Note: Descriptions are shown in the official language in which they were submitted.
CA 02543367 2011-08-16
METHOD AND APPARATUS FOR CONVERTING AND REMOVING
ORGANOSULFUR AND OTHER OXIDIZABLE COMPOUNDS FROM DISTILLATE
FUELS, AND COMPOSITIONS OBTAINED THEREBY
BACKGROUND OF THE INVENTION
Field of the Invention
One aspect of the present invention is directed to a process for reducing the
concentration of organosulfur compounds in any hydrocarbon-based fluid and a
multi-stage
system for conducting the same.
Discussion of the Background
Natural fuel stock comprises hydrocarbons and other undesirable components,
such as
organosulfur compounds. These organosulfur compounds include, but are not
limited to.
thiophenes, benzothiophenes, dibenz.othiophenes, naphthothiophenes
naphthobenzothiophenes and their substituted analogs. When combusted, these
organosulfur
compounds produce undesirable sulfur pollutants that have been generally
attributed to
societal problems such as respiratory illnesses, acid rain, etc. The sulfur
pollutants also
poison tail pipe catalytic converters. The catalytic converters are designed
to decrease other
diesel engine pollutants such as particulate matter, oxides of nitrogen and
uncombusted or
partially combusted hydrocarbons. Consequently. technologies have been
implemented in
order to remove organosulfur compounds from natural fuel stock.
At present, hydrodesulfurization (HDS) is the most commonly employed
technology
used to desulfurize natural fuel stock, said technology being capable of
reducing the amount
of sulfur to levels of about 300 to 500 ppmw (parts-per-million by weight).
However, some
of the above-mentioned organosulfur compounds are difficult to desulfurize via
HDS because
they are sterically hindered. This is especially true for the 4 or 6-mono- or
4,6-di-alkyl-
CA 02543367 2011-08-16
substituted dibenzothiophenes. Recently, newer HDS technology has been
introduced that is
capable of desulfurizing the "difficult to desulfurize" (or hard sulfur)
compounds;
consequently, this technology affords refineries with the opportunity to
reduce the sulfur
levels even further. However, this newer HDS technology requires more
demanding
desulfurization conditions, such as higher temperatures (> 650 F (343 C)) and
pressures (>
1000 prig (68.9 bars)), and reduced space velocities. Under these conditions,
unnecessary
side reactions (e.g., hydrogenation of unsaturated carbon-carbon bonds) become
kinetically
viable with respect to the sulfur-reduction reaction. Accordingly, large
amounts of hydrogen
are required for adequate desulfurization, which in turn, results in an
overall increase in
operating and capital costs. This last matter is due, in part, to the fact
that in order to operate
at the higher temperatures and pressures, a refinery must equip itself with
specialized reactors
and equipment. Therefore, this newer HDS technology is somewhat cost and space
prohibitive, and thus, may not be an economical alternative for many
refineries.
Regardless of the economics associated with this newer HDS technology, all
refineries are now facing newly promulgated governmental regulations that
limit the sulfur
content of fuels. Specifically, the United States Environmental Protection
Agency (US EPA)
will soon limit sulfur content of "on-road" diesel fuel to 15 ppmw. As noted
above, this
presents a problem for many refineries because the only available technology
capable of
producing "on-road" diesel fuel that meets this newly imposed requirement is
economically
unattractive.
Consequently, the newly introduced stringent regulations coupled with the
shortcomings of existing HDS technology have necessitated a search for
technologies that
may either supplant or complement the existing HDS technology.
Ideally, it would be convenient if the organosulfur compounds could be
separated
from the hydrocarbon liquid by distillation. Unfortunately. this is not
possible, as the
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CA 02543367 2011-08-16
physical properties of organosulfur compounds found in hydrocarbon fuels are
often very
similar to the fuel itself. For example, middle distillate fuels such as
atmospheric or vacuum
gas oils are produced via distillation. The organosulfur compounds that are
contained in
these gas oils have the same boiling range as the fuel itself. In fact,
organosulfur compounds
are found throughout the boiling range of the fuel. Therefore separation of
the organosulfur
compounds by distillation is not possible. However, an attractive avenue of
exploration is
one directed to a chemical process whereby organosulfur compounds are
converted to altered
organosulfur compounds whose physical properties are significantly different
than those of
the starting organosulfur compounds, and thus, from the overall hydrocarbon
liquid.
One possible approach that has recently received attention involves oxidative
desulfurization. Oxidative desulfurization operates at mild temperatures (<
212 F (100 C))
and pressures (< 30 psig (2.07 barg)). and several patents have been granted
describing
oxidative desulfurization processes. Some earlier U.S. Patents (2,749,284;
3,341,448;
3.413.307), describe two common themes of oxidative desulfurization, which
include, but are
not limited to, reaction of a fuel stock containing organosulfur compounds
with an oxidant
followed by separation. Other references (US Patents 5,753,102; 5,824207;
5,910,440;
5,958,224; 5.961,820; 6,160,193, 6,171,478, 6,231,755; 6.254,766; 6,274.785;
6,277,271;
6,338,794; 6.402,940; 6,402,939; and 6,406,616; and US Statutory Invention
Registration
H 1986), encompass the earlier developed themes of oxidization of unwanted
organosulfur
compounds present in hydrocarbon liquids followed by separation of the
oxidized
organosulfur compounds from the desired hydrocarbon liquid. On the whole,
these
references represent the conventional processes for reducing unwanted
organosulfur
compounds from fuel stocks; all of which involve an oxidation reaction,
wherein
organosulfur compounds are converted to their respective sulfoxides and
sulfones, followed
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by one or more separation steps. The separation steps include, but are not
limited to,
extraction and adsorption (either alone or in combination).
The themes associated with oxidative desulfurization contained in many of
these
references shows that when organosulfur compounds are oxidized, the resultant
oxidized
organosulfur compounds have significantly different physical properties that
provide an
opportunity for separating the oxidized organosulfur compounds from the
hydrocarbon liquid.
For example, when the sulfur-containing compounds contain thiophenic sulfur.
the oxidized
organosulfur compounds comprise corresponding thiophenic sulfoxides or
sulfones whose
physical properties (e.g., polarity and volatility) are significantly
different than those of the
unoxidized thiophenic compounds. These differences in the physical properties
enable the
separation of oxidized organosulfur compounds from the hydrocarbon fuel.
Separation
techniques can rely on many physical properties, and the two mentioned
properties (e.g.,
polarity and volatility) are not exhaustive but are mentioned for illustrative
purposes.
Even though the above-identified references are directed to the problem of
removing
unwanted organosulfur compounds from fuel stocks, these references do not
adequately
describe a process that may be adapted for use in middle distillate fuel
stocks that contain
about 5000 ppmw or more of organosulfur compounds. The reason for this lies in
the overall
conversion of the oxidation reaction. For example, in order to satisfy the US
EPA standard of
15 ppmw, a process that includes the oxidation reaction must be able to
consistently operate
at a reaction conversion of no lower than about 99.4%, when the organosulfur
content is
about 5000 ppmw. Ideally, it is desirable to develop a substantially
quantitative oxidative
process, in order to remove substantially all of the sulfur-containing
hydrocarbons from a
middle distillate fuel stock.
Accordingly, a problem to be solved by the present invention relates to a
process
wherein the conversion of unoxidized organosulfur compounds to oxidized
organosulfur
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CA 02543367 2011-08-16
compounds occurs substantially quantitatively. Substantially quantitative
oxidation
simultaneously allows for efficient separation and removal of organosulfur
compounds and
further recovery of hydrocarbon fuel.
This problem becomes apparent when one considers that efficiency of the above-
mentioned separation processes (i.e., extraction and adsorption) is dependent
upon the overall
oxidation conversion process. For example, when processing fuels with
approximately 5000-
ppmw sulfur content, it has been found that it is advantageous to remove most
of the sulfur
compounds utilizing a liquid-liquid extraction process. However, an extraction
step that
involves high sulfur removal leads to high solvent to feed ratios. While
recovery of the
solvent extract after the liquid-liquid extraction does not pose major
difficulties, the resultant
extract is not only rich in oxidized organosulfur compounds, but also contains
sulfur-free fuel
components, particularly aromatic compounds. The quantity of fuel lost via the
liquid-liquid
extraction step may range from 20 to 35 wt %. which leads to another problem
to be solved.
That is, liquid-liquid extraction of an oxidized fuel stock leads to a
concomitant loss of fuel.
If the overall conversion of the oxidation is not substantially quantitative,
then it becomes
difficult to recover lost fuel. While it may be possible to further process
the solvent extract
stream in other refinery units or to burn the solvent extract stream for its
energy value or use
the solvent extract stream as an asphalt modifier, the inventors found that
downgrading the
solvent extract stream, i.e., as feed to another refinery processing unit, is
not economically
advantageous.
Accordingly, the present invention provides a solution aimed at overcoming
these
difficulties. by in turn providing a new process that is attractive in that it
overcomes a
problem of fuel loss upon liquid-liquid extraction. It is noted that minimized
fuel loss is
made possible by achieving substantially quantitative oxidative conversion
during the
oxidation stage of the overall process. Consequently, the solvent extract that
contains fuel
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CA 02543367 2011-08-16
may be subjected to additional process steps that a6ord the recovery of fuel
via distillation.
This provides a higher overall recovered yield of fuel that has heretofore
never been
accomplished, as other oxidative processes cannot simultaneously achieve the
low sulfur fuel
yields made possible by the present invention.
In addition to the advantages inhered by the substantially quantitative
oxidative
conversion process. the present invention inheres additional advantages over
pre-extraction
type processes, such as those described, for example, by Gore in U.S. Patent
Nos. 6,160,193
and 6,274.785. For example, these advantages include: (1) Favors fuel recovery
over
minimizing oxidant consumption; (2) Minimizes the circulation of extraction
solvent; (3)
Eliminates the need for an extract wash step; and (4) Minimizes corrosive
catalytic acids in
downstream lines and equipment.
SUMMARY OF THE INVENTION
Accordingly. a solution to the problems presented by the above-identified
government
mandate is found in a process which comprises contacting a first liquid
comprising at least
one hydrocarbon compound with a first oxidant in a first reactor and
contacting a second
liquid comprising at least one hydrocarbon obtained from the first reactor
with a second
oxidant in a second reactor.
In this process the first liquid may be any hydrocarbon-based fluid. Both
oxidants
comprise a percarboxylic acid that is obtained by reacting carboxylic acid
with hydrogen
peroxide. The second liquid is obtained directly or indirectly from the first
reactor. For the
purpose of this disclosure, when the second liquid is obtained directly from
the first reactor,
the second liquid comprises a first reactor effluent (or first effluent). When
the second liquid
is obtained indirectly from the first reactor, the second liquid is obtained
by separating the
first effluent into two phases in a first vessel, i.e., a first light phase
comprising at least one
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CA 02543367 2011-08-16
hydrocarbon compound and a first heavy phase comprising a polar solvent;
wherein said
polar solvent comprises a carboxylic acid.
As noted above, the first liquid may be any hydrocarbon-based fluid, which may
be a
crude gas oil, a distillate of crude oil, a middle distillate comprising
hydrocarbons having
boiling points that range from 65 C to 385 C, or a crude gas oil obtained by a
hydrodesulfurization process. An attractive feature of the disclosed invention
is that the
process may be employed either prior or subsequent to an HDS process.
A key feature of said process is that the overall oxidation is achieved by
employing a
counter-current oxidation scheme. That is, the first liquid that makes contact
with the first
oxidant has a higher unoxidized sulfur content than the second liquid that
makes contact with
the second oxidant; which means that the total oxidant concentration in the
first oxidant may
equal to or lower than the total oxidant concentration in the second oxidant.
Stated in another
way. the ratio of the total oxidant concentration in the first oxidant,
[Oxt_1], to the total
oxidant concentration in the second oxidant, [Ox1,2], is less than or equal to
1, i.e.,
[Oxt,1]/[Oxt,2] < 1. In the practice of the invention, the ratio
[Ox1,l]/[Oxr,2] may range from
0.0001 to 1, preferably from 0.001 to 1, more preferably from 0.01 to 1, and
most preferably
from 0.1 to 1.
Not to he limited by theory, but application of the counter-current oxidation
scheme
may be explained in terms of the kinetics of oxidation. When the unoxidized
sulfur content is
high, then oxidant concentration need not be too high, in order to achieve an
acceptable
conversion rate. However, when the unoxidized sulfur content is lower, then
the oxidant
concentration becomes more relevant. Accordingly, the total oxidant
concentration in the
second oxidant will be higher than that of the total oxidant concentration in
the first oxidant,
as the unoxidized sulfur content of the second liquid is lower than that of
the first liquid.
These and other aspects will be explained in more detail below.
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CA 02543367 2011-08-16
While the U.S. EPA mandate is concerned with decreasing the concentration of
organosulfur compound in "on-road" diesel fuel, it is conceivable that the
disclosed process
would be applicable for decreasing the concentration of organo-nitrogen
compounds that are
present in any hydrocarbon-based fluid. Moreover, an attractive feature of the
present
invention is that it is capable of improving the storage stability of a
product gas oil obtained
by the disclosed process.
Additionally, another aspect of the present invention is achieved by a multi-
stage
system capable of reducing organosulfur compounds iin a liquid, comprising an
oxidation
stage; an extraction stage; a raffinate washing stage; a raffinate polishing
stage; a solvent
recovery stage; a solvent purification stage; and a hydrocarbon recovery
stage. A more
detailed description of the process appears below.
BRIEF DESCRIPTION OF DRAWINGS
Figure 1 A is generalized block flow diagram representing the disclosed
process.
Figure 1 B is a block flow diagram representing the seven major unit
operations of the
disclosed reactor and process.
Figure 2 is plot of Temperature ( F) versus Distillate Collected (Volume
Percentage)
of comparative distillation curves.
Figure 3 is a specific process flow diagram of the Oxidation portion of the
disclosed
process.
Figure 4 is a specific process flow diagram of the Sulfox Extraction and
Raffinate
Washing portion of the disclosed process.
Figure 5 is a specific process flow diagram of the Raffinate Polishing portion
of the
disclosed process.
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CA 02543367 2011-08-16
Figure 6 is a specific process flow diagram of the Solvent Recovery and
Solvent
Purification portion of the disclosed process.
Figure 7 is a specific process flow diagram of the Hydrocarbon Recovery
portion of
the disclosed process.
Figure 8 is a specific process flow diagram of an Improved Oxidation portion
of the
disclosed process.
DETAILED DESCRIPTION OF THE INVENTION
A schematic block flow diagram showing one preferred embodiment of the
invention
is given in Fig. I A, attached, and described in more detail below.
The invention process is particularly suitable to treat middle distillate
fuels that
contain a broad array of sulfur compounds. The sulfur compounds may be present
in per cent
level concentrations. The oxidant is a peroxycarboxylic acid. The inventors
found that the
carboxylic acid used to form the peroxycarboxylic acid is optimally used as
the solvent. If a
different solvent is chosen. then two separate "Solvent Recovery and
Purification" steps and
two separate "Hydrocarbon Recovery" steps would be needed.
1. Reactor System
The first step in the process is to combine the oxidant solution in Stream A,
the high
sulfur feed in Stream B and the carboxylic acid or an aqueous solution of the
carboxylic acid
in Stream DI in the "Reactor System". In this step, the organosulfur compounds
in the fuel
are converted to sulfoxides or sulfones. Excess water from the reactor system,
Stream C, is
directed to the "Solvent Recovery and Purification" step. The light phase
leaves the "Reactor
System" via Stream E. If the reactor conditions are chosen so that only one
phase forms then
the entire contents of the "Reactor System" leaves via Stream E.
2. Extraction
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The next step in the process is the "Extraction". The Extraction may be
carried out in
any suitable liquid/liquid-contacting device. The fuel containing oxidized
sulfur compounds
in Stream F is contacted with the solvent in Stream D2. The more polar
sulfoxides and
sulfones leave the "Extraction" step together with the solvent in Stream F.
The raffinate
leaves the "Extraction" step via Stream H. Stream H comprises fuel with less
sulfur
compounds and some solvent.
3. Water Wash
The next step in the process is a "Water Wash." The purpose of this step is to
remove
residual solvent from the fuel. This step is accomplished by contacting the
fuel with water in
any suitable liquid/liquid-contacting device. Fuel enters this step via Stream
H and Stream 0.
Water enters via Stream G. The heavy phase leaves via Stream 1. Stream I
comprises water
and solvent. Stream 1 is directed to the "Solvent Recovery and Purification"
step. The fuel,
substantially free of solvent, leaves via Stream J.
4. Adsorption
The next step in the process is "Adsorption". This step may or may not be
needed
depending on the sulfur concentration remaining after extraction. The purpose
of the
"Adsorption" step is to remove the last traces of sulfur from the fuel. The
fuel enters via
Stream J and exits this step via Stream K. A number of solids have been found
to be suitable
for this step of the process that include, but are not limited to, refiner's
clay. The regeneration
of the adsorbent may be carried out in several ways. These methods involve the
use of a
carrier fluid and changes in temperature, pressure, or concentration. These
changes alter the
equilibrium, and favor desorption of the adsorbed substance. If the extraction
solvent is used
for the regeneration, then the resultant stream may be directed to the
"Solvent Recovery and
Purification" step.
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Bed regeneration may be accomplished using the extract solvent and subsequent
recycling to the front end of the process.
5. Solvent Recovery and Purification
The next step in the process is "Solvent Recovery and Purification". The
purpose of
this step is to recover and re-use the carboxylic acid that is used as the
solvent and the
precursor for the peroxycarboxylic acid. The additional capital and operating
expense of this
step is less than the cost of purchasing fresh solvent. The "Solvent Recovery
and Purification"
step includes various unit operations, such as distillation and flash
evaporation, designed to
separate solvent from water or solvent from extract.
Solvent enters. this step primarily via Stream F, Stream C, if present, and
possibly via
a regeneration step associated with the "Adsorption" step. Recovered solvent
leaves via
Stream D and is directed to the unit operations requiring solvent. Fresh
solvent may be added
to this stream or at other convenient points in the process to make up for
losses.
Water with some solvent enters the "Solvent Recovery and Purification" step
via
Stream I. Water enters the process in Stream A and Stream G. Some water is
also formed
during the transformation of the carboxylic acid to the peroxycarboxylic acid
using hydrogen
peroxide. For example acetic acid, when reacted with hydrogen peroxide, is
transformed to
peracetic acid (PAA) with the concomitant formation of water.
Hydrogen peroxide is commercially available as aqueous solutions. For these
reasons
water must be purged from the system via Stream M to prevent an accumulation
of water.
