PERRY GAS PROCESSORS, INC. TECHNICAL BULLETIN NO. 76-002
APRIL 1, 1977)<. ODESSA, TEXAS
REVISED FEBRUARY 15, 1982
FUNDAMENTALS OF GAS TREATING
MANUAL FOR DESIGN OF GAS TREATERS
Robert S. Purgason
Perry Gas Processor~ Inc. Odessa, Texas
Presented To
The Gas Conditioning Conference, , 1976, 1978, 1979, 1980, 1981, 1982
Norman, Oklahoma
i Page i to 45
FUNDAMENTALS OF GAS TREATING
Introduction .••••.••.••...••.•
Section 1: Monoethanol/Diethanol Amine Process
A.
B.
c.
D.
Description of the Process .
Design Calculations 1. 2. 3. 4. 5.
Input Data. • ••• Circulation .•.••. Contactor ••••••••• Regeneration Vessels • • ••• Heat Loads • • • • . • •••• a. Re boiler . . . . . . . . . . . . . . . . b. Solution Exchangers .•••• c. Solution Coolers •••• d. Reflux Condenser •••••••••••.
Cost •••••..••• 1. Capital Investment • 2. Operating Expenses •
Graphs and Tables Table 1-1: Amine Unit Operating Costs Table 1-2: Regeneration Vessel Sizes • Figure 1-2: Typical Flow Sheet Figure 1-3: Contactor Capacities Figure 1-4: Cost Curve ••••
Sectoin 2: Iron Sponge Process
A. Description of Process
B. Design Calculations •• . ••••• 1. Linear Velocity ••• 2. Space Velocity •••••••••••• 3. Bed Life •••.• 4. Air Regeneration •••••
c. Cost •••••.••• 1. Capital Investment • 2. Operating Expenses .
D. Graphs and Tables
2
2
3 4 4 5 5 5 5 5 5 5
6 6 6
7 8 9
10 12
13
13
14 14 14 15 16
16 16 16
Table 2-1: Bed Capacity for Various Vessels 17 Figure 2-1: Typical Flow Sheet . • • . . • • . 18 Figure 2-2: Bed Life • • • • • • • • • • 19 Figure 2-3: Bed Capacities . • • . • • 20
ii
Section 3:
Section 4:
Section 5:
Section 6:
Section 7:
FUNDAMENTIALS OF GAS TREATING - Continued
Proprietary Liquid Processes . . . . . . . . . A.
B.
c.
D.
E.
F.
Diglycolamine (Econamine) (Description, Advantages and Applications) . Sulfinol (Description, Advantages and Applications) . Propylene Carbonate (Description, Advantages and Applications) . Selexol (Description, Advantages and Applications) . MDEA (Description, Advantages and Applications) • Typical Flow Sheets
. .
Figure 3-1: Propylene Carbonate •••••• • • Figure 3-2: Selexol ••••
Proprietary Solid Processes (Mol Sieves) . A. Description of Process • • . . .•• B. Advantages and Applications ••• C. Figure 4-1: Typical Flow Sheet •••••
Proprietary Processes with Direct Reduction to Sulfur •••••••••••• A. Townsend Process (Description) • • • B. Stretford Process (Description) C. Takahax Process (Description) •••••
Sulfur Recovery Processes ••••• A. Claus Process for Sulfur Recovery
1. Description of Process ••• 2. Capital Investment .••.•••••.•••
B. Recycle Selectox Sulfur Recovery Process 1. Description of Process . 2. Capital Investment •••
C. Graphs and Tables Table 6-1: Annual Operating Costs for
Claus Sulfur Recovery Plants with Incinerators •••••••••
Table 6-2: Annual Costs for Claus Sulfur Recovery Plants with Wellman-Lord
Figure 6-1: Typical Claus Plant for Sulfur Recovery ••.••••.
Figure 6-2: 3-Stage Claus plus Incinerator -Capital Cost ..•••••..•
Figure 6-3: 2-Stage Claus plus Wellman-Lord -Capital Cost ••••••••••••
Figure 6-4: Sulfur in Inlet Gas •••••• Figure 6-5: Typical Recycle Selectox Flow Sheet
References
iii
2 1
21
22
22
23
23 23 24 25
26 26 27 28
29 29 30 31
32 32 32 33 34 34 35
36
37
38
39
40 41 42
43
F U N D A M E N T A L S 0 F G A S T R E A T I N G
INTRODUCTION
In this paper it is our intention to review the fundamentals of gas
treating by liquid and solid processes. In Section l we describe the
Ethanolamine Process for treating gas to remove acid gas constituents.
In Section 2 Iron Sponge as a solid gas treating process is described.
In Section 3 we present a description of Proprietary liquid processes,
and in Section 4 Proprietary solid processes are discussed.
Section 5 presents proprietary processes that incorporate sulfur
recovery. A brief description of the Claus Recovery Process and the
Recycle Selectox Sulfur Recovery Process are presented in Section 6.
References and selected reading materials are listed in Section 7.
FACTORS TO CONSIDER IN SELECTING A TREATING PROCESS:
1. Treating Pressure (Inlet Gas Pressure)
2. Acid Gas Content (CO2+ H2S)
3. COS/CS2 Content
4. Treated Gas Specifications
5. Availability of Power
6. Environmental Constraints
K-1
SECTION 1
MONOETHANOL/DIETHANOL AMINE PROCESS
The Amine sweetening process has been widely accepted in removing co2 and
H2s from natural gas streams. Monoethanol Amine has been used for many
years* and lately Diethanol Amine has come into favor with the gas treating
industry. Since both processes use essentially identical equipment, we
will describe the process and then indicate differences in design calcu
lations between the two.
The reader is referred to Gas Purification, by Kohl and Riesenfield, plus
papers by Gene Goar, Charles Perry, Ward Rosen, and Allen Shell for more
information on the Amine Process.** Jefferson Chemical has published
brief design and operating considerations for amine treaters.
A. DESCRIPTION OF PROCESS
Referring to Figure 1-1, Natural Gas containing co2 and H2S is
contacted in countercurrent gas-liquid absorption process in a
trayed or packed tower to provide intimate contact for a chemical
reaction between H2S - CO2 (acid gas) and Amine. Good design
practice dictates a scrubber on the inlet gas to remove entrained
liquids including distillate and water from the gas before it
enters the contactor. Also, a scrubber on the outlet gas to
recover any Amine solution carried over from the contactor should
be provided.
