8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 16

is staffed with experienced engineers and technologistsho provide technical and practical expertise

o the upstream oil and gas industry

alconrsquos main focus areas arengineering Procurement Construction Management (EPCM)

echanical Process Electrical Civil Engineering amp DraftingutoCAD CADWorx Plant (Caesar II Stress Analysis)

rocurement amp ExpeditingLC Programming

rctic and Offshore Oil and Gas Facilities

ternational Oil and Gas Facilities Design and Evaluation

weetSour OilGas Tie-Inspecialty Services - Process Design amp De-bottlenecking

egulatory Compliance

ontact Information

elephone 403-253-2741

AX 403-252-9613

ostal Address

6 Skyline Crescent NE

algary Alberta T2K 5X7 CANADA

Web Page wwwfalcon-edfcomeneral Information infofalcon-edfcom

alcon EDF is a full service engineering company

is staffed with experienced engineers and technologists

ho provide technical and practical expertise

o the upstream oil and gas industry

alconrsquos main focus areas are

ngineering Procurement Construction Management (EPCM)

echanical Process Electrical Civil Engineering amp DraftingutoCAD CADWorx Plant (Caesar II Stress Analysis)rocurement amp Expediting

LC Programming

rctic and Offshore Oil and Gas Facilities

ternational Oil and Gas Facilities Design and EvaluationweetSour OilGas Tie-Ins

pecialty Services - Process Design amp De-bottlenecking

egulatory Compliance

Refrigeration provides economic process for recovering NGL from CO2-EOR recycle gas

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 26

T E C H N O L O G Y

Kenneth J VargasFalcon EDF LtdCalgary

Refrigeration provides economic process forecovering NGL from CO2-EOR recycle gas

Removing and sellingNGL from the producedgas stream is one of theways CO

2 enhanced oil

recovery projects can im-prove their economics

These projects involve

gas recycling into thereservoir for maintaining pressure andimproving oil mobility The recycled

gas absorbs NGL in the reservoir that

can be recovered with a refrigerationprocess

The process involves refrigeratingthe gas and then separating stabilizingand recovering the NGL A companythen can market the recovered NGL asC

3 C

3+ C

4 and C

5+ or use it to spike the

crude

This article describes the followingaspects of NGL recovery

bull Process plant required for eachalternative (C

3+ fractionated liquids or

crude spiking)bull Liquid recoveries for varying

process conditions such as chiller final

temperaturebull Process property method selectionbull Process flow diagram with mate-

rial balancebull Approximate costs for

different process alternativesbull Economics of the alter-

nativesIt is best to install this

NGL recovery and refrigera-tion facility early in the CO

2

injection projectrsquos life

Process facility A typical CO

2-EOR NGL recovery

plant consists of the basic batterywith additional equipment to handleincreased produced gas and water asfollows

bull Inlet separation

Production

PROCESS FLOW Fig 1

Solution-gasinlet separator

Low-pressureseparator

1B

1A

Btu C2

De-C2Btu A

20

Mol sieve 13x

4

5

VPCHP de-C2

Comp de-C2 Btu de-C2

Gas-gas

Mix gas-gas

Chiller

R sep 1

R comp 1

Mix-economizer

Economizer

R4

HP R2HP R1

Btu ref

Condenser

R comp 2

VE

De-C2 - chiller

VF

Comp K1 Btu K1

Cooler K1

Gas-gas K1Splitgas-gas

Gas-gas K2

Comp K2 Btu K2

HP P2

HP P1

HP K2

HP K2

Cooler K2

Inj 120 MM

Reda 2

Reda 1

Mix inj

Btu C3 cooler

Btu C3 prod

C3 cooler

C3 sales

Btu C3 reboiler

De-C3

Btu C4 cooler

Btu C4 prod

C4 cooler

C5+ cooler

C4 sales

C5+ sales

Btu C5+ prod

Btu C4reboiler

De-C4

3

5A 5B 66A

7

21

1

9

22

2

10

1

2

9

2

3

2

3

7

16

10

19

25

24

R2

R3 R5

R9

R5A

R8 R7

R6

R1 R10

11

14

19

16

17

15

23

18

12

8 13

6B

Reprinted with revisions to format from the January 18 2010 edition of OIL amp GAS JOURNAL

Copyright 2010 by PennWell Corporation

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 36

bull Treating or emulsion breaking forfurther phase separation

bull Liquid products storage (oil andwater)

bull Produced gas compression for alloil and water separation pressures

bull Water injection

Fig 1 shows the process f low dia-gram and Table 1 the material balancefor 14 of the most relevant streams Theadditional units for CO

2-EOR include

bull Solution gas gas-gas precoolingexchanger

bull Solution gas chiller with propanerefrigeration plant

bull Low-temperature separator to feedrefrigerated liquids to the fractionationplant

bull Amine or mol-sieve liquid NGL

sweetening unitbull Lean-gas pump-suction exchang-

ers The case shown in Fig 1 has two170 million cu mday (60 MMscfd)trains

bull Two compressor trains of 170million cu mday each to boost thepressure from 1256 kPa-g (182 psig)to obtain a dense phase for 92 CO

2

gas The critical point is 7500 kPa-g(1090 psig) and the pressure selectedwas 8860 kPa-g (1285 psig) which is

in the dense-phase region Fig 2 showsthe phase envelope for the solution gas

bull Two centrifugal multiphasepumps one per train to boost the pres-sure to the injection pressure of 15400kPa-g or 2230 psig

The analysis looked at the following

three alternatives for liquids recovery1 Fractionation plant with de-eth-anizer depropanizer (sell C

3 product)

debutanizer (sell C4 and C

5+ product)

and distillation towers with ancillaries(aerial condensers reflux drums reboil-ers and product coolers)

2 De-ethanizer to produce C3+

product and sell

3 De-ethanizer to produce C3+

andspike the crude

Process design The design assumed a gas with a 15sp gr (436 molecular weight) Table 2shows the gas composition

This article presents the designfor only the 92 CO

2 gas content

with a full fractionation case becausethe recoveries for this case are moreconservative due to the high CO

2

concentration The base case includesa fractionation train because all otheralternatives are subsets of it The other

two alternatives either sell or spike thecrude with the C

3+

The analysis used the Peng-Robinsonproperty method for all the simula-tions

Table 3 summarizes the key streamsimulations of the process shown in

Fig 1

Propane refrigeration The refrigeration loop is in the up-

per right-hand corner of Fig 1 A gas-gas exchanger precools the plant inletsolution gas It uses cold gas off of thelow-temperature separator

After the gas-gas exchanger a chillerrefrigerates the gas to ndash29deg C Therefrigeration feeding the chiller on theshell side is a propane loop with an

economizer on the interstage of thepropane screw compressor

The schematic simplified thepropane loop as a two-stage refrigera-tion without a second chiller on thelast stage It shows only one chiller forsimplicity

A two-stage refrigeration loop wouldreduce the compression by 19 andthe condenser duty by 8 The mate-rial balance (Table 1) ref lects this Itshows the process requires 3940 hp

for the first stage(HP R1) and 670hp for the secondstage (HP R2) Thedesign corrects thetotal 4610 hp byndash19 for two-stage eff iciencyThus refrigera-tion compressionwould total about3750 hp

The propanerefrigeration con-denser would pro-vide 339 MMbtuhr times 091 or 31MMbtuhr

Note the pro-cess requires anethylene-glycolinjection loop fordehydrating the

Fig 2PHASE ENVELOPE

ndash140 ndash120 ndash100 ndash80 ndash60 ndash40 ndash20 0 20 40

P r e s s u r e

k P a - g

8000

7000

6000

5000

4000

3000

2000

1000

0

B - cricondenbar

T - cricondentherm

C - critical point

Hydrate curve

vf 0000

vf 1000

BTC

Temperature degC

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 46

T E C H N O L O G Y

MATERIAL BALANCE

Stream 2 13 18 Btu-Ref C3-sales C

4-sales C

5+-sales

Vapor fraction 07994 10000 00000 20000 10000 00000 00000Temperature degC ndash290000 380000 320000 00000 430000 430000 430000Pressure kPa-g 13248419 87929143 87929143 00000 11948897 5743614 6088352Flow MMscfd 1200000 590204 589976 00000 03459 07178 09172Liquid flow cu mday 79048550 38248715 38233949 00000 360220 861025 1344131Liquid flow bd 497199569 240576765 240483892 00000 2265710 5415676 8454313Mole weight 436115 432465 432465 441822 578919 791946Energy btuhr 250311E+07 101560E+07 449534E+06 338568E+07 2494771803 150502347 ndash44669296Energy hp 98376048 39914675 17667385 133062406 980482 59150 ndash17556H

2 mole fraction 0000000 0000000 0000000 0000000 0000000 0000000

N2 mole fraction 0008801 0008949 0008949 0000000 0000000 0000000

CO2 mole fraction 0918692 0934119 0934119 0000286 0000000 0000000

H2S mole fraction 0009101 0009245 0009245 0000000 0000000 0000000

C1 mole fraction 0015102 0015355 0015355 0000000 0000000 0000000

C2 mole fraction 0013501 0013728 0013728 0000043 0000000 0000000

C3 mole fraction 0016002 0013220 0013220 0993553 0022790 0000000

iC4 mole fraction 0002900 0001526 0001526 0005000 0231329 0000198

nC4 mole fraction 0007401 0002945 0002945 0001118 0739641 0009998

C5+

mole fraction 0008500 0000912 0000912 0000000 0006240 0989804

GAS COMPOSITION Table 2

Component Mole fraction

H2 00000

He 00000N

2 00088

CO2 09187

H2S 00091

C1 00151

C2 00135

C3 00160

iC4 00029

C4 00074

iC5+

00085

PROCESS CONDITIONS Table 3

Inlet parametersLevel of CO

2 92

Inlet separator pressure kPa-g 1380Inlet separator temperature degC 27 Inlet gas-gas exchangerTubeside in-out temperature degC 27ndash1Shellside in-out temperautre degC ndash28 3Duty MJhr MMbtuhr 84 80 Propane refrigerationChiller in-out temperature degC ndash1ndash29Chiller duty MJhr MMbtuhr 7874Propane compressor hp 4610Propane condenser duty MJhr MMbtuhr 35831 De-ethanizerFlow in cu mday 1689Tower diameter mm 1100Reboiler duty MJhr MMbtuhr 15815Liquid produced cu mday 257Overhead gas 1000 cu mday 625Recompressor hp 100 Inlet gas-gas K1-K2 exchangersDuty MJhr MMbtuhr 40-3838-35Minimum tempererature out degC 38 (gas) 32 (liquid)

Pump maximum flow cu m day MMscfd 382459Pump flow specific gravity 04-055Pump hp 1454-1064 DepropanizerFlow in cu mday 256Tower diamater approximate mm 610Reboiler duty MJhr MMbtuhr 1918Liquid produced cu mday 220Overhead gas 1000 cu mday 98 DebutanizerFlow in cu mday 86Tower diamater approximate mm 310Reboiler duty MJhr MMbtuhr 3129Liquid produced cu mday 134Overhead gas 1000 cu mday 203

gas to avoid gas hydrates after cooling

For simplicity Fig 1 does not showEG injection

NGL stabilization The lower left-hand portion of Fig

1 shows the fractionation plant Therefrigerated gas goes to a low-tempera-ture separator which separates the liq-uids that enter the fractionation plant