Some water may be recycled via Stream L. A small hydrocarbon phase may be
generated
during solvent recovery and purification. This stream may be processed through
the "Water
Wash" to improve yield.
6. Hydrocarbon Recovery
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The next step of the process is "Hydrocarbon Recovery". Material is fed to
this step
via Stream N. Stream N is the extract (Stream F) with the solvent removed.
Stream N
contains the oxidized organosulfur compounds (sulfoxides and sulfones) and
fuel components,
and residual acetic acid. The fuel components are primarily the more polar
aromatic
compounds that boil in the diesel range. The "Hydrocarbon Recovery" step
utilizes the
volatility difference between the sulfoxides and sulfones and the aromatic
fuel compounds.
The inventors found that the boiling points of the oxidized sulfur compounds
are beyond
most of the compounds normally found in diesel. Distillation, vacuum
distillation in
particular, is a suitable unit operation for separating the fuel components
from the oxidized
sulfur compounds. The recovered fuel components are returned to the process
via Stream O.
The final extract leaves the process via Stream P.
One advantage of the present invention is realized by taking advantage of many
of the
physical property differences that are imparted to the organosulfur compounds
once they are
converted to their respective sulfoxides or sulfones. The instant invention is
economically
favorable for the removal of undesired components and maximizes the fuel yield
across the
process.
As noted above, the disclosed reactor process is made surprisingly superior,
and
consequently, economically feasible by attaining significant hydrocarbon
recovery via
distillation. This is especially true if the oxidation step is capable of
substantially complete
conversion of the organosulfur compounds to their respective polar
organosulfur compounds.
In their unoxidized form, the organosulfur compounds have the same boiling
range as the rest
of the hydrocarbons found in the distillate stream. If left unoxidized, these
organosulfur
compounds distill simultaneously with the hydrocarbons rendering distillation
ineffective as a
method to minimize yield loss. Once oxidized, the boiling points of these
compounds are
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shifted significantly higher. This increase in the boiling points allows
distillation to become a
feasible method of hydrocarbon recovery.
As noted above. the multi-stage system and process is based on a middle
distillate
considered as Light Atmospheric Gas Oil (LAGO). This middle distillate
comprises aliphatic,
cycloaliphatic (or naphthenic). olefinic, aromatic, and heteroatom-containing
derivatives
thereof. For the purpose of this disclosure, the middle distillate is that
portion of crude oil
that distills from about 150 F (65.6 C) to about 800 F (385 C). Furthermore,
in addition to
"on-road" diesel, it believed that the disclose process is capable of
producing "off-road" and
"marine" diesel having reduced sulfur content. Additionally, it is believed
that the process
disclosed herein is capable of reducing sulfur content in the following
feedstocks: Middle
Distillates; Gas Oils; Atmospheric Gas Oils; Light Atmospheric Gas Oils;
Distillate Fuel Oils;
Kerosine; Diesel Fuel; Jet Fuel; Home Heating Oil; Solvents; Hydrotreated
Middle Distillates;
Hydrotreated Gas Oils; Hydrotreated Atmospheric Gas Oils; Hydrotreated Light
Atmospheric
Gas Oils; Kerosine (ASTM D-3699); Kerosine (No. 1-K) (ASTM D-3699); Kerosine
(No. 2-
K) (ASTM D-3699); Civil Aviation Turbine Fuels (ASTM D-1655); Jet A-I Civil
Aviation
Turbine Fuel (ASTM D-1655); Jet A Civil Aviation Turbine Fuel (ASTM D-1655);
Military
Aviation Turbine Fuels; JP-5 Military Aviation Turbine Fuel: JP-8 Military
Aviation Turbine
Fuel; Diesel Fuel Oils (ASTM D-975): Diesel Fuel C)il Grade No. 1-D S500 (ASTM
D-975);
Diesel Fuel Oil Grade No. I-D S5000 (ASTM D-975); Diesel Fuel Oil Grade No. 2-
D S500
(ASTM D-975); Diesel Fuel Oil Grade No. 2-D S5000 (ASTM D-975); Diesel Fuel
Oil
Grade No. 4-D (ASTM D-975); Fuel Oils (ASTM D-396); Grade I Fuel Oil (ASTM D-
396);
Grade I Low Sulfur Fuel Oil (ASTM D-396); Grade 2 Fuel Oil (ASTM D-396); Grade
2
Low Sulfur Fuel Oil (ASTM D-396); Grade 4 Light Fuel Oil (ASTM D-396); Grade 4
Fuel
Oil (ASTM D-396); Marine Distillate Fuels; Grade DMX Marine Distillate Fuel;
Grade
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DMA Marine Distillate Fuel; Grade DMB Marine Distillate Fuel; Grade DMC Marine
Distillate Fuel.
In essence, the reaction chemistry changes the physical properties (i.e.,
polarity and
volatility) of the organosulfur compounds contained in LAGO. The process then
takes
advantage of these changes in the physical properties to separate the oxidized
organosulfur
compounds from the balance of the hydrocarbon fuel.
As highlighted below, the disclosed process is illustrated based on a
simulated gas oil
feed that comprises about 5 100 ppm of sulfur by weight. However, it is
possible to apply the
same process to other middle distillate feeds with a lower or higher sulfur
content. for
example, from _S to 100,000 ppm, which includes 5; 10; 50; 100; 500; 1000;
.2000; 3000;
4000; 5000, 6000; 7000; 8000, 9000; 10.000, 20,000; 50,000; 75.000; 100,000;
ppm by
weight and any combination thereof. In the case of hydrotreated middle
distillates (i.e., HDS-
treated middle distillates), the invention is expected to perform both
technically and
economically better than the specific example described herein. The process is
also suitable
for treating other middle distillates, since the overall concept clearly
applies.
In the case of hydrotreated middle distillates. where the overall sulfur
content is
typically below 500 pprnw, the multi-stage process is expected to perform both
technically
and economically better than the specific example described herein,.
Hydrotreated middle
distillates typically lack the lower molecular weight thiophenic compounds and
are rich in
higher molecular weight highly substituted dibenzothiophenes (i.e., the hard
sulfur
compounds). As mentioned previously, these higher molecular weight highly
substituted
dibenzothiophenes are easier to oxidize via the disclosed oxidation process
with respect to a
HDS process. For this reason, as well as the lower total sulfur content of the
feed, an overall
decrease in the consumption of oxidant is expected. In addition, it may be
possible to achieve
total oxidation with a simpler oxidation system. For example, employing a
hydrocarbon-
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CA 02543367 2011-08-16
based liquid obtained by an HDS process in which the organosulfur
concentration has been
substantially reduced. It may he possible to achieve total oxidation in a
single reactor,
wherein said reactor is a plug-flow reactor or series of continuous stirred-
reactors. Once
oxidized. these higher molecular weight highly substituted dibenzothiophenes
will have very
high boiling points. Therefore, the ease of hydrocarbon recovery should
increase, thereby
allowing an improvement in the overall process yield. Potential yields of
greater than 98
percent may be possible, which is more than adequate when one considers that
the starting
sulfur content is about 500 ppmw.
A better understanding of the overall disclosed process may be gleaned upon
reading
the following text in view of Fig. 1 B. A more detailed discussion of a
preferred embodiment
of the disclosed invention is presented below.
It should be apparent upon inspection of Fig. I B that there are, preferably,
seven
major unit operations in the invention process: (1) Oxidation, (2) Sulfox
Extraction, (3)
Raffinate Washing. (4) Raffinate Polishing, (5) Solvent Flash / Solvent
Recovery, (6) Solvent
Purification, and (7) Hydrocarbon Recovery.
In the Oxidation System, the thiophenic compounds in fuel (gas oil) are
ultimately
oxidized to sulfones. The oxidation is accomplished with hydrogen peroxide in
the presence
of recycled carboxylic acid (CA). It should be clear that the requisite
overall molar
conversion of the oxidation process is, of course, dependent upon the amount
of unoxidized
organosulfur compounds in the feed stock. However, the overall molar
conversion of the
unoxidized organosulfur to the oxidized organosulfur compounds is about 99.4
percent,
preferably 99.6 percent, more preferably 99.7 percent, and most preferably
99.8 percent;
wherein for every mole of sulfur present in the feed, about 2.5 to 5.0 moles
of oxidant,
preferably 3 moles of oxidant are required. This amount of oxidant is 50
percent more than
the stoichiometric requirement necessary for complete conversion to the
sulfone. The water
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formed by the reaction of acetic acid and hydrogen peroxide and the water that
enters the
oxidation system with the hydrogen peroxide are separated from the oxidized
gas oil and fed
to Solvent Purification for recovery of CA and purging of reaction water. The
oxidized gas
oil that is now saturated with CA is fed to the Sulfox Extraction System.
The solution chemistry that may occur in the Oxidation portion of the reactor
process
is outlined as follows.
There are many organosulfur compounds in straight run LAGO. Typically, these
organosulfur compounds have a fairly high molecular weight and belong to a
general class of
compounds called thiophenics. In most cases, these compounds are
benzothiophene,
naphthothiophene, dibenzothiophene, naphthobenzothiophene, and their
substituted
homologues. Their respective molecular structures are shown below.
~/ S \ S
R1 S
S RI R2 Rl~~ R2
R 1 ~~
These organosulfur compounds are oxidized to sulfoxides and subsequently
sulfones
via reactions with active oxygen in the form of percarboxlic acid. In the
invention process,
the reactions are typically conducted at moderate temperatures (50 F (10 C) to
250 F
(121 C), which includes 50, 75, 100, 115, 120, 122. 125, 135. 145, 155, 165,
175, 185. 195.
200, 205, 210, 212, 214, 220. 250 F, and any combination thereof) and at or
about
atmospheric pressure. In this temperature range, the reaction mixture
preferably includes two
liquid phases. The oxidation reactions could be conducted in a single-phase
mixture by
utilizing a higher temperature.
In the present application, the R, R1 and R2 groups may each independently be
any
linear or branched, cyclic or aliphatic, substituted or unsubstituted C,-C20
alkyl group,
substituted or unsubstituted C7-C3() aryl group, C7-C;o arylalkyl group, and
combinations
-16-
CA 02543367 2011-08-16
thereof. This includes those having 1, 2, 3, 4, 5. 6, 7. 8, 9, 10, 11, 12, 13,
14, 15. 16, 17, 18,
19, 20, 21, 22, 23, 24. 25, 26, 27, 28, 29, and 30 carbons, and any
combination thereof.
The heavy phase contains carboxylic acid, hydrogen peroxide, percarboxylic
acid,
water. sulfuric acid, soluble hydrocarbons, and soluble thiophenic compounds.
The dominant
species in the heavy phase is the carboxylic acid, which is a carboxylic acid
is represented by
the formula RCOOH, wherein R is an radical selected from the group consisting
of H, methyl,
ethyl. n-propyl, i-propyl, n-butyl, i-butyl, s-butyl, n-pentyl, i-pentyl, and
s-pentyl. Though
not to be limiting, the carboxylic acid that may be employed is selected from
the group
consisting of formic acid, acetic acid, propionic acid, butyric acid,
pentanoic acid, hexanoic
acid, and mixtures thereof; preferably the carboxylic acid is selected from
the group
consisting of formic acid, acetic acid, propionic acid, and mixtures thereof;
and more
preferably the carboxylic acid is selected from the group consisting of formic
acid, acetic acid,
and mixtures thereof; and most preferably the carboxylic acid is acetic acid .
The formation
of PCA primarily occurs in the heavy phase. Once formed, a portion of the PCA
migrates to
the light phase.
The light phase preferably includes mostly hydrocarbons with a significant
amount of
carboxylic acid, and relatively small amount of percarboxylic acid, hydrogen
peroxide, water
and sulfuric acid.
The oxidation of thiophenic compounds to sulfones probably occurs in both the
light
and heavy phases. The formation of sulfones may be very fast in the heavy
phase since the
concentration of PCA may be relatively high. In the light phase, oxidation
rates are slower,
especially as the concentration of unoxidized sulfur-containing compounds
approaches zero.
The reaction paths are quite complex involving both reaction kinetics and mass
transfer effects. Intimate contact between the two liquid phases in the
reaction mixture is
preferred for obtaining a sufficient rate of transfer of the PCA between the
two phases.
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CA 02543367 2011-08-16
Percarboxylic Acid (PCA) Formation (Equation 1)
Percarboxylic acid (PCA) is formed via an equilibrium reaction between
hydrogen
peroxide and carboxylic acid (CA); wherein R is selected from the group
consisting of H,
methyl, ethyl, n-propyl, 1-propyl, n-butyl, i-butyl, s-butyl, n-pentyl, i-
pentyl, and s-pentyl.
O O
R-KOH + H702 RA, O~OH + H2O (1)
CA PCA
In addition to PCA, water is formed as a byproduct. The reaction is slightly
exothermic
liberating approximately 348 calories (1.46 kJ) per g-mole of PCA formed.
At room temperature, without the aid of a catalyst, the reaction may be
extremely
slow and equilibrium concentration may take many hours to achieve. Higher
temperatures
can be utilized to accelerate the reaction rate within limits. Above 194 F (90
C),
decomposition of both the hydrogen peroxide and the resulting PCA begins to
become
significant.
Significant increases in reaction rate without significant losses due to
decomposition
are best achieved by using a catalyst. Typically, a strong acid catalyst may
be utilized. In the
invention process, sulfuric acid may be used to catalyze the formation of PCA.
At the reaction temperatures, hydrogen peroxide, CA, and sulfuric acid
concentrations
used in the invention process, near reaction equilibrium conditions are
achieved within 2 to 5
minutes and approximately 90% of the hydrogen peroxide has been converted to
PCA. The
equilibrium constant for the reaction is approximately 2.2 and may be a weak
function of the
reaction temperature. A large excess of CA may be utilized to favor the
product side of the
equilibrium reaction.
Sulfoxide Formation (Equation 2)
Oxidation of the thiophenic compounds occurs in two reaction steps. In the
first step.
thiophenic compounds react with PCA to form a sulfoxide. CA is generated as a
byproduct.
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CA 02543367 2011-08-16
This reaction is irreversible and highly exothermic. At relatively high
thiophenic
concentrations, this reaction is very fast. The reaction shown below depicts
the oxidation of a
generic dibenzothiophene. Similar reaction stoichiometry occurs for
benzothiophenes,
naphthothiophenes. and naphthobenzothiophenes.
O
11
S1 S
~. + PCA + CA (2)
RI/~ R2 R1 R2
Sulfone Formation (Equation 3)
In the presence of PCA, the sulfoxide, once formed, may be quickly oxidized to
the
sulfone (Eqn. 3). As in the formation of the sulfoxide, the formation of the
sulfone also
results in the production of CA. This reaction is also irreversible, highly
exothermic, and
very fast. The reaction shown below depicts the oxidation of a generic
dibenzothiophene
sulfoxide. Similar reaction stoichiometry occurs for benzothiophene
sulfoxides,
naphthothiophene sulfoxides, and naphthobenzothiophene sulfoxides.
O 0O
S/
\ / + PCA - ~- \ / + CA (3)
The literature on the oxidation of thiophenic compounds utilizing PCA
indicates that the
formation of the sulfoxide is the rate-limiting step when considering the
oxidation only. For
dibenzothiophene. the relative difference in reaction rate of thiophenics with
respect to
sulfoxide is approximately 1.4. Namely, the oxidation rate of dibenzothiophene
sulfoxide to
dibenzothiophene sulfone is 40% greater than the oxidation rate of
dibenzothiophene to
dibenzothiophene sulfoxide. Therefore, once formed, the sulfoxide is quickly
oxidized to the
sulfone.
In the oxidation of the thiophenic compounds contained in LAGO, many reactions
are
occurring in parallel and series. Some thiophenic species are much more
reactive than others.
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CA 02543367 2011-08-16
Laboratory studies on single model compounds indicate that the reactivity of
the thiophenic
compounds increases as the aromatic nature of the compounds increases and as
the aromatic
substitution increases. Namely, benzothiophene is less reactive than
dibenzothiophene,
which in turn is less reactive than naphthobenzothiophene and dibenzothiophene
is less
reactive than methyldibenzothiophene, which in turn is less reactive than
dimethyldibenzothiophene. The nature of these reactivity differences has been
attributed to
electronic density effects surrounding the aromatic sulfur atom. Increased
aromatic character
and aliphatic side chain substitution cause the electron density surrounding
the sulfur atom to
increase. This higher electronic density makes the sulfur atom more prone to
attack by the
PCA molecule.
In a complex mixture like LAGO, this reactivity matrix results in a near
continuous
set of reaction rates. Under these circumstances, the possibility of
minimizing the
consumption of oxidant by selectively oxidizing to the sulfoxide is
essentially futile. Kinetic
studies on systems containing just five thiophenic species clearly indicate
that the partial
oxidation approach results in a marginal benefit.
Since the partial oxidation approach requires sub-stoichiometric quantities of
oxidant
(less than 2 moles of oxidant per mole of sulfur), near complete oxidation of
the organosulfur
compounds in LAGO may be not possible under these circumstances. Without
complete
oxidation, maximizing hydrocarbon yield via distillation may be not possible.
In order to
achieve total oxidation in a reasonable residence time. a sufficient amount of
excess oxidant
is required.
Side Reactions (Equations 4 and 5)
In addition to the primary reactions, several side reactions may be occurring.
Experiments indicate that excess active oxygen above and beyond the
stoichiometric quantity
-20-
CA 02543367 2011-08-16
needed to oxidize all sulfur atoms to their corresponding sulfones does not
remain after
oxidation is complete. The nature of these side reactions remains unknown at
this point.
Although it may be possible to decompose hydrogen peroxide and/or PCA. these
side
reactions do not occur at the normal reaction temperatures anticipated for the
invention
process. If decomposition does occur, one of the byproducts would be oxygen.
Experiments
designed to capture any non-condensable gases formed by decomposition gave
negative
results.
Literature sources indicate that it is possible to oxidize light aromatic
hydrocarbons
with PCA. Typically, the byproducts are phenols, aldehydes and ketones.
Although there is
no definitive proof at this stage of the research effort,, it is believed that
these side reactions
do occur. When the unoxidized sulfur concentration is high, the oxidation of
the sulfur atom
is favored and is significantly faster. As the concentration of unoxidized
sulfur diminishes,
the side reactions become more prevalent, especially at elevated temperatures.
This behavior
results in wasting oxidant in undesirable oxidation reactions. The invention
process accounts
for this undesirable shift in reaction path by carefully controlling the
temperature of the
reaction mixture at several levels. This method allows for the most efficient
use of excess
oxidant.
Olefins present in the gas oil may be oxidized to an epoxide (Eqn. 4).
