PERRY GAS PROCESSORS, INC. TECHNICAL BULLETIN NO. 76-002
ODESSA, TEXAS APRIL 1, 1977)<.
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 2
A. Description of the Process . 2
B. Design Calculations 3 1. Input Data. • ••• 4 2. Circulation .•.••. 4 3. Contactor ••••••••• 5 4. Regeneration Vessels • • ••• 5 5. Heat Loads • • • • . • •••• 5 a. Re boiler ...... 5 b. Solution Exchangers .•••• 5 c. Solution Coolers •••• 5 d. Reflux Condenser •••••••••••. 5
c. Cost •••••..••• 6 1. Capital Investment • 6 2. Operating Expenses • 6
D. Graphs and Tables Table 1-1: Amine Unit Operating Costs 7 Table 1-2: Regeneration Vessel Sizes • 8 Figure 1-2: Typical Flow Sheet 9 Figure 1-3: Contactor Capacities 10 Figure 1-4: Cost Curve •••• 12
Sectoin 2: Iron Sponge Process 13
A. Description of Process 13
B. Design Calculations •• . ••••• 14 1. Linear Velocity ••• 14 2. Space Velocity •••••••••••• 14 3. Bed Life •••.• 15 4. Air Regeneration ••••• 16
c. Cost •••••.••• 16 1. Capital Investment • 16 2. Operating Expenses . 16
D. Graphs and Tables 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 FUNDAMENTIALS OF GAS TREATING - Continued
Section 3: Proprietary Liquid Processes ...... 2 1 A. Diglycolamine (Econamine) (Description, Advantages and Applications) . 21 B. Sulfinol (Description, Advantages and Applications) . 22 c. Propylene Carbonate (Description, Advantages and Applications) . 22 D. Selexol (Description, Advantages and Applications) . . . 23 E. MDEA (Description, Advantages and Applications) • 23 F. Typical Flow Sheets 23 Figure 3-1: Propylene Carbonate •••••• • • 24 Figure 3-2: Selexol •••• 25 Section 4: Proprietary Solid Processes (Mol Sieves) . 26 A. Description of Process • • . . .•• 26 B. Advantages and Applications ••• 27 C. Figure 4-1: Typical Flow Sheet ••••• 28 Section 5: Proprietary Processes with Direct Reduction to Sulfur •••••••••••• 29 A. Townsend Process (Description) • • • 29 B. Stretford Process (Description) 30 C. Takahax Process (Description) ••••• 31 Section 6: Sulfur Recovery Processes ••••• 32 A. Claus Process for Sulfur Recovery 32 1. Description of Process ••• 32 2. Capital Investment .••.•••••.••• 33 B. Recycle Selectox Sulfur Recovery Process 34 1. Description of Process . 34 2. Capital Investment ••• 35 C. Graphs and Tables Table 6-1: Annual Operating Costs for Claus Sulfur Recovery Plants with Incinerators ••••••••• 36 Table 6-2: Annual Costs for Claus Sulfur Recovery Plants with Wellman-Lord 37 Figure 6-1: Typical Claus Plant for Sulfur Recovery ••.••••. 38 Figure 6-2: 3-Stage Claus plus Incinerator - Capital Cost ..•••••..• 39 Figure 6-3: 2-Stage Claus plus Wellman-Lord - Capital Cost •••••••••••• 40 Figure 6-4: Sulfur in Inlet Gas •••••• 41 Figure 6-5: Typical Recycle Selectox Flow Sheet 42 Section 7: References 43
iii 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 co 2 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:
Duty-Btu/Hr. Area-Sq. Ft.
