P.1 

OAR Box 1123 

Prepped by Ollie Stewart 

Document Number: 

168) ll-B-2 

Docket Number: 

A-84-44 

Printed 4/16/2008 2:49:56 PM Sheet 168 ofl 78 
P.2 

R.JLS.050 

September 11, 1984 

DRAFT- " 

JL-H.-2-

MEMORANDUM 

TO: Docket A-80-20A 

FROM: Jeffrey L. Shumaker 

SUBJECT: Sweetening Cost Factors 

Purpose 

.Develop capital and operating cost factors for estimating sweetening 

unit costs for a range of well head gas compositions. 

Summary of Factors 

Small Large 

Facilities Facilities 

(DEA; (Sulfinol; 

<100 LT/D) >100 LT/D) 

$0.51 

Component 

Capital Cost 

Units ,of, Factor 

Annualized Capital Cost $0,064 

Operating and Maintenance: 

Regeneration Heat 

Pumping Cost 

Labor 

Maintenance Supplies 

Chemicals 

Total Operating and Maintenance: 

$127,000/year + $0.3024 0_^ 

Total Annualized Cost: 

$127,000/yea.r + $0.3664 OR. 

$0,257 

$0.0167 

$127,000 

$0.01 

$0.0187 

$0.61 SCFD Acid; Gas 

$0,069 /year per SCFD Acid Gas 

$0,158 /year per SCFD Acid Gas 

$0,012 /year per SCFD Acid Gas 

i 

$127,000 /year 

$0.01Z. /year per SCFD Acid Gas 

$0.0187 /year per SCFD Acid Gas 

$0.2'H /year per SCFD Acid Gas 

$0. /year per SCFD Acid Gas 
c 

pp 

Docket A-80-20A 

Page Two 

September 11, 1984 

j,-.-^ —=f Approach Basis for Small Facility (<100 LT/D) Factors 

Small facilities of the type likely to have adverse economic impacts 

generally have a small percentage of acid gas in the well head gas stream. 

As a result, the amine sweetening processes are commonly used for facilities 

of this type. For the purpose of generating cost factors the diethanolamine 

process is considered. 

Installed capital cost and operating cost factors where developed 

based on information learned in contacting 4 sweetening unit design firms 

(Attachments 1 through 7). 

Capital Cost 

Factor Basis: $4000/GPM of DEA circulation installed capital cost in 

mid 1984 dollars (See Attachment 1). Other factors are from Attachment 2. 

Panital rn*t - $ 400 0 v Ga 1 S °

1 > n y 1 b So1 'n v 105.14 lb 

uapitai ^osz Gal/min

 x 8.6 lb Sol'n x 0.35 lb DEA x lb mole DEA 

lb mole DEA lb mole Day 

x 0.5 lb mole acid gas 379 SCF x 24 x 60 min 

Capital Cost = $0.51/SCFD Acid Gas (mid 1984 dollars) 

Annualized Capital Cost Assuming 20 year life and 10 percent interest rate 

= $0.51/SCFD x .11246 = $0.064/year per SCFD Acid Gas. 

Operating Costs 

1. Regeneration Heat: Basis 950 BTU/Gal DEA approximate reboiler heat 

duty (See Attachment 2, pg. 15-21). Also assume a value of $3.35 per 

MMBTU of natural gas to fire the reboiler (high end of EAB's price 

range), and an 80 percent boiler efficiency. 

950 BTU . Gal Sol'n .. lb Sol'n .. 105.14 1b Heat Cost Gal Sol'n

 A 8.6 lb Sol"n A 0.35 lb DEA A lb mole DEA 

lb mole DEA lb mole _1 _ $3.35 350 Days 

x 0.5 lb mole Acid Gas x 379 SCF x .80 x MMBTU x year 

= $0.257/year per SCFD Acid Gas 
P.4 

irf " T 

Docket A-80-20A 

Page Three 

September 11, 1984 

2. Pumping Cost: The pumping cost is a function of the GPM of amine 

circulated and the pressure pf the gas in the contactor (i.e., the gas 

well pressure minus transmit!on temperature). The higher the pressure 

the higher the energy requirement and thus the cost. Normally contactors 

are not run above 1000 PSI. Well pressures higher than this generally 

are let down prior to sweetening. Lower pressures will incur lower 

pumping costs. Thus, it is assumed that the contactor operates at 

1000 PSI pressure. Further it is assumed that a pump efficiency of 

70 percent is achieved. Using the hydraulic horsepower equation 

(Perry's 5th Ed., pg. 6-3), an assumed energy cost of $0.05/KWH, and 

other appropriate factors the pumping cost is as follows: 

Pumoina Cost = Hp 1000 PSI Gal Sol'n 

Pumping Lost 17 i 4 PSI-Gal Sol'n/min

 x 1 x 8.6 lb Sol'n 

lb SoTn 

0.35 lb DEA 

Day 

24 x 60 min 

.7457 KW $0.05 105.14 lb DEA lb mole 

0.70 Hp

 x KWH x lb DEA x 379 SCF 

8400 hr 

year l..0167/year per SCFD Acid Gas 

3. Operating and Maintenance Labor: One person per shift for all cases; 

$14.50/hr labor rate = $127,000/year. 

4. Maintenance Supplies: Based on 2 percent of capital cost 

= $0.51/SCFD x 0.02 = $0.01/year per SCFD Acid Gas 

5. Replacement Chemicals: Based on -\>25 percent per year replacement rate 

and a DEA cost of $7.50/Gal a 50 GPM amine unit incurs an annual 

replacement costs of $7300/year in 1984 dollars (Attachment 1). 

$7300 1 Gal Sol'n lb SoTn 105.14 lb 

year

 x 50 Gal Sol'n/min x 8.6 lb Sol'n x 0.35 lb DEA x lb mole DEA 

lb mole DEA lb mole x 0.5 lb mole Acid Gas

 x 379 SCF x 24 x^min

 = $0-0187/year per SCFD 

Acid Gas 
P.5 

r|M"!i •-• 

Docket A-80-20A 

Page Four 

September 11, 1984 

J'J' ^ Basis for Large Facility Factors 

Large plants (>100 LT/D sulfur feed) generally employ a sweetening 

technology that requires higher capital investment than the DEA process but 

which has lower operating costs per unit volume of gas treated. Systems 

using sulfinol as the absorbent are common in this larger size range and this 

process is selected as the basis for estimating costs for sulfur feed rates 

greater than 100 LT/D. The sulfinol process, like other processes commonly 

employed in larger plants, is a proprietary process distributed under 

licensing agreements. Consequently, less information is available on this 

process in public literature and the development of general cost factors 

involves a higher degree of approximation than was needed for the DEA process 

factors discussed previously. The following assumptions are made: 

t The sulfinol solution is 33 percent solfolane, 33.3 percent 

di-isopropanol amine, and 33 percent water. 

• The acid gas pickup rate is 10.5 SCF/gallon solution (the midpoint 

of the range in Attachment 2). 

• The reboiler heat duty is 1135 Btu/gallon lean solution (the 

midpoint of the range in Attachment 2). 

• Solution density is 9.1 lb/gallon; average MWT of solvents is 

126.7. 

• The capital cost of the sulfinol process is 20 percent higher than 

the DEA process capital cost. 

With these assumptions, the cost factors can be calculated as follows: 

Capital Cost 

$0.51/SCFD x 1.2 = $0.61/per SCFD Acid Gas 

Annualized Capital Cost: 

$.61/SCFD x .11246 = $0.069/year per SCFD Acid Gas 
P.6 

-fi nn'!1 

Docket A-80-20A 

Page Five 

September 11, 1984 

Operating Costs 

1. Regeneration Heat: 

1135 Btu Gal SoTn 1 $3,^5 350_Davs 

Gal SoTn

 A 10.5 SCF Acid Gas x .80 MMBtu x year 

= $0.158/year per SCFD Acid Gas 

2. Pumping Cost: 

J B ,. 1000 PSI „ Gal SoTn .7457 KW 

1714 PSI-Gal Sol'n/min

 x 1 x 10.5 SCF x Hp - ~ 

x W" x 24 xD6^ min x ^T^ = »^2/year per SCFD Acid Gas 

3. Operating and Maintenance Labor: (Same as for DEA) 

$127,000/year 

4. Maintenance Supplies: 

$0.61 x .02 = $0.012/year per SCFD Acid Gas 

5. Replacement Chemicals: (Same as for DEA) 

$0.0187/year per SCFD Acid Gas 

JLS:dw 

Attachments (7) 
P.7 

•• r-irjji 

ATTACHMENT 1 

RADIAN 

TELEPHONC CALL RECOftO 

Qloco-11* Call 

Oouctolng tell 

Fatten Calling 

Jeffrey L. Shumaker, Radian 

Patton Calla* 

Kent Carlisle, Ortloff 

rtojMt *>. 

231-059-06 

NSS-0NSH0RE NATURAL GAS PRODUCTION 

Oaw 

8/10/84 and 8/24/84 

Tiat 

Activity 

Activity 

(915) 699-3258 

PUtiltK"l("' 

CtMTfl Sukjact 

SWEETENING COST ESTIMATES FOR SMALL PLANTS 

TOPICS DISCUSSED ANO ACTION TAKEN 

• Dominant parameter is moles of acid gas processed - They (Ortloff) use 

this approach for quick analyses. 

• $4000/GPM amine circulation is a good installed capital costs factor in 

1984 dollars. 

• 0.5 moles acid gas/mole amine maximum. ^0.35 factor recommended. 

• Fuel consumption: 1 lb steam/min per GPM circulation; 80 percent efficient 

boiler; 1000 BTU/lb steam or 15,000 BTU/hr per mole acid gas removed 

each hour. 

• Pumping costs - Dependent upon gas pressure in the contactor - use 

standard horsepower equation; stripper is at 1 atmosphere; contactor 

pressures range from <500 to 1000 PSI—if >1000 pressure usually let 

down to 1000 PSI. Use 1000 lb assumption. 

• Solvent replacement - -^25 percent per year solvent replacement. Some 

examples follow: 400 GPM unit - $340,000 initial solvent cost, $85,000/year 

replacement solvent lost; 50 GPM , 3900 gallons capacity - $29,000 

initial solvent cost, $7300/year replacement solvent cost. Solvent 

costs about $7.50 per gallon in 1984 dollars. 

• Operating labor - Some companies will not leave one operator alone in a 

sour gas processing facility; process will not need more than 1. 

• Operating costs or at least heat and pumping cost factors are appropriate 

irrespective of size. 

• The capital cost factor is appropriate for small facilities (5 LT/D), 

but will become less appropriate for smaller and smaller facility sizes--

i.e., because of the loss of economy of scale. 

• If he were estimating the costs quickly, he would use the above factors. 

• The contactor is not a dominant capital cost component. It runs ^10 to 

15 percent of the total cost. The heat exchangers and pump make up a 

large fraction of the capital cost. These vary in cost only with the 

GPM of amine solution. 
P.8 

• •" • r | IFJ 1' ! 1' " ' 

TELEPHONE CAU RECORD 

Olncoalat Call 

Oout«olas Call 

rttaon Calllai 

Ftrton Callad 

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Itam 

Activity 

Activity 

4 .T <£?? 3Z£TB 

Canatal Subjact 

CUH£-CK./*XJ ^u)e<Lr^>Jf^ (?os r Fftcro^ s 

TOPICS DISCUSSED ANO ACTION TAKEN 

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P.9 

*VrT^<K^«e-V^T £- • 

rf |"!|P|tlPrT 

fori, Bacon &Mm 

3ncorporalr& 

Engineers - Constructors 

August 8, 1984 

Radian Corporation 

P. 0. Box 1300 

Research Triangle Park, NC 27709 

Attn: Mr. Jeff Schumaker 

Dear Sir: 

Attached you will find a copy of the Gas Treating Section 

from NGPSA Engineering Data Book as was discussed today. 

You will also find a copy of Process Design of various types 

of Sulfur Removal Processes for your files. 

I regret that I could not ?r-e of further service at this time. 

Sincerely, 

FORD, BACON & DAVIS, INCORPORATED 

Texas Division 

L. L. Patterson, P. E. 

Principal Process Engineer 

JGJ10/LLP 

P.O. Box 38209 • Dallas, Texas 75238-0209 • (214) 238-6500 • Telex 79-2488 

2201 North Central Expressway • Richardson, Texas 75080 
P.10 

c 

c 

L 

2. Dingman, J. C, and LeBas, C. L., Oil Gas J.,62 (No. 

5), 75-77 (1964). b. 

3. Hales, G. E., "Molecular Sieves in Cryogenic PlantsOperating Experiences and Guidelines", Natural Gas 

Processors Association, Regional Meeting, New Orleans, Louisiana, January 21,1974. a, b, c. Mole sieve. 

4. Jacoby, Robt. H., "Calculation of Methanol Requirements to Prevent Formation of Gas Hydrates", Gas, 

Feb. 1955, p. 114. Methanol. 

5. Kessock, Jr., A., and Fris, J. P., "How Kenova Plant 

Dries Feed Gas", Hydrocarbon Processing and Petroleum Refiner, 44, No. 4, pp 123-124 (1965). c. Mole 

sieve. 

6. Kraychy, P. N. and Masuda, A., "Molecular Sieves 

Dehydrate High-Acid Gas at Pine Creek", Oil and Gas 

Journal, August 8, 1966. a, c. Mole sieve. 

7. Loomer, J. A., and Welch, J. W., Proceedings Gas 

Conditioning Conference 1961, 71-80, The University 

of Oklahoma, a, c, d. Glycol. 

8. McCarthy, E. L., Boyd, W. L., and Reid, L. S., Petroleum Transactions AIME, 189, 241 (1950). b. Water 

Content. 

9. McLeod, H. 0., and Campbell, J. M., J. Petro. Technol., 

13, 590-94 (1961). b. Hydrate equilibrium. 

10. Patterson, E. O., Jr., "Get Low Dew Points with Solid 

Desiccants", Oil & Gas Journal, March, 1969. a, c. 

Solid desiccants. 

11. Perry, C. R., Oil Gas J., 58 (No. 32), 71-74 (1960). a, c. 

Glycol. 