*The Monoethanol Amine process was first commercially applied as the "Girbitol" process in 1930.
**A complete set of references is listed in Section 7.
K-2
Rich Amine from the bottom of the contactor is flashed at reduced
pressure to remove entrained gases, including part of the acid gas
and then heated in a rich-lean Amine exchanger. The solution is
then fed to the stripper tower where the solution is regenerated
and denuded of acid gas by steam stripping. Acid gases are con
centrated in the overhead accumulator and disposed of by burning
in a flare, reboiler, or other incineration device. In cases
where sulfur exceeds 2-5 tons per day, the acid gases may be
processed for sulfur recovery. Lean amine solution from the
bottom of the reboiler is exchanged with rich amine in the
solution exchangers, then pumped in multi-stage pumps back to the
contactor to complete the process loop.
Inasmuch as a clean solution is a key to the success of a treating
system, good filtration is essential. Activated carbon filters
have been found to provide the best and most economical filtration
Also, MEA solutions may be reclaimed in a side stream reclaimer.
DEA solutions cannot be reclaimed due to the high boiling point of
DEA. (Our experience has indicated reclaiming is not necessary
for DEA solutions.)
B . DESIGN CALCULATIONS
1. Input Data
a. Gas Flow - Minimum, Normal, Maximum
b. Gas Pressure - Minimum, Normal, Maximum
c. Gas Temperature - Minimum, Normal, Maximum
d. Gas Composition - Sp Gr, H2s, CO2, cos, CS2, H2O
e. Specs for Treated Gas
f. Utilities Available
K-3
2. Circulation
Circulation of amine solution is determined by the acid gas
content of the inlet gas, strength of the solution, and the
type of amine to be used. In Figure 1-2, we have provided
quick determination of circulations required as a function of
gas volume , acid gas content, and type of amine.
Amine circulation can be calculated by the following formulas :
For MEA: GPM = 41. 0 X QX/Z
For DEA: GPM = 45.0 x QX/Z (conventional)
GPM = 32.0 x QX/Z (high-load)
For DGA: GPM = 77.0 x QX/Z
Where Q = Gas to be processed, MMscfd
X = Acid gas content, volume percent Grains H2S
= Mol % CO2+ 632
z = Amine concentration, wt. %
The above formulas are based on mol loadings of 0.33 Mol/Mol
for MEA, 0.5 Mol/Mol for conventional DEA, 0.7 Mol/Mol for
high-lo ad DEA, and 0.3 Mol/Mol for DGA.
3. Contactor
Pressure and gas volume are major influences on the contactor
sizing. Referring to Figure 1-3, the size of the contactor
may be determined as a function of pressure and gas flow.
K-4
4. Regeneration Vessels - Table 1-2
a. Still
b. Surge tank
c. Reflux Accumulator
d. Flash-Skimmer Tank
e. Carbon Filter
5. Heat Loads
Estimates of heat loads and surface areas to be required are
shown below as a function of circulation rates for amine
solutions:
a. Reboiler (Direct Fired)
Duty-Btu/Hr.
72,000 x GPM
45,000 x GPM
15,000 X GPM
Area-Sq. Ft.
-1 1 • 3 0 x GPM
11.25 x GPM
1 0. 20 x GPM
b. Solution Exchangers
c. Solution Coolers (Air Cooled)
d. Reflux Condenser (Air Cooled) 30,000 x GPM 5.20 x GPM
6. Power Required
Power requirement for the process is largely for pumping of
the amine solution. The table below gives approximate power
requirements as a function of amine solution circulation:
Solution Pumps
Booster Pumps
Reflux Pumps
Aerial Cooler
GPM x AP HP = 1713 (eff1c1ency of pump)
GPM x Psig x 0.00065
GPM x 0.06
GPM x 0.06
GPM x 0.36
K-5
= HP
= HP
= HP
= HP
C. COST
1. Capital Investment
Capital investment for plants to use either MEA or DEA pro
cess can be broken into two segments. First, the contactor
investment is a function of pressure, gas flow, and amine
circulation. The cost for the regenerator unit is a function
of amine circulation. Figure 1-4 presents investment data for
battery limit gas treating plants as a function of amine
circulation. It is assumed that the regenerator equipment
will be skid-mounted and assembled in a shop to minimize field
construction. The cost figures include transportation,
foundations, and piping required at the location as a turn key
installation.
2. Operating Expenses
Operating expenses for the amine unit will be largely
operating labor, power and fuel gas. Typical operating
expenses for a treating unit are presented Table 1-1.
K-6
TABLE 1-1
AMINE UNIT OPERATING COSTS
Basis: Treating plant with 40GPM circulation;treating 3MMcfd gas with 10% acid gas: January, 1981, cost index
¢/Mcf
Operating Labor (2@ 22,000) •.•••.•••.•..•.•.....•. 4.07
Supervision (1@ 35,000) .........•.......... . ..•... 3.24
Employee Benefits@ 35% Payroll .................... 2.59
Utilities
Chemicals & Supplies . ............................. .
2.78
1. 39
Repair Materials & Labor@ 4% Investment ........... 4.44
Direct Overhead@ 5% Investment .................... 5.57
Corporate Overhead@ 3.5% Investment ............... 3.89
Depreciation ( 10 Year Straight Line) ...•.......•.. 11.11
Interest (@ 20%, first year) ...•.....•..•......... 22.22
Insurance & Taxes@ 2% Investment .................. 2.22 63.52
Plant Investment Estimate $1,200,000
K-7
$/Year
44~000
35,000
28,000
30,000
15,000
48, 0,00
60,000
42,000
120,000
240,000
24,000 686,000
~ I
(X)
SG . Jtion (i r·cl Rate .
GPM
10
25
50
. JO
200
300
400
St ill Diameter
I
.
16
24
30
42
60
72
84
TABLE I - 2
REGENERATION VESSEL SIZES
{Inches)
Surge Tank Reflux Ac cum. Diam. I Length Diam. Length
- ,----
24 72 16 36
42 96 24 48
48 144 30 96
60 192 42 96
84 288 60 96
84 384 72 96
96 384 84 96
Elash Tank Carbon Filter Diam. Length Diam. Lenqth
24 72 16 84
42 96 24 84
48 144 36 96
60 192 48 96
84 288 60 96
84 384 72 96
96 384 84 96
~ I
I.O
I Contact or r l
~r~k Slac
' ~ ,if. '.
