Novel LNG-Based Integrated Process Configuration Alternatives for

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Novel LNG based integrated process configuration alternatives for co-production of LNG and NGL Mehdi mehrpooya, Mohammad Hossieni, and Ali Vatani Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/ie502370p • Publication Date (Web): 20 Oct 2014 Downloaded from http://pubs.acs.org on October 25, 2014

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Industrial & Engineering Chemistry Research

Novel LNG based integrated process configuration alternatives for co-production of LNG and NGL

1 2 3 4 5 6 7 8 9 10 11

Mehdi Mehrpooya ∗1, Mohammad Hossieni2 and Ali Vatani2 0F

1

Renewable Energies and Environment department, Faculty of New Sciences and Technologies, University of Tehran, Tehran, Iran.

2

School of Chemical Engineering, University College of Engineering, University of Tehran, P.O.Box: 113654563, Tehran, Iran.

12

Abstract:

13

In this study three novel process configurations for co-production of LNG and NGL is

14

introduced and analyzed. C3-MR, DMR and MFC refrigeration systems are used for supplying

15

the required refrigeration. High ethane recovery (90 %+) and low specific power (0.4

16

kWhr/kgLNG) for typical natural gas feed compositions are two of the basic characteristics of

17

the proposed configurations. The proposed processes compared to the conventional natural gas

18

liquefaction processes are simple and operable. Four or five multi stream heat exchangers and

19

one demethanizer column are utilized for co-production of LNG and NGL. Also the analysis

20

show that performance of the processes is efficient and comparable with similar cases.

21

Key words: LNG, NGL, Integration, Ethane Recovery, process design, refrigeration

22 23 24

∗

Corresponding Author: Tel: +98 21 61118564, Fax: +98 21 88617087 Email address: [email protected], [email protected]

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e

Specific flow exergy

Ex

Exergy

ΔT

thermal component

h

enthalpy

ΔP

pressure component

I

irreversibility rate

𝐦̇

Mass flow rate

N

Exergy loss number

P

Pressure

Q

Heat transfer rate

S

Entropy

T

Temperature

W

Work

Superscript

Names used for blocks in plants

Greek letters Efficiency

Δ

Gradient

Σ

Sum

αij

Relative volatility i/j

Subscripts cold

h

hot

i

inlet

o

outlet

id

ideal

ph

physical

ch

chemical

t

total

Q

heat rate

0

dead state

a

air

Compressor

TE-i

Turbo Expander

T-i

Tower

E-i

Multi stream heat exchanger

D-i

Flash Drum

V-i

Valve

AC-I

Air cooler

Abbreviation

µ

c

C-i

LNG

Liquefied natural gas

NGL

Natural gas liquids

NG

Natural Gas

MR

Mixed refrigerant

MFC

Mixed fluid cascade

DMR

Dual mixed refrigerant

SWHE

Spiral Wood Heat Exchanger

PFHE

Plate and Fin Heat Exchanger

C3-MR Propane precooling


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70 71 72 73


1. Introduction

74

One of the methods which is used for transportation of natural gas over long distances where

75

pipelines do not exist is liquefying. Natural gas can be liquefied at cryogenic temperatures as

76

Liquefied Natural Gas (LNG) [1]. In the other hand natural gas contains many desirable and

77

undesirable components that should be separated [2]. Natural gas liquids (NGL) have added

78

value and it is used in petrochemical processes as a main feed. [3]. Various kinds of process

79

configurations have been introduced for NGL recovery. There are several records in this area

80

which decreasing the capital and operating costs is the main subject of them [4-5]. Turbo

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expander is the simplest process for NGL recovery. Feed gas after pre cooling is sent to the

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expander and tower respectively. Residue Recycle (RR) was developed to achieve the NGL

83

recovery higher than 80 percent. Gas Sub cooled Process (GSP) was developed to overcome the

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problems encountered with the conventional expander process. The Cold Residue Recycle

85

(CRR) process is a modification of the GSP process to achieve higher ethane recovery levels

86

(98%). Another improvement of the turboexpander-based NGL process is the IPSI [5] Enhanced

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NGL Recovery Process. This process utilizes a slip stream from or near the bottom of the

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demethanizer as a mixed refrigerant. Mixed refrigerant stream is evaporated totally or partially

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and back to the column [4, 5]. There are several papers that have investigated the operating

90

condition of NGL recovery plants for improving the process efficiency [6-11]. Various kinds of

91

process configurations for natural gas liquefaction and NGL recovery process have been

92

introduced. An open-closed self-refrigerant system for NGL recovery was introduced in [4, 12].

93

The most important part in both of them is refrigeration system [13]. LNG processes can be

94

classified by refrigerant composition and refrigeration system. Some of them were introduced by 3 ACS Paragon Plus Environment


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Shukri [14]. Turbo expander, Dual Mixed Refrigerant (DMR), Single Mixed Refrigerant (SMR),

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cascade and Mixed Fluid Cascade (MFC) and propane precooling (C3-MR) are major applicable

97

processes for natural gas liquefaction. Mixed refrigerant cycles have better thermodynamic

98

efficiency [5, 13]. Five of the most conventional LNG processes was investigated in [15]. NGL

99

recovery and LNG production are done in cryogenic processes. In both of them refrigeration

100

system is a main part. Increasing level of the integration is a fundamental way for improving the

101

efficiency and decreasing the operating and capital costs [3]. Gas product of the NGL recovery

102

plants follows to the pipeline as sweet treated natural gas at about 55°C. However this gas leaves

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the demethanizer column at about -100°C, but it is used for supplying a portion of the required

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refrigeration for cooling the inlet feed before following to the pipeline. Using low temperature

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gas which leaves the demethanizer column directly to the liquefaction unit as a feed, is the main

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idea for design of integrated NGL/LNG process configurations. With integration efficiency

107

increases and total cost decreases [16]. Operating temperature in NGL recovery processes may

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reach to -100ºC in top of the demethanizer column [5]. Also in LNG processes operating

109

temperature reaches to near -162ºC. In non-integrated processes required refrigeration is supplied

110

from different cycles and separated heat exchangers while in integrated processes refrigeration is

111

provided by joint refrigeration cycles and in shared devices. On the other hand, it is clear that

112

LNG and NGL processes are series plants. Advantages of integration have been caused that

113

companies tend to design integrated processes. ConocoPhillips, APCI invented integrated

114

processes with especial design. For example, ConocoPhillips integrated process can produce

115

LNG approximately 7% more, while the required power is the same [3]. Fluor Technologies

116

claims that 10% energy saving is achievable by integration of LNG and NGL processes [17].

117

Weldon L. Ransbarger [18] introduced an integrated process with cascade arrangement and more


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flexibility in feed composition and operating condition. In this process propane is used for the

119

hottest cycle refrigeration, ethane for the middle cycle and an open methane cycle for

120

liquefaction and sub cooling the lean gas. A tower is used for NGL capturing after precooling

121

the feed and a recycle stream which is supplied from the liquefaction section. However this

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patent introduces some good ideas about the integrated NGL/LNG processes, but there is no

123

discussion about the numerical values of the required power in the process.

124

In [19] an integrated NGL/LNG process configuration was suggested. In this configuration NGL

125

is separated by two columns which one of them, demethanizer column, operates in high pressure

126

and this point decreases the required power for recompression of the residue gas.

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Reflux stream of the NGL recovery process column can be obtained from the NGL fractionator

128

or a liquid stream from the liquefaction unit [20-22]. Cueller at al. [16] designed an integrated

129

ethane recovery and LNG liquefaction process with adjustable C2+ concentration and HHV. The

130

results showed that the specific power and ethane recovery are 0.335 kWh/kg LNG and 87.5 %

131

respectively. Nonetheless they did not disclose the process configuration. An integrated

132

NGL/LNG process was proposed by [23, 24]. A middle heavy natural gas with 85 percent ethane

133

content was used as feed of the process. Reported specific power is 0.37 kWh/kg LNG for 90+ %

134

ethane recovery and 0.29 kWh/kg LNG for 42 % ethane recovery. However the process needs

135

hot utility but it can be removed with increasing the integration level of the configuration. One of

136

the newest ideas about the integrated processes was introduced by [25]. Results show

137

liquefaction efficiency of the process from a rich typical feed gas (methane 75%, and

138

heavier hydrocarbons 23%) is 0.414 kWh/kg LNG and it can recover ethane higher than

139

93.3%.



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In this study three novel integrated process alternatives for cogeneration of LNG and NGL with

141

reasonable energy consumption and high ethane recovery are introduced. New process

142

configurations that is introduced in this paper are LNG based. It means it was tried to invent new

143

configurations based on the conventional LNG processes. And this point will noticeable in

144

process modifications from a LNG process to a combined NGL-LNG process. NGL based

145

process configurations can be defined versus LNG based ones. In fact they are a suitable

146

alternative for modification of an existed NGL recovery process. Introduced configuration in [4,

147

26] can be classified as a NGL base process. As explained various refrigeration systems have

148

some advantages and disadvantages. In this study recognized and conventional liquefaction

149

processes are considered. In fact conventional and efficient refrigeration systems such as DMR,

150

C3-MR and MFC are used for supplying the required refrigeration in the process. Also it was

151

tried to consider all LNG and NGL process design limitations, such as temperature approach in

152

multi stream exchangers, standard and allowable operating condition in the devices

153

(compressors, air coolers, flash drums …) and processes [27].

