introducing a novel integrated ngl recovery process

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Chemical Engineering and Processing 49 (2010) 376–388

Contents lists available at ScienceDirect

Chemical Engineering and Processing:Process Intensification

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 e p

Introducing a novel integrated NGL recovery process configuration (with a

self-refrigeration system (open–closed cycle)) with minimum energy

requirement

Mehdi Mehrpooya ∗, Ali Vatani, S.M. Ali Mousavian

Department of Chemical Engineering, Faculty of Engineering, University of Tehran, P.O. Box 11365-4563, Tehran, Iran

a r t i c l e i n f o

Article history:

Received 22 October 2009

Received in revised form 9 February 2010

Accepted 5 March 2010

Available online 15 March 2010

Keywords:

Natural gas

Liquid recovery

Separation

Self-refrigeration

Integration

a b s t r a c t

In this study a novel process configuration for recovery of hydrocarbon liquids from natural gas is

proposed. The required refrigeration in this configuration is obtained by a self-refrigeration system

(open–closed cycle). High performance of the multi-stream heat exchangers, high recovery levels of the

hydrocarbon liquids and low required compression power(in the internal refrigeration section) are three

of most important characteristic of the proposed configuration. The effects of the mixed self-refrigerant

flow rate and pressure on the performance of the process are discussed. Various values for feed compo-

sition are tested and the results show that the process can work efficiently with different feeds. In order

to analyze the need of external refrigeration by a close or open cycle that is related to the composition

of the inlet gas, a configuration with external refrigeration is designed the manner that it is similar with

the purposed configuration in the separation section.

© 2010 Elsevier B.V. All rights reserved.

1. Introduction

Increase in thepriceof energysources andeconomical problems

have causedcryogenic natural gasliquid recovery plantsto be more

complex and efficient. In other words, the new generation of NGL

plants is created based on decreasing the fixed and operating costs

of the plant for a specific output.

In recent years, there has been great incentive to improve the

efficiency with which the existing capital in chemical processing

facilities is utilized. Retrofit design projects aim to find ways to

maximize theuse of existing equipment, andthus minimize expen-

ditureon newcapital,when theproduction objectives change[1,2].

Process configuration and plant operation are two important

factors which can affect the performance of the process signifi-cantly.

Numerous expansion processes are commonly used for hydro-

carbon liquids recovery in the gas processing industry, particularly

in the recovery of ethane and propane from high pressure feed gas.

In most known NGLexpander processes,feed gasis cooledto a rela-

tively low temperature to achieve partial condensation, typicallyby

heat exchange with the demethanizer overhead vapor, side reboil-

ers, and/or external propane refrigeration. In preferred plants, the

Abbreviations: NGL , natural gas liquid; BHP, break horse power; PRSV,

Peng–Robinson–Stryjek–Vera.∗ Corresponding author. Tel.: +98 21 66905037; fax: +98 21 66957784.

E-mail addresses: [email protected], [email protected] (M. Mehrpooya).

refrigeration content of the demethanizer overhead product and

subsequent expansion is used to subcool a portion of thepreferably

unprocessed feed gas to produce a low temperature reflux, while

a portion of the expander discharge is heated by the preferably

unprocessed feed gas to form a temperature-controlled column

feed [3].

Thegas subcooledprocess(GSP)was developedto overcomethe

problems encountered with the conventional process. This process

altersthe conventionalprocess in several ways. A portion of thegas

from thecoldseparator is sent toa heat exchangerwhereit istotally

condensed with the overhead stream. This stream is then flashed

to top of the demethanizer providing reflux to the demethanizer.

As with the RR (residue recycle) process, the expander feed is sent

to the tower several stages below the top of the column. Because of this modification,the cold separatoroperates at much warmercon-

ditions well awayfrom the system critical. Additionally,the residue

gas recompression is less than with the conventional expander

process. The horsepower is typically lower than the PR process at

recovery levels 92%.

A new process scheme has been developed to combine the GSP

andRR processes into an integrated process scheme. This concept is

basedon applyingthe best featuresof each process tothe integrated

design. This combination can result in higher ethane recovery effi-

ciency than can be achieved with GSP. The cold residue recycle

(CRR) process is a modification of theGSP process to achieve higher

ethane recovery levels. The process flow diagram is similar to the

GSP except to the overhead system to take a portion of the residue

gas and provide additional refluxfor the demethanizer.This process

0255-2701/$ – see front matter © 2010 Elsevier B.V. All rights reserved.

doi:10.1016/j.cep.2010.03.004




M. Mehrpooya et al. / Chemical Engineering and Processing 49 (2010) 376–388 377

is attractive for extremely high ethane recovery. Recovery levels

above 98% are achievable with this process [4].

This process has the further advantage that it can be operated

for near complete rejection of ethane while maintaining in excess

of 99% propane recovery. While the CRR process is unmatched in

terms of recovery efficiency, the recycle split-vapor process (RSV)

sometimes requires less capital investments. Like the CRR process,theRSV processusesthe split-vapor feed toprovide thebulk ethane

recovery in the tower. The methane reflux stream for the tower

is produced by withdrawing a small portion of the recompressed

residue gas, condensing and subcooling it, then flashing it down to

tower pressure andsupplying it as thetop feed [5]. The higherpres-

sure of this methane stream (compared to CRR) allows the tower

overhead gas to be used to provide the condensing and subcooling,

so that the split-vapor feed can be supplied directly to the tower

[5].

A variation of the RSV process is the recycle split-vapor with

enrichment process(RSVE). Similar toRSV, a recycle streamis with-

drawn from the recompressed residue gas, but it is mixed with the

split-vapor–vapor feed before being condensed and subcooled so

that it does not require a separate exchanger or exchanger passage.Since the ethane content of the top tower feed is richer than for

the RSV process, the ultimate ethane recovery is limited to slightly

lower levels than RSV dueto equilibrium effects, butthe lower cap-

ital investments and simplicity of RSVE relative to RSV may justify

the small loss in ethane recovery in some projects [5].

Another improvementof the turboexpander-basedNGL process

is the IPSI enhanced NGL recovery process. This process utilizes

a split stream from or near the bottom of the distillation column

(demethaniozer) as a mixed refrigerant. The mixed refrigerant is

totally or partially vaporized, providing refrigeration for inlet gas

cooling otherwise normally accomplished using an external refrig-

eration system. The vapor generated from this “self-refrigeration”

cycle is recompressed and recycled back to the bottom of the

tower where it serves as a stripping gas. The innovation not only

reduces or eliminates the need for inlet gas cooling via external

refrigeration, but also provides the following enhancements to the

demethanizer operation:

• Lowers the temperature profile in the tower, thereby permit-

ting better energy integration for inlet gas cooling via reboilers,

resulting in reduced heating and refrigeration requirement.• Reduces and/or eliminates the need for external reboiler heat,

thereby saving fuel plus refrigeration.• Enhances the relative volatility of the key components in the

tower when operated at a typical pressure, thereby improving

separation efficiency and NGL recovery; or alternatively allows

increased tower pressure at a typical recovery efficiency, thereby

reducing the residue gas compression requirements [4].