12. Polderman, L. D., Oil Gas J., 54 (No. 5), 198-199 (1955). 

b. Glycol. 

13. Russell, T. H., Oil Gas J., 60 (No. 45), 94-99 (1962). c. 

Glycol. 

14. Scauzillo, F. R., J. Petrol. Techno., 13, 697-702 (1961). 

b. Glycol. 

15. Swerdloff, W., Proceedings Gas Conditioning Conference 1969, A-l-A-25, The University of Oklahoma, a, 

c, d. Glycol, solid desiccants. 

16. Weyemuller, G. H., Harlan, J. D., and Roberts, Darrel, 

"Molecular Sieves Keep Helium Plant on Stream", 

Chemical Processing, November 1966. c, d. Mole sieve. 

17. Worley, M. S., Proceedings Gas Conditioning Conference 1967, H-l-H-15, The University of Oklahoma, a, 

c. Glycol. 

18. Wunder, J. W. J., "How to Design a Natural Gas 

Drier", Oil & Gas Journal, August 6, 1962, page 137. 

a. Solid bed dehydration. 

19. Young, Jr., Philip L., "Molecular Sieve Adsorption in 

LNG Plant Pre-Purification", Presented at the Distribution Conference American Gas Association, Houston, Texas (May 6-9,1968). a, c, d. 

20. Campbell, J. M., "Gas Conditioning and Processing", 

Third Edition 1974. 

21. Weibe, R., and Gaddy, V. L., J. Am. Chem. Soc. (1941), 

Vol. 63, 475. CO-rH.O. 

22. Selleck, F. T., Carmichael, L. T., and Sage, B. H., Ind. 

Eng. Chem. (1952), Vol. 44, 2219. H:S-H,0. 

23. Hall, G. D., and Polderman, L. D., Chem. Ene 54 

(No. 10), 52-58 (1960). TEG. 

Gas treating 

There are many methods that may be employed 

to remove acid gases (primarily H2S and C02 ) from 

natural gas. The methods may be broadly categorized as those depending on chemical reaction, physical solution, or adsorption. 

Processes employing each of these methods are 

described and a general bibliography is included. 

Safety precautions 

Hydrogen sulfide is a highly toxic gas, rapidly 

causing death when inhaled in high concentrations. 

Even at low concentrations, irritation of the eyes, 

nose and throat are possible. These points illustrate 

the hazard of varying concentrations: 

1. Threshold limit value, (TLV), for prolonged 

exposure: 10 ppm* 

2. Slight symptoms after several hours exposure: 

70-150 ppm 

3. Maximum concentrations that can be inhaled 

for one hour without serious effects: 170-300 ppm 

4. Dangerous after exposure of 30 minutes to one 

hour: 400-700 ppm 

5. Fatal in 30 minutes: 700 ppm and above 

* 1.0 grain/100 scf is equivalent to 17.1 ppm by 

weight, or 15.9 ppm by volume. 

The greatest danger from the inhalation of HoS is 

from its acute effects; it is not cumulative in action. 

Hydrogen sulfide is a flammable gas. In concentrations of 4.3 to 46.0 vol % in air, it explodes upon 

ignition. The vapors are heavier than air and may 

travel considerable distance to a source of ignition 

and flash back. 

Anyone engaged in the design or operation of a 

facility in which H2S is present should consult some 

source of information on safety precautions to be 

taken. One such source is the Manufacturing Chemists Association, Manual Sheet SD-36. 

GENERAL CONSIDERATIONS 

There are many variables involved in gas treating. This makes the precise definition of area of 

application of a given process very difficult. Some 

of the factors to be considered in making a process 

selection are: 

1. Type and concentration of impurities in. the 

sour gas. 

2. Specifications for the residue gas. 

3. Specifications for the acid gas. 

4. Temperature and pressure at which the sour 

gas is available and at which the sweet gas must 

be delivered. 

5. Volume of gas to be processed. 

6. Hydrocarbon composition of the gas. 

7. Selectivity of acid gas removal required. 

1 5 - 19 Rev. 1976 
" I 

P.11 

8. Capital costs and operating costs. 

9. H2S to CO;, ratio in the sour gas. 

Gases with high HL.S to CO., ratios are less corrosive than those having low H2S to CO* ratios. In 

general, a high H2S-to-C02 ratio gas will permit 

higher acid-gas concentration in the absorbing solution. Carbon steel may be satisfactory for much of 

the equipment in any process. However, recommendations for materials of construction are best handled 

on an individual unit basis by the contractor or 

supplier. 

All gas-sweetening units should have a good inlet 

separator. A continuous introduction of one ppm of 

foreign material entering a 100-MMscfd plant represents an accumulation of 1 ton/yr. Contacting towers 

should be equipped with differential pressure gauges 

to aid in detecting flooding and/or foaming. The 

absorber should have a mist extractor above the top 

tray and/or an efficient separator in the exit gas line. 

Reboiler heating elements should always be submerged. 

There should be adequate space allowance for vapor flow, 

and if steam is employed as the heating medium, it should 

always be saturated—never superheated. In solution heat 

exchangers, the acid-gas rich solution should be on the 

tube side in upflow. 

A flash tank between the absorber and the heat 

exchanger will reduce erosion and also minimize 

hydrocarbon content in the acid gas. Valves for 

pressure let-down are subject to severe service. 

Stainless trim with stellite facing is commonly used. 

Where applicable a solution reclaimer should be 

considered. The reclaimer will normally handle 2% 

to 3% of the total solution circulation. Solution filtration should also be provided. Common design 

filtration ranges from 5% of circulation to full 

stream. In some cases, activated carbon treatment of 

a side stream is used to adsorb impurities. 

Activated carbon cannot replace a reclaimer for 

removal of degradation products. Where a separate 

activated carbon unit is provided, it should be located downstream of the filter. 

Here are some operating precautions which experience 

has proven to be good practice: 

1. All equipment should be cleaned after installation 

and before initial operation. This cleaning may be done by 

circulating a hot 1% to 2% by weight solution of low chloride 

content soda ash. 

2. Where water make-up is used, it should come from 

steam condensate or deionized water. 

3. Do not use soap-based lubricants. 

4. A short-term (seconds to minutes) increase in differential pressure across the contactor can indicate foaming. 

A gradual increase in differential pressure over a long 

period of time may indicate fouling of trays or packing. 

5. If foaming is thought to be occurring, check for entry 

or condensation of liquid hydrocarbons, and run laboratory 

tests with and without foam inhibitors. Never add a foam 

inhibitor to the plant solution without prior testing in the 

laboratory. 

6. If a reclaimer and/or filter is provided, operate it 

from the day the plant is started. Change or clean filters 

as required. 

7. Lean solution to contactor and stripper overhead 

condenser outlet temperatures should be kept as low as 

possible to minimize solution and water losses. Lean solution temperatures should be 10°-15" above entering gas 

temperature to minimize hydrocarbon condensation which 

may cause foaming. 

8. Monitor corrosion rates with coupons or resistance 

probes. 

9. Lean and rich solutions should be analyzed routinely 

for absorbent concentration and acid gas loading. 

10. Housekeeping, particularly packing gland leaks, is 

a major factor in controlling chemical losses. 

Chemically reacting processes 

These processes are based on reversible chemical reactions. The absorbent reacts with CO, and H.S in the contactor with the direction of reaction being reversed by high 

temperature and/or low pressure in the stripper. Fig. 15-

25A lists approximate guidelines for a number of chemically reacting processes, each of which will be discussed 

below. Fig. 15-25B gives a summary of process capabilities. 

Fig. 15-26 depicts the process flow of a typical chemical 

absorption process. 

Fig. 15-28 is a table of physical properties of gas-treating 

chemicals. Figs. 15-29 through 15-32 show vapor pressures 

at various temperatures, and freezing points and specific 

gravity for some of the treating chemicals. 

AMINE PROCESSES 

Monoethanolamine, conventiona l Diethanolamine, 

SNPA-DEA and high-loading DEA processes will be considered here. 

The ethanolamine processes are particularly applicable 

when acid gas partial pressures are low and/or low levels 

of acid gas are desired in the residue gas. 

Units employing MEA have been successfully designed 

in the 50-100 psig operating range; however, care should 

be used in the design of DEA plants for low pressure gas 

sweetening. The actual CO, pickup has been reported to 

be lower than expected. In fact, a high H,S selectivity has 

been reported which indicates the process may be rate 

limiting (18). 

Because of the water content of the circulating solution, these processes are well suited for gases rich in 

heavier hydrocarbons. The aqueous amine processes have 

very little selectivity for H,S or CO,, except as noted. The 

DEA processes are particularly applicable when the gas 

stream contains carbonyl sulfide which reacts irreversibly 

with MEA. 

The SNPA-DEA process (33) is patented in Canada and 

possibly other countries, excluding the U. S. The patent 

covers a range of loading from 0.9 to 1.3 moles of DEA/ 

mole of acid gas. This process is used for high pressure, 

high acid gas content streams having a relatively high 

ratio of H,S/CO,. 

For MEA units, total acid gas pick-up is normally limited to 0.33 to 0.35 moles of acid gas/mole of MEA. The 

concentration of MEA is generally held in the 15 to 18 

c 

e 

e 

Rev. 1976 15 -2 0 
MI""" 

P.12 

c 

FIG. 15-25 A 

Approximate guidelines for sweetening processes'1' 

c 

c 

c 

MEA DEA SNPA-DEA<9> Econamine(DGA) Sulfinol 

Acid gas pickup, scf/gal @ 100°f. normal range'21 3.1 to 4.3 3.8 to 5.0 6.7 to 7.8 4.7 to 6.6 4 to 17 

Acid gas pickup, mols/mol amine, normal range131 0.33 to 0.40 0.35 to 0.65 0.72 to 1.02 0.25 to 0.3 N.A. 

Lean solution residual acid gas, mol/mol amine, normal range141 .. . 0.12+ 0.08± 0.08± 0.10+ N.A. 

Rich solution acid gas loading, mol/mol amine, normal range' 3'— 0.45 to 0.52 0.43 to 0.73 0.8 to 1.1 0.35 to 0.40 N.A. 

Solution concentration, wt. V normal range 15 to 25 25 to 35 25 to 30 50 to 70 3 components.Vanes 

Approx. reboiler heat duty. Btu/gal lean solution'51 1.000 to 1.200 900 to 1,000 900 to 1.000 1,100 to 1.300 550 to 1.720 

Steam heated reboiler tube bundle, approx. average heat flux, Q/A = 

M Btu/hr-ft2'

61

 9 to 10 9 to 10 9 to 10 9 to 10 9 to 10 

Direct fired reboiler fire tube, average heat flux, 

Q/A = M Btu/hr-ft2'

6

' 8 to 10 8 to 10 8 to 10 8 to 10 . 8 to 10 

Reclaimer, steam bundle or fire tube, average heat flux, Q/A = 

M Btu/hr-ft? <« 6 to 8 N.A'" N.A.l'l 6 to 8 N.A. 

Reboiler temp., normal operating range, "FM 225 to 260 230 to 250 230 to 250 250 to 260 230 to 280 

Heats of reaction;""'approximate: Btu/lb H2S 550 to 670"°> 500 to 600"" 511 674 Vary with loading. 

Btu/lbC02 620 to 700 580 to 650 653 850 Vary with loading. 

N.A.—not applicable or not available 

NOTES: 

1. These data alone should not be used for specific design purposes. Many design factors must be considered for actual plant design. 

2. Dependent upon acid gas partial pressures and solution concentrations. 

3. Dependent upon acid gas partial pressures and corrosiveness of solution. 

4. Varies with stripper overhead reflux ratio. Low residual acid gas contents require more stripper trays and/or higher reflux ratios yielding larger reboiler duties. 

5. Varies with stripper overhead reflux ratios, rich solution feed temperature to stnpper and reboiler temperature. 

6. Maximum point heat flux can reach 20.000 to 25.000 Btu/hr-ft2 at highest flame temperature at the inlet of a direct fired fire tube. The most satisfactory design 

ol fire tube heating elements employs a zone by zone calculation based on thermal efficiency desired and limiting the maximum tube wall temperature as required 

by the solution to prevent thermal degradation. The average heat flux. Q/A, is a result of these calculations. 

7. Reclaimers are not used in DEA systems. Activated carbon filters are used to remove degradation product. 

8. Reboiler temperatures are dependent on solution cone, flare/vent line back pressure and/or residual CO? content required. It is good practice to operate the 

reboiler at as low a temperature as possible. 

9. Canadian and Foreign Patented DEA Process. 

10. B. L. Crynes and R. N. Maddox. Oil Gas J., p. 65-67. Dec. 15 (1969). The heats of reaction tend to vary with acid gas loading and solution concentration. 

11. J. I. Lee. F D. Otto, and A. E. Mather. Can. GPA Research Seminar. Nov. 1972. The heats of reaction tend to vary with acid gas loading and solution concentration. 

FIG. 15-25B wt.% ran^e to limit corrosion. DEA concentrations up to 

Summary of process capabilities ('35 wt%jare acceptable. 

Normally Equilibrium solubility data and application for H £ and 

capable of R«m.ve> CO, in MEA and DE A solutions are given in reference s 

Meeting Mercap- 43 an d 44 

Va, Groin tan 8, Selective 

s p £ $!°!, Ab»rb.d solution Degraded by u For both prwMMses contactors m d stripper s typically 

have 18 to 22 trays. MEA units should hav e a reclaimer. 

M °"° . References: 11.14,15,24,25,26.33,34x45,46,47, and 51. 

ethanolamine Yes Partial No Yes (COS. CO,. CS,) 

Diethanolamine Yes Partial No Some (COS, CO,, CS,) 

Sulfinol Yes Yes Yes3 Some (CO,. CS,) Example Problem 

Diglycolomine Yes Portial No Yes (COS. CO,. CS,) » * i * ,^„»,» . , . . nnn • • . . 

Seiexoi Yes Slight Yes3 No A natural ga s stream of 50 MMscfd at 900 psig is to be 

Hot Pot-Benfield Yes1 No2 No No sweetened. The sourga s contains 0.50%JH2S and 2.0% C0 2 . 

Fluor Solvent No4 No No No The rich amine leaves the contactor at 140°F (60°C). 