11 k '·" .... ,
.· .: ,
Aerial Cooler
~ I .....
0
FIGURE I ~2 1
AMINE CIRCULATION REQUIRED ' '
GPM CIRCULATION REQUIRED PER MMCFD 0 2 3 5 6 7 8 9 10
·- -..-- --- ,- - - - - .... _ ,..._ - .... f--
I I I I t-~ 1-,,
-1 .... -t _J ,..._-t I t- . I -t -t- - - :- : -I I : I- : ,..._ I : I : J : I I I
I
I I I I I I '1 I I ' ' ' . ' ' ' . ' ' ' . . ' ' ' . . ' '
- ·-,--,- .-- -,- r- - - ·- ..-- ~-,--, ,-
- - """7 -T ' I 7 J T- ! r- - I 1- -, I I I- -, I - I I I
I I .1 _L I I I , I 1 I I I l 1 I I _J
I~'" I I I I I I I I I ""llil ~
I I I I I I I I ~-·-i-- - r- ~-,- -~
- --r-: " ~ s:: --- t% AJ 'H f'... :'---- r-- r-- IF
~I' i" C-.:::r----.. r-.r--...... r--r--r-- --
~ "n. "' . -- - --,- - !-- - -r,.___ r---: rr:::: ~'1il, I I". ~ I I I I I I I I I "'k_"' r-1 I -r-- .... ,o Nd r--
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'" - f---- I"-. -K r----. --.... r--... " I"--- "'
r--,.;: ,A A.I' .
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"'-i q ...____ ~ t/1 -
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" ~iy ~ I'. ......
I" .. "' ' ~ ~~u " C
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l_ i "'-.' . i "
r, , i i_
. i - ' i ' i
,, ~i I _j L- I_ I'-. _I .... -1 ,..._ J_ L' - I"\
.--I'-, - 'r--..
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__ ,_
!"'. ~ - - -
I I 1 J I I I I I I I J : I l I : I I- :
1, I'(: I I I I ! - I I I I I : -f- +- !-- I- I I !- -I --l I ,-+ I ~ -~ - - ' ' ' ' . I I I ' ' I
"' -,-- ,__ -e- ,-- --- ,___ .... ,-
"' - -
I
' : I
0 2 I
~ 4 5 6 · 7 8 9 10
I ' X ACiiO GAS I ,
]•
Jl
I I -/
V
- FIGURE 1-3 V CONTACT0R CAPACITY / - .,.
JO
I / I /
11!1 I / V
2.6 ;/ / '.r
2"4-I
' / 2l. I 7 I V
/
10 I / ~/ /
Q u. u ,a (/) I / ' / ;fa ~ ~
16 I " ,,,,.
I ~/ ' ~
>= t: u '-4, ~ 5
11
I V / ~ --·/ I / "?Jo' .~
I I V / /""
/ / 1/)
< (.!)
..J 10
< QC ::)
8 r-< z
6
... II / V V
/ /
/1 V V ~I-:-
/ ~ .,:;;,,.--
~
,_____.
'/ / ,
1...--------,,,,,,,..-
~ ~,
_,
-+ '/v ~ ~
1---
~ ~ 16,,
1 ~ ~ t.---~ __.,,, ~ ~
0 0 lOO •OO 600 800 1000 IZOO
OPERATING P~ESSURE
K-11
.,,., 200 300 400 ~00 600 700 800 900
200 - - ~~7~=t=- -- ----~ --- ------------~~-.--·----/v-/ 200
180 7 --·- ·- . 180
! ~ --- ---- --- - . --- --a---t--l---+--1---f--Y--1/ .J___J
--- - · - · - - - . - - .. ··- · - -· -· - ·--·- --;-----t----t----lf-----1-----J-/-,IL/--l------1--- 160
100 1000
,J --- -1- .------- --·--- - ·----- ·•··---i--t-t--1------1-l-7.?--t-----..1----,.__
~ 1401=~-~~~-:- ~ ==: :~---=~~---1----- --t--i----t---t-/-f-',7/-t..'---t---t--.L.----1140
~ 120
_ +- +-1 __ T I l~-- ---~------t--/---,'"l-v----l---f----¼---1-----1----1 120
~ 100 7-- --- ---r--~---1,--- -/~v~/--t-----t·--+--+---1--1---!---.--J 100
Ill 80 - ----:-v~-t---t---f---1-...1---L--1-...J.._--I 80
~ -;--~ --- ~ ~ 1--t--~-t-v--:r-t---t---f---..,I FIGURE 1-4
J ---, . V INSTALLED AMINE PLANT COSTS 60
--- ,--- --7·------- V ·, I APRIL 1977
4 0 -~-~+J-=fP~17-~ ----~-·===, ====-=--=--~---=·--1----1 40
20
- - --14-_--__ ----l -- -l--+-~-- i--;---J----.1.---l--- -- 1-- 1------~ 20
OL---1---L--L- 0 100 200 300 4 00 500 600 700 800 00 · Q
SECTION 2
IRON SPONGE PROCESS
The iron sponge process is one of the oldest and simplest processes for
treatng sour gas. However, economics normally limits its application to
gases containing less than 20 grs. H2S per 100 Scf. It may be used on
low pressure gas, but is more successful on high pressure gas. The sponge
must be moist to be reactive, and if the feed gas is dehydrated, it must be
resaturated with water before entering the sponge bed.
The iron sponge material is oak wood shavings impregnated with a hydrated
iron oxide. It is available as both the 9 lb. and 15 lb. grade (based on
iron content of 9 lbs/bushel and 15 lbs/bushel.) A bushel of sponge will
occupy one cubic foot of space when packed in a treating unit.
A. DESCRIPTION OF THE PROCESS
An iron sponge treating unit is an open vessel sized for the
proper bed volume and cross sectional area for the volume of gas
to be processed. A screen, or coarse packing is placed in the
bottom of the vessel to support the sponge. The remainder of the
vessel is filled with iron sponge, leaving only a small void space
at the top. Gas enters a top connection, passe downward through
the bed, and then out the bottom connection. Hydrogen sulfide
reacts with the iron oxide in the sponge to form iron sulfide,
thus sweetening the gas. Carbon dioxide does not react with the
sponge, and passes through the bed unchanged.