A strategy for process configuration design and debottleneck-

ing of natural gas processing plants based on turboexpansion was

presented in Ref. [6]. The approach combines a rigorous process

simulation model and a mixed-integer nonlinear programming

(MINLP) optimization methodology that embeds different expan-

sion alternatives within a superstructure. A wide range of natural

gas mixtures with 6–25% of condensable components is studied in

order to determine optimal plant topology and operating parame-

ters under different process conditions.

There aresome process parameters which affect thepower con-

sumption and product quality and quantity [7–11].

Fish [12] demonstrated low pressure NGL plant configurations.

In their configuration ethane recovery was at least 85 mol% and

propane recovery was at least 99 mol%. The term “low pressurefeed gas” refers to a pressure that was at or below about 77bar and

even less. A typical feed gas flow rate of 100 kmole/h supplied at

about 42.4 bar and 154.4◦C with a composition of typically 1% N2,

0.9% CO2, 92.35%C1, 4.25% C2, 0.95% C3, 0.2% iC4, 0.25% nC4 and 0.1%

C5+ was used. Additional cooling was provided via external ethane

and propane refrigerants.

A significant cost in the NGL recovery processes is related to the

refrigeration required to chill the inlet gas. Refrigeration for those

lowtemperature schemes is generallyprovidedby usingpropaneasrefrigerant. In some applications, mixed refrigerants and a cascade

refrigeration cycle have been used. Refrigeration is also provided

by turbo expansion or work-expansion of the compressed natural

gas feed with appropriate heat exchange [13].

Yao et al. [14] discloses an open cycle self-refrigeration scheme

which aims to improve the efficiency and economy of NGL recov-

ery processes. In this process, a portion of a hydrocarbon liquid is

withdrawn from the lower portion of a distillation column. This

withdrawn liquid hydrocarbon is expanded and heated to produce

a two-phasesystem forseparation into a heavy,liquid hydrocarbon

product and a vapor phase for recycling to the column, preferably

as an enhancement vapor. The withdrawn hydrocarbon liquid is

preferably heatedby indirect heat exchange with theinlet gas, thus

reducing or eliminating the external refrigeration requirements of the process. The expanded, heated vapor recycled to the column

increases the ethane andpropane concentration in the column, thus

reducing the traytemperature profile and increasing the separation

efficiency.

In Ref. [13] methods forimproving theefficiency of processesfor

the recovery of natural gas liquids from gas feed, e.g., raw natural

gas or a refinery or a petrochemical plant gas stream was intro-

duced. An internal refrigeration system consisting of an open cycle

refrigerant withdrawnfrom a distillationcolumn anda closedcycle

refrigerant derived from the open cycle refrigeration system was

discussed as a new configuration.

For the required refrigeration, the integrity of the process and

thenumber of cold boxes in a configurationare also very important

factor.

Combined composite cooling and heating curves for exchang-

ers show the performance of these exchangers. Most optimum

composite curves and high efficiency of cold boxes result in

better integration in the configuration. More integrated NGL

recovery plants need less fixed and operating costs for specific

products.

Cryogenic facilities have made extensive use of brazed-

aluminum plate-fin heat exchangers since the 1950s. Instead of a

shell andtube configuration, these units consist of channels formed

by a thin sheet of aluminum pressed into a corrugated pattern (the

fin) sandwiched between two aluminum plates. Each layer resem-

bles the end view of corrugated cardboard. The fin channels may

be straight or may have a ruffled or louvered pattern to interrupt

the straight flow path [15].

Tirandaziet al. [16] workedon a C2+ recovery plant.In that plantpropane refrigeration cycle supplied the required refrigeration. But

optimum designof theprocess which leads to high ethanerecovery

level (94+) is achieved only by two multi-stream heat exchangers.

Also this optimum design decreases the plant fixed costs signifi-

cantly.

In this study a novel process configurationis proposed andcom-

pared with the configuration which has been demonstrated in Ref.

[13]. This configuration uses an open–closed self-refrigeration sys-

tem. Next, the operation of the process is analyzed and the effect

of the plant performance through changing the operating parame-

ters is discussed. Also, the effect of the different feed compositions

on the process characteristics is investigated. In order to analyze

the need of external refrigeration by a close or open cycle that is

related to the composition of the natural gas feed a configurationwith external refrigeration is designed the manner that it is similar

with the purposed configuration in the separation section.




378 M. Mehrpooya et al. / Chemical Engineering and Processing 49 (2010) 376–388

Fig. 1. Process flow block diagram with internal refrigeration system.

2. Process description

2.1. Process with self-refrigeration

As illustrated in Fig. 1, a feed gas comprising a pretreated andclean natural gas or refinery gas stream is introduced into the illus-

trated process through inlet stream Feed Gas at a temperature of

about 37.7◦C and an elevated pressure of about 63bar. This stream

is cooledin themulti-stream heat exchangerMSHX-1to reduce the

temperature of the stream to about −32.5 ◦C. The output stream

from MSHX-1 follows to the D-1 flash drum for separation of the

condensed liquid, if any. A portion of the liquid is introduced into

the middle of demethanizer column for further fractionation. A J –T

valve decreases its temperature to about −52 ◦C before entering

to the tower. Another portion, stream 3, is expanded through the

expansion valve and fed to the demethanizer.

The outletvapor stream2 fromthe D-1 drumis divided intotwo

portions, the main portion 2 andthe remaining portion 6. The main

portion which is about 60%, is expanded through a work-expansionturbine Ex-1 prior to entering the demethanizer right below the

top rectifyingsection as expander discharge 8. The remaining vapor

portion 6 is cooled to substantially condensation, and in most cases

subcooling, to approximately −95 ◦C via MSHX-1. This subcooled

liquid stream 9 is expanded through the expansion valve to top of

the demethanizer as liquid reflux.

The demethanizer operated at approximately 25 bar is a distil-

lation column containing conventional kinds of trays applied in the

demethanizer towers. It is equipped with four liquid draw trays in

the lower section of the column to provide heat to the column for

stripping volatile components off from the bottom liquid product.

This is accomplished viathe useof twomulti-stream heat exchang-

ers.The sidedrawliquids14, 16, 18and20 toenter the MSHX-1and

MSHX-2 at −54, −53, −36 and −16.5◦

C respectively, and exit asstreams 15, 17, 19 and 37 at approximately −42, −41, −24 and

25 ◦C, respectively, prior to returning to the demethanizer.

