Selection Criteria for Claus Tail Gas Treating Processes
Mahin Rameshni, P.E.
Technical Director, Sulphur Technology and Gas Processing
181 West Huntington Drive, Monrovia, California 91016, USA
Mahin.Rameshni@WorleyParsons.com
Introduction
With the sulphur content of crude oil and natural gas on the increase and tightening sulphur
content in fuels, refiners and gas processors are pushed for additional sulphur recovery capacity.
At the same time, environmental regulatory agencies of many countries continue to promulgate
more stringent standards for sulphur emissions from oil, gas and chemical processing facilities. It
is necessary to develop and implement reliable and cost effective technologies to cope with the
changing requirements. In response to this trend, several new technologies are now emerging to
comply with the most stringent regulations.
Typical sulphur recovery efficiencies for Claus plants are 90-96% for a two- stage, and 95-98%
for a three- stage plant. Most countries require sulphur recovery efficiency in the range of
98.5% to 99.9+%. Therefore the sulphur constituents in the Claus tail gas need to be reduced
further.
The key parameters effecting the selection of the tail gas clean-up process are:
f Feed Gas composition, including H2S content and hydrocarbons and other contaminants
f Existing equipment and process configuration
f Required recovery efficiency
f Concentration of sulphur species in the stack gas
f Ease of operation
f Remote location
f Sulphur product quality
f Costs (capital and operating)
-1-
Various aspects and considerations when choosing the most optimum process configuration for
tail gas treating are discussed. There are several key features effecting the selection of the tail
gas clean-up process that three steps should be taken. When required recovery efficiency and
concentration of sulphur species in the stack gas is known, selection of the tail gas process is one
step closer. The first step is one the most important criteria for the selection of the tail gas treating
processes. When the required sulphur recovery is established, the selection of the tail gas
process will be limited. Table 1 represents the various tail gas clean-up process with the recovery
will be achieved. When concentration of impurities in the acid gas such as COS and CS2, H2S
content, and feed gas composition, and finally treated gas specifications are established, the type
of amine used for a particular application could be selected in step two. Finally the third step is
the evaluation between the identical process chosen for ease of operation, capital and operating
cost, and remote location. For revamp units, minimum equipment modifications and process
configuration should be considered as a main key factor.
The WorleyParsons BSR Amine process for Claus tail gas treatment clearly represents Best
Available Control Technology (BACT), potentially achieving 99.99+% overall sulphur recovery
with emissions of < 10 ppmv H2S and 30 ppmv total sulphur.
There are other processes such as direct oxidations processes, Sub dew point processes that are
able to achieve higher sulphur recovery from the conventional 3-stage Claus unit up to 99.8 %
depending on the feed compositions to the Claus unit.
WorleyParsons offer DEGSULF a sub dew point process with partnership with DEG-ITS for those
applications with the relaxed overall recovery.
Brief History
Under the leadership of David Beavon of the Ralph M. Parsons Company, Parsons and Union Oil
of California (Unocal) co-developed the Beavon Sulphur Removal Process (BSRP) in San Pedro,
California, in the late 1960s, for which US Patent 3,752, 877 was awarded to Parsons in 1973.
The fundamental process, still employed today, typically heats the Claus tail gas to 550-650°F (~
290-340°C) by inline sub-stoichiometric combustion of natural gas in a Reducing Gas Generator
(RGG) for subsequent catalytic reduction of virtually all non-H2S sulphur components to H2S.
Conversion of SO2 and elemental sulphur (Sx) is by hydrogenation:
SO2 + 3 H2 → H2S + 4 H2O + ΔH
Sx + x H2 → x H2S + ΔH
Conversion of COS and CS2 is by hydrolysis:
COS + H2O → H2S + CO2 + + ΔH
CS2 + 2 H2O → 2 H2S + CO2 + ΔH
-2-
CO is essentially hydrolyzed to yield additional H2 according to the “water gas shift” reaction:
CO + H2O → H2 + CO2 + ΔH
CO and H2 naturally present in the Claus tail gas will typically satisfy up to 70% of TGU demand,
with the balance generated in the RGG.
A cobalt-moly catalyst, similar to hydrodesulphurization catalyst, is typically employed. As
received, the catalyst is an alumina substrate impregnated with oxides of cobalt and molybdenum
which must be converted to the active sulfided state. To convert the cobalt oxide to the sulfide, a
simple exchange of the oxide with H2S is all that is necessary:
CoO + H2S → CoS + H2O + ΔH
Converting molybdenum trioxide to the active disulfide, however, requires a change in oxidation
number that also requires hydrogen:
MoO3 + 2 H2S + H2 → MoS2 + 3 H2O + ΔH
The reduced tail gas is then cooled to 90-100°F (~ 30-40°C) to condense most of the water
vapor, which accounts for ~ 35% of the stream. While Beavon recognized the potential for H2S
recovery using an alkanolamine, he was concerned about formation of heat stable thiosulfate
resulting from SO2 breakthroughs. Consequently, Parsons adopted the Stretford redox process
which employed an alkaline vanadium salt solution to oxidize absorbed H2S to elemental sulphur
particles which were subsequently removed by froth flotation, filtered and melted.
