Optimum Ethane Recovery in Natural Gas Processing Plant

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CHEMICAL ENGINCERING RESEARCH AND DESIGN 8 8 (2 0 1 0 ) 7 7 9 - 7 8 7

ELSEVIER

Contents lists available at ScienceDirect

Chemical Engineering Research and Design

j o u r n a l h o m e p a g e : w w w e l s e v i e r . c o m / l o c a t e / c h e r d

Optimum ethane recovery in conventional

turboexpander process

R. Chebbi'̂ -% N.S . AI-A moodi^ N.M. Abdel )ahhar°, G.A. Husseíni^ K.A. Al Mazroui

^ Department 0/Chemical Engineering, American uniuersity o/Sharjah, P.O. Box 26666, Sharjah, United Arab Emirates

^ Department 0/Chemical Engineering, The Petroleum Institute, P.O. ßox 2533, Abu D habi, United Arab Emirates

A B S T R A C T

Ethane recovery in a conventional turboe xpand er process is optimized considering different d eme thaniz er p ressure s

and different feeds: a lean gas and a rich one. The design variables are varied, while meeting proc ess con straints ,

in order to find the optimum conditions achieving the maximum profit. The analysis covers the whole process

including the refrigeration part, and the entire typical demethanizer pressure range. The optimum ethane recovery

is comp ared with the nnaximum possible recovery for each value ofthe dem etha nize rpre ssur e. Reco mm endations are

given regarding the selection of the level of ethan e recovery, along with the dem ethan izer pressure , and refrigeration

recovery system.

I I © 2009 The Institutio n of Chem ical Enginee rs. Published by Elsevier

B V

All rights reserved.

Keywords: Natu ral gas liquids (NGL); Ethane : Recovery; Optim ization; Tbrboex pander; Simu lation

Introduction

Natural gas l iquids are valuable components in natural gas .

Several extraction processes were proposed. Reviews about

proc esse s for NGL recovery can be found in M anning a nd

Thonnpson (1991), Arnold and Stewart (1999), GPSA (2004)

and Kidnay and Parrish (2006). A recent ove rview is given

by Chebbi et al. (2008), and briefly reviewed here. Differ-

ent opti ons include JT (Joule Tho mp son) valve e xpa nsion ,

exte rnal refrigeration (Rüssel, 1977) using propa ne as a refrig-

erant, and turboexpansion. Typically, a turboexpander is used

in com binat ion w ith JT expansion and pr opane refrigera-

t ion. Other methods include cascade refrigerat ion which is

complex and requires high compression cost (Manning and

Th om pson , 1991). Mixed refrigerant (MacKenzie and Donnelly,

1985;

Mannin g and Tho mp son , 1991) is comm only u sed in LNG

proce sses, but m uch less in NGL recovery. Refrigerated lean

oil absorpt ion is expensive in terms of equipm ent a nd energy

requ irem ents and is hard to operate (Arnold and S tewart , 1999;

GPSA, 2004),

The turboexpander process is dominat ing ethane recovery

processes. Different processes have been proposed. The evo-

lution in design for the old gen eratio n wa s sum ma rize d by

McKee (1977). Th e simple p lant co nsist s of tu rboe xpa nsion .

The other options include the use of side reboiler, refrigera-

tion, and two sta ges of expa nsion (McKee, 1977).

The next generat ion of etha ne recovery processes includes

the re sidue recycle (RR) proc ess, the gas subcoo led proce ss

(GSP) (Pitman et aL, 1998; GPSA, 2004). In the GSP, the gas

stream leaving the cold separator is spl i t in to two s treams,

one of them feeding the turboexpan der, and the second one is

subcooled by the dem ethan izer o verhead s tream, flashed in a

valve, and th en se nt to the dem etha nize r as a reflux. The next

gene ration also includes th e CRR proc ess. It ha s one additio n

to the GSP and is reported to achieve very high eth an e recovery

{Pitman et al., 1998); but with a cryogenic com press or, cost ma y

be proh ibitive (Lee et al., 1999). Th e IPSI is an ot he r mod ification

of the GSP (Lee et al., 1999; GPSA, 2004). Combining GSP with

liquid subcooled process (LSP) (Jibril et al., 2006) was found

to yield higher recoveries t ha n GSP and LSP used individually

using proc ess sim ulation for eight different feeds. RSV (recycle

split-vapor) proce ss {Pitman et al., 1998) is ano the r modifica-

tion ofthe GSP, and RSVE (recycle spli t -vapor with enrichm ent)

process is itself a modification of the RSV (Pitman et al., 1998)

found to lower capital cost compared to RSV.