154 155

2. Process description

156 157

2.1. Selection of the liquefaction refrigeration system

158

There are several types of refrigeration systems that can be used for LNG process. Refrigeration

159

systems can be classified by refrigerant, number of stages and process configurations. Fig.1

160

illustrates commercial LNG refrigeration systems classification based on the configuration and

161

working fluid. Each of them has some advantages and disadvantages. Single stage cooling cycles

162

have simple configurations and accordingly they need fewer number of components, but their


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163

capacity is limited. Cascade and multistage systems have higher energy efficiency. In such

164

systems the required refrigeration the in the low temperature stages decreases and this point

165

reduces the overall required power [28]. Single stage refrigeration system is suitable for pick

166

shaving and floating LNG processes [29]. Pure refrigerant system is reliable and simple in

167

design, but the required power in the multi stage compressors is high [28]. Pure refrigerant

168

evaporates at a constant temperature, leading to a stair way style in the composite curves of the

169

heat exchangers. Mixed refrigerant evaporates over a range of temperatures and creates more

170

efficient cooling composite curve. Small temperature difference throughout the heat exchangers

171

in the low-temperature systems decreases the required power [28].

172

Fig.1. LNG process refrigeration systems classification [30-34]

173

In some processes like PMR, C3-MR and AP-X (C3-MR-N2), both pure and mixed refrigerant

174

cycles are used [31, 35]. Aim of this study is design and analysis of the integrated NGL recovery

175

and base load LNG processes. Also decreasing the required power and increasing the ethane

176

recovery (90 %+) are two of the most important factors which are considered in these processes.

177

So, pure refrigerant and single systems are disregarded. The current market of the LNG

178

technology has been dominated by C3-MR refrigeration system with nearly 80% of installed

179

trains. C3-MR is chosen because it is reliable and efficient [30]. But capacity of the propane

180

precooling/mixed refrigerant is limited to less than 6 MTPA [31]. Dual mixed refrigerant was

181

developed for higher capacities, higher efficiency and more flexibility in the feed composition

182

and ambient conditions. MFC use three mixed refrigerant cascade cycles which is more flexible

183

and efficient than the DMR process [30]. So C3-MR, DMR and MFC processes are considered

184

because they are efficient, flexible and applicable.



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185 186

2.2. NGL recovery section

187

In a NGL recovery plant, demethanizer column is core of the process based on the union model

188

[28]. Relative volatility of the methane/ethane must be sufficiently high in order to reaching to a

189

high ethane recovery. Relative volatility depends on the pressure and temperature of the column.

190

Suitable α methane/ethane is achievable in the pressures lower than 30 bar and temperatures lower

191

than -80 ºC [5, 10]. So decreasing the pressure and temperature of the natural gas is necessary. In

192

the other hand CP of the natural gas which is a function of the pressure plays important role in

193

thermal design of the process. Outlet lean gas from top of the tower follows to the liquefaction

194

and subcooling section. Pressure of the inlet gas to liquefaction section should be increased

195

because with changing the Cp, style of the composite curve can be optimized.

196

Pure propane multi stage systems provide the required refrigeration in the NGL recovery

197

processes. In the integrated NGL-LNG process configurations propane cycle can be eliminated.

198

Feed precooling and cold recycle stream in the demethanizer column is provided by main LNG

199

refrigeration systems. Heat integration is a procedure which is used in order to eliminations of

200

the external hot utility. Conditions of the hot streams of the process, heat capacity (ṁCP) and

201

temperature, is suitable for supplying the required hot utility in the column. Thus column side

202

streams are used for removing the heat from the process in the precooling and liquefaction heat

203

exchangers. In fact required the cold utility can be saved by process integration. Table 1 shows

204

the specifications of the tower in the integrated processes. In this study the feed specifications are

205

constant for all proposed cases.

206


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2.3. Mixed Fluid Cascade (MFC) process configuration

208

Fig.1 illustrates process flow diagram (PFD) of MFC integrated process. As can be seen NGL

209

recovery and liquefaction are done in one integrated PFD. Fig.3 shows overall composite curve

210

of the process.

211

2.3.1.

NGL recovery section

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However PFD of the NGL recovery section in the proposed processes are the same but operating

213

condition is different in each of them. So NGL recovery section is described in this section only.

214

Cleaned and pretreated natural gas feed enters the plant at 37°C and 63 bar. Composition of the

215

feed is shown in Table 2. Pretreated natural gas is cooled in two steps. At first, it follows to E-1A

216

heat exchanger and is cooled to 3°C and further cooling up to -30ºC is done in the second heat

217

exchanger, E-1B. A portion of D-2 gas product, 40%, flows to E-2 and subcooled to about -88°C

218

(Stream 114). Next stream 114 follows to top of the demethanizer tower via a J-T valve as cold

219

reflux. Another portion of outlet vapor from D-2 is expanded thought a turbo expander prior to

220

entering the demethanizer right below top section of the tower. Also, liquid bottom is spilted into

221

two parts. Stream 108 is introduced to the column for fractionation through passing a J-T valve.

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Another portion, stream 107, is subcooled via E-2 to –50°C and enters the tower through a J-T

223

valve. Demethanizer tower operates at about 25 bar and contains conventional trays used in

224

demethanizer columns. This tower has three liquid draw trays to provide the required heat for

225

striping volatile component from the produced NGL. Required heat is supplied by two multi

226

stream heat exchangers, E-1A and E-1B. Side streams, 1, 2 and 3 enter the heat exchangers at 11,

227

-9, -44 °C and return to the tower at 35, 0, 0 °C, respectively. In this configuration there is no

228

need to have a reboiler and ethane recovery is 92%.



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229

Fig.2. Process flow diagram of the MFC configuration

230 231

Table1 presents typical operating condition of the column. Tray numbering is from top to bottom

232

and all stages are ideal stage.

233 234

Table 1. Demethanizer column data sheet

2.3.1.1.Liquefaction section

235

Lean gas stream that leaves the demethanizer tower at about -97°C and 25 bar enters the LNG

236

section. Stream 112 is pressurized via compressor C-100 up to about 63 bar and then is cooled in

237

E-2 to about -85°C. It is clear that gas warms up through the pressurizing, but it should be done

238

because in low pressures the required heat transfer area in heat exchangers increases due to low

239

thermal heat capacity of low pressure gas. Final cooling in LNG production in this process is

240

performed in E-3 heat exchanger. The outlet cold stream from E-2, stream 119, is cooled to

241

about -160 °C in E-3 heat exchanger. Next, outlet stream from E-3 follows to D-1 flash drum via

242

passing a J-T valve. The liquid product of D-1 is LNG at atmospheric pressure.

243

2.3.1.2. Refrigeration system

244

Due to advantages of mixed refrigerant systems such as high thermal efficiency and high

245

flexibility, MR refrigerant was used in this integrated process. Also cascade system has good

246

thermodynamic efficiency and can provide closer and uniform cold/hot composite curves in heat

247

exchangers. Overall composite curves (Fig.3) show that design of the system is optimum and

248

close hot and cold composite curves results high thermodynamic efficiency.

249

The hottest cycle (CYCLE 400)

250

Process flow diagram of this cycle is shown by red lines (Fig.1). First MR cycle provides the

251

required refrigeration for pre-cooling of the feed. Also it is a heat sink for cooler cycles, cycle 10 ACS Paragon Plus Environment

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200 and 300. The cycle 400 refrigerant is a mixture of propane and ethane. Temperature and

253

pressure of stream 400, outlet stream from air cooler AC-400, is about 36°C and 17 bar

254

respectively. This stream enters E-1A and cooled to about 9°C. Next it is divided into two

255

portions. First portion (about 56%), stream 406, returns to E-1A via an expansion valve at -1°C

256

and 670 kPa. Refrigerant stream warms up to about 31 °C in E-1A. Reminded stream, 402,

257

enters E-1B and leaves it at -22°C. Cooled stream, 403, passes a J-T valve and expands to about

258

300 kPa. In this condition refrigerant mixture is at -28 °C. Expanded stream supplies the required

259

cooling in E-1B. After passing through the heat exchanger, stream 406 temperature reaches to

260

2°C and then it is pressurized via C-400A compressor to 670 kPa. This stream and another

261

portion of the main stream follow to a mixer and next enter C-400B for pressurizing to about 17

262

bar. Then pressurized stream is cooled by AC-400 and returns to the initial state. Refrigerant

263

composition is the most important variable which can affect the thermal design of the multi

264

stream heat exchangers. Refrigerants composition are presented in Table 2.

265

Middle cycle (CYCLE 200)

266

This cycle provides a portion of the required refrigeration for liquefaction section and main

267

portion of the required refrigeration for NGL recovery unit. Also middle cycle is a heat sink for

268

the coldest cycle, cycle 300. Refrigerant of this cycle is composed of methane, ethane, propane

269

and ethylene. Refrigerant stream warms up to -30° C in E-2 heat exchanger. Next this low

270

pressure stream enters C-200A and pressurized up to 15 bar. For decreasing the required power,

271

pressurizing is performed in two stages. Stream 206 follows to AC-200A and then enters C-200B

272

and its pressure is increased to 28 bar. Stream 209 returns to E-2 after passing through E-1A and

273

E-1B at about -27°C. In this heat exchanger it is cooled to -81°C and returns via an expansion



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274

valve at -93 °C and 310 kPa. Refrigerant stream reaches to initial state after passing through E-2.