The Beavon Stretford process actually had some advantages over amine absorption:
f No acid gas recycle to the Claus unit
f No steam consumption
f < 5 ppm residual H2S, obviating incineration
f Temporary high capacity for excessive Claus tail gas H2S or SO2 resulting from off-ratio
operation
However, these were outweighed by poor sulphur quality, high chemical makeup costs, high
disposal costs from purging of byproduct thiosulfate, absorber fouling, oxidizer foaming,
inconsistent froth formation, troublesome filter operation and atmospheric corrosion. By the
1980s, Parsons essentially abandoned the Stretford process in favor of MDEA absorption/
regeneration. Today, WorleyParsons retains the BSR trademark in reference to the catalytic
reduction stage and subsequent cooling/condensation. A typical BSR Amine system is shown in
Figure 1. Figure 2 is a typical BSR amine system including the start up blower.
-3-
Reducing Gas Generator (RGG)
Many competitors use the inline burner design in Figure 3. Burner vibration is common, the
extreme temperature gradient between the combustion and tail gas mixing zones makes it difficult
to optimize skin temperatures at the transition, and a combustion zone shell leak resulting from
localized refractory failure forces a shutdown.
The proposed BSR unit comprises of three process steps:
f Reducing Gas Generation (RGG) and tail gas preheat
f Hydrogenation/Hydrolysis of SO2 and other sulphur species to H2S
f Gas cooling and waste heat recovery
f WorleyParsons proprietary RGG design provides process gas reheating and reducing gas
(H2 and CO) generation in one single process unit. No external supply of hydrogen gas is
required. This feature enhances the reliability of the process unit by eliminating the
uncertainties associated with the availability of external hydrogen supply and the quality of
hydrogen gas.
WorleyParsons design has the following advantages compare to the other licensors.
WorleyParsons Conventional BSR Section
f Start Up Blower – reduce the emission during start up to the stack
f Caustic wash – reduce breakthrough of SO2 to amine during start up and prevent
degrade of the solvent, No SO2 breakthrough to the amine unit
f Stable operation for different mode of operation
f Good Turn down
f No external hydrogen required
f Less pressure drop
f Less heat loss in RGG configuration
f Could be started up independently from SRU’s
f No vibration in RGG Burner
f No heat Loss in RGG
-4-
f No Refractory damages as the results of un-uniform heat distribution
f Proprietary RGG design provides process gas reheating and reducing gas (H2 and CO)
generation in one single process unit. External supply of hydrogen gas is not required in
most cases.
f Start up blower, to eliminate violating the emission during the start up
-5-
Figure 1 – WorleyParsons BSR Amine Flow Scheme
-6-
Figure 2 – WorleyParsons BSR Amine Flow Scheme with start up blower
-7-
Figure 3– Common TGU Feed Heater
By comparison, the WorleyParsons design (Figure 4) employs a brick-lined internal combustion
zone for stable combustion unaffected by downstream turbulence. Optimum outer-shell skin
temperatures are easily ensured, heat loss is minimized and potential leakage through the
combustion zone wall does not result in atmospheric release. Some units have been in service
for 30+ years with no major refractory repairs. The RGG is typically elevated so that minor
entrained sulphur will free-drain to the reactor (and vaporize).
Figure 4 – WorleyParsons Reducing Gas Generator (RGG)
Industry consensus is apparently lacking with regard to the optimum air/fuel ratio. Many
competitors’ units operate at stoichiometric air and rely on supplemental H2 for hydrogenation of
SO2 and Sx. Perhaps contrary to intuition, equilibrium O2 is nominally 0.6 % at stoichiometric air,
and only goes to zero at < 90% of stoichiometric. There is experience to suggest that chronic O2
leakage leads to catalyst sulfation, although there is disagreement within the industry on this
point. Nonetheless, WorleyParsons generally recommends operating at 80% of stoichiometric to
avoid, or at least minimize, O2 leakage (and also maximize H2 yield).
The advisability of supplemental H2 is also a source of controversy. Many clients consider the
availability of import H2 necessary to minimize the risk of SO2 breakthroughs, whereas in reality it
is as easy to reduce Claus combustion air (with the same effect) as increase H2 addition. In the
absence of supplemental H2, the operator quickly learns the value of monitoring residual H2 as a
-8-
sensitive indicator of Claus tail gas ratio, and arguably is more likely to routinely optimize Claus
air demand when constrained by a limited H2 supply. Three-stage Claus units clearly do not need
supplemental H2, while residual H2 may be marginal with 2-stage units, in which case
supplemental H2 may be advisable to ensure ability to optimize the Claus tail gas H2S/SO2 ratio.
H2 analyzers based on thermal conductivity measurement are very reliable, with minimal
servicing. Where the TGU is coupled to a single Claus train, the H2 analyzer can in fact supplant
the Claus air demand analyzer. Where multiple Claus trains are coupled to a single TGU,
combustion air to a Claus unit whose air demand analyzer is out of service can be temporarily
adjusted based on TGU residual H2.
LP steam injection to the burner in the nominal ratio of 1/1 lb/lb steam/fuel is generally advisable
for soot inhibition when firing sub-stoichiometrically, by virtue of the following reactions:
C + H2O → CO + H2 - ΔH
C + 2 H2O → CO2 + 2 H2 - ΔH
While modern high-intensity burners may be operable at as low as 80% of stoichiometric air
without steam injection, injection is still prudent in view of the possibly of lower air/gas ratios
resulting from meter error or localized fuel-rich zones due to burner damage or fouling. With high
intensity burners, steam injection via a dedicated steam gun is preferred. Otherwise, injection
into the combustion air is the most practical.