Studies have been carried out in order to opt imize ethane

recovery. Wang consid ered different feeds w ith low CO2

content and found that a combinat ion of turboexpansion



78 0

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E N G I N E E R I N G R E S E R C H N D D E S I G N

8 8 [ 2 0 I o 1 7 9 - l S ?

and external refrigeration achieves minimum energy require-

ments (Bandoni et al, 1989}. Bandoni et al. (1989) divide the

process into two sectors. The first one includes com pression,

recompression and atm ospheric hea t exchange while the sec-

ond sector includes separation, expansion refrigeration, and

heat exchange below ambient temperature.

A

general super-

structure embedding different features was considered. The

main optimization variables they used are: cold tank tem-

perature and pressure in addition to demethanizer column

pressure. The energy balance over sector 2 was found to

provide guidelines for the optimum ethane recovery process

selection, and the economic objective function was found

to increase with demethanizer pressure due to less energy

requirement for gas recompression. Fernandez et al. (1991)

extended the analysis to CO2-containing mixtures. One con-

straint was added: prevention of carbon dioxide precipitation

in the demethanizer. The effect of increasing the demeth-

anizer pressure was also found to enhance the net annual

benefit (calculated as sales minu s operating cost min us an nual

investment cost).

A cost-revenue comp arison was perform ed for GSP, SFR

{split-flow reflux), ISS (typical industry-standard stage piant)

a typical turboexpander process, and CRR processes by

Wilkinson and Hudson (1992) for one gas feed containing

6.76% C2 P a flow rate of lOOMMscfd, and

1 0 5 0

psig nominal

pipeline p ressure . The SFR process is based on GSP, with the

subcooled stream used to reflux the demethanizer indirectly.

The

G SP

was found to be the best choice if min imum capital

cost is needed, and CRR process was found to be the process

of choice if ethane market is expected to be favorable. SFR pro-

cess

w a s

found

t o b e

the best option if ethane price projections

are low.

L S P ,

BTEP (basic turbo expand er process ), GSP, and 2-sDP

(two-stage demethanization process) were compared by Diaz

et al. (1997) for different feed gas mix tures w ith C2+ rang-

ing from

6

to

2 5 .

The

G SP

was found to give the ma ximu m

profit for mo st m ixtures, especially the lighter on es (lower C 2+

conten t). Optimal design and operating cond itions were found

using a superstructure embedding different expansion alter-

natives together with mixed-integer n onlinear programm ing

(MINLP). The impact of CO 2 conten t was found to decrease

ethane recovery.

Konukman and Akman (2005) analyzed flexibility and oper-

ability of a H EN-integrated n atura l gas exp ande r plan t, using

HYSYS process simulation. They showed that HEN synthe-

sis design and flexibility analysis, based on nominal values,

would provide incorrect results due to significant changes

of physical properties. Konukman and Akman concluded the

difficulty or impossibility of au tom ated HEN syn thesis for a

highly energy-integrated plant like the turboexpander process

using process flowsheet s imulators.

Mehrpooya et al. (2006) simulated an existing

N G L

recov-

ery un it u sing HYSYS. Two modifications were cons idered

suitable for optimization: turboexpander, and turboexpander-

exchanger configurations to find the best revamping alterna-

tive.

A

genetic algorithm that used for optimization to ensure

the maxim um profit achieved is global and not only local.

Five different turboexpander ethane recovery processes

as we have m ention ed above. Considering four different g

feeds and four different demethanizer pressures, and varyi

all available design variables, Chebbi et a l . (2008) showed t

the conventional turboexpander process yields larger max

mum ethane recovery in most cases except when both t

feed is lean and the dem ethanizer pressure is low.

In the present

work

w e

consider the effect of dem etha ni

pressu re on optimal e than e recovery level selection to achie

the m aximum profit in a conventional turboexpander proce

Optimization is performed b y changing a ll the design variab

for each value of the demethanizer pressure, while satisfyi

process constraints. In contrast to the works in Bandoni et

(1989) and Diaz et al. (1997), the proce ss sim ulation does n

start at the cold tank, and does include the whole

N G L

reco

ery unit. Recommendations are given based on the prese

results.