The residue gas 7 exiting the upper portion of the demeth-

anizer is fed to the MSHX-2 exchanger, providing refrigeration

for condensing/subcooling the vapor splitsream 6 and subcooling

the liquid stream 3 from the D-1 drum. The residue gas exiting

the MSHX-2 is further warmed to near the feed gas temperaturevia MSHX-1. The warmed residue gas 11 leaving the MSHX-1 at

approximately 34 ◦C is sent to the suction of the expander com-

pressor C-1, where it is compressed to 29 bar by utilizing work

extracted from the expander Ex-1. Depending upon the needed

delivery pressure, a residue gas compressor C-2 may be needed to

further compress the residue gas stream 12 followed by an after-

cooler AC-1, prior to its final delivery at 62bar.

In this configuration the refrigeration provided by the residue

gas from the demethanizer, turbo expander Ex-1 and side liquid

draws is not sufficient to achieve high levels of ethane recovery.

So, a self-refrigeration system (similar to the one proposed in Ref.

[13]) is applied.

Stream 20, the open cycle refrigerant, is withdrawn from the

chimney tray of the demethanizer column; the resulting mixedrefrigerant is preferentially fed to the MSHX-1 for subcooling prior

tobeingexpandedthroughthe expansiondevice VLV-1at 9 bar. The

expanded streamis directed back to theMSHX-1 providingindirect

heat exchange with the Feed gas stream and thereafter fed to the

suction knockoutdrumD-2 where unvaporizedliquid,if any, is sep-

aratedwhile the refrigerant is used to cool the inlet gas stream. The

vapor stream24 producedin theknockoutdrumis withdrawnfrom

the topthereof to twostage recycle compressor C-3. The repressur-

ized gas stream 28 exiting compressor C-3 is cooled to 89◦C by the

MSHX-3 heat exchanger, and then it flows to the AC-2 air cooler

resulting in partial condensation. The partially condensed product

exiting theAC-2 cooleris introduced into separatorD-3 where con-

densed liquid is separated. A portion of the output liquid stream

withdrawn from separator D-3 (representing closed cycle refrig-erant) is used as refrigerant in the heat exchanger MSHX-3. In fact

this portion provides a part of the required refrigeration in the con-





Table 1

Material balance for the processes illustrated in Figs. 1 and 4.

Stream no. Fig. 1 Fig. 4

Temperature (◦C) Pressure (bar) Flow rate (kmol e/h) Temperature (◦C) Pressure (bar) Flow rate (kmole/h)

Feed gas 37.78 63.09 14942.28 37.78 63.09 14942.28

1 −30.00 63.09 14942.28 −29.00 63.09 –

2 – – 4126.94 −29.00 63.09 4181.663 – – 1330.22 −29.00 63.09 1296.35

4 – – 5116.24 – – 3689.61

5 −63.25 26.80 1330.22 −73.52 25.80 1296.35

6 −30.00 63.09 5699.11 −29.00 63.09 5774.67

7 −93.87 25.86 11349.28 −95.50 23.86 11351.04

8 −66.21 26.00 4126.94 −66.53 25.00 4181.66

9 −94.70 26.50 5699.11 −94.97 – 5774.67

10 −36.41 25.86 11349.28 −45.00 23.86 11351.04

11 35.00 25.86 11349.28 10.00 23.86 –

12 43.83 28.37 11349.28 19.79 26.60 11348.39

13 62.00 11349.28 185577.88 101.87 62.00 –

14 −54.87 26.14 2000.15 −55.29 23.87 2200.00

15 −42.00 26.14 1999.68 −43.00 – –

16 −51.11 26.53 2199.88 −51.88 23.88 2000.00

17 −40.00 26.53 2199.88 −38.00 – –

18 −20.94 26.75 3549.97 −31.34 23.89 3200.02

19 −3.00 26.75 3550.19 −20.00 – –20 2.42 26.82 3580.29 −7.70 23.89 –

21 −33.42 9.00 3580.29 35.00 – –

22 35.73 9.00 3580.29

23 6.77 9.00 450.00 10.86 6.50 672.26

24 30.61 9.00 3635.76 10.86 6.50 2527.74

25 30.61 9.00 394.54 −17.31 2.68 –

26 31.97 26.89 394.54 −17.31 2.68 2172.08

27 20.00 14.00 14.32 −17.84 2.63 355.65

28 94.20 26.89 3650.10 −17.31 2.68 355.65

29 91.42 26.89 3650.10 −18.01 2.61 858.64

30 35.00 26.89 3650.10 −38.55 1.18 –

31 35.00 26.89 2006.96 −38.52 1.18 858.64

32 35.00 26.89 1150.20 −15.65 1.17 –

33 −18.01 2.61 1669.10

34 14.70 14.00 492.94 −15.65 1.17 858.64

35 20.00 14.00 492.94 −18.00 2.61 2172.08

36 21.28 26.89 28.62 65.29 13.00 3200.00

37 39.29 26.89 3580.32 37.75 13.00 –

38 10.86 6.50 –

Pipeline gas 64.51 62.00 11349.28 64.62 62.00 11348.39

Liquid product 32.84 26.89 3591.57 25.64 23.89 3591.07

densation section of open cycle refrigeration system. The flow rate

of stream 34, pressure drop in the VLV-4 expansion device and the

temperature of stream35 areparameters which shouldbe adjusted

based on theopen refrigeration cycle performance.Also these three

parameters affect the air cooler performance and its design condi-

tion. Consequently a trade off between the compressors shaft work

and condensation costs (fixed and operating costs related to the air

cooler) will determine their optimum value.A portion of theoutlet streamfrom MSHX-3flows to D-4knock-

out drum as the closed cycle separator. The vapor product of D-4,

27 is introduced to the second stage of the C-3 compressor.

After indirect heat exchange with one or more process

streams, the heated open and closed refrigerants are preferably

combined for simplicity and introduced into suction knockout

drum D-2 where the vaporized refrigerant is separated. The

vapor stream 24 is then introduced to the first stage recycle

compressor C-3.

The D-4 liquid product pressure is increased by the P-2 pump,

next this stream, 36 is mixed with stream 26 (separator D-2 liq-

uid product). Finally stream 37 is introduced to the bottom of

the demethanizer column (stream 37). The pressure, temperature

and flow rate of the process streams are presented in Table 1.Figs. 2 and 3 show the MSHX-1 and MSHX-2 composite curves for

Fig. 1 process.

2.2. Process with external propane refrigeration

As illustrated in Fig. 4, a feed gas comprising a pretreated and

clean natural gas or refinery gas stream is introduced into the illus-

trated process through inlet stream Feed Gas at a temperature of

Fig. 2. Composite curves for MSHX-1 in Fig. 1.






about 37.7◦C and an elevated pressure of about 63bar. This stream

is cooled in the multi-stream heat exchanger MSHX-1 to reducethe temperature of the stream to about −29 ◦C. The output stream

from MSHX-1 follows to the D-1 flash drum for separation of the

condensed liquid, if any. A portion of the liquid is introduced into

the middle of demethanizer column for further fractionation. A J –T

valve decreases its temperature to about−47.65 ◦C before entering

to the tower. Another portion, stream 3, is expanded through the

expansion valve and fed to the demethanizer.