Strettord Yes Partial Yes Yes (HCN via 

thiocyanate) Solution 

Iron Sponge Yes Partial Yes — 

Mol Sieve Yes Yes Yes3 — Assuming ideal gases, the partial pressure ofH^S in the 

1. Hi-pure version. SOUr ga s will be 

2. Hydrolizes COS only Q 0Q 5 x 91 _ y (760/1 4 7 ) _ 237 m m Hg 

J. Some selectivity is exhibited by these processes. 

4. Can be met with special design features. The partial pressure of COo is 

) - ' • groin H,S/100scf 

0.02 X 915 X (760/14.7) = 946 m m Hg 

15-21 Rev. 1976. 1981 
P.13 

The ratio of acid gas partial pressure is 

Rv = (237/946) = 0.25 

The equilibrium composition of the H2S in the amine 

can then be read from the known value of Rv (0.25) and the 

H2S partial pressure of 237 mm Hg. From Fig. 15-35 we 

read the equilibrium concentration of H2S to be 0.09 moles 

ofH,S per mole of MEA. From Fig. 15-39 using the 237 mm 

Hg of H;,S partial pressure in the sour gas and the equilibrium concentration of 0.09 mols of H^S per mole of MEA 

we read an RL value of 0.17. 

RL = (mols H2S/mole MEA/mols COj/mole MEA) 

Mols COjmole MEA = (0.09/0.17) = 0.53 

The equilibrium acid gas concentration of the MEA 

solution at the estimated 140°F temperature leaving the 

absorber will be 

0.09 mole H,S + 0.53 mole C02 = 

0.62 mole acid gas total/mole MEA 

DIGLYCOLAMINE PROCESS 

This process uses Diglycolamine (DGA) brand of 2-(2-

aminoethoxy) ethanol in aqueous solution as the sweetening agent. The process can be used successfully for treating 

natural or refinery gas streams containing carbonyl sulfide 

and/or carbon disulfide. A solution concentration of 50 to 

70 wt% DGA is used to permit higher acid gas removal per 

gallon of solution. 

DGA has an advantage for plants operating in cold climates where freezing of the solution could occur. A reclaimer should be provided. References: 17, 32, and 18. 

HOT CARBONATE PROCESSES 

This process was developed at the U. S. Bureau 

of Mines and employs an aqueous solution of potassium carbonate. The absorber and regenerator both 

operate at elevated temperatures in the neighborhood 

of 230° to 240°F. The process cannot be used on gas 

streams containing only H2S. It also encounters 

difficulty in attempting to treat gas to 0.25 grains of 

HgS/lOO scf, the normal specification. 

The most popular.of the carbonate processes are 

those containing an activator. These proprietary 

processes are: 

1. Benfield: Several activators. 

2. Girdler: Alkanolamine activators. 

3. Catacarb: Alkanolamine and/or borate activators. 

4. Giammarco-Vetrocoke: Arsenic and other 

activators. 

In general, plants are constructed of carbon steel, with 

stainless steel being used at critical points. Because the 

circulating solution is a concentrated salt solution, care 

must be taken to avoid cold spots in lines and equipment. 

References: 2, 5, 23, and 50. 

STRETFORD PROCESS 

This process was developed by the North Western 

Gas Board and Clayton Company Ltd. of England. It 

utilizes a quinone solution to convert the H..S to sulfur. The process is selective for H.S and can reduce 

the H2S content well below the pipeline specification. 

Physical-solvent processes 

These processes are based on physical absorption 

and operate in much the same way that the oil absorber in a natural gas plant does. In general, a 

physical solvent process should be considered when: 

1. The partial pressure of the acid gas in the feed 

is 50 psi or higher. 

2. The concentration of heavy hydrocarbons in 

the feed gas is low. 

3.. When only bulk removal of the acid gas is 

desired. 

4. When the solvent is able to do a satisfactory 

dehydration as well as acid gas removal. 

5. When selective H2S removal is desired. 

The absorber employed is a conventional trayed 

or packed tower. Regeneration is accomplished by 

using one or more of these techniques: 

1. Multi-stage flashing to lower pressures. 

2. Relatively low temperature stripping with an 

inert gas. 

3. Heating and stripping with the solvent vapor. 

Physical-solvent processes employed in the natural gas industry include the Fluor Solvent Process, 

the Selexol Process, the Sulfinol Process, and the 

Rectisol Process. 

In general, all the physical-solvent processes have 

the limitation of a high affinity for heavy hydrocarbons. Fig. 15-27 is a typical process flow diagram 

for removal of both C02 and H2S. 

FLUOR SOLVENT 

This process is patented by the Fluor Corporation and 

is based on the use of anhydrous propylene carbonate. The 

temperature of the lean solution to the absorber is usually 

well below ambient, and some method of refrigerating the 

solvent is usually necessary. References: 8, 9, 13, 19, 37, 

and 38. 

SELEXOL 

This process uses the dimethylether of polyethylene 

glycol as a solvent. It can be used for obtaining very low 

partial pressures of acid gas by adding di-isopropanolamine. 

The process is licensed by the Allied Chemical Corp. The 

temperature of the lean solution to the absorber is approximately ambient. 

With the use of this solvent it is possible to achieve 

pipeline specification residue gas with respect to H.S, 

while rejecting as much as 85% of the CO. present in the 

original sour ga3. References: 28, 31, 39, iO, and 55. 

SULFINOL PROCESS 

The Sulfinol solvent consists of sulfolane (tetrahydrothiophene dioxide) di-isopropanolamine, and water in proportions varied to suit the application. The process can be 

designed to meet a wide range of product specifications. 

Bulk CO, removal is achieved with minimal heat requirement and without refrigeration. 

c 

c 

c 

e 

Rev. 1976. 1981 15-22 
P.M 

c 

Removal of H.S, COS, mercaptans, and organic sulfides 

to meet pipeline specifications is achieved over a wide 

range of operating conditions. The process is licensed by 

Shell Development Company. References: 20, 21, and 27. 

RECTISOL PROCESS 

This process uses methanol as a solvent, and was developed by the German Lurgi Co. and Linde A. G. Because of 

the vapor pressure of methanol, the process is normally 

applied at low temperatures of -30 to -100°F for the Rectisol solution. It has been applied for the purification of 

natural gas for LNG production. The process is best suited 

where there are limited quantities of ethane and heavier 

components. 

Fig. 15-28 is a table of physical properties of gastreating chemicals. Figs. 15-29 through 15-32 show 

vapor pressures at various temperatures, and freezing 

c 

points and specific gravity for some of the treating 

chemicals. 

Dry bed processes 

IRON-SPONGE PROCESS 

This process utilizes a hydrated iron oxide, impregnated 

on wood chips, which reacts only with H.S. It finds application only on gases with low concentrations of H.S. In 

batch-process operations, the iron sulfide can be regenerated a few times with air. In addition, the chips must 

periodically be replaced. Care musi be taken during recharging to prevent air pollution and/or fire because the 

spent iron sulfide on wood chips can ignite spontaneously. 

Successful treating requires controlled gas temperature 

and water content of the bed. Reference 56 (copies available from Connelly, Inc., 3154 South California Avenue, 

Chicago, 111. 60608) discusses this process. 

FIG. 15-26 

Typical ga< sweetening by chemical reaction 

Sweet Gas /O 

w 

Contactor 

7\ " 

(p) Cooler 

Optional equipment 

c 

Sour gat 

azD 

"B—<&-% 

Fuel 

7 

ox 

I flash tank . 

Acid gas 

Courtesy Fish Engineering & Construction, Inc. 

FIG. 15-27 

Typical gas sweetening by physical absorption 

c 

Sour gas 

e 

H,S ond 

stripping gas 

>

Cooler I 

if * / * 

rH

 l v % T 

L J • i 

• 

t 

Solvent 

recovery 

• Stripping gas 

Courtesy fish Engineering & Construction, Inc. 

Rev. 1981 15-2 3 Rev. 1976 
P.15 

Formula 

Molecular Wt. 

leiling pom 9 760 mm Hg, "I. 

Frteiing point, 'I. . 

Critical constant! 

failure, psia. 

Temperature, °F. . 

Density 9 20'C. gm/c< 

Wtight, lb/get 

Sptcific gro«ity 2O°C./20°C. 

Specific hoot 9 U'F. ITU7lb/"F 

Thermal conductivity. 

ITU/hr. FT', "F./ft.. (a! W f.... 

letlnt htot sf voportiation. ITU/lb 

Noot of rooctien. ITU/lb of tcid Coi 

H.S 

CO. 

Viuoiity, cp 

Itfractivt indti. «„ 6I°F. 

fink point, COC. 'I 

HOCiHahMi 

61.M 

JM.9 

sot 

I Oil 

1.41 9 M°F 

1.0179 

0600 9 U'F 

o.i« 

ISS O 7*0 mm Mg 

74.1 9 M°F 

I4S19 

200 

FIG. 15-28 

Physical properties ot gas treating chemicals 

Trielheeol-

—eiweeelomiee 

(HOCJH.I.HH 

105 14 

SUI 

M.4 

474.7 

1171 

IMS 

101 9 M*F 

I09I9(10/20'C) 

OUO 

0117 

2M 9 7) mm Hf 

3W<3 6I°F 

(at 90% 

ot. solution) 

l.477o 

210 

(HOClHallK 

I4II9 

UO (Oecempoies) 

77 3 

ISS 

957.7 

1.124 

? 37 9 U°F 

III M 

0.70 

210 9 7U mn 

-400 

-430 

1011 9 6I'F 

lit 95% 

wl. solution) 

I.4IS2 

16S 

oiGircouuuNce 

H(0C)H.),«H> 

I0S.I4 

430 

• s 

S47.ll 

754 4 

l.0SI@U°F 

1.12 9 U'F 

1.0571 

0.571 

0.121 

119.14 9 7Um 

-474 

-ISO 

40 a U'F 

1.4591 

2 U 

Witeprepaaelemiiie 

(H0C>H.),MH 

113.19 

4797 

107 4 

5441 

22» 

I W 9 M'C 

0.9J9 (3 45'C >20'C 

0.49 9 10'C 

"1*9 US 9 (?|mm Hg 

170 9 M°F 

I tl 9 I U'F 

16 9 llf l 

14542 9 I U'F. 

2SS 

SCUXOl* 

Oimolhylother 

of polytthylmo 

llycol. 

2M 

SU 

-20 

1.0)1 9 77'F 

I.U 9 77V 

0.49 9 41 °F 

0.lla?7»F 

-190 9 77'F 

-IU<a77V 

SJ a 77'F 

104 

Propylene 

Corbonow 

C.H.COi 

102.09 

447 

- 5 4 4 

I.20S7 

1.203 

0.33! 

0.12 @ (?)'f 

201 (3 760 mm Mg 

1.67 9 100'F 

19.4 ci 9 —40*F 

1.79 Cl <§> I00'f 

0.127 ci <3 2I0"F 

1.4209 

270 

c 

( 

Formula 

•oiling point 9 740 mm Hg, °F. 

frtuing point, °f. 

(ritkol conifonti 

Frtlluro, ptio . .. 

Ttrnporaturo, 'F 

Oomily 9 20'C, gm/ec 

Wtight, lb/gal 

Specific gravity 20'C./20'C 

Spocific htaf rp U'F. ITU/lb/"F 

Thormot conductivity. 

ITU/hr. FT.

 -F./ft. (SlU'F 

lattnl hoot sf vaporisation, ITU/lb 

IHiyWuei 

Glycol 

HOC,HaOH 

12.07 

317.0 

1.6 

1114 2 

701.4 

9 21 9 6I°F 

I.IISS 

0552 

0.144 (3 60°F 

144 (3 7*0 mm Hg 

Glycol 

H0(CjMa0)iH 

106.12 

472.6 

17.6 

476.0 

766.4 

9)1 9 Wl 

I IIM 

0 535 

0144 9 U'F 

2)2 9 7*0 mm Hg 

TriethyW 

Otycol 

HO(CiH.0|iM 

ISO 17 

$49) 

19.0 

4792 

121 

•34 9 U'F 

1.1254 

0.492 

0 1)9 9 U'F 

179 9 760 mmHg 

Glycol 

H0(C,Ha0|aH 

19412 

621.1 

201 

376) 

I I* 

9 34 9 U'F 

U4I 

0510 

141 9 760 mm Hg 

SulfoUno® 

C.H.S0) 

12017 

S4S 

11.7 

767) 

101)1 

10.62) 9 )0't /WC 

1.261 

0)6 9 30°C 

0.114 9 I00°F 

22S7 9 2I2'F 

MOHMMI 

CHIOH 

12.04 

1411 

-143 1 

11539 

464 

0.7917 

0.590 (S'-I0"C | 

0.124 

474 9 7*0 mm Hg 

c 

c 

Viicoiity. cp .. . 

lofracHvs indei, >* U'F. 

Floih point, COC. *F. 

20.9 9 U'F 

14314 

2*0 

35 7 9 U'F 

1.4472 

290 

47.9 9 U'F 

14559 

110 

«0 0 9 U'F 

1.4595 

ISO 

10 3 9 U'F 

4.1 9 I22'f 

J.S 9 JU'f 

1.4 9101'F 

0.47 9 192"F 

1.411 9 UT 

ISO 

O.t 9 U'F 

1.1216 

SI 

e 

Rev. 1976 15-2 4 
P.16 

, ,- -,r, 

( a 

FIG. 15-29 

Vapor pressures of gas treating chemicals 

1,000 

c 

100 

c 

o a 

i10 

c 

A r V 

"V *y 

- cf. 

/ / 

rfV 

fJ 

T / / ii 

is/-

/ / 

1 f 

1 

/ 1 I H 

f 

i / 

<S 

/ 

M 

f 

f

) 

/ 3 

f-i 

c 

,8 

<%J— 

S 

A 

V i 

s? 

cfi 

-

_.< 

h *' M 

1 

I t , / 

/ J " 

w 

•>?'# sit 

Q M sA g A 

V 

o 

-J/f—J /A 7** 

//<$n i 

'f/ w~ 

if f J 

OJ, 

«?/ 

.?/ 

<?/ 

si 

7 

// 

* 

i 

\f 

' / 

/ / 

f'l 

1 / •ft 

/ OV 

/ ' i i 

j 

Jeff erson 

m^M 

i 

f / 

/ 

r 

Chen 

/ 

J 

/ 

/ 

' 

/ 

/ 

r 

f j 1 

neal :o. 

