U.S. patent application number 10/115463 was filed with the patent office on 2003-09-11 for method for increasing the efficiency of a gasification process for halogenated materials.
Invention is credited to Galloway, Connie M., Henley, John P., Jewell, Dennis W..
Application Number | 20030167692 10/115463 |
Document ID | / |
Family ID | 27789288 |
Filed Date | 2003-09-11 |
United States Patent
Application |
20030167692 |
Kind Code |
A1 |
Jewell, Dennis W. ; et
al. |
September 11, 2003 |
Method for increasing the efficiency of a gasification process for
halogenated materials
Abstract
Methods for improving a gasification process for halogenated
materials and in particular for producing useful end products such
as anhydrous or highly concentrated hydrogen halides and/or
synthesis gas, the methods including recycling water/hydrogen
halide vapors and/or carbon dioxide to a gasification reactor.
Inventors: |
Jewell, Dennis W.;
(Angleton, TX) ; Henley, John P.; (Midland,
MI) ; Galloway, Connie M.; (Angleton, TX) |
Correspondence
Address: |
THE DOW CHEMICAL COMPANY
INTELLECTUAL PROPERTY SECTION
P. O. BOX 1967
MIDLAND
MI
48641-1967
US
|
Family ID: |
27789288 |
Appl. No.: |
10/115463 |
Filed: |
April 2, 2002 |
Related U.S. Patent Documents
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Application
Number |
Filing Date |
Patent Number |
|
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10115463 |
Apr 2, 2002 |
|
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09566183 |
May 5, 2000 |
|
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Current U.S.
Class: |
48/197FM ;
422/184.1; 422/187; 423/481; 423/650; 423/651; 48/199FM |
Current CPC
Class: |
C10J 2300/0973 20130101;
C10J 2300/1223 20130101; C10J 3/466 20130101; C01B 3/32 20130101;
C10J 2300/09 20130101; C10K 1/003 20130101; C01B 7/01 20130101;
C10J 2300/0969 20130101; C10K 1/10 20130101; C10J 2300/169
20130101; C10J 2300/1807 20130101; C10J 2200/09 20130101; C01B
2203/1211 20130101 |
Class at
Publication: |
48/197.0FM ;
423/481; 48/199.0FM; 423/650; 423/651; 422/184.1; 422/187 |
International
Class: |
C10J 003/02 |
Claims
What is claimed is:
1. An improved method for a gasification process for halogenated
materials, comprising: drawing water/hydrogen halide vapors from a
distillation stage of the gasification process; and recycling the
vapor as a reactant and/or moderator feed to a gasification reactor
stage of the process.
2. The method of claim 1 that includes managing the pressure,
temperature and flow rate of the water/hydrogen halide vapor to
control process water balance, to lower carbon particle soot output
and to moderate flame temperature in the gasification reactor.
3. The method of claim 1 wherein the gasification process includes
at least a gasification reactor stage in fluid communication with a
quench stage, in fluid communication with an absorber stage, in
fluid communication with a distillation stage.
4. The method of claim 1 that includes adding carbon dioxide as an
additional reactor and/or moderator gas to the gasification reactor
stage.
5. The method of claim 1 that includes heating a drawn vapor prior
to recycling the vapor.
6. An improved method for a gasification process for halogenated
materials, comprising capturing carbon dioxide from synthesis gas
produced by gasification of halogenated materials; and feeding the
carbon dioxide as a reactant and/or moderator gas to a gasification
reactor stage of the process.
7. The method of claim 6 wherein the gasification process includes
at least a gasification reactor stage in fluid communication with a
quench stage, in fluid communication with a quench stage, in fluid
communication with an absorber stage, in fluid communication with a
distillation stage.
8. Apparatus for improving a gasification process for halogenating
materials, comprising: a gasifier, the gasifier in fluid
communication with a source of halogenated materials; an absorber
in fluid communication with the gasifier; a distillation unit in
fluid communication with the absorber; and a conduit providing
fluid communication from a vapor-draw from the distillation unit to
the gasifier.
9. The apparatus of claim 8 wherein the gasifier is in fluid
communication with a source of oxygen.
10. The gasifier of claim 8 wherein the absorber is in fluid
communication with a source of liquid hydrogen halide.
11. The apparatus of claim 8 that includes a conduit attached to a
heater for heating the vapor.
12. The apparatus of claim 8 that includes a quench in fluid
communication with, and between, the gasifier and the absorber.
Description
CROSS-REFERENCE TO RELATED APPLICATIONS
[0001] This application is a continuation of U.S. patent
application Ser. No. 09/566,183, filed May 5, 2000.
FIELD OF THE INVENTION
[0002] The invention relates to apparatus and methods to be
utilized for the gasification of halogenated materials, and in
particular to apparatus and methods that efficiently produce useful
end products such as anhydrous or highly concentrated hydrogen
halide and/or synthesis gas.
BACKGROUND OF THE INVENTION
[0003] Related inventions include a prior patent application for a
Method and Apparatus for the Production of One or More Useful
Products from Lesser Value Halogenated Materials, PCT international
application PCT/US/98/26298, published 1 Jul. 1999, international
publication number WO 99/32937. The PCT application discloses
processes and apparatus for converting a feed that is substantially
comprised of halogenated materials, especially by-product and waste
chlorinated hydrocarbons as they are produced from a variety of
chemical manufacturing processes, to one or more "higher value
products" via a partial oxidation reforming step in a gasification
reactor. Other related inventions include six co-filed applications
for certain other aspects of the process for gasifying halogenated
material, such aspects including apparatus and methods for reactor
vessel designs, gasifier nozzle designs, controlling aerosols,
producing high quality acids, particulate removal and quench vessel
designs.
[0004] A gasification reaction process for halogenated materials is
a technology for consuming halogenated material byproducts and
waste streams, most likely liquid chlorinated organic byproducts
and waste streams, and for producing substantially useful products
therefrom. Successful implementation of the technology may replace
liquid thermal oxidation facilities which represent the current
industry technique for treating such waste and byproduct streams.
Gasification offers several advantages over thermal oxidation
including more economic costs, reduced emissions and the capture of
maximal chemical value from the feed stream constituents.
Gasification is also more flexible than the competing technologies
in that it has a significantly broader range of acceptable
feedstock composition.
[0005] In a gasification process for halogenated materials, a
source of oxygen (in gaseous form) is mixed with a source of one or
more halogenated material feeds (typically in liquid form and pre
treated or pre processed if necessary or desirable), the mixture
taking place in at least one gasification reactor to produce
syngas. The syngas typically comprises a hydrogen halide, CO and
H.sub.2 with residual C, CO.sub.2, H.sub.2O and trace elements.
[0006] Such gasification in a reactor occurs at partial oxidation
conditions, i.e. at oxygen to fuel ratios that are
substoichiometric with reference to complete combustion. Under such
conditions carbon particles or soot can be formed as a side
product. This soot requires additional capture and treatment steps
downstream in the process, thereby decreasing the economic
efficiency of the process as a whole. One goal of the instant
invention is to operate the gasification process more optimally
while managing parameters such as pressure, temperature, reactants
and flow rates so that the production of C (carbon particles or
soot), CO.sub.2 and H.sub.2O is minimized. Higher oxygen to fuel
ratios can reduce the formation of soot. However, oxygen to fuel
ratios are limited by permissible flame temperatures.
[0007] In various processes for gasifying essentially
hydrocarbonaceous fuels or waste products, steam is known to be
used as a gasifying agent. Under suitable conditions steam is known
to react with carbon (or carbonaceous waste products or soot) to
convert the carbon to carbon monoxide and the steam to hydrogen,
both carbon monoxide and hydrogen being desirable products. Steam
is also known to be used as a "moderator" in regard to several
functions in the environment of gasifying hydrocarbonaceous
materials. The addition of steam "moderates" flame temperatures,
allowing higher oxygen to fuel ratios to be utilized. Higher oxygen
to fuel ratios, as mentioned above, can reduce the formation of
soot due to a higher partial pressure of oxygen.
[0008] Steam is also known to be used in gasification processes for
essentially hydrocabonaceous materials for adjusting the hydrogen
to carbon monoxide ratio of a product synthesis gas to meet the
requirements of downstream customers.
