Aspen Plus® Study to Enhance Benzene Production by · PDF fileAspen Plus® Study to...

Aspen Plus® Study to Enhance Benzene Production by Hydrogen Purification at Dow’s Aromatic Plants

in The Netherlands

Joana de Oliveira Moutinho

Thesis to obtain the Master of Science Degree in

Chemical Engineering

Supervisors: Professora Doutora Maria Diná Ramos Afonso

Master Jan Tange

Bachelor Levien Everaert

Examination Committee

Chairperson: Professor Doutor Sebastião Manuel Tavares da Silva Alves

Supervisor: Professora Doutora Maria Diná Ramos Afonso

Members of the Committee: Professora Doutora Maria Rosinda Costa Ismael

May, 2015

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“Tell me and I forget

Show me and I may remember

Involve me and I understand”

by Benjamim Franklin

“Keep it simple” by Wim Kamperman

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Acknowledgments First of all, I would like to thank Jan Tange for giving me the opportunity to work in such a great

international enterprise as The Dow Chemical Company. In particular, I also would like to deeply thank

Levien Evaraert and Wim Kamperman for their availability to help me and answer my questions

anytime. I had learned a lot from them during the four months internship.

I would like to express my gratitude to Professor Diná Afonso for all her guidance, support in many

issues, friendship and incredible patience during these last 9 months.

I also would like to thank all my friends and colleagues that in many ways helped me along the MSc

degree in Chemical Engineering.

A word of gratitude especially goes to Pedro Galamba for all his love, patience and endless support

shown throughout the last few years.

Last, but certainly not least, I would like to thank my parents for their unconditional support along the

journey that has been my academic life. This thesis is dedicated to them.

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Abstract This work studied the hydrogen purification in the cold box unit (Aromatics-2 facility) at Dow´s plant in

The Netherlands, with the main aim of rising the benzene production in 35% at the hydrogenation

reactor (Aromatics-1 facility), whenever the hydrodealkylation reactor was stopped.

A model of the cold box was developed by using the Aspen Plus® V8.0 software and it was validated

based on the plant process data, revealing deviations < 10%. The cold box model was tested for its

robustness, being concluded that the model was valid for feed flow rates from 2,8 up to 30 t/h.

Subsequently, three distinct scenarios were analysed by using the developed Aspen model. Case

study 1 represented the current plant situation, in which the hydrodealkylation reactor was working.

Case Study 2 represented the main aim of this study, i.e., it considered the hydrodealkylation reactor

stopped. Additionally, Case Study 3 was analogous to Case Study 2 but it considered the washing

column working normally.

Sensitivity analyses were conducted for the Case Study 2 whereas Case Studies 1 and 3 were

analysed by single simulations. In all the cases, a cold box feed stream of 22,7 t/h and a hydrogen

purity of 90% (v/v) were set. In Case Study 1, the heavies ended up in the ethane outlet stream unlike

Case Study 3, in which the heavies ended up in the methane outlet stream. In Case Study 2, the

methane stripper was dispensable and yet the model yielded trustworthy results.

Keywords: Aspen; benzene production; cold box unit; hydrodealkylation reaction; hydrogen

purification; trickle flow reactor.

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Resumo Neste trabalho estudou-se a purificação de hidrogénio na caixa fria (unidade de Aromáticos-2) da

fábrica de produtos químicos da Dow na Holanda, com o objectivo primordial de aumentar a produção

de benzene em 35% no reactor de hidrogenação (unidade de Aromáticos-1), quando o reactor de

hidrodealquilação estiver parado.

Desenvolveu-se um modelo da caixa fria utilizando o software Aspen Plus® V8.0, que foi validado

pelos dados processuais da fábrica, revelando desvios < 10%. Efectuou-se um teste de robustez que

demonstrou que o modelo era válido para caudais de alimentação da caixa fria na gama 2,8-30 t/h.

Posteriormente, analisaram-se três cenários fabris recorrendo-se ao modelo desenvolvido. O Caso

Estudo 1 representou a situação actual da fábrica, na qual o reactor de hidrodealquilação estava em

funcionamento. O Caso Estudo 2 representou o objectivo deste projecto, i.e., considerou-se o reactor

de hidrodealquilação parado. O Caso Estudo 3, idêntico ao 2, considerou a coluna de lavagem em

funcionamento.

Realizaram-se diversas análises de sensibilidade ao Caso Estudo 2, enquanto que os Casos Estudo

1 e 3 foram analisados por simulações individuais. Em todos os casos, considerou-se a alimentação

da caixa fria de 22,7 t/h e fixou-se a pureza do hidrogénio em 90% (v/v). Nos Casos Estudo 1 e 3 os

componentes pesados surgiram nas correntes de saída da caixa fria referentes ao etano e ao

metano, respectivamente. O modelo do Caso Estudo 2 forneceu resultados fidedignos e concluiu-se

que a coluna de stripping era dispensável.

Palavras-chave: Aspen; hidrodealquilação; produção de benzeno; purificação de

hidrogénio; reactor de hidrogenação; unidade criogénica.

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Contents

Acknowledgments .................................................................................................................................. V

Abstract ................................................................................................................................................ VII

Resumo ................................................................................................................................................. IX

List of Figures ...................................................................................................................................... XIII

List of Tables ...................................................................................................................................... XVII

List of Abbreviations ............................................................................................................................ XIX

List of Symbols .................................................................................................................................... XXI

List of Compounds .............................................................................................................................. XXI

1 Introduction ...................................................................................................................................... 1

1.1. Work introduction ......................................................................................................................... 2

1.2. State of the art ............................................................................................................................. 4

1.3. Original contributions ................................................................................................................. 10

1.4. Dissertation outline .................................................................................................................... 10

2 Company ........................................................................................................................................ 11

2.1. The Dow Chemical Company .................................................................................................... 12

2.2. Light hydrocarbons (LHC) department, Terneuzen ................................................................... 13

3 Aromatics facilities ......................................................................................................................... 15

3.1. Aromatics facilities [50] ................................................................................................................. 16

3.2. Aromatics-2 facility ..................................................................................................................... 19

4 Modelling of Dow´s cold box unit ................................................................................................... 27

4.1. Materials .................................................................................................................................... 28

4.2. Flowsheet .................................................................................................................................. 29

4.3. Assumptions .............................................................................................................................. 33

4.4. Equipments specifications ......................................................................................................... 34

5 Simulations results ......................................................................................................................... 41

5.1. Model validation ......................................................................................................................... 42

5.2. Model robustness ...................................................................................................................... 51

6 Case studies .................................................................................................................................. 57

6.1. Case Study 1 ............................................................................................................................. 59

6.2. Case Study 2 ............................................................................................................................. 62

6.3. Case Study 3 ............................................................................................................................. 67

7 Conclusions and future work .......................................................................................................... 71

7.1. Conclusions ............................................................................................................................... 72

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7.2. Future work ................................................................................................................................ 73

Bibliography .......................................................................................................................................... 75

Appendixes ........................................................................................................................................... 79

Appendix A1 – Case Study 1 ............................................................................................................ 79

Appendix A2 – Case Study 2 ............................................................................................................ 86

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List of Figures Figure 1 General methodology framework…………………………………………………………………….3

Figure 2 World benzene sources. HDA – hydrodealkylation of toluene and higher molecular weight

aromatics; TDP – toluene disproportionation [1]………………….………………..……………………….…4

Figure 3 Schematic illustration of the location of trickle, mist, bubble and pulsing regimes with respect

to gas and liquid flow rates [20]………………………...………………..……………………………………....7

Figure 4 The LHC department, the core of the Dow site in Terneuzen. LHC – light hydrocarbons; E.O.

– ethylene oxide; L.D.P.E. – low density polyethylene [49]……………………………………………13

Figure 5 Schematic overview of the aromatics facilities. The blue dashed line indicates the option to

forward high purity H2 from the cold box unit at Aro.2 facility to the hydrogenation reactor at Aro.1

facility. BTX – benzene/toluene/xylenes mixture; TX – toluene/xylenes mixture; HBTX – hydrogenated

benzene/toluene/xylenes mixture; BT – benzene/toluene mixture, EB – ethylbenzene…………..……18

Figure 6 Schematic overview of the Aromatics-2 facility. TX - toluene/xylenes mixture, HDA –

hydrodealkylation. The blue dashed line indicates the option to forward high purity H2 from the cold

box unit at Aro.2 facility to the hydrogenation reactor at Aro.1 facility, whenever the HDA reactor is

stopped…………………………………………………………………………………………………………..19

Figure 7 Illustration of the pre-cooling and drying processes……………………………………………...22

Figure 8 Picture of the cold box unit………………………………………………………………………….23

Figure 9 Schematic overview of the cryogenic section (cold box unit delimited by a dashed line) based

on Dow´s flowsheet and IP.21 process data………………………………………………………………...26

Figure 10 Aspen modelling - flowsheet of the cold box unit……………………………………………….33

Figure 11 Multi stream heat exchanger………………………………………………………………………35

Figure 12 Aspen approach for the ethylene chiller, B1-heater for ethylene stream, B2-cooler for hot

stream…………………………………………………………………………………………………………....36

Figure 13 Methane stripper. …….…………………………………………………………………………….37

Figure 14 Flash separator. …………………………………………………………………………………....38

Figure 15 Other equipments: a) pump, b) valve, c) mixer, and d) splitter………………………….....….39

Figure 16 Comparison of the flow rates of the hydrogen outlet stream…………………………………..43

Figure 17 Comparison of the flow rates of the methane outlet stream………………………………..… 43

Figure 18 Comparison of the flow rates of the ethane outlet stream……………………………………..43

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Figure 19 Temperature distribution of the main process stream through the cold box.

Direction: outside → inside. …………………………………………………………………………………..46

Figure 20 Temperature distribution of the methane process stream through the cold box.

Direction: inside → outside. …………………………...……………………………………………………...46

Figure 21 Temperature distribution of the ethane process stream through the cold box.

Direction: inside → outside. …………………………………………………………………………………..47

Figure 22 Temperature distribution of the hydrogen process stream through the cold box.