Figure 5 – WorleyParsons Reducing Gas Generator (RGG) Details
-9-
Hydrogenation Reactor
With good catalyst activity and no excessive HCs in the acid gas feed to the Reaction Furnace,
organic residuals in the Absorber offgas should be as shown in Table 1:
Table 1 – Residual Sulphur with Fresh Catalyst
Contaminant
PPMV
Carbonyl sulfide (COS)
< 20
Carbon monoxide (CO)
< 200
Carbon disulfide (CS2)
0
Methyl mercaptan (CH3SH)
0
With fresh conventional catalyst, temperatures of 400-450°F (204-232°C) are typically required to
initiate the hydrogenation reactions, and 540-560°F (282-293°C) for hydrolysis. As the catalyst
loses activity with age, progressively higher temperatures may be required. Typically, activity
loss is first evidenced by (1) reduced COS, CS2 and CO conversion, and (2) potential methyl
mercaptan formed by the reaction of CS2 and H2, while hydrogenation of SO2 and Sx may still be
complete because of the lower initiation temperatures required.
The potential formation of methyl mercaptan at low temperature or impaired catalyst activity is
perhaps not widely appreciated. In cases where the TGU tail gas is discharged without
incineration, nominal mercaptan levels can result in serious nuisance odors. In Stretford units,
there is reason to expect that the mercaptan is oxidized to disulfide oil (DSO) which can impair
froth formation.
Excessive HCs in the SRU acid gas feed will tend to increase the carbon-sulphur compounds in
the Reactor effluent. In the Figure 4 example, HCs in the amine acid gas are evidenced by (1)
increased air demand per volume of gas, (2) increased tail gas volume resulting from the
additional air and HC combustion products, and (3) increased Total Reduced Sulphur (TRS) in
the Absorber offgas – predominantly COS, but also potentially including CS2 and methyl
mercaptan (RSH). (While TRS also includes H2S, the H2S content did not increase in this case.)
Low Temperature Hydrogenation Catalyst
WorleyParsons has started offering low temperature catalyst if requested by client as
applicable and meet the project emissions. Low temperature catalyst eliminate of using the
reducing gas generator and indirect heating system could be used instead. Low-temperature
TGU catalysts reportedly capable of operating at inlet temperatures of 210-240°C (410-464°F),
achievable with steam reheat, have recently become available. The primary advantage (in a
new unit) is elimination of the RGG, translating to (1) lower capital cost, (2) operating
simplicity, (3) improved turndown, (4) reduced TGU tail gas volume, (5) reduced CO2 recycle to
the SRU, and (6) elimination of risk of catalyst damage by RGG misoperation.
- 10 -
Historically, Claus tail gas treating units (TGTU) have required reactor inlet temperatures of ~
550°F for appreciable hydrolysis of COS, CS2 and CO, typically requiring preheat by inline firing
or heat exchange with hot oil or heat transfer fluid.
Vendor claims of energy savings are questionable since they tend to (1) assume the plant is long
on LP steam, and (2) disregard the cost of HP steam. Long term performance of low-temperature
catalysts is still uncertain. The following considerations should be taken into account:
f A steam reheater will limit the ability to compensate for normal catalyst activity loss with
age, potentially limiting its useful life.
f A bottom layer of titania in the first Claus converter may be required for COS/CS2
hydrolysis.
f Higher residual CO levels could mean operating the incinerator at 1500°F (~ 800°C)
instead of 1200°F (~ 650°C).
f Incomplete CS2 destruction, and hence methyl mercaptan formation, can result in serious
nuisance odors if the TGU tail gas is discharged without incineration.
Reactor inlet temperatures are only half the story; outlet temperatures are the other half. Any
catalyst will probably initiate SO2 hydrogenation at 400-450°F (~ 205-230°C) and, with sufficient
temperature rise and excess catalyst, will subsequently achieve virtually complete hydrolysis.
New catalysts by Criterion and Axens require lower activation temperatures achievable by indirect
reheat by 600# steam, thus reducing investment cost, operating complexity and, in some cases,
energy consumption. In addition, lower reactor outlet temperatures may obviate the downstream
waste heat boiler.
While reduced investment and complexity are a given, whether the claimed energy savings is real
is site-specific. Reduced feed preheat energy only constitutes a savings if the plant is already
long on low-pressure waste heat steam (40-70 psig). Otherwise, incremental heat input is fully
recovered. Furthermore, in the absence of a steam surplus, elimination of the waste heat boiler
may have forfeited recoverable BTUs.
Relative COS, CS2 and CO conversion efficiencies need to be compared. It is not necessarily
sufficient to achieve regulatory compliance.
COS, CS2 and CO Hydrolysis using low temperature catalyst
Relative COS, CS2 and CO conversion efficiencies can be critical. It is not necessarily sufficient
to achieve regulatory compliance.
f Regulations could become more stringent in the future.