The fractionation plant has three dis-tillation towers The first is a de-etha-nizer (de-C

2 in Fig 1) with a reboiler

as the bottom stage The deethanizedliquids go to the depropanizer (de-C

3) unit consisting of tower overhead

condenser reflux-drum and bottomsreboiler

The specification sales propane isStream C

3 sales in Fig 1

The depropanizer bottoms go toa similar tower reflux and reboilerdistillation column This last distillationcolumn is a debutanizer (de-C

4)

The tower overheads go to butanesales (Stream C

4 sales) The column

bottoms are the light gasoline sales orC

5+ salesThe de-ethanizer overhead gas after

recompression mixes with the refriger-ated gas off the low-temperature sepa-rator The streamthen goes to thegas-gas shell sideof the exchangerand subsequentlythe reciprocatingcompressor-pump

tandem combina-tion

Fig 1 shows anaerial cooler afterthe stabilizer over-head of the re-compressor (compde-C

2) however it

is not required

Desulfurization Fig 1 shows

the sweetening ofthe liquids witha mol-sieve unitoperation on thebottom of the de-ethanizer reboilerproduct stream A13x Grade Z10-03mol-sieve unit or aliquid amine con-tactor can sweeten

the stabilized liquid NGLPreferable is a mol-sieve unit because

it has a dry system that can be regener-ated with hot fuel gas A typical 13xGrade Z10-03 mol-sieve unit has two orthree contactors to ensure 24-hr sweet-ening of sour NGL

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 56

Gas-gas cooling The solution gas from the inlet

gas-gas exchanger enters two streamsOne is the inlet of the tube side of theK1 exchanger (Stream 8 not shown inTable 1) and the other is the tube sideof K2 exchanger (Stream 9 not shownin Table 1) About half of the 3392 mil-lion cu mday enters each exchanger

The exit gas from the tube side ofthese exchangers feeds the K1 and K2

compressors The compressor discharg-es back into the shell side of K1 and K2

exchangers Utilizing parallel K1 andK2 exchangers ensures that the suctionstreamsrsquo temperatures feeding pumpsReda 1 and 2 are as low as possible

A temperature cooler than 33deg Cis optimal to ensure the multistagecentrifugal pumpsrsquo lowest horsepowerdraw In our case this is 1010 hp at15200 kPa-g (2200 psig)

The K1 com-pressor trainis for summerconditions or am-bient temperature3deg C cooler thanthe compressor

coolerrsquos dischargeof 43deg C Thelowest achievabletemperature afterthe K1 exchangeris 38deg C

The corre-sponding pumprequires 1454hp vs the K2

exchangerrsquos discharge of 32deg C whichrequires 1064 hp This increases the

horsepower by 27Because the simulations are mainly

for refrigeration cooling and stabiliza-tion Fig 1 shows the solution-gas in-jection compressors with a single stageabove the K1 and K2 exchangers

Pump performance The process cools the streams from

NGL RECOVERIES Table 4

Liquid produc- Refrigera- Refrigeration tion increaseCase Type tion hp temp 983151C cu mday

1 Spike oil 1630 ndash23 1672 Spike oil 2200 ndash27 167 (no change)3 NGL C

3+ 1600 ndash23 176

4 NGL C3+

2200 ndash27 2155 Fractionate 4610 ndash29 C

3= 36 C

4= 86

C5+

= 134 Case 5 sum

of C3 C

4 and C

5+ 256

THEORETICAL MAXIMUM NGL RECOVERY Table 5

Maximum liquid Ideal gas NGL liquid yieldCom- Mole Volume conversion cu mdayponent fraction Mscfd cu ftgal (bd)

N2 00088 1056

CO2 09186 110232

H2S 00091 1092

C1 00151 1812

C2 00135 1620

C3 0016 1920 3637 200 (125692)

iC4 00029 348 3064 43 (27042)

nC4 00074 888 3179 106 (66508)

iC5 0002 240 2738 332 (20870)

nC5 00024 288 2767 394 (24782)

C6+

00042 504 2616 73 (45872) ndashndashndashndashndashndashndashndash ndashndashndashndashndashndashndashndashndashndashndashndashndash Total 120000 494 (310766)

From Table 4 the total theoretical NGL recovery from the recycle gas stream is 494

cu m day Our fractionation is recovering 256 cu mday or 52

FRACTIONATION PLANT COST Table 6

MillionCapital costs for 3912 million cu mday (120 MMscfd) plant $

300 cu mday de-ethanizer complete (includes recompression reboiler) 225000 hp propane refrigeration complete unit 90Two 9 GJhr heat exchangers (chiller and gas-gas) 13Two 27-40 GJhr heat exchangers (gas-gas K1 and gas-gas K2) 07NGL recovery depropanizer-butanizer complete 30Refrigeration major electrical and mechanical equipment 30NGL mol sieve sweetening 10Low-termperature separator 08Piping racks cable trays insulation buildings and other consumables 35Total all equipment and materials 245Installed costs (15 times equipment and materials) 368Contingency (20) 123 ndashndashndashndashndash Total 736

Costs from skid vendors

NGL COSTS RECOVERED Table 7

ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndash Rate ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashHydrocarbon cut $cu m times cu mday $day $ millionyear

C3 295 times 36 11000

C4rsquos 400 times 86 34000

C5+

345 times 134 46000ndashndashndashndashndashndashndash

Total 91000 3185

The crude spiking alternative gives lower yields because much of the NGL flash aftermixing The costs recovered are $345cu m times 167 = $58000day or $20165 million year

Table 1

HP-P1 HP-P2 HP-R1 HP-R2 Inj -120 MM R3 Solution gas

20000 20000 20000 20000 10000 10000 1000000000 00000 00000 00000 623072 ndash329945 27000000000 00000 00000 00000 154118830 503597 1380000000000 00000 00000 00000 1180180 402712 120000000000 00000 00000 00000 76482664 41872170 7904855000000 00000 00000 00000 481060657 263367577 497199569

432465 438165 436115370150E+06 270696E+06 100267E+07 169455E+06 210598E+07 212384E+07 543272E+07

14547471 10638752 39406609 6659849 82768284 83470270 213513939 0000000 0000000 0000000 0008949 0000000 0008801 0934119 0000000 0918692 0009245 0000000 0009101 0015355 0000000 0015102 0013728 0020000 0013501 0013220 0980000 0016002 0001526 0000000 0002900 0002945 0000000 0007401 0000912 0000000 0008500

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 66

T E C H N O L O G Y

the K1 and 2 last stage discharge intothe Reda pumps to 38deg C for the sum-mer case (K1 compressor) and 32deg Cfor the winter case (K2 compressor)

This pump suction temperaturecauses pumping problems at less than

550 kgcu m which corresponds totemperatures greater than 32deg C

NGL yields As discussed the analyses looked at

several alternatives for determining theprocess with the best economics forNGL recovery

The first case looked at oil spiking ata maximum of 167 cu mday at ndash23deg CThe next case analyzed C

3+ recovery at

ndash23deg ndash27deg and ndash29deg C

Table 4 shows the best recoveries areat ndash29deg C The horsepower (4610 vs2200) however doubles for 16 addedrecovery (256 vs 215 cu m)

Table 4 shows the scenarios simu-lated to evaluate NGL yields for a120-MMscfd recycle gas throughput

To put NGL recovery in perspectivea calculation was made to determinethe theoretical maximum liquid recoverbased on a flow of 33912 millioncu mday (Table 5) The recovery is

256494 = 52

Economics Table 6 shows the cost for installed

refrigeration of 33912 million cu mday The costs are approximate andwere obtained from equipment packag-ers and project execution experience

Table 7 shows the revenue recoveredfrom NGL sales

Economics run from these costs andrevenues indicate that the project wouldpay back in 23 years and have a presentvalue of $73 million Payments were$3185 million

year and the eco-nomics assumeda 5year inter-est and no futurevalue

Also the eco-nomic analysis in-cluded a sensitiv-ity case for a lowermol recycle gas The gas analysis inthis case was from a mid-phase CO

2 in-

jection recycle gas (Table 8) containing

82 mol CO2 This gas is much richerthan the initial 92 CO

2 case

Table 9 shows the simulations forpredicting liquid recoveries with thesame plant configuration as for the pre-vious case The table shows the muchhigher recoveries that lead to a 1-yearpayback

Observations From this evaluation several points

were noted as follows

bull Depending on the refrigera-tion requirements the process shouldinclude dense-phase pumping if pos-sible Pumping has a lower cost thancompression if the process has enoughcooling in summer conditions

bull An added advantage of the refrig-eration is ethylene glycol dehydrationfor avoiding hydrates when chilling thegas to drop out liquids The recycle gastherefore is dehydrated without theneed of exotic piping for corrosion pro-

tection or a hydrate risk when depres-surizing or compressing in a centrifugalcompressor

bull The maximum recovery of themethane and ethane is insufficient tojustify methyldiethanolamine treat-ing to recover the 85000 cu mday

(3 MMscfd) C1 and C

2120 MMscfd of

recycle gasbull A further refinement of the pro-

cess simulations found that droppingthe separatorrsquos temperatures permit-ted recoveries of 176 cu mday of NGL

for ndash23deg C 215 cu mday for ndash27deg Cand 256 cu mday for ndash29deg C This isnot the case for the oil-spiking case It

appears that the oil will not pick up ad-ditional NGL at less than ndash23deg C

bull Installation of the NGL-recoveryprocess equipment should be in modu-larized increments The towers heatexchangers compressors and chillersdo not have efficient turndowns pastplusmn25

Hence the first phase of the proj-ect would instal l 2832 million cu mday (100 MMscfd) units followed byexpansion of the refrigeration compres-

sion stabilization and associated gascompression in 2832 million cu mdayincrements

The authorKenneth J Vargas (kvargasfalcon-edfcom) is president ofFalcon EDF Ltd and has morethan 35 years of experience indesigning upstream oil and gasfacilities Vargas specializes in process engineering and project management of onshore andoffshore projects He previously worked for IBM DuPont Eldorado Nuclear andGulf Vargas is a graduate of the US Air Force Academy and a registered professional engineer inWestern Canada Northwest Territories NorthDakota and Montana

RECYCLE GAS SENSITIVITY ANALYSIS Table 8

Component Mole fraction

H2 00000

He 00000N

2 00166

CO2 08125

H2S 00119

C1 00531

C2 00304C

3 00407

iC4 00154

C4 00073

C5 00121

NGL COSTS RECOVERED 82 CO2 Table 9

ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndash Rate ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashHydrocarbon cut $cu m times cu mday $day $ millionyear

C3 295 times 170 50000

C4rsquos 400 times 222 88800

C5+

345 times 206 71000ndashndashndashndashndashndashndash

Total 209000 7343

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 26

T E C H N O L O G Y

Kenneth J VargasFalcon EDF LtdCalgary

Refrigeration provides economic process forecovering NGL from CO2-EOR recycle gas