[0009] In the process of the gasification of hydrocarbonaceous
materials, however, unlike in the instant gasification process,
excess water created by used steam can be purged as waste water
from downstream unit operations with a near negligible loss of
valued products. In the gasification process of halogenated organic
materials, the situation is otherwise. While the addition of steam
to the gasification reactor can have the same beneficial effects
mentioned above (of reducing soot and allowing higher oxygen to
fuel operating ratios and supplying additional hydrogen,) the
addition of steam can be wasteful. If the halogenated organic
gasification process includes the production of a hydrogen halide
to an anhydrous form, or even to a highly concentrated aqueous
solution, the purge of the excess water can result in the loss of
valuable product. In both processes, excess steam or water must be
purged from the system downstream to maintain a water balance. In
the case of the production of anhydrous or concentrated hydrogen
halides, the purge step contains a significant concentration of the
hydrogen halide. This loss is in proportion to the amount of steam
moderator furnished to the gasifier.
[0010] The present invention teaches a method to close the water
balance in the halogenated organic gasification process while
significantly minimizing the loss of valuable hydrogen halide
product in an aqueous purge. More particularly, a gasification
process for halogenated materials, if separated hydrogen halide is
anticipated to be sold as an anhydrous product or in a highly
concentrated solution, includes a distillation step to separate
hydrogen halide product from water (in particular from water
absorbed when hydrogen halide gas passes through an absorber
stage). The present invention teaches the use of a vapor side-draw
from the distillation stage wherein water/hydrogen halide vapor is
extracted and recycled to the gasifier as a "moderator" steam
stream. The distillation system can be run at a pressure higher
than the gasifier, thereby providing pressure to straightforwardly
feed the extracted water/hydrogen halide vapor into the gasifier.
Optionally, to help avoid liquid carryover in the "moderator"
stream to the gasifier, the water/hydrogen halide vapor stream can
be superheated with an appropriate heat source, such as steam, a
heat transfer fluid, or the like.
[0011] The recycled vapor from the distillation step is principally
water vapor but contains significant amounts of hydrogen halide.
The hydrogen halide recycles through the gasifier to be subject to
recapture again in the hydrogen halide recovery stage. The water
vapor or steam is primarily consumed via gas shift reactions and
carbon consuming reactions, discussed above. In such manner, the
water balance of the process is maintained or completed while also
achieving the desired objective of soot reduction. A combination of
steam as well as recycled vapor can be utilized in whatever ratio
needed in order to match and achieve the process water balance, as
necessary. Recycled vapor with or without steam can also be used to
adjust H.sub.2 to CO ratio of the product syngas.
[0012] Moderator streams are typically supplied to a gasification
reactor through a suitably designed burner for intimate and
appropriate mixing of all reactants. Lipp et al. describes one such
burner system in a co-filed and co-pending patent application
entitled Method and Apparatus for a Feed Nozzle for a Gasification
Reactor for Halogenated Materials.
[0013] An alternate methodology of the present invention teaches
the use of another moderator, either together with or in lieu of
the water/hydrogen halide vapor moderator, for helping to drive and
to maintain the water balance of the gasification reactor process.
As discussed above, synthesis gas created from the gasification of
halogenated organics contains carbon dioxide. Methods for the
removal and capture of carbon dioxide from synthesis gas are known.
Carbon dioxide has some of the same reforming tendencies as steam.
That is, carbon dioxide reacts with carbon and soot particles to
produce carbon monoxide at gasification conditions. It is another
aspect of the present invention that the carbon dioxide produced in
the synthesis gas reaction can be captured and recycled as an
alternate or further moderator, augmenting or displacing steam.
Some water vapors are produced due to the gas shift reaction, e.g.
CO+H.sub.2O<-->CO.sub.2+H.sub.2. The use of carbon dioxide as
a moderator and/or a combination of steam and carbon dioxide thus
further allows the process water balance to be managed without
purging or losing hydrogen halide in an aqueous discharge. It can
also be used to adjust H.sub.2 to CO ratio in the product syngas.
Depending on the operating pressure of the carbon dioxide recovery
system, carbon dioxide can be pressured back to a gasifier reactor
or a compression operation can be included for pressurizing the
CO.sub.2 stream to suitable pressures for feed to the gasifier.
Alternately, carbon dioxide can be purchased and stored as a
commodity. Carbon dioxide, thus stored can be supplied at
appropriate pressure to the gasifier.
[0014] As discussed above, while it is known in current
gasification practice for conventional hydrocarbonaceous materials
to use steam, and to a lesser extent carbon dioxide, to minimize
soot formation and to adjust hydrogen to carbon monoxide ratios in
the product syngas for intended consumers, the instant invention
improves upon the above in that the "moderator" is or can be a
recycled process fluid. Such use of the recycled process fluid
prevents loss otherwise of hydrogen halide mixed into a purged
water vapor process fluid. Using a recycled water/hydrogen halide
vapor as a moderator provides a means for controlling the water
balance of the process with the additional advantage of minimizing
the aqueous waste volume discharged from the plant and minimizing
the loss of product. As a further advantage, by providing a method
for managing water balance, using a recycled water/hydrogen halide
vapor as a moderator permits the use of higher water addition rates
to a hydrogen halide absorption column. Use of higher water
addition rates to a hydrogen halide absorption column for the
synthesis gas creates a higher recovery efficiency of hydrogen
halide.
[0015] Recycling the purged water/hydrogen halide vapor as a
moderator should have the further advantage of also permitting
efficient utilization of a wider array of feed stock compositions.
That is, feed stocks with a lower halide content can be processed
while still producing anhydrous or highly concentrated hydrogen
halide product since the loss of hydrogen halide has been lowered.
Said otherwise, without use of the water/hydrogen halide vapor as a
recycled moderator in the gasification reactor, recovered aqueous
hydrogen halide from low halide feed concentration materials might
be unsuitable for anhydrous recovery because of the otherwise
excessive halide loss through aqueous discharge.
[0016] The instant invention has a further advantage of requiring
no additional significant equipment, except perhaps a vapor
superheater. Generation of the water/hydrogen halide vapor and its
recycling can be easily integrated into the distilling system.
Whether anhydrous or aqueous hydrogen halide product is desired,
recycling water/hydrogen halide vapor from a distillation stage
allows the production of more concentrated solutions by managing
water balance without loss of product. Further, for feed stocks
lean on hydrogen, the recycled water/hydrogen halide vapor serves
as an additional source of hydrogen for converting all halide to a
hydrogen halide component.
SUMMARY OF THE INVENTION
[0017] The present invention offers improved methods for a
gasification process for halogenated materials. The improvements
include one or more of the following goals: increasing the
efficiency of the process; increasing and/or maximizing the
anhydrous hydrogen halide recovery; minimizing the aqueous
discharge; and adjusting the H.sub.2 to CO ratio.
[0018] The present invention includes apparatus and methods for
increasing the efficiency of a gasification process for halogenated
materials. The invention in one embodiment includes removing
water/hydrogen halide vapors from a distillation stage of a
gasification process and recycling the vapor as a reactant and/or
moderator feed to a gasification reactor stage of the process.
[0019] The method includes managing pressure, temperature and flow
rate of the water/hydrogen halide vapor to control water balance,
to lower carbon particle soot output and to moderate flame
temperature in the gasification reactor. The method and apparatus
include alternately or additionally capturing carbon dioxide from
synthesis gas produced by a gasification of halogenated materials,
or otherwise securing carbon dioxide, and feeding the carbon
dioxide as a reactant and/or moderator gas to a gasification
reactor stage of the process. The carbon dioxide may be added in
addition to or in lieu of a water/hydrogen halide vapor
moderator.
BRIEF DESCRIPTION OF THE DRAWINGS
[0020] A better understanding of the present invention can be
obtained when the following detailed description of the preferred
embodiment is considered in conjunction with the following
drawings, in which:
[0021] FIGS. 1A and 1B illustrate block flow diagrams for a
gasification process for halogenated materials; FIG. 1A illustrates
recycling water/hydrogen halide vapor while FIG. 1B illustrates
recycling captured CO.sub.2.
[0022] FIGS. 2A and 2B illustrate in more detail a gasifier stage
for a gasification process of FIGS. 1.
