U.S. patent application number 13/545920 was filed with the patent office on 2013-08-22 for process for the removal of sulfur compounds from gas streams.
This patent application is currently assigned to ARCHON TECHNOLOGIES LTD.. The applicant listed for this patent is Alan AYASSE, Conrad AYASSE, Ahmed M. SHAHIN. Invention is credited to Alan AYASSE, Conrad AYASSE, Ahmed M. SHAHIN.
Application Number | 20130216460 13/545920 |
Document ID | / |
Family ID | 48794309 |
Filed Date | 2013-08-22 |
United States Patent
Application |
20130216460 |
Kind Code |
A1 |
AYASSE; Alan ; et
al. |
August 22, 2013 |
PROCESS FOR THE REMOVAL OF SULFUR COMPOUNDS FROM GAS STREAMS
Abstract
A method of reducing sulfur compounds from an incoming gas
stream, comprising flowing the gas stream over a hydrolysis
catalyst to convert COS and CS.sub.2 to H.sub.2S and reduce
SO.sub.2 to elemental sulfur to form an effluent stream; providing
an acidic gas removal unit comprising an absorbent; flowing said
effluent stream over said absorbent to produce a stream free of
acidic gases; applying an acidic-gas desorption mode to said
acidic-gas rich absorbent to produce an acidic gas stream;
introducing oxygen to said acidic gas-rich stream; providing a
direct oxidation vessel containing catalyst suitable for catalyzing
the oxidation of the H.sub.2S to sulfur wherein the temperature of
the vessel is at or above the sulfur dew point at the reaction
pressure; and flowing said acidic gas-rich stream over said
catalyst to produce a processed stream having a reduced level of
sulfur compounds.
Inventors: |
AYASSE; Alan; (Calgary,
CA) ; SHAHIN; Ahmed M.; (Airdrie, CA) ;
AYASSE; Conrad; (Calgary, CA) |
|
Applicant: |
Name |
City |
State |
Country |
Type |
AYASSE; Alan
SHAHIN; Ahmed M.
AYASSE; Conrad |
Calgary
Airdrie
Calgary |
|
CA
CA
CA |
|
|
Assignee: |
ARCHON TECHNOLOGIES LTD.
Calgary
CA
|
Family ID: |
48794309 |
Appl. No.: |
13/545920 |
Filed: |
July 10, 2012 |
Current U.S.
Class: |
423/224 ;
423/242.1 |
Current CPC
Class: |
Y02P 20/151 20151101;
B01D 2253/102 20130101; B01D 2253/104 20130101; B01D 2257/308
20130101; B01D 2257/304 20130101; B01D 2252/204 20130101; B01D
2257/306 20130101; B01D 53/1462 20130101; Y02P 20/152 20151101;
B01D 2255/2092 20130101; C01B 17/0491 20130101; C01B 17/0404
20130101; C01B 17/046 20130101; B01D 53/75 20130101; B01D 53/8615
20130101; B01D 2257/302 20130101; B01D 2255/20707 20130101; B01D
53/48 20130101 |
Class at
Publication: |
423/224 ;
423/242.1 |
International
Class: |
B01D 53/48 20060101
B01D053/48; B01D 53/52 20060101 B01D053/52; B01D 53/50 20060101
B01D053/50 |
Foreign Application Data
Date |
Code |
Application Number |
Feb 17, 2012 |
CA |
2768359 |
Feb 17, 2012 |
CA |
2769060 |
Claims
1. A method of reducing the amount of sulfur compounds in an
incoming gas stream and producing a CO.sub.2 stream of high purity,
the method comprising: i. providing a guard bed containing an
hydrolysis catalyst for the conversion of COS and CS.sub.2 to
H.sub.2S and the reduction of SO.sub.2 to elemental sulfur, and
optionally also an RSH adsorbent suitable for RSH removal; ii.
flowing said incoming gas stream through said guard bed to produce
an effluent stream; iii. providing an acidic gas removal unit
comprising an absorbent suitable for acidic gas absorption; iv.
flowing said effluent stream from the guard bed over said absorbent
in the acidic gas removal unit to produce a stream that is free of
acidic gases, said absorbent becoming rich in acidic gases, v.
applying an acidic-gas desorption mode to said acidic-gas rich
absorbent to produce an acidic gas stream rich in acidic gases; vi.
introducing oxygen to said acidic gas-rich stream; vii. providing a
direct oxidation vessel containing a catalytic reaction zone
comprising a catalyst suitable for catalyzing the oxidation of the
H.sub.2S to sulfur wherein the temperature of the reaction zone is
at or above the sulfur dew point at the reaction pressure; viii.
flowing said acidic gas-rich stream over said catalyst to produce a
processed stream comprises a reduced level of said sulfur compounds
when compared to the incoming effluent; and ix. recycling at least
a portion of said processed stream of reduced sulfur compounds back
to said guard bed and acidic gas removal unit.
2. A method according to claim 1, wherein water is added to said
incoming gas stream prior to delivery of said incoming gas stream
to said guard bed.
3. A method according to claim 1 wherein said sulfur compounds
comprise one or more of COS, CS2, SO2, RSH and H2S.
4. A method according to claim 1 wherein said RSH adsorbent
comprises activated carbon.
5. A method according to claim 1 wherein said hydrolysis catalysis
includes one or more of alumina, titania or zirconia.
6. A method according to claim 1 wherein said guard bed is
maintained in the range of from 20.degree. C. to 300.degree. C.
7. A method according to claim 1 wherein said acidic-gas removal
sorbents include physical or chemical solvents for acidic-gas
removal.
8. A method of claim 6 wherein said physical or chemical solvents
are in liquid form or supported on porous support.
9. A method according to claim 1 wherein the acidic-gas absorption
or adsorption mode is conducted at a temperature below 100.degree.
C.
10. A method according to claim 1 wherein the acidic-gas absorption
or adsorption mode is conducted at a pressure of up to 1500
psig.
11. A method according to claim 1 wherein the method comprises a
desorption step wherein acidic gas is desorbed from the
sorbent.
12. A method according to claim 1 wherein the desorption step is
conducted at a temperature at least 20.degree. C. above the
absorption or adsorption temperature.
13. A method according to claim 1, wherein the desorption step is
conducted at a pressure up to 1500 psig.
14. A method according to claim 1 wherein the temperature of the
reaction zone in the direct-oxidation vessel is in the range of
150.degree. C. to 400.degree. C.
15. The method according to claim 1 wherein the incoming gas stream
to the direct oxidation vessel is at a gas hourly space velocity
between 100 to 10,000 hr.sup.-1.
16. The method according to claim 1 wherein the pressure in the
reaction zone in the direct oxidation vessel is between 15 and 500
psig.
17. The method as claimed in claim 1, wherein said sorbent suitable
for acid gas absorption comprises an amine based sorbent of high
selectivity toward H.sub.2S.
18. The method as claimed in claim 17, wherein said amine based
sorbent of high selectivity toward H.sub.2S comprises one or more
of N-methylpyrrolidone (NMP)/dodecane,
1,4-Diazabicyclo[2,2,2]-Octane and diisopropanolamine.