275

This cycle is illustrated by green lines in Fig.2.

276 277

Liquefaction cycle (CYCLE 300)

278

The main duty of this cycle is supplying the required refrigeration for liquefaction and

279

subcooling. Refrigerant in this cycle is composed of methane, ethylene and nitrogen. Refrigerant

280

stream, 303, enters E-3 at -85°C and 29 bar and is cooled to about -159°C. Stream 305 returns to

281

E-3 at -167°C and 350 kPa. After passing through E-3, refrigerant stream warms up to about -

282

87°C. Refrigerant stream pressure is increased in two stages. Pressurized and warmed stream,

283

300, enters E1-A at 35°C. The reminded cooling is done in E1-A, E1-B and E-2 heat exchangers.

284

Finally the output stream from E-2, 303, follows to E-3 for sub cooling the lean gas, stream 118.

285

Blue lines present this cycle in Fig.2.

286

Table 2. MFC configuration main streams data

287

Fig.3. Composite curves of the MFC configuration

288 289

2.4. Dual mixed refrigerant (DMR)

290

This process configuration is based on the dual mixed refrigeration system. Fig.4 illustrates

291

process flow diagram of DMR process. Red lines show precooling refrigeration cycle and blue

292

lines illustrate liquefaction one. Table 3 presents the main streams operating condition.

293

2.4.1. NGL recovery section

294

Operating condition and components of the NGL recovery section of DMR process are same as

295

the MFC.


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2.4.2. Liquefaction unit

297

Column gas product is pressurized via a compressor to 63 bar. Next, stream 113 follows to E-2A

298

and E-2B heat exchangers respectively. In fact this stream is cooled in two steps, -128 ºC and -

299

161 ºC respectively. Like the previous process, outlet stream from the cooling section, stream

300

119, throttles via a J-T valve and then enters the D-1 flash drum. LNG product leaves the plant as

301

bottom product of the D-1.

302

2.4.2.1.Refrigerant system

303

DMR is an efficient and reliable natural gas liquefaction process. In mixed refrigerant cycles

304

closer composite curves can be achieved but increasing number of the cycles increases capital

305

costs and decreases the required power. Dual mixed refrigerant systems use mixed refrigerants in

306

two different cycles. They use mixed refrigerants in optimized number of the cycles. In DMR,

307

energy loss is more than the MFC arrangement, but high total mechanical efficiency

308

compensates low thermodynamic efficiency. The hottest MR cycle provides the required cold

309

utility for E-1A and E1-B heat exchangers. Also, a portion of the required refrigeration for

310

liquefaction is provided by this cycle. Second cycle is used for liquefaction and subcooling

311

which provides the required cold utility in E-1A and E1-B heat exchangers (Fig.4.). Overall

312

composite curves of the DMR process is shown in Fig.5.

313 314

Fig.4. Process flow diagram of the DMR configuration

315

Fig.5. Overall composite curves of the DMR configuration

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Cycle 200 (precooling cycle)

318

Stream 200 is composed of ethane, propane or ethylene. Table 3 presents composition of the

319

MR. As illustrated in Fig.4, stream 200 is precooled via E-1A heat exchanger to -5ºC. Precooled

320

stream is divided to two portions. Stream 202 (About 63 percent) returns to E-1A via a J-T valve.

321

Another portion, stream 205, follows to E-1B heat exchanger and is cooled to -33ºC, and then its

322

pressure is decreased to 300kPa via a J-T valve. Expanded stream supplies the required cold

323

utility for E-1B. Outlet refrigerant from E-1B, stream 208, is pressurized to 760 kPa via C-200A

324

compressor. Next hot refrigerants, streams 204 and 208, follow to a mixer and pressurized to

325

1920 kPa via C-200B compressor. Then outlet stream, 211, is cooled by AC-200 air cooler.

326

Cycle 300 (Liquefaction cycle)

327

Table 3illustrate composition of the mixed refrigerant. Outlet stream from AC-300A is at 4860

328

kPa and 35ºC. This steam is cooled in two steps by E-1A and E-1B heat exchangers respectively.

329

Refrigerant temperature reaches to -33.15 ºC after passing through precooling section. Next

330

stream 302 follows to the D-300 drum. Bottom and top product of D-300 enter E-2 in two

331

different sides and cooled to about -128ºC. Gas outlet, stream 307, enters the E-3 heat exchanger

332

and is cooled to -160ºC, then it is expanded to 300 kPa via V-300 expansion valve and returns to

333

E-3. Expanded stream supplies the required refrigeration in E-3. Outlet refrigerant from this

334

exchanger, stream 310, is cold yet. This stream combines with bottom outlet of D-300 that

335

cooled and expanded via E-2 and V-301 respectively. Then stream 311enters E-2 and supplies a

336

part of the required cold utility. After E-2, refrigerant temperature is increased to -43ºC. Next

337

stream 312 is pressurized in two steps by compressors C-300A and C-300B to 4860 kPa.

338

2.5. Propane precooled refrigerant system 14 ACS Paragon Plus Environment

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339

C3-MR refrigeration system is a conventional and reliable liquefaction process in LNG

340

industries [1, 36]. Pure propane refrigeration system pre chills natural gas for feeding to NGL

341

recovery section. Fig.6 presents C3-MR process configuration and Fig.7 shows overall

342

composite curves of the system.

343

Fig.6. Process flow diagram of the C3-MR process

344 345

2.5.1. NGL recovery section

346

NGL recovery section is same as the previous ones. Pretreated gas is fed to the plant in the same

347

condition and quantity.

348

2.5.2.

Precooling cycle

349

Inlet stream is pre cooled by three stage pure propane refrigerant cycle. Pure propane at 35 ºC

350

and 1430 kPa is expanded and cooled via V-1 valve to 500 kPa. Next, stream 201 follows to D-1

351

drum. Liquid product of D-1 is divided to two portions. Stream 204, 48 percent of the bottom

352

outlet of D-1, supplies the required cold utility of E-1A heat exchanger. Another portion, stream

353

205, is expanded to 200 kPa and then is divided to two portions. Stream 209, about 60 percent of

354

liquid product of D-3, enters E-1B heat exchanger. Another portion supplies the required cold

355

utility of of pre cooling section at 100 kPa. Three single stage compressors supply the required

356

compression in the cycle.

357

2.5.3. MR Cycle

358

Mixed refrigerant is used for liquefaction and sub cooling the natural gas. Composition of the

359

refrigerants is illustrated in Table 4. Stream 300 enters precooling section at 35 ºC and 4900 kPa. 15 ACS Paragon Plus Environment


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

360

Refrigerant is cooled in three steps to 4,-17,-34 ºC in heat exchangers E-1A, E-1B and E-1C

361

respectively. Outlet mixed refrigerant from E-1C follows to D-4 drum. Outlet liquid from D-4,

362

stream 305, is cooled to -128 ºC. Stream 304, D-4 gas product, is cooled to -128 ºC via E-2A

363

heat exchanger. Next it is throttled to 300 kPa via V-4. Next, cooled and condensed stream,

364

stream 307, supplies the required refrigeration for sub cooling the pressurized LNG. Stream 307

365

passes through the E-2B heat exchanger and is cooled to -161 ºC. Refrigerant returns to E-2B via

366

V-5 economizer valve. Refrigerant leaves E-2B at -139 ºC and 300 kPa. This stream is able to

367

supply a portion of the required duty of E-2A. Next, stream 313 and outlet stream from V-4

368

follows to a mixer and supply the required cold utility of E-2A. Mixed refrigerant is pressurized

369

in three steps by C-300A and C-300B compressors.

370 371

Fig.7. Composite curves of the C3-MR process. 2.5.4. Liquefaction section

372

Compressed and lean gas, stream 117, enters E-2A and E-2B heat exchangers and is cooled to

373

-128 and -162 ºC respectively. Sub cooled and pressurized LNG is expanded to atmospheric

374

pressure via V-104 valve. LNG becomes two phase after passing through a throttling valve.

375

Finally LNG leaves the D-101 drum as a liquid product of the process. Table 4 shows main

376

streams condition and composition.

377 378

Table 4. C3-MR configuration main streams data 3. Process analysis

379

Power consumption per liquid product and NGL recovery are main factors which should be

380

considered through the analysis of such processes. NGL recovery level depends on the supplied


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381

refrigeration in the process and it can be increased with increasing the refrigeration load.

382

However increasing the refrigeration load increases the required power in the process. Feed

383

composition is another important factor which can affect the power consumption.

384 385

3.1. NG to LNG thermodynamic trajectory

386

In NGL recovery process natural gas is cooled before following to the fractionation column. In

387

this step a portion of the cooled gas which contains heavy hydrocarbons is separated in flash

388

drum. However gas and liquid products of the cold separator in the process are sent to the

389

demethanizer column for more separation. Output lean gas from NGL recovery process is

390

subcooled and converted to LNG. Fig. 8 shows thermodynamic trajectory of the NG to LNG in

391

P-T diagram.