[0023] FIG. 3 illustrates a quench and solids removal stage of a
gasification process of FIG. 1.
[0024] FIGS. 4A and 4B illustrate an absorber and an aqueous acid
cleanup stage of a gasification process of FIGS. 1.
[0025] FIG. 5 illustrates an anhydrous distillation stage of a
gasification process of FIGS. 1.
[0026] Tables 1A and 1B illustrate a numerical simulation of a run
of a gasification reactor for halogenated materials demonstrating
sensitivity of the outlet gas composition to varying the moderator
flow rate.
[0027] Tables 2A, 2B and 2C illustrate parameters for a
commercially available system for the capture of carbon dioxide
from syngas.
[0028] Tables 3A-3E show, from a mathematical model, heat and
material balances, demonstrating balancing of the water across a
plant by recycling of the water.
DETAILED DESCRIPTION OF PREFERRED EMBODIMENTS
[0029] An embodiment for a gasification process for halogenated
materials is indicated in block diagram form in FIGS. 1A and 1B.
FIG. 1A illustrates an embodiment of the process in which
water/hydrogen halide vapors 530 (assumed for the purpose of the
embodiment to be H.sub.2O/HCL) are recycled from a distillation
unit 500 back to a gasifier 200, in a first aspect of the present
invention. FIG. 1B illustrates an embodiment of the invention
wherein syngas produced from a gasifier 200 is finished in a gas
finishing stage 700 and is further processed in a CO.sub.2 recovery
stage 700', for carbon dioxide recovery and the carbon dioxide
stream 730 is recycled back to gasifier 200. Of course, a preferred
embodiment could provide for both recycling of carbon dioxide and
water/hydrogen halide vapors. Further, carbon dioxide could be
purchased and/or stored as opposed to, or in addition to being
captured in a recovery stage.
[0030] More particularly, FIGS. 2-5 illustrate in more detail
aspects of an embodiment of a gasification process for halogenated
materials than are indicated in block diagram form in FIGS. 1A and
1B. Elements of FIGS. 2 and 5, in particular, will be discussed in
detail to illustrate preferred embodiments of the instant invention
and to place the instant invention in perspective.
[0031] FIGS. 2A and 2B illustrate a gasifier 200 in accordance with
a preferred embodiment. The particular gasifier design of FIGS. 2A
and 2B has two stages, primary gasifier R-200 and secondary
gasifier R-210 for converting a fuel comprised substantially of
halogenated materials to reaction products including hydrogen
halide and synthesis gas components. For the purpose of this
discussion the halogenated material will be assumed to be comprised
of chlorinated hydrocarbons (RCL's). In FIG. 2A, RCL liquid stream
144 is atomized in primary reactor R-200 with a pure oxygen stream
290 and a steam stream 293, both injected through a main burner or
nozzle BL-200. In the harsh gasification environment inside
gasification reactor R-200, the RCl components are partially
oxidized and converted to synthesis gas (syngas) comprised
primarily of carbon monoxide, hydrogen, hydrogen chloride, and of
lesser amounts of water vapor and carbon dioxide, as well or of
some undesirable side products such as soot or carbon. The syngas
flows into secondary reactor R-210 to allow all reactions to
proceed to completion, thus yielding very high destruction
efficiencies of all species and minimizing undesirable side
products such as soot. The output is syngas stream 210.
[0032] Because of the corrosive nature of HCl, both as a hot, dry
gas and as a condensed liquid, the reactor shells and connecting
conduit are shown jacketed with a closed heat transfer fluid system
for wall temperature control, as indicated in FIGS. 2A and 2B in
combination, which system comprises the subject of a related
invention, filed simultaneously hereto. FIG. 2B illustrates a
temperature control fluid circulating system to control the
temperature of a pressure vessel wall of a gasifier, the fluid
flowing between a pressure vessel wall and a jacket.
[0033] The primary gasifier R-200 of the instant embodiment
functions to atomize the liquid fuel, evaporate the liquid fuel,
and thoroughly mix the fuel with oxygen, moderator, and hot
reaction products. The gasifier burner or nozzle design forms the
subject of a related invention, filed simultaneously hereto. The
gasifier R-200 of the preferred embodiment operates at
approximately 1450.degree. C. and 75 psig. These harsh conditions
insure near complete conversion of all feed components.
[0034] The reactions that take place in gasifier R-200 are many and
complex. The reaction pathways and kinetics are not completely
defined nor understood. Indeed, for the numerous species that
comprise the anticipated gasifier feeds, the multiple reactions and
their kinetics for each will be somewhat different. However,
because of the extreme operating conditions in the gasifier, the
reactions can be fairly represented by the overall reactions as
defined below, in a close approach to equilibrium for most
species.
[0035] RCl Partial Oxidation:
[0036] Chlorinated organics are partially oxidized to CO, H.sub.2
and HCl.
C.sub.vH.sub.wCl.sub.y+(v/2)O.sub.2.fwdarw.(v)CO+[(w-y)/2]H.sub.2=(y)HCl
[0037] However, since the gasifier operates with a slight excess of
oxygen above this stoichiometry, further oxidation occurs. Water
vapor and carbon dioxide can also participate as oxidizers at
gasification conditions.
C.sub.vH.sub.wCl.sub.y+CO.sub.2.fwdarw.(v+1)CO+[(w-y)/2)]H.sub.2+(y)HCl
C.sub.vH.sub.wCl.sub.y+H.sub.2O.fwdarw.(v)CO+[1+(w-y)/2]H.sub.2+(y)HCl
[0038] Further oxidation reactions:
CO+1/2O.sub.2.fwdarw.CO.sub.2
H.sub.2+1/2O.sub.2.fwdarw.H.sub.2O
[0039] The oxidation reactions with oxygen, including the reaction
C.sub.vH.sub.wCl.sub.y+(v/2)O.sub.2.fwdarw.(v)CO+[(w-y)/2]H.sub.2=(y)HCl,
are highly exothermic, and thus provide the energy for driving the
other reactions, maintaining the gasifier temperature as
desired.
[0040] Thermal Decomposition Reactions:
[0041] In local fuel rich zones resulting from the less than
perfect mixing inherent to any burner, thermal decomposition occurs
in the absence of oxygen or oxidizing species.
C.sub.vH.sub.wCl.sub.x.fwdarw.C.sub.r+(x)HCl+(v-r)CH.sub.4+[w-x-4*(v-r)/2]-
H.sub.2
[0042] where C is soot, and methane CH.sub.4 is the simplest
hydrocarbon molecule which is quite stable.
[0043] Gas Shift Reactions:
[0044] CO+H.sub.2OCO.sub.2+H.sub.2, classic gas shift reaction,
driven primarily by gas composition, pressure and temperature have
limited effect within the narrow opening range of the gasifier.
[0045] CH.sub.4+H.sub.2OCO+3H.sub.2, steam--methane reforming
driven almost completely to the right at gasifier conditions.
[0046] Soot is also subject to partial oxidation reactions as
described in paragraph 1 above, excluding the chlorine atom.
[0047] Other Reactions:
[0048] Due to the low partial pressure of oxygen in the gasifier,
essentially all halogens, including chlorine as shown above,
equilibrate to the hydrogen halide.
[0049] Operating temperature in the gasifiers R-200 and R-210
should not be allowed to drop below approximately 1350.degree. C.
Conversion efficiency is reduced at lower temperatures. Because of
accelerated corrosion attack to the refractory system, the gasifier
temperature should not be allowed to exceed 1500.degree. C.
Conversion efficiency is very high at 1450.degree. C. and only
limited gains are made at higher temperatures, not justifying the
accelerated refractory corrosion. Preferably, no RCl or liquid
fluid is introduced to the gasifier until it is preheated to an
acceptable operating temperature. Reactor temperature is actually
controlled on a cascade loop with oxygen/fuel ratio. As described
above, the oxidation reactions provide the heat to drive reactor
temperature. The 0.sub.2/fuel ratio will therefore be increased or
decreased as necessary to adjust reactor temperature to the
targeted value. This ratio must be carefully controlled because of
the sensitivity in using pure oxygen where small increments can
cause significant temperature changes. The control band must also
be limited to approximately one-half of the stoichiometric
oxygen/fuel ratio to insure that the flammable mixture (syngas)
environment in the gasifier is always maintained in a reducing
state. Hazardous deflagrations can occur if excess oxygen is
introduced to the fuel rich reactor chamber. Target oxygen to fuel
ratio for the base feedstock is 0.489 lb of oxygen per 1.0 lb of
liquid fuel. This will of course vary as the feed composition
changes and if moderator flow is varied.