19. A method of reducing the amount of sulfur compounds in an
incoming gas stream, wherein said sulfur compounds comprise one or
more of COS, CS.sub.2, SO.sub.2, and H.sub.2S, comprising the steps
of: i. providing a guard bed containing an hydrolysis catalyst for
the conversion of COS and CS.sub.2 to H.sub.2S and the reduction of
SO.sub.2 to elemental sulfur; ii. flowing said incoming gas stream
through said guard bed, to produce an effluent stream; iii.
providing an acidic gas removal unit comprising an absorbent
suitable for acidic gas absorption; iv. flowing said effluent
stream from the guard bed over said absorbent in the acidic gas
removal unit to produce a stream that is free of acidic gases, said
absorbent becoming rich in acidic gases, v. applying an acid gas
desorption condition to said acidic-gas rich absorbent to desorb
acid gases therefrom and to produce an acidic gas stream rich in
acidic gases and containing H.sub.2S; vi. introducing oxygen to
said acidic gas-rich stream; vii. providing a direct oxidation
vessel containing a catalytic reaction zone comprising a catalyst
suitable for catalyzing the oxidation of the H.sub.2S to sulfur
wherein the temperature of the reaction zone is at or above the
sulfur dew point at the reaction pressure; viii. flowing said
acidic gas-rich stream over said catalyst to produce a processed
stream comprising a reduced level of said sulfur compounds when
compared to the acidic gas-rich stream; and ix. recycling at least
a portion of said processed stream back to said guard bed and
acidic gas removal unit.
20. The method as claimed in claim 19, further comprising the steps
of: adding water to said incoming gas stream prior to passing said
incoming gas stream through said guard bed; and said step of
applying said acid gas desorption condition comprising the step of
applying heat to said absorbent to raise it to a temperature of
approximately 130.degree. C.
21. The method as claimed in claim 19, further comprising the step
of, after step v, purging the acid gas removal unit with a gas
stream to remove residual CO.
22. The method as claimed in claim 19, further comprising the step
of, at approximately the time of applying an acid gas desorption
condition to said acidic-gas rich absorbent, redirecting said
incoming gas stream to a secondary guard bed and thereafter to a
secondary acidic gas removal unit.
23. A system for of reducing the amount of sulfur compounds in an
incoming gas stream and producing a CO.sub.2 stream of high purity,
such system comprising: i. a guard bed containing an hydrolysis
catalyst, said hydrolysis catalyst adapted for the conversion of
COS and CS.sub.2 to H.sub.2S and the reduction of SO.sub.2 to
elemental sulfur, and optionally also an RSH adsorbent suitable for
RSH removal, said guard bed adapted to receive said incoming stream
and produce an effluent stream after passage of said incoming
stream through said guard bed; ii. an acidic gas removal unit
comprising an absorbent suitable for acidic gas absorption, adapted
to receive said effluent stream and produce a produced stream that
is free of acidic gases, said absorbent becoming rich in acidic
gases; iii. said absorbent adapted, when said having heat applied
thereto, to produce an acidic gas stream rich in acidic gases; iv.
oxygen supply means, adapted to supply oxygen to said acidic gas
stream; v. a direct oxidation vessel, adapted to receive said
acidic stream rich in acidic gases and oxygen, and containing a
catalytic reaction zone comprising a catalyst suitable for
catalyzing the oxidation of the H.sub.2S to sulfur wherein the
temperature of the reaction zone is at or above the sulfur dew
point at the reaction pressure, and to produce a processed stream
comprising a reduced level of said sulfur compounds when compared
to the incoming acidic stream; and vi. recycling piping to recycle
at least a portion of said processed stream back to said guard bed
and acidic gas removal unit.
24. The system as claimed in claim 23, wherein said sulfur
compounds comprise one or more of COS, CS2, SO2, RSH and H2S.
25. The system as claimed in claim 23 wherein the RSH adsorbent
comprises activated carbon.
26. The system as claimed in claim 23, wherein the hydrolysis
catalysis includes one or more of alumina, titania or zirconia.
27. The system according to claim 23 wherein said acidic-gas
removal sorbents include physical or chemical solvents for
acidic-gas removal.
28. The system according to claim 23 wherein said physical or
chemical solvents are in liquid form or supported on porous
support.
29. The system as claimed in claim 23, wherein said absorbent
suitable for acidic gas absorption comprises an amine based
sorbents.
30. The system as claimed in claim 29, wherein said amine based
sorbent comprises one or more of to 1,4-Diazabicyclo[2,2,2]-Octane,
1,5-Diazabicyclo[5,4,0]-Undec-5-ene, 1,4-dimethylpiperazin-2-one
and 1,5-Diazabicyclo[4,3,0]-non-5-ene.
Description
PRIORITY
[0001] This application claims priority from Canadian Patent
Applications 2,768,359 and 2,769,060 each filed Feb. 17, 2012.
FIELD OF THE INVENTION
[0002] The invention generally relates to a chemical processes used
in processing recovered gas and oil, and more particularly to a
process and apparatus for the removal of sulfur compounds from gas
streams.
BACKGROUND
[0003] Natural gas and refinery gas streams are commonly
contaminated with sulfur-containing compounds such as hydrogen
sulfide (H.sub.2S) and/or carbonyl sulfide (COS) and carbon dioxide
(CO.sub.2). If substantial amounts of H.sub.2S are present,
regulatory restrictions dictate special precautions must be taken
to purify the gas streams. The first step of the H.sub.2S removal
process from the H.sub.2S-containing streams is accomplished by an
acid-gas removal unit which removes substantial amounts of H.sub.2S
and CO.sub.2 from the acidic-gas containing streams. The off-gas
from the acid-gas removal unit is mainly H.sub.2S and CO.sub.2. The
sulfur from this off-gas stream is usually removed by the Claus
reaction which produces salable elemental sulfur. After a
`tail-gas` treatment to further reduce the sulphur content, the
remaining CO.sub.2 may be safely vented to the atmosphere. However,
there has been increasing concern about the damage caused by
CO.sub.2 and this has led to an increased demand to reduce the
emission of CO.sub.2 to the atmosphere.
[0004] Typically, separation of CO.sub.2 and H.sub.2S from streams
containing acidic gas is achieved by the chemical absorption
process employing liquid amine solutions, such as monoethanolamine
(MEA), diethanolamine (DEA) or methyldiethanolamine (MDEA). In this
process the CO.sub.2 reacts with the liquid amine solution to form
a carbamate, while H.sub.2S reacts with the amine solution to form
(amine)H.sup.+ and bisulfide (SH.sup.-) species. Upon heating, the
carbamate and (amine)H.sup.+ species decompose to release the
absorbed CO.sub.2 and H.sub.2S and produce a regenerated amine
solution. Disadvantageously with this process, however,
sulfur-containing compounds such as SO.sub.2, COS and/or CS.sub.2,
if present in the feed stream, react with the liquid amine
absorbent and a higher temperature is required to regenerate the
amine solution. SO.sub.2 also reacts with the amine to form
sulphates which necessitates partial replacement of the amine.
[0005] Liquid alkoxylated amines, such as diisopropanolamine, have
been used for CO.sub.2 removal from streams containing acidic
gases. U.S. Pat. No. 4,044,100 described the use of liquid mixtures
of diisopropanolamine and polyethylene glycol for acid gas removal
from gaseous streams.
[0006] There are many fields of applications in which it is
required to remove H.sub.2S and CO.sub.2 from streams containing
acidic gases. U.S. Pat. No. 4,553,984 describes a process for the
removal of CO.sub.2 and H.sub.2S, simultaneously, from streams
containing acidic gases wherein the stream is brought into counter
flow contact with an aqueous of methyldiethanolamine (MDEA) at a
pressure of 10-110 bars. Nevertheless, there are different
applications in which it is required to reduce the H.sub.2S to a
very low level without essential removal of CO.sub.2; therefore,
solvents with high H.sub.2S-absorbing power are desired. U.S. Pat.
No. 5,277,884 disclosed a process for selective removal of H.sub.2S
from streams containing both H.sub.2S and CO.sub.2 acidic gases.
The process according to that invention comprises contacting the
acidic gas containing stream with a solvent that comprises a
mixture of N-methylpyrrolidone (NMP) and dodecane.