Direction: inside → outside..…………………………………………………………………………………. 47

Figure 23 Pressure distributions through the cold box unit. …………….…………………………….…..48

Figure 24 Model robustness test concerning the flow rates of each cold box outlet stream…….……..52

Figure 25 Model robustness test concerning the temperature distribution of each cold box outlet

stream……………………………………………………………………………………………………………53

Figure 26 Model robustness test concerning the pressure distribution of each cold box outlet

stream……………………………………………………………………………………………………………54

Figure 27 Model robustness test concerning the flow rate of refrigerant supplied to the ethylene

chiller……………………………………………………………………………………………………………. 54

Figure 28 Model robustness test concerning the flow rates of the methane

stripper……………………………………………………………………………………………...……………55

Figure 29 Schematic overview of Case 1 (Aro.2 facility): green blocks working. LHC I/II: ethylene

plants; DC-MTH: methanator; GB1,2: compressors; DA-WASH: washing column; DC-HDA:

hydrodealkylation reactor; CB: cold box unit. ….……………………………………………………………59

Figure 30 Schematic overview of Case 2 (Aro.2 facility): grey blocks stopped and green blocks

working. LHC I/II: ethylene plants; DC-MTH: methanator; GB1,2: compressors; DA-WASH: washing

column; DC-HDA: hydrodealkylation reactor; CB: cold box unit….……………………………………….62

Figure 31 Case 2 - sensitivity analysis concerning the compositions of the cold box outlet

streams…..............................................................................................................................................63

Figure 32 Case 2 – sensitivity analysis concerning the H2 composition of the hydrogen outlet

stream……………………………………………………………………………………………………………64

Figure 33 Case 2 – sensitivity analysis concerning the temperature of the cold box outlet streams….65

Figure 34 Case 2 - sensitivity analysis concerning the pressure of the cold box outlet streams.……..66

Figure 35 Case 2 - sensitivity analysis concerning the refrigerant flow rates in the cold box unit……..66

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Figure 36 Case 2 - sensitivity analysis concerning the flow rates of the methane stripper…………….67

Figure 37 Schematic overview of Case 3 (Aro.2 facility): grey blocks stopped and green blocks

working. LHC I/II: ethylene plants; DC-MTH: methanator; GB1,2: compressors; DA-WASH: washing

column; DC-HDA: hydrodealkylation reactor; CB: cold box unit………………………………………….68

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List of Tables Table 1 Comparison of hydrogen purification technologies [38]…………………………………………..…8

Table 2 Composition of Dow´s cold box outlet streams (sampling on November19, 2014)…..………..30

Table 3 Mass flow rate of the cold box outlet streams (t/h)………………………………………………. 30

Table 4 Composition and mass flow rates of the cold box feed stream...…………………………….….31

Table 5 Blocks characterization according to Aspen library……………………………………………….34

Table 6 Flow rates deviations between Aspen results and IP.21 database of the hydrogen outlet

stream……………………………………………………………………………………………………………44

Table 7 Flow rates deviations between Aspen results and IP.21 database of the methane outlet

stream……………………………………………………………………………………………………………44

Table 8 Flow rates deviations between Aspen results and IP.21 database of the ethane outlet

stream……………………………………………………………………………………………………………44

Table 9 Aspen results of the ethylene chiller………………………………………...……………………...49

Table 10 Aspen results of the methane stripper…………………………………………………………….50

Table 11 Outlet results of the deep cooler no.2……………………………………………………………..50

Table 12 Results of the flow rates of each outlet stream for three distinct total feed gas flow

rates…………………………………………………………………………………………………...…………52

Table 13 Different parameters of the two simulation models………………………………………………58

Table 14 Case 1 – Feed flow rates and compositions of the cold box……………………………………60

Table 15 Case 1 – Aspen results for the flow rates, compositions, and operating conditions of the cold

box outlet streams……………………………………………………………………………………………...60

Table 16 Case 1 – Aspen results for the operating conditions and flow rates in the ethylene chiller …61

Table 17 Case 1 – Aspen results for the operating conditions, flow rates, and compositions of the

methane stripper streams……………...…………………………………………………………………….. 61

Table 18 Aspen results for the flow rates of the cold box outlet streams……………………………….. 64

Table 19 Case 3 – Feed flow rates and compositions of the cold box …………………………..………68

Table 20 Case 3 – Aspen results for the flow rates, compositions, and operating conditions of the cold

box outlet streams ……………………………………………………………………………………………..69

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Table 21 Case 3 – Aspen results for the operating conditions and flow rates of the refrigerant in the

ethylene chiller……………...…………………………………………………………………………………..70

Table 22 Case 3 – Aspen results for the operating conditions, flow rates and compositions of the

methane stripper streams……………………………………………………………………………………...70

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List of Abbreviations Aro.1 Aromatics-1 facility

Aro.2 Aromatics-2 facility

Aro.3 Aromatics-3 facility

ASPEN Advanced System for Process Engineering

BT Benzene/toluene mixture

BTX Mixture mainly composed by benzene, toluene, and xylenes

CB Cold box

DC-HDA Hydrodealkylation reactor

DC-MTH Methanator

DA-WASH Washing column

DesignSpec Design specification

EB Ethylbenzene

EO Ethylene oxide

GB1 Compressor 1

GB2 Compressor 2

HBTX Hydrogenated benzene/toluene/xylenes mixture

HDA Hydrodealkylation

IP.21 InfoPlus.21 database

LDPE Low density polyethylene

LHC Light hydrocarbons

LHC I/II Ethylene plants

LPG Liquefied petroleum gas

PSA Pressure swing adsorption

PSRK Predictive Redlich-Kwong-Soave

Pygas Pyrolisis gasoline

NRTL Non-Random Two-Liquid

TBR Trickle bed reactor

TDP Toluene disproportionation

TX Mixture mainly composed by toluene and xylenes

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List of Symbols

Symbol Variable Units

T Temperature °C

P Pressure bar, barg

C Concentration vol%, wt%

Qm Mass flow rate kg/h, t/h

Q Heat duty kW

Vf Vapour fraction -

in Inlet -

out Outlet -

vap. Vapour phase -

liq. Liquid phase -

List of Compounds

Compound Molecular formula Molecular weight (g/mol)

Benzene C6H6 78,11

Ethane C2H6 30,07

Ethylene C2H4 28,05

Hydrogen H2 2,02

Hydrogen sulfide H2S 34,08

Methane CH4 16,04

Toluene C7H8 92,14

Xylenes C8H10 106,16

Water H2O 18,01

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1

1 Introduction

Contents

1.1. Work introduction…………………………………………………………………………….2

1.2. State of the Art………………………………………………………………………………..4

1.3. Original contributions……………………………..……………………………………….10

1.4. Dissertation Outline…………………………………………….…………………………..10

2

1.1. Work introduction

Currently, in Dow´s chemical plant at Terneuzen, benzene is mainly produced from pyrolysis

gasoline (pygas). From Dow’s pyrolysis gasoline1, a mixture mainly composed of benzene, toluene

and xylenes (BTX) is obtained by extraction and further treated to recover benzene.

In a trickle bed reactor (TBR) located at the Aromatics-1 facility, a hydrogenation reaction takes

place under a trickle flow regime. This reactor is fed with a BTX liquid phase and a gas phase

consisting of 60% (v/v) hydrogen.

Benzene is separated from the TX mixture, the latter of which is subsequently sent to the Aromatics-

2 facility where it undergoes a hydrodealkylation reaction in order to produce more benzene. This

process will be further described in detail. Depending on benzene market price and its cost

production, sometimes the hydrodealkylation reactor is shut down and the TX mixture is directly sold

in the market.

Nevertheless, if the purity of the hydrogen fed into TBR would raise, it would allow to reduce the

hydrogen flow rate and to augment the BTX liquid flow rate, which might in turn lead to the increase

of the benzene production rate simultaneously.

The aim of this project was to investigate the possibility of using the cold box unit, located at the

Aromatics-2 facility, to purify the 60% (v/v) hydrogen whenever the hydrodealkylation reactor of this

facility is not working. For this plant scenario, the economics should be very favourable as the TX

mixture might be sold and additionally, the benzene production might increase.

Thus, several stages were performed, namely:

! The first stage was developing a process model for the cold box unit by using the Aspen

Tech Process Modelling V8.0 - Aspen Plus® V8.0 software;

! Then, the validation of the developed Aspen model was accomplished based on the plant

process data;

! The last stage consisted in the analysis of three different plant scenarios using the

developed Aspen model;

Figure 1 illustrates the general framework of the methodology followed comprising the four

sequential stages.

11Pyrolysis gasoline is a by-product of ethylene production. It contains a mixture of hydrocarbon compounds in

the C4-C12 boiling range and the primary constituent is benzene, which can range from 25-45% of the stream

mass.

3

Figure 1 General methodology framework.

4

1.2. State of the art

An extensive literature review was conducted in order to provide a global perspective on the

background and current situation of the benzene production industry. This literature review focused

on the technology, techniques and processes involved in this project, in terms of the worldwide

benzene production and in terms of the hydrogenation technology, as well. The latter intended to

enlighten the readers on trickle bed reactors and all the issues related to the system operation, such

as the hydrogen purity. Since the main aim of this study was purifying of the hydrogen fed into a

trickle bed reactor, the literature review also provided an overview on the typical industrial methods

for hydrogen purification.

1.2.1. Benzene production

Benzene, having the highest production in comparison to xylenes and toluene, is mainly produced

through the extraction of pyrolysis gasoline from steam-crackers also known as pygas (47%),

followed by catalytic reforming (33%). A significant percentage of benzene, ca. 15%, is obtained by

the hydrodealkylation (HDA) of toluene and higher molecular weight aromatics, and/or by toluene

disproportionation (TDP). About 5% of benzene is obtained from coke oven light oil (Figure 2).

Since benzene is ranked a co-product, the production volumes and respective economics are

constantly varying.

Pyrolysis Gasoline 47%

Reformate 33%

HDA/TDP 15%

Coke Oven Light Oil 5%

Reformate 68%




Reformate 33%

HDA/TDP 15%


Reformate 68%



6

3. Aromatics – Sources, Demand and Applications

4PVSDFT�The main sources of aromatics (benzene, toluene, xylenes) are reformate from cata-lytic reforming, pyrolysis gasoline from steam-crackers and coke oven light oil from coke oven plants (Fig. 3-01). The reformate from catalytic reforming provides the basic supply of benzene, toluene, xylenes and heavier aromatics. The majority of toluene and heavier aromatics from reformate is converted to benzene and xylenes and is mainly used for p-xylene production. The remaining supply of aromatics is produced from pyrolysis gasoline and from coke oven light oil.

Benzene, which has the highest production rate beside xylenes and toluene, is mainly produced from pyrolysis gasoline, followed by reformate. A significant percentage of Benzene, i.e. 15%, is also obtained from the hydrodealkylation (HDA) of heavier aro-matics and from toluene disproportionation (TDP). The smallest percentage of about 5% of benzene is obtained from coke oven light oil (Fig. 3-02).