[0050] Not only steam stream 293 but also, or alternately, an
HCl/water vapor mixture stream 530 from a desorber T-510 (FIG. 5)
can be used as moderator flow. The moderator flow can be used to
temper the flame temperature of the pure oxygen/fuel burner. This
moderator can also serve as a coolant flow for the burner.
Depending on the heating value of the liquid fuel, pure oxygen and
the fuel can operate at the target gasifier temperature with
insufficient oxygen to complete the partial oxidation reactions.
This results in decreased conversion efficiency, increased soot. To
correct this deficiency, moderator flow can be increased, thus
permitting additional oxygen while maintaining the target gasifier
temperature within limits. Moderator flow can be increased until
sufficient oxidant is present to complete the desired reactions. In
practice this can be defined by the concentration of fully oxidized
species in the exit gas. For example, CO.sub.2 and H.sub.2O may be
targeted to be no less than 1.0 volume % each in the exit gas, and
values as high as 10-15% vol. may be acceptable for heavy sooting
or poor converting feedstocks. Steam as a moderator flow should be
limited as possible because it does put additional load on the
plant water balance and decreases the concentration of aqueous HCl
absorbed downstream.
[0051] The burner BL-200 is an integral and vital component of a
primary gasifier. The discharge jet from the burner provides a
momentum source for mixing in a primary gasifier. The main burner
should atomize the liquid into this mixing jet. Target atomization
performance might be defined as where 99% of the liquid volume is
of a droplet size of 500 microns or smaller. This should provide
for a sufficient liquid surface area enabling rapid evaporation of
the fuel. Two mechanisms play a role in the atomization in
preferred embodiments. The preferred embodiments form the subject
of a co-filed patent application. In preferred embodiments, liquid
is injected through an annular arrangement of orifices centered
around a central oxygen discharge. Pressure drop through these
orifices initiates coarse atomization of the discrete liquid jets.
The orifices, and thus the liquid jets, are directed to intersect
out in front of the face of the burner, or more specifically, along
the axis of the oxygen discharge, and so intersect with the oxygen
discharge jet. The oxygen discharge jet provides a primary energy
source for atomization. Static pressure of the oxygen is converted
to kinetic energy through the burner nozzle. Preferably the burner
provides a supersonic nozzle and so achieves a maximal velocity.
The velocity differential between gas and liquid provides an
atomization energy which reduces the liquid jet to fine, discrete
droplets. Moderator steam may also be mixed with the oxygen
upstream of the burner in this particular operating mode. Oxygen to
the gasifier is preferably preheated to 120.degree. C. to offset
the temperature drop as oxygen is expanded through a supersonic
atomizing nozzle, thus increasing atomization efficiency.
[0052] To avoid induction of hot reaction chamber products into a
near pure oxygen jet immediately at a burner face, and to avoid the
extreme temperature conditions which result, moderator, or some
portion thereof, can be jetted into the gasifier as an annular film
surrounding the oxygen/fuel jet. This "inert" layer tends to move
the hot oxidizing zone out away from the face of the burner, thus
reducing the heat flux and resulting temperatures on the burner
face.
[0053] FIG. 2B, as mentioned above, illustrates a temperature
control fluid system for reactor vessel wells. This system forms
the subject of a separate co-filed patent application. The system
can operate to control the wall temperature of the pressure vessels
to approximately 200.degree. C., or safely above the dew point of
HCl to avoid condensation and resulting in increased corrosion of
the pressure vessel wall.
[0054] FIGS. 3, 4A and 4B illustrate a quench and solids removal
stage 300 of a preferred embodiment of a gasification process and
an absorber 400 and aqueous acid 450 cleanup stage of a preferred
embodiment of a gasification process. The quench, solids removal
absorber and cleanup stages of the preferred embodiment lead to an
anhydrous distillation stage 500 of FIG. 5, which is of particular
significance to the instant invention. The disclosures of FIGS. 3,
4A and 4B are included for background purposes and
clarification.
[0055] FIG. 5 illustrates features of a preferred embodiment for an
anhydrous distillation process for halogenated materials. The
anhydrous distillation area 500 in general consists basically of a
distillation system, including desorber T-510, with auxiliary
equipment to desorb a hydrogen halide stream, treated herein as an
HCl stream, from an aqueous (hydrogen halide) HCl stream. A
desorber overheads stream 503 in the preferred embodiment of FIG. 5
should comprise essentially a saturated HCl stream (+99 vol. %
HCl). This HCl stream 503 can be further processed in one or more
condensors, E-515 and E-520, and in an anhydrous HCl drying and
compression area 600, including an HCl drying tower T-620. Desorber
bottoms from desorber T-510, stream 501, should comprise an
azeotropic (.about.22 wt. % HCl) aqueous HCl stream which can be
recycled to an HCl recovery absorber, illustrated as stream 554,
where it can be reconcentrated to target aqueous acid strength.
[0056] A hydrogen chloride--water system is a highly non-ideal
mixture. It forms an azeotrope at approximately 20.0 wt. % HCl at
atmospheric pressure. Water has a higher activity coefficient above
this concentration. The azeotrope shifts with pressure, decreasing
(HCl concentration reference) as pressure increases. The azeotrope
is approximately 16.6 wt. % at 59 psig. When an absorber bottoms
stream, 483, 500', enters a desorber T-510 above the azeotropic
concentration in the desorber, HCl is a volatile species and is
fractionated overhead.
[0057] In the preferred embodiment of FIG. 5, aqueous acid from
storage illustrated as stream 483 and referenced in FIG. 4, can be
cross exchanged with the bottoms stream 510 and fed to the HCl
desorber T-510. The feed is preferably introduced between an upper
and lower packed section. The HCl desorber can fractionate HCl
overhead while discharging a weak aqueous HCl stream from the
bottoms. At preferred base design conditions (100 psig, 45.degree.
C. from the secondary condenser E-520) the overheads gas should be
about 96 vol. % HCl, 0.12 vol. % H.sub.20, with small amounts of
noncondensibles--primarily CO.sub.2 and to a lesser extent N.sub.2.
Essentially all of the noncondensibles should be driven overhead in
the desorber. Column bottoms may operate at approximately
175.degree. C., and an acid concentration of about 22 wt. % HCl
could be expected. Condensed liquid from both a primary E-515 and a
secondary E-520 condenser can be collected in a reflux drum D-515
and pumped back as, column reflux. A knock-out drum D-520 after the
secondary condenser can also remove free liquid to help prevent its
carryover into the anhydrous HCl drying system. The column reboiler
E-510 can be driven by 235 lb. steam. Condensate level on the
stream (shell) side of the reboiler can be controlled to manipulate
heat transfer surface area, and thus reboiler duty for the
column.
[0058] When producing anhydrous HCl, as per the present invention,
the water balance is preferably closed by using a sidedraw vapor
514 from a desorber as a moderator for the gasifier. This vapor may
be, for instance, about 59 wt % H.sub.20 and 41 wt. % HCl. When
operating in this mode, the delivery pressure to a gasifier
dictates the operating pressure of the desorber, which is about 100
psig. If no sidedraw vapor is required for the gasifier, operating
column pressure can be reduced to 65-75 psig. The advantage of a
lower operating pressure is cooler bottoms temperature which
results in lower corrosion and permeation rates for the equipment.
Boiling HCl as may exist at the bottoms of the desorber can be very
aggressive, and milder operating conditions are more favorable to
equipment reliability. Bottoms temperature is preferably not
allowed to exceed 185.degree. C. due to limitations of the typical
impregnated graphite materials of reboiler tubes and the typical
Teflon linings for towers and piping.
[0059] The bottoms liquid stream 510, which is cross exchanged with
a desorber feed, can be further cooled to approximately 40.degree.