a. Reboiler (Direct Fired) 72,000 x GPM -1 1 • 3 0 x GPM
b. Solution Exchangers 45,000 x GPM 11.25 x GPM
c. Solution Coolers 15,000 X GPM 1 0. 20 x GPM (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: GPM x AP HP = 1713 (eff1c1ency of pump)
Solution Pumps GPM x Psig x 0.00065 = HP
Booster Pumps GPM x 0.06 = HP
Reflux Pumps GPM x 0.06 = HP Aerial Cooler GPM x 0.36 = HP
K-5 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 $/Year Operating Labor (2@ 22,000) •.•••.•••.•..•.•.....•. 4.07 44~000 Supervision (1@ 35,000) ...... •...... •... 3.24 35,000
Employee Benefits@ 35% Payroll ...... 2.59 28,000
Utilities 2.78 30,000
Chemicals & Supplies ...... 1. 39 15,000
Repair Materials & Labor@ 4% Investment ...... 4.44 48, 0,00
Direct Overhead@ 5% Investment ...... 5.57 60,000
Corporate Overhead@ 3.5% Investment ...... 3.89 42,000
Depreciation ( 10 Year Straight Line) ...•...... •.. 11.11 120,000
Interest (@ 20%, first year) ...•.....•..•...... 22.22 240,000
Insurance & Taxes@ 2% Investment ...... 2.22 24,000 63.52 686,000
Plant Investment Estimate $1,200,000
K-7 TABLE I - 2 REGENERATION VESSEL SIZES {Inches)
SG . Jtion ~ I (i r·cl . Rate St ill Surge Tank Reflux Ac cum. Elash Tank Carbon Filter (X) GPM Diamete r Diam. I Length Diam. Length Diam. Length Diam. Lenqth
I - ,----
. 10 16 24 72 16 36 24 72 16 84 25 24 42 96 24 48 42 96 24 84 50 30 48 144 30 96 48 144 36 96 .JO 42 60 192 42 96 60 192 48 96 200 60 84 288 60 96 84 288 60 96 300 72 84 384 72 96 84 384 72 96 400 84 96 384 84 96 96 384 84 96 Ir Contact or Aerial l Cooler
' ~ ,if. '. k ....'·" , 11 .· .: ,
~ I I.O
~r~k Slac 1 FIGURE I ~2 GPM CIRCULATION REQUIRED PER MMCFD AMINE CIRCULATION REQUIRED 0 3 5 6 7 8 9 10 ' ' 2
,- .... _ f-- ·- -..------,..._ - ....
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--- ,___ .... ,- -,-- ,__ -e- ,- - I -- : ' I I "' 0 2 ~ 4 5 6 · 7 8 9 "'10 I ' X ACiiO GAS I , I I - V ]• / - FIGURE 1-3 V CONTACT0R CAPACITY Jl - / .,. I / JO I / V 11!1 I / '.r
2.6 ;/ / I 2"4- ' / I V 2l. I 7 / / 10 I / ~/ Q u. ' / u ;fa (/) ,a I / ,,,,. ~ " ~ 16 ' I I ~/ ~ >= I V -- t: / ~ u '-4, ·/ I / "?Jo' .~ ~ V /"" 5 / 11 I / / 1/) I ... / V (.!)< 10 / V ..J II / ~I- < V .,:;;,,.--:- QC V ~ ::) ,_____. 8 /1 / ~ r- , ,,,,,,,..- z< 1...------~, _, ~ 6 '/ / 1--- ~ ~ ~ 16,, -+ '/v ~ ~ t.---~ __.,,, ~ 1 ~ ~
0 0 lOO •OO 600 800 1000 IZOO
OPERATING P~ESSURE
K-11 .,,., 100 200 300 400 ~00 600 700 800 900 1000 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
,J --- -1- .------·--- - ·----- ·•··---i--t-t--1------1-l-7.?--t-----..1----,.__
14 ~ 01=~-~~~-:- ~ ==: :~---=~~---1------t--i----t---t-/-f-',7/-t..'---t---t--.L.----1140
120
_ T 120 ~ +- +-1__ Il~-- ---~------t--/---,'"l-v----l---f----¼---1-----1----1 ~ 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 20 1 - - -- 4-_--__ ----l -- -l--+-~-- i--;---J----.1.---l--- -- 1-- 1------~ 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, it is necessary to determine if the bed
life will be satisfactory. The following is a description of
these steps.