100 150 200 

C 

250 

Temperature, °f. 

300 350 400 450 500 

15-25 Rev. 1976 
P.17 

a 

E 

a 

5> 

FIG. 15-30 

Freezing points of aqueous amine solutions 

90 

80 

70 

60 

50 

40 

30 

20 

10 

0 

- 1 0 

- 2 0 

- 3 0 

- 4 0 

- 5 0 

FIG. 15-31 

Specific gravity of aqueous glycol solutions 

eff 

1 ers 

1 _ 

on Ch 

1__ 

em 

1— 

ica 

1 

C 

1 _ 

\ v\ 

• 

o. -

[\\n 

\\ 

1 g 

> 

i i 

-$ 

. Q 

r

/L 

/r^i 

111 

iri 

a 

U'Af. 

nu\ ; 

In 1 mil ' 

III 

W i 

P\§ *"*(# 

jM 

1 

C 

J / 

/ 

/ 

/ 

/ 

| 

/ 

Sl i?r 

20 40 60 80 100 

FIG. 15-32 

Specific gravity of aqueous amine solutions 

1.14 

1.12 

1.10 

1.08 

1.06 

1.04 

1.02 

1.00 

1 

1 

1 

A 

- -

' 

n 

\ 

TEA ^ 

!

 D§A ^ 

1 -Jr 

n 

~r 

1 

1 

DIPA 

0* 

^^ 

(•fs 

1 

"1"" 

| 

1 I 

'"J err ers 

i 

' 

PG> 

j 

1 

j 

I 1 

i 

i ^ 

1 1 

k 

1 

on cr. em ca Cc 

-

. 

o o 

o 

0> 

1.14 

1.12 

1.10 

1.08 

1.06 

1.04 

1.02 

1.00 

tp> 

*/& 

f< t 

r 

I I I 

'si 

^ ^ 

11.11 . 

20 40 60 80 100 

Glycol, wt% 

20 40 60 

Amine, wt% 

80 100 

MOLECULAR SIEVE 

Molecular-sieve processes are being used for removal 

of sulfur compounds from natural gas streams (see Fig. 

15-33). The grade of molecular sieves to be used is selected 

on the basis of the operating conditions. Hydrogen sulfide 

can be removed to meet 0.25 grains/100 scf specification 

with selective removal of H.S. The sieve bed can be designed to dehydrate and sweeten simultaneously. 

In general, the concentrations of acid gas are such that 

cycle times are in the order of 2 hours. To operate properly 

the sieves must be held at a temperature close to 600°F. 

for a long enough period of time to remove all adsorbed 

materials. Exact arrangement of the regeneration cycle 

depends upon process conditions. 

The problem of COS formation during processing according to the reaction of 

H,S + C0,=.C0S + H.O 

has been extensively studied. The molecular sieve can catalyze the reaction; however, new molecular sieve products 

have been developed to retard significantly COS formation. The central zone in the regeneration cycle is most 

favorable to COS formation. 

Operation of adsorbent plants is quite simple, but design is complex. Manufacturers must be consulted for 

potential applications. Refer to gas drying and liquid 

sweetening sections for additional details. References: 12, 

JH, JK), .5.J, 57. 5M, 5!l. Cll and til. 

c 

e 

Rev. 1976 15 - 26 
P.18 

c 

FIG. 15-33 

Typical COj-rkh natural gas treater 

90 MMscfd 

200 MMscfd 

6% CO, 

10-100 gr HiS 

7 Ib HjO'MMscf 

110 MMscfd 

J^\ 

Molecular 

sieve 

Adsorbing 

Molecular 

sieve 

regenerating 

w 

Courtesy Davison Chemical 

10 MMscfd 

Liquid sweetening 

and 

dehydration process 

100 MMscfd 

67. CO, 

0.25 gr H,S 

Bone dry 

100 MMscfd 

Nil CO, 

0.25 gr H,S 

7 Ib H,0/MMscf 

200 MMscfd 

3% CO, 

0.25 gr H,S 

3 Ib H,0 MMscf 

c 

c 

C 

Equilibrium data for 

amine-sour gas systems 

One of the peculiarities of amine treating systems is 

the interactive effects of one acid gas constituent with 

amine upon the equilibrium partial pressure of the other 

constituent. The most commonly encountered sour gas 

constituents are H,S and CO,. The capacity of a given 

amine for either one of the acid gas constituents alone is 

much greater than when the two occur together. 

Jones et al (63) have presented data to confirm the interactive effect of H,S and CO, in monoethanolamine solutions. More recently, Lee et al (64) have presented similar 

data for diethanolamine solutions. These data provide the 

basis for predicting the equilibrium concentration of MEA 

and DEA solutions when in contact with sour gases containing both H,S and CO,. 

McCoy and Maddox have plotted the data of Jones et 

al for MEA in nomograph form for ease of use. The McCoy 

and Maddox charts are shown in Figs. 15-34 to 15-41. The 

Lee, et al, charts for DEA are shown in Figs. 15-42 to 15-49. 

Neither the data on MEA nor the data on DEA cover 

all possible temperatures that might be encountered in an 

amine treating unit. Consequently, some technique is 

needed for interpolating acid-gas partial pressures and 

equilibrium concentrations to temperatures intermediate 

between those shown in the equilibrium data charts. Approximate values at other temperatures may be calculated 

based on the assumption that vapor pressure is a function 

of temperature only and follows the rule: 

ln P<,~- (D 

Where: 

P" - pure component vapor pressure 

T = absolute temperature 

For reacting systems, where the amount of one component reacted affects the amount of another component that 

can react, this is only an approximation. However, it provides the best rule available for interpolation between 

temperatures on either the McCoy and Maddox or the Lee 

et al charts. • 

Applying Equation 1 to the interpolation between two 

temperatures to determine partial pressure for a given 

acid gas loading in the amine yields: 

ln P° = ln P° + 

Where: 

yr, T,; 

yr, T,/ 

In -^ (2) 

1 = the lower temperature 

2 = the higher temperature 

i = a temperature intermediate between 1 and 2 at 

which the partial pressure of the particular acid 

gas over the amine solution is desired 

Example: 

A 2.0 NDEA solution contains 0.2 mole H S/DEA and 

0.4 mole CO/DEA. 

a. What will be the equilibrium partial pressures of 

CO, and H.S over this solution at 50°C? 

From Figure 15-44 

ppCO =4.2 psia 

From Figure 15-42 

ppH.S = 2.8 psia. 

b. What will be the equilibrium partial pressures of 

H.S and CO, over the solution at 100°C? 

From Figure 15-45 

ppCO, = 47 psia 

From Figure 15-43 

ppH S = 15 psia. 

15 - 27 Rev. 1976 
P.19 

c. What will be the equilibrium partial pressures of H.S 

and CO, over the DEA solution at t = 60°C? 

Interpolation for temperatures intermediate between 50° and 100°C makes use of the fact that the 

log of the vapor pressure of a material varies inversely with absolute temperature. Assuming that 

the vapor pressure of an acid gas constituent over 

the DEA solution plots as a straight line when log 

vapor pressure is plotted against the reciprocal of 

the absolute temperature gives rise to the following 

interpolation equation: 

l n

 PPT = 1" PPTM + 

For CO, at 60°C 

ln pp C Q = ln 42 + 

= 1.9761 

PPCQ = 7.2 psia 

For H,S at 60° C 

ln ppH.S = ln 2.8 + 

= 1.4056 

PPH g = 4.1 psia 

1 

323 

1 

1 

t + 273 

1 

323 373 

1 

323 

1 

1 

333 

1 

323 373 

1 

323 

1 

1 

333 

1 

323 373 

. / PP \ 

'"(-PPT; 

In 

47 

4.2 

ln 

15 

2.8 

3. Ballard, D., Hydrocarbon Processing, 44 (No. 4), 131-

138 (1965). c, d. Silica gel. Alumina. 

4. Beck, J. E., "Operational Problems in CO Removal at 

the Mitchell Plant," Proc. Gas Cond. Conf. 1960,129-

40. MEA, DEA. 

5. Benson, H. E.; Field, J. H.; Haynes, W. P.: "Improved 

Process for CO, Absorption Uses Hot Carbonate Solutions," Chemical Engineering Progress, 52, pp 433-8 

(Oct., 1956). HKC. 

6. Benson. H. E. and Parrish, R. W., "Hi-Pure Process 

Removes CO/H,S", Hyd. Proc. p. 81. (April 1974). 

Benfield HKC. 

7. Boston, F. C, "Adsorption Process Uses Multistaging", 

Hyd. Proc. and Pet. Ref. 43, No. 8, pp. 141-145 (August 1964). a. Solid bed. 

8. Buckingham, P. A., "Fluor Solvent Process Demonstrated in EPNG Plant", NGAA Regional Meeting, 

Odessa, Texas, May 1961. Fluor. 

9. Buckingham, P. A., Hydrocarbon Processing and 

Petroleum Refiner, 43, No. 4, p 113 (April 1964). Fluor. 

10. Comeaux, R. V., Hydrocarbon Processing, 41 (No. 5), 

141-143 (1962). d. MEA corrosion. 

11. Connors, J. S., Oil Gas J., 56 (No. 9), 100-110 (1958). 

a, c, d. MEA. 

12. Conviser, S. A., "Molecular Sieves Used to Remove 

Mercaptans from Natural Gas", Oil and Gas Journal, 

pp. 130-133 (Dec. 6,1965). c. Mol sieve 

13. Cook, T. P. and R. N. Tennyson, Chem. Eng. Progr., 

65, No. 11, 61-64, (November 1969). Fluor. 

14. Dailey, L. W., Oil Gas J., p. 136 (May 4, 1970). a, c. 

• SNPA-DEA. 

15. Dingman, J. C, Hydrocarbon Processing, 42 (No. 9), 

189-191 (1963). b. MEA. 

16. Dingman, J. C; Allen, D. L.; Moore, T. F., Hydrocarbon Processing, 45, (No. 9) 285 (1966). a, c, d. MEA 

corrosion. 

c 

BIBLIOGRAPHY-GAS TREATING 

Subject is indicated by code following reference. 

a—design 

b—physical and chemical properties 

c —operation 

d—corrosion and materials of construction 

HKC—-hot potassium carbonate 

Fluor—propylene carbonate 

Selexol—polyglycol ether 

General—considers many processes 

1. Allensandria, A. B. and Jaggard, Norton, "Refiners 

Report New Cases of Stainless Steel Failure," Petrol. 

Refiner 39, No. 5, 151-156 (1960). c, d. 

2. Allied Chemical Corp., Solvay Process Division: "The 

Hot Potassium Carbonate Process for Acid Gas Absorption," Technical Service Report No. 6.61 (March 

30, 1961). HKC. 

17. Dingman, J. C. and Moore, T. F., "Compare DGA and 

MEA Sweetening Methods", Hydrocarbon Proc. 47, 

No. 7,138-140 (1960). MEA, DGA. 

18. Donnelly, R. F. and D. R. Henderson, Can. NGPA 

Third Quarterly Meeting, Calgary, Sept. 13, 1974. 

DEA. 

19. Dow Chemical Company, The, "Propylene Carbonate 

for Natural Gas Purification", Form No. 164-86, September 1960. Fluor. 

20. Dunn, C. L.; Freitas, E. R.; Goodenbour, J. W.; Henderson, H. T., and Papadopoulos, M. N., "New Pilot 

Plant Data on Sulfinol Process", Hydrocarbon Processing and Petroleum Refiner, Vol. 43, No. 3, pp 150-

154, March 1964. Sulfinol. 

21. Dunn, C. L., Freitas, E. R., Hill, E.'S., and Sheeler. 

J. E. R., "Sulfinol Process", Oil and Gas Journal, 

March, 1965, pp 89-92. Sulfinol. 

L 

<r 

Rev. 1976 15-2 8 
P.20 

c 

c 

C 

22. Ebdon, J. Fred, "Two El Paso Natural Plants", Gas 

45, No. 6, 52-58 (1969). DGA. 

23. Eickmeyer, A. G., "Catalytic Removal of CO.", Chemical Engineering Progress, 58, #4, pp 89-91 (April, 

1962). HKC, Catacarb. 

24. Ellis, G. C; Leachman, G. S.; Formaina, R. E.; Hazelton, R. F.; Smith, W. C, Proceedings Gas Conditioning Conference 1963, B-l-B-19, The University of 

Oklahoma, b. MEA. 

25. Fitzgerald, K. J., and Richardson, J. A., Hydrocarbon 

Processing, 45 (No. 7), 125 (1966). a, b, c. MEA. 

26. Fitzgerald, K. J. and Richardson, J. A., Oil Gas J, 64 

(No. 43), 110 (1966). a, b, c. MEA. 

27. Goar, B. G., "Sulfinol Process Has Several Key Advantages", Oil and Gas Journal, June 30, 1969, p 117-

120. Sulfinol. 

28. Greene, P. A. and Kutsher, G. S., "Carbon Dioxide 

Removal and Recovery—New Polyglycol Ether Processes", Proceedings of the Gas Conditioning Conference, University of Oklahoma Extension Division, 

Norman, Oklahoma (1965). Selexol. 

29. Gustafson, K. J., and Healey, Jr., M. J., "Removal of 

Hydrogen Sulfide from Natural Gas/Carbon Dioxide 

Mixtures with Molecular Sieves", Presented at the 

18th Annual Gas Conditioning Conference, Norman, 

Oklahoma (April 3, 1968). a, b, c. Mol sieve. 

30. Harris, T. B., and Victory, R. M., "Recent Molecular 

Sieve Products and Process Developments in Natural 

Gas Processing", Presented at the 20th Annual Gas 

Conditioning Conference, University of Oklahoma, 

Norman, Oklahoma (April 1, 1970). a, b, c. Mol sieve . 

31. Hegwer, A. M., and Harris, R. A., "Selexol Solves High 

H,S/CO, Problem", Hydrocarbon Processing, Volume 

49, No. 4, April 1970, p 103. Selexol. 