C. (or by using cooling tower E-550, which may include use of even
sea water) and directed on to a dilute acid drum D-550. This drum
can serve as a surge volume for the weak acid, which can be pumped
back to a middle section of an HCl absorber, illustrated as stream
554, where it absorbs additional HCl. A small blowdown to an
environmental area, illustrated as to neutralizer R-810, can be
used to control contaminant concentrations if these undesirables
(salts, metals, etc.) build up to unacceptable levels.
[0060] The following example, produced by computer model,
illustrates typical parameters of a gasification reactor process
for halogenated materials.
EXAMPLE 1
[0061] The following feeds streams are fed to a gasifier through an
appropriate mixing nozzle:
1 Chlorinated Organic Material 9037 kg/hr Oxygen (99.5% v purity):
4419 kg/hr Recycle Vapor or moderator: 4540 kg/hr [58.8 wt % water
vapor, 41.2 wt % hydrogen chloride]
[0062] The resulting gasification reactions result in a synthesis
gas stream rich in hydrogen chloride.
[0063] In a preferred embodiment of the present invention,
referencing the above example, this stream would be cooled or
quenched and passed through an absorption step where the hydrogen
chloride is recovered in an aqueous solution. This aqueous solution
would be forwarded to a distillation system whose principal purpose
is to distill nearly water free hydrogen chloride as an overhead
product. The distillation tower is preferably operated at a
pressure sufficient to flow side-draw vapor through a superheater,
through a control valve, and through a gasifier mixing nozzle. A
vapor side-draw is preferably extracted from a "reboiler section"
of a distillation tower at a flowrate to complete the plant water
balance. For the above example this would be per the flowrate and
composition described for a gasifier feed. The vapor is preferably
passed through a superheating exchanger imparting typically
10-20.degree. C. superheat to the vapor, to insure that no liquid
droplets remain. This vapor would then be fed to a gasifier mixing
nozzle as a moderator stream.
[0064] Alternatively and/or in addition to the above system, a
synthesis gas which has been absorbed free of bulk hydrogen
chloride, as described above and illustrated as stream 418 in FIG.
4, passes through a finishing system 700, FIG. 1B, where
essentially all hydrogen chloride and other contaminants are
recovered. This clean synthesis gas can then be fed to a
commercially available carbon dioxide removal system, illustrated
as unit 700' in FIG. 1B. Carbon dioxide can be absorbed, as is
known, from the syngas, liberated from any solvent or sorbent,
compressed if necessary, and fed back to a gasifier feed nozzle as
stream 730 in FIG. 1B, also as a moderator.
[0065] FIG. 1B, discussed initially, illustrates in block flow
diagram form the addition of a carbon dioxide recovery unit 700'
after syngas finishing unit 700. Tables 2A, 2B and 2C illustrate a
mathematical model run of a prior art carbon dioxide recovery unit
and illustrate that it is known to recover CO.sub.2 from syngas
streams. FIG. 1B also illustrates a CO.sub.2 recycle stream 730
recycled back and fed to a gasifier 200. The CO.sub.2 would
preferably be fed through a nozzle or burner in a passageway
provided for an inert gas moderator, such as steam.
[0066] Tables 1A and 1B illustrate the mole fractions of exit gas
from the secondary reactor of FIG. 2 in a model run upon varying
the moderator flow rate. The tables chart the breakdown of stream
210 when using a hydrogen halide/steam recycle moderator. The flow
rate in lbs/hr of the moderator stream was varied from 2,000 lbs/hr
to 20,000 lbs/hr. Results by mathematical model were computed with
and without a nitrogen purge. Note the increased oxygen content as
evidenced by decreasing CO.sub.2 and H.sub.2O concentrations as
recycled vapor moderator flow is increased. The higher
concentrations support the formation of soot. Another key factor to
note for the operation is the decreasing fraction of HCN and MCBZ,
for the various moderator flows, indicating more complete
destruction of undesirable species as the moderator flow
increased.
[0067] Tables 3A-3E illustrate, from a mathematical model, the
composition of various streams indicated in FIG. 2--FIG. 5 for a
sample run. The heat and material balances demonstrate balancing of
the water across the plant by recycling of the water.
[0068] The foregoing disclosure and description of the invention
are illustrative and explanatory thereof, and various changes in
the size, shape, and materials, as well as in the details of the
illustrated system may be made without departing from the spirit of
the invention. The invention is claimed using terminology that
depends upon a historic presumption that recitation of a single
element covers one or more, and recitation of two elements covers
two or more, and the like.
2 Exit Gas From Secondary Reactor (Stream #210) using HCI/Steam
recycle Moderator Basis: 6-30-99 RCI feed slate HCI/steam recycle
from Desorber as moderator fluid O2/Fuel varied to control primary
gasifier to 1450 C 1.5 MM Btu/hr heat loss from primary 1.5 MM
Btu/hr heat loss from secondary 1000 #/hr N2 purge flow (except if
noted) 75 psig operating pressure Aspen File Folder:
.backslash.Gasifier with recycle HCI vapor.backslash. 2,000 5,000
No 2 No N2 Moderator Flow (#/hr) 2,000 purge 3,000 5,000 purge
10,000 15,000 20,000 Substream: MIXED Mole Frac WATER 9.43E-06
3.51E-06 1.07E-02 4.54E-02 4.36E-02 1.24E-01 0.190369 0.246732 N2
0.022592 0.000138 0.027186 0.025582 0.001352 0.0223 0.019777
0.017778 O2 0.00E+00 9.39E-21 5.87E-14 1.25E-12 1.03E-12 1.38E-11
4.60E-11 1.06E-10 CO2 6.76E-06 2.41E-06 7.10E-03 2.72E-02 2.60E-02
6.13E-02 0.081936 0.095095 CO 0.496939 0.50605 0.479269 0.429943
0.442537 0.336036 0.2695 0.219925 H2 0.211813 0.219081 0.220143
0.216676 0.224961 0.201889 0.184427 0.166906 HCL 0.249677 0.259451
0.249226 0.249262 0.25546 0.249337 0.2494 0.249446 CL2 6.83E-08
8.51E-08 6.74E-08 7.22E-08 7.20E-08 8.53E-08 9.95E-08 1.15E-07 CL
1.54E-05 1.94E-05 1.56E-05 1.68E-05 1.66E-05 1.94E-05 2.18E-05
2.42E-05 CH4 6.59E-03 6.98E-03 6.30E-03 5.92E-03 6.07E-03 5.15E-03
0.0045531 0.004081 HCN 0.011664 2.60E-03 1.15E-05 2.30E-06 6.00E-07
5.51E-07 2.36E-07 1.21E-07 NH3 3.72E-06 2.90E-07 4.29E-06 4.00E-06