O
R \~ R2 + PCA + CA (4)
R1 R'
In straight run gas oils, the quantity of olefins is usually very small.
However, if light cycle
oils are blended with the straight run gas oil, significant quantities of
olefins may be present.
The presence of'sulfuric acid in the reaction mixture also creates an
environment for
the possible formation of sulfonates (Eqn. 5). Normally, sulfonations are
conducted at
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CA 02543367 2011-08-16
R1 R'
V / + H2SO4 nZ SO;H + H2O (5)
moderate temperatures (I76 F (80 C) to 320 F (160 C)) with high sulfuric acid
concentrations. In the present oxidation system, the sulfuric acid may be
typically below
10,000 ppm and the temperature may be typically at or below 176 F (80 C).
Therefore, the
extent of sulfonation is believed to be minor. However, at one point in the
present oxidation
system, temperatures are as high as 392 F (200 C) and sulfuric acid
concentrations are
approximately 20,000 ppm. In this environment, the sulfonation reactions may
become more
likely. These sulfonation reactions are most likely to occur in the heavy
phase. Due to the
water content of the heavy phase, most, if not all of the sulfuric acid used
to catalyze the
formation of PCA may be present in the this phase. As mentioned previously,
the heavy
phase contains a significant quantity of CA as well. The presence of CA in the
heavy phase
causes a significant increase in the solubility of both monocyclic and
polycyclic aromatic
compounds.
As noted above. PCA present in the system may be destroyed. The solution
chemistry
associated with this destruction is outlined as follows.
Destruct Reaction (Equations 6 and 7)
After oxidation is complete, the light phase leaving the oxidation system may
still
contain small amounts of excess active oxygen that should be removed. By
elevating the
temperature at specific points, the invention process forces the decomposition
of both
hydrogen peroxide and percarboxylic acid. The reaction stoichiometry for each
of these
decompositions is shown below.
2 1HH2O2 0, + 2 H2O (6)
2 PCA 2 CA + 0, (7)
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CA 02543367 2011-08-16
As noted above. upon exiting the oxidation portion of the reactor process, the
oxidized gas oil comprising polar organosulfur compounds may be saturated with
carboxylic
acid. This oxidized gas oil is then fed to the Sulfox Extraction System.
In the Sulfox Extraction System, the residual PCA in the oxidized gas oil is
first
destroyed by heating to 230 F (l 10 C) for a period of time. At this
temperature, PCA in the
gas oil undergoes decomposition to oxygen and carboxylic acid. The resulting
gas oil is then
fed to an extraction column where most of the oxidized organosulfur compounds
are removed
by contacting with recycled carboxylic acid. The extraction temperature is
about 1 13 F
(45 C). The recycle solvent is mostly CA and contains about 0.6 wt% water and
about 5.4
wt% of hydrocarbon. Given a starting sulfur content of 5 100 ppm,. in the
feed, a sulfur
removal of greater than 99 percent is obtained in this extraction step. The
resulting extract
that contains most of the oxidized organosulfur compounds is fed to the
Solvent Flash /
Solvent Recovery System. The gas oil raffinate that is still saturated with CA
and contains
small amounts oforganosulfur compounds is fed to the Raffinate Wash System.
The gas oil raffinate that exits the Sulfox Extraction portion of the reactor
process
may be saturated with CA and may comprise a small amount of polar
organosulfirr
compounds. Accordingly, a stage in the process designed to remove these
impurities is
denoted as the Raftinate Wash System, and is discussed briefly as follows.
In the Raffinate Wash System. CA is removed from the gas oil by contacting
with
water in a mechanically agitated extraction column. The extraction is
conducted at about
1 13 F (45 C) and the resulting gas oil raftinate contains approximately 5800
ppm by weight
of acetic acid. The extract is fed to the Solvent Purification System for
recovery of the
extracted CA and the purification of the water. The gas oil raffinate is fed
to the Raffinate
Polishing System.
-23-
CA 02543367 2011-08-16
In the Raffinate Polishing System, the remaining organosulfur compounds and CA
are
removed from the raffinate gas oil in a solid bed adsorption column.
Currently, the design of
the adsorption beds is based on refinery clay. The ability of this material to
adsorb sulfones
has been demonstrated in the laboratory. A purpose of this portion of the
invention is to
obtain a product gas oiI which comprises less than 10 ppm by weight sulfur and
essentially no
acetic acid.
The heavy phase extract obtained from the Sulfox Extraction portion of the
process is
transported to the Solvent Flash / Recovery System, in which CA may be removed
from the
extract produced in the Sulfox Extraction System. First, most of the CA may be
removed in a
single stage flash. The resulting extract, comprising approximately 15 wt% CA
is then fed to
a small distillation column. In this column, the CA content of the extract is
reduced to
approximately 2 wt% before being fed to the Hydrocarbon Recovery System. The
recovered
CA from the single stage flash and the distillation column may be combined.
This recovered
CA comprises light hydrocarbons that form minimum boiling homogeneous
azeotropes with
the acetic acid. Most of the recovered CA is recycled to the Oxidation System
and to the
Sulfox Extraction System. However, in order to control the build up of
azeotropic
hydrocarbons in these recycle loops; a portion of the recovered CA is fed
forward to the
Solvent Purification System.
In the Solvent Purification System, a distillation column is utilized to
separate CA
from water and azeotropic hydrocarbons. The feed streams to this distillation
column
comprise a stream comprising CA and water generated in the Oxidation System, a
stream
comprising CA and water generated in the Raffinate Wash System, and a stream
comprising
CA and hydrocarbon generated in the Solvent Flash/Recovery System. Due to the
high water
content and reduced CA content, the distillate resulting from this column is a
heterogeneous
azeotrope. Upon condensing, two liquid phases result. The hydrocarbon rich
phase is
-24-
CA 02543367 2011-08-16
combined with the gas oil feed to the Raffinate Wash System for recovery of
the hydrocarbon
and recovery of the carboxylic acid. The water phase that contains small
quantities of CA
and small quantities of hydrocarbon is split into two streams. One stream is
purged from the
system. This stream preferably comprises the water that enters the process
with hydrogen
peroxide and the water produced from the formation of PCA (Eqn. 1). The other
water
stream is recycled to the Raffinate Wash System as the extraction solvent.
In the Hydrocarbon Recovery System. the concentrated extract from the Solvent
Flash/Recovery System is distilled under vacuum to recover the hydrocarbon
content.
Vacuum distillation is necessary due to the high boiling points of the
sulfones contained in
this extract stream. The overhead product from this distillation is
hydrocarbon with 2.7 wt%
CA. This material is combined with the gas oil feed to the Raffinate Wash
System for
recovery of the hydrocarbon and the recovery of the CA. The material leaving
the bottom of
the vacuum distillation is a combination of hydrocarbon and sulfones. The
sulfone content is
approximately 32 wt%. This vacuum distillation recovers approximately 80
percent of the
hydrocarbon in the feed to this system. As a result, the overall hydrocarbon
yield for the
entire process is about 90 percent. Theoretically, the overall hydrocarbon
yield could be as
high as 97 percent. Experimentation on extract distillation followed by
additional process
engineering optimization is necessary to determine the feasibility of higher
hydrocarbon
yields, for example, higher steam pressures in the reboiler or deeper vacuum
levels in the
distillation column would allow additional hydrocarbon recovery.
-25-
CA 02543367 2011-08-16
Neutralization Reaction (Equations 8 and 9)
The process includes a section where wastewater is treated,. This wastewater
stream
contains CA and sulfuric acid that must be neutralized before disposal. The
neutralization
may be accomplished by utilizing sodium hydroxide. The products of this
neutralization are
sodium carboxylate (NaC) and sodium sulfate. The use of other neutralizing
bases may be
possible.
(8)
CA + NaOH NaC + 14,0
H2SO4 + NaOH - Na2SO4 + 2 H2O (9)
The disclosed process may be achieved by a reactor design, which is as
follows. The
design is particularly suitable for a typical small to medium petroleum
refinery that has
limited or no hydrodesulfurization (HDS) capability or has limited
availability of hydrogen.
For the purpose of the disclosed invention, one of ordinary skill would
understand that
the process and design comprises all equipment necessary for desulfurization
that normally
does not exist in a typical refinery.
It is noted that a feed capacity can range from as low as 5 Barrel Per Stream
Day
(BPSD) to as much as 50,000 BPSD, which includes 10, 15, 20, 25, 30, 40, 50,
100, 250, 500,
750, 1000. 2500, 5000, 7500, 10000, 15000, 20000, 25000, 30000, 35000, 40000,
and 45000
BPSD and range therein between and combination thereof.
Initial pilot plant studies were conducted using a middle distillate (Marine
Diesel)
obtained from Petro Star Inc. In particular, an ASTM D-86 Distillation Curve
was measured
for the Petro Star Inc. Marine Diesel (12/07/99), the results of which are
shown in Figure 2.
A simulated process is outlined below; wherein the feed employed for this
simulated
process was modeled to mimic a typical straight run LAGO derived from the
crude
atmospheric distillation unit in a typical petroleum refinery, i.e., Petro
Star Inc. Marine Diesel
-26-
CA 02543367 2011-08-16
(12/07/99). The measured and simulated distillation curves for the actual feed
is shown in
Fig. 2 and the components in the simulated feed are listed in Table 1.
-27-
CA 02543367 2011-08-16
c o 0o r
O Or O~ ~O a N N CO 00 CO N r
s 0 N N- N r rn C r N N- 0 00
d L. 0 N r N rn r a N N r
r r c, 0 0
o0 00 00 00 a ON a
CC
a=
u c
c O
a) 0. 00 00 00 -. r o r c
p bD o N- N oo 0 00 oc r r- C~
O O N 00 N M in O N NI 00 N
In In
F- ~
"~ s c
' aU t o a c
s O 75 s c: c: C põ 0 s
= C Q, s O 0 u 0 O
c y s `~ O. -c N O C N R. L:. r- s
y Ia. -a N 0 y O
0 sO .O
o
s s N N C
o s'' a~ c s c s
Y E , Q Ca 2
W'p'b
fl,
E N o0
N O 0 0 0
G) ) ) N C s
c v c a s U
N r, s y cu y ,
v .s t (U >> N N
CQ J u C C C id ¾. ~. 0
C o c c ,- s
o ti o :~ o s ~' s a ac
~ 0 CO r" .t rYj G ~ L y ¾' U.. O X i. 0 - 0..
p ry N
27 p N C C ~ N.
N
0 0 0
O 0 S 0 V O O N C G' (0 O C O O G 0 0
O 0 0 o- O O(0_ 0 0 0 0 C O C C
~, C) G fl- N O > >. ~ ca ~ ~ U U U 0 0 0 c3 n ca ca
7E V U V -0 -7; 7:;
rU >, 0 "O 'O 'O ca ca ca O O
ca U u C.)
,4+ Z dU 0 J C O ~, '~ Y, ¾, 0 O' C O 'i,
Q o C ~ v r. '~, =,_ >, .~ p f- ~ a) a) a) C O W a) L
U
N 0 4 n ^..a O O O C C= C
N rr: O
p N M
a o0 rl) In In `O o0 00 10 N n "0
GD In O -t 00 N In o0 - 7 r C. N ~r, r C,
C N N 111 cl) N =r n In 'O 'C =.D 'C 'c r
v o I I I I r
.O a 0o r7 ~n In o0 O0 N .o - et In
m
LT. ~n O ~n O '7 00 N In oo -- - r ^ N In r a
N N N N N -t d' -t n In In
of
CA 02543367 2011-08-16
The feed contains about 5,100 ppm by weight of sulfur in the form of
thiophenic
compounds including benzothiophene, dibenzothiophene, naphthobenzothiophene,
and
several of their substituted homologues. This corresponds to a thiophenic
composition of
2.89 wt %. The aliphatic content of the feed is about 66.4 wt % while the non-
sulfur
containing aromatic content of the feed is about 30.7 wt %.
DETAILED DISCUSSION OF A SIMULATED REACTOR AND PROCESS
A better appreciation of the disclosed invention may be made without limiting
the
scope of the invention by inspecting the details associated with a simulated
process, which is
represented pictorially in Figs. 3-7. and described in the following text. In
the following text,
numerical ranges are presented showing the range of values in which the
process may occur.
Next to the numerical ranges, preferred values are shown in parentheses.
As a guide for better understanding the figures, it should be noted that solid
lines
indicate continuous flow, while dashed lines indicate intermediate flow.
Streams flowing
throughout the process are designated numerically (Stream Nos. 1-50) - these
numbers being
enclosed within hexagons and located proximal to the stream in question. The
simulated
material balances and properties of the streams are tabulated in Tables 2-14
and appear below.
Reactors, columns, vessels. tanks, heat exchangers, pumps, and the like, are
represented
numerically (100-172). When different from the data shown in the tables,
stream physical
properties are presented as numbers within various geometrical shapes; e.g.,
stream
temperature (number in F enclosed in a rectangle), stream pressure (number in
psia enclosed
in oval), and stream mass flow (number in lb/hr enclosed in curved rectangle
(D)). Other
representations will be recognized by one of ordinary skill. For convenience,
streams that
lead to reactors, vessels, and the like that appear in separate figures are so
labeled along the
-29-
CA 02543367 2011-08-16
periphery of the figure with a directional indication of flow and a numerical
designation
showing the source/destination of the stream.
In this illustrated embodiment, the first liquid comprises a middle distillate
(Marine
Diesel) obtained from Petro Star Inc. The selected carboxylic acid is acetic
acid, which
means that reaction of acetic acid (AA) with hydrogen peroxide results in the
formation of
peracetic acid (PAA) as shown in eqn. (1).
Oxidation Stage (Figure 3)
The organosulfur compounds in the gas oil feed (first liquid) are
substantially
completely oxidized to polar organosulfur compounds via reactions with active
oxygen in the
form of PAA. As noted above, PCA may be formed in situ by reacting hydrogen
peroxide
with acetic acid. The overall conversion ofthiophenic sulfur to sulfones is
99.8 %. A total of
ranging between 2.5 to 5.0 (3.0) moles of hydrogen peroxide per mole of sulfur
are used in
the oxidation.
In the discussion concerning the solution chemistry of the oxidation process,
the
reaction mixtures in the Oxidation System comprise two liquid phases. The
formation of
PCA occurs in the heavy phase while the oxidation of organosulfur compounds to
polar
organosulfur compounds occurs in both phases. Sulfuric acid, hydrogen
peroxide, and water
primarily reside in the heavy phase. AA. PAA, thiophenics, and sulfones
distribute between
both phases. Hydrocarbons primarily stay in the light phase, although some of
the aromatic
compounds and, to a lesser extent, some of the aliphatic compounds in the gas
oil are soluble
in the heavy phase.
Fig. 3 shows a detailed depiction of the oxidation system. In particular, the
Oxidation
System utilizes two reactors (100A and 104A), two decanters (101A and 106), a
reboiled
flash vessel (108A), and three heat exchangers (102A, 105A, and 109A).
-30-
CA 02543367 2011-08-16
Fresh gas oil (Stream No. 1) may be introduced at a temperature of about 68 F
where
it may be first partially heated in a heat exchanger (105A) by a higher
temperature
downstream process fluid (Stream 7). The temperature of the fresh gas oil
stream upon
departure from the heat exchanger (105A) may be increased before introduction
to the reactor
(100A) by introducing said stream to a second heat exchanger (102A, which
employs 150-
psig steam) prior to the introduction of recycled acetic acid. The
introduction of recycled AA
from the Solvent Flash/Recovery System (Stream No. 29) which may be at a
temperature of
about 300 F (148.9 C) to the fresh gas oil stream occurs prior to entry into
the First Stage
Oxidizer (100A). Approximately, one pound of recycled AA is used for every
five pounds of
gas oil; wherein the combined stream has a temperature of about 176 F (80 C)
(Stream No.
5). The combined gas oil/AA stream is then fed to the First Stage Oxidizer
(100A).
Recycled oxidant (Stream No. 16) from the Second Stage Oxidizer Oil Decanter
(106) is also
fed to the First Stage Oxidizer (100A). This recycled stream comprises
approximately 1.8 to
3.0 moles of oxidant per mole of sulfur in the gas oil feed to the First Stage
Oxidizer (100A);
preferably about 2.5 moles ofoxidant per mole of suilfur in the gas oil feed
to the First Stage
Oxidizer (100A). In addition to oxidant, this recycle stream comprises the
catalyst
comprising sulfuric acid. As noted above, the temperature of the combined feed
(Stream No.
5) to the First Stage Oxidizer (100A) may range from about 140 F (60 C) to
about 194 F
(90 C). preferably (176 F (80 C)). Obviously, the precise temperature may be
dependent
upon the temperatures of both the heated feed gas oil and the recycled acetic
acid.
With an aim not to be limited by theory, it is believed that addition of AA to
the gas
oil prior to contacting with oxidant is important for maintaining a relatively
high
concentration of PAA in the heavy phase within the First Stage Oxidizer
(100A). Due to the
relatively high AA distribution coefficient, if the gas oil does not comprise
sufficient acetic
acid, redistribution may occur when the oxidant solution contacts the gas oil.
This
-31-
CA 02543367 2011-08-16
redistribution may cause a decrease in the AA concentration in the heavy
phase. This in turn
may cause some of the PAA in the heavy phase to revert back to AA and hydrogen
peroxide
in order to satisfy the reaction equilibrium conditions. Due to a less
favorable distribution
coefficient, hydrogen peroxide is not as effective as PAA, and therefore, an
overall decrease
in reaction rate would result.
The presence of sulfuric acid in the First Stage Oxidizer (100A) is also
important.
When the oxidant solution contacts the gas oil, PAA will distribute between
the two phases.
In the heavy phase, compensation for departure from reaction equilibrium
conditions can best
occur if the rate of PAA formation is relatively fast. Rapid PAA formation is
best obtained in
the presence of a strong acid catalyst like sulfuric acid.
In the First Stage Oxidizer (100A), the bulk of the organosulfur compounds may
be
converted to sulfones. Approximately. 96 to 99 percent conversion (98 percent)
may be
obtained within a residence time of about 5 to 30 minutes (20 minutes). On the
whole, the
reactor is designed to operate under adiabatic conditions at a pressure of 17
pounds per
square inch absolute (psia). The two liquid phases flow concurrently upward
through the
reactor, yet as the reaction proceeds, the heat generated by oxidation causes
the temperature
of the reaction mixture to increase. An outlet temperature may range from 145
F (62.8 C) to
200 F (93.3 C) (18 1 F (82.8 C)). The first stage oxidizer serves to provide
enhanced contact
between the two liquid phases. Mass transfer of PAA from the heavy phase to
the light phase
may dictate the overall reaction rate.