K-13
The iron sponge process is a batch process, thus requiring
multiple towers, on interruption of flow when the bed is changed
or regenerated.
B. DESIGN CALCULATIONS
An iron sponge treating vessel must be designed based both on
linear velocities through the bed and contact time. Once these
parameters are determined,
life will be satisfactory.
these steps.
1. Linear Velocity
it is necessary to determine if the bed
The following is a description of
The superficial linear velocity through the bed should not
exceed 10 ft. per minute based on the empty vessel cross
sectional area and the gas volume corrected to ACFM at flowing
temperature and pressure. Pressure drop through the bed
should be approximately 1 - 2 Psi per foot of bed depth.
2. Space Velocity (or Contact Time)
The sponge bed should be of sufficient size to allow a maximum
space velocity of 6.0 ACFH per cubic foot of bed for each
grain of H2s in the inlet gas.
Figure 2-3 may be used to estimate the capacity of sponge
vessels at various pressures.
K-14
Space Velocity (cont'd)
In a typical sponge treating unit, the gas channels to some extent
along the shell of the vessel. The H2S front moves through the
bed as an inverted cone, and the greater the linear and space
velocities, the steeper will be this cone. Thus, if too high
linear and space velocities are used, when the H2S breaks
through around the vessel shell, there will be a large mass of
unreacted sponge in the center.
3. Bed Life
Iron sponge theoretically can pickup 0.56 lbs. of sulfur per pound
of iron oxide. For design purposes, this must be aerated by about
25%. The following forumula can be used to determine bed life:
MMcf/Change Bu. Per Charge
= 4.69 Gr./100 of H2S
The above formula is for 15 lb. sponge. For 9 lb. sponge, the
constant 2.80 should be substituted for 4.69.
K-15
4. Air Regeneration
Iron Sponge may be regenerated with air, either batch wise or
continuously with some limitations . In batch regeneration, a
spent sponge bed is isolated and depressured. A blower is
used to circulate gas through the bed at atmospheric pressure .
Air is injected into the gas slowly over a 24 hour period,
allowing the oxygen content in the gas to slowly increase to 8% .
The bed life after each regeneration will be approximately 60%
of the previous bed life , resulting in only about 2 regenerations
being feasible. After a bed has been regenerated two or three
times, elemental sulfur formed on regeneration builds around
the sponge making it extreme l y difficult to remove the spent
sponge.
-~ C . COST
L Capital Investment
Capital costs of approximately $15 , 000.00 per Mmcfd capacity
( up to 20 MMcfd) to $10,000.00 per MMcfd capacity (above 30
MMcfd) may be used to estimate installed costs for 1000 Psi
units.
~ Operating Expenses
For operating costs, 0.12/Mcf for each 5 gr./100 Scf in the
inlet gas will cover the costs of sponge, labor and amortization
of cost of equipment .
K-16
Vessel Dia., Inches
24 II OD
30 11 ID
36 II ID
42 11 ID
48 11 ID
54 11 ID
60" ID
66 11 ID
72 II ID
84 11 ID
96 II ID
Table 2-1
BED CAPACITIES FOR COMMON SIZE SPONGE TREATERS
Bed Capacity, Bushels (or Cu. Ft) For Bed Depth Indicated
10' 12' 15' 18' 20'
30 34 44 52 58
50 60 74 88 98
70 84 104 118 140
96 116 144 174 192
126 150 190 226 252
160 190 240 286 320
196 236 294 354 392
238 286 356 428 474
284 340 424 410 568
384 462 578 692 768
500 600 752 902 1004
K-17
E INLET )I----
LDA\DIN~ "PLATf"ORrw1
SPONGE' ll?C'ATE:R VESSELS
< GAS OUTLET 11,--""'--:!-+~---=----------......c:"-----'--4
TOC
s'
FIGURE" .:l.-1
FLOW DIAGRAIYI FOR IJc>ON SPONGE T ~EATE K' PLJ>...N1 .1
F1GURE 2·1 CAPACJTY OF IRON OXIDE FOR SUUH GAS
BETWEc.N RECHARGES
BASED ON : I 5 LB. F E.i O, I BUSHEL 0.56LB. S PICKUP/ LB. FE4 o~ 75 % BED SATURATION
IRON SPONGE - BUSHELS PER. CHARGE
FORMULA FOR GRAPH: MMC F/ CHARGE= 4 69 _s_u_. ___ _
. GR.~l~cP SCF CRP
•
FIGURE 1-3 LINEAR ANO SPACE VELOCITtES FOR IRON SPONGE TREATERS
BASED ON : MAXIMUM LINEAR VELOCITY : IOFT./MIN. MAXIMUM SPACE VELOCITY: 60 CU. FT.
PER 'HOUR PER CU. FT. OF BED FO IO GR./ IO O H.2 S CONT ENT
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TREATER OPERATING PRESSURE- PSI G
K-20
I
SECTION 3
PROPRIETARY LIQUID PROCESSES
A. DIGLYCOLAMINE (Econamine)
The Diglycolamine (DGA) process is carried out in equipment very
stmilar to the equipment for Ethanolamine. The major difference being
in the concentration of the Diglycolamine which is normally in the fifty
to sixty percent concentration range. The DGA process does require a
specialized purification unit on the side stream of the circulation.
Major advantages for the DGA process over the MEA process is a higher
allowable concentration of DGA that is up to 60% with resulting higher
acid gas loadings which leads to substantial savings in equipment size
and cost, and utility cost. Optimum applicability is in the range of
acid gas concentration from 1.5% to 8.0%.
K-21
B. SULFINOL (Shell)
The Sulfinol process is unique in that it combines characteristics of a
solvent process and an amine process. The high solvency of sulfinol for
acid gases at high acid gas concentrations results in a greatly reduced
circulation rate which is the basis of the economic advantages. The
advantage disappears at low acid gas concentration. The sulfinol
process has limitations which must also be considered. It has a great
affinity for aromatics, therefore pre-treatment facilities must be
installed ahead of the acid gas if it is being fed to a sulfur recovery
plant. Initial cost for the sulfinol process is quite high due to high
sulfinol solution cost and royalties paid to Shell Development Company.