- 11 -
f Some plants must also buy emission credits per pound of SO2 discharged.
f Excessive CO residuals could require higher incinerator temperatures, or require
incineration otherwise obviated in units able to achieve TGTU absorber H2S emissions <
10 ppm by the use of acid-aided MDEA.
Hydrolysis of COS, CS2 and CO typically requires higher temperatures than hydrogenation of SO2
and Sx. Perhaps accordingly, COS, CS2 and CO conversion efficiencies are the first to suffer as
conventional catalysts lose activity with age. Higher reactor inlet temperatures will tend to
compensate for deactivation, thus extending catalyst life considerably. Depending on the design
limits, temperatures can generally be increased by 50-150°F (28-83°C).
Assuming the same holds true for the low temperature catalysts, a steam reheater will
substantially limit the extent to which temperatures can be increased, in effect potentially
shortening catalyst life. The lower initiation temperature of the Criterion 734 at start-of-run is thus
significant, as it affords the greatest margin for increase.
At 464°F (240°C) – generally the limit of a 600# steam reheater – hydrolysis of CO, COS and CS2
approaches that of conventional high temperature catalysts. At 428°F (220°C), however, Axens
concedes that COS/CS2 conversion must be accomplished in the 1st Claus stage by (1)
supplementing the alumina bed with a bottom layer of expensive titania catalyst, or (2) increasing
the inlet temperature to 550-600°F (288-316°C). The latter will nominally
f reduce Claus recovery efficiency from
f increase SRU tail gas rate
f increase TGTU sulphur load
However, the 1st stage will not effect CO conversion.
Conventional cobalt-moly catalyst will generate minor, but significant, levels of methyl mercaptan
by the reaction of CS2 and hydrogen at 480°F (249°C) when in good condition, and at much
higher temperatures if the catalyst is aged or damaged. While the manufacturers claim no
residual mercaptans with the low temperature catalysts, there is some uncertainty – in the
author’s view – as to whether that will remain true a few years into the run.
Hydrogen Balance using low temperature catalyst
Compared with firing the feed heater at stoichiometric air and importing H2, a steam reheater will
of course have no impact on the H2 balance. However, many plants avoid the need for
supplemental H2 by the use of a reducing gas generator (RGG), typically burning natural gas substoichiometrically to generate H2 and CO.
- 12 -
In the absence of an RGG, the alternative is to operate the SRU more air-deficient as necessary
to maintain, say, 2% residual H2 downstream of the TGTU reactor. This will nominally
f reduce Claus recovery efficiency
f increase SRU tail gas rate
f increase TGTU sulphur load
CO2 Balance using low temperature catalyst
Eliminating the inline burner has the benefit of reducing the TGTU tail gas volume (for the
assumed basis with an RGG). Assuming 85% CO2 slip, the acid gas load on the TGTU amine is
reduced.
Energy Balance using low temperature catalyst
A steam reheater will not only eliminate the following natural gas required by the RGG, but will
also reduce incinerator fuel by virtue of the reduced tail gas rate:
f RGG fuel savings
f Incinerator fuel savings
Assuming H2S/SO2 = 2 in the SRU tail gas, of supplemental H2 will be required to maintain a 2%
residual in the TGTU tail gas. As a rule-of-thumb, the value of relatively pure (non-reformer) H2 is
four times that of natural gas.
Figure 6 represents WorleyParsons BSR/amine with the low temperature catalyst.
- 13 -
SRU TAIL GAS
STARTUP
BLOWER
H2
RECYCLE
WATER
HP STEAM
CONTACT
CONDENSER
SOUR WATER
BLOWDOWN
HYDROGENATION
REACTOR
DESUPERHEATER
REDUCED TAIL GAS
10% NaOH
TREATED TAIL GAS TO ATMOSPHERE OR INCINERATOR
ABSORBER
ACID GAS
RECYCLE
TO SRU
REFLUX
INTERMITTENT
PURGE TO SWS
REGENERATOR
RICH AMINE
PROCESS
STEAM
LEAN AMINE
Figure 6 – WorleyParsons BSR Amine Flow Scheme with Low Temperature Catalyst
- 14 -
Figure 7 – WorleyParsons Impact of Hydrocarbons in Acid Gas to SRU
In the event of a burner trip, there is usually ample time to relight the RGG before the reactor bed
cools to the point of SO2 breakthrough. In the Figure 7 example, relight was delayed by a
plugged pilot fuel gas restriction orifice, and the main burner was down ~ one hour (65 minutes).
At all times at least one point in the bed was 510ºF or higher, which likely explains the absence of
an SO2 breakthrough. By the end of, say, a 2-hour outage, all temperatures would have been <
400ºF, and it is possible that serious SO2 breakthrough would thus start to occur within 1½-2
hours.
The reactor contained 37.5 Mlb of Criterion 534 cobalt-moly catalyst, a 2-Mlb top layer of ½”
alumina and a 4.5-Mlb support layer of ceramic balls.
- 15 -
Reactor Temperatures
800
700
Temperature, F
600
500
400
300
11:00
10:55
10:49
10:44
10:39
10:34
10:29
10:24
10:20
10:15
10:09
9:59
10:05
9:54
9:49
9:44
9:39
9:34
9:29
9:25
9:20
9:14
9:10
9:05
9:00
8:55
8:50
8:45
8:40
8:35
8:30
8:25
8:20
8:14
8:09
8:05
8:00
200
Figure 8 – WorleyParsons Hydrogenation Reactor Bed Temperatures During RGG Outage
The total tail gas rate is shown in Figure 9. There are actually two identical reactors in parallel,
with only half of the indicated flow through each.