2 . Simulation and optimiza tion

A s

described

in

Chebbi et

a l.

(2008), in the conventional proc

considered

( F i g . 1

in Chebbi et

a l.

(2008)), the feed is first coo

while providing the reboiler duty, then cooled further by he

exchange with the residue gas, followed by heat exchan

with propane in a chiller to reach a temperature of -31

After separation from the liquid, the gas leaving the sep

rator is cooled by heat exchange with the overhead strea

leaving the demethanizer, then separated from the liquid

the cold separator, expanded in a turboexpander and sent

the dem ethanizer. The liquids leaving the two separators a

expanded through

JT

expansion and sent to the demethani

at lower levels. The turboexpander provides part of the pow

needed to recompress the residue gas. The other portion

provided by a recompressor to raise the residue gas pressu

to882psia.

The tem peratures after cooling by heat exchange w ith t

residue gas leaving the demethanizer are design variab

that are changed in the optimization process to maximi

ethane recovery for different demethanizer pressures. T

constraints are conditions preventing temperature cross

these two heat exchangers.

Fig. 1 also shows the refrigeration loop. The chiller is co

mon to both the turboexpand er process and the refrigerati

unit. A propane econom izer cycle is selected (Manning a

Thompson, 1991). The compression cost is smaller compar

to propane refrigeration simple cycle since part of the vapor

obtained at an intermediate pressure, called economizer pr

sure after the first valve expansion (Manning and T homp so

1991); therefore the economizer cycle requires less power

recompression.

The optimum economizer pressure achieving minimu

total compression power selected in our process is given

Fraser (Manning and Thompson, 1991):

cond.out

Pc h

in which Pcord,out ar^d Pd, are the pressures of propane af

condensation (achieved by air cooler in our case) and ins



C H E M I C A L E N G I N E E R I N G R E S E A R C H A N D D E S I G N 8 8 3 0 I o )

761

Feed

NG L

R FRIG R TION

LOOP

Fig.

Conventional ethane recovery process with the propanerefrigeration oop

a t

low demethanizer pressure.

and

4 5 0

psia as in Chebbi et

a l.

2008). In all case s, the feed gas

is at 100 F and 8 8 2

psia.

the residue gas is recompressed to

882psia. and in the

N G L

stream , the molar ratio of Ci to

C2

is

set at 0.02 Manning and T hom pson, 1991). Also in all cases

the feed gas rate is 10 980 Ihmol/h as in Chebhi et al. 2008).

2 . 1 .

Lou; deme thonizer pressu re

The demethanizer pressure considered here is

10 0 psia. The

outlet temperature from the first gas-to-gas heat exchanger

is a design variable that affects th e cooling partition betwe en

the ftrst gas-to-gas hea t exchange an d refrigeration.

As explained in the next section, the simulation and opti-

mization study for this case shows that the profit is enhanced

when the first gas-to-gas heat exchanger is discarded, with

refrigeration providing the full cooling duty required to lower

the feed gas tem peratu re to - 3 1 F. The process flowsheet

shown in Fig. 1 reflects the abovementioned process change.

Table Feed gas co mposition.

Component

Nitrogen

Methane

0.01

0,93

0.01

0.69

2

Intermediate demeth nizer pressure I

The pressure considered as intermediate is

2 1 5 psia. A s

will be

mentioned in the next section, it is more economical in this

case to keep the first gas-to-gas heat exchanger to cool the

feed gas before refrigeration to recover refrigeration from the

residue gas leaving the dem ethaniz er, after it exits the second

gas-to-gas heat exchanger operating at low temperature. The

process at intermediate demethanizer pressure corresponds

to the typical

flowsheet

n which two gas-to-gas heat exchang-

ers are used Fig,

1

in Chebbi et al. 2008)).

2 . 3 .