[0007] The acid gas removal process utilizing liquid amine
solutions is costly and energy-intensive because the liquid amine
solution has a limited life time due to its degradation through
oxidation. Furthermore, the high corrosivity of the utilized amine
makes it prohibitive to use high concentrations of the amine
solutions. Therefore, new acidic gas capture technology utilizing
thermally stable solid sorbents has increasingly received attention
due to its potential for reducing corrosion and energy cost and
improving mass/heat transfer efficiency. Such technology is based
on the ability of a porous solid sorbent to reversibly adsorb the
CO.sub.2 and H.sub.2S from the acidic gas containing streams at
high pressure.
[0008] U.S. patent application Ser. No. 13/399,911 filed Feb. 17,
2012 relates to a process for a acidic gas recovery from acidic gas
containing streams employing a class of novel thermally stable
amine adducts (sorbents). The regenerable sorbents described in
that process had high CO.sub.2 and H.sub.2S absorption capacity and
comprised a porous solid support, a cross-linked amine and a polyol
reactive toward the utilized amine. The sorbents according to this
invention enable acidic gas absorption/desorption cycles at various
temperatures and pressures. Advantageously, the
absorption/desorption cycles could be conducted at a pressure of
1500 psig and a temperature of 130.degree. C., so that the CO.sub.2
at this condition was ready for direct downhole storage or
pipelining at greatly reduced compression costs. In addition the
adsorption could take place at low pressure with desorption at high
pressure.
[0009] Typically, the desorbed gas stream from an acid-gas removal
unit is mainly H.sub.2S and CO.sub.2 and the sulfur is usually
removed by the Claus process. In the first step in the Claus
process, one third of the hydrogen sulfide present in the feed
stream is oxidized to sulfur dioxide, SO.sub.2, by the reaction as
follows:
H.sub.2S+O.sub.2=SO.sub.2+H.sub.2
In the second step, the remaining H.sub.2S and the SO.sub.2 are
reacted in the presence of a Claus catalyst to form elemental
sulfur in a Claus reactor according to Reaction 1:
2H.sub.2S+SO.sub.2=2H.sub.2O+3S Claus reaction 1
[0010] The Claus reaction is limited by thermodynamic equilibrium
and only a portion of the total sulfur can be produced. Therefore,
multiple stages with sulfur condensation between the stages are
needed in order to increase the sulfur recovery factor. The
effluent gas from a series of reactors in a Claus plant contains
varying amounts of different compounds including sulfur vapor,
sulfur dioxide, un-reacted H.sub.2S, carbonyl sulfide (COS), and/or
carbon disulfide (CS.sub.2). Carbon disulphide is formed according
to Reaction 2:
CH.sub.4+4S.fwdarw.CS.sub.2+2H.sub.2S High temperature Claus
furnace or combustion reaction 2
[0011] Removal of the sulfur content of the off-gas streams from
the Claus process is accomplished by catalytic reduction with
hydrogen to convert the sulfur compounds to H.sub.2S, absorption of
the H.sub.2S produced with an additional amine system and then
recycling the desorbed gas to the Claus plant. This process is
operable as long as the concentration of the CO.sub.2 is up to 15%
and H.sub.2S is above 50% by volume in the feed stream. However, if
the H.sub.2S/CO.sub.2 feed gas stream to Claus process contains
less than 40% by volume H.sub.2S, the Claus plant becomes difficult
to operate with respect to the thermal zone and special
considerations have to be taken when combusting part of H.sub.2S to
SO.sub.2 as required for the Claus reaction. These operational
difficulties mainly arise from the fact that the required
temperatures for the combustion of H.sub.2S cannot be reached in
the thermal zone. Therefore, the off-gas stream from the Claus
plant is burned with air to convert all sulfur-containing compounds
in the stream to SO.sub.2 before discharge into the atmosphere. As
the environmental requirements are becoming stricter, the SO.sub.2
emission limit is being lowered, giving rise to the challenge of
how to reduce or completely eliminate SO.sub.2 emissions.
Consequently, another sulfur removal process is needed that can
handle H.sub.2S/CO.sub.2 feed gas streams containing CO.sub.2 of
concentrations greater than 15% and H.sub.2S of a concentration
less than 40% by volume.
[0012] The direct oxidation of H.sub.2S to elemental sulfur using
oxidation catalysts has gained broad acceptance for achieving high
sulfur removal efficiency. U.S. Pat. No. 4,197,277 describes a
process for the oxidation of H.sub.2S to elemental sulfur by the
following H.sub.2S Oxidation Reactions 3 and 4:
H.sub.2S+0.5O.sub.2.fwdarw.S+H.sub.2O H.sub.2S Partial oxidation
3
H.sub.2S+1.5O.sub.2.fwdarw.SO.sub.2+H.sub.2O H.sub.2S Complete
oxidation 4
[0013] According to U.S. Pat. No. 4,197,277, the
H.sub.2S-containing gas is passed with an oxygen-containing gas
over a catalyst which comprises iron oxide and vanadium oxide as
active materials and aluminum oxide as a support material. The
catalyst described in that patent gives rise to at least a partial
Claus equilibrium, so that SO.sub.2 formation cannot be prevented.
Similarly, U.S. Pat. No. 5,352,422 describes a process for
oxidizing the un-reacted H.sub.2S in the Claus tail gas to
elemental sulfur. The patent describes a catalyst prepared by
impregnation of an iron containing solution or an
iron/chromium-containing solution into several carriers followed by
calcinations in air at 500.degree. C.
[0014] U.S. Pat. No. 4,818,740 disclosed a catalyst for the
H.sub.2S oxidation to elemental sulfur, the use of which prevents
the reverse Claus reaction to a large extent. The catalyst
according to that patent comprises a support of which the surface
exposed to the gaseous phase does not exhibit any alkaline
properties under the reaction conditions, while a catalytically
active material is applied to this surface. An improvement of the
method disclosed in '740 is disclosed in European Patent 409,353.
This patent relates to a catalyst for the selective oxidation of
sulfur-containing compounds to elemental sulfur, comprising at
least one catalytically active material and optionally a support.
The described catalyst exhibits substantially no activity towards
the reverse Claus reaction under the reaction conditions.
[0015] The H.sub.2S direct oxidation to elemental sulfur is
suitable for gas streams comprising high concentrations of CO.sub.2
and low concentrations of H.sub.2S. Nevertheless, the total sulfur
removal efficiency decreases if carbon monoxide or COS gases are
present in the feed stream. Carbon monoxide, if present in the feed
gas streams, undergoes side reactions during the H.sub.2S direct
oxidation to form COS. In addition, CO.sub.2 may also react with
H.sub.2S to form COS during direct oxidation reaction:
CO+S.fwdarw.COS 5
CO+H.sub.2S.fwdarw.COS+H.sub.2 6
3CO+SO.sub.2.fwdarw.COS+2CO.sub.2 7
H.sub.2S+CO.sub.2.fwdarw.COS+H.sub.2O 8
[0016] U.S. patent application Ser. No. 13/399,710 filed Feb. 17,
2012 entitled "Removal of Sulfur Compounds from a Gas Stream"
relates to a process for simultaneously oxidizing H.sub.2S to
elemental sulfur and hydrolyzing COS to H.sub.2S in the presence of
an oxidation catalyst and a feed gas stream containing CO of a
concentration greater than 1% by volume and CO.sub.2 of a
concentration greater than 14% by volume of the total feed gas
flow. In this process, an H.sub.2S-containing stream was mixed with
a molecular oxygen containing gas and then passed over an oxidation
catalyst at a temperature of 220.degree. C., a gas hourly space
velocity of 1000 hr.sup.-1 and a pressure of 100 psig. The
concentration of the COS produced decreased from 1900 ppm, using a
dry gas stream, to 316 ppm upon using a feed stream containing
greater than 10% water. The oxygen in the feed gas stream was
adjusted to achieve the highest conversion of H.sub.2S to elemental
sulfur and to deliberately produce an off-gas stream containing
H.sub.2S/SO.sub.2 ratio of 2:1 which is ready as a feed gas stream
for other sulfur removal units such as Crystasulf.TM..sup.1.