- 16 -
TGTU Tail Gas Rate
1400
1350
Tail gas rate, MSCFH
1300
1250
1200
1150
1100
1050
Figure 9 – WorleyParsons TGU Tail Gas Rate During RGG Outage
Resultant TRS (measured at the absorber outlet) and SOx emissions are shown in Figure 10.
- 17 -
11:00
10:54
10:49
10:44
10:39
10:34
10:29
10:24
10:19
10:15
10:09
10:04
9:55
10:00
9:50
9:44
9:40
9:35
9:29
9:24
9:19
9:14
9:09
9:04
8:59
8:54
8:49
8:45
8:40
8:35
8:29
8:24
8:19
8:15
8:09
8:05
8:00
1000
Emissions
400
350
absorber TRS
F-754
PPM, corrected to air-free basis
300
250
200
150
100
50
10:59
10:54
10:49
10:44
10:39
10:35
10:29
10:24
10:19
10:14
10:09
9:59
10:04
9:54
9:49
9:45
9:39
9:34
9:30
9:25
9:20
9:15
9:10
9:05
9:00
8:55
8:50
8:45
8:40
8:34
8:30
8:25
8:19
8:14
8:10
8:04
8:00
0
Figure 10–WorleyParsons Impact of RGG Outage on Emissions
Contact Condenser (2-Stage Quench)
Common industry practice is to cool the reduced tail gas from the reactor by the generation of LP
waste heat steam followed by direct quench with a recirculating water stream to cool it to 90100°F (~ 30-40°C), thus condensing most of the water vapor which accounts for ~ 35% of the
stream.
WorleyParsons utilizes a unique 2-stage tower comprised of a bottom Desuperheater section and
top Contact Condenser.
f The contact condenser has 2 sections, the first section de-superheats the gas and scrub
any SO2 may breakthrough from hydrogenation reactor, and the second section cools the
gas and condensate the water, therefore there is no need for make up water to maintain
the caustic concentration. The condense water will provide the water to maintain the
caustic concentration. We do not have continuous purge, but we provide water make up
for the water is evaporated, just like any other quench system.
- 18 -
f Tail gas is desuperheated in the lower section of the contact condenser by a circulating
water stream. This water is maintained alkaline to protect against any SO2 breakthrough
from the reactor. In the upper packed section of the tower, most of the water vapor in the
tail gas is condensed by direct contact with a circulating stream of cooled water. A pH
analyzer with a low-pH alarm is installed in the quench water circulation line and will
indicate when the pH of the quench water is reducing, from either a breakthrough of SO2,
or incomplete reduction of the sulphur compounds in the gas stream from the
Hydrogenation Reactor.
(Figure 1 and 2) A 10 %-wt NaOH solution is recirculated through the Desuperheater to capture
SO2 potentially resulting from a process upset, while also cooling it to its dewpoint of ~ 165°F (~
75°C). The only cooling is by vaporization. The gas is further cooled to 90-100°F (~ 30-40°C) by
direct contact with an externally cooled recycle water stream in the upper Contact Condenser
section. A recycle water slipstream is returned to the Desuperheater on Desuperheater levelcontrol via two bubble-cap wash trays to capture entrained caustic.
A blowdown slipstream of recycle water is purged, usually to sour water, on Contact Condenser
level-control. While the recycle water is usually classified as sour water, the H2S content is
typically < 50 ppmv by virtue of CO2 saturation. In situations where the increased load on the
plant sour water stripper is undesirable, a simple blowdown stripper is occasionally provided at
the TGU. This typically involves LP stripping steam injection (as opposed to a reboiler) and
return of the uncondensed overhead stream to the Desuperheater.
Startup Blower
WorleyParsons provide a start up blower on the contact condenser overhead to eliminate flaring
large quantities of H2S to atmosphere and to prevent violation of the emission. For those cases
that a booster blower required then booster blower will have dual function as a start up blower
and as a booster blower.
Booster Blower
Many of the Claus units that are in operation do not have enough pressure to handle a new tail
gas unit in other words the provision of operating the Claus unit at the higher pressure was not
considered, if the source pressure changed the existing amine unit requires higher reboiler duty
and in most cases required significant changes in the amine unit. WorleyParsons have been
offering a booster blower in the tail gas unit to overcome the pressure limitation.
Retrofit Tail Gas Units will typically require a booster blower downstream of the Contact
Condenser to overcome the additional pressure drop. The blower is located after the Contact
Condenser to minimize the actual volume (by virtue of cooling and condensation), and before the
Absorber to take advantage of the higher pressure.
With proper design and operation, booster blowers are inherently very reliable, requiring minimal
maintenance. Typically, the case is cast iron or carbon steel, with an aluminum impellor. N2-
- 19 -
purged tandem shaft seals (typically carbon rings) eliminate process leakage to atmosphere on
the discharge end as well as air aspiration into the process on the suction end, which is typically
at a vacuum.