High deme thanizer press ures

The pressures considered here are 335 and

4 5 0 psia. A s

for the

intermediate pressure case, the results show th at a first gas-

to-gas heat exchanger

is

also economically justified and yields

a higher profit. Another modification, dictated by thermody-

namics, is required at high demethanizer pressures. As the

demethanizer pressure increases, the temperature profile in

the column is shifted up. The simulation results show that, at

335 and 4 5 0 psia, the feed gas temperature is not high enough

to provide heat in the reboiler. Thus, an external heat source

has to be used in this case. The process flowsheet shown in

Fig. 1 was modified to reflect the two above mention ed process





C H E M I C A L E N G I N E E R I N G R E S E A R C H A N D D E S I G N 8 8 ( 2 0 I 0 )

3.1.2. Maximum profit

leíble 2 clearly shows that selecting lower demethanizer pres-

sures yields higher profit for feed

D ,

which is rich. The higher

NGL recovery offsets the additional cost of compression at

low pressure. For feed A, which is lean, there is an optimum

demethanizer pressure at which the ohjective function, and

therefore the profit, is at a maxim um (see Table 3). Compared

to feed A , the profits are substantially higher in the cas e of feed

D (due

to

the substantially higher

N G L

content) with compa-

rable capital costs at low and intermediate (100 and 215psia)

demethanizer pressures.

3.1.3.

Ouerall capital cost

F or feed A . as the demethan izer pressure dec reases, the overall

capital cost increases, which is expected due to higher recom -

pression cost needed to raise the residue gas pressure for the

purpose of transportation. For feed

D ,

the overall capital cost

increases as pressure increases from 2 1 5 to 33 5 psia. Although

the recompression cost is lower at higher demethan izer pres-

sure, the refrigeration load is higher due to a lower potential of

cooling for the residue gas which leaves the demethanizer at

a higher temperature. Another reason for the increase of the

capital cost at higher demetha nizer p ressures is the necessity

of an external heat source.

3.1.4. Rc/rigeration costs

The results in lïihle 3 clearly show that refrigeration is less

costly (both in term s of capital cost and utility) in the 21 5 psia

case compared to both 335 and

4 5 0

psia cases, which is

expected since the residue g a s reaches the irst gas-to-gas hea t

exchanger at a lower temperature, and has a larger potential

to cool the feed g as, and as a result, the load on the refriger-

ation unit is reduced. We would expect the refrigeration cost

to be lower in the 33 5 psia case compared to the 4 50 psia one

for the same reason; nevertheless another factor should be

taken into account. At these high demethanizer pressures,

the tem peratures of the residue gas leaving the demethanizer

become closer, making the difference in potential for cooling

less;

however, the h eat cap acity of the residue

g a s .

is larger for

the

4 50

psia case due to the fact that more ethane and heavier

components are left in the residue gas, and this tends to give

more potential for cooling. Simulations show that the second

effect (higher heat capacity) slightly surpasses the first effect

(lower residue gas temperature) at these high demethanizer

press ures (335 and 45 0 psia).

The refrigeration cost (both in terms

of

capital cost an d util-

ity) is the h ighest in the 10 0 psia case, due to the fact that the

first gas-to-gas hea t exchanger is eliminated, which adds the

corresponding cooling load to the refrigeration unit through

the chiller. .

3.1.5. Main process capito and utility costs

Fo r

both feeds

A

and

D .

the main process capital cost decreases

as the demethanizer pressure increases due to lower recom-

pression cost (see Tables 2 and 3). The same trend is noticed

for the utilities cost for all demethanizer pressures in the case

of feed A and for low and intermediate demethanizer pres-

Table 4 - Results for maximum ethane recovery in the

cas e of feed D high NGL to gas price ratio).

Maximum ethane recovery (%)

Objective

function (MM )

91.0

85.3

76.3

64.3

657.1

653.9

637.3

619.1

ble

5 - Results for maximum ethane recovery in the

ca se of feed A high NGL to gas price ratio).

M axim um eth a n e recovery (%) Objective

function (MM )

94.1

82.4

65.4

45.9

353.1

354.9

349.9

344.7

3.1.6. Comparison between optimum and maximum

recouery cases

Using the design variables values obtained to get the max-

imum ethane recovery for each demethanizer pressure in

Chebbi et al. (2008) yields the profit at maximum recovery.

Results are show n in Tables 4 and

5

for feeds

A

and

D .

respec-

tively.

Comparing results from Tables 2 and 4 clearly show s th at

the optimum occurs at the maximum ethane recovery for feed

D .

which is rich. For feed A, which is lean, the same con-

clusion is reached at intermediate and high demethanizer

pressu res (see Tobies 3 and

5) ;

however, even at low demeth-

anizer pres sure. 100psia. the difference is not significant; and

operating at the highest ethane recovery level would give

a

profit nearly equal to the optim um , along with a higher ethane

recovery

(1.5

difference at the lowest demethan izer pressure

1 00 psia).