Therefore, the process was operated at a relatively low sulfur
yield of 78.1% and a total H.sub.2S conversion of 90.4%. .sup.1
Trademark of URS CORPORATION CORPORATION for sulfur removal
units.
[0017] In summary, high sulfur removal efficiency can be achieved
by utilizing a multi-stage Claus process and off-gas post
treatment. Importantly, however, this process is limited by the
concentration of the CO.sub.2 in the gas stream and necessity of
employing an H.sub.2S enrichment unit. Therefore, other sulfur
recovery processes, such as the H.sub.2S direct oxidation process,
have gained worldwide attention. In fact, the H.sub.2S direct
oxidation to elemental sulfur process has become the cornerstone of
the high sulfur recovery upon coupling with Claus process.
Disadvantageously, however, the H.sub.2S direct oxidation process
is still limited due to the process conditions and feed gas
composition. As mentioned, a considerable amount of COS is produced
when operating the H.sub.2S direct oxidation process, in a
once-through mode, with sulfur-containing gas streams comprising CO
and CO.sub.2 at a temperature above the sulfur dew point and a high
pressure. Consequently, a robust sulfur removal process that can
overcome the aforementioned difficulties is still needed.
SUMMARY OF THE INVENTION
[0018] The present invention provides a robust process for the
efficient carbon dioxide recovery and desulfurization of feed
stream gases comprising sulfur constituents as well as a
considerable amount of carbon dioxide at elevated pressure,
including but not limited to CO.sub.2 of a concentration greater
than 14% by volume of the total feed gas flow.
[0019] The process according to this invention not only converts
the sulfur-containing compounds to elemental sulfur but also
produces a high pressure CO.sub.2 stream of high purity. This
process will remarkably reduce the size of the reactor required for
the desulfurization of the feed streams and will also provide a
significant energy consumption advantage when the CO.sub.2 gas
stream is compressed for pipelining or deep well disposal. The feed
streams suitable for the process according to the present invention
comprise but are limited to sulfur containing compounds, such as
H.sub.2S, SO.sub.2, COS, CS.sub.2; oxidizable constituents such as,
hydrogen, carbon monoxide, light hydrocarbons, e.g. methane, ethane
or propane; natural gas; associated gas from oil production; gases
produced from oilsand refining, e.g. coker gas; gases produced from
Toe-to-Heel-Air-Injection process (THAI.TM.); or other in situ
combustion gas; coal or oil gasification processes; inert gases,
such as nitrogen, helium or carbon dioxide and any combination
thereof.
[0020] The approach utilized in the present invention is to
selectively remove and concentrate the H.sub.2S and/or CO.sub.2
from the gas streams, and then oxidize the H.sub.2S to salable
elemental sulfur. More particular, this invention comprises a
process for the removal of H.sub.2S and/or CO.sub.2 from the sour
gases at room temperature and elevated pressure by contacting the
sour gas with a suitable acid gas absorbent. Then, subjecting the
absorbent to a desorption mode at a pressure similar to the
absorption pressure but at an elevated temperature. The produced
gas stream from the desorption mode contains mainly H.sub.2S,
CO.sub.2 and/or N.sub.2. Subsequently, the product gas from the
desorption mode is mixed with a stream containing molecular oxygen
and is then passed to an H.sub.2S direct oxidation reactor to
partially oxidize the H.sub.2S to elemental sulfur.
[0021] Accordingly, in one broad aspect of the method of the
present invention, such method comprises a method of reducing the
amount of sulfur compounds in an incoming gas stream comprising:
[0022] a. providing a guard bed containing an hydrolysis catalyst
for the conversion of COS and CS.sub.2 to H.sub.2S and the
reduction of SO.sub.2 to elemental sulfur, and optionally also an
RSH adsorbent suitable for RSH removal; [0023] b. flowing said
incoming gas stream through said guard bed to produce an effluent
stream; [0024] c. providing an acidic gas removal unit comprising
an absorbent suitable for acidic gas absorption; [0025] d. flowing
said effluent stream from the guard bed over said absorbent in the
acidic gas removal unit to produce a stream that is free of acidic
gases, said absorbent becoming rich in acidic gases, [0026] e.
applying an acidic-gas desorption condition to said acidic-gas rich
absorbent to desorb acid gases from said absorbent and produce an
acidic gas stream rich in acidic gases; [0027] f. introducing
oxygen to said acidic gas-rich stream; [0028] g. providing a direct
oxidation vessel containing a catalytic reaction zone comprising a
catalyst suitable for catalyzing the oxidation of the H.sub.2S to
sulfur wherein the temperature of the reaction zone is at or above
the sulfur dew point at the reaction pressure; [0029] h. flowing
said acidic gas-rich stream over said catalyst to produce a
processed stream comprises a reduced level of said sulfur compounds
when compared to the incoming effluent; and [0030] i. recycling at
least a portion of said processed stream of reduced sulfur
compounds back for passage through to said guard bed and acidic gas
removal unit.
[0031] In a preferred embodiment, such process produces a
pressurized stream of high CO.sub.2 purity.
[0032] In a further embodiment, water is added to said incoming gas
stream prior to delivery of said incoming gas stream to said guard
bed.
[0033] In a further preferred embodiment, the sulfur compounds
comprise one or more of COS, CS.sub.2, SO.sub.2, RSH and
H.sub.2S.
[0034] In a further preferred embodiment, the RSH adsorbent
comprises activated carbon.
[0035] In a further preferred embodiment, the hydrolysis catalysis
includes one or more of alumina, titania or zirconia.
[0036] In a preferred embodiment of the above method, the guard bed
is maintained in the range of from 20.degree. C. to 300.degree.
C.
[0037] In a preferred embodiment of the above method, according to
claim 1 wherein said absorbent suitable for acidic-gas removal
includes physical or chemical solvents.
[0038] In a preferred embodiment of the above method, the physical
or chemical solvents used as absorbents are in liquid form or
supported on porous support.
[0039] In a still-further embodiment of the above method, the
acidic-gas absorption or adsorption mode is conducted at a
temperature below 100.degree. C.
[0040] In a still further embodiment of the above method, the
acidic-gas absorption or adsorption mode is conducted at a pressure
of up to 1500 psig.
[0041] In a still-further preferred embodiment of the above method,
the method comprises a desorption step wherein acidic gas is
desorbed from the acidic gas absorbent.
[0042] In a still further preferred embodiment, the desorption step
is conducted at a temperature at least 20.degree. C. above the
absorption or adsorption temperature.
[0043] In a still further preferred embodiment, the desorption step
is conducted at a pressure up to 1500 psig.
[0044] In a still further preferred embodiment, the temperature of
the reaction zone in the direct-oxidation vessel is in the range of
150.degree. C. to 400.degree. C.
[0045] In a still further preferred embodiment, the incoming gas
stream to the direct oxidation vessel is at a gas hourly space
velocity between 100 to 10,000 hr.sup.-1.
[0046] In a still further preferred embodiment, the pressure in the
reaction zone in the direct oxidation vessel is between 15 and 500
psig.