The process has numerous advantages, especially its ability to
selectively remove H2S leaving up to 50% of CO2 in the gas.
C. PROPYLENE CARBONATE
The Propylene Carbonate process has advantages over the Amine processes
in that it has high acid gas loading and corresponding decrease in
capital cost. Regeneration of the solvent is by pressure reduction
requiring no steam regeneration. Disadvantages for the Propylene
Carbonate process are similar to those of the Sulfinol process in that
the solvent will retain heavy hydrocarbons so that acid gas must be
treated to remove these hydrocarbons before being processed for sulfur
recovery.
K-22
D. SELEXOL
Selexol process is a solvent process developed by Allied. Major
advantages of the Selexol process are high solubility of H2S relative
to CO2 making it possible to remove all H2S, leaving a certain
amount of CO2 in the residue gas. Selexol removes carbonyl sulfide,
carbon disulfide, and mercaptan without degradation. The solvent will
dehydrate natural gas if water is removed in the regeneration step.
Disadvantages of the Selexol process is typical of other solvent
processes, i.e., absorption of heavy hydrocarbons, expensive solvents,
and royalty charges.
E. MDEA
The ethanolamine amine process utilizing n-methyldiethanol amine "MDEA"
is carried out in equipment very similiar to the equipment used for
primary and secondary ethanolamines. The major difference being the
concentration of the MDEA, the number of trays in the stripper and
absorber, and the heat load for the process.
Major advantages of the MDEA process over MEA, or DEA, is the capacity
of higher selectivity with respect to H2S in the presence of CO2.
This can result in substantial savings in equipment size and cost, and
utility cost depending on the desired and achievable selectivity.
K-23
ABSORBE:R
. -TREATED E:A5
RECYCLE GAS
LEAN ~OLVENT
ACID GAS FLASH DRUMS
FIGURE:' 3-1 PROPYLENE CARBONATE PROCESS
~ I
l'v U1
ABSOPBE~ HIGH PR[SS.UR[
f"LA5M
r-<-- ____ R_ E:_CY __ c ___ L_E: _____ -1
INTE RME'DIAT£
PQf!,~UJ?E FLA~H
LOW
PRE.S.SUl.'E f"LASH
V~NT
STRIPPER
VENT
AIROR i--:s---0-lNERT &A :
F"IGURE 3 -1 StLEXOL PROCESS
SECTION 4
PROPRIETARY SOLID PROCESS
(Molecular Sieve)
A. DESCRIPTION OF THE PROCESS
In the molecular sieve process, the sour gas passes through one of two,
three, or four fixed beds of molecular sieves where the hydrogen
sulfide, along with water and organic sulfur compounds, are removed from
the gas by a process similar to adsorption. Whenever the bed becomes
saturated with hydrogen sulfide, the main gas flow is switched to
. another bed which is freshly regenerated. This cycling of beds
continues on a regular cycle, based either on time, or on breakthrough
of hydrogen sulfide.
A side stream of the sweet gas (usually approximately 20%) is heated to
approximately 600°F. to 700°F. and passed through the last fouled bed to
regenerate it. The hot regeneration gas is then cooled and processed by
a wet process (usually amine or sulfinol) to remove the hydrogen sulfide
from the regeneration gas. After the regeneration gas is sweetened, it
rejoins the main gas stream downstream of the sieve beds.
K-26
I
B. ADVANTAGES AND APPLICATIONS
The molecular sieve process may be used to selectively remove hydrogen
sulfide in the presence of carbon dioxide. Thus, if it is not necessary
to remove carbon dioxide from a gas stream, the molecular sieve process
may offer a lower capital cost, lower fuel requirements, and/or lower
operating expenses. Molecular sieves will also remove organic sulfur
compounds.
Offsetting these advantages, the necessity of a separate regeneration
gas tr~ating unit makes this process somewhat complicated and less
applicable to small streams. While the sieves can selectively remove
hydrogen sulfide, the presence of carbon dioxide does adversely affect
the capacity of the sieves for H2S. Also, sieves are subject to
attack by contaminants occasionally found in gas, and care muit be taken
to see that the incoming gas is free of contaminants.
K-27
sou~ GA.S •).-.--------------iA..l'-------------,-----------e-'NLCT ~£.4Tt:D GAS It-: __ _,
:.
1-------------------i--------------------------~-- -------ADSORBING COOLING- DESORBING- AMtNE SCRu!!>e.t:J~
F"I GU
G-LrtoL scR"uBoFJ
SECTION 5
PROPRIETARY PROCESSES WITH DIRECT REDUCTION TO SULFUR
A. TOWNSEND PROCESS
In the Townsend Process, an SO2 rich organic solvent, such as
triethylene glycol, contacts the sour gas, and the hydrogen sulfide
reacts with the SO2 to form elemental sulfur and water. The glycol
solvent also absorbs additional water vapor from the gas. The solvent,
containing elemental sulfur and water, is heated to melt the sulfur and
distill out the water. The liquid sulfur is decanted, and a portion is
burned to form SO2, which is absorbed in the lean solvent, and
recyled to the contactor.
At present, various types of catalysts may be used to catalyze the
H2S - SO2 reaction, and other solvents are being investigated.
B. STRETFORD PROCESS
The Stretford Process is an absorption process, in which the H2s rich
solution is air blown to reduce the H2S to finely divided elemental
sulfur particles, which are removed by centrifuge or filtration. The
air blowing and reduction of the H2S regenerates the solvent for
recirculation.
The solvent used in the Stretford Process is an aqueous solution of
sodium carbonate and anthraquinone disulfonic acid with a sodium
metavanadate activator.
The chemicals in the solvent are all quite stable, but the presence of
air and sulfur does cause the formation of a small amount of
thiosulfate in the form of a sodium salt . The process produces good
purity sulfur ) and can be used for low H2S: CO2 ratio streams.
K-30
C. TAKAHAX PROCESS
The Takahax Process is similar to the Stretford Process, in that a
special solvent removes the H2S from the gas, and the rich solution is
air blown to reduce the H2S to elemental sulfur and regenerate the
solvent.
In the Takahax Process, the solvent is an alkaline solution of
1,4-naphtaquinone, 2-sulfonate sodium, and a catalyst. Sulfur formed in
the solution during regeneration is removed by filtration.