Removing and sellingNGL from the producedgas stream is one of theways CO

2 enhanced oil

recovery projects can im-prove their economics

These projects involve

gas recycling into thereservoir for maintaining pressure andimproving oil mobility The recycled

gas absorbs NGL in the reservoir that

can be recovered with a refrigerationprocess

The process involves refrigeratingthe gas and then separating stabilizingand recovering the NGL A companythen can market the recovered NGL asC

3 C

3+ C

4 and C

5+ or use it to spike the

crude

This article describes the followingaspects of NGL recovery

bull Process plant required for eachalternative (C

3+ fractionated liquids or

crude spiking)bull Liquid recoveries for varying

process conditions such as chiller final

temperaturebull Process property method selectionbull Process flow diagram with mate-

rial balancebull Approximate costs for

different process alternativesbull Economics of the alter-

nativesIt is best to install this

NGL recovery and refrigera-tion facility early in the CO

2

injection projectrsquos life

Process facility A typical CO

2-EOR NGL recovery

plant consists of the basic batterywith additional equipment to handleincreased produced gas and water asfollows

bull Inlet separation

Production

PROCESS FLOW Fig 1

Solution-gasinlet separator

Low-pressureseparator

1B

1A

Btu C2

De-C2Btu A

20

Mol sieve 13x

4

5

VPCHP de-C2

Comp de-C2 Btu de-C2

Gas-gas

Mix gas-gas

Chiller

R sep 1

R comp 1

Mix-economizer

Economizer

R4

HP R2HP R1

Btu ref

Condenser

R comp 2

VE

De-C2 - chiller

VF

Comp K1 Btu K1

Cooler K1

Gas-gas K1Splitgas-gas

Gas-gas K2

Comp K2 Btu K2

HP P2

HP P1

HP K2

HP K2

Cooler K2

Inj 120 MM

Reda 2

Reda 1

Mix inj

Btu C3 cooler

Btu C3 prod

C3 cooler

C3 sales

Btu C3 reboiler

De-C3

Btu C4 cooler

Btu C4 prod

C4 cooler

C5+ cooler

C4 sales

C5+ sales

Btu C5+ prod

Btu C4reboiler

De-C4

3

5A 5B 66A

7

21

1

9

22

2

10

1

2

9

2

3

2

3

7

16

10

19

25

24

R2

R3 R5

R9

R5A

R8 R7

R6

R1 R10

11

14

19

16

17

15

23

18

12

8 13

6B

Reprinted with revisions to format from the January 18 2010 edition of OIL amp GAS JOURNAL

Copyright 2010 by PennWell Corporation

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 36

bull Treating or emulsion breaking forfurther phase separation

bull Liquid products storage (oil andwater)

bull Produced gas compression for alloil and water separation pressures

bull Water injection

Fig 1 shows the process f low dia-gram and Table 1 the material balancefor 14 of the most relevant streams Theadditional units for CO

2-EOR include

bull Solution gas gas-gas precoolingexchanger

bull Solution gas chiller with propanerefrigeration plant

bull Low-temperature separator to feedrefrigerated liquids to the fractionationplant

bull Amine or mol-sieve liquid NGL

sweetening unitbull Lean-gas pump-suction exchang-

ers The case shown in Fig 1 has two170 million cu mday (60 MMscfd)trains

bull Two compressor trains of 170million cu mday each to boost thepressure from 1256 kPa-g (182 psig)to obtain a dense phase for 92 CO

2

gas The critical point is 7500 kPa-g(1090 psig) and the pressure selectedwas 8860 kPa-g (1285 psig) which is

in the dense-phase region Fig 2 showsthe phase envelope for the solution gas

bull Two centrifugal multiphasepumps one per train to boost the pres-sure to the injection pressure of 15400kPa-g or 2230 psig

The analysis looked at the following

three alternatives for liquids recovery1 Fractionation plant with de-eth-anizer depropanizer (sell C

3 product)

debutanizer (sell C4 and C

5+ product)

and distillation towers with ancillaries(aerial condensers reflux drums reboil-ers and product coolers)

2 De-ethanizer to produce C3+

product and sell

3 De-ethanizer to produce C3+

andspike the crude

Process design The design assumed a gas with a 15sp gr (436 molecular weight) Table 2shows the gas composition

This article presents the designfor only the 92 CO

2 gas content

with a full fractionation case becausethe recoveries for this case are moreconservative due to the high CO

2

concentration The base case includesa fractionation train because all otheralternatives are subsets of it The other

two alternatives either sell or spike thecrude with the C

3+

The analysis used the Peng-Robinsonproperty method for all the simula-tions

Table 3 summarizes the key streamsimulations of the process shown in

Fig 1

Propane refrigeration The refrigeration loop is in the up-

per right-hand corner of Fig 1 A gas-gas exchanger precools the plant inletsolution gas It uses cold gas off of thelow-temperature separator

After the gas-gas exchanger a chillerrefrigerates the gas to ndash29deg C Therefrigeration feeding the chiller on theshell side is a propane loop with an

economizer on the interstage of thepropane screw compressor

The schematic simplified thepropane loop as a two-stage refrigera-tion without a second chiller on thelast stage It shows only one chiller forsimplicity

A two-stage refrigeration loop wouldreduce the compression by 19 andthe condenser duty by 8 The mate-rial balance (Table 1) ref lects this Itshows the process requires 3940 hp

for the first stage(HP R1) and 670hp for the secondstage (HP R2) Thedesign corrects thetotal 4610 hp byndash19 for two-stage eff iciencyThus refrigera-tion compressionwould total about3750 hp

The propanerefrigeration con-denser would pro-vide 339 MMbtuhr times 091 or 31MMbtuhr

Note the pro-cess requires anethylene-glycolinjection loop fordehydrating the

Fig 2PHASE ENVELOPE

ndash140 ndash120 ndash100 ndash80 ndash60 ndash40 ndash20 0 20 40

P r e s s u r e

k P a - g

8000

7000

6000

5000

4000

3000

2000

1000

0

B - cricondenbar

T - cricondentherm

C - critical point

Hydrate curve

vf 0000

vf 1000

BTC

Temperature degC

8102019 Refrigeration for NGL Recovery

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T E C H N O L O G Y

MATERIAL BALANCE

Stream 2 13 18 Btu-Ref C3-sales C

4-sales C

5+-sales

Vapor fraction 07994 10000 00000 20000 10000 00000 00000Temperature degC ndash290000 380000 320000 00000 430000 430000 430000Pressure kPa-g 13248419 87929143 87929143 00000 11948897 5743614 6088352Flow MMscfd 1200000 590204 589976 00000 03459 07178 09172Liquid flow cu mday 79048550 38248715 38233949 00000 360220 861025 1344131Liquid flow bd 497199569 240576765 240483892 00000 2265710 5415676 8454313Mole weight 436115 432465 432465 441822 578919 791946Energy btuhr 250311E+07 101560E+07 449534E+06 338568E+07 2494771803 150502347 ndash44669296Energy hp 98376048 39914675 17667385 133062406 980482 59150 ndash17556H

2 mole fraction 0000000 0000000 0000000 0000000 0000000 0000000

N2 mole fraction 0008801 0008949 0008949 0000000 0000000 0000000

CO2 mole fraction 0918692 0934119 0934119 0000286 0000000 0000000

H2S mole fraction 0009101 0009245 0009245 0000000 0000000 0000000

C1 mole fraction 0015102 0015355 0015355 0000000 0000000 0000000

C2 mole fraction 0013501 0013728 0013728 0000043 0000000 0000000

C3 mole fraction 0016002 0013220 0013220 0993553 0022790 0000000

iC4 mole fraction 0002900 0001526 0001526 0005000 0231329 0000198

nC4 mole fraction 0007401 0002945 0002945 0001118 0739641 0009998

C5+

mole fraction 0008500 0000912 0000912 0000000 0006240 0989804

GAS COMPOSITION Table 2

Component Mole fraction

H2 00000

He 00000N

2 00088

CO2 09187

H2S 00091

C1 00151

C2 00135

C3 00160

iC4 00029

C4 00074

iC5+

00085

PROCESS CONDITIONS Table 3

Inlet parametersLevel of CO

2 92

Inlet separator pressure kPa-g 1380Inlet separator temperature degC 27 Inlet gas-gas exchangerTubeside in-out temperature degC 27ndash1Shellside in-out temperautre degC ndash28 3Duty MJhr MMbtuhr 84 80 Propane refrigerationChiller in-out temperature degC ndash1ndash29Chiller duty MJhr MMbtuhr 7874Propane compressor hp 4610Propane condenser duty MJhr MMbtuhr 35831 De-ethanizerFlow in cu mday 1689Tower diameter mm 1100Reboiler duty MJhr MMbtuhr 15815Liquid produced cu mday 257Overhead gas 1000 cu mday 625Recompressor hp 100 Inlet gas-gas K1-K2 exchangersDuty MJhr MMbtuhr 40-3838-35Minimum tempererature out degC 38 (gas) 32 (liquid)

Pump maximum flow cu m day MMscfd 382459Pump flow specific gravity 04-055Pump hp 1454-1064 DepropanizerFlow in cu mday 256Tower diamater approximate mm 610Reboiler duty MJhr MMbtuhr 1918Liquid produced cu mday 220Overhead gas 1000 cu mday 98 DebutanizerFlow in cu mday 86Tower diamater approximate mm 310Reboiler duty MJhr MMbtuhr 3129Liquid produced cu mday 134Overhead gas 1000 cu mday 203

gas to avoid gas hydrates after cooling

For simplicity Fig 1 does not showEG injection

NGL stabilization The lower left-hand portion of Fig

1 shows the fractionation plant Therefrigerated gas goes to a low-tempera-ture separator which separates the liq-uids that enter the fractionation plant

The fractionation plant has three dis-tillation towers The first is a de-etha-nizer (de-C

2 in Fig 1) with a reboiler

as the bottom stage The deethanizedliquids go to the depropanizer (de-C

3) unit consisting of tower overhead

condenser reflux-drum and bottomsreboiler

The specification sales propane isStream C

3 sales in Fig 1

The depropanizer bottoms go toa similar tower reflux and reboilerdistillation column This last distillationcolumn is a debutanizer (de-C

4)

The tower overheads go to butanesales (Stream C

4 sales) The column

bottoms are the light gasoline sales orC

5+ salesThe de-ethanizer overhead gas after

recompression mixes with the refriger-ated gas off the low-temperature sepa-rator The streamthen goes to thegas-gas shell sideof the exchangerand subsequentlythe reciprocatingcompressor-pump

tandem combina-tion

Fig 1 shows anaerial cooler afterthe stabilizer over-head of the re-compressor (compde-C

2) however it

is not required

Desulfurization Fig 1 shows

the sweetening ofthe liquids witha mol-sieve unitoperation on thebottom of the de-ethanizer reboilerproduct stream A13x Grade Z10-03mol-sieve unit or aliquid amine con-tactor can sweeten

the stabilized liquid NGLPreferable is a mol-sieve unit because

it has a dry system that can be regener-ated with hot fuel gas A typical 13xGrade Z10-03 mol-sieve unit has two orthree contactors to ensure 24-hr sweet-ening of sour NGL