9.76E-07 3.26E-06 2.63E-06 2.12E-06 FORMHYDE 4.77E-07 4.97E-07
4.77E-07 4.20E-07 4.49E-07 3.04E-07 2.22E-07 1.64E-07 NAPTHALN
1.30E-08 1.37E-08 1.24E-08 1.16E-08 1.19E-08 1.01E-08 8.95E-09
8.03E-09 C2HCL5 2.12E-22 2.08E-21 1.53E-28 7.04E-30 0.00E+00
6.03E-31 1.69E-31 6.85E-32 C2H2CL4U 3.43E-18 3.01E-17 2.54E-24
1.11E-25 0.002+00 8.39E-27 0.00E+00 0.00E+00 PERCHLOR 3.65E-16
3.52E-15 2.65E-22 1.23E-23 0.00E+00 1.06E-24 3.01E-25 1.23E-25 CCL4
1.41E-16 5.04E-16 1.17E-19 2.63E-20 0.00E+01 8.77E-21 5.30E-21
3.83E-21 C2H2CL4S 8.06E-18 7.07E-17 5.96E-24 2.61E-25 3.44E-25
1.97E-26 0.00E+00 0.00E+00 TCE 5.20E-12 4.58E-11 3.87E-18 1.72E-19
0.00E+00 1.33E-20 0.00E+00 0.00E+00 C2H3CL3 5.16E-14 4.11E-13
3.91E-20 1.64E-21 0.00E+00 1.10E-22 0.00E+00 0.00E+00 C2CL6
1.17E-26 1.27E-25 8.25E-33 0.00E+00 0.00E+00 0.00E+00 1.23E-35
5.65E-35 PDC 4.48E-13 8.20E-12 3.12E-22 2.46E-24 3.94E-24 3.36E-26
0.00E+00 0.00E+00 EPI 9.57E-19 7.38E-18 7.54E-25 2.71E-26 3.86E-26
1.20E-27 0.00E+00 0.00E+00 DCIPE 5.12E-30 6.03E-28 0.00E+00
0.00E+00 0.00E+00 0.00E+00 0 0 2M2P 1.65E-16 1.88E-14 9.24E-32
0.00E+00 0.00E+00 0.00E+00 0.00E+00 0 123TCP 3.20E-16 6.33E-15
2.17E-25 1.78E-27 0.00E+00 2.70E-29 0.00E+00 0.00E+00 PROPANAL
1.31E-11 8.57E-11 1.04E-17 3.53E-19 0.00E+00 1.34E-20 0.00E+00
0.00E+00 ACETONE 1.77E-11 1.11E-10 1.41E-17 4.70E-19 0.00E+00
1.74E-20 0.00E+00 0.00E+00 BENZENE 2.13E-08 2.25E-08 2.03E-08
1.91E-08 1.91E-08 1.66E-08 1.47E-08 1.32E-08 33DCPENE 1.73E-11
3.48E-10 1.18E-20 9.87E-23 1.51E-22 1.55E-24 0.00E+00 0.00E+00
13DCPENE 9.42E-11 1.87E-09 6.43E-20 5.34E-22 0.00E+00 8.32E-24
0.00E+00 0.00E+00 HXCLBZ 9.29E-23 4.33E-20 0.00E+00 0.00E+00
0.00E+00 0.00E+00 0 0 NCL3 0 7.62E-27 6.08E-26 7.07E-26 1.59E-26
9.71E-26 1.26E-25 1.59E-25 3CLACN 1.65E-13 3.10E-13 1.16E-22
1.05E-24 0.00E+00 2.06E-26 0.00E+00 0.00E+00 1-3BCH 3.76E-21
6.31E-20 2.67E-30 1.76E-32 0.00E+00 0.00E+00 0.00E+00 0 1-2BCH
1.51E-20 2.55E-19 1.08E-29 7.10E-32 0.00E+00 0.00E+00 0.00E+00 0
2-3DCBUT 9.16E-16 4.18E-14 5.70E-28 8.79E-31 1.65E-30 0.00E+00
0.00E+00 0 C4CL6 8.80E-30 6.65E-28 0.00E-00 0.00E+00 0.00E+00
0.00E+00 0 0 SIO2 0 0 0 0 0 0 0 0 SOOT 0 0 0 0 0 0 0 0 BO 1.13E-18
2.02E-17 8.20E-28 5.57E-30 0.00E+00 5.06E-32 0.00E+00 0.00E+00
BGLYCOL 7.88E-26 4.80E-25 6.40E-32 1.79E-33 0.00E+00 0.00E+00
0.00E+00 0 BUTANAL 8.23E-15 1.35E-13 5.88E-24 3.88E-26 6.63E-26
3.37E-28 0.00E+00 0.00E+00 IPROPCL 7.01E-11 1.18E-09 5.02E-20
3.81E-22 0.00E+00 4.63E-24 0.00E+00 0.00E+00 PROPCL 9.59E-11
1.63E-09 6.88E-20 5.24E-22 8.54E-22 6.41E-24 4.56E-25 5.62E-26
[0069]
3 2,000 5,000 No N2 No N2 Moderator Flow (#/hr) 2,000 purge 3,000
5,000 purge 10,000 15,000 20,000 Substream: MIXED Mole Frac 2CLPENE
3.95E-08 7.24E-07 2.78E-17 2.22E-19 0.00E+00 3.08E-21 0.00E+00
0.00E+00 PHENPROP 9.45E-19 1.67E-15 0.00E+00 0.00E+00 0.00E+00
0.00E+00 0 0 MPK 1.18E-11 8.41E-09 4.79E-33 0.00E+00 0.00E+00
0.00E+00 0.00E+00 0 PHENOL 5.82E-10 6.36E-08 3.03E-25 8.48E-29
0.00E+00 5.19E-32 0.00E+00 0.00E+00 AMS 1.25E-08 6.07E-05 4.13E-35
0.00E+00 0.00E+00 0.00E+00 0 0 ODCB 1.12E-08 3.72E-06 5.00E-27
3.43E-31 0.00E+00 0.00E+00 0.00E+00 0.00E+00 PARADOW 1.02E-08
3.35E-06 4.52E-27 3.10E-31 0.00E+00 0.00E+00 0.00E+00 0.00E+00 MCBZ
1.60E-05 4.89E-03 7.33E-24 4.84E-28 0.00E+00 0.00E+00 0.00E+00
0.00E+00 PYRENE 8.21E-09 8.70E-09 7.85E-09 7.38E-09 7.56E-09
6.42E-09 5.67E-09 5.09E-09 133TCPEN 3.64E-14 7.80E-13 2.41E-23
2.08E-25 0.00E+00 3.61E-27 0.00E+00 0.00E+00 ALLYL-CL 4.41E-08
1.24E-07 3.11E-17 2.50E-19 3.88E-19 3.51E-21 0.00E+00 0.00E+00 CBE
0 0 0 0 0 0 0 0 PHENBUTE 8.33E-12 1.01E-07 0.00E+00 0.00E+00
0.00E+00 0.00E+00 0 0 DPHENMP 0.00E+00 7.46E-15 0.00E+00 0.00E+00
0.00E+00 000E+001 0 0 HBR 0 0 0 0 0 0 0 0 BR2 0 0 0 0 0 0 0 0 SICL4
6.53E-04 6.91E-04 2.33E-07 1.34E-09 1.58E-08 1.88E-09 8.25E-10
5.03E-10 FE2O3 0 0 0 0 0 0 0 0 FECL3 5.06E-06 5.36E-06 4.84E-06
4.55E-06 4.66E-06 3.96E-06 3.50E-06 3.14E-06 AL2O3 0 0 0 0 0 0 0
ALCL3 7.93E-06 8.40E-06 7.58E-06 7.13E-06 7.30E-06 6.20E-06
3.70E-06 2.60E-06 CAO 0 0 0 0 0 0 0 0 CACL2 7.21E-06 7.63E-06
6.89E-06 6.48E-06 6.64E-06 6.63E-06 4.98E-06 4.46E-06 BR 0 0 0 0 0
0 0 0 23DCPENE 1.99E-10 3.92E-09 1.36E-19 1.12E-21 0.00E+00
1.74E-23 0.00E+00 0.00E+00 22DCP 6.83E-14 1.24E-12 4.75E-23
3.74E-25 6.00E-25 5.08E-27 0.00E+00 0.00E+00 C5CL5N 4.79E-24
6.55E-23 0.00E+00 0.00E+00 0.00E+00 0.00E+00 0 0 C5H2CL3N 6.57E-16
7.58E-15 3.57E-31 0.00E+00 0.00E+00 0.00E+00 0.00E+00 0 C5HCL4N
2.94E-20 3.71E-19 1.55E-35 0.00E+00 0.00E+00 0.00E+00 0 0 DICLCNPY
3.25E-15 7.59E-15 1.72E-33 0.00E+00 0.00E+00 0.00E+00 0 0 BISETHER
0 0 0 0 0 0 0 0 ACROLEIN 1.27E-08 8.94E-08 9.94E-15 3.52E-16
0.00E+00 1.52E-17 0.00E+00; 0.00E+00 CL2PNOL 4.86E-20 3.50E-19
3.76E-26 1.31E-27 1.93E-27 5.49E-29 0.00E+00 0.00E+00 CH2CL2
2.72E-08 7.96E-08 2.37E-11 4.89E-12 5.73E-12 1.28E-12 6.00E-13
3.38E-13 Total Flow lb mol/hr 1321.261 1247.963 1381.906 1470.352
1434.68 1691.44 1912.531 2133.624 Total Flow lb/hr 32529.66
31126.74 33617.81 35871.65 34762.16 41528.93 47199.24 52887.04
Total Flow gal/min 58921.15 56368.69 61749.02 65956.6 64290.43
76400.18 86786.61 97143.3 Temperature C. 1376.724 1397.981 1380.072
1386.49 1384.792 1397.981 1405.718 1411.342 Pressure psig 74.5 74.5
74.5 74.5 74.5 74.5 74.5 74.5 Vapor Frac 1 1 1 1 1 1 1 1 Liquid
Frac 0 0 0 0 0 0 0 0 Solid Frac 0 0 0 0 0 0 0 0 Enthalpy Btu/lb mol
-14478.5 -14979.1 -16364 -20516.2 -21031.3 -29004.8 -35508.6
-40677.4 Enthalpy Btu/lb -588.075 -600.556 -672.665 -840.944
-867.99 -1181.34 -1438.82 -1641.05 Enthalpy Btu/hr -1.91E+07
-1.87E+07 -2.26E+07 -3.02E+07 -3.02E+07 -4.91E+07 -6.8E+07 -8.7E+07
Entropy Btu/lb mol-R 22.80366 22.93713 22.34239 21.53223 21.63093
19.6209 18.02903 16.70795 Entropy Btu/lb-R 0.926219 0.919617
0.918414 0.88259 0.892736 0.799144 0.730543 0.67405 Density lb
mol/cu ft 2.80E-03 2.76E-03 2.79E-03 2.78E-03 2.78E-03 2.76E-03
0.002747 0.002738 Density lb/cu ft 0.068832 0.068846 0.067877
0.067807 0.061413 0.06777 0.067805 0.067876 Average MW 24.62016
24.94203 24.32713 24.39665 24.22991 24.55241 24.67894 24.78743
[0070]
4 GLOBAL GAS/SPEC TECHNOLOGY GROUP AMINE PLANT PROGRAM DATE: 16
DEC. 1999 SALES: STC NUMBER: 99438UU RUN BY: DUPART COMPANY: DOW
PLANT NAME: SYN GAS PLANT LOCATION: MIDLAND RESULTS GIVEN TO:
HENLEY SOLVENT TYPE: MEA TREATED GAS REQUESTED. 95+% PURITY
HYDROGEN INLET GAS 12.800 MMSCFD,(DRY) 0.000% H2S 100 Deg F.