The outlet vapor stream 2 from D-1 drum is divided into two

portions, the main portion 2 and the remaining portion 6.The main

portion which is about 42%, is expanded through a work-expansion

turbine EX-1 prior to entering the demethanizer right below the

top rectifyingsection as expander discharge 8. The remaining vapor

portion 6 is cooled to substantially condensation, and in most cases

subcooling, to approximately−71.5 ◦C via MSHX-1. This subcooled

liquid stream 9 is expanded through the expansion valve to top of

the demethanizer as liquid reflux.

The demethanizer operated at approximately 23 bar is a dis-

tillation column containing conventional kinds of trays applied in

the demethanizer towers. It is equipped with four liquid draw

trays in the lower section of the column to provide heat to the

column for stripping volatile components off from the bottom liq-uid product. This is accomplished via the use of two multi-stream

heat exchangers. The side draw liquids 14, 16, 18 and 20 to enter

the MSHX-1and MSHX-2 at −55.29, −51.88, −31.34 and −7.70 ◦C

respectively, andexit as streams 15, 17, 19 and21 at approximately

−43, −38, −20 and 35 ◦C, respectively, prior to returning to the

demethanizer.

The residue gas 7 exiting the upper portion of the demeth-

anizer is fed to the MSHX-2 exchanger, providing refrigeration

for condensing/subcooling the vapor splitsream 6 and subcooling

the liquid stream 3 from the D-1 drum. The residue gas exiting

the MSHX-2 is further warmed to near the feed gas temperature

via MSHX-1. The warmed residue gas 11 leaving the MSHX-1 at

approximately 10 ◦C is sent to the suction of the expander com-

pressor C-1, where it is compressed to 26.6 bar by utilizing workextracted from the expander Ex-1. Depending upon the needed

delivery pressure, a residue gas compressor C-2 may be needed to

further compress the residue gas stream 12 followed by an after-

cooler AC-1, prior to its final delivery at 62bar.

In this configuration the refrigeration provided by the residue

gas from the demethanizer, turbo expander EX-1 and side liquid

draws is not sufficient to achieve high levels of ethane recovery.

So, a three stage propane refrigeration cycle is applied in order

to compare the performance of the process with the configura-

tion explained in Section 2.1. It should be noted that based on the

potentialof the proposed configuration, distribution of the propane

refrigerant in the evaporators (MSHX-1 and MSHX-2) was done

in the manner that the maximum cold recovery is gained. Also

integrity of the process was considered though using the external

Fig. 4. Process flow block diagram with external refrigeration system.






refrigeration system. Figs. 5 and 6 show the MSHX-1and MSHX-2

composite curves for Fig. 4 process.

Outletstreamfromthe aircoolerAC-2(condenser ofthe cycle)inthe temperature of 37.7 ◦C and pressure of 13 bar follows to VLV-4

and its pressure is decreased to 6.5 bar. Stream38 is sent to the D-2

flash drum. Gas product of the D-2 enters the 3rd stage of a three

stage compressor C-3. The outlet liquid propane from D-2 is split

into twoparts after passing an expansiondevice. A portion of it,85%

follows to MSHX-1 in and provides the required refrigeration for

cooling the feed gas. Outlet stream from MSHX-1 enters D-4 flash

drum at −18 ◦C and 2.61bar. Another portion, 28 follows to D-4

after passing VLV-7 expansion valve. The gas product of D-4 enters

the 2nd stage of C-3 compressor and the liquid product which is

the1st stage refrigerant of thecycle follows to theMSHX-2 through

VLV-6expansionvalve at−38.55 ◦C and1.18bar.The outletstream,

31 is not vaporized totally in this heat exchanger so it is sent to the

MSHX-1 as second evaporator. Finally the vaporized refrigerant 32with minimum pressure enters the D-3 before following to the 1st

stage of C-3 compressor at −15.65 ◦C and 1.17 bar. The pressure,

temperature and flow rate of the process streams are presented in

Table 1.

A comparison between Figs. 3 and 5 and Figs. 5 and 6 shows

that the multi-stream heat exchangers in Fig. 1 process are more

efficient than the ones which used in the process with external

refrigeration system (Fig. 4).

3. Analysis of the process operation

Table 2 shows a comparison between the configuration sug-

gested in Ref. [13] and the configuration presented in Fig. 1.


Table 2

Overall performance of the configuration (Fig. 1).

Ref. [13] This work

% Ethane recovery 90.1 91.41

% Propane recovery 99.4 99.4

Self-refrigeration compression, kW 4113 3475.36

Propane refrigeration, kW 0 0

Residue gas compression, kW 8904 9042

As can be seen, self-refrigeration compression power is 15.5%

less than [13]. But residue gas compression power is 1.5% more.

In the case of the percentage of ethane recovery we can reach to

ethane recoveries 1.45% higher than [13].

There are two multi-stream heat exchangers in this config-

uration. This was the minimum number of the required heat

exchangers which can be used. In [13], 3–5 heat exchangers were

applied in the flow sheet. With reviewing the literature it can

be deduced that a new generation of compact heat exchang-

ers turn the cryogenic liquid recovery processes to integrated

ones. Also with decreasing the number of heat exchangers the

problems related to the heat exchanger networks design is

decreased.

In this configuration feed gas is cooled in two steps by MSHX-

1 and MSHX-2. In other words there are two heat exchangers for

two kinds of cold streams: the cold streams whose temperatures

are very low (streams which exit from the top section of the col-

umn)and the coldstreamswhose temperatures are higher(streams

which exit from thelower sections of thecolumn).This methodhas

twoadvantages: first it made it possible to use the cold process and

cycle streams in themanner that maximum heat recoveryis gained.

Second, it enables us to have the most efficient composite curves

for heat exchangers.

It should be noted that adjusting the operating condition in the

process is a very important point which can affect the exchanger’s

performance significantly. But more heat exchangers create a big-

ger optimization problem.It may be possible to design a configuration with one multi-

streamheat exchanger butcontrol of thecold streams temperature

level for reaching the most optimum performance will be limited.

The shape of the combined composite cooling and heating

curves for exchangers is the determining factor in the performance

of such equipments. Also, the pinch point in these curves should

be avoided. Physical abilities of the heat exchanger determine the

minimum allowable temperature difference. In this work the min-

imum temperature difference in most of the under consideration

cases was 2 ◦C.

MSHX-1 reduces the temperature of the feed to −30 ◦C. Fig. 2

shows the composite curve of this device. As can be seen, the

required refrigeration in this device is gained from two sources

(open refrigeration system and demethanizer side streams). Thisdevice uses the warmer cold streams in the process and provides

the feed stream for entering the MSHX-2.

MSHX-2 uses stream 7 as the coolest stream in the process. Side

draws 14 and 15 are the other cold streams which enter MSHX-2.

Fig. 3 shows the composite curve of this device.