[0047] In another aspect of the present invention, the present
invention relates to a system for reducing the amount of sulfur
compounds in an incoming gas stream and producing a CO.sub.2 stream
of high purity. Accordingly, such system of the present invention,
in a broad aspect thereof, comprises: [0048] i. a guard bed
containing an hydrolysis catalyst, said hydrolysis catalyst adapted
for the conversion of COS and CS.sub.2 to H.sub.2S and the
reduction of SO.sub.2 to elemental sulfur, and optionally also an
RSH adsorbent suitable for RSH removal, said guard bed adapted to
receive said incoming stream and produce an effluent stream after
passage of said incoming stream through said guard bed; [0049] ii.
an acidic gas removal unit comprising an absorbent suitable for
acidic gas absorption, adapted to receive said effluent stream and
produce a produced stream that is free of acidic gases, said
absorbent becoming rich in acidic gases; [0050] iii. said absorbent
adapted, when said having heat applied thereto, to produce an
acidic gas stream rich in acidic gases; [0051] iv. oxygen supply
means, adapted to supply oxygen to said acidic gas stream; [0052]
v. a direct oxidation vessel, adapted to receive said acidic stream
rich in acidic gases and oxygen, and containing a catalytic
reaction zone comprising a catalyst suitable for catalyzing the
oxidation of the H.sub.2S to sulfur wherein the temperature of the
reaction zone is at or above the sulfur dew point at the reaction
pressure, and to produce a processed stream comprising a reduced
level of said sulfur compounds when compared to the incoming acidic
stream; and [0053] vi. recycling piping to recycle at least a
portion of said processed stream back to said guard bed and acidic
gas removal unit.
[0054] In a preferred embodiment of the above system the sulfur
compounds comprise one or more of COS, CS.sub.2, SO.sub.2, RSH and
H.sub.2S.
[0055] In a further preferred embodiment of the system where an RSH
adsorbent is used, such RSH adsorbent comprises activated
carbon.
[0056] In a further preferred embodiment the hydrolysis catalysis
includes one or more of alumina, titania or zirconia.
[0057] In a still further preferred embodiment, the absorbent
suitable for acidic gas absorption comprises physical or chemical
solvents for acidic-gas removal, and further wherein said physical
or chemical solvents are in liquid form or supported on a porous
support.
BRIEF DESCRIPTION OF THE DRAWINGS
[0058] FIG. 1 is a schematic diagram of the desulfurization process
of the present invention;
[0059] FIG. 2 is a schematic graph of the variation of H.sub.2S
concentration in the recycle gas from the absorber column;
[0060] FIG. 3 is a schematic graph of the variation of H.sub.2S
concentration in the recycle gas from the H.sub.2S direct oxidation
reactor;
[0061] FIG. 4 is a schematic graph of the variation of COS
concentration in the recycle gas from the H.sub.2S direct oxidation
reactor; and
[0062] FIG. 5 is a schematic graph of the variation of SO.sub.2
concentration in the recycle gas from the H.sub.2S direct oxidation
reactor
DETAILED DESCRIPTION OF THE INVENTION
[0063] Referring to FIG. 1, the overall chemical process of the
present invention is shown as a flow diagram in which the
components of the acid gas removal system apparatus 100 are
shown.
[0064] According to the first step of the process, a
sulfur-containing gas stream 2, 3, typically a sour gas stream
comprising CO.sub.2 and H.sub.2S, is fed to a primary absorber
column 7 comprising an amine-based acid gas absorbent to remove the
CO.sub.2 and H.sub.2S from the sour gas stream.
[0065] Notably, however, different sulfur containing compounds such
as COS, SO.sub.2 and/or RSH, if present in the sour gas stream,
will react with the amine-based absorbent, and reduce its CO.sub.2
and H.sub.2S absorption capacity. Therefore, a protective guard bed
6 containing alumina and/or activated carbon at a temperature of
120.degree. C. is placed on the feed gas stream prior to the
primary amine absorber column 7 [and also prior to the secondary
amine absorber column 7a--see below]. The main function of the
protective guard bed 6, 6a is to remove the RSH from the sour gas
stream 2,3 and to catalyze the reaction of the H.sub.2S with
SO.sub.2, if present, to produce elemental sulfur which can
eventually be recovered by regenerating the guard bed 6, 6a at a
temperature of 220.degree. C. in a flow of a N.sub.2 sweep gas 4.
Moreover, the alumina guard bed 6, 6a will catalyze the hydrolysis
of the COS and/or CS.sub.2 to H.sub.2S and CO.sub.2 prior to the
respective primary (or secondary) amine-based absorber 7, 7a.
[0066] In a commercial application, a single stream containing acid
gases will normally be treated. But in the laboratory demonstration
unit of FIG. 1, for ease of operation, two streams, 2 and 3
comprise the feed gas stream. Stream 2 is a mixture of nitrogen and
hydrogen sulfide and stream 3 is a mixture of the other components:
CO.sub.2, H.sub.2, CO, CH.sub.4 and N.sub.2. Streams 2 and 3 are
mixed to produce a synthetic sour gas mixture containing CO.sub.2,
H.sub.2S, H.sub.2, CO, CH.sub.4 and N.sub.2. The flow rates of the
inlet gas streams 2 and 3 are controlled via mass flow controllers
and the pressure of the guard bed 6, 6a and absorber column 7, 7a
is regulated by a back pressure control valve 10. The pressure of
the inlet feed stream 2, 3 is about 130 psig and temperature is
about 20.degree. C. The synthetic sour gas stream 2,3 is initially
passed through valve 5 to a guard bed 6 comprising alumina and/or
activated carbon at 120.degree. C. Under these conditions, the COS
and/or CS.sub.2 is hydrolyzed to H.sub.2S and CO.sub.2, while
SO.sub.2, if present, is converted to elemental sulfur by the
reaction with the H.sub.2S present in the feed gas stream.
Subsequently, the effluent gas from the protective guard bed 6 is
cooled down and then fed to a primary absorber column 7 containing
an amine-based absorbent 32 to selectively remove the H.sub.2S
and/or CO.sub.2 from the sour gas stream. The H.sub.2S and CO.sub.2
are absorbed immediately, and a purified produced gas containing
H.sub.2, CO, CH.sub.4 and N.sub.2 leaves the absorption bed 7
through valve 8. During the absorption mode, valves 8 and 9 are
employed to direct the de-sulfurized gas from the absorber column 7
to a micro gas chromatograph 11 equipped with an automated stream
selection means (not shown) to determine the moment of breakthrough
of the acidic gas, and when detected, to adjust valve 8 to direct
flow from absorber column to pump 12 during the desorption phase
(see below).
[0067] The acidic gas absorption mode is performed at room
temperature and a pressure of 100 psig, while the desorption mode
is conducted at a temperature of 130.degree. C. using a sweep gas
such as N.sub.2 or CO.sub.2.
[0068] Upon the acidic gas breakthrough (ie upon saturation of the
amine-based absorbent 32 in primary absorbent column 7 and when
detected by gas chromatograph 11 or other similar device--), the
sour feed gas stream 2,3 is switched via valve 35 to secondary
guard bed/absorber column system B, and valve 8 redirects the
produced gas stream from secondary system B to gas analyzer 11.
Secondary system B has a secondary protective guard bed 6a, and
secondary amine absorber column 7a. At such time the primary
absorber column 7 is converted to a desorption mode. Specifically,
the loaded or rich absorbent 32, i.e. absorbent containing the
absorbed H.sub.2S and CO.sub.2 within amine absorber column 7 is
heated to 130.degree. C. to free the H.sub.2S and CO.sub.2 from the
absorbent. Therefore, the pressure of the absorber column 7
increases from 100 psig (at room temperature) to 150 psig. At this
point, the rich gas stream leaving the absorber 7 is composed of
H.sub.2S, CO.sub.2 and N.sub.2 (sweep gas). If CO.sub.2 is used as
a sweep gas, the resultant gas stream cannot be processed in Claus
plant because the ratio of the H.sub.2S to CO.sub.2 would be too
low. Conversely, this stream is suitable for the H.sub.2S direct
oxidation to elemental sulfur process. The process according to
this invention, therefore, provides a subsequent batch process for
the partial oxidation of the H.sub.2S present in this stream to
elemental sulfur. The sulfur removal efficiency of the batch
process according to this invention is greater than 99% by
volume.