32. Holder, Howard L., "Diglycolamine — A Promising 

New Acid-Gas Remover", Oil Gas J., 64, No. 18, 83-

86 (1966). DGA. 

33. Jean-Pierre, R., and Guillo, R., Canadian Patent, 651, 

379. a, c. SNPA, DEA. 

34. Jones, J. H.; Froning, H. R.; and Claytor, E. E., Jr., J. 

Chem. Eng. Data, 4 (No. 1), 85-92 (1959). b. MEA. 

35. Katz, Donald L., "Handbook of Natural Gas Engineering", McGraw-Hill Book Co., Inc. (1959). a, b, c, d. 

General. 

36. Kohl, Arthur L. and Reisenfeld, Fred C, "Gas Purification", McGraw-Hill Book Co., Inc. (1960) (Chemical 

Engineering Series). General. 

37. Kohl, A. L. and P. A. Buckingham, Petroleum Refiner, 

39, No. 5, p 193 (May 1960). Fluor. 

38. Kohl, A. L. and F. E. Miller, U. S. Patent 2926751, 

March 1, 1960. Fluor. 

39. Kutsher, G. S. and Smith, G. A., "CO, Removal and 

Recovery—Pegasol-A Solvent Process", Proceedings 

of the Gas Conditioning Conference, University of 

Oklahoma Extension Division, Norman, Oklahoma 

(1966). Selexol. 

40. Kutsher, G. S.; Smith, G. A.; Greene, P. A., "Sour Gas 

Scrubbing—Allied Chemical Solvent Process", Proceedings of the Forty-Sixth NGPA Annual Convention, March 14-16, 1967, Houston, Texas. Selexol. 

41. Lang, Frances L. and Mason, J. F., Jr., "Corrosion in 

Amine Gas Treating Solutions", Corrosion 14, No. 2, 

105+ -108+ (1958). d. Amine. 

42. Leachman, George S. and Ellis, Garland C, "TroubleFree Ethanolamine Gas-Treating Plants", Oil Gas J. 

50, No. 30, 84-88 (1961). Amine. 

43. Lee, J. I., Otto, F. D., and Mather, A. E., "The Solubility of Mixtures of CO, and HS in DEA Solutions", 

Canadian J. Chem. Eng., 52, p. 125 (1974). b. DEA. 

44. Maddox, R. N., "Gas and Liquid Sweetening", J. M. 

Campbell Petroleum Series, a, b. General. 

45. Maddox, R. N., and Burns, M. D., Oil Gas J., 65 (No. 

33), 131-133 (1967): Ibid., 65 (No. 34), 83-85 (1967): 

Ibid., 65 (No. 35), 116-117 (1967). a, c. General. 

46. Maddox, R. N., and Burns, M. D., Oil Gas J., 65 (No. 

38), 112-121 (1967): Ibid, 65 (No. 40), 110-111 (1967). 

a, c. General. 

47. Maddox, R. N., Proceedings 49th NGPA Annual Convention, 93-103, March 17-19, 1970, Denver, a, c. 

General. 

48. Mason, J. R., and Griffith, T. E., "MEA-DGA Switch 

Saves Steam", Oil Gas J. 67, No. 23,67-69 (1969). DGA. 

49. Moore, Kenneth L., "Corrosion Problems in a Refinery 

Diethanolamine System", Corrosion 16, No. 10, 503+ -

506+ (1960). c, d. Amine. 

50. Mottley, J. R., and Fincher, D. R., "Inhibition of 

Monoethanol Amine Solutions", Materials Protection 

2, No. 8, 26-30 (1963). c, d. HKC. 

51. Muhlbauer, H. G., and Monaghan, P. R., Oil Gas J., 

55 (No. 52), 139-145 (1957). b. MEA, Amine equil. 

52. Perry, C. R., Proceedings, the Gas Conditioning Conference 1967, C-l-C-15, The University of Oklahoma, 

a, c. MEA. 

53. Polderman, L. D., and Steele, A. B., Oil Gas J., 54 

(No. 65), 206-214 (1956). b, c. MEA. 

54. Schoofs, J. R., "Molecular Sieves for Natural Gas 

Sweetening", Address at Western Gas Processors and 

Oil Refiners Association, Anaheim, California (October 6, 1966). b, c. Mol sieve. 

55. Sweeney, John L., and Brooks, Gerald E., "Selexol at 

Fort Stockton", Proceedings of the Gas Conditioning 

Conference, University of Oklahoma Extension Division, Norman, Oklahoma (1970). Selexol. 

56. Taylor, D. K., "How to Desulfurize Natural Gas", Oil 

and Gas Journal, Nov.-Dec. (1956). Iron Sponge. 

57. Thomas, T. L., and Clark, E. L., "Molecular Sieves 

Reduce Economic Operational Problems", Oil and Gas 

Journal, 65, No. 12, pp 112-115 (1967). c. Mol sieve. 

58. Thomas, T. L., and Clark, E. L., "Molecular Sieve 

Process Developments and Their Impact on the Natural Gas Industry", Reprinted from Proceedings for 

Forty-Sixth Annual Convention of the Natural Gas 

Processors Association, c. Mol sieve. 

59. Thomas, T. L., "Molecular Sieves in Petroleum and 

15-29 Rev. 1976 
P.21 

Natural Gas Processing", Adapted from a paper presented at the Sixth World Petroleum Congress, Frankfurt/Main, c. Mol sieve. 

60. Turnock, P. H. and Gustafson, K. J., "Advances in 

Molecular Sieve Technology for Natural Gas Sweetening", Proc. Gas Cond. Conf., Univ. of Otolahoma, 1972. 

b. Mol sieve. 

61. Young, Jr., Philip L., "Molecular Sieve Adsorption in 

LNG Plant Pre-Purification", Presented at the Distribution Conference American Gas Association, Houston, Texas (May 6-9, 1968). a, c, d. Mol sieve. 

62. Zapffe, Fred, "Sour Gas Treating: A Design Problem", 

Proc. Gas Cond. Conf., Univ. of Oklahoma, 1969. a. 

General. 

63. Jones, J. H.; Froning, H. R.; and Claytor, E. E., Jr.: 

Chem. Eng. Data, 4, p. 85 (Jan. 1959). 

64. Lee, J. I.: Otto, F. D.; and Mather, A. E.: Gas Processing/Canada, p. 26, March-April, 1973. 

65. Lee, J. I.; Otto, F. D.; and .Mather, A. E.: "Solubility of 

CO, in Aqueous DEA Solutions at High Pressures", J. 

Chem. Eng. Data 17, 465 (1972). b. 

66. Lee, J. I.; Otto. F. D.; and Mather, A. E.: "CO. and H S 

in 2.5N MEA Solution", Can. J. Chem. Eng. 52, 803 

(1974). b. 

67. Isaacs, E. E.; Otto, F. D.; and Mather, A. E.: "Solubility of Acid Gases in Aqueous DIPA Solutions", 68th 

Annual AIChE meeting, Los Angeles, Nov. 16-20,1974. 

b. HmS, CO,. 

c 

Fig. 15-34 

Hydrogen Sulfide and Carbon Dioxide Partial Pressures Over 15 wt% 

Monoethanolamine Solutions Containing Hydrogen Sulfide 

and Carbon Dioxide at 40°C. 

M 

X 

E 

E 

</» 

z 

*o 

2 

3 

tfl 

tfl 

2 

o. 

To •s 

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Rev. 1976 15-30 Rev. 1977 
9Z6I

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T — T 

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N

 Hydrogen Sulfide Partial Pressure Over 15 wt% Monoethanolamine 

Solutions Containing Carbon Dioxide and Hydrogen Sulfide at 60°C 

c 

E 

E 

e 

I 

ioo -

30 -

10 : 

2 

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t. N. ModdoR. 

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300 

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Fig. 1540 

Hydrogen Sulfide Partial Pressure Over 15 wt% Monoethanolamine 

Solutions Containing Carbon Dioxide and Hydrogen Sulfide at 100'C 

Rev. 1977 15-33 Rev. 1976 
X 

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tfl 

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1000 c 

300 -

100 -

30 -

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Fig. 15-41 

- Hydrogen Sulfide Partial Pressure Over 15 wt% Monoethanolamine 

~ Solutions Containing Carbon Dioxide and Hydrogen Sulfide at 120°C. 

_ R. N. Maddox, 

Oklahoma Slate University 

_ 

-

— 

— 

_ 

-

a— 

~ 

~ 

120°C. 

4 "i 

/ r 3 f iy ,i t 

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0.6 0.8 1.0 

RATIO IN LIQUID.. H 2S/0E A C 

15-34 Rev. 1981 
,r

 r""' 

c 

0* 06 08 1.0 

MOLE RATIO IN LIQUID. H2S/DEA 

c 

c 

c 

MOLE 

16 0 8 10 

flO IN LIQUID. C0 2 /DEA 

15-35 Rev. 1981 
0 4 0 6 0.8 10 

MOLE RATIO IN LIQUID. C0 2 /DE A 

0 6 0.8 1.0 

MOLE RATIO IN LIQUID. H2S/DEA 

c 

c 

c 

e 

Rev. 1981 15-36 
c 

lOOOr 

Rg. 15-47 

Partial Pressure of Hydrogen Sulfide over 

Diethanolamine Solutions Containing 

Hydrogen Sulfide and Carbon Dioxide 

M 11 1111 1 .• 11 

0.6 0 8 

MOLE RATIO IN LlQUIO. H2S/DEA 

c 

c 

v. 0.4 0.6 0 8 1.0 

MOLE RATIO IN LIQUID. C02 /DEA 

15-37 Rev. 1981 
Partial Pressure of Carbon Dioxide over i 

Diethanolamine Solutions Containing 

Carbon Dioxide and Hydrogen Sulfide 

I I ! I I I I I I I I I I I I I I I I I I I I I I I I I I I I I I I I 

0.2 0.4 0.6 0 8 1 0 

MOLE RATIO IN LIQUID. C02/DEA 

1.2 

c 

C 

Gasoline and LP-gas treating 

Most gasoline and LP-gas streams contain sulfur 

in various forms. Those that are especially objectionable are hydrogen sulfide, mercaptans and elemental sulfur. Gasoline containing hydrogen sulfide 

has an objectionable odor and is corrosive. Mercaptans give a very objectionable odor to the gasoline 

and elemental sulfur makes the gasoline very corrosive. A product containing the objectionable elements can be treated to remove the hydrogen sulfide 

and elemental sulfur and to either remove the mercaptans or convert them to less objectionable compounds. 

Hydrogen sulfide removal 

Hydrogen sulfide can be removed from LP-gas 

and gasoline by liquid-liquid contacting processes 

using a caustic solution or ethanolamines. The 

ethanolamine processes are described in the gastreating section earlier in this chapter. In this application, precautions must be taken to minimize the 

introduction of hydrocarbons into the regeneration 

section of the amine unit. 

Sulfur removal 

Elemental sulfur is removed from the gasoline by 

contacting it with a polysulfide wash solution. A 

sulfide solution is made up by using the following 

amounts of chemicals per 1,000 gallons of water: 

1,000 lb of caustic (NaOH), 800 lb of commercial 

Na-jS and 20 lb of sulfur. The flake sodium sulfide 

(Na^S) is melted in a vat by use of a steam lance. 

Add"the sulfur to the melted Na2S. The sulfur must 

be completely dissolved in the liquid sulfide, and 

c 

then this mixture is added to the 10% (1,000 gal 

water and 1,000 lb NaOH) caustic solution. (Protective clothing and goggles should be worn when 

handling these chemicals). 

Mercaptan treating 

Mercaptans are converted to disulfides or removed from gasoline by several methods to provide 

a satisfactory product. The method or combination of 

methods that can be used depends on the mercaptan 

content of the product to be treated and the specification that must be met. 

Regenerative caustic process 

The regenerative caustic method of removing 

methyl and ethyl mercaptans from LP-gas or butane 

employs countercurrent contacting of the liquids 

with 10% sodium hydroxide solution in a packed 

column. The caustic is regenerated in a stripping 

column by the addition of open steam or by steam 

internally generated by a column heating element. 

A condenser may be used to condense water vapors 

that go overhead with the liberated mercaptan 

vapors. The condensate is returned to the stripping 

column to maintain caustic concentration. An LPgas stream containing 0.03% mercaptan sulfur can 

be treated doctor sweet with volume ratios of hydrocarbon to caustic as low as 33 to 1. Approximately 

six pounds of process steam per gallon of caustic are 

required for regeneration. The following ratios may be 

more desirable because of plant economics: 

Lb steam/gal NaOH 4 5 6 7 8 10 12 14.5 17 £ZL 

Gal hydrocarbon/ \__ 

gal NaOH 15 25 33 39 43 50 55 60 62 

C 

Rev. 1981 

15-38 
P.30 

c 

c 

c 

c 

Highly corrosive conditions that can exist in a 

unit for this process are minimized by material selection or by careful attention to operating conditions. 

The latter is done by minimizing the contamination 

of the regenerative caustic system with sodium sulfide and by operating the reactifying column at low 

pressure and temperature. The maximum desired 

steam temperature to the kettle element is 260°F. 

The sulfide contamination is controlled by caustic 

washing the LP-gas ahead of the mercaptan absorber and by blowdown of the circulating solution. 

A maximum NajS content of 0.2% is suggested. 

The diameter of the mercaptan absorber, along 

with other liquid-liquid treaters, is based upon the 

hydrocarbon throughput. A typical design uses flow 

rates of 10-15 gpm/sq ft of cross-sectional area for 

a column packed with 1% in. raschig rings. Other 

tower packings may be used and, if so, the above 

throughput would be changed.. The packed height 

is normally about 30 ft, broken up into two or three 

packed beds to optimize dispersion of the hydrocarbon phase An additional 8-10 ft of column height 

is required for distribution and settling areas. 

Merox process 

A variant of. the regenerative caustic process has been 

developed by UOP Process Division. Sour product is treated 

with caustic containing Merox catalyst to extract the mercaptans. The Merox solution is regenerated by mixing with 

air and oxidizing. Disulfides and excess air are removed 

overhead. Existing extraction equipment with steam regeneration can be adapted to the Merox process. 