The reaction mixture (Stream No. 6) that leaves the First Stage Oxidizer
(100A) is fed
to the First Stage Oxidizer Oil Decanter (101A) where the two liquid phases
(light and heavy
phases) may be separated by gravity settling. In this particular portion of
the overall process,
the light phase is referred to as the first Stage Light Phase (Stream No. 7)
and the heavy
phase is referred to as the first Stage Heavy Phase (Stream No. 8).
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The First Stage Oxidizer Decanter (101A) operates at a pressure of about 17
psia.
The light phase comprises mostly hydrocarbon and acetic acid, sulfones, and
about 100 ppm
by weight sulfur in the form of unoxidized thiophenics. The heavy phase
comprises mostly
AA and water. However, this phase may further comprise sulfuric acid,
sulfones, and some
hydrocarbon. Due to the extended time at elevated temperatures, the amount of
active
oxygen either in the form of hydrogen peroxide or in the form of PAA is
expected to be close
to zero in both phases. The temperature of the light phase upon departure of
the First Stage
Oxidizer Oil Decanter (101A) is about 181 F (82.8 C).
The light phase is pumped (103A) to the Second Stage Oxidizer (104A). The
heavy
phase is fed forward by gravity to the Water Flash Vessel (108A).
In the Water Flash Vessel (108A), a portion of the heavy phase from the outlet
of the
First Stage Oxidizer (100A) is vaporized and sent as a vapor to the Solvent
Purification
Column (139; Stream No. 9). The Water Flash Vessel (108A) operates at about 18
psia. The
heat required for vaporization is supplied by the Water Flash Vessel Reboiler
(109A) by way
of medium pressure (MP) steam, but high pressure (HP) steam may be used as
well or a
combination of the two. Vaporization may be conducted at about 18 psia and a
temperature
of 240 F (1 15.6 C) to 410 F (210 C) (249 F (120.6 C)). The resulting vapor
stream
comprises mostly AA and about 2 to 20 wt% of water (9 wt%). The liquid
remaining after
vaporization comprises primarily AA, sulfones, hydrocarbon, a small amount of
water, and
about 2 wt% sulfuric acid. Most of this liquid (Stream No. 11) is pumped
(110A) to the inlet
of the Second Stage Oxidizer (104A). A portion (Stream No. 12) is purged from
the
Oxidation System and sent to the Wastewater Neutralization Vessel (167). The
AA lost in
this stream represents approximately 43 percent of the overall AA loss.
The water entering the system with the fresh hydrogen peroxide feed (Stream
No. 4)
and the water generated within the system during the formation of PAA is
removed via
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partial vaporization of the heavy phase leaving the first stage reactor as
described above.
Although water generated during the formation of PAA is primarily formed
within the
Second Stage Oxidizer (104A), removal of this water from the Oxidation System
can not be
accomplished until after contact in the First Stage Oxidizer (100A). The high
temperatures
used for vaporization would cause rapid and total decomposition of the active
oxygen.
The sulfuric acid used to catalyze the formation of PAA is theoretically
unused during
the reaction sequence. Therefore, total recycle of the sulfuric acid catalyst
is theoretically
possible. However, the fresh hydrogen peroxide entering the Oxidation System
comprises
stabilizers in the form of non-volatile salts. These salts are soluble in
water and tend to
remain in the heavy phase circulating in the Oxidation System. Total
recirculation of the
heavy phase. after water removal via vaporization, would therefore result in
an unchecked
accumulation of the stabilizers. A heavy phase purge is therefore required to
limit the
accumulation of stabilizers. Unfortunately, this heavy phase purge also
results in a loss of
sulfuric acid from the Oxidation System. Therefore, fresh sulfuric acid must
be added to
negate these sulfuric acid losses, and any losses due to side reactions of
sulfuric acid.
The gas oil feed to the Second Stage Oxidizer (104A) may be first cooled to
about
122 F (50 C) to about 158 F (70 C) (130 F (54.4 C)); so that upon introduction
of an
aqueous feed (Stream No. 11) coming from the Water Flash Vessel (108A) the
combined
feed will be about 140 F (60 C). In addition to this feed, fresh oxidant from
storage (Stream
No. 4) and fresh catalyst from a pipeline (Stream No. 2) may be added to the
gas oil feed at
some point prior to the introduction to the Second Stage Oxidizer (104A). In
addition, the
heavy phase is fed forward from the Water Flash Vessel to the inlet of the
Second Stage
Oxidizer (104A).
In the Second Stage Oxidizer (104A), the solvent comprising acetic acid and
fresh
oxidant comprising hydrogen peroxide come in contact to form PAA in situ;
wherein most of
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the unoxidized thiophenic compounds in the feed are converted to sulfones.
Approximately
88 to 95 percent (90 percent) conversion based on the unoxidized sulfur
content of the second
stage feed may be obtained with a residence time of about 15 to 80 minutes (20
minutes).
The reactor may operate under adiabatic conditions at a pressure of 17 psia.
The two liquid
phases may move concurrently in a pipe flow reactor. The temperature rise in
this reactor is
expected to be near zero, since the heat of reaction for the formation of PAA
is very small
and the amount of oxidation compared to the total mass flow is also very
small. The Second
Stage Oxidizer (104A) may provide enhanced contact between the two liquid
phases. Mass
transfer of PAA from the heavy phase to the light phase is again crucial to
the overall
reaction rate.
The reaction mixture that leaves the Second Stage Oxidizer (104A; Stream No.
14) is
fed to the Second Stage Oxidizer Oil Decanter (106) where the two liquid
phases are
separated by gravity settling. This decanter (106) operates at a pressure of
about 17 psia.
The light phase comprises mostly hydrocarbon, AA, and smaller amounts of PAA,
sulfones
and approximately 10 ppm by weight of unoxidized thiophenics. The heavy phase
comprises
mostly AA and water, and smaller amounts of hydrogen peroxide, PAA, sulfuric
acid,
sulfones, and some hydrocarbon.
Efficient use of oxidant is accomplished by first feeding fresh oxidant to the
Second
Stage Oxidizer (104A) and then recycling the unused oxidant from the outlet of
the Second
Stage Oxidizer (104A) to the inlet of the First Stage Oxidizer (100A). This
flow path for the
oxidant provides a high concentration of active oxygen in the Second Stage
Oxidizer (104A)
where the concentration of unoxidized organosulfur compounds is very low. The
Second
Stage Oxidizer (104A) operates at low temperature to minimize the consumption
of oxidant
in undesirable side reactions. Therefore, the heavy phase leaving the Second
Stage Oxidizer
(104A) comprises a substantial amount of unused oxidant. This makes the heavy
phase from
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the Second Stage Oxidizer (104A) an ideal candidate for recycling back to the
First Stage
Oxidizer (100A).
The light phase from the Second Stage Oxidizer Oil Decanter (106) is fed via
gravity
to the Sulfox Extraction System (Stream No. 15). The heavy phase from the
Second Stage
Oxidizer Oil Decanter (106) is recycled (Stream No. 16) via 107 to the inlet
of the First Stage
Oxidizer (100A).
Sulfox Extraction and Raffinate Washing (Figure 4)
In Sulfox Extraction and Raffinate Washing, small amounts of oxidant may be
removed from the raffinate by heat treatment and then most of the organosulfur
compounds
and AA may be removed from the gas oil via liquid-liquid extraction. Besides
the gas oil fed
forward from the Oxidation System, the recovered gas oil from the Solvent
Purification
System and the Hydrocarbon Recovery System are also treated in this system.
The gas oil
leaving this system contains approximately 50 ppm by weight of sulfur and
approximately
6000 ppm by weight of acetic acid.
A better understanding of the Sulfox Extraction and Raffinate Washing System
may
be gleaned by inspecting a pictorial depiction of a preferred embodiment shown
in Fig. 4. In
this representation, the Sulfox Extraction and Raffinate Washing System may
utilize a stirred
tank reactor (112). a packed extraction column (119), a mechanical extraction
column (122),
heat exchangers (114 - 118, and 120), and pumps (113, 121, 123, and 125). Gas
oil hold up
is provided at the end of this system by a simple vertical vessel (124).
Fresh gas oil enters this system (Stream No. 15) may range between 122 F (50
C) to
158 F (70 C) (140 F (60 C)) from the Oxidation System via gravity from the
Second Stage
Oxidizer Oil Decanter (106). Prior to entering the Destruct Reactor (112), the
gas oil may be
heated in a heat exchanger (115), by interchanging heat with the discharge
stream from the
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Destruct Reactor and in heat exchanger (114) by interchanging heat with the
recycle solvent
stream from the Solvent Recovery/Solvent Purification System. This heat
recovery system
raises the temperature of the gas oil to the desired Destruct Reactor (112)
temperature that
ranges from 212 F (100 C) to 250 F (121 C) (230 F (110 C)).
In the Destruct Reactor (112), any small amounts of oxidant may be decomposed
to
oxygen and acetic acid (see Eqn. 7). The residence time in the reactor may
vary from about 5
to about 20 minutes (10 minutes). An agitator (111) may be provided, for
example, to
maintain a homogeneous mixture. For startup purposes, the Destruct Reactor
(112) may be
equipped with a jacket serviced by 150 psig steam. Under steady state
conditions, steam
heating is not required. That is, the heat duty of the Destruct Reactor (112)
may be about 0
MMBtu/hr; consequently, the temperature of the stream exiting the Destruct
Reactor (112) is
about the same temperature as the stream that enters the reactor.
The gas oil (Stream No. 17) leaving the Destruct Reactor may be pumped (113)
to the
Sulfox Extraction Column (119), but is cooled by successively passing through
three heat
exchangers (115. 117, and 120). Before entering the extraction column, the gas
oil is cooled
from a temperature of about 230 F (110 C) to about 189 F (87.2 C) via a heat
exchanger
(115). by interchanging heat with the feed stream (Stream No. 15) to the
Destruct Reactor
(112). (As noted above, the temperatures obtained during the simulated reactor
process are
shown as numbers enclosed by rectangles.) Further downstream, the gas oil is
cooled (about
189 F (87.2 C) to about 147 F (63.9 C)) further via heat exchanger (117),
which in turn may
be accomplished by interchanging heat with the extract stream from the Sulfox
Extraction
Column (119). Finally, prior to the introduction of the gas oil to the Sulfox
Extraction
Column (119), the gas oil is cooled further (about 147 F (63.9 C) to about 113
F (45 C)) by
way of a heat exchanger (120), which may be cooled by cooling water (see
utilities above).
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The solvent used in the Sulfox Extraction Column is a combination of crude AA
(Stream No. 30) from the Solvent Flash Vessel Distillate Receiver (134) and
clean AA
(Stream No. 38) from the bottom of the Solvent Purification Column (139). This
combined
solvent is cooled to extraction temperature by successively passing through
three heat
exchangers (114. 116, and 118). (The temperatures obtained during the
simulated reactor
process are shown as numbers enclosed by rectangles.) The first heat exchanger
(114) cools
by interchanging heat with the feed stream to the Destruct Reactor (112). The
second heat
exchanger (116) cools by interchanging heat with the extract stream from the
Sulfox
Extraction Column (119). Finally, the third heat exchanger (118) cools by
circulated cooling
water (see utilities above). The extract (Stream 19) leaves via pump 121
through heat
exchangers 117 and 116 and is combined with Stream 24 before being delivered
to flash
evaporator 136 (Fig. 6).
In the Sulfox Extraction Column (119), more than 99 percent of the polar
organosulfur compounds comprising sulfones may be removed from the gas oil.
'There are three key process parameters associated with the Sulfox Extraction
Column:
(i) extraction temperature, (ii) water content of the extraction solvent, and
(iii) the solvent-to-
feed ratio. The current design is based on an extraction temperature that may
range from
about I00 F (37.8 C) to 150 F (65.6 C) (I 13 F (45 C)); solvent water content
that may
range from about 0.4 to 3.0 wt% (0.6 wt%); and a solvent-to-feed ratio that
may range from
about I to 2 (1.25). Of course, any combination of values for the three
parameters may be
realized for optimal performance of the extraction column.
Higher extraction temperatures and higher solvent-to-feed ratios would favor
the
removal of sulfones. Increased sulfone removal may result in a smaller
Raffinate Polishing
System. Unfortunately, these same higher temperatures and higher solvent-to-
feed ratios
simultaneously increase the amount of hydrocarbons that may be removed from
the gas oil,
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thereby reducing yield in this system and increasing the capacity of the
Hydrocarbon
Recovery System. In addition, higher solvent-to-feed ratios also increase the
capacity and
energy requirements of the solvent recovery system. Lower temperatures may be
undesirable
since special utility fluids such as chilled water would be necessary for
cooling the feeds to
the extraction column.
Higher water content may decrease the amount of hydrocarbon to be extracted
from
the gas oil, thereby decreasing the amount of hydrocarbon processed in the
Hydrocarbon
Recovery System. Obviously the interplay of many factors, including the
precise effect of
water content, will determine the ability of the solvent to extract sulfones.
The extract leaving the bottom of the Sulfox Extraction Column is pumped (121)
to
the Solvent Recovery / Solvent Purification System (Fig. 6). Before leaving
the Sulfox
Extraction and Raffinate Washing System, this relatively cold stream is used
to cool the gas
oil feed and the solvent feed to the Sulfox Extraction Column (119).
The raffinate (Stream No. 18) leaving the top of the Sulfox Extraction Column
may
be combined with the azeotropic hydrocarbon (Stream No. 36) and recovered
hydrocarbon
(Stream No. 49) streams from the Solvent Recovery and Solvent Purification
System (Fig. 6)
and the Hydrocarbon Recovery System (Fig. 7), respectively. In addition, the
spent gas oil
(Stream No. 25) used to rinse AA from the adsorption beds in the Raffinate
Polishing System
(Fig. 5) may also be added to this stream.
The combined gas oil (Stream No. 20) obtained from the Sulfox Extraction
Column
(119), the Solvent Recovery and Solvent Purification System (Fig. 6), the
Hydrocarbon
Recovery System (Fig. 7) and the Raffinate Polishing System (Fig. 5) may be
fed to the
bottom of the Raffinate Wash Column (122). This treatment serves to remove any
unwanted
AA from the gas of I feed.
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In the Raffinate Wash Column (122), most of the AA may be removed from the gas
oil by washing with substantially pure water (e.g., tap water with low mineral
content,
deionized water, distilled water, recycled water from solvent purification or
combinations
thereof). When this wash water is recycled from the Solvent Recovery and
Solvent
Purification System (Fig. 6), it comprises approximately 0 wt % to 5 wt % (1.5
wt %) acetic
acid.
There is one key process parameter associated with the Raffinate Wash Column
(122).
This key parameter is the solvent-to-feed ratio. The simulated design is based
on a solvent to
feed ratio of about 0.05, however, this ratio may range from 0.025 to 0.1;
wherein a higher
solvent-to-feed ratio results in higher AA recovery. Unfortunately, a drawback
of having too
high of a solvent-to-feed ratio necessitates a higher energy requirements in
the Solvent
Recovery and Solvent Purification System.
The washing temperature may range from about I00 F (37.8 C) to about 125 F
(51.7 C) (I 13 F (45 C))); and may primarily depend on the temperature of the
gas oil leaving
the Sulfox Extraction Column.
The extract leaving the bottom of the Raffinate Wash Column (122) is pumped
(123;
Stream No. 21) to the Solvent Purification Column (139) where the AA is
recovered and the
water is purified for recycle.
The raffinate leaving the top of the Raffinate Wash Column (122) flows via
gravity to
the Raffinate Hold Vessel (124). This vessel provides about 20 minutes of
surge time. From
the Raffinate Hold Vessel (124), the gas oil may be pumped (125) to the
Raffinate Polishing
System (126; Stream No. 22).
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Raffinate Polishing (Figure 5)
In the Raffinate Polishing System, small amounts of sulfur containing
compounds and
small amounts of AA are removed by adsorption onto a solid bed adsorbent. The
sulfur
content of the gas oil may be reduced to 10 ppm or less. It is estimated that
the AA content
may be reduced to 10 ppm or less.
The current design of this system is based on an observation that refinery
clay serves
generally as an effective adsorbent for polar organic compounds, particularly
polar organic
compounds and acetic acid. A particular type of refinery clay, also known as
Fuller's Earth,
may be used. However, it is believed that other forms of adsorbent material
may be used,
such as zeolites in general, silica, diatomaceous earth, natural adsorbents,
unnatural
adsorbents, mixtures thereof. or combinations thereof. Obviously many
parameters may
influence the manner in which polar organic compounds are adsorbed onto the
column
material: this may lead to a variety of adsorption system process parameters
that may be
optimized, e.g., type and/or amount of adsorbent material, temperature and/or
pressure of the
adsorption process and regeneration methods, etc.
The Raffinate Polishing System utilizes two parallel adsorption columns (126
and
129). one holding tank (127), two holding vessels (130 and 132), and three
pumps (128, 131,
and 133). One of the adsorption columns serves to polish the gas oil while the
other
adsorption column is being regenerated. The overall cycle may be about 12
hours.
For example. gas oil (Stream No. 22) from the Raffinate Holding Vessel (124)
is fed
to one of the Raffinate Polishing Columns (126). Organosulfur compounds and AA
are
adsorbed onto the solid bed as the gas oil flows through the column for about
a 6-hour period.
Upon exiting the column, the purified gas oil flows via gravity to the Product
Hold Tank
(127). After checking the quality, the gas oil (Stream No. 23) is pumped
intermittently (128)
to storage that may be outside the battery limits of the inventive process.
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During the same time period, the other Raffinate Polishing Column (129) is
being
regenerated. First, clean recycled AA is pumped through the bed. Organosulfur
compounds
left on the solid bed adsorbent by the crude gas oil are now desorbed by the
acetic acid.
Upon exiting the top of the column, the spent AA flows to the Spent AA Hold
Vessel (130).
This operation requires about 3 hours. Then. desulfurized gas oil from the
Product Hold
Tank (127) is pumped by 128 upward through the bed. The clean gas oil desorbs
AA left on
the bed from the previous step. Upon exiting the top of the column, the spent
gas oil flows to
the Spent Gas Oil Hold Vessel (132). This operation also requires about 3
hours.
The spent AA in the Spent AA Hold Vessel (130) may be continuously pumped
(131)
to the Solvent Recovery and Solvent Purification System (Fig. 6) where the AA
is recovered
and the polar organosulfur compounds removed from the gas oil via adsorption
join the
balance of the sulfur extract. The spent gas oil in the Spent Gas Oil Hold
Vessel (132) is
continuously pumped (133) to the Sulfox Extraction Column and Raffinate Wash
System
(Fig. 4) where the AA and gas oil are recovered.