K-31
SECTION 6
SULFUR RECOVERY PROCESS
A. CLAUS PROCESS FOR SULFUR RECOVERY
The most widely accepted process for recovering sulfur from acid
gas streams is the modified Claus Process. The basic Claus
reaction was disclosed in 1893:
For convenience in analyzing the steps in the plant operations,
this reaction can be broken down into two steps:
H2S + 3/202
2H2S = SO2
--->~ H2O + SO2
--->~ 2H2O + 3S
The reader is referred to Gas Purification by Kohl and Riesenfeld
plus papers by Gene Goar. A series of articles by Dr. Richard K.
Kerr, Director of Applied Research and Development, Western
Research and Development of Calgary, Alberta, has appeared in
Energy Processing/ Canada, which describes the Claus Process
Thermodynamics and Kinetics.
1. DESCRIPTION OF PROCESS
Referring to Figure 6-1, acid gas from the amine treater
containing H2S and CO2 is pre-heated and injected into the
reaction furnace along with sufficient air to burn 1/3 of the
H2S. Sufficient resident t i me is provided in the reaction
furnace to permit a part of the reaction between H2S and
SO2 to proceed. Indeed, approximately 65% of the recovered
sulfur is formed in the reaction furnace. The reactants from
the furnace are cooled in a waste-heat recovery boiler to
approximately 1100uF.
K-32
The partially cooled reactants are split with a portion of the
1100°F. stream being cooled to 275°F. where molten sulfur is con
densed. Unreacted H2S and SO2 join a portion of the hot gas
by-pass material for re-heat to approximately 600°F. for entry into
the first catalytic reactor where a bauxite (alumina) catalyst is
used to improve the reaction kinetics. The effluent from the No. l
reactor is cooled in No. 2 condenser to drop out liquid sulfur, with
unreacted H2S and SO2 passing on in similar manner to reactors
No. 2 and No. 3. Tailgas from the No. 4 condenser is normally sent
to an incinerator where the last traces of H2S are burned to SO2
and discharged in a tall stack.
Where very stringent environmental restrictions dictate a 99+%
recovery of sulfur, the tailgas will be further processed in a
"tailgas clean-up unit" such as Shell's SCOT Process or other other
tailgas clean-up processes.
2. CAPITAL INVESTMENT
Capital investment for Claus Sulfur Plants is somewhat dependent on
the percent of H2S in the acid gas from the sulfur plant. Figure
6-2 presents capital costs for a three-stage Claus unit plus
incinerator presented in a recent publication by the EPA.*
* "An Investigation of the Best Systems of Emission Reduction for Sulfur Compounds from Crude Oil and Natural Gas Field Processing Plants"; W.O. Herring and R.E. Jenkins, Research Triangle Park, North Carolina.
K-33
Figure 6-3 presents investment costs for a two-stage Claus unit plus
tailgas clean-up unit employing the Wellman-Lord process. Table 6-1
presents annual operating costs for Claus Sulfur Plants with
incinerators. Table 6-2 presents operating costs for Claus Sulfur
Recovery Unit with Wellman-Lord tailgas clean-up unit. These
operating costs are also taken from the EPA publication.
B. RECYCLE SELECTOX SULFUR RECOVERY PROCESS
A new technology recovering sulfur from gas streams has been developed
based on a new remarkable active Selectox catalyst which selectively
oxidizes H2S to sulfur using air at low temperatures without forming
SO2 or oxidizing either hydrogen or light hydrocarbons. The process
developed jointly by the Science and Technology Division of Union Oil
Company of California and The Ralph M. Parsons Company has application
in a wide range of H2S feed compositions, but especially shines in
concentrations from 1 to 40% H2S.
1. DESCRIPTION OF PROCESS
Referring to Figure 6-5 the acid gas is fed directly to the Selectox
catalyst after being mixed with a stoichemitric amount of air. Feed
temperature is about 350°F against 450°F for a standard Claus
catalyst. The reaction heats the gases which are cooled to condense
elemental sulfur. Part of the cool gas is recycled to the feed in
order to control the reaction temperature. The recycled gas dilutes
K-34
the H2S content to limit the temperature rise. Gas that is not
recycled continues to a second and third stage which may use
Selectox or less expensive conventional alumina catalyst. After the
third reactor the tail gas may be either thermally or catalytically
incinerated.
Because the Selectox process is entirely catalytic, no thermal
reaction furnace is used, instrumentation and process controls have
been greatly simplified resulting in simpler and more reliable
operation.
The Selectox catalyst is also in the BSR/Selectox I tail gas cleanup
process should the environmental restrictions dictate higher than
96% recovery.
2. CAPITAL INVESTMENT
Capital investment for Recycle Selectox Sulfur Plants should be
competitive with an equivalent size three stage Claus Sulfur Plant.
K-35
~ I
w ·a-.
Tc.hle 6-1 . ANNUAL OPERATING COSTS FOR CLAUS SULFUR RECOVERY PLANTS WITH INCTI:E~TORSa /
Plant Size , Mq/d (1 td )
Component 5 10 100
~power per Shift 1/ 2 1/2 1/ 2
Direct Labor~ $6.41/hr.b/ $ 28,100 $ 28,100 $ 28,100
Supv. & Fringes@ 40\ D~ 11,200 11,200 11,200
Electric Power@ $0.030/Kwhc/ 1,800 3,500 35,300
Fuel@ $2l.20/lO3m3 {$0.60/MCF)d/ 2,100 4,100 41,300
Maintenance@ 3\ Inv. 22,700 27,200 83,500
Steam Credit @ $1 • l 0/Mg ( $0. 50/~'.LB) c/ (5,400) (10 , 800) (108,400)
Sulfur Credit@ $24 ~61 /Mq ($25/L T) f / (~1,500) (83,000) (830,000)
Net .Operating Cost 19,000 (19,700) (739,000)
~/ 350 operating days p e~ year, 90~ a2 s gas stream. h/ U.S. Department of ~r,"Monthly Labor Review",October, 1975, p. 96. C/ 1'~unt from Ford, Bacon, and Davis, "Sulfur Recovery Plants", 1971, page 80 1 price from "'Typical
Electric Bills, 1974, FPCR-83",Federal Power Commission. d/ Process Research, Inc., EPA Contract 68-02--0242, Task 2, 1973, page 75 (LTR. L.L. Zuber, Stauffer
Chemical. to W.D. Beers, dtd. 3-21-72) . eJ Ford, Bacon, & Davis, .QE_. Cit. ~/ Assumes 95% recovery with tail gas . treatment.