Caustic wash 

A caustic wash may be used to remove hydrogen 

sulfide or methyl and ethyl mercaptans from gasoline. Equal volumes of gasoline and recirculated 

caustic are mixed by means of a pump and discharged 

to a caustic settling tank where the liquids separate. 

The size of the tank required is based on the retention time necessary for complete separation of the 

gasoline-caustic emulsion. The retention time varies 

from one and a half hours for heavy gasolines to 

three-fourths of an hour for propane. Since the 

separating tank is one-third full of caustic solution, 

the gasoline retention time is based on two-thirds of 

the tank volume. 

SOUD POTASSIUM HYDROXIDE TREATMENT 

Solid potassium hydroxide may be used to remove small 

amounts of sulfur compounds from liquid propane and LPgas streams. 

The reaction for this process is: 

H,S + 2KOH • K.S + 2H.O 

Theoretical consumption of KOH is 3.29 lb/lb of 

H2S. Actual consumption has been reported to be about 

4.45 lb/lb, representing an efficiency of 74%. 

Although water is formed as a reaction product, 

it is necessary to add small quantities of water to 

the bed in order for the reaction to progress 

properly. 

This treating method is normally applied to dry 

propane streams containing small amounts of hydrogen sulfide and mercaptan contaminant. The product 

to the treater meets cobalt bromide specification, as 

does the treated stream. The water added plus that 

formed is apparently held by the chemical and 

settles to the bottom of the vessel as a solution. This 

aqueous solution should be drained from the vessel 

daily. 

The required water addition is about one quart 

for 20,000 gal, but may vary depending upon the 

H.S content. A mixture of methanol and water can 

be added in winter to avoid freezing. An effective 

mixing device must be installed downstream of the 

point of water addition to insure that it is dissolved 

in the propane. Otherwise, the water is ineffective. 

Poor mixing is evidenced by a low solids content of 

the drained solution. 

Construction of the solid caustic treater is very 

simple—requiring only a vessel with the necessary 

internal supports for the caustic, and the various 

nozzles for the fluids. Low propane velocity (1.5 

ft/min superficial velocity) is used in order that 

the reaction products can settle and be drained off. 

At higher velocities it becomes necessary to stop 

the flow for a few minutes to permit settling. No 

regeneration is required. No utilities are required. 

Walnut-sized KOH is recommended. A screen 

should be placed over the outlet line to prevent 

carry-over of caustic in case of flashing of propane 

on startup. 

Methanol or CO, in the liquid feed stream will cause 

additional caustic consumption. CO,, reacts with KOH and 

consumes approximately the same amount of KOH as does 

H.S. If in sufficient quantities, the combined H.S, CO.., 

and methanol can dictate considering another process because of shorter bed life and cost of KOH replacement. 

Molecular-sieves treating process 

Propane, butane and mixed LP-gas streams can be 

treated to meet GPA specifications for dryness, 

volatile sulfur and corrosion with molecular sieves 

in a one step process. Most units treat propane or 

butane streams but the process principles are the same 

for mixed LP-gas streams and may be extended to 

selected natural gasoline streams. 

PROCESS DESCRIPTION 

A typical process flowsheet for a two bed system 

appears in Fig. 15-33. While one bed is on adsorption (treating) the second bed is regenerated. A 

typical process cycle consists of: 

1- Treating—Liquid feed passes (either upflow 

or downflow) through one bed of type 13X molecular 

sieve. Sufficient feed pressure is used to maintain 

liquid phase flow. Propane treaters usually operate 

at 200 to 250 psig with butane treaters at 100 to 150 

psig. The molecular sieve selectively removes first 

water and then sulfur compounds. The amount of molecular sieve required depends primarily on quantity and 

type of sulfur compounds present in the feed stream. 

15-39 Rev. 1977. 1979. 1981 
P.31 

For streams containing less than 15 grains/100 scf 

sulfur compounds the treater is about the same size 

as the alumina dehydrator required downstream 

from a caustic scrubber. 

2. Drain—After treating (to 24 hours) the beds 

are switched. Liquid is drained downflow from the 

bed taken off treating using a gas pad with 10 to 20 

psi driving force. Drained liquid may be sent to sweet 

product storage if downflow treating is used or recycled 

to feed if upflow treating is used. 

3. Depressurize—Since heating gas is usually at 

a lower pressure than the treating step, it is necessary to depressurize the vessel to the heating gas 

pressure. This is done the same time the heating 

step is started and will therefore usually be in the 

same flow direction as heating. Bed lifting may 

occcur if upflow depressurization is too rapid. Downflow depressurization is preferred at rates of 50 

psi/minute or less. During the depressure step liquid 

held up on the bed flashes. The vessel and exit lines 

may become noticeably cold. 

4. Heat—A sweet gas at moderate pressures is 

used to regenerate the molecular sieve in the opposite 

flow direction as treating. Temperature and flow rate 

depend on the quantities of water and sulfur compounds to be removed, the hydrocarbon stream treated 

and the quality of the regeneration gas. Typical operating ranges are: pressure 50 to 200 psig, temperature 

550° to 650° F., flow rate 40 to 80 scf/lb molecular sieve. 

Changes in any of the regeneration variables can significantly affect the performance of the treater. Effluent regeneration gas is cooled and returned to a gas 

system. 

5. Repressurize—This step is required after heating if liquid product is used for cooling, but would 

follow the cooling step if ambient temperature purge 

gas is used. Repressurizing with product liquid 

actually occurs as the first part of the cooling step. 

A slipstream of product liquid to be used for cooling 

is slowly bled into the hot molecular sieve bed and as 

the liquid flashes, the vessel repressures to the product pressure. A repressure rate of less than 50 psi/ 

minute is preferred. 

6. Cool and fill—Product liquid from the onstream treater is normally used for cooling in the 

same flow direction as treating. Cooling is accomplished as the liquid flashes on contact with the hot 

vessel and molecular sieve, and continues as a liquid 

level begins to build in the bottom of the vessel. Effluent cooling vapor is condensed in a cooler before 

blending into the product stream. When the treating 

direction is downflow the entire vessel must be 

filled with liquid prior to switching to drive all gas 

from void spaces to insure no channeling from two 

phase flow. Typical liquid cooling rates are 3 to 6 

gpm, 1000 lb of molecular sieve. 

Ambient temperature purge gas can also be used 

for cooling after which product liquid is bled into 

the bed to repressure and fill. When the cool and 

fill step is complete the beds are switched and treating is started on the cooled bed. 

DESIGN CONSIDERATIONS 

Two bed systems are most common for LP-gas f 

treating but three, four or multiple bed operation is 

also possible. Inlet and outlet nozzles should be 

equipped with baffle plates or equivalent distributors. To prevent loss of molecular sieve during an 

upset condition, nozzles should be covered with a 

fine mesh screen. Three-way valves may result in 

investment economies but two way valves are preferred to minimize leakage problems. 

Molecular sieve type 13X can last from 3 to 6 

years in treating service. Useful life and optimum 

performance are enhanced by proper choice of regeneration conditions and guarding against these 

problem areas: A" 

1. During start-ups, shut downs and turnarounds care should be taken to avoid carryover bumps 

of absorber oil, glycol or excessive amounts of water. 

A simple by-pass line is especially useful during 

start-ups. 

2. If corrosion inhibitors are injected upstream 

or absorber oil carryover is suspected, coalescers 

and/or a layer of activated carbon on the inlet end 

of the bed should be added. 

3. Iron sulfide and iron flake carryover may be 

eliminated by adding a filter on the feed inlet line. 

4. Sulfur compounds loaded on type 13X may 

react via the Claus reaction when exposed to air f 

during unloading. The bed should be thoroughly I 

regenerated and also purged with inert gas prior to 

dumping. 

5. Molecular sieve unloaded during maintenance 

turnarounds should be stored in airtight containers 

if re-use is planned. 

6. Minimize glycol, amine or absorber oil in the 

purge gas which could cause molecular sieve degradation. 

MOLECULAR SIEVE TREATING FEATURES 

1. Units may be designed for flowrates from only 

a few GPM to hundreds of GPM. 

2. Units may operate efficiently at throughputs 

well below the design condition if there is a decrease 

in requirements. 

3. Reduces total sulfur content by removing 

hydrogen sulfide, mercaptans and other organic sulfur compounds in the same adsorber vessel. 

c 

uct 

4. Will produce copper-strip, doctor sweet prod5. Will dry to zero or trace water levels in the ^ ~ 

same step without additional equipment. 

Rev. 1976. 1977, 1981 15-40 
P.32 

,..,.,... 

c 

c 

c 

c 

6. Non-corrosive process minimizes maintenance 

costs, downtime and disposal problems. 

7. Sulfur compounds removed by purge gas may 

be burned within fuel system. 

Carbonyl sulfide removal 

As concern for minimizing total sulfur emissions has 

increased, there has been a growth of interest in removing 

carbonyl sulfide (COS) from propane and LPG streams. It 

can, of course, be removed sacrificially by MEA. In recent 

years several alternative processes have been proposed. 

Mick1 has reported successful use of a combination of 

potassium hydroxide and methyl alcohol. This process is 

also sacrificial. The Malaprop2-

34

 process uses Diglycolamine but requires unspecified modifications in the process flow. The ADIPS

 process utilizes diisopropanolamine. 

Molecular sieves can also be used for removing COS. The 

Malaprop, ADIP and mole sieve processes are regenerative. 

1. Mick, M. B., "TVeat Propane for COS Removal," Hydrocarbon Processing, p. 137, July, 1976. 

2. McClure, G. P. and D. C. Morrow, "Malaprop Process 

Removes COS," Hydrocarbon Processing, p. 231, 

Jan. 1979. 

3. McClure, G. P. and D. C. Morrow, "Amine Process Removes COS from Propane Economically," The Oil and 

Gas Journal, p. 107, July 2, 1979. 

4. Weber, S. and G. McClure, "New Amine Process for FCC 

Desulfurizes Light Liquid Streams," The Oil and Gas 

Journal, p. 161, June 8, 1981. 

5. Klein, J. P., "Developments in Sulfinol and ADIP Processes Increase Uses," Oil and Gas International, 

September, 1970. 

Perco solid copper chloride sweetening process 

The copper chloride sweetening process is used for 

sweetening gasoline blending stocks. It is a continuous process in which the mercaptan sulfur in the 

gasoline is converted to disulfides in a reaction with 

copper chloride. 

The copper chloride reagent is impregnated in a 

stationary bed of fullers earth. As manufactured, this 

material contains approximately 16% water but loses 

its reactivity below 7%. To restore this reactivity, 

water should be added by (a) injecting a small quantity of steam into the gasoline feed stream to the 

treater, or (b) bubbling a sidestream of gasoline 

through a small pot containing water and returning 

it to the main stream. Above 30%. the copper chloride 

tends to go into solution and thus be extracted from 

the fullers earth. If the gasoline feed stream is water 

saturated, it will be necessary to heat it to prevent 

the formation of free water which would tend to 

dissolve copper chloride from the fullers earth. Any 

hydrogen sulfide in the gasoline must be removed by 

a caustic wash before contact with the conversion bed. 

In the sweetening reaction mercaptans are oxidized to the disulfide; 

(1) 4RSH + 4CuCL 2RSSR + 4CuCl + 4HC1 

The cuprous chloride is reactivated simultaneously 

by adding air to the inlet gasoline stream: 

(2) 4 CuCl + 4 HCI + O.. 4 CuCL + 2H..0 

Standard solid copper reagent towers 

Charge Rate 

bbl/doy 

187 

424 

752 

1177 

1500 

Tewer 

>iu , ft 

2 X 10 

3 X 10 

4 X 10 

5 X 10 

5 X 12 

Bed vel 

cu ft 

21.9 

49.5 

88.0 

137.5 

176.5 

Reagent 

Ib 

1090 

2475 

4400 

6870 

8830 

The standard bed depth is 7 ft with a maximum 

of 9 ft and a minimum of 4 ft if standard towers are 

not used. 

The charge rate is based on two volumes of gasoline/volume of reagent/hour. If the mercaptan content of the gasoline is low, these charge rates can be 

exceeded. 

Air is added to the sour charge by means of a diffusing disc at the rate of 1 cu ft of air (STP)/bbl of 

charge per 0.01% mercaptan sulfur. 

If the mercaptan content of the gasoline is high, 

the operating pressure must be increased to hold the 

additional air in solution. 

The rate of air injection should be maintained as 

near to this figure as practical since insufficient air 

will prevent complete reactivation of the reagent and 

excess air may result in airbinding of the treater 

vessel; an extremely hazardous condition. 

General notes 

There are many licensed processes for product treating. 

Information on both processes and suppliers is available in 

Hydrocarbon Processing special issues: September even 

years—"Refining Handbook" and April odd years—"Gas 

Processing Handbook". 

15-41 Rev. 1976. 1977. 1981 
P.33 

Mixing 

Very good mixing of gasoline and treating solutions is 

accomplished by means of a low efficiency (40%) centrifugal pump. The gasoline stream and solution recycle stream 

join just ahead of the pump in equal volumes and are discharged to the solution settling tank. The size of a mixing 

pump is based on four times the gasoline flow rate. 

Sand towers 

Sand towers are usually placed downstream of a liquid 

wash of gasoline or LP-gas to remove any entrained water 

or solution. The gasoline flows downward through the 

tower. The design of the tower is based on a space velocity 

of 4 volumes of gasoline/volume of sand/hour. 

Distribution headers 

Distribution headers should be provided in vessels used 

for liquid-liquid contacting. 

TREATING SELECTION 

As a guide in the selection of the method of treating to 

be used, the following characteristics of each are given: 

1. Regenerative caustic 

a. Handle large volume of hydrocarbon. 

b. Primarily for removing methyl and ethyl mercaptans. 

c. Capable of producing a "doctor sweet" product. 

d. Reduces the total sulfur content of treated product 

2. Perco solid copper chloride process 

a. Can treat gasoline streams with relatively high mercaptan content. 

b. Suited for small flow rate. 

c. Sulfur content not reduced. 

d. Water content must be only that of saturation. 

e. Hydrogen sulfide must be removed ahead of contact 

with the bed. 

f. If the product is fractionated after treating, the overhead product will be essentially sulfur-free. 

g. Excess regeneration oxygen may cause corrosion 

downstream of bed. Gasoline with components that 

may be affected by oxygen, such as olefins, should 

not be treated in this process. 