3.1. Mixed refrigerant systems

Using mixtures as refrigerants in the design of refrigeration

systems offers significant opportunities in the search for more

energy-efficient and compact designs. However, the design of

mixed-refrigerant systems is extremely difficult, and few suc-

cessful design methods are available. As a result, many existing

operations can be far from the optimal conditions.The difficulty in designing MR systems mainly stems from two

aspects: one is the expensive andhighly nonlinear nature of compu-





Table 3

Effect of changing stream 21 pressure.

Stream 21 pressure (bar)

10 9.5 9 8.5 8 7.5

% Ethane recovery 89.13 90.68 90.63 93.11 93.97 93.74

% Propane recovery 99.1 99.04 99.4 99.6 99.48 99.54

Self-refrigeration compression, kW 3109.986 3344.913 3594.756 3885.618 4166.039 4458.392Residue gas compression, kW 8816.848 8902.615 9094.285 9126.355 9198.697 9312.805

MSHEX1 minimum temperature approach,◦C 2.16 2.2 2.1 2.16 2.19 2.63

MSHEX2 minimum temperature approach,◦C 1.98 2 2 1.96 2 2

Table 4

Effect of changing stream 20 flow rate.

Refrigerant pressure drop (bar)

17.82 17.82 17.82 17.82 17.82 17.82 19.32

3700

(kmole/h)a

3650

(kmole/h)a

3580

(kmole/h)a

3520

(kmole/h)a

3480

(kmole/h)a

3450

(kmole/h)a

3450

(kmole/h)a

% Ethane recovery 90.67 90.63 90.93 88.6 90.2 89.65 90.04

% Propane recovery 99.763 99.4 99.51 99.1 99.72 99 99.65

Self-refrigeration compression, kW 3651.4 3594.8 3534 3436.6 3472 3381 4101.9

Residue gas compression, kW 9156.2 9094.3 9161 8837 8789 8897 8994.3MSHEX1 minimum temperature approach,◦C 2.2 2 1.77 2.6 2.42 1.77 2.77

MSHEX2 minimum temperature approach,◦C 2.4 2.01 2 2.2 2.23 1.78 2.4

a Stream 82 flow rate (kmole/h).

tation, and the other is the sensitivity of the systems to operating

changes, especially changes in the composition of the refrigerant

mixture [17].

The required refrigeration in this configuration is supplied by

a self-refrigeration system. Consequently because of the refriger-

ant composition, it is a MR cycle and all problems in such cycles

should be considered for finding the optimum performance. More-

over in theheat exchangers, theprocess is affected directly through

changing the variables of the MR system itself. In this process it is

not possible to control the refrigerant composition but other effec-tive parameters such as the refrigerant flow rate and refrigerant

pressure drop are discussed in next sections.

3.1.1. Effect of the refrigerant pressure drop

Table 3 shows the process characteristics when stream 21 pres-

sure changes by an expansion device (VLV-1). High amount of

ethane recoveries is attainable when refrigerant pressure drop is

increased. But more compression power will be needed for the

self-refrigeration system and residue gas.

3.1.2. Effect of the refrigerant flow rate

As shown in Table 4, the self-refrigeration compression power

and residue gas compression powerdecreaseas the refrigerant flow

rate decreases. But in such situations for having high amounts of liquid recoveries it is necessary to use more efficient multi-stream

heat exchangers which can perform in minimum temperature

approaches lower than 2 ◦C.

3.2. Effect of the column side streams flow rate

As can be seen in Fig. 1 there are three side streams in this con-

figuration. Adjusting the flow rate and output temperature from

heat exchangers of these streams are very important parameters.

The shapes of the heat exchangers composite curves are a function

of temperature and flow rate. Also performance of the column andseparation efficiency is affected when flow rate and temperature

of the side streams change. For adjusting the operating condi-

tion in the process an optimization problem should be solved. A

multi-objective function can be defined in which all important

parameters like liquid recoveries and needed compression power

enters. Table 5 presents the optimum values of the abovesaid vari-

ables for various defined conditions.

As the methane mole fraction decreases stream 14 tempera-

ture increases and its flow rate should be increased. It can be said

that with increasing themole fraction of heavy hydrocarbonsin the

feed stream, separation in higher temperatures is conceivable. But

as the mole fraction of heavy components increases, the flow rate

of the liquid product increases and consequently the flow rate of

the side streams which supply a part of the required refrigerationshould be increased. Accordingly, the flow rate of the streams 16

Table 5

Column side streams operating conditions for different feed compositions.

Stream no. Feed 1 Feed 2 Feed 3 Feed 4 Feed 5

14 Temperature −80.74 −65.76 −61.59 −58.9157 −53.58

Flow rate 2100.01 1999.98 2000.11 2499.99 2499.99

15 Temperature −39 −42 −42 −52.00 −49

16 Temperature −64.81 −54.54 −53.79 −52.9535 −56.31

Flow rate 2200.02 1500.13 1999.94 2200.14 3500.05

17 Temperature −38 −35 −35 −45.00 −52

18 Temperature −

21.65 −

19.45 −

18.88 −

24.48 −

26.28Flow rate 2400.12 2779.97 2899.98 4150.11 4500.02

19 Temperature 2.00 3.00 1.00 −5.00 −5.00





and18 shouldbe increasedas themethane mole fraction decreases.

In this regard the temperature of the side streams which exit cold

boxes decreases. The reason is that in such situations (heavier feed)

high liquid recoveries can be gained in higher temperature lev-

els. Finally adjusting the side streams flow rate and temperature

depends on the constraints which designers define in order to opti-

mize their objective function. The operating condition in Table 5 isselected regarding to the performance of the process with external

refrigeration (Fig. 4).

3.3. Effect of the feed composition

The gas composition has a major impact on the economics of

NGL recovery and the process selection. In general, gas with a

greater quantity of liquefiable hydrocarbons produces a greater

quantity of products and hence greater revenues for the gas pro-

cessing facility. Richer gas also entails larger refrigeration duties,

larger heat exchange surfaces and higher capital cost for a given

recovery efficiency. Leaner gases generally require more severe

processing conditions (lower temperatures) to achieve high recov-

ery efficiencies [4].

Propane refrigeration is often required to help in condensing theheavy components for a rich gas. Using the external refrigeration

in an ethane recovery plant depends on the feed gas composition.

In the plants which the feed gas is rich, in order to obtain high

ethane recoveries (90+%) applying an external refrigeration sys-

tem is necessary. So ethane recovery is a determining factor which

provokes using the external refrigeration in the case of rich gases.

Whereas there is no need to use the external refrigeration for lean

gases andhigh ethane recoveries (99%) will be possible forthem. In

Ref. [5] four configuration (GSP, CRR, RSVE, RSV) were discussed

which they can recover ethane higher than 90% with no exter-

nal refrigeration. Also they showed the ethane recovery versus the

compression power for each of the configurations in a figure. But

the feed composition in the patents which CRR [18], RSV [19] and

RSVE [20] configurations was introduced contain 92.5% methane.In GSP [21], two different feed gases were considered as samples,

one with 93.82 mol% of methane and another with 87.86mol%.