The trends are better illustrated In Figs. 2-5, in which the

difference in profit and ethane recovery are plotted versus

demethanizer pressure.

720

6 8

2 640-

o

u 600-

5 2

48 0

FeedD

o - - - - - _ _ ,

50 100 150 200 250 300 350 400 450 500

Demethanizer Pressure, psia



7 8 4

C H E M I C A L E N G I N E E R I N G R E S E A K C H A N D D E S I G N

8 8

( 2 0

I

O ) 1 1 S 1 Z 7

39

34

5 1 15 2 25 3 35 4 45 5

Demethanizer Pressure psia

Fig. 3 - Objective function as a function of demethanizer

pressure for feed A circle: optim um va lue, triangle: value

corresponding to maximum ethane recovery, continuous

curve: high N G L to gas price ratio, das he d curve: low N G L to

gas price ratio).

40

50 100 150 200 250 300 350 400


45 0

Fig. 5 - Percent

C 2

recovery as a function of dem eth aniz e

pressure for feed A circle: optim um va lue, triangle:

m axim um ethan e recovery, contin uou s curve: high NGL t

gas price ratio, dash ed curve: low

N G L

to ga s price ratio).

100

50 100 150 200 250 300 350 400 450 500


Fig. 4 - Percent C2 recovery as a function of demethan izer

pressure for feed D {circle: optimum value, triangle:

ma xim um ethan e recovery, continuou s curve: high NGL to

gas price ratio, das hed curve: low N G L to gas price ratio.

The dashed curve coincides with the con tinuous one).

3 2 Loiu NGL to gas price ratio

The s imulat ions were repeated with the fol lowing prices:

9.60 /MMBtu and 1.00 /US gallon. The ratio in this case is

low: 0.104

( / U S

gallon)/( /MMBtu), Only differ ence s or impo r-

tant similarities with respect to the previous case (high NGL

to gas price ratio) will be discussed in this part. The compared

results can be seen in Figs, 2-5 introduced previously.

3.2.1.

Pro/it

The results obtained are shown in Tables 6 and 7, In all cas

the objective function is seen to show a peak, with an op

mum pressure at which the profi t is maximum. This was n

the case in part

1

(high NGL to gas price ratio) for rich feed

D ,

which producing the m axim um NGL, by operat ing at the lo

est pressure

1 0 0

psia, yields the highest profit, a likely re

at high NGL prices for a rich gas. At low NGL to gas price ra

the values of the objective function, and therefore the pro

are sma ller in the case of feed D, wh ich is expected since t

gas is rich, and NGL price is low. On the contrary, the prof

are high er in the ca se of feed A , which is also ant icipated s i

the feed is lean (with a significantly smaller NGL product

rate), and gas prices are high . Considering feeds A and D se

arately, the objective function values are relatively close

the different dem ethan izer pressu res at low NGL to gas pr

ratio.

3.2.2. Main process and refrigeration utilities costs

Compared to the high N G L to sales gas price ratio case, the u

ities costs are higher in all cases due to the higher gas pr

excep t for feed A at low dem eth an ize r p res sur e (100 psia),

which the refrigeration ut i l ity cost increas es but the ma in p

cess utility cost declines. Optimization yields an increase

the out let temperature of the feed gas exi t ing the cold g

to-gas heat exchanger shown in Fig . 1 (only one such h

exchanger is included at 100psia, for the reasons mention

above), which enhances the power of the turboexpander, a

reduces the load on the booster, and therefore the utility c

on the process side.

Table 6 - Resu lts from optim ization for feed D low

N G L

to ga s price ratio).

P (psia)

CUT

(MM )

FCI

(MM ) Total FCI

(MM )

Objective

function (MM )

Ethane

recovery



C H E M I C A L E N G I N E E R I N G R E S E A R C H A N D D E S I G N 8 8 ( 2 0 1 0 I

Table 7 - R esults from o ptim izatio n for feed A low NGL to gas price ratio).