392

Fig.8.NG to LNG thermodynamic trajectory

393 394

3.2.Composite Curves Analysis

395

Overall composite curves present temperature driving force in the heat transfer devices of the

396

process. Required cold and hot utilities in the process are detectable in the composite curves.

397

Exergy is defined as amount of energy that can be converted to work in reversible process to(in)

398

standard condition [37]. The quality of energy is measured by exergy. Driving force of the heat

399

transfer is temperature difference. In design step, balance between deriving force and minimum

400

temperature approach must be considered. In this study minimum temperature approach was

401

supposed to be 2 ºC. Closer composite curves mean closer Tc and Th and lower LMTD [26]. 17 ACS Paragon Plus Environment


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

402

Table 5 presents minimum temperature approach and LMTD of each heat exchanger in the

403

processes.

404 405

Table 5. LMTD and minimum temperature approach of the heat exchangers

406 407

Fig.3 illustrates composite curves of the MFC process. Mixed refrigerant and three cycles result

408

close composite curves. As a result, MFC configuration has the highest efficiency between the

409

introduced processes. MFC has good efficiency, but three cycles need more equipment and more

410

capital cost. Dual mixed refrigerant was developed for elimination a cycle and using advantages

411

of mixed refrigerants. Each interval in Fig. 4 shows one heat exchanger. Using MR refrigerant

412

and two cycles results such composite curves. Fig.7. shows composite curves of C3MR process.

413

Three stage propane pre cooled cycle can be detected in the cold composite curve, based on its

414

stair way style. As expected, C3-MR process has the lowest efficiency between the introduced

415

processes. Reliability and simple operation of this configuration are two of the most important

416

advantages. However its efficiency is lower than the DMR and MFC configurations.

417 418

3.3. Effect of the integration on the utility consumption and composite curves

419

Changing the side streams mass flow rate affect the total heat capacity (∑ṁCP) of the cold

420

composite curves. So when flow rate of the side streams decreases temperature of the outlet

421

streams increases and accordingly the cold composite goes up and a temperature cross may be

422

occurred. So if mass flow rate of the side streams is decreased cold composite nears hot

423

composite. Effect of the side streams operating condition on the style of the composite curves in 18 ACS Paragon Plus Environment

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424

MFC configuration heat exchangers (E-1A and E-1B) is illustrated in Fig. 9. Left side composite

425

curve shows effect of the side 1R temperature on the position of the cold composite. Right side

426

diagram presents cold composite curve variations when flow rate of the sides 2R and 3R are

427

reduced up to 20% and their temperature is decreased up to 10 ˚C (heat exchanger outlet)

428

towards the initial value. In the integrated configurations changing the operating condition of the

429

effective streams cause different parts of the processes to be affected. Reducing the side streams

430

outlet temperature decreases a portion of the cold load which is consumed in the heat

431

exchangers. Accordingly they return to the tower in a lower temperature. Such changing reduces

432

the striping section temperature and affect the demethanizer performance.

433

extracted refrigeration load from the side streams decreases the cycles load should be increased

434

to compensate it.

Also when the

435

Fig.9. Effect of the side streams operating condition on the style of the composite curves

436

NGL section supplies a portion of the required cold utility in the precooling section. However

437

demethanizer tower needs cold utility in low temperature for separation of the C2+. As described

438

a part of the required cold utility is supplied by about 37% of the gas outlet from the cold

439

separator (stream 105 in MFC configuration). This stream returns to the top section of the

440

column as recycle stream. Flow rate and temperature of the recycle stream are main parameters

441

for the ethane recovery level. In the liquefaction multi stream heat exchanger, recycle stream is a

442

hot stream that is cooled by refrigeration system. Increasing the recycle stream flow rate

443

increases the total specific heat and consequently reduces slope of the hot composite curve as

444

illustrated in Fig. 10. Effect of the recycle flow rate is investigated in MFC configuration (Fig.

445

10). In such condition refrigerant flow rate must be increased to avoid the temperature cross in

446

the heat exchanger. 19 ACS Paragon Plus Environment


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

447 448

Fig. 10. Effect of the recycle flow rate on the hot composite curves style 3.4. Refrigeration systems analysis

449

T-S and P-H diagrams provide valuable information about the thermal design of the refrigeration

450

cycles. Fig.11 illustrates T-S and P-H diagrams of the C3-MR process refrigeration cycles. Three

451

compression steps can be detected in the P-H diagram of the pure propane refrigerant cycle.

452

Three pressure level of the methane pre cooling cycle decreases the required power [38].

453

Fig.11. P-H and T-S diagrams of the C3-MR configuration

454

Numbers in the figures refer to the name of the streams in the specified configurations. P-H and

455

T-S diagram of DMR configuration refrigeration cycles is shown in Fig. 12.

456

Fig.12.T-S and P-H diagrams of the DMR configuration

457

The hottest cycle in MFC and DMR configurations is two pressure level refrigeration system.

458

Like the precooling section in C3-MR configuration, the required power decreases by using two

459

pressure levels towards the one stage system. T-S and P-H diagrams of MFC configuration are

460

illustrated in Fig. 13. The middle and liquefaction cycles are the simplest cycles that applied in

461

the integrated processes.

462

Fig. 13. T-S and P-H diagrams of the MFC configuration

463 464

T-S diagram is a common and helpful tool for analysis of the refrigeration systems. Area of the

465

T-S diagram shows the required work for cooling. So based on this concept multi stage

466

refrigeration systems need less work. Three stage propane refrigeration cycle and two or three

467

cycle system consume lower energy. Another important factor which can be used for 20 ACS Paragon Plus Environment

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468

refrigeration cycle analysis is coefficient of performance (COP). Table 6 shows COP of all

469

cycles and total COP of liquefaction systems. COP is defined as ratio of the transferred heat from

470

the cold source to the required power. High COP means the required power for providing certain

471

refrigeration is low. MFC configuration cycles have the best COP between the other ones

472

because three independent mixed refrigerant liquefaction systems able to transfer heat closer to

473

the reversible condition. As expected C3-MR has the lowest COP. Three stage propane

474

precooling cycle generate entropy more than the other cycles as described in the composite

475

curves analysis.

476

Table 6. COP of the refrigeration cycles

477 478

3.5. Feed composition effect

479

Feed composition is one of the most important factors which can affect the NGL recovery level

480

and economic evaluations. Heavier feed needs more refrigeration load, larger heat exchangers

481

and higher capital and operating cost for a given NGL recovery level. In the other hand leaner

482

gases need more severe condition for high ethane recovery level [4]. Mehrpooya et al [1] studied

483

feed composition effect on the ethane recovery and power consumption in a novel ethane

484

recovery process. Table 7 shows effect of the feed composition on specific power of the process

485

configurations.

486

Table 7. Specific power versus methane content of the feed

487

All constrains such as temperature approach and % 90+ ethane recovery are met. As can be seen

488

total specific power increases with methane content. Heavier feeds need more cooling in the



1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

489

precooling refrigeration cycles and produce fewer LNG than the feeds with high methane

490

content. Low methane content feeds send leaner gas to the liquefaction section and accordingly

491

LNG section needs lower cooling duty, while precooling section consumes higher cooling duty.

492

It means lower portion of the feed is cooled to very low temperatures and total power

493

consumption decreases. But lighter feeds produce more LNG and consequently denominator of

494

the specific power index increases. But total power consumption increases with a higher rate.

495

Finally effects of these two factors decrease the specific power and increase the total specific

496

power.

497

3.6. Effect of the cold recycle specifications on the ethane recovery and specific power

498

A portion of the precooled natural gas is liquefied after precooling as illustrated in Figures 1,

499

3&5. Separated and precooled gas after D-2 (Figures 1&3) and D-100 (Fig.6.) is divided to two

500

parts. In a specified condition, %40 of the discharged gas from the pre separation drum (D-2 and

501

D-100) is cooled and follows to top tray of the NGL recovery column, T-100. Cold recycle flow

502

rate is a main factor for NGL recovery level. Recycle stream temperature is another important

503

design factor. It is clear that, higher amount and colder recycle result higher NGL recovery.

504

Table 8 shows effect of the cold recycle stream temperature on the ethane extraction and specific

505

power. For increasing the NGL recovery recycle stream temperature and flow rate should be

506

decreased and increased respectively. As shown in Table 9 increasing the recycle ratio increases

507

the hydrocarbon recovery. But a little decrease in the ethane recovery is observed by increasing

508

the reflux ratio after the maximum recovery. Increasing the reflux ratio causes temperature of the

509

rectifying section in the tower decreases and a little decrease in α C1/C2 in constant pressure is the

510

reason of such behavior. So, there is an optimum reflux ratio for achieving 90%+ ethane

511

recovery [5]. 22 ACS Paragon Plus Environment

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512

Higher flow rate and cooler recycle stream need more cold utility and more power consumption

513

in the refrigeration cycles. As illustrated in Tables 8&9, specific power (kWh/ kg LNG)

514

increases by NGL recovery level.