In Ref. [22] the inlet gas composition was divided into three

groups. Lean with 87.66% methane and 6% ethane, normal with

85.9% methane and 7% ethane and rich with 83.9% methane and

8% ethane. Two different configurations (GSP and RRP) were used

for analyzing the cryogenic process design alternatives for a new

project. Compression power, ethane recovery and inlet composi-

tion were criteria for their comparisons. In a fixed compression

power(6800HP) ethanerecoverywas changed from 93% to78% and

99% to 79% for GSP and RRP respectively when the inlet feed com-

position was shift from the normal to the rich. It should be noted

that GSP and RRP does not use the external refrigeration. Chebbi et

Table 6

Composition of feed stream.

Feed gas composition mole fraction

C1 C2 C3 C4+ Non-hydrocarbons

Feed 1 0.830 0.067 0.051 0.044 0.007

Feed 2 0.800 0.080 0.060 0.053 0.008

Feed 3 0.780 0.097 0.061 0.054 0.007Feed 4 [13] 0.750 0.110 0.070 0.061 0.008

Feed 5 0.700 0.130 0.083 0.073 0.010

Feed 6 0.670 0.145 0.092 0.008 0.010

al. [23] didtheir analysis with twodifferent feed gases. Feed A with

93% methane and 3%ethane and feed D with 69% methane and 15%

ethane.

Lee et al. [13] introduced their configuration in which the inlet

gas was containing 75% methane and 11% ethane (feed 4). So it can

be said that this composition is very richer than the previous cases.

In this study our concern was the rich gases, but because there is

not an exact division for the richness of the inlet gases an analysis

was performed about this point. Tables 7 and 8 show the required

refrigeration (external andinternal) foreach of inlet gases.As it wasexpected the richer inlet gas the more refrigeration is required. Six

different feed gases were presented in Table 6.

In this part the performance of the process versus six feed gases

with different values of hydrocarbon and non-hydrocarbon com-

ponents is analyzed. Tables 7 and 8 present the characteristics of

the processes illustrated in Figs. 1 and 4 respectively.

In Table 6 methane, ethane and propane mole fractions change

from 0.83 to 0.67, 0.067 to 0.145 and 0.051 to 0.092 respectively.

This wide range covers most conventional rich feed gas composi-

tions.

Tables 7 and 8 show that how process characteristics change

when inlet composition changes both in the processes which use

internal and external refrigeration. Thus a comparison between

these two kinds of configurations (Figs. 1 and 4) will be possiblewith elaboration. As it was explained, Fig. 4 was designed because

we wanted to have an exact collation between the performance of

the internal and external refrigeration system in a configuration.

Accordingly based on the operating condition (temperature, pres-

sure andflowrate)it ispossibleto assumethatsizeof theseparation

system equipments; (demethanizer column), C-1, C-2, EX-1, AC-1

and D-1 are equal in both processes. Also MSHX-1 and MSHX-2

in Fig. 4 only have one more side towards Fig. 1. Consequently it

can be said that the fixed cost of the abovesaid equipments are

equal.

In the refrigeration system, AC-2 air cooler was designed with

equal size (UA in both configurations has the same value) in both

processes. C-3in Fig.1 is a twostagecompressor whereas in Fig.4 it

Table 7

Overall performance of the process for the different feed compositions (Fig. 1).

Feed 1 Feed 2 Feed 3 Feed 4 Feed 5 Feed 6

% Ethane recovery 94.00 93.91 93.36 91.41 92.41 90.34


% Methane recovery 0.99 99.13 99.12 98.95 98.55 98.17



Self-refrigeration compression, kW 2272.57 2664.42 3095.82 3475.36 4105.62 4385.56

Residue gas compression, kW 9784.74 9334.59 9213.96 9042.51 8608.91 8321.15

Duty of refrigeration system air cooler, kW −6271.61 −8201.20 −8803.4394 −7874.10 −11656.53 −12315.83

Duty of residue gas air cooler, kW −8836.40 −8507.09 −8232.67 −9956.04 −7464.29 −7134.98

Heat flow of feed stream, kW −334912.74 −339307.05 −341035.43 −345611.94 −352546.70 −358046.23

Heat flow of liquid product, kW −81909.12 −97429.99 −104636.88 −118508.52 −140981.39 −156654.16

Heat flow of gas product, kW −256054.33 −246586.35 −241124.89 −232415.69 −217971.60 −208136.17

Fan power in the residue gas air cooler, kW 79.86 76.88 74.40 71.92 67.46 64.48Fan power in the air cooler of refrigeration system, kW 55.06 67.46 79.37 89.98 104.17 111.11

Overall required power, kW 12192.22 12143.35 12463.54 12679.78 12886.16 12882.30





Table 8

Overall performance of the process for the different feed compositions (Fig. 4).

Feed 1 Feed 2 Feed 3 Feed 4 Feed 5 Feed 6

% Ethane recovery 94 93.89 93.52 91.6 92.5 90.15


% Methane recovery 99.2 98.95 99 99 97.91 97.05


MSHEX2 minimum temperature approach,◦C 2.14 1.95 2.71 3 2.28 2.16

Propane refrigeration, kW 2972.356 3839.229 4152.029 4272.29 4852.752 5078.086

Residue gas compression, kW 9783.394 9442.384 9314.823 9020.75 8763.665 8591.18

Duty of refrigeration system air cooler, kW −9946.377 −12659.62 −14348.12 −14558 −16476.32 −16974.5

Duty of residue gas air cooler, kW −6476.368 −6001.163 −5323.794 −4946.4 −5434.938 −5927.524

Heat flow of feed stream, kW −334912.7 −339307.1 −341035.4 −345612 −352546.7 −358046.2

Heat flow of liquid product, kW −82823.76 −98315.92 −106064.6 −119402 −144425.5 −161563.7

Heat flow of gas product, kW −255756 −246370.3 −241175.9 −232421 −216416 −205715.3

Fan power in the residue gas air cooler, kW 58.53175 49.60317 48.11508 46.131 49.60317 53.57143

Fan power in the air cooler of refrigeration system, kW 94.24603 114.4144 126.5476 131.071 148.9087 177.5849

Overall required power, kW 12908.53 13445.63 13641.51 13470.2 13814.93 13900.42

is a three stage compressor. Aboutflash drums,thereare three ones

(D-2, D-3 and D-4) in both configurations whereas the size of them

in Fig.4 is greater than that shownin Fig.1.Thereforethere are equal

number of equipments in the internal and external refrigerationsystems but the size of them are smaller in the internal system.

In such condition the operating costs is determining factor

which should be considered in order to compare the performance

of Figs. 1 and 4 when the inlet gas composition changes. Fig. 7

illustrates the required compression power versus the feed com-

position for the internal and external refrigeration systems. As can

be seenthe compression power increase almostlinearlyas theinlet

gas being richer. Also the residue gas compression power and fan

power in air cooler AC-1 decreases with increasing the richness of

the feed (Tables 7 and 8). It is because the methane content of the

inlet gas decreases and consequently the less residue gas is gained.