8PSME�"SPNBUJDT�4PVSDFT 8PSME�#FO[FOF�4PVSDFT

Fig. 3-01 Source: HPP Science Fig. 3-02 Source: HPP Science

HDA: Hydrodealkylation of heavier aromatics

TDP: Toluene disproportionation

Aromatics (benzene, toluene and xylenes) rank amongst the most important intermediate products in the chemical industry and have a wide range of applications.

Figure 2 World benzene sources. HDA – hydrodealkylation of toluene and

higher molecular weight aromatics; TDP – toluene disproportionation [1].

5

Pyrolysis gasoline is mainly used as a source of recovering benzene and other hydrocarbons. A

small amount is sometimes used as refinery feedstock to boost octane in motor fuels [2]. Producing

mixed xylenes from pyrolysis gasoline is uneconomical due to the low content of xylenes and the

high content of ethylbenzene in pyrolysis gasoline.

Bearing in mind the relative imbalance of benzene demand against a limited supply, it is relevant to

understand how to rise benzene production. Certainly, extensions of the existing sources (Figure 2)

can be used to produce additional benzene if the operation costs of these operations technologies

are favourable. Some factors influence the sources of benzene [1], for example:

" Hydrodealkylation remains a swing source due to the variable economic margins between

benzene and the higher molecular weight aromatics. Considering the favourable price

differentials that occur periodically, the HDA operation will continue to be an option for

benzene production, though relatively inefficient.

" Pygas sources for benzene recovery are diminishing because several of the new ethylene

crackers are based on lighter feedstocks, producing less pygas, hence less benzene. All

benzene produced in steam crackers is essential to be recovered as chemical product.

If pure aromatic compounds are to be obtained, such as benzene, the respective feed must undergo

a pretreatment before the compounds can be separated by extractive distillation. Hydrogenation

processes have become the most appropriate technologies for removing impurities such as di-

olefins, oxygen, sulphur and nitrogen components [3].

Crude pygas from steam-crackers tends to polymerize and form gum, due to its high content of di-

olefins, even when stored in tanks under nitrogen blanketing. In Dow, there is a hydrogenation

reactor responsible for converting the di-olefins into olefins and styrene into ethylbenzene.

1.2.2. Trickle bed reactor (TBR)

Dow’s hydrogenation reactor operates as a trickle bed reactor (TBR). The latter is the most widely

used type of reactors in the process industry, in which a gas and a liquid phases flow co-currently

downwards over a packed catalyst bed. TBRs have achieved widespread commercial acceptance in

many gas–liquid–solid industrial applications. They are mainly employed in the petroleum industry

(for hydrocracking, hydrotreating, alkylation, etc.), the petrochemical industry, and the chemical

industry (for aldehydes hydrogenation, reactive amination, liquid-phase oxidation, etc.). TBRs are

also used in wastewater treatment, bio- and electro-chemical processing [4-8].

With ever-increasing demand for light oil and other products based on TBR technology, any

enlightenment of the complex phenomena taking place inside TBR may lead to major technical

and/or economical breakthroughs. A significant number of review articles have been published

during the past several decades, covering many aspects of three-phases TBRs (such as chemical

kinetics, mass and heat transfer, hydrodynamics). Some of these review articles have also focused

on the hydrodynamic and reactor modelling of TBR [9-14].

6

Usually, the commercial trickle bed reactors operate adiabatically, at high temperatures and

pressures, and often involve hydrogen and organic liquids. Industrial trickle beds have typically

diameters of 2–4 m and heights of 15–25 m. These reactors are often applied to perform strong

exothermic reactions, such as the hydrogenation of unsaturated hydrocarbons [14].

One of the major disadvantages of trickle bed reactors is their poor capability to dissipate the heat

released by the reaction. Considering the low heat capacity of the gas, this heat is most often

removed by the liquid phase, although sometimes evaporation of the liquid phase is used. Whether

the generated heat is not adequately removed, hot spots may be formed and catalyst deactivation

may occur. The formation of hot spots must be prevented for safety reasons, e.g. no leakage should

occur. The catalyst deactivation prejudices the selectivity, production capacity and operation

flexibility.

The knowledge of the flow regime in which the reactor will operate is very important because other

hydrodynamic parameters, especially the mass transfer rates, are differently affected by

hydrodynamics in each regime. In TBR, four flow regimes may occur, depending on the flow rates of

the gas and liquid phases, the fluids properties and the size and shape of the stationary phase [15-19].

The flow regime boundaries with respect to the gas and liquid flow rates are schematically shown in

Figure 3. Each flow regime corresponds to a specific gas–liquid interaction, thus having a great

influence on parameters, such as the pressure drop, liquid hold-up, and mass and heat transfer

rates.

The trickle flow regime occurs at relatively low gas and liquid flow rates. The liquid trickles over the

packing and the gas is the continuous phase occupying the remaining voids. This regime features

partially or totally wet catalyst particles, depending on the flow rate and structure of the liquid flow. At

high gas and low liquid flow rates, transition to mist flow regime occurs. The liquid mainly travels

down the column as droplets entrained by the continuous gas phase. The bubble flow regime occurs

at high liquid flow rates and low gas flow rates. In this case, the liquid is the continuous phase and

the gas moves in the form of dispersed bubbles. At moderately high gas and liquid flow rates, the

pulsing flow regime is attained. This regime is characterized by the successive passage of liquid-rich

and gas-rich regions through the bed.

7

In the chemical industry, the trickle bed reactors are usually operated in the trickle flow regime for

hydrogenation processes. This trickling flow allows the gas phase to be dissolved in the liquid phase.

Conversely, the pulsing flow means “pulsed flows” with gas followed by liquid, i.e., the liquid

periodically blocks the small channels between catalyst particles and forms liquid plugs that prevent

the gas flow.

It is worth to remark that if the hydrogen does not reach the catalyst surface, the hydrogenation

reaction does not take place. On the other hand, if only gas reaches the catalyst surface, an excess

of hydrogen may be available leading to hot spots formation since a local reaction of any component

present at the catalyst surface occurs, which in turn can lead to the catalyst deactivation.

Whereas the existence of the various flow regimes in trickle bed reactors is well known and many

efforts have been done to establish a theorytical rule to determine the regimes boundaries, none is

able to accomplish such a complex task at present. Numerous attempts have been done to model

hydrodynamics of trickle bed reactors. Reviews on published models in this area can be found [21-23].

A number of flow maps for regime transition are available in the literature. Weekman and Myers

(1964), Sato et al. (1973), and Satterfield (1975) used G (gas flow rate) versus L (liquid flow rate) as

the axis of the flow rates map [4,24,25], whereas Turpin and Huntington (1967) represented the map as

L=G versus G [26]. Different investigators have reported the effects of the fluids physical properties on

regime transitions. Chou et al. (1977) observed that with the viscosity increase or surface tension

decrease the transition shifted to lower gas and liquid superficial velocities. On the other hand, Chou

et al. (1977) and Sai and Varma (1988) observed that at a constant gas flow rate, the transition from

trickling to pulsing regime occurred at higher flow rates as the bed void fraction increased, but the

data of Sato et al. (1973) did not show any significant variation with respect to variations of the void

fraction [25,27,28].

1. Introduction

the trickle flow regime. So, knowledge of the flow regimes inside the packed beds isimportant in the investigation of hot spots formation.

Figure 1.2: Different flow regimes in three phase packed beds as a function of flowrate of gas and liquid [10].

1.2.2 Wetting efficiency

External catalyst wetting efficiency is an important parameter in trickle bed react-ors. Internal wetting efficiency is less important than the external wetting efficiencyin trickle bed reactors because the internal wetting efficiency is almost one dueto the capillary effects and external wetting efficiency is less than one. Part of theliquid is more or less stagnant and makes a high or low resistance to mass transfer(depending on the thickness of the liquid layer on the catalyst surface), whichleads to non-uniform reactant concentration around the catalyst, in case of highintrinsic reaction rates. Another phenomenon, which can reduce reactor efficiency,is incomplete wetting of the catalyst particles. In principle, incomplete wetting cantake place on a reactor scale as well as on a particle scale. Normally, incompletewetting of large catalyst zones can be one of the reasons for the formation of hotspots in trickle bed reactors [1].

4

Figure 3 Schematic illustration of the location of trickle, mist, bubble and pulsing regimes with

respect to gas and liquid flow rates [20].

8

1.2.3. Hydrogen purification

As aforementioned, the hydrogen has a crucial role on the hydrogenation reaction, thus its partial

pressure requires a constant control.

Generally, hydrogen recovery is cheaper than its generation [29,30]. The selection of a hydrogen

separation process depends on process features (e.g., primary purification requirements, feed

composition, feed pressure, product flow rate, purity, and by-product recovery) and project

considerations (e.g. process flexibility, feed turndown ratio, reliability, and ease of future

development).

Hydrogen may be recovered by purification if this is economically viable. Three purification

technologies are used industrially for hydrogen recovery: pressure swing adsorption (PSA),

membrane permeation and cryogenic separation. In general, for high flow rates, cryogenic

separation or PSA are most economical, whereas membrane separation is preferred for low flow

rates and high feed pressures. Abrado and Khurana [31], Wilcher et al. [32], Spill-man [33], and Miller

and Stoecker [34] identified the factors determining the selection between these three hydrogen

purification technologies, compared in Table 1. Absorption [35,36] and selective surface flow

permeation [37] have also been proposed for hydrogen purification. However, these techniques are

not widespread as commercial applications in oil refining.

Liao et al. (2011) [39] presented a rigorous systematic approach to determine the optimal placement

of the purifiers in a hydrogen system. The process identified hydrogen network targets with one

purifier, taking into account minimum hydrogen utility consumption and the purifier feed

concentration.

Hallale and Liu (2001) [40] have developed a process that interpreted the work done by Alves (1999)

[41] by considering the purifiers as interception units that upgrade the purity of the hydrogen source,

taking into account the pressure constraints ,as well as existing compressors.