515 516

Table 8. Effect of the recycle stream temperature on the ethane recovery and specific power at

517

the constant flow rate (40% of the gas product from D-2 and D-100)

518

Table 9. Effect of the recycle ratio at constant temperature (-88ºC) on the ethane recovery and

519

specific power

520 521

4. Numerical implementation

522

Graphical user interface (GUI) simulators with wide range of data bank are perfect choice for

523

modeling and solve the chemical processes [7]. Selection of Equations of state (EOS) is another

524

important factor which can affect the results significantly. There are many EOSs in data bank of

525

the chemical process simulators like HYSYS and Aspen Plus. Recommended EOS for natural

526

gas cryogenic processes is PR (Peng Robinson) and modified PR or PRSV EOS [8, 39]. Previous

527

studies in the related areas confirm using PR or PRSV for natural gas liquefaction [4, 8]. Also

528

using HYSYS and Aspen simulator is prevalent for simulation of LNG and NGL processes [40,

529

41].

530

5. Energy consumption analysis

531

In this study integrated NGL/LNG processes are introduced. Energy efficiency of these

532

processes is high and they can recover C2+ higher than 90%. Specific power and ethane recovery 23 ACS Paragon Plus Environment


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

533

are about 0.36-0.4 kWhr/Kg LNG and 92% respectively. Also in design step it was tried to

534

decrease the complexity of the configurations as simple as possible. Main streams mass flow rate

535

and main equipment power consumption are presented in Table 10. There are a few data about

536

this type of processes. Cuellar at al. [16, 23, 42] introduced an integrated process. The results

537

showed that maximum specific power and ethane recovery were 0.335 kWhr/Kg LNG and 87.5

538

% respectively. Nonetheless they did not disclose the process configurations and feed

539

composition. Introduced processes in this study shows the specific power is between 0.35-0.38

540

kW-hr/kg LNG and ethane recovery is upper than 92 %. The newest integrated process in this

541

area is introduced by Vatani et al [25]. They worked on a NGL based process configurations.

542

Table 11 presents a comparison between introduced and previous processes. Simpler process

543

configuration needs simpler control system and less capital cost. Also increasing the efficiency

544

of the processes reduces the energy consumption and operating costs. Ethane recovery is a

545

function of the column temperature and reflux ratio.

546

Table 10. Main equipment power consumption and specific power

547

Some devices such as compressor and multi stream heat exchangers can be removed by

548

integration. Also it is expected that heat integration decreases the required cold and hot utility.

549

Required hot utility can be covered by integration between the tower and precooling cycle.

550

Reported specific power for such processes is 0.5 to 0.7 kWhr/kgLNG. Multi stage compressors

551

can be replaced with singe stage compressors in each refrigeration cycle. Table 11 compares

552

processes with this assumption.

553

Table 11. Comparison between the introduced configurations and other processes.

554 24 ACS Paragon Plus Environment

Page 24 of 58

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555


6. Exergy analysis

556

Exergy is the maximum attainable useful work in a process that brings the system into the

557

reference state [43, 44]. By exergy analysis method, second law efficiency and amount of the lost

558

work can be calculated for a system. [37]. Exergy can be used for design and performance

559

evaluation of the chemical processes [45, 46].

560

Exergy balance and irreversibility: considering the control volume at the steady state condition

561

the exergy balance can be expressed as

562

𝐸𝑥𝑖 + 𝐸𝑥𝑄𝑖 = 𝐸𝑥𝑜 + 𝐸𝑥𝑄𝑜 + 𝑊𝑠ℎ + 𝐼

563

(2)

Where

564

Exi and Exo are exergy flow of inlet and outlet material streams respectively.

565

ExQi and ExQo are exergy flow of inlet and outlet energy streams respectively.

566

Wsh is shaft work and I represent irreversibility.

567

6.1. Exergy efficiency

568

Exergy efficiency can be calculated by two different methods: input-output and fuel-product.

569

Input – output method is defined by Eq.3 . But there are various relations for calculation of

570

exergetic efficiency by fuel-produced method [6, 47].

571

𝜂=

572

Table 12 demonstrates the exergetic efficiency definitions for main components of the processes.

573

Second law analysis of the tower is done by exergy balance around it.

∑ 𝐸𝑥𝑒𝑟𝑔𝑦 𝑜𝑢𝑡 ∑ 𝐸𝑥𝑒𝑟𝑔𝑦 𝑖𝑛

=1−

∑ 𝐸𝑥𝑒𝑟𝑔𝑦 𝑑𝑒𝑠𝑡𝑟𝑜𝑦𝑒𝑑

(3)

∑ 𝐸𝑥𝑒𝑟𝑔𝑦 𝑖𝑛



1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

574

Table 12. Definitions for exergetic efficiency of the process components

575

Fuel – product method is used for calculation of the exergy efficiency of the processes. Results

576

indicate the exergy efficiency of the under consideration processes is acceptable. Table 13 shows

577

the overall exergy analysis results for the processes and their cycles.

578

Table 13. Exergy efficiency of the processes and their cycles

579 580

Table 14 presents exergy distraction and exergy efficiency of the main components such as

581

compressors, multi stream heat exchangers, air cooler and expansion valves. The greatest

582

distraction is related to the compressors and multi stream heat exchangers. But J-T valves have

583

the least exergy efficiency because expansion process is inherently irreversible. Exergy

584

efficiency of the heat exchangers increases by operating temperature. The reason is that entropy

585

generation at lower temperatures for a constant heat transfer rate is higher (Eq.4 and Eq.5).

586

𝑆𝑔𝑒𝑛 = ∑

587

𝐸𝐷𝑒𝑠𝑡𝑟𝑜𝑦𝑒𝑑 = 𝑇0 𝑆𝑔𝑒𝑛

588

𝑄𝑖

(4)

𝑇𝑖

(5)

A part of the exergy distraction in a component can be removed by increasing the components

589

performance. Also exergy distraction of the process can be decreased by some of the

590

modification procedures. For example replacement of the throttling valves with expanders can

591

recover some exergy losses as a shaft work.

592

Table 14. Results of exergetic efficiency of the processes components.

593 594

7. Conclusion 26 ACS Paragon Plus Environment

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595

Three reliable and efficient novel process configurations for co-production of LNG and NGL are

596

introduced and analyzed. Ethane recovery in all of them is above 90%. Conventional

597

refrigeration systems like MFC, DMR and C3-MR are used for supplying the required

598

refrigeration. Specific power of the processes for typical feed gases is lower than the 0.4

599

kWhr/kgLNG. In order to discuss about the advantages of the introduced processes their

600

performance should be compared with the similar cases. Specific power, ethane recovery,

601

number of stages and process components are the parameters which can be considered. But

602

specific power and ethane recovery depend on the feed composition. So new processes outputs

603

were checked with the feed condition of the other similar cases and the results show that their

604

performance is considerably better. Simple design and using reliable configuration in refrigerant

605

systems and NGL recovery section are other advantages of these processes.

606 607 608

References:

609 610 611 612 613 614 615 616 617 618 619 620 621 622 623

[1] Barclay, M. and N. Denton, Selecting offshore LNG processes, 2005, LNG Journal. [2] Arthur J. Kidnay, W.R.P., Fundamentals of Natural Gas Processing2006: Taylor and Francis Group, LLC. [3] Doug Elliot, W.R.Q., Shawn Huang,Jong Juh (Roger) Chen,R. J. Lee,Jame Yao,Ying (Irene) Zhang, BENEFITS OF INTEGRATING NGL EXTRACTION AND LNG LIQUEFACTION TECHNOLOGY, in AICHE2005. [4] Mehrpooya, M., A. Vatani, and S.M. Ali Mousavian, Introducing a novel integrated NGL recovery process configuration (with a self-refrigeration system (open–closed cycle)) with minimum energy requirement. Chemical Engineering and Processing: Process Intensification, 2010. 49(4): p. 376-388. [5] ENGINEERING DATA BOOK (GPSA). Vol. 1,2. 2004. [6] Tirandazi, B., et al., Exergy analysis of C2+ recovery plants refrigeration cycles. Chemical Engineering Research and Design, 2011. 89(6): p. 676-689. [7] Mehrpooya, M., A. Jarrahian, and M.R. Pishvaie, Simulation and exergy-method analysis of an industrial refrigeration cycle used in NGL recovery units. International Journal of Energy Research, 2006. 30(15): p. 1336-1351.