Solvent Recovery and Solvent Purification (Figure 7)
In the Solvent Recovery System, the bulk of the AA is separated from the
sulfur
extract for immediate recycle. In the Solvent Purification System, mixtures of
acetic acid,
water. and hydrocarbons from several sources within the process are purified
for recycle and
purging.
The Solvent Flash and Solvent Purification System utilizes a single stage
flash vessel
(136) with accompanying heat exchangers (137 and 138); and a packed
distillation column
(139) with a vessel (142) and heat exchangers (141, 143, and 145).
The combined stream comprising the Sulfox Extraction Column (119) bottom
extract
(Stream No. 19) and the spent AA (Stream No. 24) from the Spent AA Hold Vessel
(130)
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may be fed to the Solvent Flash Vessel (136). The Solvent Flash Vessel
Reboiler (138) may
be used to vaporize a large portion of the feed with 300-psig steam.
The resulting bottoms stream (Stream No. 28) comprising sulfur extract and
approximately 15 wt% (10 to 50 wt%) AA may be sent forward to the Solvent
Recovery and
Hydrocarbon Recovery System.
The flashed vapor (Stream No. 27) is condensed in the Solvent Flash Vessel
Overhead
Condenser (137) and then may flow via gravity to the Solvent Flash Vessel
Distillate
Receiver (134). The condensed distillate comprises mostly AA with about 3 to
12-wt%
hydrocarbon (7-wt%). This hydrocarbon is mostly light boiling aliphatic and
aromatic
compounds that form minimum boiling homogeneous azeotropes with acetic acid.
The Solvent Flash System operates between a range from about 17 to about 75
psia
(45 psia). An elevated pressure may be utilized to establish a higher
condensing temperature
in the Solvent Flash Vessel Overhead Condenser (137). This elevated
temperature allows
heat integration with the bottoms of the Solvent Purification Column (139) by
providing most
5 of the reboiler heat duty required.
The condensed liquid from the Solvent Flash Vessel Overhead Condenser (137)
flows
via gravity to the Solvent Flash Vessel Distillate Receiver (134). This vessel
provides about
minutes of surge capacity for the unit operations that receive recycle
solvent. In addition,
this vessel is used to monitor the AA inventory within the process unit.
Utilizing on/off level
control, fresh AA from storage is added to this vessel periodically to make up
for AA losses
from streams leaving the process.
The crude AA from the Solvent Flash Vessel Distillate Receiver (134) may be
pumped (135) to the First Stage Oxidizer (100A; Stream No. 29), to the Sulfox
Extraction
Column (119; Stream No. 30), and to the Solvent Purification Column (139;
Stream No. 31).
The streams flowing to 100A and 119 are recycle streams, while the stream
leading to 139
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acts as a hydrocarbon purge for the main solvent recycle loops. Without the
purge stream to
the Solvent Purification Column (139), azeotropic hydrocarbon accumulation
would remain
unchecked in this recycle loop causing potential problems in the Oxidation
System (Fig. 3)
and in the Sulfox Extraction System (Fig. 4). The material balance herein is
based on a
recycle-to-purge weight ratio of about 5Ø but the recycle-to-purge weight
ratio may range
from about 4 to about 10. Employing a recycle-to-purge ratio of about 5, the
crude AA
recycle loop comprises hydrocarbon composition that is approximately 7.0 wt%.
A higher
recycle-to-purge ratio may result in some energy savings in the Solvent
Purification Column
(139). However, this higher recycle-to-purge ratio also causes a higher
hydrocarbon
concentration in the recycle streams. Clearly an optimum recycle-to-purge
ratio depends
upon many factors and conditions.
The Solvent Purification Column (139) receives vapor feed from the Water Flash
Vessel (108A; Stream No. 9). liquid feed from the Raffinate Wash Column (122;
Stream No.
21). and liquid feed from the Solvent Flash Vessel Distillate Receiver (134;
Stream No. 31).
The feed stream composition determines the ordering of the feed location, and
consequently
the respective introduction of each feed stream to the column. There may be at
least two feed
locations. The column may operate at about 17 psia. In the lower portion of
this distillation
column, water and light hydrocarbons are stripped from acetic acid. In the
upper portion of
this distillation column, AA is removed from water and hydrocarbon. The
separation
accomplished in this column is relatively difficult since the relative
volatility between water
and AA is low. The heat and material balance for this column is based on a
reflux ratio of 3.8
by weight and by employing a total of 38 theoretical stages. Obviously, the
optimum
configuration and operating conditions depend upon many conditions and
factors.
Approximately 90 percent of the heat required by the Solvent Purification
Column is
supplied by the Solvent Flash Vessel Overhead Condenser (137). Due to layout
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considerations, forced circulation is utilized for this reboiler. This
arrangement allows the
liquid condensate on the hot side of this reboiler to flow via gravity to the
Solvent Flash
Vessel Distillate Receiver (134). The balance of the heat requirement for the
Solvent
Purification Column is supplied in the Solvent Purification Column Trim
Reboiler (143) by
employing 150-psig steam. This is a thermosiphon reboiler and is used to
control the water
content of the streams leaving the bottom of the column.
The stream leaving the bottom of the Solvent Purification Column (139) may be
pumped (140) to either the Sulfox Extraction Column (119; Stream No. 38) or to
the Solvent
Hold Tank (Stream No. 39), or internally recycled to the main reboiler (143).
The streams
comprise mostly AA with approximately 0.5 wt% water and 1.5 wt% hydrocarbon.
The two
streams (Stream Nos. 38 and 39) represent the net bottoms output from the
distillation
column. One of these streams is recycled to the Sulfox Extraction Column (119;
Stream No.
38). The other stream is sent to the Solvent Hold Tank (Stream No. 39) where
it becomes the
recycle stream used to regenerate the adsorption beds in the Raffinate
Polishing System (Fig.
5).
Most of the hydrocarbon and water in the feeds to the Solvent Purification
Column
(139) may be driven overhead. At the top of the column, overhead vapors are
condensed in
the Solvent Purification Column Overhead Condenser (141). This is a total
condenser
utilizing cooling tower water as a heat sink. It may be necessary to vent non-
condensable
gases formed in the Water Flash Vessel (108A) by the decomposition of active
oxygen
containing species. Since the light hydrocarbons and water form minimum
boiling
heterogeneous azeotropes, two liquid phases are formed upon condensation (a
light phase and
a heavy phase). The immiscibility in this condensed stream is due to the high
concentrations
of water present. The two liquid phases are separated by gravity settling in
the Solvent
Purification Column Reflux Decanter (142), which operates at a pressure of
about 17 psia.
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The light phase comprises about 99.6 wt% hydrocarbon and may be recycled via
gravity to the bottom of the Raffinate Wash Column (122; Stream No. 36) where
the
recovered azeotropic hydrocarbon joins the main gas oil stream. The heavy
phase which is
water rich and contains approximately 1.4 wt% AA is pumped (144) to the
Solvent
Purification Column as reflex, to the top of the Raffinate Wash Column (122;
Stream No. 35)
as wash water recycle, and the Wastewater Neutralization Vessel (167; Stream
No. 34) as
purge. The recycle water to the Raffinate Wash Column (122; Stream No. 36) is
cooled in
the Solvent Purification Column Water Distillate Cooler (145) by cooling tower
water. The
purge stream leaving this distillation column represents most the water fed to
the Oxidation
System (Fig. 3) with the fresh hydrogen peroxide feed and most of the water
formed by
reaction in the Oxidation System. The AA leaving in this stream represents
approximately 30
percent of the total AA losses.
Solvent Recovery and Hydrocarbon Recovery (Figure 7)
In the Solvent Recovery System, additional AA is separated from the sulfur-
rich
extract for recycle. In the Hydrocarbon Recovery System, the remaining AA and
a large
portion of the hydrocarbons in the sulfur-rich extract are recovered for
recycle.
The Solvent Recovery and Hydrocarbon Recovery System utilizes a relatively
small
packed distillation column (149) accompanied by two heat exchangers (146 and
150) and a
solvent flash vessel (147), which may operate at atmospheric pressure.
Additionally, the
Solvent Recovery and Hydrocarbon Recovery System comprises a relatively large
packed
distillation column (152), three heat exchangers (154, 156, and 157) a
Hydrocarbon Recovery
Column Reflux Drum (158) all of which may operate under reduced atmospheric
pressure
(i.e., vacuum). Vacuum may be generated by the Hydrocarbon Recovery Column
Vacuum
System (166), which is a steam jet package. Condensate from the vacuum system
is
-46-
CA 02543367 2011-08-16
processed through the Wastewater Neutralization Vessel (167), which acts
simultaneously as
the seal for the vacuum system barometric legs and as the neutralization point
for all
wastewater streams leaving the process.
The Solvent Recovery Column (149) receives the sulfur extract stream from the
bottom of the Solvent Flash Vessel (136; Stream No. 28). This stream comprises
hydrocarbons. approximately 7.3 wt% sulfones, and 15 wt% AA. Above the feed
point of the
distillation column, hydrocarbon is removed from AA. Below the feed point of
the
distillation column, AA is stripped from the hydrocarbons and the sulfones.
The separation
accomplished in this column is relatively easy since the relative volatility
between AA and
I 0 hydrocarbons is high. The heat and material balance for this column is
based on a reflux ratio
of about 0.5 by weight and a total of 8 theoretical stages. Obviously, optimum
configuration
and operating conditions depend upon many factors.
Approximately 88 percent of the AA in the feed to the Solvent Recovery Column
(149) is sent overhead. Some hydrocarbon and substantially all the water in
the feed is also
sent overhead. The atmospheric bubble point of the liquid stream leaving the
bottom of the
distillation column limits additional recovery of acetic acid.
At the top of the column, overhead vapors (Stream No. 40) are condensed in the
Solvent Recovery Column Overhead Condenser (146). This is a total condenser
utilizing
cooling tower water as a heat sink. The condensed liquid flows to the Solvent
Recovery
Column Reflux Drum (147), which provides approximately 7.5 minutes of liquid
surge
capacity. The liquid leaving the Solvent Recovery Column Reflux Drum (147) may
be
pumped (148) to the top of the Solvent Recovery Column (149) as reflux and to
the Solvent
Flash Vessel Distillate Receiver (134) as recycle.
The heat required by the Solvent Recovery Column (149) may be supplied in the
Solvent Recovery Column Reboiler (150) by 300 psig steam. If desired, higher
pressure
-47-
CA 02543367 2011-08-16
steam could be used to increase AA recovery. Forced circulation is used for
this reboiler
since there is a significant increase in the bubble point of the liquid as
vaporization occurs.
The net liquid leaving the bottom of the column may be pumped (151) to the
Hydrocarbon
Recovery Column (152). This stream (Stream No. 43) may comprise mostly
hydrocarbon
and approximately 2 wt% AA and 8.5 wt% sulfones.
The Hydrocarbon Recovery Column (152) receives the stream (Stream No. 43)
comprising sulfur compounds from the bottom of the Solvent Recovery Column
(149).
Above the feed point of the distillation column, sulfones are removed from AA
and
hydrocarbons. Below the feed point of the distillation column, AA and
hydrocarbon are
stripped from the sulfones. The heat and material balance tabulated herein for
this column is
based on a reflux ratio of about 0.15 by weight and a total of 8 theoretical
stages. Obviously,
the optimum configuration of this column depends upon many conditions and
factors.
The top of the Hydrocarbon Recovery Column (152) operates at a pressure that
ranges
from about 5 mm Hg to about 15 mm Hg; preferably about 7 mm Hg to about 13 mm
Hg;
more preferably about 9 mm Hg to about 1 I mm Hg; most preferably about 10 mm
Hg. The
bottom of this column operates at a pressure that ranges from about 10 mm Hg
to about 20
mm Hg (about 15 mm Hg). The pressures utilized in this column were chosen
based on a
balance between the complexity of the vacuum system, the recovery of
hydrocarbon
overhead, and the bubble point of the resulting bottom stream. The current
process
configuration results in a column hydrocarbon recovery of about 80 percent by
weight, which
increases the overall hydrocarbon recovery for the entire process to at least
90 percent by
weight. Deeper vaccum levels and/or higher steam pressures in the reboiler may
be used to
increase hydrocarbon recovery.
At the top of the column, overhead vapors (Stream No. 44) are condensed in the
Hydrocarbon Recovery Column Overhead Condenser (157), which may be cooled
using
-48-
CA 02543367 2011-08-16
cooling tower water; and the Hydrocarbon Recovery Column Vent Condenser (156),
which
may be cooled using a 10 F (-12.2 C) aqueous brine solution or an aqueous
solution
comprising 25 wt% ethylene glycol. The vent condenser (156) minimizes losses
of AA to the
vacuum system. The liquid from both condensers flows to the Hydrocarbon
Recovery
Column Reflux Drum (158), which provides approximately 7.5 minutes of liquid
surge
capacity. The liquid leaving the Hydrocarbon Recovery Column Reflux Drum is
pumped
(159) to the top of the Hydrocarbon Recovery Column (152) as reflux and to the
bottom of
the Raffinate Wash Column (122; Stream No. 49) as recovered hydrocarbon and
acetic acid.
The heat required by the Hydrocarbon Recovery Column (152) is supplied in the
Hydrocarbon Recovery Column Reboiler (154), which may employ 300-psig steam. A
falling film reboiler is used for this application due to the potential
thermal sensitivity of the
bottom product. The net liquid leaving the bottom of the column may be pumped
(153)
through the Hydrocarbon Recovery Column Bottoms Cooler to storage (Stream No.
50).
This stream comprises approximately 68 wt% hydrocarbon and 32 wt% sulfones.
The vapor leaving the column vent condenser flows (Stream No. 45) to
Hydrocarbon
Recovery Column Vacuum System (166), which comprises a three-stage vacuum
package
utilizing 150-psig steam as the motive fluid and preferably comprising 3 jets
(160, 162, and
164) and 3 after-condensers (161, 163, and 165). The net gas leaving the
vacuum system is
sent to offgas treatment (Stream No 46). Very little AA is lost in this
stream. The condensed
process liquid and condensed steam from each after condenser flows via gravity
through two
separate barometric legs and a separate atmospheric leg to the Wastewater
Neutralization
Vessel (167). The AA lost in this stream represents approximately 26 percent
of the overall
AA loss.
The Wastewater Neutralization Vessel receives feed from the Water Flash Vessel
(108A; Stream No. 12) in the Oxidation System, from the Solvent Purification
Column
-49-
CA 02543367 2011-08-16
Reflux Decanter (142; Stream No. 34) in the Solvent Purification System. and
the
Hydrocarbon Recovery Column Vacuum System (166). These streams comprise
sulfuric
acid and/or AA which should be neutralized before purging to a wastewater
treatment system.
The neutralization may be accomplished by feeding 25 wt% caustic material
(Stream No. 3)
to this vessel. For example, when caustic material comprises sodium hydroxide,
sulfuric acid
is converted to sodium sulfate and AA is converted to sodium acetate. The
sensible heat in
the warm feed streams and the heat of neutralization may be removed by
recirculation
through the Wastewater Neutralization Vessel Cooler (169), which is serviced
by cooling
tower water. The net wastewater leaving the Wastewater Neutralization Vessel
is pumped
(168) to a Wastewater Treatment Plant (Stream No. 47).
-50-
CA 02543367 2011-08-16
Table 2. Material Balance and Properties of Streams - - 4
Stream Number 1 2 3 4
Hydrogen
Stream Description Catalyst 25 wt% Caustic Peroxide
Gas Oil Feed Makeup to Neutralization Feed
Temperature F 68.0 68.0 68.0 77.0
Pressure Asia 29.39 29.39 14.70 29.39
Total Flow lb-mol/hr 335.40 0.01 6.01 59.18
Total Flow lb/hr 62842.5 0.9 125.5 1457.3
Total Flow gpm 149.5 0.0 0.2 2.3
Total Flow hpsd 5125.8 0.0 6.8 77.9
Mass Flow lb/hr _
01 0.0 0.0 0.0 0.0
N, 0.0 0.0 0.0 0.0
1-110 0.0 0.0 94.1 437.2
1-1101 0.0 0.0 0.0 1020.1
H_SO, 0.0 0.9 0.0 0.0
Acetic Acid 0.0 0.0 0.0 0.0
Aliphatics 41733.7 0.0 0.0 0.0
Aromatics 19292.6 0.0 0.0 0.0
Thiophenes 1816.1 0.0 0.0 0.0
Sulfones 0.0 0.0 0.0 0.0
Sodium Hydroxide 0.0 0.0 31.4 0.0
Sodium Sulfate 0.0 0.0 0.0 0.0
Sodium Acetate 0.0 0.0 0.0 0.0
Mass Fraction lb/lb
01 0.0000 0.0000 0.0000 0.0000
N, 0.0000 0.0000 0.0000 0.0000
1110 0.0000 0.0200 0.7500 0.3000
H2O2 0.0000 0.0000 0.0000 0.7000
11_S0. 0.0000 0.9800 0.0000 0.0000
Acetic Acid 0.0000 0.0000 0.2500 0.0000
Aliphatics 0.6641 0.0000 0.0000 0.0000
Aromatics 0.3070 0.0000 0.0000 0.0000
Thiophenes 0.0289 0.0000 0.0000 0.0000
Sulfones 0.0000 0.0000 0.0000 0.0000
Sodium Hydroxide 0.0000 0.0000 0.2500 0.0000
Sodium Sulfate 0.0000 0.0000 0.0000 0.0000
Sodium Acetate 0.0000 0.0000 0.0000 0.0000
Sulfur Content ppmw
Actual 5099.3 320339.5 0.0 0.0
Fuel Basis 5099.3 na na na
Physical Properties
Density (liquid) lb/gal 7.00 14.39 10.59 10.67
Density (vapor) lb/113 na na na na
Heat Capacity btu/Ib-R 0.467 0.217 0.820 0.666
Viscosity cP 3.639 20.519 1.032
Viscosity (light phase) cP na na na na
Viscosity (heavy
phase) cP na na na na
-51 -
CA 02543367 2011-08-16
Table 3. Material Balance and Properties of Streams 5-8.