The growing demand for refinery hydrogen has been addressed by modification or development of

new production and recovery strategies. However, the hydrogen content and the operating pressure

Towler et al. (1996) illustrated the cost correlation of hydrogenrecovery from refinery off-gas using purification units. The meth-odology proposed in their work facilitates the cost of hydrogenrecovery by PSA and membrane permeation using a linear pro-gramming (LP) model to supplement the information in a graphical

display. Work performed by Zhou et al. (2002) represents a novelPSA technology that is fitting to both the operating conditions oflow pressure and avoiding the other limitations of PSA technolo-gies, which they consider to be the purging ratio and separationperformance at low operating pressure. Rabiei (2012) discussed thedifference between the three most common purification and re-covery units for hydrogen in the oil refinery, which are PSA,membrane processes and cryogenic processes. This work in-vestigates the cost for process recovery selection and the capitalinvestment of PSA and membrane processes in terms of hydrogencontent and operating pressure of the refinery.

Hallale and Liu (2001) have produced a valuable process thatinterprets the work done by Alves (1999) by considering the puri-fiers as interception units that upgrade the purity of the hydrogensource taking into account the pressure constraints as well asexisting compressors. In order to gain the highest recovery frompurification, Liu and Zhang (2004) have proposed a hybrid plan anddiscussed a detailed strategy for purifier selection that includesboth PSA and a membrane. The process deploys a mixed integerlinear programming (MILP) model to integrate purification unitswith refinery hydrogen. The optimized superstructure allows en-gineers to define the available options for better purification ofrefinery off-gases and placement of purifiers. Liao et al. (2011b)presented a rigorous, systematic approach to the problem of theoptimal placement of purifiers in a hydrogen system. The processidentifies hydrogen network targets with one purifier, taking intoaccount minimum hydrogen utility consumption and the purifierfeed concentration. Liu et al. (2013a) have presented a conceptualmethod using a quantitative relationship diagram for a hydrogenconservation strategy with purification reuse/recycle, to identifythe optimal Purification Feed Flow-rate (PFFR) and the corre-sponding maximum Hydrogen Utility Saving (HUS).

The growing demand for refinery hydrogen is being addressedby modification or development of new production and recoveryroutes. However, the hydrogen content and operating pressure ofrefinery off-gases have a large influence on both the choice ofprocess recovery and capital investment. Foo and Manan (2006)have made an extensive study of gas resources allocation targets,evaluating both membrane and PSA as purifiers for a conventionalhydrogen system in an oil refinery. However, in their approachusing the Gas Cascade Analysis method, which is practical to

Table 2Conference proceedings.

Advanced control of chemical processes

2011 International symposium on Advanced Control of Industrial Processes (ADCONIP)AIChE 2001 Spring meeting, April, Houston TXEuropean Symposium on Computer Aided Process EngineeringInternational Conference on Information and Automation 2009 ICIA0092005 IEEE International Symposium on Intelligent Control and 13th Mediterranean Conference on Control and Automation

Fig. 1. Trend of related hydrogen management publications.

Fig. 2. Hydrogen management research topics.

Table 3Hydrogen purification technology comparison source (Sabram et al., 2001).

Features Adsorption Membranes Cryogenics

H2 Purity 99.9%þ 90e98% 90e96%H2 Recovery 50e92% 85-95% 90e99%Feed Pressure 150e600psig 300e2300psig ˃75e1100psigFeed H2 Product ˃40% ˃25e50% ˃10%H2 Product Pressure Feed ˂1/3 Feed P Feed/Low PressureH2 Capacity 1-200 MM scfd 1-50 MM scfd 10-75 MM scfdPretreatment Requirement None Minimum CO2,H2O RemovalMultiple Products No No Liquid HC sPower Requirements None/Fuel H2/Feed None/H2/RefrigerationCapital Cost Medium Low HigherScale Economics Moderate Modular Valid

M. Elsherif et al. / Journal of Natural Gas Science and Engineering 24 (2015) 346e356348

Table 1 Comparison of hydrogen purification technologies [38].

9

of refinery off-gases have a large influence on both the selection of process recovery and capital

investment.

Shahraki et al. (2005) [42] have addressed the minimization of refinery hydrogen consumption by

making use of pressure constraints, all possible site-wide piping connections and optimum

placement of the purifiers. This approach claimed to be suitable for revamping industrial systems.

Cryogenic separation method

The term “cryogenic” stands for all the processes, phenomena, and equipments that are arbitrarily

used for temperatures lower than -74°C. The liquefaction of gases has always been important

because the refrigeration of many important scientific and engineering processes depend on

liquefied gases (e.g., separation of oxygen and nitrogen from air, study of material properties at low

temperatures, study of phenomena such as superconductivity and preparation of liquid propellants

for rockets). Until 1940 liquefied gases were mainly produced in laboratories [43].

In cryogenic equipments, the compounds boiling point difference at very low temperatures (i.e., their

relative volatilities difference) is used for their separation. The hydrogen relative volatility with respect

to hydrocarbons is high. The simplest and most common process for hydrogen separation is partial

condensation. In this method, the hydrogen impurities are condensed by cooling the exiting gas in

multi-passes heat exchangers [44]. The required cooling is granted by the Joule-Thompson effect,

thus the condensed hydrocarbons are suddenly depressurized. If required, the cooling duty may be

supplied by independent chilling systems or turbo-expansion of the hydrogen product.

As aforementioned, depending on the hydrogen flow rate and purity required, either cryogenic

separation or PSA are more economical [45]. In the cryogenic method, the initial investment and

energy consumption are relatively low. Due to flutuactions of the feed constituents, it is wise to

prevent optimistic reliance on the plant design, the Joule-Thomson chilling effect, and the turbo-

expanders efficiencies. The cryogenic process is the most efficient and economical for the

separation of gaseous mixtures, especially for large recoveries [46], and can be applied through two

distinct methods:

" Cryogenic process without distillation. The major advantages of this process consists of high

product recoveries, intermediate purity of light products (H2 purity up to 98% (v/v)), high

pressure operation, low operating costs and low pressure drops of light products. A

disadvantage is the impossibility of attaining very high purities of light products.

" Cryogenic process with distillation. It is similar to the previous method but comprises a

distillation tower for product recovery with higher purity. Its advantages consists of high

product recoveries, high purities of light products (H2 purity up to 99,5% (v/v)), high pressure

operation, moderate purities for heavy products, and low pressure drops of light products. A

disadvantage consists of high energy consumption.

10

1.3. Original contributions

According to the state of the art, it appears that much progress has been done in the benzene

production and its recovery. Nevertheless, as there is an imbalance of benzene demand against a

limited supply, chemical industries should test different methods to enhance benzene production.

This work studies the possibility to increase benzene production in Dow´s chemical plant in The

Netherlands, whenever the hydrodealkylation reactor system is stopped. This new approach involves

the existent equipment from the Aromatics-2 facility to purify the hydrogen from the ethylene plant,

and then feed it into the hydrogenation reactor at the Aromatics-1 facility.

In sum, a plant scenario never tested before was analysed throughout this thesis, consisting of a

unique challenge.

1.4. Dissertation outline

This thesis was organized as follows. In Chapter 2, a short introduction to Dow´s chemical plant was

conducted. Chapter 3 presents a background review covering the process description at the

aromatics facilities involved in this study. Chapter 4 describes in detail the implementation of the cold

box model by using Aspen Plus V8.0® software. In Chapter 5, the results from the developed model

were presented and validated based on the plant process data. Furthermore, a robustness test of the

model was performed. Chapter 6 presents the analysis of three different plant scenarios using the

developed model (case studies). Finally, Chapter 7 presents the main conclusions of this thesis and

a few suggestions for future work.

11

2 Company

Contents

2.1. The Dow Chemical Company…………………….……...……………………....…...…..12

2.2.Light hydrocarbons (LHC) department, Terneuzen…………………………………...13

12

2.1. The Dow Chemical Company

On May 18,1897, Herbert H. Dow founded The Dow Chemical Company, in Midland, Michigan.

Initially focused on the production of bleach on a commercial scale, the company quickly expanded

and diversified its product range.

In 1913, Dow left the bleach industry and switched to the production of styrene and saran. However,

it was in 1935 that The Dow Chemical Company entered the plastics industry with Ethocel

ethylcellulose resins and Styron polystyrene resins. In the following 80 years the company has

further grown to become the second largest chemical producer in the world, in between BASF and

Bayers, employing over 53.000 people on 201 different production sites in 36 countries [47].

In the Netherlands, Dow started the Terneuzen site construction in 1961. It has since expanded to

include 17 production facilities and 1700 employees, becoming the second largest production facility

of Dow worldwide and the largest outside the United States. The main focus of Dow Terneuzen is

the cracking of hydrocarbons and production of plastics and chemicals, most of which are

intermediates for consumer products [48]. Dow has formulated the essential elements of the company [47]:

Vision: To be the most profitable and respected science-driven chemical company worldwide.

Mission: To passionately innovate what is essential to human progress by providing sustainable

solutions to their customers.

Corporate strategy: Preferentially invest in a portfolio of technology-integrated, market-driven

performance business that creates value for shareholders and growth for customers.

Values: Integrity, respect for people and protecting the Earth. Nowadays, Dow is leading in the production of a wide range of plastics, including polyethylene and

polypropylene, as well as base chemicals. The products find application in the automotive,

construction; packaging; pharmaceutical; paint; coating and water purification industry. Furthermore,

Dow also produces energy and hydrocarbons from oil and a variety of intermediates for consumer

products [48]. Dow not only focuses on the production and sales of its products, but also actively

works on safety, the environment and health. These activities are not limited to the production sites

and employees, but extend to the surrounding areas and communities.

13

2.2. Light hydrocarbons (LHC) department, Terneuzen

There are three light hydrocarbons (LHC) plants located in Terneuzen, where this internship took

place. These naphtha crackers are the core of the Dow site in Terneuzen. In these plants, naphtha2

and liquefied petroleum gas (LPG) are converted into large amounts of valuable products, such as

like ethylene, propylene, benzene and butadiene. In this steam cracking process, a hydrocarbon

feed like naphtha or LPG is diluted with steam and shortly heated in a furnace, without the presence

of oxygen, to temperatures of around 850°C.

The main products – ethylene, propylene, benzene and butadiene – obtained in the LHC plants can

be converted into more valuable products in other plants on the Dow site in Terneuzen, as

schematically shown in Figure 4. Therefore, the LHC department may be envisaged as the heart of

the site. These three plants deliver the starter products for more valuable products, which are

synthesized in other integrated processes on the Dow site in Terneuzen. One example is the

production of styrene out of benzene and ethylene. Another well-known example is the

polymerization of ethylene to low-density polyethylene (LDPE). Furthermore, ethylene can be

converted into ethylene oxide (EO), which is a starter product for more valuable Dow products.