The fan power of AC-2 air cooler should be increased as the

inlet gas being richer. It is because the richer inlet gas needs the

more refrigerant flow rates and consequently the higher air flowrate should be consumed in an air cooler. The difference between

the refrigerant flow rate in the external and internal refrigeration

systems has been caused that the required fan power in the exter-

nal refrigeration system to be greater. Accordingly the required fan

power decreases as the inlet gas being richer in the AC-1 air cooler.

Butthe temperature of theinputstreamto theAC-1in Fig.1 process

is the higher than Fig. 4, so the more air flow should be consumed

in this process.

In all cases the ethane recovery almost was fixed for both inter-

nal and external systems in the manner that the temperature

approach in MSHX-1and MSHX-2for theinternalrefrigeration sys-

tem was adjusted to the least value bigger than 2 ◦C.Thus ascan be

Fig. 7. Required refrigeration for the different inlet gases.

seen with increasing the portion of heavy hydrocarbons the ethane

recovery decreases.

In mixtures with a high percent of methane (more than 0.8),

the performance of the processis very good, and a high percentageof ethane recoveries and low self-refrigeration compression power

is accessible. The reason for increase in residue gas compression

power is the methane content of the feed gas.

In mixtures with high amounts of C2+ hydrocarbons, the

required self-refrigeration compression power is increased as

methane mole fraction decreases. Accordingly the required

residue gas compression power decreases. Nonetheless ethane and

propane recoveries can be controlled in desired values (90+). Also

the required compression power in the process lies in similar

levels.

However the minimum temperature approach is fixed in [2 2.2]

but the shape of these curves changes as the feed composition

changes. Figs. 8 and 9 show the composite curves for feed 1 and

Figs. 10 and 11 show the composite curves for feed 6.

4. Numerical implementation

All aforementionedanalyses were tested by conventional chem-

ical process simulators (like HYSYS and Aspen). Also with the data

in Table 1 it is possible to simulate this configuration and evaluate

the process characteristics by chemical process simulators.

4.1. Selection of the equation of state

The more similar the character of the mixture molecules, the

more orderly their behavior. A single component system composed

Fig. 8. Composite curves for MSHX-1 (feed 1).







entirely of a simple, relatively spherical molecule like methane

behaves in a very predictable, correctable manner. The more dis-

similar the molecules, the less accurate the prediction becomes

[4].

HYSYS provides enhanced equations of state (PR and PRSV) for

rigorous treatment of hydrocarbon systems; semiempirical and

vaporpressure models for the heavier hydrocarbon systems; steam

correlations for accurate steam property predictions; and activity

coefficient models for chemical systems. For oil, gas and petro-

chemical applications, the Peng–Robinson EOS (PR) is generally

the recommended property package. The PR equation of state has


been enhanced to yield accurate phase equilibrium calculations for

systems ranging from low temperature cryogenic systems to high

temperature, high pressure reservoir systems. As an alternate, the

PRSV equation of state should also be considered. It can handle the

same systems as the PR equation with equivalent, or better accu-

racy, plus it is more suitable for handling moderately non-ideal

systems. PR is ideal for VLE calculations as well as calculating liq-uid densities for hydrocarbon systems. Several enhancements to

the original PR model have been made to extend its range of appli-

cability and to improve its predictions for some non-ideal systems.

PRSV is a two-fold modification of the PR equation of state that

extends the application of the original PR method for moderately

non-ideal systems [24].

Mehrpooya et al. showed that with PRSV equation of state it is

possible to simulate a currently in operation NGL recovery plant

which was fed by a treated natural gas mixture, with acceptable

accuracy [7–9]. Also PR was selected as equation of state for simu-

lationof the NGL recovery plants in Ref. [23]. Based on themixtures

discussed in Table 6, PRSV equation of state was selected in order

to calculate the thermodynamic properties and based on the feed

composition it can be said that the simulator can estimate thermo-dynamic properties with good accuracy so theresults of simulation

are acceptable for this process.

4.2. Algorithm and basic equation used for simulation

4.2.1. Demethanizer column

An absorber was defined for simulating the demethanizer col-

umn.

Column specs: Four specs shouldbe definedfor theabsorber with

side streams until the degree of freedom becomes zero. The flow

rates of side streams were set as column specs.

Solving method: Inside-out: With the “inside-out” based algo-

rithms, simple equilibrium and enthalpy models are used in the

inner loop to solve theoverall componentand heat balances as wellas any specifications. The outer loop updates the simple thermo-

dynamic models with rigorous model calculations. The inside-out

algorithmhas becomeone of the most popular methods because of

its robustness and its ability to solve a wide variety of columns. The

concept of this method and details about the algorithm and basic

equations can be found in Ref. [25].

Damping factor : The damping factor controls the step size used

in the outerloop whenupdatingthe simple thermodynamic models

used in the inner loop. For the vast majority of hydrocarbon-

oriented towers, the value of 1.0 is appropriate, which permits a

full adjustment step.

4.2.2. Air coolers

The air cooler uses the same basic equation as the heatexchanger unit operation. However, the air cooler operation can

calculate the flow of air based on the fan rating information. The

air cooler calculations are based on an energy balance between the

air and process streams. For a cross-current air cooler, the energy

balance is shown as follows:

M air(H out −H in) air =M process(H in −H out) process (1)

where:

M air = air stream mass flow rate

M process = process stream mass flow rate

H = enthalpy

The air cooler duty, Q , is defined in terms of the overall heat

transfer coefficient; the area available for heat exchange and the





log mean temperature difference:

Q = UA DTLM Ft (2)

where:

U = overall heat transfer coefficient

A = surface area available for heat transferDTLM= log mean temperature difference (LMTD)

Ft = correction factor

The LMTD correction factor, Ft, is calculated from the geometry

and configuration of the air cooler.

The air flow through the fan is calculated using a linear relation:

fanairflow =speed

design speed × design flow (3)

Each fan in the air cooler contributes to the air flow through the

cooler. The total air flow is calculated as follows [24]:

total airflow =

fan air flow (4)

The optimum air temperature rise across the tubes may be esti-

mated by the equation [21]:

t 2 − t 1 = .005U

T 2 + T 1

2 − t 1

(5)

where:

t 2 = outlet air temperature

t 1 = inlet air temperature

T 2 = temperature of process fluid out

T 1 = temperature of process fluid in

U = overall heat transfer coefficient

The fan power requirements can be estimated from the equa-

tion:

kW =P a Q a

efficiency (6)

where:

Q a = air flow rate

P a = air pressure drop in cooler

In this study air flow rate is calculated by the HYSYS software.

The value of .7% was selected for efficiency based on the process

and recommended values in Ref. [21]. Inlet air temperature was

selected 30 ◦C and outlet air temperature is calculated by HYSYS.