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 56

Gas-gas cooling The solution gas from the inlet

gas-gas exchanger enters two streamsOne is the inlet of the tube side of theK1 exchanger (Stream 8 not shown inTable 1) and the other is the tube sideof K2 exchanger (Stream 9 not shownin Table 1) About half of the 3392 mil-lion cu mday enters each exchanger

The exit gas from the tube side ofthese exchangers feeds the K1 and K2

compressors The compressor discharg-es back into the shell side of K1 and K2

exchangers Utilizing parallel K1 andK2 exchangers ensures that the suctionstreamsrsquo temperatures feeding pumpsReda 1 and 2 are as low as possible

A temperature cooler than 33deg Cis optimal to ensure the multistagecentrifugal pumpsrsquo lowest horsepowerdraw In our case this is 1010 hp at15200 kPa-g (2200 psig)

The K1 com-pressor trainis for summerconditions or am-bient temperature3deg C cooler thanthe compressor

coolerrsquos dischargeof 43deg C Thelowest achievabletemperature afterthe K1 exchangeris 38deg C

The corre-sponding pumprequires 1454hp vs the K2

exchangerrsquos discharge of 32deg C whichrequires 1064 hp This increases the

horsepower by 27Because the simulations are mainly

for refrigeration cooling and stabiliza-tion Fig 1 shows the solution-gas in-jection compressors with a single stageabove the K1 and K2 exchangers

Pump performance The process cools the streams from

NGL RECOVERIES Table 4

Liquid produc- Refrigera- Refrigeration tion increaseCase Type tion hp temp 983151C cu mday

1 Spike oil 1630 ndash23 1672 Spike oil 2200 ndash27 167 (no change)3 NGL C

3+ 1600 ndash23 176

4 NGL C3+

2200 ndash27 2155 Fractionate 4610 ndash29 C

3= 36 C

4= 86

C5+

= 134 Case 5 sum

of C3 C

4 and C

5+ 256

THEORETICAL MAXIMUM NGL RECOVERY Table 5

Maximum liquid Ideal gas NGL liquid yieldCom- Mole Volume conversion cu mdayponent fraction Mscfd cu ftgal (bd)

N2 00088 1056

CO2 09186 110232

H2S 00091 1092

C1 00151 1812

C2 00135 1620

C3 0016 1920 3637 200 (125692)

iC4 00029 348 3064 43 (27042)

nC4 00074 888 3179 106 (66508)

iC5 0002 240 2738 332 (20870)

nC5 00024 288 2767 394 (24782)

C6+

00042 504 2616 73 (45872) ndashndashndashndashndashndashndashndash ndashndashndashndashndashndashndashndashndashndashndashndashndash Total 120000 494 (310766)

From Table 4 the total theoretical NGL recovery from the recycle gas stream is 494

cu m day Our fractionation is recovering 256 cu mday or 52

FRACTIONATION PLANT COST Table 6

MillionCapital costs for 3912 million cu mday (120 MMscfd) plant $

300 cu mday de-ethanizer complete (includes recompression reboiler) 225000 hp propane refrigeration complete unit 90Two 9 GJhr heat exchangers (chiller and gas-gas) 13Two 27-40 GJhr heat exchangers (gas-gas K1 and gas-gas K2) 07NGL recovery depropanizer-butanizer complete 30Refrigeration major electrical and mechanical equipment 30NGL mol sieve sweetening 10Low-termperature separator 08Piping racks cable trays insulation buildings and other consumables 35Total all equipment and materials 245Installed costs (15 times equipment and materials) 368Contingency (20) 123 ndashndashndashndashndash Total 736

Costs from skid vendors

NGL COSTS RECOVERED Table 7

ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndash Rate ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashHydrocarbon cut $cu m times cu mday $day $ millionyear

C3 295 times 36 11000

C4rsquos 400 times 86 34000

C5+

345 times 134 46000ndashndashndashndashndashndashndash

Total 91000 3185

The crude spiking alternative gives lower yields because much of the NGL flash aftermixing The costs recovered are $345cu m times 167 = $58000day or $20165 million year

Table 1

HP-P1 HP-P2 HP-R1 HP-R2 Inj -120 MM R3 Solution gas

20000 20000 20000 20000 10000 10000 1000000000 00000 00000 00000 623072 ndash329945 27000000000 00000 00000 00000 154118830 503597 1380000000000 00000 00000 00000 1180180 402712 120000000000 00000 00000 00000 76482664 41872170 7904855000000 00000 00000 00000 481060657 263367577 497199569

432465 438165 436115370150E+06 270696E+06 100267E+07 169455E+06 210598E+07 212384E+07 543272E+07

14547471 10638752 39406609 6659849 82768284 83470270 213513939 0000000 0000000 0000000 0008949 0000000 0008801 0934119 0000000 0918692 0009245 0000000 0009101 0015355 0000000 0015102 0013728 0020000 0013501 0013220 0980000 0016002 0001526 0000000 0002900 0002945 0000000 0007401 0000912 0000000 0008500

8102019 Refrigeration for NGL Recovery

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T E C H N O L O G Y

the K1 and 2 last stage discharge intothe Reda pumps to 38deg C for the sum-mer case (K1 compressor) and 32deg Cfor the winter case (K2 compressor)

This pump suction temperaturecauses pumping problems at less than

550 kgcu m which corresponds totemperatures greater than 32deg C

NGL yields As discussed the analyses looked at

several alternatives for determining theprocess with the best economics forNGL recovery

The first case looked at oil spiking ata maximum of 167 cu mday at ndash23deg CThe next case analyzed C

3+ recovery at

ndash23deg ndash27deg and ndash29deg C

Table 4 shows the best recoveries areat ndash29deg C The horsepower (4610 vs2200) however doubles for 16 addedrecovery (256 vs 215 cu m)

Table 4 shows the scenarios simu-lated to evaluate NGL yields for a120-MMscfd recycle gas throughput

To put NGL recovery in perspectivea calculation was made to determinethe theoretical maximum liquid recoverbased on a flow of 33912 millioncu mday (Table 5) The recovery is

256494 = 52

Economics Table 6 shows the cost for installed

refrigeration of 33912 million cu mday The costs are approximate andwere obtained from equipment packag-ers and project execution experience

Table 7 shows the revenue recoveredfrom NGL sales

Economics run from these costs andrevenues indicate that the project wouldpay back in 23 years and have a presentvalue of $73 million Payments were$3185 million

year and the eco-nomics assumeda 5year inter-est and no futurevalue

Also the eco-nomic analysis in-cluded a sensitiv-ity case for a lowermol recycle gas The gas analysis inthis case was from a mid-phase CO

2 in-

jection recycle gas (Table 8) containing

82 mol CO2 This gas is much richerthan the initial 92 CO

2 case

Table 9 shows the simulations forpredicting liquid recoveries with thesame plant configuration as for the pre-vious case The table shows the muchhigher recoveries that lead to a 1-yearpayback

Observations From this evaluation several points

were noted as follows

bull Depending on the refrigera-tion requirements the process shouldinclude dense-phase pumping if pos-sible Pumping has a lower cost thancompression if the process has enoughcooling in summer conditions

bull An added advantage of the refrig-eration is ethylene glycol dehydrationfor avoiding hydrates when chilling thegas to drop out liquids The recycle gastherefore is dehydrated without theneed of exotic piping for corrosion pro-

tection or a hydrate risk when depres-surizing or compressing in a centrifugalcompressor

bull The maximum recovery of themethane and ethane is insufficient tojustify methyldiethanolamine treat-ing to recover the 85000 cu mday

(3 MMscfd) C1 and C

2120 MMscfd of

recycle gasbull A further refinement of the pro-

cess simulations found that droppingthe separatorrsquos temperatures permit-ted recoveries of 176 cu mday of NGL

for ndash23deg C 215 cu mday for ndash27deg Cand 256 cu mday for ndash29deg C This isnot the case for the oil-spiking case It

appears that the oil will not pick up ad-ditional NGL at less than ndash23deg C

bull Installation of the NGL-recoveryprocess equipment should be in modu-larized increments The towers heatexchangers compressors and chillersdo not have efficient turndowns pastplusmn25

Hence the first phase of the proj-ect would instal l 2832 million cu mday (100 MMscfd) units followed byexpansion of the refrigeration compres-

sion stabilization and associated gascompression in 2832 million cu mdayincrements

The authorKenneth J Vargas (kvargasfalcon-edfcom) is president ofFalcon EDF Ltd and has morethan 35 years of experience indesigning upstream oil and gasfacilities Vargas specializes in process engineering and project management of onshore andoffshore projects He previously worked for IBM DuPont Eldorado Nuclear andGulf Vargas is a graduate of the US Air Force Academy and a registered professional engineer inWestern Canada Northwest Territories NorthDakota and Montana

RECYCLE GAS SENSITIVITY ANALYSIS Table 8

Component Mole fraction

H2 00000

He 00000N

2 00166

CO2 08125

H2S 00119

C1 00531

C2 00304C

3 00407

iC4 00154

C4 00073

C5 00121

NGL COSTS RECOVERED 82 CO2 Table 9

ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndash Rate ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashHydrocarbon cut $cu m times cu mday $day $ millionyear

C3 295 times 170 50000

C4rsquos 400 times 222 88800

C5+

345 times 206 71000ndashndashndashndashndashndashndash

Total 209000 7343

8102019 Refrigeration for NGL Recovery

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bull Treating or emulsion breaking forfurther phase separation

bull Liquid products storage (oil andwater)

bull Produced gas compression for alloil and water separation pressures

bull Water injection

Fig 1 shows the process f low dia-gram and Table 1 the material balancefor 14 of the most relevant streams Theadditional units for CO

2-EOR include

bull Solution gas gas-gas precoolingexchanger

bull Solution gas chiller with propanerefrigeration plant

bull Low-temperature separator to feedrefrigerated liquids to the fractionationplant

bull Amine or mol-sieve liquid NGL

sweetening unitbull Lean-gas pump-suction exchang-

ers The case shown in Fig 1 has two170 million cu mday (60 MMscfd)trains

bull Two compressor trains of 170million cu mday each to boost thepressure from 1256 kPa-g (182 psig)to obtain a dense phase for 92 CO

2

gas The critical point is 7500 kPa-g(1090 psig) and the pressure selectedwas 8860 kPa-g (1285 psig) which is

in the dense-phase region Fig 2 showsthe phase envelope for the solution gas

bull Two centrifugal multiphasepumps one per train to boost the pres-sure to the injection pressure of 15400kPa-g or 2230 psig

The analysis looked at the following

three alternatives for liquids recovery1 Fractionation plant with de-eth-anizer depropanizer (sell C

3 product)

debutanizer (sell C4 and C

5+ product)

and distillation towers with ancillaries(aerial condensers reflux drums reboil-ers and product coolers)

2 De-ethanizer to produce C3+

product and sell

3 De-ethanizer to produce C3+

andspike the crude

Process design The design assumed a gas with a 15sp gr (436 molecular weight) Table 2shows the gas composition