2 Naphtha is a generic term referring to a class of colourless, volatile, flammable liquid hydrocarbon mixtures. It

is produced through fractionation of petroleum, natural gas, or coal tar and it is one of the highest volume liquid

fractions of crude. Light naphtha is a major raw material for the production of gasoline, petrochemicals and

solvents.

Figure 4 The LHC department, the core of the Dow site in Terneuzen.

LHC – light hydrocarbons; E.O. – ethylene oxide; L.D.P.E. – low density polyethylene [49].

14


15

3 Aromatics facilities

Contents

3.1. Aromatics facilities……………………………..……………………………………..…………………………..16

3.2. Aromatics-‐2 facility....……………………………………………………………...……………………………..19

16

Among the generated cracking products, the pyrolysis gasoline (pygas), is the most relevant product

for this study. The use of pygas yields a BTX mixture (B-benzene; T-toluene; X-xylene), which is

separated ahead at the aromatics facilities in order to recover one of the main products of the LHC

plant, the valuable benzene. Although BTX stands for benzene, toluene and xylenes, BTX also

contains many other compounds (e.g. ethylbenzene; non-aromatics components; heavy compounds

and others).

3.1. Aromatics facilities [50]

After the cracking process, the pygas product undergoes several pre-distillations in order to remove

the C4+- C9

+ fractions. Afterwards, it flows through a depantanizer where the C4+ fractions are

removed by the top, and then recycled to the crackers. The bottom product is fed to a

prefractionator, in which the C9+ fractions are removed by the bottom and then stored. Moreover, a

BTX product is obtained at the top of the prefractionator. Figure 5 represents a schematic overview

of the aromatics facilities.

Afterwards, at the Aromatics-1 facility (Aro.1), the BTX product is fed to a hydrogenation reactor, in

which the di-olefins are converted into olefins and the styrene into ethylbenzene. The reactor uses a

palladium catalyst and the hydrogenation reaction occurs by feeding an excess of hydrogen about

60% (v/v) from the ethylene unit, to ensure that the reaction takes place. Moreover, this reactor must

operate in the trickle flow regime, for instance by increasing the hydrogen partial pressure in the

hydrogen stream. It is possible to reduce the gas superficial flow rate whilst the liquid superficial flow

rate remains constant, thus a trickle flow regime is guaranteed.

In the trickle bed reactor the flows direction is downwards. The reactor is filled with hydrogen, and

the BTX feed passes through it as a finely divided liquid (exothermic process). The product flows

through a train of heat exchangers, towards the second reactor. It is essential that no liquid enters

into the second reactor, otherwise the polymerization reaction might proceed causing a high

pressure drop, and demanding the catalyst replacement. In the second reactor, the olefins are

converted into paraffins, by removing all the double bounds, and the sulfur components are

converted into hydrogen sulfide. The BTX product, which leaves the second reactor, is hydrogenated

(HBTX) and it is fed to a stabilizer, where H2, H2S and light components are eliminated.

At the Aromatics-3 facility (Aro.3), the hydrogenated BTX flows through a distillation tower, where the

crude benzene is separated from the TX mixture. The crude benzene leaves the top of the distillation

tower and flows to an extractive distillation followed by the benzene.

At the Aromatics-2 facility (Aro.2), the TX product is mixed with methanated hydrogen already

processed in the methanator3. This mixture passes through a chain of heat exchangers followed by a

furnace and then it is fed to the hydrodealkylation reactor (HDA), in which the hydrodealkylation 3 Methanator is a reactor fed with hydrogen, which comes from the ethylene plant. The purpose of the methanator is to eliminate the poisonous compounds, such as carbon monoxide (CO) and carbon dioxide (CO2). Using a nickel catalyst, the methanator converts CO and CO2 into methane (CH4) and water (H2O) leading to a methanated hydrogen stream, rich in CH4 and H2.

17

reaction takes place under hydrogen atmosphere. In the furnace, the mixture of TX and H2 flows

through four coils and it is heated up by radiant heat to reach the temperature required for an optimal

efficiency of the HDA reactor. The purpose of the hydrodealkylation reaction is the removal of one or

more alkyl groups from the aromatic ring to convert all alkylbenzenes (toluene, xylenes and

ethylbenzene) into benzene; besides, ethane is also produced. Afterwards, the product that leaves

the reactor bottom is cooled down through a chain of heat exchangers and a condensed product is

obtained. Thus, a gas/liquid mixture joins in a flash separator, where the liquid phase is separated

and subsequently fed to the final distillation tower to eliminate the heavy components from the

benzene product. Meanwhile, the gas phase that leaves the flash separator flows towards the

purification unit.

To prevent cloggings in the piping, some of the components present in the gas phase, such as

benzene, toluene and water, must be removed before entering into the cryogenic section. The

purification unit consists of a washing unit. Since the toluene is suited as washing agent for the

removal of the components previously mentioned, a BTX mixture is used in the washing column.

A cryogenic section is located downstream of the washing unit. In the former, the feed gas is

completely dried and cooled down until the hydrocarbons contained in the gas become liquid, to

separate them from H2. Thus, the hydrogen concentration of the feed gas increases allowing H2 to

be recycled to the HDA reactor in Aro.2 to maintain the hydrodealkylation reaction in progress.

Otherwise, H2 may be sent to the hydrogenation reactor in Aro.1.

18

Figure 5 Schematic overview of the aromatics facilities. The blue dashed line indicates the option to

forward high purity H2 from the cold box unit at Aro.2 facility to the hydrogenation reactor at Aro.1 facility.

BTX – benzene/toluene/xylenes mixture; TX – toluene/xylenes mixture; HBTX – hydrogenated

benzene/toluene/xylenes mixture; BT – benzene/toluene mixture, EB – ethylbenzene.

19

3.2. Aromatics-2 facility

As aforementioned, this study focused on the aromatics plants and the interactions between them.

This work intended to analyse the performance of the cold box, located at Aro.2 facility, and predict

its impact on the benzene production depending on the amount of BTX fed into the hydrogenation

reactor, located at Aro.1 facility. Therefore, a sketch of the main processes surrounding the cold box

unit was depicted for a better insight over the Aromatics-2 facility (Figure 6).

Figure 6 Schematic overview of the Aromatics-2 facility. TX - toluene/xylenes mixture, HDA –

hydrodealkylation. The blue dashed line indicates the option to forward high purity H2 from the cold box

unit at Aro.2 facility to the hydrogenation reactor at Aro.1 facility, whenever the HDA reactor is stopped.

20

3.2.1. Hydrodealkylation process

The hydrodealkylation (HDA) is a commercially viable process used for the production of benzene

from toluene and higher molecular weight aromatics. This process consists of removing one or more

alkyl groups (-CH3) from the aromatic ring, and replacing them by hydrogen atoms. It requires an

excess of hydrogen, high temperature and moderately high pressure.

The Dow Chemical Company has three HDA plants, located in:

Terneuzen, Netherlands (1971);

LaO (1980); Plaquemine, Louisiana

DCG (1998); Dow Central Germany

The hydrodealkylation process has an important role on Dow’s aromatics business since it is the

largest benzene producer worldwide, setting the ceiling to benzene market and ensuring an outlet to

the TX stream (sold in the market or internally reused).

In Terneuzen site, the HDA reactions that produce benzene are the following:

Obviously, for the HDA reactions to take place it is necessary to feed certain amounts of TX and

hydrogen into the Aro2 facility. This hydrogen comes from the ethylene plant towards the

methanator, located at Aro.2 facility. The HDA reactor is the final step of the aromatics facilities

where it is possible to increase the benzene production by means of the hydrodealkylation reactions.

Depending on the daily market situation and the benzene production costs, the HDA reactor may

stop running temporarily, meaning that the TX product may be directly sold in the market instead of

used at Aro.2 facility.

The product that leaves the HDA reactor is cooled down through a chain of heat exchangers,

allowing the separation of a liquid phase rich in benzene. Simultaneously, the gas phase is treated in

a purification unit comprising a washing column and a cryogenic section. Before reaching the

cryogenic section, the gas phase undergoes a pre-cooling and a drying process as described in

section 3.2.2.

C7H8 + H2 à C6H6 + CH4 Eq.1 (toluene) (hydrogen) (benzene) (methane)

C8H10 + 2 H2 à C6H6 + 2 CH4 Eq.2 (xylene) (hydrogen) (benzene) (methane)

C8H10 + H2 à C6H6 + C2H6 Eq.3 (ethylbenzene) (hydrogen) (benzene) (ethane)

21

3.2.2. Pre-cooling and drying process [49]

The washing process is conducted at 23,8°C due to its heat of dissolution. The gas phase containing

benzene and toluene purified, leaves the top of the washing column and flows towards the heat

exchanger cooler no.1. Figure 7 illustrates the pre-cooling and drying processes.

The pre-cooling process starts by cooling down the feed gas to 8,9°C in counter-current flow with

respect to the cold product streams rich in hydrogen, methane and ethane, respectively. A by-pass is

installed in the methane stream to prevent a temperature decrease below the water freezing point.

Due to this cooling a condensation product containing water and traces of all feed gas compounds is

obtained.

The condensation product is separated in a separator, at the outlet of which it expands and is fed to

the rich toluene drum for separation.

On the other hand, the feed gas flows to one of the dryers to remove its humidity. It enters at the top

of the dryer and passes through a molecular sieve whereby the liquid droplets of benzene and

toluene are deposited on the sieve material. Whenever the sieve material becomes saturated, the

dryer must be regenerated, and automatically replaced by the other dryer.

The dryers are changeable and, after their loading time of 24 hours, their fillings have to be

regenerated by means of heated methane (≈185°C). Then, compressed dry methane is taken from

the plant to cool down the dryer. The operating temperature of 10°C is reached by connecting in

parallel the dryer to be cooled down to the dryer that is already in operation.

After leaving the dryers, the gas stream flows towards the next heat exchanger, the cooler no.2,

where the true cooling process begins, and which is the start point of the cryogenic section.

22

3.2.3. Cryogenic section

The cryogenic section contains all the cooling equipments that lead to negative temperatures of the

process streams. It comprises the cooler no.2 and the cold box.

The cold box unit consists of a box containing a package of equipments, such as piping; distillation

columns; heat exchangers and instrumentation used to maintain liquids at extremely cold

temperatures (Figure 8). The space between the equipment and the inner walls of the cold box is

filled with perlite insulation, usually applied in cryogenic insulation. To prevent the formation of

explosive gas mixtures and/or humidity, the cold box is purged with nitrogen at low pressure. This

procedure prevents humidity deposits, such as ice, that would impair the insulating effect.