38.013% CO2 74.7 Psia 0.000% CH4 0.000% C2H6 0.000% C3H8 0.000%
i-C4H10 0.000% n-C4H10 0.000% i-C5H12 0.000% n-C5H12 0.000% C6H14
0.000% C7H16+ 0.000% Ar 0.000% CO 0.000% N2 60.723% H2 1.264% H2O
RUN CONDITIONS TREATED GAS H2S CONCENTRATION, WET BASIS 0.00 PPMV
TREATED GAS CO2 CONCENTRATION, WET BASIS 0.1476 VOL % BAROMETRIC
PRESSURE 14.70 Psia ABSORBER BOTTOM PRESSURE 74.70 Psia ABSORBER
LEAN AMINE FEED TEMPERATURE 110 F. ABSORBER OVERHEAD TEMPERATURE
218 F. ABSORBER BOTTOMS TEMPERATURE 136 F. STRIPPER BOTTOMS
PRESSURE 12.00 Psig STRIPPER REFLUX RATIO 2.00 Mol/Mol STRIPPER
OVERHEAD TEMPERATURE 218 F. STRIPPER FEED TEMPERATURE 205 F.
STRIPPER BOTTOMS TEMPERATURE 246 F. ACID GAS TEMPERATURE EXITING
CONDENSER 120 F. LEAN COOLER INLET TEMPERATURE 176 F. COOLING WATER
INLET/OUTLET 90 F. 110 F. AIR COOLING INLET/OUTLET 90 F. 110 F.
[0071]
5 GLOBAL GAS/SPEC TECHNOLOGY GROUP PLANT PROGRAM AMINE VARIABLES
CO2 LEAN SOLVENT LOADING 0.100 Mol/Mol CO2 NET SOLVENT LOADING
0.251 Mol/Mol CO2 GROSS SOLVENT LOADING 0.351 Mol/Mol H2S LEAN
SOLVENT LOADING 0.000 Mol/Mol H2S NET SOLVENT LOADING 0.000 Mol/Mol
H2S GROSS SOLVENT LOADING 0.000 Mol/Mol SOLVENT CONCENTRATION 15 wt
% SOLVENT CIRCULATION RATE 1750.0 USGPM GAS FLOW/PERFORMANCE DATA
INLET CO2 PARTIAL PRESSURE 1487.29 mmHg INLET H2S PARTIAL PRESSURE
0.00 mmHg NET CO2 REMOVAL 539.91 lbmole/hr NET H2S REMOVAL 0.00
lbmole/hr PERCENT CO2 REMOVED/SLIPPED 99.76 / 0.24 ENERGY BALANCE
REBOILER DUTY 77.684 MMBTU/hr STRIPPER CONDENSER DUTY 22.069
MMBTU/hr CROSS EXCHANGER DUTY 60.428 MMBTU/hr LEAN SOLVENT COOLER
DUTY 56.351 MMBTU/hr UNIT REBOILER DUTY 143884 BTU/lbmole A G.
ESTIMATED EQUIPMENT SIZES ABSORBER DIAMETER (ESTIMATE FOR TRAYED)
7.6 ft ABSORBER TRAY SPACING/No. OF TRAYS 24 in / 20 ABSORBER
DIAMETER: 1.5 IN. METAL PALL RINGS 5.6 ft CARBON: (% Slip/Vol.
(ft{circumflex over ( )}3)/Dia. (ft)) 10 / 467.88 / 7.46 STRIPPER
DIAMETER (ESTIMATE FOR TRAYED) 13.0 ft STRIPPER TRAY SPACING/No. OF
TRAYS 24 in / 20 STRIPPER DIAMETER: 1.5 IN. METAL PALL RINGS 9.5 ft
HEAT EXCHANGE EQUIPMENT: AREA (ft{circumflex over ( )}2) LMTD (F)
Uo (BTU/hr*ft{circumflex over ( )}2*F) CROSS EXCHANGER 12434 40.5
120 REBOILER 9996 51.8 150 LEAN COOLER (WATER) 11278 38.4 130 LEAN
COOLER (AIR) 13329 38.4 110 REFLUX CONDENSER (WATER) 3297 60.8 110
REFLUX CONDENSER (AIR) 4836 60.8 75 LEAN COOLER H2O (USGM) 5640.7
LEAN COOLER FAN (hp) 435.0 REFL. COND. H2O (USGPM) 2204.6 REFL.