P (psia)

C U T ( M M )

F C (MM$) Total FCI

(MM$)

Objective

function (MM$)

Ethane

recovery {%)

Main Refrigeration Main Refrigeration

proce ss unit cycle proce ss uni t cycle

1

2 1 5

3 3 5

4 5

4 1 6

4 5 2

3 6 4

2 4 4

2 4 3

8 4

1 4 9

1 4 7

2 3 3

1 4 4

1 8

8 2

8 3 1

4 1 8

6 1

6 5

3 1 6

1 8 6

1 6 9

1 4 3

3 7 4 2

3 7 8

3 7 6 6

3 7 6 3

8 3 3

8 2 4

6 5 4

4 5 9

Table 8 - Results for maximum ethane recovery in the

cas e of feed D low NGL to gas price ratio).

Conclusions

Maximum ethane recovery (%)

Objective

function (MM$)

91.0

85.3

76.2

64.3

535.7

S38.7

530.6

525.0

Table 9 - Re sults for maxim um ethan e recovery in the

case of feed A low N L to gas price ratio).

M axim um e th a n e recovery (%) Objective

function (MM$)

94.1

82.4

65.4

45.9

374.0

378.0

376.6

376.3

3.2.3.

Capital costs

The main process and refrigeration unit capital costs are the

same as for the high

N G L t o

s les gas price r tio case, excep t for

feed A at demethanizer pressure equal to 10 0

psia.

This special

case

is

the only case in which the design variable (temperature

of the feed gas exiting the cold gas-to-gas heat exchanger) at

optimum conditions does not correspond to the limiting case

where th e design variable i s set just to avoid temperature cross

in the gas-to-gas heat exchanger. In all other cases, equipm ent

sizing is the same, and equipment prices are considered not

affected

b y

the relative costs of

N G L

and sales

g a s . A s

indicated

in the previous part, there is less load on the booster in this

special case, and this yields not only

a

de crease of the utility

o n

the process side, but also a drop in the capital cost compared

to the high

N G L

to sales gas price ratio case.

3.2.4. Ethane recouery

The results are shown in Tables 6-9. As for the high NGL

to sales gas price case, for a given demethanizer pressure,

the optimum is found to occur at maximum ethane recov-

ery except for the lean feed

A

at tow demethanizer pressure,

10 0

psia.

The deviation from maximum ethane recovery is sig-

The present work shows th at in all studied cases except one,

there is an optimum demethanizar pressure at which the

profit is maximum. The exception corresponds to rich gas

D and high NGL to gas price ratio. The optimum is reached

in this special case at the lowest demethanizer pressure,

1 0 0

psia.

In contras t to the w orks in B andoni et al. (1989) and

Diaz et al. (1997) the process simulation does not start at the

cold tank, and does include the whole NGL recovery unit in

our analysis. In addition, the demethanizer pressure range

is wider in our work and does cover the typical range indi-

cated in Manning and Thompson (1991). Recovery is found

to be adversely affected at higher demethanizer pressures as

expected. For a given demethanizer pressure, the recovery at

optimum conditions is found to be equal to the maximum

recovery at the specified pressu re, excep t for feed A at low

demethanizer pressure

( 1 0 0

psia). In addition, the maximum

profit (optimum) is found equal to the profit at maximum

etha ne recovery except for the special ca se of feed A at low

demethanizer pressure for which the optimum profit and

value at maximum ethane recovery are found not equal hut

very close.

Optimization shows that the use of a gas-to-gas heat

exchanger before the chiller to recover refrigeration yields

lower profit at low demethanizer pressure (100psia), and thus

should be discarded at low pressure. At high demethanizer

pressures (335 and 4 5 0

psia),

the feed gas cannot be used to

provide the reboiler duty. Therefore, less cooling (refrigera-

tion recovery) can be achieved, and an e xterna heat source

is required. At intermediate pressure ( 2 1 5 psia). the main pro-

cess is a conventional turboexpander unit with two gas-to-gas

heat exchangers, and where the demethanizer reboiler duty

is provided by the feed gas, which results in further cooling

before refrigeration. The costing structure is relatively com-

plex and acco unts for the overall cost including th e interac tion

between the main process part and the refrigeration unit

through the chiller. Details are given to explain the costing

structure. Due to fluctuations in the prices of N G L and sales

g a s ,

it is esse ntial, for the purpose of optim ization, to have

a control system capable of adapting to changes needed in

the design variables (including the dem ethan izer p ressure) for

optimization, along with a flexi le process that encompasses

possible required changes in terms of refrigeration recovery

as the demethanizer pressure changes as discussed in the

present paper.