In this study, the main purpose of the cold box section is to cool down the feed gas until the

hydrocarbons, CH4 and C2H6, become liquid to separate them from H2 in three distinct outlet

streams. Moreover, the hydrogen purity of the feed gas increases such that it might be reused in the

process reactors, HDA reactor (Aro.2) and hydrogenation reactor (Aro.1), respectively.

Figure 7 Illustration of the pre-cooling and drying processes.

23

After leaving the dryer at 10°C and 53 bar, the feed gas stream flows towards the cooler no.2, where

it is cooled down to -60°C in counter-current with respect to the cold product streams rich in

hydrogen, methane and ethane. Figure 9 presents a schematic overview of the cryogenic section.

The next cryogenic exchanger is called deep cooler no.1 because it is located inside the cold box.

However, it works similarly to the coolers no.1 and no.2 previously described. This equipment is

responsible for cooling down the feed gas to -91°C in counter-current with respect to the cold

product streams. After passing through the deep cooler no.1, a condensate product is obtained,

containing all the compounds of the feed gas in different concentrations depending on their physical

properties.

Afterwards, the amount of condensate product is slightly increased by using an ethylene chiller,

containing ethylene as refrigerant. A level control device regulates the ethylene supply to the chiller.

The ethylene chiller is the cold source of the plant responsible for the refrigeration needs due to

insulation heat gains throughout the process, and also to heat gains in cooling down the dryer.

Thereby, ethylene evaporation is used as cold source for the cooling and liquefaction of the process

streams. In this apparatus, the ethylene refrigerant is totally vaporized by removing sensible heat

from the feed gas stream. Then, the vaporized ethylene is discharged to the suction vessel of the

first stage of the ethylene refrigeration compressor in the ethylene facility.

Figure 8 Picture of the cold box unit.

24

The separator no.1 is located downstream of the ethylene chiller, inside which a flash separation

takes place. The liquid product that leaves the bottom of the separator is rich in ethane and most

hydrogen sulfide is dissolved in this stream. This liquid stream is expanded and then stripped in the

methane stripper. It is noteworthy that the hydrogen sulfide in the feed gas stream does not entail

troubles in the cryogenic section since the quantity remaining after the drying process is dissolved in

the liquefied ethane and does not cause cloggings during the re-evaporation.

Meanwhile, the gas stream that leaves the top of the separator no.1 is rich in methane and

hydrogen. Its temperature has to be reduced in order to raise the concentration of H2 up to 90%

(v/v). Thus, another cryogenic exchanger, named deep cooler no.2, is used to cool down the feed

gas stream to -136°C by means of the cold products streams rich in methane and hydrogen. At this

temperature, part of the methane in the feed gas condenses and forms a new cooling agent.

Thereby, the product stream rich in liquid methane is collected in the separator no.2 where a flash

separation takes place, resulting in two different products.

The gas stream that leaves the top of the separator is rich in hydrogen with a purity in the range 90-

95% (v/v). The purified hydrogen flows back into the system through the several heat exchangers,

being heated up until it leaves the cryogenic section at 2°C. In the meantime, the liquid product that

leaves the bottom of the separator, rich in methane, is expanded by a control valve undergoing a

pressure drop from 53 bar to 5,6 bar, helped by the pipe diameter increase from 3” to 10”. Then, the

methane product stream flows back to the deep cooler no.2, which is the main responsible for the

cooling effect. It is noteworthy that the outlet temperature of the cryogenic exchanger deep cooler

no.2 is determined by the quantity of the expanded methane. In case a heavy refrigeration is needed

in the plant, a greater quantity of methane should be expanded. Further, the methane cold product

flows back through the several heat exchangers, being heated up until it leaves the cryogenic section

at -1,8°C. Moreover, the recovered methane either from the methane stripper (stream 15) and from

the separator no.3 (stream 23) are added to the methane product stream (streams 13 and 21),

respectively.

On the other hand, the liquid product that leaves the bottom of the separator no.1 is expanded and

supplied to the head of the methane stripper. In this equipment, CH4/C2H6 separation takes place as

the feed stream flows in counter-current with respect to vapour, via 5 trays, such that a fraction rich

in ethane is formed in the sump and a fraction rich in methane is obtained on the top of the column,

where a pressure control valve maintains the rectification system pressure. A fraction of the main

feed gas stream is used in the reboiler to deliver the heat for the rectification system.

25

The liquid product that leaves the bottom of the methane stripper is rich in ethane and flows towards

the deep cooler no.1, where it plays the role of cold fluid occurring its partial vaporization in counter-

current with respect to the main feed gas stream. After partially vaporized, this stream undergoes the

final separation in separator no.3. As aforementioned, the final flash separation of CH4/C2H6 occurs,

whereas the remaining methane (stream 23) is added to the methane outlet stream (stream 21).

From the separator no.3, a bottom liquid product rich in ethane is pumped towards the cooler no.2

(outside the cold box) where it is vaporized, leaving the cryogenic section at 5°C. The ethane outlet

stream heads to the ethane cracking furnaces in the ethylene plant, to be cracked.

Although the cryogenic section comprises the cold box unit and the cryogenic exchanger cooler no.2,

right before it (feed stream of the cold box), for the sake of simplicity the designation of “cold box

unit” stands for the whole cryogenic section throughout most text.

26

Figure 9 Schematic overview of the cryogenic section (including the cold box unit – limited by dashed line) based on Dow´s IP.21 process data.

27

4 Modelling of Dow’s cold box unit

Contents

4.1. Materials………………………………………………………………………………………28

4.2. Flowsheet ………...………………………………………………………………………….29

4.3. Assumptions ………………………………….…………………………………………….33

4.4. Equipments specifications………………………………………………………………..34

28

Chapter 4 describes how the model of Dow´s cold box unit was implemented referring the program

tools and the process data used to design the flowsheet diagram. This chapter also presents the

assumptions made throughout the simulation and describes in detail the cold box equipments.

4.1. Materials

The model for Dow´s cold box unit was developed by using the Aspen Tech software. The acronym

“ASPEN” stands for “Advanced System for Process Engineering”. Aspen is a process simulation

software widely used in industry and known as an important tool for chemical engineers, thus it is

frequently used to perform simulations of chemical processes. Given a process design and an

appropriate selection of thermodynamics models, Aspen uses mathematical models to predict the

process performance. Furthermore, this software performs mass, energy and economic analyses.

Among the wide variety of Aspen packages, Aspen Tech Process Modelling V8.0 - Aspen Plus®

V8.0 was chosen to perform the current simulations since it is a steady state process simulator and

the most adequate for the level of complexity of this specific process [51].

The model was named “Aro.2 cold box simulation”, and some tasks had to be carried out to build

the process simulation of Dow´s cold box unit, namely:

Selection of the property method;

Selection of the convergence tolerance and flash iterations;

Identification of the relevant components involved in the process;

Design of the flowsheet diagram;

4.1.1. Property method

Aspen Plus has a wide variety of property methods, which are selected based on the desired

process. The property method was properly selected. The simulations were conducted in DowLHC

template and used DowNGO as the default property method, recommended for Dow´s cold box and

based upon NRTL modifications of the PSRK-Predictive Redlich-Kwong-Soave equation of state.

Only two phases (vapour and liquid) were considered in the phase equilibrium calculations.

29

4.1.2. Convergence tolerance and flash iterations

A convergence tolerance of 10-7 was established for flash convergence and maximum iterations of

75 in order to have a safe calculation and to ensure that the program succeeded.

4.1.3. Components list

The DowLHC template provided a wide list of components. However, these simulations only took

into account the following: hydrogen (H2); methane (CH4); ethane (C2H6); ethylene (C2H4); hydrogen

sulfide (H2S); toluene (C7H8) and benzene (C6H6).

4.2. Flowsheet

A bunch of process information was required to perform realistic simulations of Dow´s cold box unit

by using the Aspen Plus software. However, some desired information was not immediately obtained

since there was a wide lack of available information. In order to start building the process flow

diagram of the cold box, some input parameters, such as the temperature, pressure, and

composition of the cold box feed stream, were required by the software. To get the maximum

relevant plant process data, it was essential to consult the tool Aspen Tag Browser, which contained

the majority of the daily plant process data, named as InfoPlus.21 (IP.21) [52]. Consulting the Aspen

Tag Browser, the temperature of 10°C and the pressure of 53 bar were specified for the feed gas

stream.

On the other hand, the process data for the current composition of the cold box feed stream was not

available since the process did not have a composition analyser in this zone neither inside the cold

box. The only available record of this composition was found in documents from 1980, which were

not certainly updated and for that reason it should not be used as the current process data, though a

comparison would be interesting, latter on.

Therefore, some samples were taken from the cold box outlet streams, as the inside of the box was

not physically accessible for this kind of operations. The samples were taken under a stable

operation while the process was running smoothly, between 10 a.m. and 12 p.m. on November 19,

2014. The composition of each outlet stream was analysed and the data are listed in Table 2.

30

Table 2 Composition of Dow´s cold box outlet streams (sampling on November19, 2014).

Hydrogen outlet stream Methane outlet stream Ethane outlet stream

C (vol%) C(wt%) C (vol%) C (wt%) C (vol%) C(wt%)

H2 90,6 53,5 8,42 1,05 0,04 <0,01

CH4 9,09 42,6 82,0 81,1 7,14 3,74

C2H6 0,01 0,09 9,32 17,3 89,8 88,2

Others 0,29 3,75 0,24 0,63 3,02 8,02

Total 100 100 100 100 100 100

The composition required for the simulations input was the composition of the feed stream of the

cold box, which was obviously related with the outlet streams data, shown on Table 2.

The total mass flow rate of each outlet stream was well known (IP.21 database). The data were

selected based on the day and time the samples were taken. Table 3 presents the mass flow rate of

each outlet stream of Dow´s cold box.

Table 3 Mass flow rates of the cold box outlet streams (t/h).

Qm (t/h) Hydrogen outlet stream Methane outlet stream Ethane outlet stream Total

H2 3,2 0,16 ≈ 0 3,3

CH4 2,5 12,7 0,06 15,3

C2H6 ≈0 2,7 1,4 4,1

Others 0,2 0,1 0,1 0,4

Total 5,9 15,6 1,6 23

As a reminder, these simulations run on steady state, thus assuming that there were no mass losses

in the process, i.e., the mass flow rate of the fluid entering in the cold box was equal to the total

mass flow rate of the fluids going out. Hence, the hydrogen, methane and ethane mass flow rates

(and respective compositions) of the cold box feed stream were calculated through their partial mass

balances, resulting in Table 4.