[0069] In the second step of the process, and with continued
reference to FIG. 1, the H.sub.2S-rich gas from the absorber column
7 is sent to an H.sub.2S direct oxidation system 30 to partially
oxidize the H.sub.2S to elemental sulfur. Accordingly, the
CO.sub.2/H.sub.2S desorbed gas stream at a pressure of 150 psig is
passed through valve 8 and then mixed with a molecular oxygen
containing stream 1 to produce a gas mixture containing mainly
CO.sub.2, H.sub.2S, O.sub.2 and/or N.sub.2. Typically, small
amounts of COS and SO.sub.2 byproducts are produced during the
H.sub.2S direct oxidation reaction. Therefore, the flow rate of the
molecular oxygen-containing stream is adjusted such that the
molecular oxygen to H.sub.2S ratio is less than 0.5. The resultant
gas mixture at a pressure of about 150 psig is then sent to a gas
circulating pump 12 to supply the gas mixture to the H.sub.2S
direct oxidation system 30 having a H.sub.2S oxidation reactor 16.
The feed gas flow rate for the H.sub.2S direct oxidation reactor 16
is controlled via a mass flow controller 13, and its pressure is
monitored by a pressure gauge 14. The feed gas stream of the
H.sub.2S oxidation reactor 16 which forms part of H.sub.2S direct
oxidation system 30 is firstly passed through a pre-heating coil 15
to bring the feed gas mixture to the desired temperature. H.sub.2S
oxidation reactor 16 in the form of a down flow reactor is utilized
for the oxidation of H.sub.2S to elemental sulfur. The down flow
reactor 16 is packed with an oxidation catalyst, and located in an
oven 17 and operated at a temperature slightly greater than the
sulfur dew point at the oxidation reaction pressure. Initially, the
pressure of the H.sub.2S direct oxidation reactor 16 is adjusted to
60 psig via the back pressure control valve 10 and then increased
to a pressure of 100 psig upon mixing with the gas mixture during
the oxidation process. As a result, the overall pressure of the
H.sub.2S direct oxidation system 30 is about a 100 psig.
[0070] The product effluent 25 from the H.sub.2S direct oxidation
reactor 16 comprises un-reacted H.sub.2S, H.sub.2, CO, CO.sub.2,
CH.sub.4, N.sub.2, sulfur vapor and a very small amount of COS
and/or SO.sub.2. Consequently, the produced fluid from the
oxidation reactor 16 is cooled to separate the produced sulfur from
the gas phase in sequential initial and secondary separators 18, 19
respectively, and the effluent gas from the secondary sulfur
separator 19 is then recycled back to the H.sub.2S direct oxidation
system 30 to increase the overall sulfur recovery factor. The
product gas from secondary separator 19 is passed through valve 9,
micro filter 20, valve 5 and then to_the guard bed 6. The
temperature of the protective guard bed 6, 6a and amine-based
absorber 7, 7a are maintained fairly constant during the effluent
gas recycling process at temperatures of 120.degree. C. and
130.degree. C., respectively. As indicated earlier, the produced
COS is hydrolyzed in the guard bed 6, 6a to H.sub.2S, and the
produced SO.sub.2 is removed by the reaction with the H.sub.2S
present in the stream producing elemental sulfur. The effluent gas
recycling procedure according to the second step of this process is
repeated until the H.sub.2S in the recycle gas is less than 50 ppmv
and the overall H.sub.2S conversion to elemental sulfur is greater
than 99%.
[0071] According to the third step of the process, the primary
absorber column 7 at a temperature of 130.degree. C. and a pressure
of 100 psig is purged with a gas free of CO.sub.2 and H.sub.2S to
avoid the re-adsorption of CO.sub.2 and H.sub.2S upon cooling down
the absorber to room temperature. A N.sub.2 gas stream or a
fraction of the off-gas stream from the secondary guard bed 6a and
absorber column 7a (CO.sub.2 and H.sub.2S free gas) is employed
until no CO.sub.2 is detected in the outlet gas stream.
Subsequently, the primary absorber column 7 is cooled to room
temperature, and valve 35 is then adjusted to prevent incoming
stream flow to secondary system B, and simultaneously allowing
incoming stream to flow to then be re-directed back to guard bed 6
and absorber column 7 then being used in a new CO.sub.2/H.sub.2S
absorption cycle, with absorber column 7a in secondary amine
absorber system B then undergoing the desorption process earlier
conducted on absorbent column 7. When using CO.sub.2 as the sweep
gas and pure O.sub.2 as the oxygen source, the off gas will be
99.9% pure. Meanwhile, the pressure of the oxidation reactor 16, if
not being supplied with desorbed gas from secondary system B, is
reduced to 60 psig.
[0072] In one particular first preferred embodiment and with
continued reference to FIG. 1, the acidic gas containing stream at
a pressure up to 1500 psig is passed through a humidifier (not
shown) at a temperature in the range from 30.degree. C. to
90.degree. C. and then through a protective guard bed 6 comprising
an RSH absorbent and/or a catalyst 32 at a temperature in the range
from 30.degree. C. to a temperature slightly greater than the
sulfur dew point at the process pressure. The RSH absorbent 32
includes but is not limited to activated carbon and silica gel
impregnated with Cu(II) and Mn(IV). The catalyst component thereof
comprises but is not limited to alumina, titania and supported
metal oxide catalyst. The use of the guard bed 6, 6a is
advantageous in the case of feed gas streams comprising CO, CO2,
RSH, COS and SO.sub.2. The metal oxide/s catalyst included in the
guard bed 6, 6a hydrolyzes the COS and CS.sub.2 to H.sub.2S and
CO.sub.2 and reduces the SO.sub.2, if present in the feed stream or
produced as a byproduct during the H.sub.2S direct oxidation, to
elemental sulfur. Therefore, H.sub.2S is the only sulfur
constituent in the off-gas stream from the guard bed 6, 6a. The
off-gas stream from the guard bed 6, 6a is then passed through
valve 8 and directed to initial and secondary separators 18, 19 and
therein cooled down. The H.sub.2S and CO.sub.2 are simultaneously
removed from the off-gas stream by a primary acidic gas removal
unit 7, 7a. The acidic gas removal units 7, 7a may contain any of
the available technologies based on the liquid or solid absorbents
which are selective toward both H.sub.2S and CO.sub.2 gases. Once
the acidic gases have broken through in either acidic gas removal
unit 7, as detected by the gas analyzer 11, the feed gas stream is
switched to a secondary guard bed/acidic gas removal unit B, and
the primary acidic removal unit 7 is conducted to a desorption
process at a temperature higher than the absorption temperature.
Carbon monoxide, if present in the feed gas, tends to react with
the H.sub.2S to form COS (equation 6) in the amine based acidic gas
removal units. Typically, the COS produced in the acid removal
units 7, 7a reacts with the amine based sorbents and a higher
energy is required to regenerate the amine based sorbents.