Table 6-2. ANNUAL COSTS FOR CLAUS SULFUR RECOVERY PIANTS WITH WELLMAN-LORD
Co nent
Hanpo-,.-er Per Shifta/
1J/ Direct Labor@ $6.41/hr.
Supv. & Fringes@ 40\ DL
Electric ~o~er@ $0.03/Kwhl:y
Fuel @ $21.20/103~ 3 {~0.(0/MCF)a/ 7 .1-'..aintenance @ 3\ INV(Claus)
~ 5\ INV (W-L)
Chemicals, @ $200/ton 50\ NaOHo/
Soft Water@ $1.14/m3 ($0.?.0/MGAtf/
steam Credit@ $1.10/Ma ($0.50/MLBS)a/
sulfur Credit@ $24.61/Mgd (~25/LT)f/ Net Operating Cost
.
5
2
$112,400
45,000
3,800
2,100
1,900
18,300
3,500
100
(4,800)
. {43 ~700)
$158,600
Claus Plant Size Mg/d (ltd) 10 100
2
$112,400
45,000
7,600
4,100
. 26,100
22,500
7,000
200
(9,600)
(87,500)
$ _:~27 ,BOO
2
$112,400
45,000
75,600
41,300
82,000
89,400
70,000
1,700
(96,.400)
(875,000)
$ (454,000)
a/ Trip Report: Standard Oil -Company of California El Segundo Refinery,10/15/73 1 C.B, Seaan, EPA.
b/ 350 operat~g days per ye~.
c/ "Hydrocarbon Processing .. ,_ . ~riL 1973, p. 116 • . d/ Scaling Exponent, 0.3.
e/ Scal.ing Exponent, 0.6.
f/ ~..ssume 100% sulfur recove::y ~-~- ~ail gas unit.
:,:: I ACID
w CX) GAS
• I '
COMBUSTION AIR BLIII/CP.
I I
nASiC H!£i ROILlR
HOT GAS !IY-P/ISS
1100 · r
COND . 11
s s s
N!"-1--
........ .. -·
TAIL GAS
s
_,64 .. • •··- . -......... --FIGURE 6-1
FLOW DIAGAAH
_.._,.._ ..... .......,ru~ ,..,..
TYPICAL CLAUS PLANT FOR SULFUR RECOVERY
-'° C> ~
>< ~ -V, J--
~ V, ,a
I ·U w '-0
__. q;: J--....... a..
·<X: u
40 .0
30.0
20.0
18:8 8.0 7.0 6.0 5.0
4.o·
3.0
2.0
1.0 o.~ 0. 0. 7 0.6 0.5 0.4
0.3
0.2
O. l
I I I I I I
A 15% H2S in acid gas
0 50% ,-i2s in acid gas ·-
(:) 90% H2S in acid gas
.__ _ _.___.'--,__,_~1,_I I II .1 . 3 .4 . 5 . 6 J B 91, 0
I 2
I I 11111( 3 4 5 6 7 8 9 10
\ SULFU.R RECOVERY PLANT CAPACITY:V.g/_d,(LT/d}
I I 30.0
20.0
10.0 9.0 8.0 7.0 6.0 5.0
4.0
3.0
2.0
LO 0.9 0.8 0.7 0.6 0.5
0.4
0.3
0.2 ·
·1 I I I I I o. 1 20 30 40 60 80 l 00
Figure 6~2 . Claus (3- stage) plus inci~erator - c~pital cost.
---\0 0 .-X ----~
I . V)
~ . .-
0 V)
0 . u _.J c::c l--0... c::c u
40.0.------T
30.0
20.0
10. 9 . 8. 7.f 6. 5.
4.cL 3.
2.
1. 0.
0.8 -0.7 0. 6 0.5
0.4 -·0 •. 3
0.2
-0~ l .1 .2
r1 I I 11 , -- ,-·11Tm 30.0
20.0
A 15% H2S in acid g?s
50% H2S in acid gas 10.0
0 9.0
90% H2S in acid gas 8.0
0 7.0 6.0 5.0
3.0
2.0
l.O 0.9 0.8 0.7 0.6 0.5
0.4
0.J .:.
0.2
I I J -I t I I • 3 • 4 • 5 • 6._7 -. 89 l ,. 0 2
I .3 4 5 6 7 89 10
SULFUR RECOVERY PU\NT CAPACITY, Mg/d,(LT/d)
2-sta~ lus Wellman-lord - capital cost.
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: : · I . - - · I . · T , ~ - - . - ,- -·_·_-_..c.;.....:...!..i..l . .c.....:.• __ ,...1-~--- · -'-'__,_ ____ , _ _ __ ....._ ___ ....._ _ _ _ _,_ ____ =r....1.. _ __ _i...;_ _ _ _ .L ____ _ J..L..:__:__· __ ...!._.:...._~_...!._...:.....'.._~!....l
Figure 6-5
THREE STAGE SELECTOX PROCESS WITH RECYCLE
HTM " HTM"
AIR-----
HTM" SELECTOX CATALYST
SELECTOX OR
ALUMINA ------=-.IL-- CATALYST
----.::::: SELECTOX OR
ALUMINA __ CATALYST
---- TAIL GAS TO
ACID GAS ----t FEED 1 TO 40% t-12S
RECYCLE BLOWER
"HTM = HEAT TRANSFER MEDIUM APPROXIMATE SULFUR YIELD
HTM"
s 82%
I
s 12%
s 2%
IITM"
TREATING OR INCINERATOR
SECTION 7
REFERENCES
1. Beaven, D.K., et al, "Developments in Selectox Technology," paper presented at the XX National Convention of AIChE (joint Mexican & American AIChE Meeting), October 15-17, 1980, Acapulco, Mexico.
2. Beaven, D.K., et al, "High Recovery, Lower Emissions Promised for Claus-Plant Tail Gas," The Oil and Gas Journal, March 12, 1979, pp. 76-80.