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624 625 626 627 628 629 630 631 632 633 634 635 636 637 638 639 640 641 642 643 644 645 646 647 648 649 650 651 652 653 654 655 656 657 658 659 660 661 662 663 664 665 666 667 668 669

[8] Mehrpooya, M., F. Gharagheizi, and A. Vatani, Thermoeconomic analysis of a large industrial propane refrigeration cycle used in NGL recovery plant. International Journal of Energy Research, 2009. 33(11): p. 960-977. [9] Mehrpooya, M., A. Vatani, and S.M.A. Mousavian, Optimum design of integrated liquid recovery plants by variable population size genetic algorithm. The Canadian Journal of Chemical Engineering, 2010. 88(6): p. 1054-1064. [10] Lily Bai, R.C., Jame Yao, Douglas Elliot, RETROFIT FOR NGL RECOVERY PERFORMANCE USING A NOVEL STRIPPING GAS REFRIGERATION SCHEME, in IPSI LLC2006: Houston, Texas U.S.A. [11] Mehrpooya, M., F. Gharagheizi, and A. Vatani, An Optimization of Capital and Operating Alternatives in a NGL Recovery Unit. Chemical Engineering & Technology, 2006. 29(12): p. 14691480. [12] Mehrpooya, M.V., A., System and method for recovering natural gas liquids with outo refrigeration system, 2013. [13] Finn A.J, J.G.L.T.T.R., developments in natural gas liquefaction. Hydrocarbon Processing, 1999: p. 4759. [14] Shukri, T., LNG technology selection, 2004: HYDROCARBON ENGINEERING. [15] Vatani, A., M. Mehrpooya, and A. Palizdar, Advanced exergetic analysis of five natural gas liquefaction processes. Energy Conversion and Management, 2014. 78: p. 720-737. [16] Cuellar, K.T., et al., CO-PRODUCING LNG FROM CRYOGENIC NGL RECOVERY PLANTS, in 81st Annual Convention of the Gas Processors Association2002. [17] Brostow, A.A. and M.J. Roberts, Integrated NGL Recovery In the Production Of Liquefied Natural Gas, 2013, Google Patents. [18] I.Ransharger, W., Intermediate pressure LNG refluxed NGL revovery process, 2008, Google Patent [19] Wesley Qualls, W.L.R., Shawa S.Huang, Jame Yao, Doug Elliot, Jong Juh Chen, Rong-Jwyn Lee, LNG facility with integrated NGL extraction technology for enhanced NGL recovery and production flexibility, 2007, ConocoPhilips Company, Google Patent [20] Mark Julian Robert, A.A.B., Integrated NGL Recovery And Liquefied Natural Gas, 2010, Google Patent [21] Mark Julian Robert, H.C.R., Integrated high pressure NGL recovery in the Production of Liquefied Natural Gas 2003, Google Patent [22] Adam Adrian Brostow, M.J.R., Integrated NGL recovery in the production of liquafaction natural gas, 2008, Google Patent. [23] John D. Wilkinson, H.M.H., Kyle T.Cuellar, Natural Gas Liquefaction, 2007, Ortloff Engineering Ltd. [24] John D. Wilkinson, J.T.L., Hank M. Hudson, Kyle T.Cuellar, Natural Gas Liquefaction, 2005, ElkCorp. [25] Vatani, A., M. Mehrpooya, and B. Tirandazi, A novel process configuration for co-production of NGL and LNG with low energy requirement. Chemical Engineering and Processing: Process Intensification, 2012. [26] Mehrpooya, M., A. Vatani, and S.M.A. Moosavian, Introducing a new parameter for evaluating the degree of integration in cryogenic liquid recovery processes. Chemical Engineering and Processing: Process Intensification, 2011. 50(9): p. 916-930. [27] Mohammad Hosseini, Mehdi Mehrpoya, Ali vatani, Integrated processes for co-generation LNG and NGL, 2013: Iran. [28] Smith, R., Chemical Process Design and Integration 2 nd ed2005: John Wiley & Sons, Ltd. [29] Khan, M.S., et al., Energy saving opportunities in integrated NGL/LNG schemes exploiting: Thermalcoupling common-utilities and process knowledge. Chemical Engineering and Processing: Process Intensification, 2014. 82: p. 54-64.


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670 671 672 673 674 675 676 677 678 679 680 681 682 683 684

[30] Luis Castillo, R.N., Camilo González, Carlos A. Dorao, Alfredo Viloria, TECHNOLOGY SELECTION FOR LIQUEFIED NATURAL GAS (LNG) ON BASELOAD PLANTS, in International Gas Convention2010: Venezuela. [31] Paul Bosma, R.K.N., Liquefaction Technology; Developments through History, in 1st Annual Gas Processing Symposium2009. [32] Justin Bukowski, Y.N.L., . Stephen Boccella, Leo Kowalski, INNOVATIONS IN NATURAL GAS LIQUEFACTION TECHNOLOGY FOR FUTURE LNG PLANTS AND FLOATING LNG FACILITIES, in International Gas Union Research Conference2011. [33] He, T. and Y. Ju, Design and Optimization of a Novel Mixed Refrigerant Cycle Integrated with NGL Recovery Process for Small-Scale LNG Plant. Industrial & Engineering Chemistry Research, 2014. 53(13): p. 5545-5553. [34] Rivera, V., A. Aduku, and O. Harris, Evaluation of LNG Technologies, 2008. [35] Mark Pillarella, Y.-N.L., Joseph Petrowsk, THE C3MR LIQUEFACTION CYCLE: VERSATILITY FOR A FAST GROWING, EVER CHANGING THE C3MR LIQUEFACTION CYCLE: VERSATILITY FOR A FAST GROWING, EVER CHANGING LNG INDUSTRY LNG INDUSTRY, Air Products and Chemicals, Inc.

685 686 687 688 689 690 691 692 693 694 695 696 697 698 699 700 701 702 703 704 705 706 707 708 709 710 711 712 713 714

[36] Alabdulkarem, A., et al., Optimization of propane pre-cooled mixed refrigerant LNG plant. Applied Thermal Engineering, 2011. 31(6-7): p. 1091-1098. [37] Kotas, T.J., The Exergy Method of Thermal Plant Analysis1995: KRIEGER PUBUSHINO COMPANY MALABAR, FLORIDA. [38] Yunus A. Çengel, M.A.B., Thermodynamics. 7 ed2006: McGraw-Hil. [39] Hyprotech HYSYS v3.2, User Guide, Aspen Technology Inc., 2003,. [40] Mortazavi, A., et al., Performance enhancement of propane pre-cooled mixed refrigerant LNG plant. Applied Energy, 2012. 93: p. 125-131. [41] Aspelund, A., et al., An optimization-simulation model for a simple LNG process. Computers & Chemical Engineering, 2010. 34(10): p. 1606-1617. [42] Martinez, T.L., et al., Liquefied natural gas and hydrocarbon gas processing, 2013, Google Patents. [43] Chang, H.-M., H.S. Lim, and K.H. Choe, Effect of multi-stream heat exchanger on performance of natural gas liquefaction with mixed refrigerant. Cryogenics, 2012. 52(12): p. 642-647. [44] Ali Vatani, M.Mehrpooya., Ali Palizdar, Energy and exergy analyses of five conventional liquefied natural gas processes. International Journal of Energy Research, 2014. doi: 10.1002/er.3193. [45] Gharagheizi, F. and M. Mehrpooya, Prediction of standard chemical exergy by a three descriptors QSPR model. Energy Conversion and Management, 2007. 48(9): p. 2453-2460. [46] Meratizaman, M., et al., Energy and exergy analyses of urban waste incineration cycle coupled with a cycle of changing LNG to pipeline gas. Journal of Natural Gas Science and Engineering, 2010. 2(5): p. 217-221. [47] Marmolejo-Correa, D. and T. Gundersen, A comparison of exergy efficiency definitions with focus on low temperature processes. Energy, 2012. 44(1): p. 477-489. [48]J.D. Seader , E.J.H., Seperation Process Principles ed. s. edition1998: John Wiley & Sons, Inc. [49]Brostow, A.A. and M.J. Roberts, Integrated ngl recovery and liquefied natural gas production, 2006, Google Patents. [50]Curt, G. and J. Mak, Configurations and methods of integrated ngl recovery and lng liquefaction, 2007, Google Patents. [51] Mak, J. and C. Graham, Configurations and methods of integrated ngl recovery and lng liquefaction, 2013, Google Patents. [52] Smith, E.M., Advances in thermal design of heat exchangers2005: John Wiley & Sons, Ltd.

715 716



1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Table caption:

Table 1. Demethanizer tower data sheet. Table 2. MFC configuration main streams data Table 3. DMR configuration main streams data Table 4. C3-MR configuration main streams data Table 5. LMTD and minimum temperature of the heat exchangers Table 6. COP of the refrigeration cycles Table 7. Specific power versus methane content of the feed Table 8. Effect of the recycle temperature at constant flow (40% of gas outlet from D-2 and D-100) Table 9. Effect of the recycle ratio on the gas outlet from D-2 and D-100 at constant temperature (-88ºC) Table 10. Main equipment power consumption and specific power Table 11. Comparison between the introduced configurations and other processes. Table 12. Definitions for exergetic efficiency of the process components Table 13. Exergy efficiency of the processes Table 14. Results of the exergetic efficiency of the processes components.