3. Caustic wash 

a. Use a single wash or two or three washes in series. 

b. Best suited for streams with low mercaptan content. 

c. Primarily for removing hydrogen sulfide or methyl 

and ethyl mercaptans. 

d. Disposing of spent caustic is a major consideration. 

e. Relatively high caustic consumption per 1,000 gal of 

gasoline. 

4. Solid potassium hydroxide treatment 

a. Low installation and operating costs. 

b. Acts as a desiccant as well as removing the sulfur 

compounds. 

5. Molecular sieves treating 

a. Can handle large or small streams with equal facility. 

b. Reduces total sulfur content by removing hydrogen 

sulfide, mercaptans, or other organic sulfur in the 

same adsorber vessel. 

c. Will produce copper-strip, doctor sweet product. 

d. Will dry in the same step without additional equipment. 

e. Requires that plant fuel supply be made slightly sour f 

by use of this stream to strip sulfur compounds from V 

the molecular sieve. 

6. Merox treating 

a. Can handle large or small streams with equal facility. 

b. Eliminate problem of disposal of spent phenolic caustic. 

c. Product is sweet when it leaves unit; no holding period is required. 

d. Adds problem of sulfur compounds from regenerated 

air leaving the regenerator. 

LP-gas dehydration 

It is usually necessary to dehydrate LP-gas (es- y^ 

pecially propane) following treating to prevent th« 

formation of hydrates as the LP-gas is cooled. Hydrates can form in propane, for instance, at 42°F. 

Normally, a solid desiccant dehydrator, similar to 

the type discussed in the gas dehydration section, is 

employed to dehydrate LP-gas. Desiccants which are 

suitable for drying natural gas are generally also 

satisfactory for drying LP-gas, except that unsaturated compounds, such as propylene, are polymerized 

by desiccants which contain a large proportion of 

silica. 

LP-gas mixtures which contain appreciable quantities of such materials should be carefully studied to f 

determine the most suitable desiccant. It may be f 

necessary to purge the desiccant bed with natural ^ -

gas before regeneration with LP-gas (containing only 

saturated compounds) or gas. 

Olefin polymerization can be avoided by using 

molecular sieve Type 3A which selectively adsorbs 

water, only. 

Regeneration of the water-saturated desiccant bed 

is usually accomplished by withdrawing a small volume of LP-gas from the inlet stream and passing it 

through a vaporizer to superheat it before it enters 

the desiccant bed. The water-bearing LP-gas vapor 

from the bed is cooled in a heat exchange to condense 

the water and then flows into a water-LP-gas separator. The LP-gas phase is returned to the inlet LP- S 

gas stream downstream of a back-pressure valve v_ 

which provides the pressure differential required to 

propel the regenerant LP-gas through the circuit. 

The heat load includes the heat of desorption of 

the water, the heat of desorption of the adsorbed 

hydroarbons (usually assumed to be equal to the 

heat of desorption of the water), sensible heat to 

bring the water vapor, desiccant, and steel vessel up 

to temperature, and atmospheric heat loss. As in the 

case of the gas dehydrator, the total heat load is set 

at twice the calculated heat load to compensate for 

variation in the rate of heat exchange during the 

heating period. 

The total quantity of LP-gas required for regeneration is established by dividing the total heat load 

by the sensible heat of the vaporized LP-gas. For an 

e 

Rev. 1976, 1977, 1981 15-42 
FIQURE 2 

C02 

ABSORBER 

GIftMMflRCO-VETROCOKE PROCESS 

AIR OUT 

CO2 OUT • 

SOLUTION PUHP 

AIR 

BLOUER 

URTER 

CIRCULATION PUMP 
FIOURE 3 

TYPICAL GAS SWEETENING BY PHYSICAL ABSORPTION 

r

2

^* 

H28 AND 

STRIPPING 0R8 

X 

POWER 

RECOVERY 

TURBINE 

I 

I 

I 

I 

f 

SOLVENT 

RECOVERY 

8TAIPPIN0 ORS 
P.36 

.,,,, • -, , 

PROCESS DESIGN OF VARIOUS TYPES OF 

SULFUR REMOVAL PROCESSES 

By 

Leon L. Patterson 

Principal Process Engineer 

FORD, BACON & DAVIS, INCORPORATED 

Texas Division 

Dallas, Texas 

June 1983 

4fbro, Bacon A^avis. 3ncorporalpo 
P.37 

„ f ;,,-

Page 1 

PROCESS DESIGN OF VARIOUS TYPES OF 

SULFUR REMOVE PROCESSES 

Prior to the 1960's, the easiest and most economical way to handle 

sour gas was to "plug the well" because the price of gas was in the 

15C/MCF range. As the demand increased by both industry and the public, 

the gas reserves decreased. In order to keep up with the demand,; deeper 

wells have been drilled which have produced more sour gas which has to be 

processed. Presently, associated gas from wells in Wyoming and other 

states contains as much as 42% hydrogen sulfide and 29% carbon1 dioxide. 

As the value of the crude oil increased, treating of this gas was 

required, not for the gas itself but for the crude oil production. This 

increased volume of sour o.a<s and the higher cost of energy have placed hew 

emphasis on the treating processes. 

Acid gas removal may be required for a number of reasons: 

1. Specification of Gas 

(Almost all of the sales gas contracts limit the amount of hydrogen 

sulfide to no more than 0.25 grain per 100 cubic feet of gas.) 

2. Specifications of Liquid Products 

(LPG has limitations on hydrogen sulfide and organic sulfur content.) 

3. Water saturated gas containing hydrogen sulfide and carbon dioxide 

can be corrosive to carbon steel. 

4. Safety & Environmental Protection 

(Hydrogen sulfide is a dangerous-toxic gas; i.e., 700 ppm in air is 

considered fatal in exposures of 30 minutes or less.) 

5. Carbon dioxide, an inert, lowers the heat content of the gas. 

(Almost all gas contracts are made on a BTU content.) 

6. Carbon dioxide can cause problems at cryogenic plant temperatures. 

4foro, Bacon &3tovie. 3ncorporalro 
P.38 

Page 2 

A study should be made to determine the process for the specific application. Questions that need to be answered include: 

1. What impurities will there be? 

2. What is the volume of gas to be processed? 

3. What is the end use—for gas or liquids? 

4. What chemical or physical process will be used? 

5. What are the causes of degradation? 

6. Can it meet the required gas specification? 

7. Can it meet the required specification of the acid gas? 

8. Is selective hydrogen sulfide removal a viable solution? 

9. Is reclaiming required? 

10. What are the heat and energy requirements for the process? 

11. What royalty is required? 

The processes that should be considered are: 

1. Hot Carbonate Processes (Figure 1) . 

Employs an aqueous solution of potassium carbonate. The absorber and 

regenerator both are at elevated temperatures. 

A. Proprietary processes include: 

(1) Benfield - several activators 

(2) Girdler - alkanolamine activators 

(3) Catacarb - alkanolamine and/or borate activators 

(4) Giammarco-Vetrocoke (Figure 2) - arsenic and other 

activators 

B. The advantages of the hot carbonate processes are: 

(1) Continuous circulating system employing an inexpensive 

chemical. 

(2) Iso-thermal system where both absorption and desorption of 

the acid gases are nearly at uniform high temperature, 

therefore no heat exchange equipment is required in the 

fluid circulating system. 

(3) The desorption of stripping is accomplished with a smaller 

steam rate than required for an amine plant. 

,iford, .Bacon A^tovis, 3ncorporalro 
P.39 

Page 3 

C. The disadvantages of the hot carbonate processes are: 

(1) Further treatment is required by a conventional amine 

system to economically produce pipeline specification 

product. 

(2) The units are prone to corrosion as with other acid gas 

removal processes. 

(3) Problems are encountered in suspended solids and foaming as 

with other processes. 

Physical Processes (Figure 3) that are available include: 

A. Fluor Solvent - uses anhydrous propylene carbonate. 

(1) The temperature of the solution is well below ambient, 

which requires refrigeration. 

B. Selexol - uses dimethylether of polyethylene glycol as a 

solvent. 

(1) Absorption-removal of sour gas fraction in feed gas in a 

conventional countercurrent absorption column. 

(2) The solvent is regenerated by a series of flashes and 

possibly a stripper. 

(A variety of stripping gas includes air, nitrogen, 

combustion gas, or steam.) 

C. Sulfinol Process - uses Sulfolane (Tetrahydrothiophene dioxide), 

Di-isopropanolamine (DIPA) and water, or Methyldiethanolamine 

(MDEA) in the place of DIPA. 

(1) Bulk removal of carbon dioxide is achieved with minimal 

heat requirements and with no refrigeration. 

(2) Removal of H S, COS, mercaptans and other organic sulfur 

compounds is accomplished to meet pipeline specifications. 

(3) The flow scheme is similar to an amine system. 

D. Rectisol Process - uses methanol as the solvent applied at 

extremely low temperatures, i.e., -30°F to -100°F. 

4foro, Bacon &3Javis, incorporated 
P.40 

I! VJl'J 

Page 4 

These physical absorption processes operate in a comparable way 

as oil absorption. They should be considered when: 

(1) The partial pressure of the acid gas in the feed is above 

50 psi. 

(2) The concentration of heavy hydrocarbons is low, as the 

physical absorption also picks up heavy hydrocarbons. 

(3) When bulk H S removal is desired. 

Dehydration of the treated gas is not required in the Rectisol 

or Selexol processes. 

Slurrisweet Process 

Removes small amounts of hydrogen sulfide (up to 100 grains/100 cubic 

feet), with the spent solution to be disposed of a's waste. 

Preliminary tests indicate the spent solution •./ be a very good 

fertilizer. 

Iron Sponge Process 

Utilizes a hydrated iron oxide which is impregnated on wood chips. 

A. This system can be regenerated a few times with air. 

B. Care must be exerted in the disposal of the spent chips as the 

iron sulfide is pyrophoric. 

Dry Bed Process 

Molecular sieves are being used for removal of sulfur compounds. 

A. The operation of adsorbent plants is quite simple but the design 

is very complex. Manufacturers must be consulted for potential 

applications. 

Amine Type Units - Chemical Absorption 

Mines that may be used are: MEA, DEA, MDEA and DGA. These are the 

most common types of units. The specific design details for a 

typical amine process are as follows. 

.ford, Bacon d^avie, incorporated 
P.41 

lf|l'IV 

Page 5 

Prior to any type of gas-liquid contact, whether it is a sour gas 

treatment process, gas dehydration, or a natural gas liquids extraction 

unit, an effective inlet gas separator must be designed to remove 

entrained droplets or slugs of liquid prior to entering the system. 

Typical contaminants in the natural gas are liquid hydrocarbons, salt 

water, sands, well treating chemicals, pipeline treating chemicals, and 

compressor oils. 

These contaminants can cause a number of operational problems once 

they get into the sweetening system. These problems may include foaming, 

equipment fouling, and high corrosion rates, which result in solvent loss 

and difficulty in meeting sweet gas specifications. 

There are four major types of separators available: vertical, 

horizontal single tube, horizontal double tube, and spherical. 

The advantage of the vertical separator is the capability of handling 

greater slugs of liquid without carryover to the gas outlet. Some of the 

disadvantages of the vertical separator are that it is more difficult and 

expensive to fabricate and ship in skid mounted assemblies. It takes a 

larger diameter separator than a horizontal separator. 

Spherical separators are inexpensive and compact. They do, however, 

have limited surge space and liquid settling. The location and action of 

the liquid level control is very critical. 

The horizontal separator has several advantages over the vertical 

separator. It is easier to skid mount, less piping is required, and a 

smaller diameter is required for a given gas capacity. It has a larger 

interface area between the liquid and gas phases, which aids in 

separation. However, the liquid level control placement is more critical 

and the surge space is somewhat limited. The double tube has a slight 

advantage over the single tube but obviously is more expensive. 

4ford, Bacon ditovie, incorporated 
P.42 

Page 6 

The gas capacity of the oil-gas separators has been designed for many 

years by the use of the Souders-Brown equation: 

Where: 

V .£ 

V = 

A = 

Q = 

G 

K 

P L - P G Then A = 2 

v 

Superficial gas velocity based on total cross-sectional 

area of vessel, ft/sec 

Cross-sectional area of separator, sq.ft. 

Gas flow rate at operating conditions, ft3/sec 

Density of liquid at operating conditions, lb/ft3 

Density of gas at operating conditions, lb/ft3 

Imperical factor 

For vertical separators: 

5 feet high - K 

10 feet high - K 

0.12 to 0.24 (avg. 0.167) 

0.18 to 0.35 (avg. 0.21) 

For horizontal separators: 

10 feet long - K 

Other lengths - K 

= 0.40 to 0.50 (avg. 0.45) 

" 0.45 *, '•« 

For spherical separators: 

K = 0.35 

DESIGN OF AMINE SYSTEMS (Figure 4 and Figure 5) 

The absorption and desorption of H S and CO are complex mass 

transfer and chemical reactions. The H S and CO react with the amine 

with the following reactions: 

Primary and Secondary Amines 

RR'NH + 

RR'NCOO" 

RR'NH + 

co2 

+ H20 

H2S 

-• 

-• 

- * • 

RR'NCOO~ 

RR'NH + 

RR'NH * 2 

+ 

+ H 

HCO ~ 

+ HS~ 

(Fast) 

(Slow) 

(Very fast) 

,1/ord, Bacon &"5>avi$. incorporated 
P.43 

irpii!'!' :•! 

Page 7 

Tertiar y Amines 

R2NR« 

R2NR' 

R2NR' 

+ 

+ 

+ 

C0? 

co2 

H2S 

+ H20 

-•• 

->• 

-• 

No reaction 

R NHR,+ + HCO " 

+ -

R NHR' + HS 

(Slow) 

(Very ti 

1 

The tertiary amines cannot react directly with CO to form carba2 i 

mates. The C02 must first react with water, then combine with the amine 

via the acid base reaction. 