Though often viewed as a liability by clients, booster blowers arguably improve operability in
several ways:
f By recirculating tail gas, the TGU can be started up and shut down independent of the
SRUs.
f Tail gas recycle ensures process stability at high SRU turndown by (1) avoiding undue
RGG burner turndown potentially conducive to sooting due to poor mixing or air/gas
flowmeter inaccuracy, and (2) diluting potentially high SO2 levels often typical of high SRU
turndown. With advance warning, tail gas recycle can avoid RGG shutdown in the event of
an SRU trip.
f By routing the SRU and TGU tail gas to the incinerator via a common header, a vacuum
can be maintained at the RGG without risk of leaking air from the incinerator back into the
TGU, thus potentially further increasing SRU capacity. In the event that the tail gas
bypass valve leaks, clean TGU tail gas is recycled to the RGG rather than SRU tail gas
bypassing the TGU (as when the RGG pressure is positive). Any such reverse flow will
improve bypass valve reliability by excluding sulphur vapor, and the valve can be partially
stroked periodically to verify operability without increased emissions.
Figure 11– WorleyParsons RGG Vacuum Operation
In the absence of a booster blower, a single startup blower recycle is usually provided for tail gas
recycle. While these machines tend to be less sophisticated, N2-purged tandem shaft seals are
still required.
- 20 -
The overall configuration of using the booster blower is shown in the Figure 10. This configuration
could be used with low temperature catalyst and indirect reheater instead of the RGG.
HYDROGENATION
REACTOR
RGG
SRU TAIL GAS
REACTOR
EFFLUENT
COOLER
TO
INCINERATOR
HY-250
PC
HC
HY-251
CONTACT
CONDENSER
ABSORBER
XY-292
FC
BOOSTER
BLOWER
DESUPERHEATER
WATER
WASH
Figure 12 – WorleyParsons BSR-TGU with Booster Blower configuration
- 21 -
Solvent Selection Criteria in the Tail Gas Unit
The most common solvent is 40-45 %-wt MDEA, (HS-101, or similar) designed for a maximum
rich loading of 0.1 mol acid gas (H2S + CO2) per mol amine with typical emission reduction to ~
100 ppmv H2S. Cooling of the lean amine to at least 100°F (38°C) is important for minimization of
emissions and amine circulation rate. Specialty TGU amines are essentially pH-modified MDEA
to facilitate stripping to lower residual acid gases for treatment to < 10 ppm H2S, potentially
obviating incineration. CO2 slip is also improved. These products are variously marketed as
f Dow UCARSOL HS-103
f Ineos Gas/Spec TG-10
f Huntsman MS-300.
An alternative to MDEA is ExxonMobil’s Flexsorb SE, a proprietary hindered amine patented by
Exxon in partnership with the Ralph M. Parsons Company. The main advantage is a 20-30%
reduction in circulation rate. The solvent is much more stable than MDEA, but is also more
expensive. Flexsorb SE Plus is also available for treatment to < 10 ppmv H2S. Both solvents
require a license agreement with ExxonMobil.
It used to be assumed that TGU carbon filtration was not required in view of the absence of
hydrocarbons. For MDEA-based solvents, at least, this has proven untrue, presumably due to
the generation of surfactant amine degradation products.
f Solvent Applications
f FLEXSORB® SE Selective removal of H2S
f FLEXSORB® SE Plus Selective removal of H2S to less than 10 ppm
f FLEXSORB® SE Hybrid Removal of H2S, CO2, and sulphur compounds (mercaptans and
COS)
f In sulphur plant tail gas applications, FLEXSORB® SE solvents can use as little as one
half of the circulation rate and regeneration energy typically required by MDEA based
solvents. CO2
f Rejection in TGTU applications is very high, typically >90% rejection.
f Flexsorb solvents offer other advantages compare to the other amine solvents for
instance, most of applications requires no reclaiming, have good operating experience, low
corrosion, and low foaming due to low hydrocarbon absorption, by providing water wash of
treated gas at low pressure system amine losses are minimum.
- 22 -
DEGSulf Sub Dew Point process by WorleyParsons& DEG-ITS
DEGSulf-SDP is a sulphur recovery process of the Claus type. A plant consists typically of a
Claus furnace plus downstream just 2 catalytic reactors and sulphur condensers. The reactors
contain a heat exchanger which keeps the operating temperature for each reactor at its optimum.
This simple system, described in detail below, allows reaching up to 99.85% sulphur recovery
rate.
Gas containing hydrogen sulfide (=H2S) is sent to the Claus furnace. There it is burned with a
stoichiometric deficiency of air so that one third of the H2S is converted to SO2. The residual H2S
and the SO2 react to elemental sulphur according to the Claus reaction (I): (I) 2 H2S + SO2 3/x
Sx + 2 H2O x = 2,4,6,8 indicates the different sulphur modifications Typically a recovery rate of
over 60 % is realized in the furnace. Gas from the waste heat boiler and sulphur condenser of the
Claus furnace is reheated by a hot gas bypass. It then flows via 4-way valve to the adiabatic part
of the first reactor, which is filled with a catalyst of high COS and CS2 conversion capability.