COND. FAN (hp) 170.3
[0072]
6 GLOBAL GAS/SPEC TECHNOLOGY GROUP AMINE PLANT PROGRAM PUMPING
EQUIPMENT ESTIMATES AMINE CHARGE PUMP (P.D.) 68.07 hp AMINE CHARGE
PUMP (CENTR.) 102.10 hp AMINE BOOSTER PUMP 105.00 hp REFLUX WATER
PUMP 2.26 hp STREAM CONDITIONS INLET SALES ACID GAS STREAMS
(lbmole/hr) GAS GAS GAS TEMPERATURE, Deg F. 100 110 120 PRESSURE,
Psia 74.70 72.70 22.70 Ar 0.00 0.00 0.00 H2 864.52 862.92 1.60 N2
0.00 0.00 0.00 CO 0.00 0.00 0.00 CH4 0.00 0.00 0.00 C2H6 0.00 0.00
0.00 C3H8 0.00 0.00 0.00 C4H10+ 0.00 0.00 0.00 CO2 541.21 1.30
539.91 H2S 0.00 0.0000 0.00 H2O 18.00 15.41 43.34 TOTAL (lbmole/hr)
1423.72 879.64 564.84 TOTAL (M lb/hr) 25.89 2.07 24.55 DENSITY
(lb/ft{circumflex over ( )}3) 0.222 0.029 0.154 ACTUAL
ft{circumflex over ( )}3/min 1942.10 1210.90 2664.27 LEAN RICH
SOLVENT STREAMS (lbmol/hr) AMINE AMINE TEMPERATURE, Deg F. 110 136
PRESSURE, Psia 72.7 74.7 Ar 0.00 0.00 H2 0.00 1.60 N2 0.00 0.00 CO
0.00 0.00 CH4 0.00 0.00 C2H6 0.00 0.00 C3H8 0.00 0.00 C4H10+ 0.00
0.00 CO2 214.71 754.62 H2S 0.00 0.00 H2O 41254.34 41256.92 MEA
2147.09 2147.09 TOTAL (lbmole/hr) 43616.14 44160.23 TOTAL (M lb/hr)
883.80 907.63 DENSITY (lb/ft{circumflex over ( )}3) 62.96 64.20
USGPM 1750.00 1762.54 USGPM MAKE-UP RE FLUX H2O FLOW 1.47 37.72
[0073]
7 Stream Number: 144 200 210 280 281 282 285 290 291 293 295 296
299 530 733 To: R-3200 R-3210 P-3280 R-4 E-3280 D-3280 E-3200
R-3200 R-3200 R-3200 R-3206 E-3280 R-3200 R-3200 From E-31-10
R-3200 D-3280 P-3280 Q3200 E-3280 E-3280 LIQUID VAPOR VAPOR LIQUID
LIQUID LIQUID LIQUID VAPOR VAPOR LIQUID VAPOR VAPOR LIQUID VAPOR
VAPOR Substream: MIXED Mole Frac WATER 0 0.1706242 0.1683706 0 0 0
0 0 0 1 0 0 0 0.743 0.01 N2 3.29E-04 1.17E-03 1.17E-03 0 0 0 0
5.00E-03 5.00E-03 0 1 0 0 0 3.00E-03 O2 0 7.05E-11 2.25E-11 0 0 0 0
0.995 0.995 0 0 0 0 0 0 CO2 0 0.0700694 0.0723234 0 0 0 0 0 0 0 0 0
0 0 0.116 CO 0 0.3104558 0.3082029 0 0 0 0 3 0 0 0 0 0 0 0.34 H2 0
0.2112548 0.2139049 0 0 0 0 0 0 0 0 0 0 0 0.33 CL2 0 1.05E-07
5.86E-08 0 0 0 0 0 0 0 0 0 0 0.257 0 CL 0 2.86E-05 1.75E-05 0 0 0 0
3 0 0 0 0 0 0 0 CH4 0 1.35E-03 1.35E-03 0 0 0 0 3 0 0 0 1 0 0
1.00E-03 HCN 0 9.96E-08 9.25E-05 0 0 0 0 0 0 0 0 0 0 0 0 NH3 0
8.11E-07 8.13E-07 0 0 0 0 0 0 0 0 0 0 0 0 FORMHYDE 0 2.89E-07
2.95E-07 0 0 0 0 0 0 0 0 0 0 0 0 BENZENE 1.49E-03 1.53E-07 1.53E-07
0 0 0 0 0 0 0 0 0 0 0 0 SOOT 0 4.50E-03 4.50E-03 0 0 0 0 0 0 0 0 0
0 0 0 FECL3 4.00E-05 3.59E-06 3.9E-06 0 0 0 0 0 0 0 0 0 0 0 0
DOWTH-RP 4.00E-05 3.59E-06 3.89E-06 0 0 0 0 0 0 0 0 0 0 0 0 Mass
Flow LB/HR 0 0 0 1 1 1 1 0 0 0 0 0 1 0 0 WATER 0 5115.789 5048.189
0 0 0 0 0 0 1.00E-03 0 0 0 5802.188 1201.463 N2 1.250002 54.3275
54.32505 0 0 0 0 41.92633 41.02633 0 1.00E-03 0 0 0 7937.243 O2 0
3.75E-06 1.20E-06 0 0 0 0 9530.3 9530.3 0 0 0 0 0 0 CO2 0 5132.189
5297.33 0 0 0 0 0 0 0 0 0 0 0 462.1573 CO 0 14472.71 14367.61 0 0 0
0 0 0 0 0 0 0 0 1428.547 H2 0 709.7647 716.3099 0 0 0 0 0 0 0 0 0 0
0 62.82928 HCL 0 13989.56 13990.23 0 0 0 0 0 0 0 0 0 0 4137.012 0
CL2 0 13989.56 13990.23 0 0 0 0 0 0 0 0 0 0 4107.012 0 CL 0
1.085736 1.033508 0 0 0 0 0 0 0 0 0 0 0 0 CH4 0 33.10844 36.10644 0
0 0 0 0 0 0 0 1.00E-00 0 0 1.51507 HCN 0 4.03E-03 4.16E-03 0 0 0 0
0 0 0 0 0 0 0 0 NH3 0 0.0201496 0.0230315 0 0 0 0 0 0 0 0 0 0 0 0
FORMHYDE 0 0.0144203 0.0147523 0 0 0 0 0 0 0 0 0 0 0 0 NAPTHALN 0
0.0158562 0.0158562 0 0 0 0 0 0 0 0 0 0 0 0 BENZENE 16.37100
0.0190582 0.0190582 0 0 0 0 0 0 0 0 0 0 0 0 SOOT 0 90.03299
90.03299 0 0 0 0 0 0 0 0 0 0 0 0 FECL3 1.050234 1.050234 1.050234 0
0 0 0 0 0 0 0 0 0 0 0 DOWTH-RP 0 0 0 5.00E+05 5.00E+05 5.00E+05
5.00E+05 0 0 0 0 0 5 0 0 ACI Components 2003.8033 Total flow LB/
140.6467 1664.297 1664.289 2115.487 2115.767 2115.446 2115.467
299.3297 299.3297 5.55E-05 3.57E-05 6.25E-08 0.0211544 439.4496
94.44552 MOL/HOUR Total Flow LB/HR 18005.1 39802.32 39802.32
5.00E+05 5.00E+05 5.00E+05 5.00E+05 9572.227 9572.227 1.00E-03
1.00E-03 1.00E-06 5 10000 2000 Total Flow GAL/ 35.60107 77079.35
75149.24 1108.471 1108.515 1114.251 1108.471 1897/831 2793.588
2.31E-06 2.27E-04 2.19E-07 9.75E-03 4667.781 334.4008 MIN
Temperature C. 150 1450 1397.446 134.7998 200.0458 205.9968 200 30
150 204.4441 30 30 30 200 195 Pressure PSIG 239 75 74.5 15 50 135
15 100 95 235 100 160.3569 85.30405 90 300 Vapor frac 0 1 1 0 0 0 0
1 1 0 1 1 0 1 1 Enthalpy BTU/HR 0.09E+06 5.12E+07 5.27E+07 6.71E+07
6.71E+07 7.01E+07 6.71E+07 11335.69 4.88E-05 -0.513492 1.54E-03
-2.01E-03 -9.871647 -3.74E+07 -4.18E+00 Density LB/CU FT 63.05406
0.0640565 0.0657017 56.23808 56.23505 55.04578 56.23807 0.6208323
0.4287346 53.90761 0.5485749 3.5086655 63.90579 0.2670977 0.7456641
Average Mwe 128.0186 23.79522 23.79538 236.3569 236.3568 236.3568
236.3568 31.97887 31.97887 10.01526 28.01348 16.04276 236.3568
22.75574 21.17623
[0074]
8TABLE 3B Stream Number: 210 310 TO: 0. 0.3310 From: VA MIXED
LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID LIQUID VAPOR VAPOR LIQUID
Mole Frac H2O CH4 CL2 0 0 0 CO2 CO 0 0 BENZENE H HCLO 0 0 0 0 0 HCL
HLN N2 0 0 0 0 PY 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0
0 CL 0 0 0 O 0 0 0 0 0 OH 0 0 0 0 0 0 0 0 HCO 0 0 0 0 0 0 NH2CO2 0
0 0 0 0 0 0 Flow HCO 0 0 0 0 CO2 CO CYAN 0 0 BENZENE H2 HCLO 0 0 0
0 0 HCL HCN N2 NAPH 0 0 0 0 0 0 0 0 0 0 0 0 0 PY 0 0 0 0 0 0 0 0 0
0 0 0 0 0 OH 0 0 0 CL 0 0 0 0 0 0 0 0 0 0 0 0 0 COO 0 0 0 0 HCO 0 0
0 H 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 Total Flow
Total Flow Total Flow Temperature Pressure PSI Vapor 0 0 0 0 0 0 0
0 1 0 0 E Density QUID PHAS
[0075]
9TABLE 3C
[0076]
10TABLE 3D
[0077]
11TABLE 3E
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