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Page 31 of 58

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Figure caption:

Fig.1. LNG refrigeration systems classification Fig.2. Process flow diagram of the MFC configuration Fig.3. Composite curve of the MFC configuration Fig.4. Process flow diagram of the DMR configuration Fig.5. Composite curve of the DMR configuration Fig.6. Process flow diagram of the C3-MR configuration Fig.7. Composite curves of the C3-MR configuration Fig.8. NG to LNG thermodynamic trajectory

Fig.9. Effect of side streams on the MFC composite curve Fig. 10. Effect of recycle flow rate on the hot composite curve Fig.11. P-H and T-S diagrams of the DMR configuration Fig.12.T-S and P-H diagrams of the C3-MR configuration Fig. 13. T-S and P-H diagrams of the MFC configuration



1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 32 of 58

Table 1. Typical demethanizer tower data Inlet streams Name Utility Cold reflux Expanded Gas Condensed liquids Condensed cold liquid Side Stream Side Stream Side Stream Lean Gas NGL Side Stream Side Stream Side Stream

MFC

DMR

115 109 117

115 109 117

110

110

115

14628

2500

-64

Side 1R Side 2R Side 3R



87000 83865 75020

2500 2500 2500

35 0 0

18 30 27 19

-97 28.5 14 -11 -48

1 30 29 26 18

112 NGL Side 1 Side 2 Side 3

112 NGL Side 1 Side 2 Side 3

Rate (kg/h) 120359 180538 34131

Specification Pressure Temperaturen (kPa) (˚C) 2500 -97 2600 -65 2550 -48

C3MR 113 108 111

Outlet streams 116 253281 NGL 96381 Side 1 87000 Side 2 83865 Side 3 75020


2500 2500 2500 2500 2500

Tray Number 1 9 14

Page 33 of 58

1 2 3 4 5 6 7 8 9 10 11 12 13 Composition 14 15 16 17 18 19 20 21 operation 22 condition 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Table 2. MFC configuration main streams data Feed gas

Hot MR

Middle MR

Cold MR

NGL

NG to liquefaction

LNG

methane

0.83

0.00

0.14

0.40

0.01

0.99

0.99

ethane

0.07

0.20

0.35

0.00

0.39

0.01

0.01

propane

0.05

0.64

0.21

0.00

0.32

0.00

0.00

n-butane

0.00

0.16

0.00

0.00

0.00

0.00

0.00

ethylene

0.00

0.00

0.30

0.42

0.00

0.00

0.00

nitrogen

0.01

0.00

0.00

0.18

0.00

0.01

0.00

Dioxide carbon

0.00

0.00

0.00

0.00

0.01

0.00

0.00

C4+

0.04

0.00

0.00

0.00

0.28

0.00

0.00

temperature

37.00

36.00

35.00

35.00

28.13

-37.18

-162.78

pressure molar flow (kmole/hr)

6309.00

1700.00

2790.00

2900.00

2500.00

6300.00

101.30

18000.00

24000.00

18500.00

12000.00

2882.01

15118.87

14843.23



1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 34 of 58

Table 3. DMR configuration main streams data

composition

operation condition

Feed gas

Hot MR

Middle MR

NGL

NG to liquefaction

LNG

methane

0.83

0.00

0.42

0.01

0.99

0.99

ethane

0.07

0.25

0.30

0.39

0.01

0.01

propane

0.05

0.64

0.21

0.32

0.00

0.00

n-butane

0.00

0.11

0.00

0.00

0.00

0.00

ethylene

0.00

0.00

0.00

0.00

0.00

0.00

nitrogen Dioxide carbon C4+

0.01

0.00

0.07

0.00

0.01

0.00

0.00

0.00

0.00

0.01

0.00

0.00

0.04

0.00

0.00

0.27

0.00

0.00

temperature

37.00

35.00

35.00

28.05

-37.31

-162.54

pressure

6309.00

4860.00

4860.00

2500.00

6300.00

101.32

molar flow (kmole/hr)

18000.00

25000.00

25000.00

2884.59

15114.94

14622.64


Page 35 of 58

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Table 4. C3-MR configuration main streams data

composition

operation condition

Feed gas

C3-MR

Middle MR

NGL

NG to liquefaction

LNG

methane

0.83

0.00

0.42

0.01

0.99

0.99

ethane

0.07

0.00

0.30

0.39

0.01

0.01

propane

0.05

1.00

0.21

0.32

0.00

0.00

n-butane

0.00

0.00

0.00

0.00

0.00

0.00

nitrogen

0.00

0.00

0.07

0.00

0.01

0.00

Dioxide carbon

0.00

0.00

0.00

0.01

0.00

0.00

C4+

0.04

0.00

0.00

0.27

0.00

0.00

temperature

37.00

35.00

35.00

27.23

-37.35

-162.86

pressure

6309.00

1430.00

4900.00

2500.00

6300.00

100.00

molar flow (kmole/hr)

18000.00

25500.00

25000.00

2892.45

15108.41

14773.58



1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 36 of 58

Table 5. LMTD and minimum temperature approach of the heat exchangers Process

C3-MR

DMR

MFC

Heat exchanger name

E-1A

E-1B

E-1C

E-2A

E-2B

E-1A

E-1B

E-2

E-3

E-1A

LMTD (⁰C) Min. Approach (⁰C)

9.3 2

15.4 4

14.6 3

4.7 2

5.6 3.3

5.4 2

3.7 2

4.8 2.1

6.8 4.5

4.9 2


E-1B E-2

E-3

5.2 2.1

3.6 2

4.7 2.2

Page 37 of 58

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Table 6. COP of the refrigeration cycles Process MFC DMR C3-MR

Cycle

COP

200 300 400 200 300 C3 MR

5.15 2.49 2.31 5.20 1.72 3.03 1.98


Total COP 3.34 2.85 2.34


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 38 of 58

Table 7. Specific power versus methane content of the feed MFC configuration MC*

S.P**

T.S.P***

DMR configuration

C3-MR configuration

S.P

T.S.P

S.P

T.S.P

0.162 0.184 0.214 0.231 0.251

0.389 0.384 0.377 0.376 0.375

0.167 0.195 0.220 0.242 0.264

0.75 0.361 0.157 0.375 0.8 0.353 0.178 0.370 0.85 0.348 0.204 0.365 0.88 0.345 0.221 0.363 0.9 0.344 0.235 0.361 * Methane Content ** Specific Power (kWh/kg LNG) ***Total Specific Power (kWh/kg produced liquids)


Page 39 of 58

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Table 8. Effect of recycle temperature at constant flow rate (40% of gas outlet from D-2 and D-100) DMR configuration

MFC configuration Temperature (⁰C) -60 -70 -80 -88

E.R* 57 77 88 91

S.P** 33.95 34.7 35.03 35.22

E.R 59.5 78.9 90.6 92.1

S.P 34.4 35.4 36.0 36.8

*ethane recovery (percent) **specific power (kW h/kg LNG)*100 (without air coolers power consumption)


C3-MR configuration E.R 60.53 79.8 91.0 92.4

S.P 34.6 36.1 37.4 38.1


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 40 of 58

Table 9. Effect of recycle ratio on the gas outlet from D-2 and D-100 at constant temperature (-88ºC) MFC configuration recycle ratio 0.3 0.35 0.4 0.45 0.5

E.R* 76.7 86.8 92.3 92.3 92.3

S.P** 34.4 34.7 35.3 35.7 36.2

DMR configuration E.R 82.3 90.3 92.1 92.4 92.1

S.P 35.5 36.1 36.7 37.3 37.8

*ethane recovery (percent) **specific power (kW h/kg LNG)*100


C3-MR configuration E.R 82.4 91.2 92.4 92.2 92.3

S.P 36.9 37.5 38.1 38.3 38.7

Page 41 of 58

Table 10. Main equipment power consumption and specific power Configuration

MFC Component Name

C3-MR

Power(kW)*

Compressors

C-100

Air Coolers

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


DMR

Component Name

Power(kW)*

Component Name

Power(kW)*

C-100

5834.73

C-100

5839.84

6065.05

C-300A

16812.07

C-300A

31189.65

C-300

59800.44

C-300B

1664.72

C-300B

15902.8

C-200A

6202.73

C-200A

24093.41

C-300C

10082.51

C-200B

17389.5

C-200B

10203.73

C-200A

2408.9

TE-100

2126.28

C-400A

7147.9

C-200B

7379.48

C-400B

21974.9

C-200C

20461.62

TE-100

2314.61

TE-100

2099.48

AC-200A

0.725

AC-200

2081.63

AC-200

1537.12

AC-2002B

164.9

AC-300A

123.1

AC-300

225.5

AC-300A

51.6

AC-300B

289.12

AC-300B

165.925

AC-300C

15

AC-400

1585.9

Feed

440069.66

440069.66

440069.66

LNG

240715.5

239486.19

236974.34

NGL Specific power (kWh /kg LNG)

194758.37

194902.86

194736.52

0.364

0.391

0.375

Mass flows (kg/hr)

* Mechanical efficiency =0.75



Table 11. Comparison between the introduced configurations and other processes.

New design Vatani et al. design [23] Commercial plant

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 42 of 58

Refrigeration system

Number of Compressor

Number of Tower

Number of Heat exchangers

Ethane Recovery

Specific Power (kW h/kg LNG)

Comparison

MFC

4

1

4

92%

0.364

*

DMR

3

1

4

92%

0.375

*

C3-MR

4

1

5

92%

0.391

*

DMR

3

1

4

93%

0.42

≈ 0.37

APCI [17, 49] ConocoPhillips Design [18] ConocoPhillips Design [19]

C3-MR

̶

1

̶

̶

̶

̶

Cascade

4

1

9

̶

̶

̶

Cascade

3

2

9

̶

̶

̶

Ortloff [16, 23, 42]

̶

5

2&1

5

42-95%

0.28 0.43, 0.5

≈ 0.35

Pure-MR

̶

2-3

3

25-85 %

̶

̶

Fluor Technologies alternatives [50, 51] ̶ not available data


Page 43 of 58

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Table 12. Definitions for exergetic efficiency of the process components

component

Exergy distraction

Compressors and Expanders

𝐼 = 𝐸𝑥𝑖 − 𝐸𝑥𝑜 = ∑𝑚̇𝑖 . 𝑒𝑖 + 𝑊 − ∑𝑚̇𝑜 . 𝑒𝑜 [6]

Throttling valves[37]