This article presents the designfor only the 92 CO

2 gas content

with a full fractionation case becausethe recoveries for this case are moreconservative due to the high CO

2

concentration The base case includesa fractionation train because all otheralternatives are subsets of it The other

two alternatives either sell or spike thecrude with the C

3+

The analysis used the Peng-Robinsonproperty method for all the simula-tions

Table 3 summarizes the key streamsimulations of the process shown in

Fig 1

Propane refrigeration The refrigeration loop is in the up-

per right-hand corner of Fig 1 A gas-gas exchanger precools the plant inletsolution gas It uses cold gas off of thelow-temperature separator

After the gas-gas exchanger a chillerrefrigerates the gas to ndash29deg C Therefrigeration feeding the chiller on theshell side is a propane loop with an

economizer on the interstage of thepropane screw compressor

The schematic simplified thepropane loop as a two-stage refrigera-tion without a second chiller on thelast stage It shows only one chiller forsimplicity

A two-stage refrigeration loop wouldreduce the compression by 19 andthe condenser duty by 8 The mate-rial balance (Table 1) ref lects this Itshows the process requires 3940 hp

for the first stage(HP R1) and 670hp for the secondstage (HP R2) Thedesign corrects thetotal 4610 hp byndash19 for two-stage eff iciencyThus refrigera-tion compressionwould total about3750 hp

The propanerefrigeration con-denser would pro-vide 339 MMbtuhr times 091 or 31MMbtuhr

Note the pro-cess requires anethylene-glycolinjection loop fordehydrating the

Fig 2PHASE ENVELOPE

ndash140 ndash120 ndash100 ndash80 ndash60 ndash40 ndash20 0 20 40

P r e s s u r e

k P a - g

8000

7000

6000

5000

4000

3000

2000

1000

0

B - cricondenbar

T - cricondentherm

C - critical point

Hydrate curve

vf 0000

vf 1000

BTC

Temperature degC

8102019 Refrigeration for NGL Recovery

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T E C H N O L O G Y

MATERIAL BALANCE

Stream 2 13 18 Btu-Ref C3-sales C

4-sales C

5+-sales

Vapor fraction 07994 10000 00000 20000 10000 00000 00000Temperature degC ndash290000 380000 320000 00000 430000 430000 430000Pressure kPa-g 13248419 87929143 87929143 00000 11948897 5743614 6088352Flow MMscfd 1200000 590204 589976 00000 03459 07178 09172Liquid flow cu mday 79048550 38248715 38233949 00000 360220 861025 1344131Liquid flow bd 497199569 240576765 240483892 00000 2265710 5415676 8454313Mole weight 436115 432465 432465 441822 578919 791946Energy btuhr 250311E+07 101560E+07 449534E+06 338568E+07 2494771803 150502347 ndash44669296Energy hp 98376048 39914675 17667385 133062406 980482 59150 ndash17556H

2 mole fraction 0000000 0000000 0000000 0000000 0000000 0000000

N2 mole fraction 0008801 0008949 0008949 0000000 0000000 0000000

CO2 mole fraction 0918692 0934119 0934119 0000286 0000000 0000000

H2S mole fraction 0009101 0009245 0009245 0000000 0000000 0000000

C1 mole fraction 0015102 0015355 0015355 0000000 0000000 0000000

C2 mole fraction 0013501 0013728 0013728 0000043 0000000 0000000

C3 mole fraction 0016002 0013220 0013220 0993553 0022790 0000000

iC4 mole fraction 0002900 0001526 0001526 0005000 0231329 0000198

nC4 mole fraction 0007401 0002945 0002945 0001118 0739641 0009998

C5+

mole fraction 0008500 0000912 0000912 0000000 0006240 0989804

GAS COMPOSITION Table 2

Component Mole fraction

H2 00000

He 00000N

2 00088

CO2 09187

H2S 00091

C1 00151

C2 00135

C3 00160

iC4 00029

C4 00074

iC5+

00085

PROCESS CONDITIONS Table 3

Inlet parametersLevel of CO

2 92

Inlet separator pressure kPa-g 1380Inlet separator temperature degC 27 Inlet gas-gas exchangerTubeside in-out temperature degC 27ndash1Shellside in-out temperautre degC ndash28 3Duty MJhr MMbtuhr 84 80 Propane refrigerationChiller in-out temperature degC ndash1ndash29Chiller duty MJhr MMbtuhr 7874Propane compressor hp 4610Propane condenser duty MJhr MMbtuhr 35831 De-ethanizerFlow in cu mday 1689Tower diameter mm 1100Reboiler duty MJhr MMbtuhr 15815Liquid produced cu mday 257Overhead gas 1000 cu mday 625Recompressor hp 100 Inlet gas-gas K1-K2 exchangersDuty MJhr MMbtuhr 40-3838-35Minimum tempererature out degC 38 (gas) 32 (liquid)

Pump maximum flow cu m day MMscfd 382459Pump flow specific gravity 04-055Pump hp 1454-1064 DepropanizerFlow in cu mday 256Tower diamater approximate mm 610Reboiler duty MJhr MMbtuhr 1918Liquid produced cu mday 220Overhead gas 1000 cu mday 98 DebutanizerFlow in cu mday 86Tower diamater approximate mm 310Reboiler duty MJhr MMbtuhr 3129Liquid produced cu mday 134Overhead gas 1000 cu mday 203

gas to avoid gas hydrates after cooling

For simplicity Fig 1 does not showEG injection

NGL stabilization The lower left-hand portion of Fig

1 shows the fractionation plant Therefrigerated gas goes to a low-tempera-ture separator which separates the liq-uids that enter the fractionation plant

The fractionation plant has three dis-tillation towers The first is a de-etha-nizer (de-C

2 in Fig 1) with a reboiler

as the bottom stage The deethanizedliquids go to the depropanizer (de-C

3) unit consisting of tower overhead

condenser reflux-drum and bottomsreboiler

The specification sales propane isStream C

3 sales in Fig 1

The depropanizer bottoms go toa similar tower reflux and reboilerdistillation column This last distillationcolumn is a debutanizer (de-C

4)

The tower overheads go to butanesales (Stream C

4 sales) The column

bottoms are the light gasoline sales orC

5+ salesThe de-ethanizer overhead gas after

recompression mixes with the refriger-ated gas off the low-temperature sepa-rator The streamthen goes to thegas-gas shell sideof the exchangerand subsequentlythe reciprocatingcompressor-pump

tandem combina-tion

Fig 1 shows anaerial cooler afterthe stabilizer over-head of the re-compressor (compde-C

2) however it

is not required

Desulfurization Fig 1 shows

the sweetening ofthe liquids witha mol-sieve unitoperation on thebottom of the de-ethanizer reboilerproduct stream A13x Grade Z10-03mol-sieve unit or aliquid amine con-tactor can sweeten

the stabilized liquid NGLPreferable is a mol-sieve unit because

it has a dry system that can be regener-ated with hot fuel gas A typical 13xGrade Z10-03 mol-sieve unit has two orthree contactors to ensure 24-hr sweet-ening of sour NGL

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 56

Gas-gas cooling The solution gas from the inlet

gas-gas exchanger enters two streamsOne is the inlet of the tube side of theK1 exchanger (Stream 8 not shown inTable 1) and the other is the tube sideof K2 exchanger (Stream 9 not shownin Table 1) About half of the 3392 mil-lion cu mday enters each exchanger

The exit gas from the tube side ofthese exchangers feeds the K1 and K2

compressors The compressor discharg-es back into the shell side of K1 and K2

exchangers Utilizing parallel K1 andK2 exchangers ensures that the suctionstreamsrsquo temperatures feeding pumpsReda 1 and 2 are as low as possible

A temperature cooler than 33deg Cis optimal to ensure the multistagecentrifugal pumpsrsquo lowest horsepowerdraw In our case this is 1010 hp at15200 kPa-g (2200 psig)

The K1 com-pressor trainis for summerconditions or am-bient temperature3deg C cooler thanthe compressor

coolerrsquos dischargeof 43deg C Thelowest achievabletemperature afterthe K1 exchangeris 38deg C

The corre-sponding pumprequires 1454hp vs the K2

exchangerrsquos discharge of 32deg C whichrequires 1064 hp This increases the

horsepower by 27Because the simulations are mainly

for refrigeration cooling and stabiliza-tion Fig 1 shows the solution-gas in-jection compressors with a single stageabove the K1 and K2 exchangers

Pump performance The process cools the streams from

NGL RECOVERIES Table 4

Liquid produc- Refrigera- Refrigeration tion increaseCase Type tion hp temp 983151C cu mday

1 Spike oil 1630 ndash23 1672 Spike oil 2200 ndash27 167 (no change)3 NGL C

3+ 1600 ndash23 176

4 NGL C3+

2200 ndash27 2155 Fractionate 4610 ndash29 C

3= 36 C

4= 86

C5+

= 134 Case 5 sum

of C3 C

4 and C

5+ 256

THEORETICAL MAXIMUM NGL RECOVERY Table 5

Maximum liquid Ideal gas NGL liquid yieldCom- Mole Volume conversion cu mdayponent fraction Mscfd cu ftgal (bd)

N2 00088 1056

CO2 09186 110232

H2S 00091 1092

C1 00151 1812

C2 00135 1620

C3 0016 1920 3637 200 (125692)

iC4 00029 348 3064 43 (27042)

nC4 00074 888 3179 106 (66508)

iC5 0002 240 2738 332 (20870)

nC5 00024 288 2767 394 (24782)

C6+

00042 504 2616 73 (45872) ndashndashndashndashndashndashndashndash ndashndashndashndashndashndashndashndashndashndashndashndashndash Total 120000 494 (310766)

From Table 4 the total theoretical NGL recovery from the recycle gas stream is 494

cu m day Our fractionation is recovering 256 cu mday or 52

FRACTIONATION PLANT COST Table 6

MillionCapital costs for 3912 million cu mday (120 MMscfd) plant $

300 cu mday de-ethanizer complete (includes recompression reboiler) 225000 hp propane refrigeration complete unit 90Two 9 GJhr heat exchangers (chiller and gas-gas) 13Two 27-40 GJhr heat exchangers (gas-gas K1 and gas-gas K2) 07NGL recovery depropanizer-butanizer complete 30Refrigeration major electrical and mechanical equipment 30NGL mol sieve sweetening 10Low-termperature separator 08Piping racks cable trays insulation buildings and other consumables 35Total all equipment and materials 245Installed costs (15 times equipment and materials) 368Contingency (20) 123 ndashndashndashndashndash Total 736

Costs from skid vendors

NGL COSTS RECOVERED Table 7

ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndash Rate ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashHydrocarbon cut $cu m times cu mday $day $ millionyear

C3 295 times 36 11000

C4rsquos 400 times 86 34000

C5+

345 times 134 46000ndashndashndashndashndashndashndash

Total 91000 3185

The crude spiking alternative gives lower yields because much of the NGL flash aftermixing The costs recovered are $345cu m times 167 = $58000day or $20165 million year