1. Linear Velocity
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:
Bu. Per Charge MMcf/Change = 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 extremel 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 Table 2-1
BED CAPACITIES FOR COMMON SIZE SPONGE TREATERS
Bed Capacity, Bushels (or Cu. Ft) Vessel Dia., For Bed Depth Indicated Inches 10' 12' 15' 18' 20'
24 II OD 30 34 44 52 58
30 11 ID 50 60 74 88 98
36 II ID 70 84 104 118 140
42 11 ID 96 116 144 174 192
48 11 ID 126 150 190 226 252
54 11 ID 160 190 240 286 320
60" ID 196 236 294 354 392
66 11 ID 238 286 356 428 474
72 II ID 284 340 424 410 568
84 11 ID 384 462 578 692 768
96 II ID 500 600 752 902 1004
K-17 LDA\DIN~ "PLATf"ORrw1
SPONGE' ll?C'ATE:R VESSELS
E INLET)I----
GAS OUTLET 11,--""'--:!-+~---=------...... c:"-----'--4 < s'
TOC
FIGURE" .:l.-1
FLOW DIAGRAIYI FOR IJc>ON SPONGE T ~EATEK' 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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' , II " 1,T I i I I ]1 .,Ji·'h lh :::: ·,i:: :: <: :!' 1~: :ii! I : J., I I 11 Ii ,.].. if:I :·i ;j: ! !::;:1ii :i,. :::::::,2;:;;~1~;kji' ~ . if W:::!:-1111 :1, ! l; . r i\ I~ r ~,ti I !~Ii!: :1~; 1 1 1 1 1 i: ;;t l!!!: :i;i ?/:::; ·.- .:: :;/~;: ::·1:} 1 1ii!i !•: :V • 1 ! ·!!1 llu : 1 I: ..!:::! 1! :: 1 1 111 11 11 1 :111!;:1:,:!: :'.!:::.l.•·: ·:.:!·:::1.·,: ~i: l ,1 1, · 11 • Pif " r µ .... ,,1 .. " 1 i:: .;: ::;! :::::!:::::: .. :. .;: .:.. : :,: ;I:~ !I 1 1 ,11, 1:i!i i 11, 1 iw, I 11 ,,1 ~ ,1 , l 1 1. , !!ii: :: .:,:: _ .... __._._~_...... ,...... ,_..~ _..,J....UA.,l.."-"-1.-Jw.....LI-'~...... l-.:.W.~~LU.il-=.u.LU,,LO=:-,u ...u.,,-,__,__....,...... , 1.~ u J 6 7 a ~ 100 lt!O 20 0 al'!l'9 300 _.oo ~QC 660 700 OOO t Od TREATER OPERATING PRESSURE- PSI G K-20 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%. I 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 ACID GAS ABSORBE:R FLASH DRUMS . - TREATED E:A5 RECYCLE GAS LEAN ~OLVENT FIGURE:' 3-1 PROPYLENE CARBONATE PROCESS ABSOPBE~ HIGH INTE RME'DIAT£ LOW STRIPPER PR[SS.UR[ PQf!,~UJ?E PRE.S.SUl.'E f"LA5M FLA~H f"LASH r-<-- ____R _ E:_CY__c ___ L_E: ___ __-1 VENT ~ I l'v V~NT U1 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 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. I K-27 sou~ GA.S •).-.------iA..l'------,------e-• 'NLCT ~£.4Tt:D GAS It-: ___, :. 1------i------~------ADSORBING COOLING- DESORBING- AMtNE SCRu!!>e.t:J~ G-LrtoL scR"uBoFJ F"I GU 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 --->~ H2O + SO2 2H2S = 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 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) ~ I w Sulfur Credit@ $24 ~61 /Mq ($25/LT) f / (~1,500) (83,000) (830,000) ·a-. 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 Claus Plant Size Mg/d (ltd) Co nent 5 10 100 Hanpo-,.-er Per Shifta/ 2 2 2 Direct Labor@ $6.41/hr. 1J/ $112,400 $112,400 $112,400 Supv. & Fringes@ 40\ DL 45,000 45,000 45,000 Electric ~o~er@ $0.03/Kwhl:y 3,800 7,600 75,600 3 3 Fuel @ $21.20/10 ~ {~0.