Stream Number 5 6 7 8
First Stage First Stage
Stream Description Oxidizer Oxidizer First Stage First Stage
Feed Effluent Light Phase Heavy Phase
"temperature F 176.0 181.4 181.4 181.2
Pressure psia 14.70 17.00 17.00 17.00
Total Flow lh-mol/hr 933.94 939.39 582.59 351.35
blat l1' lb/hr 96619.7 96619.7 76941.4 19518.3
'Dotal Flow gpm 218.8 553.3 179.9 38.5
total Flow bpsd 7500.1 18971.4 6167.2 1319.2
Mass Flow Ib/hr
01 0.0 160.0 0.0 0.0
N . 0.0 0.0 0.0 0.0
1110 772.0 1297.9 139.6 1158.4
H10, 993.0 0.0 0.0 0.0
112SO4 156.0 156.0 0.0 156.0
Acetic Acid 26754.5 26754.5 12750.3 14004.2
Aliphatics 42381.8 42381.8 42070.3 311.5
Aromatics 21469.5 21469.5 19625.1 1844.4
Thiophenes 1816.1 72.6 58.7 14.0
Sulfones 2276.7 4327.2 2297.4 2029.8
Sodium Hydroxide 0.0 0.0 0.0 0.0
Sodium Sulfate 0.0 0.0 0.0 0.0
Sodium Acetate 0.0 0.0 0.0 0.0
Mass Fraction lb/lb
01 0.0000 0.0017 0.0000 0.0000
N, 0.0000 0.0000 0.0000 0.0000
1110 0.0080 0.0134 0.0018 0.0593
1110, 0.0103 0.0000 0.0000 0.0000
H2SO4 0.0016 0.0016 0.0000 0.0080
Acetic Acid 0.2769 0.2769 0.1657 0.7175
Aliphatics 0.4386 0.4386 0.5468 0.0160
Aromatics 0.2222 0.2222 0.2551 0.0945
Thiophenes 0.0188 0.0008 0.0008 0.0007
Sulfones 0.0236 0.0448 0.0299 0.1040
Sodium Iydroxide 0.0000 0.0000 0.0000 0.0000
Sodium Sulfate 0.0000 0.0000 0.0000 0.0000
Sodium Acetate 0.0000 0.0000 0.0000 0.0000
Sulfur Content ppmw
Actual 7397.9 7397.9 4616.0 18424.9
Fuel Basis 10520.2 10472.9 5545.0 85630.9
Physical Properties
Density (liquid) lb/gal 7.36 7.39 7.12 8.45
Density (vapor) lb/ft') na na na na
I leat Capacity btu/Ib-R 0.527 0.587 0.518 0.595
Viscosity cP na na 0.867 0.544
Viscosity (light
phase) cP 0.827 na na
Viscosity (heavy
phase) cP 0.559 na na
-52-
CA 02543367 2011-08-16
Table 4. Material Balance and Properties of Streams 9-12
Stream Number 9 10 11 12
Recycle Acid Reactor
Stream Description Flash Drum Recycle Acid to Second Recycle
Vapor to First Stage Stage _ Purge
Temperature F 249.3 249.1 249.1 249.3
Pressure psia 18.00 18.00 18.00 18.00
Total Flow Ih-mol/hr 247.96 0.00 102.88 0.52
Total Flow lb/hr 11350.5 0.0 8127.8 40.8
Total Flow gpm 8892.6 0.0 15.4 0.1
Total Flow bpsd 304890.4 0.0 528.0 2.7
Mass Flow lb/hr
01 0.0 0.0 0.0 0.0
N, 0.0 0.0 0.0 0.0
FO 1050.3 0.0 107.6 0.5
11,0, 0.0 0.0 0.0 0.0
FI,SO4 0.1 0.0 156.0 0.8
Acetic Acid 10003.9 0.0 3980.3 20.0
Aliphatics 69.6 0.0 240.7 1.2
Aromatics 226.5 0.0 1609.8 8.1
Thiophcnes 0.2 0.0 13.7 0.1
Sulfones 0.0 0.0 2019.6 10.1
Sodium Ilvdroxide 0.0 0.0 0.0 0.0
Sodium Sulfate 0.0 0.0 0.0 0.0
Sodium Acetate 0.0 0.0 0.0 0.0
Mass Fraction lb/lb _
01 0.0000 0.0000 0.0000 0.0000
N2 0.0000 0.0000 0.0000 0.0000
H2O 0.0925 0.0132 0.0132 0.0132
11202 0.0000 0.0000 0.0000 0.0000
112S04 0.0000 0.0192 0.0192 0.0191
Acetic Acid 0.8814 0.4897 0.4897 0.4898
Aliphatics 0.0061 0.0296 0.0296 0.0296
Aromatics 0.0200 0.1981 0.1981 0.1981
Thiophenes 0.0000 0.0017 0.0017 0.0017
Sultones 0.0000 0.2485 0.2485 0.2485
Sodium Hydroxide 0.0000 0.0000 0.0000 0.0000
Sodium Sulfate 0.0000 0.0000 0.0000 0.0000
Sodium Acetate 0.0000 0.0000 0.0000 0.0000
Sulfur Content ppmw
Actual 5.0 44052.1 44052.1 44022.5
Fuel Basis 190.0 44052.1 92187.5 92115.6
Physical Properties __
Density (liquid) lb/gal na 8.79 8.79 8.79
Density (vapor) Ib/f13 0.159 na na na
I teat Capacity btu/Ib-R 0.966 0.572 0.572 0.572
Viscosity cP 0.012 0.497 0.497 0.497
Viscosity (light phase) cP na na na na
Viscosity (heav
phase) cP na na na na
-53-
CA 02543367 2011-08-16
Table 5. Material Balance and Properties of Streams 13-16
Stream Number 13 14 15 16
Second Stage Second Stage
Stream Description Oxidizer Oxidizer Second Stage Second Stage
Feed Effluent Light Phase I Leavy Phase
'I empcrature F 139.6 140.0 140.0 140.0
Pressure psia 17.00 17.00 17.00 17.00
Total Flow lb-mol/hr 744.65 744.65 514.53 230.12
Total Flow ib/hr 86526.5 86526.5 73545.4 12981.1
Total Flow gpm 196.0 196.0 170.4 24.3
Total Flow bpsd 6719.5 6719.9 5843.9 834.2
Mass Flow lb/hr
01 0.0 0.0 0.0 0.0
N, 0.0 0.0 0.0 0.0
1-110 684.3 698.7 58.2 640.4
11-01 1020.1 993.0 0.0 993.0
HLSO4 156.0 156.0 0.0 156.0
Acetic Acid 16730.6 16730.6 9205.7 7524.9
Aliphatics 42311.0 42311.0 42179.1 131.9
Aromatics 21234.9 21234.9 19976.7 1258.3
Thiophenes 72.4 0.0 0.0 0.0
Sullones 4317.1 4402.2 2125.6 2276.6
Sodium I Iydroxide 0.0 0.0 0.0 0.0
Sodium Sulfate 0.0 0.0 0.0 0.0
Sodium Acetate 0.0 _ 0.0 0.0 0.0
Mass Fraction Will
01 0.0000 0.0000 0.0000 0.0000
N, 0.0000 0.0000 0.0000 0.0000
11,0 0.0079 0.0081 0.0008 0.0493
11,0, 0.0118 0.0115 0.0000 0.0765
H,S04 0.0018 0.0018 0.0000 0.0120
Acetic Acid 0.1934 0.1934 0.1252 0.5797
Aliphatics 0.4890 0.4890 0.5735 0.0102
Aromatics 0.2454 0.2454 0.2716 0.0969
Thiophenes 0.0008 0.0000 0.0000 0.0000
Sultbnes 0.0499 0.0509 0.0289 0.1754
Sodium Hydroxide 0.0000 0.0000 0.0000 0.0000
Sodium Sulfate 0.0000 0.0000 0.0000 0.0000
Sodium Acetate 0.0000 0.0000 0.0000 _ 0.0000
Sulfur Content ppmw
Actual 8242.7 8242.7 4335.9 30376.6
Fuel Basis 10498.4 10496.4 4960.8 _ 107539.3
Physical Properties _ _ _
Density (liquid) lb/gal 7.35 7.35 7.19 8.88
Density (vapor) Ib/ft3 na na na na
I teat Capacity btu/Ib-R 0.486 0,486 0.485 0.487
Viscosity cP na na 1.278 0.745
Viscosity (light
phase) cP 1.235 1.235 na na
Viscosity (heavy
phase) cl' 0.765 0.762 na na
-54-
CA 02543367 2011-08-16
Table 6. Material Balance and Properties of Streams 17-20
Stream Number 17 '18 19 20
Destruct Sulfox Sulfox
Stream Description Reactor Extraction Extraction Wash
Effluent Raffinate _ Extract Column Feed
Temperature F 230.0 113.0 113.0 110.4
Pressure psia 17.00 29.39 29.39 0.19
Total Flow lh-mol/hr 514.53 315.22 1840.21 523.27
Total Flow lb/hr 73545.4 46294.6 119179.4 75270.0
Totall'low gpm 178.7 1112.0 229.7 176.0
Total Flow bpsd 6128.4 3839.3 7875.5 6035.4
Mass Flow lb/hr
01 0.0 0.0 0.0 0.0
N, 0.0 0.0 0.0 0.0
1110 58.2 1.7 604.4 16.5
I110, 0.0 0.0 0.0 0.0
II,SO4 0.0 0.0 0.1 0.0
Acetic Acid 9205.7 6315.6 89301.6 9159.8
Aliphatics 42179.1 35598.5 8336.0 46985.8
Aromatics 19976.7 4373.0 18816.1 19101.8
I'hiophenes 0.0 0.0 0.0 0.2
Sultuncs 2125.6 5.8 2121.2 5.9
Sodium Hydroxide 0.0 X0.0 0.0 0.0
Sodium Sulfate 0.0 0.0 0.0 0.0
Sodium Acetate 0.0 0.0 0.0 0.0
Mass Fraction lb/lb
01 0.0000 0.0000 0.0000 0.0000
N, 0.0000 0.0000 0.0000 0.0000
ILO 0.0008 0.0000 0.0051 0.0002
1110, 0.0000 0.0000 0.0000 0.0000
11,SO4 0.0000 0.0000 0.0000 0.0000
Acetic Acid 0.1252 0.1364 0.7493 0.1217
Aliphatics 0.5735 0.7690 0.0699 0.6242
Aromatics 0.2716 0.0945 0.1579 0.2538
Thiophcnes 0.0000 0.0000 0.0000 0.0000
Sullimes 0.0289 0.0001 0.0178 0.0001
Sodium Ilvdroxidc 0.0000 0.0000 0.0000 0.0000
Sodium Sulfate 0.0000 0.0000 0.0000 0.0000
Sodium Acetate 0.0000 0.0000 0.0000 0.0000
Sulfur Content ppmw
Actual 4335.9 18.9 2670.5 12.2
Fuel Basis _ 4960.8 21.9 10872.3 13.9
Physical Properties
Density (liquid) Ib/gal 6.85 6.88 8.64 7.12
Density (vapor) lb/ft3 na na na na
I feat Capacity btu/Ib-R 0.536 0.491 0.435 0.473
Viscosity cP 0.677 1.447 0.922 1.565
Viscosity (light
phase) cP na na na na
Viscosity (heavy
phase) cl' na na na na
-55-
CA 02543367 2011-08-16
Table 7. Material Balance and Properties of Streams 21-24
Stream Number 21 22 23 24 _
Wash
Stream Description Column Gas Oil to Product Gas Spent Acetic
[xtract Polishing Oil to Storage Acid -
Temperature F 113.0 113.0 113.0 110.7
Pressure psia 14.70 14.70 14.70 73.48
Total Flow lb-mol/hr 384.85 363.49 270.32 152.33
Total Flow lb/hr 12588.8 66437.9 56414.6 10023.3
Total Flow gpm 25.8 164.1 139.4 20.2
'T'otal Flow bpsd 885.4 5624.7 4779.9 691.5
Mass Flow lb/hr
02 0.0 0.0 0.0 0.0
N, 0.0 0.0 0.0 0.0
1110 3701.1 16.5 0.0 53.2
1110, 0.0 0.0 0.0 0.0
I1 SO4 0.0 0.0 0.0 0.0
Acetic Acid 8824.5 388.7 0.0 7459.9
Aliphatics 2.3 46984.8 40143.7 1742.6
Aromatics 59.2 19043.6 16270.9 762.8
Thiophenes 0.0 0.1 0.0 0.0
Sulfones 1.7 4.2 0.0 4.7
Sodium I lvdroxide 0.0 0.0 0.0 0.0
Sodium Sulfate 0.0 0.0 0.0 0.0
Sodium Acetate 0.0 0.0 0.0 0.0
Mass Fraction lb/lb
01 0.0000 0.0000 0.0000 0.0000
N2 0.0000 0.0000 0.0000 0.0000
1110 0.2940 0.0002 0.0000 0.0053
H202 0.0000 0.0000 0.0000 0.0000
I1=SO4 0.0000 0.0000 0.0000 0.0000
Acetic Acid 0.7010 0.0058 0.0000 0.7443
Aliphatics 0.0002 0.7072 0.7116 0.1739
Aromatics 0.0047 0.2866 0.2884 0.0761
'I hiophenes 0.0000 0.0000 0.0000 0.0000
Sultunes 0.0001 0.0001 0.0000 0.0005
Sodium I lvdroxide 0.0000 0.0000 0.0000 0.0000
Sodium Sulfate 0.0000 0.0000 0.0000 0.0000
Sodium Acetate 0.0000 0.0000 0.0000 0.0000 _
Sulfur Content ppmw
Actual 21.2 9.8 0.0 71.2
Fuel Basis _ 4220.0 9.9 0.0 _ 284.3
Physical Properties
Density (liquid) lb/gal 8.12 6.74 6.74 8.28
Density (vapor) lb/t13 na na na na
Heat Capacity btu/Ib-R 0.697 0.470 0.470 0.454
Viscosity cP 0.701 1.981 1.981 0.923
Viscosity (light phase) cP na na na na
Viscosity (heavy
phase) cP na na na na
-56-
CA 02543367 2011-08-16
Table 8. Material Balance and Properties of Streams 25-28.
Stream Number 25 26 27 28
Feed to
Stream Description Solvent Vapor from Liquid from
Recovery Solvent Solvent Recovery
Spent Gas Oil Flash Recovery Flash Flash
Temperature F 105.7 185.3 342.9 342.9
Pressure psia 73.48 44.09 44.09 44.09
Totall'low Ib-mol/hr 82.47 1992.54 1771.74 220.80
Total Flow lb/hr 9613.8 129202.7 100021.1 29181.5
Total Flow gpm 21.6 262.4 29118.3 66.5
Total Flow bpsd 739.0 8996.4 998340.1 2281.3
Mass Flow lb/hr
01 0.0 0.0 0.0 0.0
N, 0.0 0.0 0.0 0.0
1110 12,0 657.6 650.0 7.6
11,0, 0.0 0.0 0.0 0.0
Fl2SO4 0.0 0.1 0.0 0.1
Acetic Acid 2353.1 96761.5 92403.0 4358.5
Aliphatics 5175.5 10078.6 2508.6 7570.0
Aromatics 2073.0 19578.9 4459.3 15119.6
'1'hiophcnes 0.0 0.1 0.0 0.1
Sulfones (l.1 2126.0 0.2 2125.7
Sodium Hydroxide 0.0 0.0 0.0 0.0
Sodium Sulfate 0.0 0.0 0.0 0.0
Sodium Acetate 0.0 0.0 0.0 _ 0.0
Mass Fraction _ lb/lb
01 0.0000 0.0000 0.0000 0.0000
N, 0.0000 0.0000 0.0000 0.0000
11,0 0.0013 0.0051 0.0065 0.0003
F 110- 0.0000 0.0000 0.0000 0.0000
11'S04 0.0000 0.0000 0.0000 0.0000
Acetic Acid 0.2448 0.7489 0.9238 0.1494
Aliphatics 0.5383 0.0780 0.0251 0.2594
Aromatics 0.2156 0.1515 0.0446 0.5181
'l'hiophenes 0.0000 0.0000 0.0000 0.0000
Sulfones 0.0000 0.0165 0.0000 0.0728
Sodium Hydroxide 0.0000 0.0000 0.0000 0.0000
Sodium Sulfate 0.0000 0.0000 0.0000 0.0000
Sodium Acetate 0.0000 0.0000 0.0000 0.0000
Sulfur Content ppmw _
Actual 2.1 2468.8 0.4 10929.6
Fuel Basis 2.8 10036.1 5.4 12852.7
Physical Properties
Density (liquid) Ib/gal 7.43 8.20 na 7.30
Density- (vapor) lb/I13 na na 0.428 na
Heat Capacity btu/Ib-R 0.477 0.491 1.050 0.541
Viscosity cP 1.353 0.567 0.013 0.435
Viscosity (light phase) cl' na na na na
Viscosity (heavy
phase) ell na na na na
-57-
CA 02543367 2011-08-16
Table 9. Material Balance and Properties of Streams 29-32.
Stream Number 29 30 31 32
Recovered Recovered Acid Recovered Ovhd Vapor from
Stream Description Acid to First to Sulfox Acid to Purification
Stage Extractor Purification Column
Temperature F 299.9 299.9 299.9 222.3
Pressure psia 44.09 44.09 44.09 18.00
lb-
Total l'lovamol/hr 368.41 1181.35 295.34 1468.03
Total Flow lb/hr 20796.0 66684.0 16671.0 25355.6
Total Flow gpm 46.7 149.8 37.5 67112.5
Total Flow bpsd 1601.8 5136.1 1284.0 2300998.9
Mass Flow lb/hr
0, 0.0 0.0 0.0 0.0
N, 0.0 0.0 0.0 0.0
I12O 131.5 421.6 105.4 23994.6
11,0, 0.0 0.0 0.0 0.0
11,SO4 0.0 0.0 0.0 0.0
Acetic Acid 19229.7 61661.7 15415.4 347.8
Aliphatics 516.2 1655.3 413.8 354.4
Aromatics 918.5 2945.2 736.3 658.8
Thiophenes 0.0 0.0 0.0 0.1
SulIimes 0.0 0.1 0.0 0.0
Sodium hydroxide 0.0 0.0 0.0 0.0
Sodium Sulfate 0.0 0.0 0.0 0.0
Sodium Acetate 0.0 0.0 0.0 0.0
Mass Fraction lb/lb
01 0.0000 0.0000 0.0000 0.0000
N2 0.0000 0.0000 0.0000 0.0000
1110 0.0063 0.0063 0.0063 0.9463
11,0, 0.0000 0.0000 0.0000 0.0000
I12S04 0.0000 0.0000 0.0000 0.0000
Acetic Acid 0.9247 0.9247 0.9247 0.0137
Aliphatics 0.0248 0.0248 0.0248 0.0140
Aromatics 0.0442 0.0442 0.0442 0.0260
Thiophenes 0.0000 0.0000 0.0000 0.0000
Sullones 0.0000 0.0000 0.0000 0.0000
Sodium I Iydroxide 0.0000 0.0000 0.0000 0.0000
Sodium Sulfate 0.0000 0.0000 0.0000 0.0000
Sodium Acetate 0.0000 0.0000 0.0000 0.0000
Sulfur Content ppmw
Actual 0.4 0.4 0.4 0.0
Fuel Basis 5.2 5.2 5.2 0.0
Physical Properties _
Density (liquid) lb/gal 7.41 7.41 7.41 na
Density (vapor) lb/ft3 na na na 0.047
leat Capacity btu/Ib-R 0.653 0.653 0.653 0.522
Viscosity cI' 0.299 0.299 0.299 0.013
Viscosity (light
phase) cP na na na na
Viscosity (heavy
phase) cP na na na na
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CA 02543367 2011-08-16
Table 10. Material Balance and Properties of Streams 33-36.