Table 1

HP-P1 HP-P2 HP-R1 HP-R2 Inj -120 MM R3 Solution gas

20000 20000 20000 20000 10000 10000 1000000000 00000 00000 00000 623072 ndash329945 27000000000 00000 00000 00000 154118830 503597 1380000000000 00000 00000 00000 1180180 402712 120000000000 00000 00000 00000 76482664 41872170 7904855000000 00000 00000 00000 481060657 263367577 497199569

432465 438165 436115370150E+06 270696E+06 100267E+07 169455E+06 210598E+07 212384E+07 543272E+07

14547471 10638752 39406609 6659849 82768284 83470270 213513939 0000000 0000000 0000000 0008949 0000000 0008801 0934119 0000000 0918692 0009245 0000000 0009101 0015355 0000000 0015102 0013728 0020000 0013501 0013220 0980000 0016002 0001526 0000000 0002900 0002945 0000000 0007401 0000912 0000000 0008500

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 66

T E C H N O L O G Y

the K1 and 2 last stage discharge intothe Reda pumps to 38deg C for the sum-mer case (K1 compressor) and 32deg Cfor the winter case (K2 compressor)

This pump suction temperaturecauses pumping problems at less than

550 kgcu m which corresponds totemperatures greater than 32deg C

NGL yields As discussed the analyses looked at

several alternatives for determining theprocess with the best economics forNGL recovery

The first case looked at oil spiking ata maximum of 167 cu mday at ndash23deg CThe next case analyzed C

3+ recovery at

ndash23deg ndash27deg and ndash29deg C

Table 4 shows the best recoveries areat ndash29deg C The horsepower (4610 vs2200) however doubles for 16 addedrecovery (256 vs 215 cu m)

Table 4 shows the scenarios simu-lated to evaluate NGL yields for a120-MMscfd recycle gas throughput

To put NGL recovery in perspectivea calculation was made to determinethe theoretical maximum liquid recoverbased on a flow of 33912 millioncu mday (Table 5) The recovery is

256494 = 52

Economics Table 6 shows the cost for installed

refrigeration of 33912 million cu mday The costs are approximate andwere obtained from equipment packag-ers and project execution experience

Table 7 shows the revenue recoveredfrom NGL sales

Economics run from these costs andrevenues indicate that the project wouldpay back in 23 years and have a presentvalue of $73 million Payments were$3185 million

year and the eco-nomics assumeda 5year inter-est and no futurevalue

Also the eco-nomic analysis in-cluded a sensitiv-ity case for a lowermol recycle gas The gas analysis inthis case was from a mid-phase CO

2 in-

jection recycle gas (Table 8) containing

82 mol CO2 This gas is much richerthan the initial 92 CO

2 case

Table 9 shows the simulations forpredicting liquid recoveries with thesame plant configuration as for the pre-vious case The table shows the muchhigher recoveries that lead to a 1-yearpayback

Observations From this evaluation several points

were noted as follows

bull Depending on the refrigera-tion requirements the process shouldinclude dense-phase pumping if pos-sible Pumping has a lower cost thancompression if the process has enoughcooling in summer conditions

bull An added advantage of the refrig-eration is ethylene glycol dehydrationfor avoiding hydrates when chilling thegas to drop out liquids The recycle gastherefore is dehydrated without theneed of exotic piping for corrosion pro-

tection or a hydrate risk when depres-surizing or compressing in a centrifugalcompressor

bull The maximum recovery of themethane and ethane is insufficient tojustify methyldiethanolamine treat-ing to recover the 85000 cu mday

(3 MMscfd) C1 and C

2120 MMscfd of

recycle gasbull A further refinement of the pro-

cess simulations found that droppingthe separatorrsquos temperatures permit-ted recoveries of 176 cu mday of NGL

for ndash23deg C 215 cu mday for ndash27deg Cand 256 cu mday for ndash29deg C This isnot the case for the oil-spiking case It

appears that the oil will not pick up ad-ditional NGL at less than ndash23deg C

bull Installation of the NGL-recoveryprocess equipment should be in modu-larized increments The towers heatexchangers compressors and chillersdo not have efficient turndowns pastplusmn25

Hence the first phase of the proj-ect would instal l 2832 million cu mday (100 MMscfd) units followed byexpansion of the refrigeration compres-

sion stabilization and associated gascompression in 2832 million cu mdayincrements

The authorKenneth J Vargas (kvargasfalcon-edfcom) is president ofFalcon EDF Ltd and has morethan 35 years of experience indesigning upstream oil and gasfacilities Vargas specializes in process engineering and project management of onshore andoffshore projects He previously worked for IBM DuPont Eldorado Nuclear andGulf Vargas is a graduate of the US Air Force Academy and a registered professional engineer inWestern Canada Northwest Territories NorthDakota and Montana

RECYCLE GAS SENSITIVITY ANALYSIS Table 8

Component Mole fraction

H2 00000

He 00000N

2 00166

CO2 08125

H2S 00119

C1 00531

C2 00304C

3 00407

iC4 00154

C4 00073

C5 00121

NGL COSTS RECOVERED 82 CO2 Table 9

ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndash Rate ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashHydrocarbon cut $cu m times cu mday $day $ millionyear

C3 295 times 170 50000

C4rsquos 400 times 222 88800

C5+

345 times 206 71000ndashndashndashndashndashndashndash

Total 209000 7343

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 46

T E C H N O L O G Y

MATERIAL BALANCE

Stream 2 13 18 Btu-Ref C3-sales C

4-sales C

5+-sales

Vapor fraction 07994 10000 00000 20000 10000 00000 00000Temperature degC ndash290000 380000 320000 00000 430000 430000 430000Pressure kPa-g 13248419 87929143 87929143 00000 11948897 5743614 6088352Flow MMscfd 1200000 590204 589976 00000 03459 07178 09172Liquid flow cu mday 79048550 38248715 38233949 00000 360220 861025 1344131Liquid flow bd 497199569 240576765 240483892 00000 2265710 5415676 8454313Mole weight 436115 432465 432465 441822 578919 791946Energy btuhr 250311E+07 101560E+07 449534E+06 338568E+07 2494771803 150502347 ndash44669296Energy hp 98376048 39914675 17667385 133062406 980482 59150 ndash17556H

2 mole fraction 0000000 0000000 0000000 0000000 0000000 0000000

N2 mole fraction 0008801 0008949 0008949 0000000 0000000 0000000

CO2 mole fraction 0918692 0934119 0934119 0000286 0000000 0000000

H2S mole fraction 0009101 0009245 0009245 0000000 0000000 0000000

C1 mole fraction 0015102 0015355 0015355 0000000 0000000 0000000

C2 mole fraction 0013501 0013728 0013728 0000043 0000000 0000000

C3 mole fraction 0016002 0013220 0013220 0993553 0022790 0000000

iC4 mole fraction 0002900 0001526 0001526 0005000 0231329 0000198

nC4 mole fraction 0007401 0002945 0002945 0001118 0739641 0009998

C5+

mole fraction 0008500 0000912 0000912 0000000 0006240 0989804

GAS COMPOSITION Table 2

Component Mole fraction

H2 00000

He 00000N

2 00088

CO2 09187

H2S 00091

C1 00151

C2 00135

C3 00160

iC4 00029

C4 00074

iC5+

00085

PROCESS CONDITIONS Table 3

Inlet parametersLevel of CO

2 92

Inlet separator pressure kPa-g 1380Inlet separator temperature degC 27 Inlet gas-gas exchangerTubeside in-out temperature degC 27ndash1Shellside in-out temperautre degC ndash28 3Duty MJhr MMbtuhr 84 80 Propane refrigerationChiller in-out temperature degC ndash1ndash29Chiller duty MJhr MMbtuhr 7874Propane compressor hp 4610Propane condenser duty MJhr MMbtuhr 35831 De-ethanizerFlow in cu mday 1689Tower diameter mm 1100Reboiler duty MJhr MMbtuhr 15815Liquid produced cu mday 257Overhead gas 1000 cu mday 625Recompressor hp 100 Inlet gas-gas K1-K2 exchangersDuty MJhr MMbtuhr 40-3838-35Minimum tempererature out degC 38 (gas) 32 (liquid)

Pump maximum flow cu m day MMscfd 382459Pump flow specific gravity 04-055Pump hp 1454-1064 DepropanizerFlow in cu mday 256Tower diamater approximate mm 610Reboiler duty MJhr MMbtuhr 1918Liquid produced cu mday 220Overhead gas 1000 cu mday 98 DebutanizerFlow in cu mday 86Tower diamater approximate mm 310Reboiler duty MJhr MMbtuhr 3129Liquid produced cu mday 134Overhead gas 1000 cu mday 203

gas to avoid gas hydrates after cooling

For simplicity Fig 1 does not showEG injection

NGL stabilization The lower left-hand portion of Fig

1 shows the fractionation plant Therefrigerated gas goes to a low-tempera-ture separator which separates the liq-uids that enter the fractionation plant

The fractionation plant has three dis-tillation towers The first is a de-etha-nizer (de-C

2 in Fig 1) with a reboiler

as the bottom stage The deethanizedliquids go to the depropanizer (de-C

3) unit consisting of tower overhead

condenser reflux-drum and bottomsreboiler

The specification sales propane isStream C

3 sales in Fig 1

The depropanizer bottoms go toa similar tower reflux and reboilerdistillation column This last distillationcolumn is a debutanizer (de-C

4)

The tower overheads go to butanesales (Stream C

4 sales) The column

bottoms are the light gasoline sales orC

5+ salesThe de-ethanizer overhead gas after

recompression mixes with the refriger-ated gas off the low-temperature sepa-rator The streamthen goes to thegas-gas shell sideof the exchangerand subsequentlythe reciprocatingcompressor-pump

tandem combina-tion

Fig 1 shows anaerial cooler afterthe stabilizer over-head of the re-compressor (compde-C

2) however it

is not required

Desulfurization Fig 1 shows

the sweetening ofthe liquids witha mol-sieve unitoperation on thebottom of the de-ethanizer reboilerproduct stream A13x Grade Z10-03mol-sieve unit or aliquid amine con-tactor can sweeten

the stabilized liquid NGLPreferable is a mol-sieve unit because

it has a dry system that can be regener-ated with hot fuel gas A typical 13xGrade Z10-03 mol-sieve unit has two orthree contactors to ensure 24-hr sweet-ening of sour NGL

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 56

Gas-gas cooling The solution gas from the inlet

gas-gas exchanger enters two streamsOne is the inlet of the tube side of theK1 exchanger (Stream 8 not shown inTable 1) and the other is the tube sideof K2 exchanger (Stream 9 not shownin Table 1) About half of the 3392 mil-lion cu mday enters each exchanger

The exit gas from the tube side ofthese exchangers feeds the K1 and K2

compressors The compressor discharg-es back into the shell side of K1 and K2

exchangers Utilizing parallel K1 andK2 exchangers ensures that the suctionstreamsrsquo temperatures feeding pumpsReda 1 and 2 are as low as possible

A temperature cooler than 33deg Cis optimal to ensure the multistagecentrifugal pumpsrsquo lowest horsepowerdraw In our case this is 1010 hp at15200 kPa-g (2200 psig)