𝐼 = 𝐸𝑥𝑖 − 𝐸𝑥𝑜 = ∑𝑚̇𝑖 . 𝑒𝑖 − ∑𝑚̇𝑜 . 𝑒𝑜 [37]

Heat exchangers

𝐼 = 𝐸𝑥𝑖 − 𝐸𝑥𝑜 = ∑𝑚̇𝑖 . 𝑒𝑖 − ∑𝑚̇𝑜 . 𝑒𝑜 [52]

Separators, drums & mixers Tower [48]

Aircoolers Process (Overall)

Exergy efficiency

𝜂=

𝜂 = 1 − ��

∑(𝑚̇ 𝚤 .△𝑒𝑖 )

𝐼 = 𝐸𝑥𝑜 + 𝐸𝑥𝑄𝑜 + 𝑊𝑠ℎ − 𝐸𝑥𝑖 + 𝐸𝑥𝑄𝑖 [15]


∆𝑒 Δ𝑇 ∆𝑒 Δ𝑃

� +�

∑(𝑚̇ 𝚤 .△ℎ𝑖 ) ℎ

𝜂=

𝑊𝑚𝑖𝑛 = 𝐸𝑥𝑖 − 𝐸𝑥𝑜 = ∑𝑚̇𝑖 . 𝑒𝑖 − ∑𝑚̇𝑜 . 𝑒𝑜

𝐼 = 𝐸𝑥𝑖 − 𝐸𝑥𝑜 = ∑𝑚̇𝑖 . 𝑒𝑖 + 𝑚̇𝑎𝑖 . 𝑒𝑎𝑖 + 𝑊 − ∑𝑚̇𝑜 . 𝑒𝑜 − 𝑚̇𝑎𝑜 . 𝑒𝑎𝑜 [15]

𝑊

η=

𝐼 = 𝐸𝑥𝑖 − 𝐸𝑥𝑜 = ∑𝑚̇𝑖 . 𝑒𝑖 − ∑𝑚̇𝑜 . 𝑒𝑜 [6, 7] 𝐿𝑜𝑠𝑡 𝑊𝑜𝑟𝑘(𝐿𝑊) = 𝑇0 (� 𝑚̇𝑜 . 𝑠𝑜 − � 𝑚̇𝑖 . 𝑠𝑖 )

∑𝑚̇.𝑒𝑖 −∑𝑚̇.𝑒𝑜

𝜂= 𝜂=

∑𝑚̇𝑜 .𝑒𝑜 ∑𝑚̇𝑖 .𝑒𝑖

[6]

∑(𝑚̇ 𝚤 .△𝑒𝑖 )

[6]

𝑊𝑚𝑖𝑛 𝐿𝑊 + 𝑊𝑚𝑖𝑛

∑𝑚̇𝑜 .𝑒𝑜 +𝑚̇𝑎𝑜 .𝑒𝑎𝑜

∑𝑚̇𝑖 .𝑒𝑖 +𝑚̇𝑎𝑖 .𝑒𝑎𝑖 +𝑊

𝜂=

� � [52]

∑(𝑚̇ 𝚤 .△ℎ𝑖 ) 𝑐

∑ 𝐸𝑥𝑒𝑟𝑔𝑦 𝑜𝑢𝑡 ∑ 𝐸𝑥𝑒𝑟𝑔𝑦 𝑖𝑛

[15]

[37]


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 44 of 58

Table 13. Exergy efficiency of the designed processes and cycles process

Exergy net in(kW)

Exergy net out(kW)

efficiency

MFC

136829.5

80560.4

0.59

DMR

136392.6

76089.7

0.56

C3-MR

140420.0

76836.1

0.55


Cycles efficiency 400 0.53 300 0.75 200 0.73 300 0.72 200 0.73 300 0.74 200 0.65

Page 45 of 58

Table 14. Result of exergetic efficiency of the process components.

Heat exchangers

Compressor and expander

Valves

MFC

Air Coolers

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


C3-MR

Equip.

Efficiency

Loss(kW)

V-100

0.3

584

V-101

0.4

451

V-3

V-200

0.9

1480

V-4

V-300

0.86

1316

V-5

V-1

0.54

619

V-100

V-2

0.53

159

V-101

V-3

0.69

1014

V-102

V-4

0.31

2415

C-200A

0.775

5340

DMR

Equip.

Efficiency

Loss(kW)

V-1

V-103

0.15 0.55 0.98 0.55 0.89 0.21 0.39 0.21 0.72

2930 587 82 2712 1027 412 592 168.2 2338

C-200A

0.74

634

C-200B

0.75

1832

V-2

Equip.

Efficiency

Loss(kW)

V-100

0.21

408

V-101

0.83

593

V-102

0.26

167

V-103

0.93

2351

V-200

0.89

382

V-300

0.9

2615

V-301

0.92

1071

C-100

0.67

1870

C-200B

0.78

2188

C-200C

0.77

4724

TE-100

0.5

1047

C-300A

0.76

3959

C-300A

0.77

7270

C-200A

0.76

1485

C-300B

0.76

387

C-300B

0.79

3348

C-200B

0.79

3667

C-100

0.68

1944

C-300C

0.78

2195

C-300

0.8

11141

TE-100

0.52

1142

C-100

0.68

1873

TE-100

0.5

1038

E1-A

0.97

2070 2246

E-1A

0.97

1687

E-1A

0.97

2355

E-1B

0.92

E-1B

0.98

1460

E-1C

0.92

1742

E-1B

0.98

1027

E-2

0.97

2702

E-2A

0.74

5609

E-2A

0.94

5983

E-3

0.94

2416

E-2B

0.88

1413

E-2B

0.85

1599

7803.5 AC-200

0.93

3251.5

AC-300A

0.97

962.6

AC-200

0.9

4594.75

AC-300B

0.97

1867

AC-300

0.92

5220.26

AC-300C

0.99

682.5

AC-400

0.86

AC-200A

0.98

741

AC-200B

0.98

1028.5

AC-300A

0.99

173.3

AC-300B

0.99

238.1

Demethanizer Tower Second Law Analysis Efficiency : 0.51 , Lost Work: 1905 kW

demethanizer



1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 46 of 58

Fig.1.

Refrigeration systems

Multi stage

Cascade

Single stage

Pure refrigerant

Mixed refrigerant

Pure refrigerant

Mixed refrigerant

Pure refrigerant

Mixed refrigerant

APCI (AP-N) N2 expansion

APCI(SMR) Black & Veath (PRICO II) BHP(cLNG)

ConocoPhillips (POCP) Technip‐TEALARC

LindeStatoil(MFC)

propane Precooling refrigeration cycle

Shell (DMR) Axens-IFP

APCI (C3-MR) Shell (C3-MR) APCI (AP-X) Shell(PMR)

Two cycles Three cycles


Page 47 of 58

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48


Fig.2. 307

308

309 C-300B

AC-300B

306 C-300A

AC-300A

206 207

208

200

C-200B AC-200A

AC-200B

300

411

409

C-200A

410 C-400A

AC-400 C-400B

408 400

205

407

V-100

405 406

401 201 301

Feed gas E-1A

404

204

V-101

V-200

V-300

305

203

121

403

402

202

302 118

101

105 E-1B

107

303 V-4

E-2

E-3

E

119

120

D-1

102 LNG

108

D 114

103

C

A

D-2

V-1 109

112 115

B

V-2 104

side3 C-100

116 TE-100

113

108 V-3

110

117 Side3R side2R

side2 side1R

side1


T-101

NGL


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48

Fig.3.


Page 48 of 58

Page 49 of 58

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48


Fig.4. 313

C-300

AC-300 312 211

210

209

AC-200

308

310

C-200A 302

C-200B

408 200

306

203

208

V-200

202

Feed gas

301 E-1A

311 207

309

V-300

121

V-301

V-201

304

206

205

201

D-300

308

303

307 V-4

101 E-1B

105

102

119

118 E-2

120

D-1

E-3 LNG

108 113 114

107 103

C-100

116

TE-100

side3 115

D-2

112

V-1 109 V-2

104

108 V-3 Side3R

110

117

side2R

side2 side1R

side1


T-101

NGL


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48

Fig.5.


Page 50 of 58

Page 51 of 58

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48


Fig.6. 318 317

AC-300C 300

220

216 215

202

C-200A 222

210 D-1

201

304

214

D-2

D-3

D-4 310

205 V-1

V-2

203 206

207

211

V-3

212 309

209

204

213 302

101 E-1A

102 E-1B

306

120 119 V-104 E-2B

118

E-1C 112

105 107

116

117

D-100 Side2R

109

113 108

Side3R

Side1R

104

C-100

V-101

114 V-102

115

110 111

T-100

V-100 Side2 NGL Side1

side3


121

D-101

LNG

TE-100 103

V-5 311

P-39 307

V-4

106

312

313

E-2A

305

303

301 Feed gas

C-300A

218

C-200B

222

AC-300A C-300B

217

AC-200

314

AC-300B

C-300C

219

221 C-200C

315

316


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48

Fig.7.


Page 52 of 58

Page 53 of 58

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48


Fig.8.



1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48

Fig.9.


Page 54 of 58

Page 55 of 58

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48


Fig.10.



1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48

Fig.11.


Page 56 of 58

Page 57 of 58

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48


Fig.12.



1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Fig. 13.


Page 58 of 58

Novel LNG-Based Integrated Process Configuration Alternatives for

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