Residual traces of free oxygen from the Claus furnace are eliminated in this layer. The gas enters
the cooled section of the reactor bed at a temperature of between 300 and 350°C. Cooling takes
place by evaporating boiler feed water or hot oil. Here the Claus reaction continues further close
to the equilibrium at appr. 260 °C, which is slightly above the sulphur dew point at outlet
conditions. The gas leaves the reactor and passes via the second 4-way valve to the only
sulphur condenser of the catalytic part. The sulphur condenser operates at gas outlet
temperatures of between 135 °C and 150 °C and produces low pressure steam. The process gas
leaves the condenser through a mist eliminator. Total sulphur recovery up to this point exceeds
95 %. The gas is reheated again before entering the second reactor which can be regarded as
the tail gas treatment. In the steam jacketed pipe the temperature is raised by appr. 20°C in order
to be safely above the sulphur dew point. In the adiabatic zone of the second reactor the Claus
reaction proceeds. Claus gas then enters the cooled part of the second reactor, where the
reaction temperature is lowered to 100 - 125 °C by the cooling coils. Elemental sulphur from the
adiabatic zone and formed in the cooled zone is adsorbed by the aluminum-based catalyst. The
coils keep the reactor outlet temperature at constant level throughout the complete adsorption
period. The evenly low temperature throughout the bed causes a substantial increase of the
sulphur recovery rate compared to state-of-the-art processes.
- 23 -
E-201
WASTE HEAT
BOILER
F-201
REACTION
FURNACE
PC
E-202
SULFUR
CONDENSER
E-203
NO. 1
REHEATER
R-201
NO. 1
REACTOR
V-203
NO. 1
REACTOR
STEAM DRUM
E-205
NO. 1
REACTOR STEAM
CONDENSER
E-204
NO. 2
REHEATER
R-202
NO. 2
REACTOR
V-204
NO. 2
REACOR
STEAM DRUM
E-206
NO. 2
REACTOR STEAM
CONDENSER
PA-201
DEGASSING
PACKAGE
T-201
SULFUR PIT
ED-201
SULFUR PIT
VENT
EDUCTOR
FC
ACID GAS FROM
KO DRUM
-
X
FC
TC
TC
X
E-205
MP
STEAM
E-206
MP
STEAM
M
+
X
PC
MP
STEAM
F-201
PC
TAIL GAS TO
INCINERAOR
E-204
E-203
X
PC
R-201
R-202
LI
E-201
LI
V-203
T
CONDENSATE
V-204
FC
BFW
LC
NH3 GAS FROM
KO DRUM
T
BFW
CONDENSATE
TC
REACTOR
SWITCHING
CONTROLS
TC
AC
M
H2S/SO2
E-202
BLOWDOWN
VENT
FC
BFW
FC
VENT GAS TO
INCINERATOR
B-201
FC
ED-201
MP STEAM
PA-201
SULFUR
M
COMBUSTION AIR
FROM SPARE
BLOWER
M
T-201
P-203A/B
B-201
COMBUSTION AIR
BLOWER
P-203A/B
SULFUR
DEGASSING PUMP
PROCESS FLOW DIAGRAM - SULFUR RECOVERY UNIT – TRAIN 2
Figure 13, WorleyParsons /DEG-ITS Sub Dew Point process (DEGSULF)
- 24 -
P-204A/B
P-204A/B
SULFUR TRANSFER
PUMP
AIR
Ammonia Destruction in a TGU (RACTM)
The general industry consensus is that the amount of ammonia that can be conventionally processed in
the SRU is limited to 30-35 %-vol on a wet basis. With what appears to be a trend toward higher-nitrogen
crudes, refiners are increasingly faced with the need for alternative processing schemes, as well as SRU
debottlenecking. With sour water stripping schemes such as Chevron’s Waste Water Treatment (WWT)
process for separating H2S and NH3, producing a pure marketable NH3 product is relatively difficult
compared with bulk separation of NH3 containing minor H2S.
WorleyParsons’ Rameshni Ammonia Conversion (RACTM) process, for which a patent is pending, substoichiometrically combusts a high-NH3 H2S-contaminated stream in the RGG. (Figure 14) Typically, the
NH3-gas heat release will exceed that required to reheat the Claus tail gas, thus necessitating a waste
heat boiler prior to the TGU reactor. A supplemental natural gas fire ensures process stability in the
event of NH3-gas curtailment. Sub-stoichiometric combustion of the NH3-gas generates supplemental H2
for the hydrogenation reactor and minimizes NOx. Most of any NOx that is made is reduced in the
reactor. Minor unconverted NH3 is automatically recycled to the sour water stripper via the Contact
Condenser blowdown.
Table 2 defines the nominal feed bases for two hypothetical cases, where Case 1 involves a pure NH3
stream, and Case 2 a high-NH3 low-H2S stream. Table 2 compares the nominal impact on key
parameters of routing those NH3 streams to the TGU (Cases 1b and 2b) as opposed to the SRU reaction
furnace (Case 1a and 2a).