The K1 com-pressor trainis for summerconditions or am-bient temperature3deg C cooler thanthe compressor

coolerrsquos dischargeof 43deg C Thelowest achievabletemperature afterthe K1 exchangeris 38deg C

The corre-sponding pumprequires 1454hp vs the K2

exchangerrsquos discharge of 32deg C whichrequires 1064 hp This increases the

horsepower by 27Because the simulations are mainly

for refrigeration cooling and stabiliza-tion Fig 1 shows the solution-gas in-jection compressors with a single stageabove the K1 and K2 exchangers

Pump performance The process cools the streams from

NGL RECOVERIES Table 4

Liquid produc- Refrigera- Refrigeration tion increaseCase Type tion hp temp 983151C cu mday

1 Spike oil 1630 ndash23 1672 Spike oil 2200 ndash27 167 (no change)3 NGL C

3+ 1600 ndash23 176

4 NGL C3+

2200 ndash27 2155 Fractionate 4610 ndash29 C

3= 36 C

4= 86

C5+

= 134 Case 5 sum

of C3 C

4 and C

5+ 256

THEORETICAL MAXIMUM NGL RECOVERY Table 5

Maximum liquid Ideal gas NGL liquid yieldCom- Mole Volume conversion cu mdayponent fraction Mscfd cu ftgal (bd)

N2 00088 1056

CO2 09186 110232

H2S 00091 1092

C1 00151 1812

C2 00135 1620

C3 0016 1920 3637 200 (125692)

iC4 00029 348 3064 43 (27042)

nC4 00074 888 3179 106 (66508)

iC5 0002 240 2738 332 (20870)

nC5 00024 288 2767 394 (24782)

C6+

00042 504 2616 73 (45872) ndashndashndashndashndashndashndashndash ndashndashndashndashndashndashndashndashndashndashndashndashndash Total 120000 494 (310766)

From Table 4 the total theoretical NGL recovery from the recycle gas stream is 494

cu m day Our fractionation is recovering 256 cu mday or 52

FRACTIONATION PLANT COST Table 6

MillionCapital costs for 3912 million cu mday (120 MMscfd) plant $

300 cu mday de-ethanizer complete (includes recompression reboiler) 225000 hp propane refrigeration complete unit 90Two 9 GJhr heat exchangers (chiller and gas-gas) 13Two 27-40 GJhr heat exchangers (gas-gas K1 and gas-gas K2) 07NGL recovery depropanizer-butanizer complete 30Refrigeration major electrical and mechanical equipment 30NGL mol sieve sweetening 10Low-termperature separator 08Piping racks cable trays insulation buildings and other consumables 35Total all equipment and materials 245Installed costs (15 times equipment and materials) 368Contingency (20) 123 ndashndashndashndashndash Total 736

Costs from skid vendors

NGL COSTS RECOVERED Table 7

ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndash Rate ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashHydrocarbon cut $cu m times cu mday $day $ millionyear

C3 295 times 36 11000

C4rsquos 400 times 86 34000

C5+

345 times 134 46000ndashndashndashndashndashndashndash

Total 91000 3185

The crude spiking alternative gives lower yields because much of the NGL flash aftermixing The costs recovered are $345cu m times 167 = $58000day or $20165 million year

Table 1

HP-P1 HP-P2 HP-R1 HP-R2 Inj -120 MM R3 Solution gas

20000 20000 20000 20000 10000 10000 1000000000 00000 00000 00000 623072 ndash329945 27000000000 00000 00000 00000 154118830 503597 1380000000000 00000 00000 00000 1180180 402712 120000000000 00000 00000 00000 76482664 41872170 7904855000000 00000 00000 00000 481060657 263367577 497199569

432465 438165 436115370150E+06 270696E+06 100267E+07 169455E+06 210598E+07 212384E+07 543272E+07

14547471 10638752 39406609 6659849 82768284 83470270 213513939 0000000 0000000 0000000 0008949 0000000 0008801 0934119 0000000 0918692 0009245 0000000 0009101 0015355 0000000 0015102 0013728 0020000 0013501 0013220 0980000 0016002 0001526 0000000 0002900 0002945 0000000 0007401 0000912 0000000 0008500

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 66

T E C H N O L O G Y

the K1 and 2 last stage discharge intothe Reda pumps to 38deg C for the sum-mer case (K1 compressor) and 32deg Cfor the winter case (K2 compressor)

This pump suction temperaturecauses pumping problems at less than

550 kgcu m which corresponds totemperatures greater than 32deg C

NGL yields As discussed the analyses looked at

several alternatives for determining theprocess with the best economics forNGL recovery

The first case looked at oil spiking ata maximum of 167 cu mday at ndash23deg CThe next case analyzed C

3+ recovery at

ndash23deg ndash27deg and ndash29deg C

Table 4 shows the best recoveries areat ndash29deg C The horsepower (4610 vs2200) however doubles for 16 addedrecovery (256 vs 215 cu m)

Table 4 shows the scenarios simu-lated to evaluate NGL yields for a120-MMscfd recycle gas throughput

To put NGL recovery in perspectivea calculation was made to determinethe theoretical maximum liquid recoverbased on a flow of 33912 millioncu mday (Table 5) The recovery is

256494 = 52

Economics Table 6 shows the cost for installed

refrigeration of 33912 million cu mday The costs are approximate andwere obtained from equipment packag-ers and project execution experience

Table 7 shows the revenue recoveredfrom NGL sales

Economics run from these costs andrevenues indicate that the project wouldpay back in 23 years and have a presentvalue of $73 million Payments were$3185 million

year and the eco-nomics assumeda 5year inter-est and no futurevalue

Also the eco-nomic analysis in-cluded a sensitiv-ity case for a lowermol recycle gas The gas analysis inthis case was from a mid-phase CO

2 in-

jection recycle gas (Table 8) containing

82 mol CO2 This gas is much richerthan the initial 92 CO

2 case

Table 9 shows the simulations forpredicting liquid recoveries with thesame plant configuration as for the pre-vious case The table shows the muchhigher recoveries that lead to a 1-yearpayback

Observations From this evaluation several points

were noted as follows

bull Depending on the refrigera-tion requirements the process shouldinclude dense-phase pumping if pos-sible Pumping has a lower cost thancompression if the process has enoughcooling in summer conditions

bull An added advantage of the refrig-eration is ethylene glycol dehydrationfor avoiding hydrates when chilling thegas to drop out liquids The recycle gastherefore is dehydrated without theneed of exotic piping for corrosion pro-

tection or a hydrate risk when depres-surizing or compressing in a centrifugalcompressor

bull The maximum recovery of themethane and ethane is insufficient tojustify methyldiethanolamine treat-ing to recover the 85000 cu mday

(3 MMscfd) C1 and C

2120 MMscfd of

recycle gasbull A further refinement of the pro-

cess simulations found that droppingthe separatorrsquos temperatures permit-ted recoveries of 176 cu mday of NGL

for ndash23deg C 215 cu mday for ndash27deg Cand 256 cu mday for ndash29deg C This isnot the case for the oil-spiking case It

appears that the oil will not pick up ad-ditional NGL at less than ndash23deg C

bull Installation of the NGL-recoveryprocess equipment should be in modu-larized increments The towers heatexchangers compressors and chillersdo not have efficient turndowns pastplusmn25

Hence the first phase of the proj-ect would instal l 2832 million cu mday (100 MMscfd) units followed byexpansion of the refrigeration compres-

sion stabilization and associated gascompression in 2832 million cu mdayincrements

The authorKenneth J Vargas (kvargasfalcon-edfcom) is president ofFalcon EDF Ltd and has morethan 35 years of experience indesigning upstream oil and gasfacilities Vargas specializes in process engineering and project management of onshore andoffshore projects He previously worked for IBM DuPont Eldorado Nuclear andGulf Vargas is a graduate of the US Air Force Academy and a registered professional engineer inWestern Canada Northwest Territories NorthDakota and Montana

RECYCLE GAS SENSITIVITY ANALYSIS Table 8

Component Mole fraction

H2 00000

He 00000N

2 00166

CO2 08125

H2S 00119

C1 00531

C2 00304C

3 00407

iC4 00154

C4 00073

C5 00121

NGL COSTS RECOVERED 82 CO2 Table 9

ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndash Rate ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashHydrocarbon cut $cu m times cu mday $day $ millionyear

C3 295 times 170 50000

C4rsquos 400 times 222 88800

C5+

345 times 206 71000ndashndashndashndashndashndashndash

Total 209000 7343

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 56

Gas-gas cooling The solution gas from the inlet

gas-gas exchanger enters two streamsOne is the inlet of the tube side of theK1 exchanger (Stream 8 not shown inTable 1) and the other is the tube sideof K2 exchanger (Stream 9 not shownin Table 1) About half of the 3392 mil-lion cu mday enters each exchanger

The exit gas from the tube side ofthese exchangers feeds the K1 and K2

compressors The compressor discharg-es back into the shell side of K1 and K2

exchangers Utilizing parallel K1 andK2 exchangers ensures that the suctionstreamsrsquo temperatures feeding pumpsReda 1 and 2 are as low as possible

A temperature cooler than 33deg Cis optimal to ensure the multistagecentrifugal pumpsrsquo lowest horsepowerdraw In our case this is 1010 hp at15200 kPa-g (2200 psig)

The K1 com-pressor trainis for summerconditions or am-bient temperature3deg C cooler thanthe compressor

coolerrsquos dischargeof 43deg C Thelowest achievabletemperature afterthe K1 exchangeris 38deg C

The corre-sponding pumprequires 1454hp vs the K2

exchangerrsquos discharge of 32deg C whichrequires 1064 hp This increases the

horsepower by 27Because the simulations are mainly

for refrigeration cooling and stabiliza-tion Fig 1 shows the solution-gas in-jection compressors with a single stageabove the K1 and K2 exchangers

Pump performance The process cools the streams from

NGL RECOVERIES Table 4

Liquid produc- Refrigera- Refrigeration tion increaseCase Type tion hp temp 983151C cu mday

1 Spike oil 1630 ndash23 1672 Spike oil 2200 ndash27 167 (no change)3 NGL C

3+ 1600 ndash23 176

4 NGL C3+

2200 ndash27 2155 Fractionate 4610 ndash29 C

3= 36 C

4= 86

C5+

= 134 Case 5 sum

of C3 C

4 and C

5+ 256

THEORETICAL MAXIMUM NGL RECOVERY Table 5

Maximum liquid Ideal gas NGL liquid yieldCom- Mole Volume conversion cu mdayponent fraction Mscfd cu ftgal (bd)

N2 00088 1056

CO2 09186 110232

H2S 00091 1092

C1 00151 1812

C2 00135 1620

C3 0016 1920 3637 200 (125692)

iC4 00029 348 3064 43 (27042)

nC4 00074 888 3179 106 (66508)

iC5 0002 240 2738 332 (20870)

nC5 00024 288 2767 394 (24782)

C6+

00042 504 2616 73 (45872) ndashndashndashndashndashndashndashndash ndashndashndashndashndashndashndashndashndashndashndashndashndash Total 120000 494 (310766)