31

Although there was a wide variety of components present in the process, the amount of hydrogen

(H2), methane (CH4) and ethane (C2H6) represented more than 99% (w/w) of the total mass flow rate,

meaning that the concentrations of the others compounds were negligible. Therefore, a simplification

was assumed for the simulations, viz. that the cold box feed gas stream only contained the main

compounds in Table 4.

Table 4 Composition and mass flow rates of the cold box feed streams.

Cold box feed stream

Qm (t/h) C (wt%) C (vol%)

H2 3,3 14,6 60,2

CH4 15,3 67,2 34,8

C2H6 4,1 18,2 5

Total 22,7 100 100

As the main input parameters were obtained, Aspen was then able to start building the cold box

flowsheet with all the streams and blocks, which are individually described in section 4.4. The

proposed overall process is presented in Figure 10.

The inlets specifications set throughout the simulations, were based on the process operating

conditions by the sampling time.

32

Figure 10 Aspen modelling - flowsheet of the cold box unit.

33

4.3. Assumptions

Due to the lack of available data to build a robust model that consisted of a good approach of the

real cold box unit, several parameters had to be specified and an extensive list of assumptions had

to be made to run the simulations and taken into account in the process analysis.

4.3.1. Simulations

The following hypothesis were assumed:

" 100% capacity for the cold box unit;

" Only vapour and liquid phases in the equilibrium calculations;

" Tolerance error of 10-4 for the results of the sensitivity analyses. Whenever a sensitivity

analysis is performed, the results may be slightly different from the standard single

simulation because the tolerance error is different. The sensitivity analyses do not use a

robust mathematical method to perform the calculations order to yield fast results.

4.3.2. Control system

Control valves, involving not only the equipments level control but also the temperature and pressure

regulation, were responsible for the dynamic process control of the cold box unit.

The model was developed under steady state based only on mass and energy balances, i.e., there

was no accumulation in the system, thus it was assumed that all equipments were under an ideal

level control. In order to build a robust model, it was preferable to minimize the number of blocks in

the flowsheet. Moreover, many operation parameters and entrance variables were specified in most

blocks. Therefore, the implementation of a control system was dispensable in this model.

34

4.4. Equipments specifications

As previously described in Chapter 3, the cold box unit contains a package of equipments used for

the hydrocarbons separation. Due to the lack of available data of the equipments characteristics, it

was mandatory to do some simplifications and set some specifications required by the Aspen

software to accomplish a realistic approach of the cold box unit.

Table 5 summarizes the Aspen library used to build the cold box flowsheet. Each block is described

in section 4.4.1 emphasizing the assumed specifications and simplifications.

Name Quantity Type Description Specifications

Cooler no.2 1 MHeatX Multi stream heat

exchanger

ΔPmain stream = 0,3486 bar ; ΔPH2 stream = 0,3486

bar ; ΔPCH4 stream = 0,16319 bar ; ΔPC2H6 stream=

0,16319 bar ; Tmain stream= -60,5°C; TCH4 stream = -

1,75°C; TC2H6 stream = 4,98°C

Deep cooler no.1 1 MHeatX Multi stream heat

exchanger


bar ; ΔPCH4 stream = 0,16319 bar ; ΔPC2H6 stream=

0,16319 bar ; Tmain stream= -91,05°C;

TCH4 stream = -65°C ; VfC2H6 stream = 0,628

Deep cooler no.2 1 MHeatX Multi stream heat

exchanger


bar ; ΔPCH4 stream = 0,16319 bar ; Tmain stream = -

140°C (Design Spec)

Ethylene chiller 1 Heater Thermal and phase

state changer

ΔPmain stream = 0,3486 bar; Tout= -95°C (Design

Spec)

Reboiler 1 Heater Thermal and phase

state changer ΔP = 0 bar ; Q = -88,21 kW

Methane stripper 1 RadFrac Rigorous 2-‐phase fractionation for single columns

No trays = 5 ;PTop= 5,92 barg; PBottom= 5,72 barg

Flash separator 3 Flash2 Two outlet flash ΔP = 0 bar; Q = 0 kW

Pump 1 Pump Pump Pout = 7 bar

Valve 1 Valve Valve Pout = 4,6 barg

Mixer 3 Mixer Stream mixer ΔP = 0 bar

Splitter 1 FSPLIT Stream splitter ΔP = 0 bar ; Limit flow rate of S9 = 1136,1 kg/h

Table 5 Blocks characterization according to Aspen library.

35

4.4.1. Cryogenic heat exchangers

General

Figure 11 presents the cryogenic heat exchangers of these simulations.

" These simulations comprised simple energy and mass balances to the heat exchangers, not

taking into account their geometries neither design parameters;

" For each heat exchanger, all the streams calculations converged simultaneously inside a single

block instead of converging sequentially in distinct blocks, thus extra assumptions on the heat

duty distribution were not required.

Cooler no.2:

" The IP.21 plant database allowed the temperature specification of the cold box outlet streams.

Nevertheless, the temperature of the hydrogen outlet stream was considered the only process

variable since it was the only process stream that was not significantly affected by variations of

the feed flow rate of the cold box.

Deep cooler no.1:

" In order to achieve an accurate mass balance, a vapour fraction of 0,63 was assumed for the

ethane outlet stream;

" A vapour fraction of 0,95 was assumed for the hot stream based on IP.21 process database;

" The temperature of the hydrogen outlet stream was considered the sole process variable for the

same reason mentioned in the cooler no.2.

Figure 1

Figure 11 Multi stream heat exchanger.

36

Deep cooler no.2:

" In order to achieve an accurate mass balance, a vapour fraction of 0,7 in the hot stream was

assumed, leading to a hydrogen purity of 91,2% (v/v) as a simulation result.

4.4.2. Ethylene chiller

" As a simulation simplification, the ethylene chiller was envisaged as two blocks (Figure 12)

interacting with each other through a DesignSpec that was implemented for this equipment;

" To build a robust simulation it was mandatory to create a DesignSpec, that yielded a stable

response to variations of the ethylene chiller heat duty. The conditions of the ethylene outlet

stream were maintained by varying the ethylene feed flow rate. In sum, the DesignSpec operated

on block B1 by varying the ethylene feed flow rate until the required heat duty of block B2 was

reached.

" Regarding the ethylene supply, a realistic range of the plant feed flow rates from 0 up to 6000

kg/h was assumed.

" The ethylene liquid feed stream was assumed, to be at the boiling point, i.e., -73°C and 4,5 bar,

thus when it flowed through the heat exchanger, it immediately flashed close to the atmospheric

pressure and became a vapour stream at -101°C and 1,2 bar.

Figure 12 Aspen approach for the ethylene chiller, B1-heater for ethylene stream, B2-cooler for hot stream.

37

" A vapour fraction of 0,92 was specified in block B2. As there were no available data for the

temperature and vapour fraction of this stream the value of 0,92 was assumed to increase the

amount of condensate without a significant variation of the outlet temperature, which must be

lower than the refrigerant temperature and also close to the temperature of the ethylene outlet

stream.

4.4.3. Methane stripper

Figure 13 represents the methane stripper composed by a stripping column and the respective

reboiler.

Stripper

" Since there was no available data for the composition of the stripper outlet streams, 40% (w/w) of

CH4 that entered in the methane stripper was assumed to leave at the bottom of the equipment in

order to achieve an accurate mass balance for the cold box unit. The heat duty was then

calculated and fixed.

" The simulation of this block did not account for the design parameters of the column, but 5 trays

were specified based on plant process data. The pressure drop was set to 0,04 barg per tray.

" Based on available data, the pressure at the top of the stripper was specified as 5,9 barg. This

value should never exceed 6,4 barg, which is set a safety pressure recommended.

Figure 13 Methane stripper.

38

Reboiler

" A kettle reboiler was selected for the reboiler simulation;

" The reboiler consisted of a coil-type tube bundle, which guaranteed the liquid flow around the

tubes by means of installed spacers. To simplify the construction of the apparatus, the bundle

was installed in the stripper sump.

" Firstly, the heat duty of the methane stripper was calculated to get a realistic separation of

CH4/C2H6, and in turn to an accurate mass balance. The calculated heat duty was specified in the

reboiler, allowing to fix the inlet gas flow rate to get a realistic outlet temperature, which was set

as -104,8°C.

4.4.4. Separators

Figure 14 presents the flash separators of the cold box flowsheet.

" For the flash separations, that took place in vessels, adiabatic processes (Q=0) were assumed.

Obviously, any vapour/liquid separation has an efficiency inferior to 100%.

" No pressure drops were assumed through the separator vessels, as the pressure drops

throughout the cold box were counted for in the heat exchangers and methane stripper.

Figure 14 Flash separator.

39

4.4.5. Other equipments (Figure 15)

a) Due to the lack of available process data, the outlet pressure was set at 7 bar.

b) Right before the deep cooler no.1 there was a splitter that separated a fraction of the main

process stream towards the reboiler of the methane stripper. This stream had a control valve, the

opening percentage of which was constant, thus the reboiler heat duty was constant, as well.

c),d) No pressure drop neither heat transfer were assumed through the mixers and the splitter.

Figure 15 Other equipments: a) pump, b) valve, c) mixer, and d) splitter.

40


41

5 Simulations results

Contents

5.1. Model validation……………………………………………………………………………..42

5.2. Model robustness…………………………………………………………………………...51

42

Chapter 5 presents the model simulations results, which were validated by comparison with the IP.21

online process data. A robustness test was also performed to the developed model. The present

results analyses cover the composition of the cold box outlet streams and the temperature/pressure

distributions throughout the simulations. Moreover, the results of the ethylene chiller, methane

stripper, and deep cooler no.2 are also presented in this chapter.

5.1. Model validation

As aforementioned in Chapter 4, a model of the cold box unit was developed by the implementation

of its process diagram (flowsheet). To assess if this model was a reliable approach of the cold box

unit, a validation procedure was performed based on the current plant process data. The model

validation involved the analysis of the deviations between the IP.21 process database and the Aspen

results.

5.1.1. Cold box mass balances

To achieve a reliable approach of the cold box unit, the first step was trying to reach accurate mass

balances. As aforementioned in section 4.2, some samples were taken from the cold box outlet

streams, and their compositions were analysed. These data combined with IP.21 process data of the

total flow rate of each outlet stream, allowed to predict the cold box feed gas composition that was

inserted as a simulation input. Several parameters were duly adjusted to approach Aspen results to

the plant data as much as possible.