Stream Number 33 34 35 36
Rellux to Water Purge from Solvent Water
Stream Description Purification Purification to Wash Distillate from
Column Column Column Purification
Temperature f 206.8 113.0 113.0 206.8
Pressure psia 18.00 18.00 18.00 18.00
Total Flow lb-mol/hr 1174.42 59.51 225.07 9.02
"Total Flow lb/hr 19603.6 993.3 37569 1001.7
"Total Flow gpm 42.6 2.0 7.7 2.6
Total Flow bpsd 1460.4 69.9 264.3 __90.3_
Mass Flow Ib/hr
01 0.0 0.0 0.0 0.0
N, 0.0 0.0 0.0 0.0
11,0 19312.1 978.6 3701.1 2.8
0.0 0.0 0.0 0.0
IT'S04 0.0 0.0 0.0 0.0
Acetic Acid 279.1 14.1 53.5 1.1
Aliphatics 6.8 0.3 1.3 345.9
Aromatics 5.6 0.3 1.1 651.8
Thiophenes 0.0 0.0 0.0 0.1
Sulfones 0.0 0.0 0.0 0.0
Sodium Hydroxide 0.0 0.0 0.0 0.0
Sodium Sulfate 0.0 0.0 0.0 0.0
Sodium Acetate 0.0 0.0 0.0 0.0
Mass Fraction lb/lb
01 0.0000 0.0000 0.0000 0.0000
N, 0.0000 0.0000 0.0000 0.0000
1110 0.9851 0.9851 0.9851 0.0028
1110, 0.0000 0.0000 0.0000 0.0000
II SO2 0.0000 0.0000 0.0000 0.0000
Acetic Acid 0.0142 0.0142 0.0142 0.0011
Aliphatics 0.0003 0.0003 0.0003 0.3453
Aromatics 0.0003 0.0003 0.0003 0.6507
Thiophenes 0.0000 0.0000 0.0000 0.0001
Sulfones 0.0000 0.0000 0.0000 0.0000
Sodium IIydroxidc 0.0000 0.0000 0.0000 0.0000
Sodium Sulfate 0.0000 0.0000 0.0000 0.0000
Sodium Acetate 0.0000 0.0000 0.0000 0.0000
Sulfur Content ppmw
Actual 0.0 0.0 0.0 10.5
Fuel Basis 5.6 _ 5.6 5.6 10.5
Physical Properties
Density (liquid) lb/gal 7.67 8.12 8.12 6.33
Density (vapor) lb/tt3 na na na na
I leat Capacity btu/Ib-R 0.987 0.926 0.926 0.520
Viscosity cP 0.289 0.612 0.612 0.361
Viscosity (light phase) cl' na na na na
Viscosity (heavy
phase) cP na na na na
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CA 02543367 2011-08-16
Table 11. Material Balance and Properties of Streams 37-40.
Stream Number 37 .38 39 40
Purified Acid Vapor
Stream Description Acetic Acid to Sulfox Purified Acid Distillate Acid
to Rcboiler Extraction to Storage Rec Column
Temperature P 255.1 255.1 255.1 255.9
Pressure psia 44.09 44.09 18.00 18.00
Total flow lb-mol/hr 6371.22 459.54 175.00 108.68
Total Flow lb/hr 350000.0 25244.6 9613.8 6122.4
Total Flow gpm 745.9 53.8 20.5 3447.4
Total Flow bpsd 25572.6 1844.5 704.0 118198.0
Mass Flow lb/hr
01 0.0 0.0 0.0 0.0
N, 0.0 0.0 0.0 0.0
11,0 1749.9 126.2 48.1 11.4
1110, 0.0 0.0 0.0 0.0
1 L504 0.9 0.1 0.0 0.0
Acetic Acid 343139.3 24749.8 9425.3 5784.8
Aliphatics 1387.7 100.1 38.1 115.1
Aromatics 3703.3 267.1 101.7 211.2
Thiophenes 1.0 0.1 0.0 0.0
Sulfones 17.8 1.3 0.5 0.0
Sodium Ilvdroxide 0.0 0.0 0.0 0.0
Sodium Sulfate 0.0 0.0 0.0 0.0
Sodium Acetate 0.0 0.0 0.0 0.0
Mass Fraction lb/lb
01 0.0000 0.0000 0.0000 0.0000
N, 0.0000 0.0000 0.0000 0.0000
H,0 0.0050 0.0050 0.0050 0.0019
11,0, 0.0000 0.0000 0.0000 0.0000
H2S04 0.0000 0.0000 0.0000 0.0000
Acetic Acid 0.9804 0.9804 0.9804 0.9448
Aliphatics 0.0040 0.0040 0.0040 0.0188
Aromatics 0.0106 0.0106 0.0106 0.0345
Thiophenes 0.0000 0.0000 0.0000 0.0000
Sul f ones 0.0001 0.0001 0.0001 0.0000
Sodium Hydroxide 0.0000 0.0000 0.0000 0,0000
Sodium Sulfate 0.0000 0.0000 0.0000 0.0000
Sodium Acetate 0.0000 0.0000 0.0000 0.0000
Sulfur Content ppmw
Actual 9.1 9.1 9.1 0.0
Fuel Basis 626.2 626.2 626.2 0.0
Physical Properties
Density (liquid) Ib/gal 7.81 7.81 7.80 na
Density (vapor) lb/ft') na na na 0.221
1leat Capacity btu/Ih-R 0.555 0.555 0.559 0.959
Viscosity eP 0.372 0.372 0.367 0.012
Viscosity (light phase) eP na na na na
Viscosity (heavy
phase) cP na na na na
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CA 02543367 2011-08-16
Table 12. Material Balance and Properties of Streams 41-44.
Stream Number 41 42 43 44
Reliux to Rec Acid from Bottoms from Ovhd Vapor
Stream Description Acid Rec Acid Rec Acid Rec from Hyd Rec
Column Column Column Column
Temperature F 253.1 253.1 396.5 294.1
Pressure psia 18.00 18.00 18.00 0.19
1lotal Flow lb-mol/hr 36.23 72.45 148.35 134.78
"total Flow lb/hr 2040.8 4081.6 25099.9 21140.8
Total Flow gpm 4.4 8.8 60.6 643009.1
Total Flow bpsd 150.3 300.6 2078.7 22046027.1
Mass Flow lb/hr
01 0.0 0.0 0.0 3.0
N, 0.0 0.0 0.0 10.0
lC0 3.8 7.6 0.0 0.0
1110, 0.0 0.0 0.0 0.0
Ii2SO4 0.0 0.0 0.1 0.0
Acetic Acid 1928.3 3856.5 502.0 575.5
Aliphatics 38.4 76.7 7493.3 6746.8
Aromatics 70.4 140.8 14978.8 13805.5
Thiophenes 0.0 0.0 0.1 0.0
Sulfones 0.0 0.0 2125.7 0.0
Sodium Hydroxide 0.0 0.0 0.0 0.0
Sodium Sulfate 0.0 0.0 0.0 0.0
Sodium Acetate 0.0 0.0 0.0 0.0
Mass Fraction lb/lb
01 0.0000 0.0000 0.0000 0.0001
N, 0.0000 0.0000 0.0000 0.0005
11:0 0.0019 0.0019 0.0000 0.0000
11:0: 0.0000 0.0000 0.0000 0.0000
Il,S0.1 0.0000 0.0000 0.0000 0.0000
Acetic Acid 0.9448 0.9448 0.0200 0.0272
Aliphatics 0.0188 0.0188 0.2985 0.3191
Aromatics 0.0345 0.0345 0.5968 0.6530
Thiophenes 0.0000 0.0000 0.0000 0.0000
Sultones 0.0000 0.0000 0.0847 0.0000
Sodium Hydroxide 0.0000 0.0000 0.0000 0.0000
Sodium Sulfate 0.0000 0.0000 0.0000 0.0000
Sodium Acetate 0.0000 0.0000 0.0000 0.0000
Sulfur Content ppmw
Actual 0.0 0.0 12706.9 0.7
Fuel Basis _ 0.0 0.0 12966.3 0.7
Physical Properties
Density (liquid) lb/gal 7.84 7.75 6.89 na
Density (vapor) lb/113 na na na 0.00410
Heat Capacity btu/Ib-R 0.512 0.524 0.550 0.428
Viscosity cP 0.390 0.369 0.456 0.007
Viscosity (light
phase) eP na na na na
Viscosity (heavy
phase) cP na na na na
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CA 02543367 2011-08-16
Table 13. Material Balance and Properties of Streams 45-48.
Stream Number 45 416 47 48
Vapor to Offgas to Wastewater
Stream Description Vacuum Thermal to Treatment Rellux to Hyd
System Oxidizer Plant Rec Column
Temperature F 20.0 113.0 115.5 112.4
Pressure psia 0.19 17.40 17.40 0.19
Total Flow lb-mol/hr 0.73 0.50 86.20 17.48
'total Flow lb/hr 27.7 14.2 1487.8 2753.1
"total Flow gpm 1927.6 20.2 2.8 6.0
'total Flow bpsd 66090.5 693.1 96.9 205.4 _
Mass Flow lb/hr
01 3.0 2.6 0.0 0.0
N, 9.9 9.1 0.0 0.0
H10 0.0 0.7 1402.6 0.0
11,0, 0.0 0.0 0.0 0.0
H_SO00.0 0.0 0.0 0.0
Acetic Acid 12.0 0.0 0.0 73.5
Aliphatics 1.3 0.9 2.0 879.6
Aromatics 1.4 0.9 8.9 1800.0
Thiophenes 0.0 0.0 0.1 0.0
Su!Ibnes 0.0 0.0 10.1 0.0
Sodium Hydroxide 0.0 0.0 0.0 0.0
Sodium Sulfate 0.0 0.0 1.1 0.0
Sodium Acetate 0.0 0.0 63.0 _ 0.0
Mass Fraction lb/lb
01 0.1084 0.1817 0.0000 0.0000
N, 0.3586 0.6427 0.0000 0.0000
110 0.0000 0.0466 0.9427 0.0000
1110, 0.0000 0.0000 0.0000 0.0000
II,SO00.0000 0.0000 0.0000 0.0000
Acetic Acid 0.4338 0.0008 0.0000 0.0267
Aliphatics 0.0484 0.0638 0.0013 0.3195
Aromatics 0.0507 0.0645 0.0060 0.6538
Thiophenes 0.0000 0.0000 0.0000 0.0000
Sullones 0.0000 0.0000 0.0068 0.0000
Sodium Iydroxide 0.0000 0.0000 0.0000 0.0000
Sodium Sulfate 0.0000 0.0000 0.0008 0.0000
Sodium Acetate 0.0000 0.0000 0.0423 0.0000
Sulfur Content ppmw
Actual 0.0 0.0 1208.4 0.7
Fuel Basis 0.0 X0.0 85330.7 0.7
Physical Properties
Density (liquid) lb/gal na na 8.11 7.66
Density (vapor) lh/113 0.00179 0.087 na na
I lcat Capacity btu/lh-R 0.427 0.268 0.922 0.417
Viscosity cP 0.013 0.018 na 2.177
Viscosity (light
phase) cP na na 0.599 na
Viscosity (heavy
phase) cP na na 2.706 na
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CA 02543367 2011-08-16
Table 14. Material Balance and Properties of Streams 49-50.
Stream Number 49 50
Stream Description Byproduct
Hyd Rec Extract to
Distillate Product Storage
Temperature F 112.4 131.0
Pressure psia 0.19 17.00
"total Flow lb-mol/hr 116.56 31.54
"total Flow lb/hr 18360.0 6725.3
Total Flow gpm 39.9 12.8
Total Flow hpsd 1369.5 439.3
Mass Flow lb/hr
01 0.0 0.0
N, 0.1 0.0
110 0.0 0.0
H O~ 0.0 0.0
It's(), 0.0 0.1
Acetic Acid 490.0 0.0
Aliphatics 5865.8 1626.1
Aromatics 12004.0 2973.3
Thiophenes 0.0 0.0
Sullones 0.0 2125.7
Sodium Ilydroxide 0.0 0.0
Sodium Sulfate 0.0 0.0
Sodium Acetate 0.0 0.0
Mass Fraction lb/lb
01 0.0000 0.0000
N, 0.0000 0.0000
1110 0.0000 0.0000
H1O2 0.0000 0.0000
I1'SO00.0000 0.0000
Acetic Acid 0.0267 0.0000
Aliphatics 0.3195 0.2418
Aromatics 0.6538 0.4421
Thiophenes 0.0000 0.0000
Sullones 0.0000 0.3161
Sodium Ilydroxide 0.0000 0.0000
Sodium Sulfate 0.0000 0.0000
Sodium Acetate 0.0000 0.0000
Sulfur Content ppmw
Actual 0.7 47422.9
Fuel Basis 0.7 47423.4
Physical Properties
Density (liquid) lb/gal 7.66 8.75
Density (vapor) Ib/tt3 na na
I teat Capacity btu/Ib-R 0.417 0.343
Viscosity cP 2.177 6.093
Viscosity (light
phase) eP na na
Viscosity (heavy
phase) eP na na
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CA 02543367 2011-08-16
Detailed Equipment Description
Thus far, the bulk of the disclosure is directed toward the process and its
generalized
unit operations. At this point a discussion of the specific equipment is
warranted. These
equipment details can be used in a specific embodiment of the inventive
process. They are
based on a U.S. Gulf coast facility designed to process 5100 bbls/day of light
atmospheric gas
oil at a sulfur content in the feed of about 5100 ppmw, and produce a product
gas oil with a
sulfur content of about 10 ppmw.
Reactors
In the present design. there are three reactors in the invention process: the
First Stage
Oxidizer (100), the Second Stage Oxidizer (104), and the Destruct Reactor
(112).
The simulated process described herein employed a mechanically agitated
contactor
for the First Stage Oxidizer. Normally, this type of contactor is used in
countercurrent liquid-
liquid extraction. This device may be preferably utilized as a co-current
upflow liquid-liquid
contactor. With this flow pattern, this device mimics the effects of a plug
flow reactor with
minimal back mixing. The agitation enhances mass transfer by creating
dispersed heavy
phase droplets within the continuous light phase. In addition, the agitation
minimizes the
difference in the velocity of the phases in order to give approximately equal
residence time
for each phase.
A pilot scale mechanically agitated contactor achieved approximately 96
percent
conversion of the sulfur containing compounds in the gas oil. The volume of
the commercial
mechanically agitated contactor is based on a 20-minute residence time used in
the pilot
process and the dimensions were scaled according to the hydraulic capacity of
the test
apparatus.
It should be noted, however, that utilizing a mechanically agitated contactor
for the
First Stage Oxidizer is a very expensive option. In addition, this apparatus
is generally
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CA 02543367 2011-08-16
speaking unattractive to operating personnel due its mechanical nature and
probable need for
intensive maintenance. It is conceivable that this apparatus may be replaced
with a less
expensive type of column that does not have moving parts (100B, Fig. 8). Also
see concepts
under Improved Oxidation Schemes.
The Second Stage Oxidizer (104A) is a pipe reactor, equipped with static mixer
elements. The volume of the reactor is based on the 10-minute residence time
used in the
laboratory experiments. The diameter of the pipe is based on the minimum
velocity
necessary for creating coarse heavy phase droplets that are dispersed in the
continuous light
phase. Also see concepts under Improved Oxidation Schemes.
For the Destruct Reactor (111), a continuous stirred vessel was chosen. The
operating
temperature is about 230 F (110 C). During steady state operation,
interchangers transfer
sufficient heat to the feed from other process streams. A conventional jacket
is preferably
provided for startup purposes only and uses 150 psig steam when heating is
necessary. The
working liquid volume provides approximately 10 minutes of residence time. The
dimensions were chosen for maximizing agitator performance.
Extraction Columns
There are two liquid-liquid extraction columns in the invention process. These
are the
Sulfox Extraction Column (119) and the Raffinate Wash Column (122).
The Sulfox Extraction Column (119) is a countercurrent packed bed liquid-
liquid
contactor. The column is equipped with structured packing.
The Raffinate Wash Column (122) is a countercurrent mechanically agitated
liquid-
liquid contactor. During the same testing program, the trials utilizing a
packed bed extractor
in this application revealed poor dispersion of the phases. Additional energy
input was
necessary to overcome the high interfacial surface tension between the two
liquid phases. As
stated earlier, the solvent to feed ratio in this column is very low. This low
solvent to feed
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CA 02543367 2011-08-16
ratio also decreases the mass transfer efficiency. The commercial column was
scaled from
the pilot tests based on the hydraulic capacity needed for the larger
throughput. The
commercial column contains 36 agitated stages. The heavy phase is dispersed,
while the light
phase is continuous.
Utilizing a mechanically agitated contactor for the Raffinate Wash Column
(122) is a
very expensive option. In addition, this apparatus is generally speaking
unattractive to
operating personnel due its mechanical nature and probable need for intensive
maintenance.
There are potential process improvements aimed at replacing this apparatus
with a less
expensive type without moving parts (Fig. 8).
Distillation Columns
There are three distillation columns in the invention process. These are the
Solvent
Purification Column (139). the Solvent Recovery Column (149), and the
Hydrocarbon
Recovery Column (152). In all three cases, conventional packed columns were
utilized.
The Solvent Purification Column (139) is relatively large with an estimated
height of
82 feet (tangent to tangent) and a diameter of 7 feet. The separation is
difficult due to the low
relative volatility between water and acetic acid. A total of 38 theoretical
stages are
necessary to complete the separation. High efficiency packing is utilized to
minimize the
column height. The column operates slightly above atmospheric pressure at 17
psia. The
column includes three packed sections so that the optimum feed location is
used for the
various streams entering the column.
The Solvent Recovery Column (149) is relatively small with an estimated height
of 25
feet (tangent to tangent) and a diameter of 18 inches. The separation of AA
from extract is
relatively easy. A total of 8 theoretical stages are necessary to complete the
separation.
Standard packing is utilized to minimize cost. The column operates slightly
above
atmospheric pressure at 18 psia.