From Table 4 the total theoretical NGL recovery from the recycle gas stream is 494

cu m day Our fractionation is recovering 256 cu mday or 52

FRACTIONATION PLANT COST Table 6

MillionCapital costs for 3912 million cu mday (120 MMscfd) plant $

300 cu mday de-ethanizer complete (includes recompression reboiler) 225000 hp propane refrigeration complete unit 90Two 9 GJhr heat exchangers (chiller and gas-gas) 13Two 27-40 GJhr heat exchangers (gas-gas K1 and gas-gas K2) 07NGL recovery depropanizer-butanizer complete 30Refrigeration major electrical and mechanical equipment 30NGL mol sieve sweetening 10Low-termperature separator 08Piping racks cable trays insulation buildings and other consumables 35Total all equipment and materials 245Installed costs (15 times equipment and materials) 368Contingency (20) 123 ndashndashndashndashndash Total 736

Costs from skid vendors

NGL COSTS RECOVERED Table 7

ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndash Rate ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashHydrocarbon cut $cu m times cu mday $day $ millionyear

C3 295 times 36 11000

C4rsquos 400 times 86 34000

C5+

345 times 134 46000ndashndashndashndashndashndashndash

Total 91000 3185

The crude spiking alternative gives lower yields because much of the NGL flash aftermixing The costs recovered are $345cu m times 167 = $58000day or $20165 million year

Table 1

HP-P1 HP-P2 HP-R1 HP-R2 Inj -120 MM R3 Solution gas

20000 20000 20000 20000 10000 10000 1000000000 00000 00000 00000 623072 ndash329945 27000000000 00000 00000 00000 154118830 503597 1380000000000 00000 00000 00000 1180180 402712 120000000000 00000 00000 00000 76482664 41872170 7904855000000 00000 00000 00000 481060657 263367577 497199569

432465 438165 436115370150E+06 270696E+06 100267E+07 169455E+06 210598E+07 212384E+07 543272E+07

14547471 10638752 39406609 6659849 82768284 83470270 213513939 0000000 0000000 0000000 0008949 0000000 0008801 0934119 0000000 0918692 0009245 0000000 0009101 0015355 0000000 0015102 0013728 0020000 0013501 0013220 0980000 0016002 0001526 0000000 0002900 0002945 0000000 0007401 0000912 0000000 0008500

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 66

T E C H N O L O G Y

the K1 and 2 last stage discharge intothe Reda pumps to 38deg C for the sum-mer case (K1 compressor) and 32deg Cfor the winter case (K2 compressor)

This pump suction temperaturecauses pumping problems at less than

550 kgcu m which corresponds totemperatures greater than 32deg C

NGL yields As discussed the analyses looked at

several alternatives for determining theprocess with the best economics forNGL recovery

The first case looked at oil spiking ata maximum of 167 cu mday at ndash23deg CThe next case analyzed C

3+ recovery at

ndash23deg ndash27deg and ndash29deg C

Table 4 shows the best recoveries areat ndash29deg C The horsepower (4610 vs2200) however doubles for 16 addedrecovery (256 vs 215 cu m)

Table 4 shows the scenarios simu-lated to evaluate NGL yields for a120-MMscfd recycle gas throughput

To put NGL recovery in perspectivea calculation was made to determinethe theoretical maximum liquid recoverbased on a flow of 33912 millioncu mday (Table 5) The recovery is

256494 = 52

Economics Table 6 shows the cost for installed

refrigeration of 33912 million cu mday The costs are approximate andwere obtained from equipment packag-ers and project execution experience

Table 7 shows the revenue recoveredfrom NGL sales

Economics run from these costs andrevenues indicate that the project wouldpay back in 23 years and have a presentvalue of $73 million Payments were$3185 million

year and the eco-nomics assumeda 5year inter-est and no futurevalue

Also the eco-nomic analysis in-cluded a sensitiv-ity case for a lowermol recycle gas The gas analysis inthis case was from a mid-phase CO

2 in-

jection recycle gas (Table 8) containing

82 mol CO2 This gas is much richerthan the initial 92 CO

2 case

Table 9 shows the simulations forpredicting liquid recoveries with thesame plant configuration as for the pre-vious case The table shows the muchhigher recoveries that lead to a 1-yearpayback

Observations From this evaluation several points

were noted as follows

bull Depending on the refrigera-tion requirements the process shouldinclude dense-phase pumping if pos-sible Pumping has a lower cost thancompression if the process has enoughcooling in summer conditions

bull An added advantage of the refrig-eration is ethylene glycol dehydrationfor avoiding hydrates when chilling thegas to drop out liquids The recycle gastherefore is dehydrated without theneed of exotic piping for corrosion pro-

tection or a hydrate risk when depres-surizing or compressing in a centrifugalcompressor

bull The maximum recovery of themethane and ethane is insufficient tojustify methyldiethanolamine treat-ing to recover the 85000 cu mday

(3 MMscfd) C1 and C

2120 MMscfd of

recycle gasbull A further refinement of the pro-

cess simulations found that droppingthe separatorrsquos temperatures permit-ted recoveries of 176 cu mday of NGL

for ndash23deg C 215 cu mday for ndash27deg Cand 256 cu mday for ndash29deg C This isnot the case for the oil-spiking case It

appears that the oil will not pick up ad-ditional NGL at less than ndash23deg C

bull Installation of the NGL-recoveryprocess equipment should be in modu-larized increments The towers heatexchangers compressors and chillersdo not have efficient turndowns pastplusmn25

Hence the first phase of the proj-ect would instal l 2832 million cu mday (100 MMscfd) units followed byexpansion of the refrigeration compres-

sion stabilization and associated gascompression in 2832 million cu mdayincrements

The authorKenneth J Vargas (kvargasfalcon-edfcom) is president ofFalcon EDF Ltd and has morethan 35 years of experience indesigning upstream oil and gasfacilities Vargas specializes in process engineering and project management of onshore andoffshore projects He previously worked for IBM DuPont Eldorado Nuclear andGulf Vargas is a graduate of the US Air Force Academy and a registered professional engineer inWestern Canada Northwest Territories NorthDakota and Montana

RECYCLE GAS SENSITIVITY ANALYSIS Table 8

Component Mole fraction

H2 00000

He 00000N

2 00166

CO2 08125

H2S 00119

C1 00531

C2 00304C

3 00407

iC4 00154

C4 00073

C5 00121

NGL COSTS RECOVERED 82 CO2 Table 9

ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndash Rate ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashHydrocarbon cut $cu m times cu mday $day $ millionyear

C3 295 times 170 50000

C4rsquos 400 times 222 88800

C5+

345 times 206 71000ndashndashndashndashndashndashndash

Total 209000 7343

8102019 Refrigeration for NGL Recovery

httpslidepdfcomreaderfullrefrigeration-for-ngl-recovery 66

T E C H N O L O G Y

the K1 and 2 last stage discharge intothe Reda pumps to 38deg C for the sum-mer case (K1 compressor) and 32deg Cfor the winter case (K2 compressor)

This pump suction temperaturecauses pumping problems at less than

550 kgcu m which corresponds totemperatures greater than 32deg C

NGL yields As discussed the analyses looked at

several alternatives for determining theprocess with the best economics forNGL recovery

The first case looked at oil spiking ata maximum of 167 cu mday at ndash23deg CThe next case analyzed C

3+ recovery at

ndash23deg ndash27deg and ndash29deg C

Table 4 shows the best recoveries areat ndash29deg C The horsepower (4610 vs2200) however doubles for 16 addedrecovery (256 vs 215 cu m)

Table 4 shows the scenarios simu-lated to evaluate NGL yields for a120-MMscfd recycle gas throughput

To put NGL recovery in perspectivea calculation was made to determinethe theoretical maximum liquid recoverbased on a flow of 33912 millioncu mday (Table 5) The recovery is

256494 = 52

Economics Table 6 shows the cost for installed

refrigeration of 33912 million cu mday The costs are approximate andwere obtained from equipment packag-ers and project execution experience

Table 7 shows the revenue recoveredfrom NGL sales

Economics run from these costs andrevenues indicate that the project wouldpay back in 23 years and have a presentvalue of $73 million Payments were$3185 million

year and the eco-nomics assumeda 5year inter-est and no futurevalue

Also the eco-nomic analysis in-cluded a sensitiv-ity case for a lowermol recycle gas The gas analysis inthis case was from a mid-phase CO

2 in-

jection recycle gas (Table 8) containing

82 mol CO2 This gas is much richerthan the initial 92 CO

2 case

Table 9 shows the simulations forpredicting liquid recoveries with thesame plant configuration as for the pre-vious case The table shows the muchhigher recoveries that lead to a 1-yearpayback

Observations From this evaluation several points

were noted as follows

bull Depending on the refrigera-tion requirements the process shouldinclude dense-phase pumping if pos-sible Pumping has a lower cost thancompression if the process has enoughcooling in summer conditions

bull An added advantage of the refrig-eration is ethylene glycol dehydrationfor avoiding hydrates when chilling thegas to drop out liquids The recycle gastherefore is dehydrated without theneed of exotic piping for corrosion pro-

tection or a hydrate risk when depres-surizing or compressing in a centrifugalcompressor

bull The maximum recovery of themethane and ethane is insufficient tojustify methyldiethanolamine treat-ing to recover the 85000 cu mday

(3 MMscfd) C1 and C

2120 MMscfd of

recycle gasbull A further refinement of the pro-

cess simulations found that droppingthe separatorrsquos temperatures permit-ted recoveries of 176 cu mday of NGL

for ndash23deg C 215 cu mday for ndash27deg Cand 256 cu mday for ndash29deg C This isnot the case for the oil-spiking case It

appears that the oil will not pick up ad-ditional NGL at less than ndash23deg C

bull Installation of the NGL-recoveryprocess equipment should be in modu-larized increments The towers heatexchangers compressors and chillersdo not have efficient turndowns pastplusmn25

Hence the first phase of the proj-ect would instal l 2832 million cu mday (100 MMscfd) units followed byexpansion of the refrigeration compres-

sion stabilization and associated gascompression in 2832 million cu mdayincrements

The authorKenneth J Vargas (kvargasfalcon-edfcom) is president ofFalcon EDF Ltd and has morethan 35 years of experience indesigning upstream oil and gasfacilities Vargas specializes in process engineering and project management of onshore andoffshore projects He previously worked for IBM DuPont Eldorado Nuclear andGulf Vargas is a graduate of the US Air Force Academy and a registered professional engineer inWestern Canada Northwest Territories NorthDakota and Montana

RECYCLE GAS SENSITIVITY ANALYSIS Table 8

Component Mole fraction

H2 00000

He 00000N

2 00166

CO2 08125

H2S 00119

C1 00531

C2 00304C

3 00407

iC4 00154

C4 00073

C5 00121

NGL COSTS RECOVERED 82 CO2 Table 9

ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndash Rate ndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashndashHydrocarbon cut $cu m times cu mday $day $ millionyear

C3 295 times 170 50000

C4rsquos 400 times 222 88800

C5+

345 times 206 71000ndashndashndashndashndashndashndash

Total 209000 7343