3. Berry, R.I., "Treating Hydrogen Sulfide: When Claus Is Not Enough," Chemical Engineering, October 6, 1980, pp. 92-93.
4. Dingman, J.C., and Moore, T.F., "Compare DGA and MEA Sweetening Methods," Hydrocarbon Processing, Vol. 47, No. 7, July 1968.
5. Dingman, J.C., and Moore, T.F., "Gas Sweetening with Diglycolamine," Proceedings of the 1968 Gas Conditioning Conference, University of Oklahoma, Norman, Oklahoma.
6. Ducksworth, G.L., and Geddes, J.H., "Natural Gas desulfurized by the Iron Sponge Process," The Oil and Gas Journal, September 13, 1965.
7. Dunn, C.L., Freitas, E.R., Hill, E.S., and Sheeler, J.E.R., "Shell Reveals Commercial Data on Sulfinol Process,: The Oil and Gas Journal, March 29, 1965.
8. Fails, J.C., and Harris, W.D., "Practical Way to Sweeten Natural Gas," · The Oil and Gas Journal, July 11, 1960.
9. Goar, B.G., "Sulfinol Process Has Several Key Advantages," The Oil and Gas Journal, June 30, 1969.
10. Goar, B.G., "Today's Gas Treating Processes," Proceedings of the 1971 Gas Contitioning Conference, University of Oklahoma, Norman, Oklahoma.
11. Goar, B.G., "Today's Sulfur Recovery Processes," Hydrocarbon Processing, Vol. 47, No. 9, September, 1968.
1 2. Gustaf son, D. J. and Healy, M. J., "Removal of Hydrogen Sulfide from Natural Gas/Carbon Dioxide Mixtures with Molecular Sieves," Proceedings of the 1968 Gas Conditioning Conference, University of Oklahoma, Norman, Oklahoma.
13. Haas, R.H., Ingalls, M.N., Trinker, T.A., Goar, B. Gene, Purgason, R.S., "Packaged Selectox Units - A New Approach to Sulfur Recovery." Paper presented at the 60th Annual Gas Processors Association Convention, San Antonio, Texas, March 23 - 25, 1981.
K-43
14. Hegwer, A. M., arrl Harris, R. A., 11Selexol Solves High H2S/C02 Problens, 11
Hydrocarlxm Processing, Vol. 49, No. 4, April, 19 70 .
15. Holder, H. L., 11Diglycolamine - A Pranising New Acid-cas Rerrover," rn,,e Oil and Gas Journal, May 2, 1966.
16. Kool, A.. L., and Riesenfeld, F. C., Gas Purification, McGraw-Hill Book Conpmy, Inc. , New York City, New York, 1960.
17. Maddox, R. N., and Burns, M. D., "Designing a Hot Carbonate Process," The Oil and Gas Journal, Noverrber 13, 1967.
18. Maddox, R. N., and Burns, M. D., "Hot Carbonate - Another Possibility," 'Ihe Oil and Gas Journal, October 9, 196 7.
19 . Maddox, R. N., and Burns, M. D., "Iron-Oxide Process Design calculations, " rn,,e Oil and Gas Journal, August 12, 1968 .
20. Maddox, R. N. , and Burns, M. D., "MFA Process to be Considered First, 11
Th= Oil and Gas Journal, August 21, 196 7.
21. Maddox, R. N., and Burns, M. D., "Physical Solution is the Key to rn,,ese Treating Processes," 'Ihe Oil and Gas Journal, Janua:ry 8 , 1968 .
22. Mason, J. R. , and Griffith, T. E. , "A Case History - Conversion of a sweetening Solution fran MFA to OOA, 11 Proceedings of the 1969 Gas Corrlitioning Conference, University of Oklahoma, Noman, Oklahana .
23 • Nicklin, T., and Bnmner, E. "Hew Stretford Process is Working," Hydrocarbon Processing, Decenber, 1961.
24. Perry, C.R., "A New IDok at Iron Sponge Treatrrent of Sour Gas," Proceedings of the 1970 Gas Conditioning Conference, University of Oklahorra, No.rman, Oklahana.
25 . Perry, C. R., "Basic Design and Cost Data on MEA Treat ing Units, 11
Portable Treaters, Inc., OOessa, Texas.
26. Per:ry, C. R., "Design and Operating Arnire Units for Trouble Free Unattended Operations," Proceedings of the 1969 Gas Conditioning Conference, University of Oklahoma, Nonrian, Cklahana.
27. Rcsen, W. D. , "Here's a Quick Design Method for Amine sweetening Plants," 'Ihe Oil and Gas Journal, March 18, 1968.
28. Rushton, W. w., and Hays, w., "Selective Adsorption to Renove HiS," The Oil and Gas Journal, Septerrber 18, 1961.
29 • Scheinrian, W. L., "Selection of Total Sulfur Recovery Systens Errploying the Stretford Process," Proceedings of the 1972 Gas Conditioning Conference., University of Oklahorra, No.rman, Oklahana.
30. Shell, A. D., "Consider Basic Prd:>lans in Irrproving MFA Filtration," 'Ihe Oil and Gas Journal, February 12, 1968.
K-44
31. 8wai.m, C. D., "Gas Sweetening Processes of the 1960's," Hydrocarbon Processing, Vol. 49, N:>. 3, March, 19 70 •
32. 8wai.m, C. D., "Takahax Desulfurization Process, 11 Proceedings of the 1972 Gas Conditioning Conference, University of Oklahana, Norman, Oklahana.
33 • 'Iaylor, D. K., "HCM to Desulfurize Natural Gas, 11 A series in '!he Oil and Gas Journal, N:>verrber 5 and N:>venber 19, Decenber 3 and Decerrber 10, 19 56 •
34. 'Ihonas, T. L., and Clark, E. L., "Molecular Sieves Reduce Econcrni.c Operational Process Problans, 11 'Ihe Oil and Gas Jow:nal, August 12, 1968
35. Tavnsend, F. M., arrl Reid, L. s., "N:!west Sulfur Recovery Process, 11
Petroleum Refiner, Novanber, 1958.
36. zapffe, F., "Gas sweetening," Proceedings of the 1965 Gas Conditioning Conference, University of Oklahorra, N:>nnan, Oklahana .
37. zapffe, F., "Iron Sponge Renoves M:rcaptans, 11 'llle Oil and Gas Journal, August 19, 1963 •
K-45