(0/MCF)a/ 2,100 4,100 41,300 7 .1-'..aintenance @ 3\ INV(Claus) 1,900 . 26,100 82,000 ~ 5\ INV (W-L) 18,300 22,500 89,400 Chemicals, @ $200/ton 50\ NaOHo/ 3,500 7,000 70,000 Soft Water@ $1.14/m3 ($0.?.0/MGAtf/ 100 200 1,700 steam Credit@ $1.10/Ma ($0.50/MLBS)a/ (4,800) (9,600) (96,.400) sulfur Credit@ $24.61/Mgd (~25/LT)f/ . {43 ~700) (87,500) (875,000) Net Operating Cost $158,600 $ _:~27 ,BOO $ (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. HOT GAS !IY-P/ISS TAIL GAS 1100 · r nASiC H!£i ROILlR :,:: I ACID w COND . 11 CX) GAS •I ' s s s s COMBUSTION AIR BLIII/CP. _,64 .. • •...··-...... - _.._,.._ ..... --...... ,ru~ ,..,.. N!"-1-- FIGURE 6-1 FLOW DIAGAAH ...... TYPICAL CLAUS PLANT .. -· FOR SULFUR RECOVERY I I 40 .0 I I I I I I I I 30.0 30.0 20.0 20.0 A 15% H S in acid gas 2 10.0 1 0 50% ,-i s in acid gas ·- 9.0 8:8 2 8.0 8.0 in acid gas (:) 90% H2S 7.0 7.0 6.0 6.0 5.0 5.0 4.o· 4.0 '°-C> 3.0 ~ 3.0 >< ~ 2.0 2.0 -V, J-- V, ~ ,a I ·U w __. '-0 q;: LO ...... J-- 1.0 0.9 a.. o.~ 0.8 · 0.2 0.2 · O. l .___ _.___.'--,__,_~1,_I I II I I I 11111( ·1 I I I I I o. 1 .1 . 3 .4 . 5 . 6 J B 91, 0 2 3 4 5 6 7 8 9 10 20 30 40 60 80 l 00 \ SULFU.R RECOVERY PLANT CAPACITY:V.g/_d,(LT/d} Figure 6~2 . Claus (3- stage) plus inci~erator - c~pital cost. 40.0.------T I I 11 , -- ,-·11Tm r1 30.0 30.0 20.0 20.0 15% H S in acid g?s A 2 10.0 9.0 10. in acid gas 9. 0 50% H2S 8.0 8. 90% H S in acid gas 7.0 0 2 6.0 7.f6. 5.0 5. 4.cL 3.0 \0--- .-0 3. X 2.0 2. ~ --- -V) I . ..- ~ V) 0 . 0u l.O _.J 0.9 c::c 1. 0.8 l- 0. 0.7 0... 0.8 -c::c - u 0.7 0.6 0. 6 0.5 0.5 0.4 0.4 - 0.J ·0 •.3 .:. 0.2 0.2 I I J -I t I I I -0~l .3 4 5 6 7 89 10 .1 .2 • 3 • 4 • 5 • 6._7 -. 89 l ,. 0 2 SULFUR RECOVERY PU\NT CAPACITY, Mg/d,(LT/d) 2-sta~ lus Wellman-lord - capital cost. 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I : t I '· I ! ·1 ; I I I I I i I ; I _! 0 ! .l..!..! l .:_1_ , j - J.....!_ ' ..!...t:~...!...~-~ ! : J . ' I I ' : I I I J I I . ' l I ' I ' • I i ' ! ' ' . I ; I I : I ! ; i ,....,....-J ' ' 4 : ' ' I ' ' ! . ' . ; . , __1_ ' .. I ·-.-7 I I I ~=J-=-----:r:-: ~ FIGURE 6-4. d+±J s_y(j ....~i :w:...t: '""t"" iv"(J ;-i~-Gis·-·_( ·<·]}JcJ~~~~Ff7n~v--- +·7 -'_- , ~-~~- : I : : : -.--•. - ' ,,- - ~ - ...:.::~ I -== -~--- ·r-, ::· ·~----c-i.· .. .--.~i-·, ' '._j'! -1--+-I ', ,• · ,1 • ,' ; I I • I _·_. ~-rp!.! _.:. ---1 .!.~e-: q·· ...T .-- . 7 r - --,-,--.-. + . -; -;- : J ! I I I .• - -i ..J ! I I . .... J....!..J I _; . . -~~ ~_!_ I . +- -·.' _,I _:' ,7 ·. I______; . I --r~.·. ·---:-o ,l--;--·'. •...... ' •, Ti"",, I : : · 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----- SELECTOX ----.::::: SELECTOX SELECTOX OR OR HTM" CATALYST ALUMINA ALUMINA ------=-.IL-- CATALYST __ CATALYST ---- TAIL GAS TO TREATING OR INCINERATOR ACID GAS ----t FEED IITM" 1 TO 40% HTM" t-12S RECYCLE BLOWER s s s "HTM = HEAT TRANSFER MEDIUM APPROXIMATE SULFUR YIELD 82% 12% 2% I 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. 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