But it is controlled by the simulator through internal calculations

and the value of air outlet is selected a maximum of 30 ◦C morethan the air inlet.

4.2.3. LNG heat exchangers

The LNG (liquefied natural gas) exchanger model solves heat

and material balances for multi-stream heat exchangers and heat

exchanger networks. The solution method can handle a wide vari-

ety of specified and unknown variables. For the overall exchanger,

you can specify various parameters, including heat leak/heat loss,

UA or temperature approaches. Two solution approaches are

employed; in the case of a single unknown, the solution is cal-

culated directly from an energy balance. In the case of multiple

unknown variables, an iterative approach is used to determine the

solution which satisfies not only the energy balance, but also any

constraints, such as temperature approach or UA. For the weightedmethod, the heating curves are broken into intervals, which then

exchange energy individually. An LMTD and UA are calculated for

each interval in the heat curve and summedto calculate the overall

exchanger UA.

Intervals: The number of intervals, applicable only to the

weighted rating method, may be specified. For nonlinear tempera-

ture profiles, more intervals will be necessary.

The Solver group box includes the solving parameters used for

LNG’s:Tolerance: You may set the calculation error tolerance (toler-

ance = .0001).

Current error : When thecurrent error is less than thecalculation

tolerance, the solution is considered to have converged (current

error= 2.528×10−6).

Iterations: The current iteration of the outer loop is displayed. In

the outer loop, the heat curve is updated and the property package

calculations are performed. Nonrigorous property calculations are

performed in the inner loop. Any constraints are also considered in

the inner loop (iterations = 10) [24].

4.2.4. Compressor/expander

For an adiabatic compressor or expander, HYSYS calculates the

compression (or expansion) rigorously by following the isentropicline from the inlet to outlet pressure. Using the enthalpy at that

point, as well as the specified efficiency, HYSYS then determines

the actual outlet enthalpy. From this value and the outlet pressure,

the outlet temperature is determined.

Depending on whether the process is an expansion or compres-

sion, the work determined for the mechanically reversible process

is multiplied or divided by an efficiency to give the actual work. In

this study 75% adiabatic efficiency was selected [24].

For the compressor:

powerrequiredactual = heat flow outlet− heat fl ow i nlet (7)

For the expander:

powerproducedactual = heatflowinlet− heat flow outlet (8)

4.2.5. Recycle

The capability of any flow sheet simulator to solve recycles reli-

ably andefficiently is critical.The recycle installs a theoretical block

in the process stream. The feed into the block is termed the calcu-

latedrecyclestream,and the product is the assumed recycle stream.

The following steps take place during the convergence process:

1. HYSYS uses the conditions of the assumed stream and solves the

flow sheet up to the calculated stream.

2. HYSYS then compares the values of the calculated stream to

those in the assumed stream.

3. Based on the difference between the values, HYSYS modifies the

values in the calculated stream and passes the modified values

to the assumed stream.4. The calculation process repeats until the values in the calcu-

latedstreammatch thosein the assumed stream within specified

tolerances [24].

In this work we use the minimum number of required recycles

(four) for running the simulator. Increasing the number of recy-

cles may decrease the possibility of convergence of the simulation.

Also, the accuracy of the calculations may be affected by recycles

in the flow sheet. This point will be important when a simula-

tion is used for function evaluation in the optimization algorithms

because optimization algorithms create different values for deci-

sion variables and the simulator should converge for each of them

many times. In this regard with an additional recycle, the simula-

tor can give wrong answers. So the minimum number of recyclesand their right positions should be found. This point is a function of

kind and position of the decision variables in the flow sheet and it





Table 9

Characteristics of the optimization problem.

Fig. 1 Fig. 4

Decision variables Decision variables

Flow rate, kmole/h Flow rate, kmole/h

x1: stream 14 molar flow x1: stream 14 molar

flow


flow


flow


flow


flow

Temperature, ◦C Temperature, ◦C

x6: stream 1 x6: stream 1







x13: stream 21 x13: stream 21x14: stream 22 x14: stream 31

Pressure, bar Pressure, bar

x15: demethanizer operating

pressure

x16: demethanizer

operating pressure



Constraints Constraints

C2 recovery, ≥91% C2 recovery, ≥91%

C3 recovery, ≥99% C3 recovery, ≥99%

2≤MSHEX1 minimum temperature

approach, ◦C ≤2.2

2≤MSHEX1 minimum

temperature approach, ◦C

≤3

2≤MSHEX2 minimum temperature

approach, ◦C ≤2.2

2≤MSHEX2 minimum

temperature approach, ◦C

≤3

Output air temperature in air coolers,≤60 ◦C

Output air temperaturein air coolers, ≤60 ◦C

Residue gas temperature, ◦C ≤64.5 Residue gas

temperature, ◦C ≤64.5

Objective function, kW Objective function, kW

Required power in the refrigeration

system

Required power in the

refrigeration system

will be found by trial and error. In this process the best position for

the minimum required recycles was found to be for column side

streams.

5. Optimal design of the plant

Moreover the process configurations there are many operat-

ing parameters which should be tuned in other to finding the

most optimal design. In this study an optimization problem was

solvedfor adjusting the optimal operating condition.Table 9 shows

the decision variables, constraints and objective function of the

optimization problem. Flow rate, pressure and temperature of

the process fluid are the parameters which can affect the perfor-

mance of the process significantly [26]. The objective function was

the required power in the refrigeration system, because the liq-

uid recoveries are controlled by the constraints. Also the residue

gas compression power is a function of liquid recoveries and the

required power in the refrigeration system. Integrity of the pro-cess can be control by the minimum temperature approach in the

multi-stream heat exchangers.

6. Conclusion

In thisstudy a novelNGL recovery configuration was introduced.

The required refrigeration is suppliedby a self-refrigeration system.

The results show that the self-refrigeration compression power is

15.5% lessthan [13]. Inthe case ofthe percentage ofethanerecovery

we can reach to ethane recoveries 1.45% higher than [13].Effectof themixedself-refrigerant flowrate andpressureon the

performance of the process was discussed. Higher liquidrecoveries

are possible with higher amounts of refrigerant flow rate and pres-

sure drop. High liquid recoveries (90+) are conceivable with this

configuration. For such processes the optimum number of multi-

stream heat exchangers is two. The combined composite cooling

and heating curves for these two heat exchangers show that they

perform efficiently. Also these curves show that the integrityof the

process is high. Various values for feed composition were tested

and the results demonstrated that the process can work efficiently

with different feeds.

Appendix A. Nomenclature

A-1 air cooler

C-1 compressor

C-2 compressor

C-3 compressor

C-4 compressor

MSHX-1 multi-stream heat exchanger

MSHX-2 multi-stream heat exchanger

VALV1 valve

VALV2 valve

VALV3 valve

VALV4 valve

D-1 flash drum

D-2 flash drumD-3 flash drum

D-4 flash drum

D-5 flash drum

References

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introducing a novel integrated ngl recovery process

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