents
10.86 wt % (or 108600 wppm O-atoms), resulting in a
hydrodeoxygenation conversion of 99.89%.
Example 2
[0126] n-Paraffin's and conventional naphtha have been steamcracked
under different severity conditions. Table 6 gives the results. It
is evident from the results that such-obtained bio-naphtha are
better feedstock for steamcracking compared to fossil naphtha.
[0127] Significant higher ethylene and propylene yields can be
obtained whereas the methane make and the pyrolysis gasoline make
is reduced with at least about 20%. The ultimate yield of HVC (High
value Chemicals=H2+ethylene+propylene+butadiene+benzene) is above
70 wt %. Ethylene/Methane weight ratio is always above 3.
TABLE-US-00006 TABLE 6 Naphtha n-Decane n-C15 n-C20 Naphtha
n-Decane n-C15 n-C20 P/E 0.59 0.44 0.50 0.49 0.50 0.39 0.44 0.44
COT 812 812 812 812 832 832 832 832 S/HC 0.35 0.35 0.35 0.35 0.35
0.35 0.35 0.35 Summary wt % (dry) wt % (dry) wt % (dry) wt % (dry)
wt % (dry) wt % (dry) wt % (dry) wt % (dry) Hydrogen 0.87 0.66 0.59
0.57 0.96 0.76 0.69 0.67 Methane 14.79 11.67 10.65 10.00 16.25
12.80 11.80 11.15 Acetylene 0.25 0.25 0.25 0.25 0.36 0.37 0.37 0.37
Ethylene 25.39 38.87 36.24 35.82 26.91 39.67 36.93 36.47 Ethane
4.09 6.58 6.07 5.84 3.89 6.10 5.62 5.42 Methyl-Acetylene 0.29 0.21
0.22 0.22 0.36 0.26 0.27 0.27 Propadiene 0.21 0.15 0.16 0.16 0.25
0.18 0.19 0.19 Propylene 15.10 17.29 18.08 17.63 13.48 15.59 16.28
15.91 Propane 0.51 0.73 0.69 0.66 0.44 0.62 0.59 0.57
Vinyl-Acetylene 0.04 0.04 0.04 0.04 0.05 0.06 0.07 0.07 Butadiene
4.61 5.96 6.88 7.30 4.41 5.79 6.49 6.79 Butene (sum) 4.86 2.99 3.34
3.43 3.67 2.12 2.34 2.38 Butane (sum) 0.08 0.14 0.12 0.12 0.06 0.11
0.09 0.09 Total C5-C9's 23.69 12.48 14.65 15.75 22.30 13.14 15.33
16.42 Total C10+ 5.17 1.93 1.96 2.15 6.53 2.38 2.86 3.18 Carbon
Oxide 0.05 0.05 0.05 0.05 0.07 0.07 0.07 0.07 Carbon Dioxide 0.00
0.00 0.00 0.00 0.01 0.00 0.00 0.00 Ultimate Ethylene 28.67 44.14
41.09 40.49 30.02 44.55 41.43 40.80 C2= + C3= 43.77 61.43 59.17
58.12 43.51 60.14 57.71 56.70 BENZENE 8.27 5.35 6.46 7.05 9.42 6.55
7.77 8.39 HVC's 54.25 68.14 68.24 68.37 55.18 68.35 68.16 68.23
Ultimate HVC's 57.52 73.40 73.10 73.04 58.29 73.23 72.66 72.56
Naphtha composition wt % Normal paraffins 31.26 Iso paraffins 33.48
Naphtenics 28.1 Aromatics 7.16 Olefins 0 Others 0 P/E is the
propylene/ethylene ratio COT is the coil outlet temperature S/HC is
the ratio steam/hydrocarbon
* * * * *