US20030167692A1

US20030167692A1 - Method for increasing the efficiency of a gasification process for halogenated materials

Info

Publication number: US20030167692A1
Application number: US10/115,463
Authority: US
Inventors: Dennis Jewell; John Henley; Connie Galloway
Original assignee: Individual
Current assignee: Individual
Priority date: 2000-05-05
Filing date: 2002-04-02
Publication date: 2003-09-11

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.

Description

CROSS-REFERENCE TO RELATED APPLICATIONS

This application is a continuation of U.S. patent application Ser. No. 09/566,183, filed May 5, 2000.[0001]

FIELD OF THE INVENTION

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

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.

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.

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 ₂with residual C, CO₂, H₂O and trace elements.

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 ₂and H₂O is minimized. Higher oxygen to fuel ratios can reduce the formation of soot. However, oxygen to fuel ratios are limited by permissible flame temperatures.

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.

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.

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.

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.

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 ₂to CO ratio of the product syngas.

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.

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 ₂O<—>CO₂+H₂. 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 H₂to 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 CO₂stream 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.

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.

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.

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

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 ₂to CO ratio.

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.

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

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: [0020]
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[0021] ₂.
FIGS. 2A and 2B illustrate in more detail a gasifier stage for a gasification process of FIGS. 1. [0022]
FIG. 3 illustrates a quench and solids removal stage of a gasification process of FIG. 1. [0023]
FIGS. 4A and 4B illustrate an absorber and an aqueous acid cleanup stage of a gasification process of FIGS. 1. [0024]
FIG. 5 illustrates an anhydrous distillation stage of a gasification process of FIGS. 1.[0025]
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. [0026]
Tables 2A, 2B and 2C illustrate parameters for a commercially available system for the capture of carbon dioxide from syngas. [0027]
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. [0028]

DETAILED DESCRIPTION OF PREFERRED EMBODIMENTS

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 [0029] 530 (assumed for the purpose of the embodiment to be H₂O/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 CO₂recovery 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.
More particularly, FIGS. [0030] 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.
FIGS. 2A and 2B illustrate a [0031] 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.
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. [0032]
The primary gasifier R-[0033] 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° C. and 75 psig. These harsh conditions insure near complete conversion of all feed components.
The reactions that take place in gasifier R-[0034] 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.
RCl Partial Oxidation: [0035]
Chlorinated organics are partially oxidized to CO, H[0036] ₂and HCl.
C_vH_wCl_y+(v/2)O₂→(v)CO+[(w−y)/2]H₂=(y)HCl
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. [0037]
C_vH_wCl_y+CO₂→(v+1)CO+[(w−y)/2)]H₂+(y)HCl
C_vH_wCl_y+H₂O→(v)CO+[1+(w−y)/2]H₂+(y)HCl
Further oxidation reactions: [0038]
CO+½O₂→CO₂
H₂+½O₂→H₂O
The oxidation reactions with oxygen, including the reaction C[0039] _vH_wCl_y+(v/2)O₂→(v)CO+[(w−y)/2]H₂=(y)HCl, are highly exothermic, and thus provide the energy for driving the other reactions, maintaining the gasifier temperature as desired.
Thermal Decomposition Reactions: [0040]
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. [0041]
C_vH_wCl_x→C_r+(x)HCl+(v−r)CH₄ +[w−x−4*(v−r)/2]H₂
where C is soot, and methane CH[0042] ₄is the simplest hydrocarbon molecule which is quite stable.
Gas Shift Reactions: [0043]
CO+H[0044] ₂O⇄CO₂+H₂, classic gas shift reaction, driven primarily by gas composition, pressure and temperature have limited effect within the narrow opening range of the gasifier.
CH[0045] ₄+H₂O⇄CO+3H₂, steam—methane reforming driven almost completely to the right at gasifier conditions.
Soot is also subject to partial oxidation reactions as described in [0046] paragraph 1 above, excluding the chlorine atom.
Other Reactions: [0047]
Due to the low partial pressure of oxygen in the gasifier, essentially all halogens, including chlorine as shown above, equilibrate to the hydrogen halide. [0048]
Operating temperature in the gasifiers R-[0049] 200 and R-210 should not be allowed to drop below approximately 1350° 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° C. Conversion efficiency is very high at 1450° 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 0₂/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.
Not only steam [0050] 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, CO₂and H₂O 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.
The burner BL-[0051] 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° C. to offset the temperature drop as oxygen is expanded through a supersonic atomizing nozzle, thus increasing atomization efficiency.
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. [0052]
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° C., or safely above the dew point of HCl to avoid condensation and resulting in increased corrosion of the pressure vessel wall. [0053]
FIGS. 3, 4A and [0054] 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.
FIG. 5 illustrates features of a preferred embodiment for an anhydrous distillation process for halogenated materials. The [0055] 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 (˜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.
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, [0056] 483, 500′, enters a desorber T-510 above the azeotropic concentration in the desorber, HCl is a volatile species and is fractionated overhead.
In the preferred embodiment of FIG. 5, aqueous acid from storage illustrated as [0057] 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° C. from the secondary condenser E-520) the overheads gas should be about 96 vol. % HCl, 0.12 vol. % H₂₀, with small amounts of noncondensibles—primarily CO₂and to a lesser extent N₂. Essentially all of the noncondensibles should be driven overhead in the desorber. Column bottoms may operate at approximately 175° 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.
When producing anhydrous HCl, as per the present invention, the water balance is preferably closed by using a [0058] sidedraw vapor 514 from a desorber as a moderator for the gasifier. This vapor may be, for instance, about 59 wt % H₂₀and 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° C. due to limitations of the typical impregnated graphite materials of reboiler tubes and the typical Teflon linings for towers and piping.
The bottoms [0059] liquid stream 510, which is cross exchanged with a desorber feed, can be further cooled to approximately 40° 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.
The following example, produced by computer model, illustrates typical parameters of a gasification reactor process for halogenated materials. [0060]

EXAMPLE 1

The following feeds streams are fed to a gasifier through an appropriate mixing nozzle: [0061]

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]
The resulting gasification reactions result in a synthesis gas stream rich in hydrogen chloride. [0062]
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° 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. [0063]
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 [0064] 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.
FIG. 1B, discussed initially, illustrates in block flow diagram form the addition of a carbon [0065] 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 CO₂from syngas streams. FIG. 1B also illustrates a CO₂recycle stream 730 recycled back and fed to a gasifier 200. The CO₂would preferably be fed through a nozzle or burner in a passageway provided for an inert gas moderator, such as steam.
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 [0066] 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 CO₂and H₂O 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.
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. [0067]

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.



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:	\Gasifier with recycle HCI vapor\

		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



		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



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.



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/hrft{circumflex over ( )}2F)

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



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



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

TABLE 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

TABLE 3C

TABLE 3D

TABLE 3E

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.