Moreover, a considerable amount of COS will be produced during the
oxidation of H2S to elemental sulfur which in turn will reduce the
sulfur selectivity per each cycle. Although the produced COS will
be hydrolyzed to H.sub.2S in the guard bed, the overall sulfur
removal process will be too long (Example 2). Therefore, once the
acidic gases have broken through, the off-gas stream from the
primary acidic gas removal unit 7 is mixed with a molecular oxygen
containing stream 1 and the resultant mixture is then sent to an
H.sub.2S direct oxidation reactor 16 comprising a suitable
oxidation catalyst at a temperature slightly greater than the
sulfur dew point at the reaction pressure. The present invention
employs any catalyst suitable for the oxidation of H.sub.2S to
elemental sulfur. Typically, the oxidation catalyst comprises an
oxide and/or sulfide form of one or more metals deposited or mixed
with one or more refractory metal oxides. The metal oxides and/or
sulfides include, but are not limited to oxides and/or sulfides of
V, Cr, Mn, Fe, Co, Ni, Cu, Nb, Mo, Tc, Ru, Rh, Pd, Hf, Ta, W, Re,
Os, Ir, Pt, Au, La, Ce, Pr, Nd, Pm, Sm, Eu, Gd, Tb, Dy, Ho, Er, Tm,
Yb, Lu, Bi or any combinations thereof. The refractory metal oxides
include, but are not limited to Al, Ti, Si, Zr and any combinations
thereof.
[0073] According to one embodiment of the present invention, the
high desulfurization level of the resultant mixture is achieved by
utilizing a batch process, which is accomplished by recycling the
effluent gas from the H.sub.2S direct oxidation unit 16 to the
primary guard bed 6 at a temperature in the range of from
30.degree. C. to a temperature slightly greater than the sulfur dew
point, carrying out acidic gas removal at a temperature greater
than the acidic gas absorption temperature and then directing such
stream flow to the H.sub.2S direct oxidation unit 16. Interstage
cooling between recycling is accomplished via initial and secondary
separators 18, 19 which are provided to remove the produced sulfur
from the recycle stream. The effluent gas recycling process is
repeated until the H.sub.2S concentration in the recycle gas is
about 10 ppmv. Before cooling down to room temperature, the primary
acidic gas removal unit 7 is purged with an H.sub.2S and
CO.sub.2-free gas such as N.sub.2 (stream 4, by adjusting valve 35
to permit flow thereof) and the off gas stream from the purging
process is mixed with the feed gas stream of the secondary guard
bed/acidic gas removal unit B. Meanwhile, the pressure of the
direct oxidation reactor 16 is reduced to 60 psig, producing a
CO.sub.2 stream of purity greater than 99.9% by volume.
[0074] In a second embodiment, the acidic gas removal unit 7, 7a
according to the present process comprises amine based sorbents
suitable for the removal of the acidic gases from acidic gases
containing streams and for the hydrolysis of COS to H.sub.2S and
CO.sub.2 at low temperatures. These amines include but are not
limited to 1,4-Diazabicyclo[2,2,2]-Octane,
1,5-Diazabicyclo[5,4,0]-Undec-5-ene, 1,4-dimethylpiperazin-2-one
and 1,5-Diazabicyclo[4,3,0]-non-5-ene. These amines can be in the
liquid form or supported on any type of the porous solid support
systems known in the art. The use of these amines is advantageous
in the case of using feed streams of high CO content because it
eliminates the necessity of the purging step required for the
removal the CO from the acidic gas removal units.
[0075] In a third embodiment, one or both of the acidic gas removal
units 7, 7a according to the present process comprise amine based
sorbents of high selectivity toward H.sub.2S. The amines suitable
for manufacturing the sorbents according to the present process
include but are not limited to one or more of N-methylpyrrolidone
(NMP)/dodecane, 1,4-Diazabicyclo[2,2,2]-Octane and
diisopropanolamine. These amines can be in the liquid form or
supported on any type of the porous solid support systems known in
the art. The benefits of utilizing the high H.sub.2S selective
amine sorbents is that it can handle a large volume of the acidic
gas containing streams and increase the concentration of the
H.sub.2S in the off-gas stream from the acidic gas removal unit 7,
7a.
[0076] In a fourth embodiment according to the present invention,
an H.sub.2S and/or CO.sub.2 containing stream is supplied to a
primary acidic gas removal unit 7 without pretreatment. The acidic
gas removal unit comprises amine based sorbents suitable for the
COS hydrolysis to H.sub.2S and of high H.sub.2S absorption
selectivity. Once the acidic gases have broken through, the primary
acidic gas removal unit 7 is purged at room temperature with
N.sub.2 gas to remove the residual CO gas, if present in the feed
gas stream, and is then conducted to a desorption mode at a
temperature higher than the absorption temperature. The effluent
stream from the primary acidic gas removal unit 7 is mixed with a
continuous flow of a molecular oxygen containing stream 1 and the
oxygen to H.sub.2S ratio in the resultant gas mixture is
deliberately adjusted to a ratio less than 0.5 to avoid the
oxidation of the H.sub.2S to SO.sub.2. The resultant gas mixture is
then supplied to an H.sub.2S direct oxidation system 30 having an
H.sub.2S direct oxidation reactor 16 containing any H.sub.2S
oxidation catalyst known in the art to partially oxidize the
H.sub.2S in the gas mixture to elemental sulfur. Similarly, the
high desulfurization level of the gas mixture can be achieved in a
batch process by recycling the off-gas stream from the H.sub.2S
direct oxidation reactor 16 to the acidic gas removal unit 7, 7a at
a temperature greater than the acidic gas absorption
temperature.
Example 1
[0077] This example illustrates the first embodiment. In this
example, the acidic gas removal unit 7 comprises a porous
solid-supported amine sorbent to remove the acidic gases from the
feed stream. The supported amine sorbent utilized in this example
has a high absorption capacity for H.sub.2S ad CO.sub.2.
Synthesis of the Sorbent
[0078] The supported amine sorbent was synthesized similarly to
reported procedure (see, U.S. patent Ser. No. 13/399,911 filed Feb.
17, 2012). The absorbent was manufactured in small fractions which
were combined. The surface physical characteristics of the support
utilized are shown in Table 1.
TABLE-US-00001 TABLE 1 Physical characteristics of the absorber
supports Absorber support Examples 1 and 2 Example 3 and 4 Support
Code Degussa 4041 Alcoa LD-5 Surface Area, m.sup.2/g 155 300 min
Pore Volume, cc/g 0.9-1.0 0.63 Bulk Density, g/cc 0.4400-0.460
0.465 A1.sub.20.sub.3, % wt <500 ppm 99 S102, % wt >99.8 0.40
max Fe203, % wt, max <30 ppm 0.04
[0079] Approximately 500 ml of the synthesized sorbent particles
were enclosed between two glass wool zones and loaded into a down
flow stainless steel absorber column. The absorber column was
pretreated with a N.sub.2 gas stream at a temperature of
130.degree. C. for 2 hours. The acidic gas absorption mode was
conducted at room temperature and a pressure of 100 psig. Two
different gas streams were used to prepare a synthetic feed gas of
a composition shown in Table 2, which is similar to the composition
of the gas produced from the THAI.TM. process.
TABLE-US-00002 TABLE 2 Synthetic feed gas composition. Component %
by volume H.sub.2 1.83 O.sub.2 00 N.sub.2 75.41 CH.sub.4 5.49 CO
1.04 CO.sub.2 15.73 H.sub.2S 0.50
[0080] The stream 2,3 containing acidic gases was passed through
the absorber column 7 with a flow of 330 ml/min and the
breakthrough time of the acidic gases was determined by a micro gas
chromatograph 11 equipped with an automated stream selection valve.