78 6

C H E M I C A L E N G I N E E R I N G R E S E A R C H A N D D E S I G N 8 8 ( 2 0 0 ) 7 7 9 - 7 8 7

Appendix A. Appendix

The relevant bare module equipment costs (CBM) (TUrton et

a l . , 2003), utility cost s, and total annu al costs are comp ared

in Tables A1-A3. The CBM values for case 2 can be o btain ed

directly as for case 1; however most of those values can be

obtained by s imply using the eq uat ion s given below.

The ut i l i t ies in Table Al are propo rt ional to the c omp ressor

duties , therefore we have

Th e ma ss flow rate of the refrigerant, propa ne, is obtained

(A

.4 .

The intensive properties inside the propane loop remain t

same. The utilities cost on the refrigeration side can be simp

obtained as

4

Utilityj - fTT^Jtilityi (refrigeration loop)

(A

Duty2

The s izes of the com presso rs and their drives in the refrige

tion loop can be found as

(A

l^ble Al - Compared annual utility cost estimates for two different cooling partitions between gas-to-gas heat exchange

and refrigeration.

Equipmen t

Refrigeration part: 50%

Duty (Btu/h) Duty ( /year)

Refrigeration par t: 40%

Duty (Btu/h) D uty ( / ye

Refrigeration

First compre ssor

Second compressor

Tbtal {refrigeration)

7,635,000

7,298,000

14,933,000

Main Process (equipmen ts affected by the partit ion chang e)

Co mp ress or (booster) {2) 70,437,000

Total 85,370,000

535,100

511,400

1 46 5

4,936,200

5,982,700

6,108,000

5,838,000

11,946,000

75,224,000

87,170,000

428,100

409,100

837,200

5,272,000

6,109,000

Table A2 - CBM estim ates compared for two different cooling partitions bet we en gas -to-ga s heat exch ang e and

refrigeration the only equ ipm ent includ ed in the total cost are those affected by the partition change).

Equipment (number of

units if more than 1)


Size per unit

CBM


Size per unit

CBM ( )

250.6 ft^

820 ft=

1 6 6 Btu/h

2,294,000 Btu/h

1 535

Btu/h

2,192,000 Btu /h

3189 ft̂

Refrigeration

First separa tor

Second separator

Chiller (2)

First com pressor

Drive for the ñrst

compressor

Second compressor

Drive for the second

compressor

Air cooler

Total CBM {refrigeration)

Main process {equipment affected by the partit ion change)

First gas-to gas hea t 786 ft̂

exchanger

Co mp res sor (booster) (2) 7,936,000 Btu/h

Drive for th e bo os ter (2) 10,581,000 Btu/h

Air coo ler 3278 ft'

Total CBM (process

equipment affected by the

partit ion change)

Total

13,500

95,500

1 461 4

1 454 4

343,700

680,900

331,400

241,300

4,622,100

183,200

4,610,900

2,319,100

261,300

7,374,500

11,996,600

14.3 ft̂

179.4ft'

750 ft̂

1 285 Btu/h

1 835 Btu/h

1 228

Btu/h

1 754 Btu/h

2552 ft^

1054 ft=

8 475

Btu/h

11,300,000 Btu /h

352 5 ft=

11,800

77,800

1 342 5

1 21 8

287,300

566,400

277,000

217,100

3,990,700

212,800

4,817,400

2,442,600

270,700

7,743.500

11,734,200



C H E M I C A L E N G I N E E R I N G R E S E A R C H

A N D

D E S I G N

8 8 ( 3 0 I O )

The volumes of the two separators in the refrigeration loop

can be obtained using

A2 p2

AiDi

VQi/Ua

3 3

3

V 5

(A5)

where V , A, D and L represent the volume, the cross-sectional

area, diameter and length of the vessel, respectively, and Ua

the maximum allowable vapor velocity.

The sizes of the booster and the first gas-to-gas heat

exchanger cannot be calculated directly from the values

obtained for case 1 since there are changes in relevant inten-

sive properties. The same is also valid for the chiller since it

has one stream on the process side.

The size of the equipme nt being found, and the bare mod-

ule factors being the same for cases 1 and 2, the bare module

equ ipm ent cost, CBM2,

can

be obtained using

CBM

(A6}

where the K values, for the corresponding equipment, are

given in nirton et al. (2003).

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Optimum Ethane Recovery in Natural Gas Processing Plant

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