- 25 -
Figure 14 – WorleyParsons Ammonia Destruction in TGU (RAC TM)
- 26 -
Table 2 – WorleyParsons RACTM Hypothetical Feed Streams
Fresh Feed Gas, Mol %
Component
Case 1
Acid Gas
Case 2
NH3 Gas
Acid Gas
H2S
80
80
CO2
16
16
NH3 Gas
5
96
NH3
H2O
Total
65
4
4
4
30
100
100
100
100
95
5
100
Fresh feed, LTPD S
29
NH3 / total fresh feed, mol %
28
Table 3 – WorleyParsons RACTM Impact on Key Parameters
Comparison
Case 1 NH3 Gas Route
Key Parameter
Case 2 NH3 Gas Route
Case 1A
SRU
Case 1B
TGU
∆
%
Case 2A
SRU
Case 2B
TGU
∆
%
Claus tail gas, MSCFH
689
364
-47
760
346
-54
Claus recovery, %
92.7
96.5
92.3
96.5
RGG fuel, MMBTU/hr
10.4
0.5
-95
11.4
0.5
-96
TGU amine AG, MSCFH
17.5
10.1
-42
18.1
15.3
-15
BSR Selectox
Selectox catalyst is a proprietary catalyst patented by WorleyParsons for low-temperature H2S-oxidation
and Claus-reaction catalyst development by the Ralph M. Parsons Company and Unocal. Reduced tail
gas from the BSR Contact Condenser is steam-reheated to about 400°F (~ 200°C) and combined with a
stoichiometric quantity of air in the reactor to produce elemental sulphur, which is subsequently
condensed. (Figure 15) Overall recoveries of 98.5-99.5% are achievable. The reactor inlet is limited to 5
%-vol H2S, above which recycle dilution (or inter-bed heat removal) is necessary to limit the exothermic.
- 27 -
SRU TAIL GAS
RECYCLE
WATER
NATURAL GAS
COMBUSTION AIR
CONTACT
CONDENSER
RGG
SOUR WATER
BLOWDOWN
HYDROGENATION
REACTOR
DESUPERHEATER
STEAM
REHEATER
REDUCED TAIL GAS
10% NaOH
AIR
SELECTOX
REACTOR
LP STEAM
SULFUR
CONDENSER
TAIL GAS TO
INCINERATOR
SULFUR
Figure 15 – WorleyParsons BSR Selectox
- 28 -
WorleyParsons Current Case Histories
The following are the case histories of the different projects have been recently designed by
WorleyParsons.
Project Case 1
WorleyParsons designed a new tail gas unit for a US refinery to meet the emission requirements. The
following were the key elements of the project.
f The existing sulphur plant did not have adequate pressure to handle the tail gas pressure
f Total H2S of less than 100 ppm
f COS, CS2 hydrolysis
f SO2 concentration at reactor outlet
f Hydraulic and unit capacity
WorleyParsons evaluated this project and the final design was based according to the following criteria.
f Reducing gas generator (RGG) was selected to achieve high temperature in the hydrogenation
reactor for COS and CS2 hydrolysis without changing any catalyst in the existing SRU.
f Flexsorb solvent was selected because it requires less circulation rate compare to the other tail
gas amine solvent Therefore, the capital cost reduced.
f A booster blower is provided to boost the pressure in the tail gas unit downstream of the quench
section. The booster blower has dual function where it will be used as a start up blower to
eliminate large volume of H2S to the flare and recycle back to the unit and the booster blower will
boost the pressure in the unit.
f The final design is according to the Figure 12 that is provided in this paper.
Project Case 2
WorleyParsons has designed two new tail gas units one for a US refinery and one for a Canadian refinery
with the following configuration.
f The existing sulphur plant did not have adequate pressure to handle the tail gas pressure
f Total H2S of less than 100 ppm
f COS, CS2 hydrolysis
- 29 -
f Hydraulic and unit capacity
WorleyParsons evaluated this project and the final design was based according to the following criteria.
f Reducing gas generator (RGG) was selected to achieve high temperature in the hydrogenation
reactor for COS and CS2 hydrolysis without changing any catalyst in the existing SRU.
f MDEA solvent was selected simply they do not need to deal with two different solvent for the
amine and tail gas unit and there was no cost saving to use other solvent.
f A booster blower is provided to boost the pressure in the tail gas unit downstream of the quench
section. The booster blower has dual function where it will be used as a start up blower to
eliminate large volume of H2S to the flare and recycle back to the unit and the booster blower will
boost the pressure in the unit.
f The final design is according to the Figure 2 that is provided in this paper.
Project Case 3
WorleyParsons has designed a new sulphur recovery and BSR/MDEA tail gas unit for a refinery in South
America with the following configuration.
f Total H2S of less than 100 ppm
f COS, CS2 hydrolysis
f Hydraulic and unit capacity
WorleyParsons evaluated this project and the final design was based according to the following criteria.
f Low temperature catalyst is selected in the tail gas unit and the first reactor bed in the Claus unit
will contain some Ti catalyst
f MDEA solvent was selected simply they do not need to deal with two different solvent for the
amine and tail gas unit and there was no cost saving to use other solvent.
f A n start up blower is provide only for start up purposes and will not be used at normal operation
except where is the very low turn down and may be used to boost the pressure.
f The final design is according to the Figure 6 that is provided in this paper.
- 30 -
References
1. Ammonia Destruction in a Claus Tail Gas Treating Unit, by M. Rameshni, presented at British
Sulphur Conference, Canada, 2007
2. Operating experience of a 2-reactor Claus plant for up to 99.85% sulphur recovery, by J. Kunkel,
P.M. Heisel, LINDE AG, Ulf Nilsson, Peter Eriksson, NYNÄS AB
- 31 -