Once the acidic gases broke though, the feed gas stream 2,3 was
switched to a secondary absorber column system B and the primary
absorber column 7 was purged with a N.sub.2 gas stream 4 to remove
the residual CO. Subsequently, the temperature of the absorber
column 7 was increased gradually to 130.degree. C. to free the
adsorbed H.sub.2S and CO.sub.2, and the pressure of the absorber
column 7 increased from 100 psig to about 150 psig. The temperature
of the absorber column 7 was kept fairly constant at a temperature
of 130.degree. C. to avoid the re-adsorption of the H.sub.2S and
CO.sub.2 during the circulation of the off-gas stream from the
absorber column 7. The off-gas stream from the absorber column 7
was mixed with a continuous flow of air and the resultant gas
mixture was then fed to an H.sub.2S direct oxidation reactor 16 via
a gas circulating pump 12. The air flow was adjusted such that the
ratio of oxygen to H.sub.2S was less than 0.5.
[0081] FIG. 2 (Line 1) shows the variation in the H.sub.2S
concentration in the recycle gas stream from the absorber column 7
during the gas circulation step. At this point, the recycle gas
stream from the absorber column 7 became the feed gas stream of the
H.sub.2S direct oxidation reactor 16.
[0082] The H.sub.2S oxidation reactor 16 was loaded with 20 ml of
an alumina-supported bismuth/copper oxidation catalyst and the
H.sub.2S oxidation reaction was conducted at a temperature of
220.degree. C. and a pressure of 100 psig. The flow rate of the
feed gas stream of the oxidation reactor 16 was adjusted via a mass
flow controller 13 mounted on the recycle gas stream from the
circulating pump 12 to supply the feed gas stream to the H.sub.2S
direct oxidation reactor 16 at a gas hourly space velocity of 1000
hr.sup.-1. The produced fluid from the H.sub.2S oxidation reactor
16 entered a sulfur knockout separator 18 to remove the sulfur from
the product gas stream. A 2.mu. stainless steel filter was also
employed to capture the trace of the sulfur.
[0083] FIGS. 3, 4 and 5 (Line 1) illustrate the variation in the
H.sub.2S, COS and SO.sub.2 respectively in the product gas from the
H.sub.2S direct oxidation reactor 16 during the circulation step.
From FIGS. 4 and 5 (Line 1) small amounts of SO.sub.2 and COS were
produced as byproducts from the H.sub.2S oxidation reactor 16. The
produced gas from the sulfur knockout separator 18 was passed
through a humidifier comprising water at a temperature of
80.degree. to increase the water partial vapor pressure as required
for the hydrolysis of the COS present in the product gas. The
humidified product gas was then recycled to the protective guard
bed 6.
[0084] The productive guard bed 6 was loaded with 10 ml of pure
alumina catalyst and operated at the same system pressure (about
100 psig) and at a temperature of 120.degree. C. The outlet stream
from the protective guard bed 6 was cooled down and then fed to the
absorber column 7. The small amount of the COS produced during the
H.sub.2S oxidation was hydrolyzed to H.sub.2S, while SO.sub.2 was
reduced to elemental sulfur in the guard bed 6. Therefore, no COS
or SO.sub.2 was detected and H.sub.2S was the only sulfur compound
in the off-gas stream from the guard bed 6. The off-gas steam from
the protective guard bed 6 was then recycled to the absorber column
7. The gas circulation process was repeated until the H.sub.2S in
the recycle gas was 10 ppm. Subsequently, the absorber column 7 at
a temperature of 130.degree. C., was purged with a N.sub.2 gas
stream 4 to avoid the re-adsorption of CO.sub.2.
Example 2
[0085] As a further illustration of the First Embodiment, this
Example is identical to Example 1 except the absorber column 7 was
not purged to remove the residual CO after the acidic gas
absorption step. FIGS. 2 and 3 (Line 2) respectively show the
variation in the H.sub.2S concentration in the recycle gas stream
from the absorber column 7 and from the oxidation reactor 16 during
the gas circulation step. As a consequence of the presence of CO, a
considerable amount of COS was produced during the H.sub.2S direct
oxidation reaction, FIG. 4 (Line 2). In addition, a sudden increase
in the SO.sub.2 concentration was detected in the outlet gas stream
of the H.sub.2S direct oxidation reactor toward the completion of
the oxidation cycle, FIG. 5 (Line 2). This can be attributed to the
sudden increase in the oxygen-to-H.sub.2S ratio toward the
completion of the oxidation cycle. Nevertheless, no COS or SO.sub.2
was detected in the recycle gas from the guard bed and the overall
desulfurization process duration increased significantly due to the
low sulfur selectivity during the H.sub.2S direct oxidation
reaction to elemental sulfur.
Example 3
[0086] This is an illustration of the Second and Third Embodiments.
This example is identical to Example 2 except that the acidic gas
removal unit contained an amine based sorbent of high H.sub.2S
selectivity and is suitable for COS hydrolysis to H.sub.2S and
CO.sub.2.
Synthesis of the Sorbent
[0087] The synthesis of the absorber was conducted by ordinary
methods as practiced by those knowledgeable in the art. The amine
based absorber support was Alumina spheres (LD-5) obtained from
Alcoa. The physical characteristics of the support are shown in
Table 1. Approximately, 25.5 g of 1,4-Diazabicyclo[2,2,2]-Octane
was dissolved in acetone and the solution was added to 427.2 gm of
the alumina support by the method of incipient wetness to achieve
5.6 wt. % amine in the final sorbent. The absorbent was left in the
air to dry over night. Subsequently, the absorbent was loaded in
the absorber column 7 and then conditioned at a temperature of
105.degree. in a flow of nitrogen for 3 hours. The acidic gas
absorption mode was conducted at room temperature and a pressure of
100 psig utilizing a gas stream of a composition similar to the gas
stream employed in Examples 1 and 2. After the H.sub.2S has broken
through, the absorber column 7 was conducted to a desorption mode
at a temperature of 120.degree. C.
[0088] The breakthrough time of the H.sub.2S from the acidic gas
removal unit 7 increased significantly upon using the hindered
amine based sorbent and therefore, the desulfurization step of the
desorbed gas from the acidic gas removing unit 7 was expected to be
longer than the acidic gas removal step. However, for a continuous
sulfur removal process, the desulfurization step of the desorbed
gas from the primary absorber column 7 was operated at low overall
desulfurization efficiency and was deliberately terminated when the
H.sub.2S in the recycle gas stream from the primary absorber column
7 was about 1750 ppm FIG. 2 (Line 3). Subsequently, the primary
absorber column 7 was cooled down and therefore, the overall
pressure of the system decreased 60 psig. The inlet feed stream of
the H.sub.2S direct oxidation reactor 16 was then switched to the
outlet gas stream from the secondary absorber column 7a. Meanwhile
the primary absorber column 7 was cooled down further to room
temperature and then conducted to a new acidic gas removal cycle.
Similarly, no COS or SO.sub.2 was detected in the recycle gas
stream from the protective guard bed 6.
Example 4
This is an Illustration of the Fourth Embodiment
[0089] This example is identical to Example 3, except that the
protective guard bed 6 (and 6a) was eliminated from the process.
The oxygen to H.sub.2S ratio in the feed gas stream of the H.sub.2S
oxidation reactor was adjusted to a ratio less than 0.5 to prevent
the oxidation of H.sub.2S to SO.sub.2. Therefore, no SO.sub.2 was
detected in the outlet gas stream during the desulfurization step.
However, a considerable amount of the COS was detected in the
recycle gas stream from the H.sub.2S direct oxidation reactor 16
FIG. 4 (line 4). The recycle gas stream from the oxidation reactor
16 was cooled down to a temperature of 50.degree. C. to condense
the produced sulfur and the moistened off-gas stream from the
sulfur secondary separator was then recycled to the primary
absorber column 7 to hydrolyze the produced COS to H.sub.2S and
CO.sub.2. Typically, the oxidation of H.sub.2S to elemental sulfur
produces water (reaction 3), therefore, no additional water was
required for the hydrolysis of the produced COS to H.sub.2S in the
primary absorber column 7.
* * * * *