To say it in another way, the rate of reaction of CO with amines is 

slower than the rate of reaction and transfer of H S. Significant 

differences are experienced in the rate of absorption of CO from1 one type 

of amine to another. 

Tertiary amines exhibit the greatest selectivity for H S; all amines 

show some selectivity. Selectivity increases with decreasing temperature 

and decreasing pressure. 

The design of an amine gas. treating plant includes determining the 

solvent circulation rate, reboiler heat duty, and tower diameters. The 

amine circulation rate is determined by an iterative calculation of heat 

and material balances around the absorber, based on an estimated :approach 

to equilibrium of the acid gas in the rich solvent at the bottom of] the 

absorber and the temperature of the exiting sweet gas stream at the top. 

To prevent hydrocarbon condensation, the lean solvent temperature should 

be around 10°F warmer than the feed gas temperature. The product gas 

temperature is assumed to be the same as the temperature of the lean 

solvent. 

Several articles contain equilibria data on the solubility of H S, 

(12 3) C0„, and H S-CO mixtures in MEA and DEA. ' ' The use of the'se data, 2 2 2 (3 4) 

as well as the heat of reaction of H S and C00 ' , enable a rigorous 

determination of tho required strength and circulation rates. 

.iford. Bacon &"$avie, incorporated 
P.44 

- M r j,, r . , —> 

Page 8 

Computer programs have been developed for a rigorous solution to 

amine design. Some of these programs are: Simulation Science, ChemShare, 

TSWEET, AMSIM, as well as Oklahoma State University. 

Absorber-Contactor 

. Overall tray efficiency will be about 25 to 40%; therefore, 16 to 20 

actual trays are generally used. Some of the aforementioned computer 

programs model the absorber with the use of Murphree Tray Efficiency. 

After contact in the absorber, the rich amine solution passes into a 

hydrocarbon flash tank. The function of the flash tank is to remove most 

of the dissolved hydrocarbons. Data on the solubility of methane and 

ethane in aqueous MEA and DEA solutions, either free of acid gas or with 

H S and CO , were reported by Lawson and Garst. Because of the 

molecular structure, DEA has the poorest affinity for hydrocarbons. The 

next level is MEA and DIPA, with DGA having the strongest affinity for 

hydrocarbons. A conservative approach would be to use hydrocarbon 

solubilities in water at various temperatures and pressures. 

A slipstream of lean amine to a small packed scrubber will reabsorb 

the acid gases flashed out of solution with the hydrocarbon vapors. A 

residence time of approximately 3 to 4 minutes is required for this flash 

tank. Some of the flash tanks have a weir provided to enable periodical 

removal of liquid hydrocarbons that may be entrained with the rich amine 

solution. 

The rich amine passes from the flash tank through lean-rich amine 

heat exchangers. Solution heat exchanger corrosion problems occur when 

acid gases are liberated at the hot unit. General practice is not to 

preheat the rich amine solvent above 190°F. These exchangers should have 

impingement baffles at shell entry nozzles to minimize tube erosion and 

corrosion. 

,iford, Bacon &3>avis. incorporated 
P.45 

?age 9 

These exchangers should be stacked vertically with rich amine 

solution passing through the tubes and entering the lower part of the 

bottom exchanger in order that the liberated gases will be removed in 

i • ' 

upward flow of amine solution. The maximum velocity to minimize corrosion 

in the tubes is 2 ft/sec. 

Regenerator (Stripper) 

The stripper is either a tray or packed column with approximately 20 

trays or the packing equivalent. To reduce amine vaporization lo|ss, there 

is generally a wash section at the top of the vessel with an ajdditional 

four to six trays. The preheated, rich amine enters near the top of the 

vessel and flows down countercurrent to a gas stream consisting of steam, 

H S and CO . The reboiler generates the steam, which reduces the partial 

pressure of H S and CO in the gas strengthening the driving force of the 

i acid gases from the amine solution. The overhead gas passes through a 

condenser to recover water and the small amount of amine which is 

vaporized in the regenerator. 

The reboiler heat duty includes the sensible heat required to raise 

the temperatures of the rich amine feed, the reflux, and the make--up water 

to the temperature of the reboiler, the heat of reaction to break the 

chemical bonds between the acid gas and the amine as well as the, heat of 

vaporization of water to produce a stripping vapor—steam. The reflux 

ratio (moles of steam to moles of acid gas in the overhead gas, up stream 

of the condenser) ranges from 1.5 to 4 to 1 ratio, depending upon the 

required purity of the lean amine. 

A filtration scheme of mechanical and activated carbon filters is 

important in maintaining good solution control. Mechanical filters 

(cartridge or precoat) remove particulate material, while the carbon 

filter removes the chemical contaminants, i.e., entrained hydrocarbons and 

surface active compounds. 

.iford, Bacon &"3>avi$, incorporated 
P.46 

Page 10 

Filters are located in the rich line in some plants; in other, plants 

they are located on the lean line. Locating the filters in the (rich line 

up stream of the lean/rich exchanger will protect both the heat exchanger 

and the stripper from plugging by reducing the erosion/corrosion irate in 

the exchanger. However, in terms of safety, it is far more hazardous to 

locate the filters in the rich line than the lean line. The mechanical 

filter should be located in front of the carbon filter as the carbon 

filter would plug considerably faster due to solid material. A 10 . to 

20 micron mechanical filter should be adequate for particulate removal. 

Circulation rates through the mechanical filter varies by preference from 

5% to full flow. Circulation rates are similar for the carbon filter. 

The carbon filter should be changed when a high pressure drop exists 

across the bed as with the mechanical filter, or when a color comparison 

can be made between a sample taken from filter effluent and a sample run 

through a fresh carbon filter in the lab. An increase in foaming tendency 

will also require a filter change. 

A side stream reclaimer operation for MEA or DGA (usually 0.5% to 3% 

of circulated volume) is used to effect the separation of the amine from 

its heat stable salts. The flow is controlled by heat input rate and a 

level controller. Initially the reclaimer boiling temperature rises as 

the concentration increases. Temperature cutoffs in the MEA are in the 

280°to 300°F range to prevent carryover. A high shutdown temperature 

control is desirable to prevent carryover with subsequent foaming. 

Some companies request a means of injecting caustic soda to the 

reclaimer. A study made by the Societe Nationale Elf Aquitaine of 

(7) France indicates that the caustic soda reacts in an irreversible way, 

first with the strongest acids, then if it is in excess, with H S and CO 

to form sodium sulfide and carbonate. 

The heat flux on a steam reclaimer should be between 6 to 

8 MBTU/hr.ftJ. The heat flux on a steam heated reboiler should be between 

9 to 10 MBTU/hr.ft2. The normal reboiler temperature for MEA is 225 to 

260°F, for DEA it is 230 to 250°F, and for DGA it will range from 250 to 

260°F. 

.ifard. Bacon &"^avi6. incorporated 
P.47 

Page 11 

MATERIALS OF CONSTRUCTION (Figure 6) 

Absorber-Contactor 

Type 304 or 410 stainless steel trays are recommended to prevent 

sticking of the valves. The vessel should be constructed with 1/8" 

corrosion allowance carbon steel. 

The maximum velocity of the lean-rich exchanger should be limited to 

2.0 ft/sec. Preference of 304 stainless steel tubes has been indicated by 

some contractors or operating companies. If the pressure break (control 

valve) is at the stripper, this may not be justified. Carbon steel with a 

corrosion allowance of 1/4" has been used satisfactorily. Chahging to 

stainless steel tubes has increased the life approximately double. If 

carbon steel is used, stress relieving the bundle is necessary. 

Reboiler tubes of carbon steel have been used successfully with a 

corrosion allowance of 1/4" on the shell, with a stress relieved bundle 

and shell. Do not exceed 350°F on the steam side. Overhead condensers 

have been constructed with carbon steel with no apparent disadvantage. A 

small amount of amine in the overhead condensate will provide corrosion 

inhibition. Experience of good operators has shown that about 0.5% amine 

in the condensate is a happy medium. Allowable velocities in the piping 

system should be limited to 3 ft/sec. Control valves should have 

velocities between 15 and 20 ft/sec. 

Regenerator (Stripper) 

Type 304 stainless steel trays should be used in the regenerator. 

The corrosion problem in this tower appears to be more severe than that in 

the absorber discussed earlier. The vessel should be constructed out of 

carbon steel with 1/4" corrosion allowance. The reflux drum should be 

stress relieved. Corrosion is high here as attested by the need for the 

1/4" corrosion allowance. 

/ord, Bacon d^tovis, incorporated 
P.48 

Page 12 

CONCLUSION 

The proper selection of the process for acid gas removal is]necessary 

to reduce the excessive costs, expenses, and headaches we ,have all 

\ experienced in the past. More innovative processes must be developed to 

| I- J 

be able to efficiently process the very sour gas produced from the deeper 

formations and the associated gas from the oil wells in Wyoming. 

.iford. Bacon &3tovis. incorporated 
P.49 

•«V-»rrr'-'" " 

Page 13 

REFERENCES 

1. Lee, Otto and Mather, "Solubilities of Acid Gases in Amine Treating 

Solutions", CNGPA, November 30, 1973. 

Ibid, "Solubility of Acid Gas in MEA Solutions", CNGPA, November 29, 

1974. 

Ibid, "Solubility of H S and CO in Aqueous MEA Solution", Canadian 

Journal of Chemical Engineering, Vol. 52, December 1974. 

2. Lawson, J.D, and A.W. Garst, "Gas Sweetening Data; Equilibrium 

Solubility of H S and CO in Aqueous MEA and Aqueous DEA Solutions", 

Journal of Chemical Engineering Data, Vol. 21, No. 1, 1976. 

3. Gas Processors Suppliers Association, Engineering Data Book, GPSA, 

Tulsa, Oklahoma, 1981. 

4. Kohl, A.L. and F.C. Riesenfeld, Gas Purification, 3rd Edition, Gulf 

Publishing Company, Houston, Texas, 1979. 

5. Lawson, J.D. and A.W. Garst, Journal of Chemical and Engineering 

Data, Vol. 21, No. 1, pp. 30-32, 1976. 

6. McKetta, J.J., et al., Petroleum Refiner, various volumes, 1951, 

1958, 1961, 1962. 

7. Blanc, C , M. Grail, G. Demarais, Society Nationale Elf Aquita'ine, 

Lacq, Franc, "The Part Played by Degradation Products in the 

Corrosion of Gas Sweetening Plants Using DEA and MDEA", Proceeding ofGas Conditioning Conference, 1982, The University of Oklahoma. 

SUGGESTED REDIN G 

"Lease-Gas Sweetening", 1-11, The Oil and Gas Journal, August 14, 1967 

May 13, 1968. 

,iford, Bacon &3tovis. incorporated 
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FICURE 3 

TYPICAL GAS SWEETENING BY PHYSICAL ABSORPTION HoS ANO 

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CONTACTOR 

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FIQURE 4 

TYPICAL GAS SWEETENING BY CHEMICAL REACTION 

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W , COOLER 

CHARCOAL MECHANICAL 

FILTER FILTER 

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RECLAIMER 

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SPLIT-STREAM AMINE PROCESS 

^.flCID GRS 

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FIQURE 6 

TYPICAL GAS SWEETENING BY CHEMICAL REACTION 

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AUGUST 24 

SUMMARY OF COST FACTORS 

CAPITAL COST (BASED ON $4000/6PM OF AMINE SOLUTION) $ PER SCFD ACID GAS ,0.480 

ANNUALIZED CAPITAL COST .20 YEAR EQUIP. LIFE.: 107. INT. RATE;CRF=. 11246) $/YR PER SCFD ACID 6AS 0.054 

REGENERATION HEAT f/YR PER SCFD ACID GAS 0.240 

PUMPING COST W R PER SCFD ACID 6AS 0.016 

LABOR COST (IMAN/SHIFT, 3SHIFTS/DAY, 365DAYS/YR; 114.50 PER HOUR! $/YR 127000 

MAINTENANCE SUPPLIES (27. OF CAPITAL COST) W R PER SCFD ACID GAS 0.001 

CEMICAL COST (BASED QN 25% PER YEAR REPLACEMENT) W R PER SCFD ACID GAS 0.018 
jr,i'|'T", 

CASE 

GPA 

1 

2 

SULFUR 

LT/D 

1 

5 

5 

ACID 

GAS 

XH2S 

0.13 

0.5 

0.2 

FACTOR 

CAPITAL 

COST 

it) 

101875 

127344 

318360 

FACTOR 

CAPITAL 

COST 

(WR) 

11461 

14326 

35BI6 

FACTOR 

HEAT 

COST 

!WR) 

50938 

63672 

159130 

FACTOR 

PUMPING 

COST 

!WR) 

3396 

4245 

10612 

FACTOR FACTOR 

LABOR MAINTENANCE 

COST 

(WR) 

127000 

127000 

127000 

SUPPLIES 

($/YR) 

212 

265 

663 

FACTOR 

CHEMICALS 

COST 

(WR) 

3820 

4775 

11939 

FACTOR 

TOTAL 

O&M 

(WR) 

185366 

199958 

309394 

FACTOR 

, TOTAL 

COST 

; (WR) 

196827 

214284 

' 345209 

10 

11 

12 

5 

10 

10 

0.8 

0.5 

0.2 

79590 

2546BS 

636720 

8954 

28652 

71631 

39795 

127344 

318360 

2653 

3490 

21224 

127000 

127000 

127000 

166 

531 

1327 

2985 

9551 

23877 

172598 

272915 

491788 

181552 

301567 

563419 

21 

?? 

' 100 

100 

0.2 

0.5 

6367200 

2546880 

716310 

286524 

3133600 

1273440 

212240 

84896 

127000 

127000 

13265 

5306 

238770 

95503 

3774375 

15S6150 

4491185 

1872674 

27 1000 0.3 15913000 1790775 7959000 530600 127000 

23 1000 0.5 25463300 2365240 12734400 343960 127000 

29 1000 0.2 63672000 7163100 31836000 2122400 127000 

33163 

53060 

132650 

596925 

955030 

2337700 

9246638 

14718500 

• 36605750 

11037463 

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