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CA 02543367 2011-08-16
The Hydrocarbon Recovery Column (152) is relatively short with an estimated
height
of 28 feet (tangent to tangent) but has a relatively large diameter at 7 feet.
The separation of
hydrocarbons from sulfones is relatively easy. However, the column operates at
a pressure of
about 0. 19 psia, which creates considerable volumetric vapor traffic. A total
of 8 theoretical
stages are necessary to complete the separation. Standard packing is utilized
to minimize
cost.
Liquid-Li uid Decanters
There are three decanters in the invention process. These are the First Stage
Oxidizer
Oil Decanter (101), the Second Stage Oxidizer Oil Decanter (106), and the
Solvent
Purification Column Reflux Decanter (142).
Conventional horizontal gravity separators with internal baffles are utilized.
Generally speaking, the materials being separated have a low viscosity (< 2
cP) and the
density ratio between the heavy phases and light phases are approximately 1.2.
Therefore,
separations are relatively easy. Conservative methods were utilized for
sizing, and therefore.
a reduction in the dimensions of these decanters is definitely possible.
Efficient separation in the Second Stage Oxidizer Oil Decanter (106) is
preferred
since carryover of heavy phase to the Destruct Reactor (112) would compromise
its 316 SS
materials of construction.
Vapor-L,i uid Separators
There are two primary vapor-liquid separators in the invention process. These
are the
Water Flash Vessel (108A), the Solvent Flash Vessel (136). The Water Flash
Vessel (108A)
and the Solvent Flash Vessel (136) are conventional vertical separators with
mist eliminators.
Generally speaking, the vapor-liquid separation in these vessels is relatively
easy. Efficient
separation in the Water Flash Vessel (108A) is preferred since carryover of
the liquid phase
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CA 02543367 2011-08-16
to the Solvent Purification Column (139) would compromise its 316 SS materials
of
construction.
Adsorption Columns
There are two adsorption columns in the invention process. These are the
Raffinate
Polishing Columns (126 and 129). The columns are identical with an estimated
height of 42
feet (tangent to tangent) and a diameter of 5 feet. Each column contains two
15-foot beds of
refiner's clay; however, other adsorbent material may be possible. Both
columns are used for
polishing the gas oil by removing small amounts of sulfur-containing compounds
and small
amounts of acetic acid.
Heat Exchangers
There are a total of 25 heat exchangers in the invention process. Shell and
Tube
exchangers are utilized in all cases. Generally speaking, a horizontal
orientation was used for
condensing applications and a vertical orientation was used for vaporizing
applications.
There are five traditional reboilers. The Water Flash Vessel Reboiler (109A),
the
Solvent Flash Vessel Reboiler (138), and the Solvent Purification Column Trim
Reboiler
(143) are thermosiphons. The Solvent Recovery Column Reboiler (150) utilizes
forced
circulation since bubble point variation along the boilling path is large. The
Hydrocarbon
Recovery Column Reboiler (154) is based on falling film technology to minimize
the hot wall
contact time for the concentrated sulfone stream.
The Solvent Flash Vessel Overhead Condenser (137) has a dual function. The
exchanger is used to condense vapors from the Solvent Flash Vessel while
vaporizing liquid
from the Solvent. Purification Column. A vertical orientation is utilized with
the vaporization
on the tube side and the condensation on the shell side. Forced circulation is
utilized on the
vaporizing tube side to allow gravity flow liquid return to the Solvent Flash
Vessel Distillate
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CA 02543367 2011-08-16
Receiver for the condensing shell side. Where present, non-condensable gases
were
considered in the design of condensers.
Potential Oxidation System Improvement
The oxidation system described above comprises the following concepts: (1) two
stage addition of oxidant - this assures that the lowest concentrations of
unoxidized sulfur
compounds are in contact with the highest concentrations of oxidant; (2) Water
Removal
between Oxidation Stages - this allows recycle of the heavy phase leaving the
second Stage
Oxidation to the First Stage Oxidation without substantial loss of oxidant.
This also
eliminates any water dilution affects on the fresh oxidant added to the second
Stage
Oxidation, and therefore. promotes maximum mass transfer of oxidant from the
heavy phase
to the light phase: and (3) second Stage Oxidation at reduced temperature -
this serves to
minimize unwanted side reactions in the second Stage Oxidation, and therefore,
preserves the
oxidant for recycle to the First Stage Oxidation.
Based on these concepts and the simulated commercial process shown in Figs. 3-
8
and described above, and extensive laboratory experimentation showed that the
proposed
reactor and process will be capable of consistently producing gas oil
comprising a sulfur
content less than 10 ppm by weight.
However, it is believed that improvements may be possible. For example, it
should
be noted that the residence time required for oxidation in the first stage is
relatively short.
Conversions greater than 98 percent were obtained in less than 5 to 10
minutes. However.
experimental data indicated that a relatively long residence time would be
required in the
second Stage Oxidation. Long residence time in the second Stage Oxidation
would result in a
large expensive reactor. In addition, a long residence time would also cause
excessive
depletion of oxidant via side reactions in the heavy phase.
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CA 02543367 2011-08-16
Based on measurements of the active oxygen concentration in the light phase
during
the second Stage Oxidation, it became apparent that the reaction mechanism at
low
concentrations of unoxidized sulfur compounds is kinetically controlled rather
than mass
transfer controlled. In addition. it was discovered that the solubility of PAA
in the light phase
is great enough to provide a stoichiometric excess for completing the
oxidation.
With this information in mind, a revised oxidation concept was developed, in
which a
pictorial depiction is shown in Figure 8 and is described as follows.
Improved Oxidation Scheme -Part 1 (Figure 8)
A two-stage oxidation with water removal between stages and a lower second
stage
oxidation temperature is still employed. However, in this revised scheme, the
Second Stage
Oxidizer is a single liquid phase plug flow reactor (104B). Mass transfer of
the oxidant to the
light phase is accomplished in a short residence time static mixer placed
immediately up-
stream of the single liquid phase plug flow reactor. A residence time of one
to two minutes is
sufficient to transfer sufficient oxidant from the heavy phase into the light
phase. The lower
portion of the plug flow reactor is used to separate the two phases via
gravity settling. The
heavy phase is immediately recycled to the First Stage Oxidizer. Immediate
removal of the
heavy phase minimizes the extent of side reactions, and therefore, maximizes
the amount of
recycle oxidant. The isolated light phase flows through the plug flow reactor,
where
residence times can be made arbitrarily long without an excessive cost impact.
Since the residence time in the Second Stage Oxidizer is shorter than the time
required for sufficient in situ conversion of hydrogen peroxide to PAA, a
continuously stirred
tank reactor (CSTR) is added to the oxidation system. The PAA Reactor (171) is
used to pre-
form the PAA from fresh 70 wt% hydrogen peroxide and recycle acetic acid. The
fresh
catalyst necessary to replace the sulfuric acid purged from the oxidation
system via Water
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CA 02543367 2011-08-16
Flash Vessel (108B) is introduced through the PAA Reactor (171). Although a
CSTR was
chosen for the application. it is technically feasible to utilize a simple
plug flow reactor
instead. The CSTR is expected to cost more than the plug flow reactor, but it
does offer an
easier mode of operation, especially with respect to startup of the oxidation
system.
Based on the short residence time requirements for the First Stage Oxidation,
the
relatively expensive, high maintenance, mechanically agitated (e.g., 100A,
Fig. 3) is replaced
with a plug flow pipe reactor equipped with an internal static mixer. This is
expected to
reduce capital costs and maintenance costs.
A pilot reactor systern was employed to conduct a continuous flow pilot
testing that
serves as the basis for the revised oxidation concept and commercial extension
described
above. The results of this pilot testing. although not optimized, indicate
that the oxidation
system can consistently produce gas oil with less than 25 ppmw of unoxidized
sulfur
compounds.
Improved Oxidation Scheme - Part 2
To obtain gas oil with less than 10 ppm,v of unoxidized sulfur compounds, it
is
proposed that a three-stage oxidation system be used.
The residence time in the Second Stage Oxidizer shown in Figure 8 would be
divided
appropriately into two reactors. Each of these reactors would be equipped with
a mixing
zone at the inlet, followed by a separation zone where the heavy phase would
be removed.
Finally, each of these reactors would have a single liquid phase pipe flow
segment where the
sulfur containing compounds in the light phase continue to be oxidized.
The gas oil would flow through these two reactors in series. The fresh PAA
solution
from the Peracetic Acid Reactor (171) would be split into two parallel
streams. Each of these
two PAA streams would be fed to the inlet mixing sections of a Second Stage
Oxidizer and a
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third Stage Oxidizer. The heavy phase from the settling zone of each of these
reactors would
be recycled to the First Stage Oxidizer.
Other Process Improvements
In the process described thus far, the spent acetic acid used to regenerate
the Raffinate
Polishing Columns is sent to the Solvent Recovery and Solvent Purification
System (Fig 6)
Purification Column (139). This spent acetic acid contains sulfone compounds
and possibly
small amounts of unoxidized sulfur compounds. A more energy efficient approach
is to
recycle the spent acetic acid directly to the beginning of the process in
order to partially
saturate the gas oil feed. This reduces the heat load requirements for the
Solvent Flash
Vessel (136) by approximately 10 percent. In addition, if the spent acetic
acid from the
Raffinate Polishing Columns (126 and 129) comprise unoxidized sulfur
compounds,
recycling this stream to the oxidation system would increase conversion and
eliminate a
potential buildup of these materials in this recycle loop.
It is possible to add an extraction column to the oxidation system where heavy
phase
from the discharge of the First Stage Oxidizer Oil Decanter (101B) is
contacted with fresh
gas oil. Acetic acid in the heavy phase will be extracted into the gas oil.
This reduces the
amount of acetic acid that must be vaporized in the Water Flash Vessel (108B),
thereby
reducing steam consumption. In addition, the amount of acetic acid processed
through the
Solvent Purification Column (139) will be reduced. An added benefit could be
the recovery
of a portion of the unreacted peracetic acid leaving the First Stage Oxidizer.
In the simulated design, crude solvent from the Solvent Flash Vessel (136) is
used to
saturate the fresh gas oil feed. This crude solvent contains a substantial
amount (about 4.4
wt%) of aromatic hydrocarbons. These aromatic hydrocarbons are susceptible to
chemical
attack by the oxidant, and therefore, could cause additional oxidant and gas
oil losses. The
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acetic acid from the bottom of the Solvent Purification Column (139) has a
lower aromatic
content (about I .I wt%). Therefore, it is possible to use the acetic acid
from the bottom of
the Solvent Purification Column to saturate the fresh gas oil feed, and it is
possible to use all
the crude acetic acid from the Solvent Flash Vessel (136) for feeding the
Sulfox Extraction
Column (119).
An additional improvement may be possible by replacing the stream jet system
of the
Hydrocarbon Recovery Column (152) with a liquid ring vacuum pump for the
Hydrocarbon
Recovery Column (152). If fresh gas oil can be used as the vacuum pump cooling
fluid, it
may be possible to reduce the refrigeration requirements for an alternative
Chiller System and
simultaneously reduce the loses of acetic acid. Instead of utilizing 0 F (-
17.8 C) -10 F (-
12.2 C) brine, 40 F (4.4 C) chilled water may be possible. In the best case, a
Chiller System
and the Hydrocarbon Column Vent Condenser (156) would be eliminated entirely.
It is
hoped that the gas oil absorbs the acetic acid from the vent stream. Once this
acetic acid is
absorbed, it can then be feed to the front of the process and recovered. The
maximum
possible acetic acid recovery is 12.0 lb/hr, which is worth about 0.028 USD
per bbl of
product. In addition, steam consumption and wastewater production is reduced.
However,
additional electricity may be necessary.
A mechanically agitated contactor is utilized for the Raffinate Wash Column
(122) in
the simulated process described above. The mechanically agitated contactor is
expensive and
will probably require substantial maintenance. Therefore, it is possible to
replace the
mechanically agitated contactor with a series of mixer/settlers. This should
reduce capital
requirements. With mixer/settlers, it may also be possible to decrease the
wash water
requirements.
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If unoxidized thiophenes co-distill in the Hydrocarbon Recovery Column (152),
the
recycle distillate should be sent to the oxidation system rather than the
Raffinate Wash
Column (122).
One may consider adding a feed vaporizer to the Solvent Flash Vessel (136) due
to
large difference in bubble points between feed and bottoms liquid.
It may be possible to delete the Solvent Recovery Column Reflux Drum (147) and
rflux the top of the Solvent Recovery Column (149) directly from the Solvent
Recovery
Column Overhead Condenser (146). A hydraulic study is necessary to determine
the
feasibility of this cost savings idea. It may also be possible to delete 148.
It may be possible to delete the Destruct Reactor (112). During the continuous
flow
pilot testing, the oxidant level leaving the second Stage Oxidation was
monitored. The
concentration of active oxidant was very low. If the Destruct Reactor (112) is
removed, most
of the active oxygen remaining in the gas oil should be removed in the Sulfox
Extraction
Column (119) by the acetic acid extraction solvent. The solvent stream leaving
the Sulfox
Extraction Column (119) flows to the Solvent Flash Vessel (136) where the high
temperature
will certainly destroy any remaining active oxygen. However, prior to deleting
the Destruct
Reactor (112), a complete safety study is necessary.
One may consider adding steam to the bottom of the Hydrocarbon Recovery Column
(152). This could allow a higher operating pressure and/or increased recovery
of
hydrocarbon.
There are currently five heat exchangers that cool process liquids with
cooling water.
These are: 118 (3.331 mmbtu/hr (3.5 14 MJ/hr)), 120 (1.216 mmbtu/hr (1.283
MJ/hr)), 145
(0.368 mmbtu/hr (0.3882 MJ/hr)), 155 (0.819 mmbtu/hr (0.864 MJ/hr)), and a
solvent hold
tank cooler (0.857 mmbtu/hr (0.9041 MJ/hr)). The heat duties for these five
exchangers sum
to a total load of 6.6 mmbtu/hr (6.963 MJ/hr). This is worth approximately
0.21 usd/bbl of
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feed or 0.23 usd/bbl of product. It may be desirable to utilize additional
process/process
interchanger, in order to recover some of the wasted energy.
The Solvent Purification Column Overhead Condenser (141) has a heat duty of
26.6
mmbtu/hr (28.063 MJ/hr). It may be possible to recover a large portion of this
energy by
increasing the operating pressure of the Solvent Purification Column (139).
However,
increasing this pressure would either increase the size of the Solvent Flash
Vessel Reboiler
(138) or increase the steam pressure requirements for this exchanger.
Potential Advantages
to Based on the disclosure contained herein, it should be apparent that
potential
advantages include:
(1) Two Stage Addition of Oxidant - This assures that the lowest
concentrations
of unoxidized sulfur compounds are in contact with the highest concentrations
of oxidant.
(2) Water Removal between Oxidation Stages - This allows recycle of the heavy
phase leaving the second stage oxidation to the first stage oxidation without
loss of oxidant.
This also eliminates any water dilution affects on the fresh oxidant added to
the second stage
oxidation, and therefore, promotes maximum mass transfer of oxidant from the
heavy phase
to the light phase.
(3) Second Stage Oxidation at Reduced Temperature - This minimizes the
unwanted side reactions in the second stage oxidation, and therefore,
preserves the oxidant
for recycle to the first stage oxidation.
Based on these concepts and the simulated process shown in Figs. 3-7,
extensive
laboratory experimentation proved the viability and repeatability of these
oxidation concepts.
The experiments consistently produced gas oil with a sulfur content less than
10 ppm by
weight.
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It also became apparent that the residence time required for oxidation in the
first stage
is relatively short. Conversions greater than 98 percent were obtained in less
than 5 to 10
minutes. However, the experimental data also indicated that a relatively long
residence time
is required in the second stage oxidation. Long residence time in the second
stage oxidation
results in a large expensive reactor. In addition. this long residence time
also causes
excessive depletion of oxidant via side reactions in the heavy phase.
Based on measurements of the active oxygen concentration in the light phase
during
the second stage oxidation, it became apparent that the reaction mechanism at
low
concentrations of unoxidized sulfur compounds is kinetically controlled rather
than mass
transfer controlled. In addition, it was discovered that the solubility of PAA
in the light phase
is great enough to provide a stoichiometric excess for completing the
oxidation.
With this information, an improved oxidation scheme is disclosed herein (Fig.
8).
A two-stage oxidation with water removal between stages and a lower second
stage
oxidation temperature is still employed. However, in this revised scheme, the
Second Stage
Oxidizer is a single liquid phase plug flow reactor. Mass transfer of the
oxidant to the light
phase is accomplished in a short residence time static mixer placed
immediately up-stream of
the single liquid phase plug flow reactor. A residence time of one to two
minutes is sufficient
to transfer sufficient oxidant from the heavy phase into the light phase. The
lower portion of
the plug flow reactor is used to separate the two phases via gravity settling.
The heavy phase
is immediately recycled to the First Stage Oxidizer. Immediate removal of the
heavy phase
minimizes the extent of side reactions, and therefore. maximizes the amount of
recycle
oxidant. The isolated light phase flows through the plug flow reactor, where
residence times
can be made arbitrarily long without an excessive cost impact.
Since the residence time in the Second Stage Oxidizer is shorter than the time
required for sufficient in situ conversion of hydrogen peroxide to peracetic
acid, a
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continuously stirred tank reactor (CSTR) is added to the oxidation system. The
PAA Reactor
(171) is used to pre-form the PAA from fresh 70 wt% hydrogen peroxide and
recycle acetic
acid. The fresh catalyst necessary to replace the sulfuric acid purged from
the oxidation
system via Water Flash Vessel (108A) is introduced through the PAA Reactor.
Although a
CSTR was chosen for the application. it is technically feasible to utilize a
simple plug flow
reactor instead. The CSTR is expected to cost more than the plug flow reactor,
but it does
offer an easier mode of operation. especially with respect to startup of the
oxidation system.
Based on the short residence time requirements for the first stage oxidation,
the
relatively expensive, high maintenance mechanically agitated contactor is
replaced with a
plug flow pipe reactor equipped with an internal static mixer. This is
expected to reduce
capital costs and maintenance costs.
The continuous flow pilot testing results indicate that the oxidation system
in the
invention process can consistently produce gas oil with less than 25 ppm by
weight of
unoxidized sulfur compounds; although, it is possible to achieve a gas oil
with a lower sulfur
content.
Obviously, numerous modifications and variations on the present invention are
possible in light of the above teachings. The scope of the claims should not
be limited by the
preferred embodiments set forth in the examples, but should be given the
broadest
interpretation consistent with the description as a whole.
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