In order to validate the model, the Aspen results were compared to the plant data (Figures16-18) and

their deviations were calculated, as shown in Tables 6-8.

Besides, some process data of the cold box were found in documents from October 1980. Although

these data were not updated ever since, the compositions of the outlet streams were predicted by

using the composition ratios in 1980 (inlet/outlet), and then compared with the Aspen results. This

validation process was not so accurate as the direct comparison with IP.21 database whenever

available. Anyway, it was rather interesting to check that there were only significant variations in the

mass balance of the ethane outlet stream (Figure 18) since 1980 until the present.

43

Tables 6-8 correspond to Figures 16-18 and present the individual flow rates deviations between the

Aspen results and IP.21 process data of each outlet stream. They also show the individual

concentrations of H2, CH4 and C2H6 in each outlet stream, calculated by Aspen.

Figure 16 Comparison of the flow rates of the hydrogen outlet stream.

Figure 17 Comparison of the flow rates of the methane outlet stream.

Figure 18 Comparison of the flow rates of the ethane outlet stream.

Qm Total (t/h) Qm H2 (t/h) Qm CH4 (t/h) Qm C2H6 (t/h) 1980 5,30 3,05 2,25 0 IP21 5,68 3,16 2,51 0 Aspen 5,72 3,23 2,48 0

Qm Total (t/h) Qm H2 (t/h) Qm CH4 (t/h) Qm C2H6 (t/h) 1980 14,1 0,270 13,0 0,845 IP21 15,6 0,164 12,7 2,70 Aspen 15,4 0,088 12,7 2,57

Qm Total (t/h) Qm H2 (t/h) Qm CH4 (t/h) Qm C2H6 (t/h) 1980 3,28 0 0,017 3,27 IP21 1,48 0 0,060 1,42 Aspen 1,61 0 0,047 1,56

≈ ≈ ≈

≈ ≈ ≈

44

Table 6 Flow rates deviations between Aspen results and IP.21 database of the hydrogen outlet stream.

Hydrogen outlet stream IP.21 deviation (%) CAspen (wt %)

H2 2,41 56,6

CH4 1,37 43,4

C2H6 49,5 ≈ 0

Total flow rate ≈ 0 100

Table 7 Flow rates deviations between Aspen results and IP.21 database of the methane outlet stream.

Table 8 Flow rates deviations between Aspen results and IP.21 database of the ethane outlet stream.

Methane outlet stream IP.21 deviation (%) CAspen (wt %)

H2 46,5 ≈ 0

CH4 ≈ 0 82,8

C2H6 5,05 16,7

Total flow rate 1,07 100

Ethane outlet stream IP.21 deviation (%) CAspen (wt %)

H2 100 ≈ 0

CH4 21,5 2,94

C2H6 9,82 97,1

Total flow rate 8,55 100

45

This analysis shows that accurate mass balances were estimated by Aspen for the cold box unit

since most of the current deviations are inferior to 10%. However, a few deviations are much higher

than 10%. Anyway, the components showing huge deviations corresponded to minor fractions (< 5%

(w/w)) of the streams total flows, thus they did not have a significant impact on the mass balances. In

sum, the model was validated as it yielded a very good estimation of the cold box mass balances.

5.1.2. Temperature distributions

The analysis of the temperature distributions throughout the cold box unit was also a validation

criterion of the developed model. Therefore, a comparison was carried out between the Aspen

results and IP.21 process data concerning the temperature of each process stream. The feed

process stream was analysed (Figure 19) up to the temperature at which the hydrocarbons

separated in three different outlet streams, that were individually analysed, as well (Figures 20-22).

Please note, that not every temperature calculated by Aspen had a corresponding temperature in

IP.21 database for comparison purposes.

The temperature results of the feed process stream (Figure 19) shows a perfect match with the IP.21

process data, as expected, since this temperature was specified based on the IP.21 database. On

the other hand, the temperature results of the methane and ethane outlet streams (Figures 20-21

respectively) represent an excellent match with the current plant situation since the maximum

temperature difference between the Aspen results and IP.21 process data was solely about 13°C for

the methane outlet stream, and 7°C for the ethane outlet stream.

Finally, the temperature of the hydrogen outlet stream revealed the highest deviations, which

reached 20°C, at the cold box outlet (Figure 22). In fact, the temperature of the hydrogen outlet

stream was the only process variable, i.e., it represented the sole degree of freedom in the

simulation. The errors associated to the running simulation merged on that variable, thus implying a

variation of the hydrogen stream temperature. Moreover, the measurement instruments bear random

and systematic errors, which were reflected in the IP.21 database and in the operating conditions set

in Aspen from this database. Still, the temperature variation of 20°C was acceptable and the

temperature distributions were considered valid.

46

Figure 20 Temperature distribution of the methane process stream through the cold box.

Direction: inside → outside.

Figure 19 Temperature distribution of the main process stream through the cold box.

Direction: outside → inside.

-160

-140

-120

-100

-80

-60

-40

-20

0

Stre

am te

mpe

ratu

re (°

C)

Equipments path

IP21 Aspen

-120

-100

-80

-60

-40

-20

0

20 S

tream

tem

pera

ture

(°C

)

Equipments path

IP21 Aspen

Dryer (out)

Cooler no.2 (out)

Ethylene chiller (in)

Deep cooler no.1 (out)

Separator no.2 (out)


Methane stripper (out)


Cooler no.2 (in)

Cooler no.2 (out)

47

-120

-100

-80

-60

-40

-20

0

20

Stre

am te

mpe

ratu

re (°

C)

Equipments path

IP21 Aspen

-160

-140

-120

-100

-80

-60

-40

-20

0

20

Stre

am te

mpe

ratu

re (°

C)

Equipments path

IP21 Aspen

Figure 22 Temperature distribution of the hydrogen process stream through the cold box.


Figure 21 Temperature distribution of the ethane process stream through the cold box.


Methane stripper Deep cooler

no.1 (in)


Cooler no.2 (in)

Cooler no.2 (out)

Separator no.2 (out)



Cooler no.2 (out)

48

5.1.3. Pressure distributions

In the IP.21 database there was a wide lack of pressure data throughout the cold box unit. For that

reason, the calculation of the pressure drops of the main and hydrogen process streams, was based

on the only two available pressures in the IP.21 pressure data, 54,5 and 51,9 bar, respectively,

measured outside the cold box. Moreover, the calculation of the pressure drops of the methane

process streams was also based on the two available pressures in the IP.21 pressure data, 4,9 and

5,6 bar, measured outside and inside of the cold box, respectively. On the other hand, there was not

a single available pressure of the ethane process stream. Thus, this stream was assumed to have

absolute pressures and pressure drops similar to the corresponding ones in the methane process

stream.

As there was not enough pressure data in IP.21 database, the pressure distributions throughout the

cold box unit were plotted only on Aspen results, which did not reveal significant variations in each

process stream (Figure 23). Obviously, the pressure evolution of each process stream was

predictable because this variable was specified throughout the model simulations. The small

variations observed were due to the pressure drops occurring throughout the equipments of the cold

box.

Figure 23 Pressure distributions through the cold box unit.

0

10

20

30

40

50

60

Stre

am p

ress

ure

(bar

)

Equipments path

Main process stream

Hydrogen stream

Methane stream

Ethane stream

49


As aforementioned in Chapter 3, the cold box used ethylene as refrigerant. Currently, according to

the IP.21 database, liquid ethylene is supplied to the ethylene chiller at -73°C and 4,5 bar. This utility

stream was assumed to operate at the boiling point, thus when the ethylene entered in the chiller it

was totally vaporized at the atmospheric pressure by removing the heat from the hot process stream.

Since there was no available data for the ethylene flow rate, a reasonable value had to be assumed.

As the ethylene flowmeters presented a maximum of 6000 kg/h and there was a control valve to

regulate the ethylene flow rate into the ethylene chiller with a current opening of 61%, the ethylene

flow rate was estimated as 3693 kg/h, assuming that the control valve had a linear behaviour. Table

9 shows the Aspen results of ethylene chiller.

Table 9 Aspen results of the ethylene chiller.


Table 10 shows the Aspen results of the methane stripper. The temperature/pressure in the Aspen

results were relatively close to the IP.21 process data.

Inlet Outlet

Qm (kg/h) 3693 3693

Tprocess fluid (°C) -91 -100

Trefrigerant (°C) -73 -101

P (bar) 4,5 1,2

Vf process fluid 0,95 0,92

Vf refrigerant ≈ 0 ≈ 1

50

Table 10 Aspen results of the methane stripper.

Feed Bottom Top

Qm (kg/h) 4918 3762 1156

CH2 (wt%) 0,30 9,3x10-9 1,3

CCH4 (wt%) 38 20 95

CC2H6 (wt%) 62 80 4

TAspen (°C) -100 -108 -117

TIP.21 (°C) - -103 -110

P (bar) 53,5 6,7 6,9

Vf ≈ 0 ≈0 ≈1

Column

No. trays 5

Condenser Absent

Reboiler (kW) 88

5.1.6. Deep cooler no.2

As aforementioned in section 4.2, the deep cooler no.2 was the cryogenic heat exchanger

responsible for H2/CH4 separation and the outlet temperature must not be much lower than -136°C.

Furthermore, the calculated hydrogen purity should be about 90% (v/v). Table 11 shows the principal

Aspen results, which were much close to the IP.21 process data.

Table 11 Outlet results of the deep cooler no.2

Aspen result IP.21 process data

H2 (vol%) 91,2 90,6

Toutlet (°C) -150 -136

51

5.2. Model robustness

The model that represented the cold box unit was implemented and validated based on the plant

current data. As the developed model was built for steady state, it was crucial to perform some

simulation tests to confirm if it was robust enough to proceed with further variations of the feed

stream composition. Therefore, variations of the feed gas flow rate were made, whereas the feed

gas composition remained constant, to observe if the flow rate variations had a significant impact in

the cold box unit, especially in the outlet streams. In such a case, the model would not be robust

enough to further vary the composition of the feed gas stream.

The specification of the process main stream, leaving the heat exchanger deep cooler no.1, was

varied and a vapour fraction of 0,95 was set to guarantee the desired amount of condensate

irrespective to the flow rate of the feed gas stream. To analyse the behaviour of the cold box unit,

sensitivity analyses were performed by varying the total flow rate of the cold box inlet, from 0 up to

30 t/h, whilst its composition remained constant. The following sections focused within the range

where the developed model was valid, which was selected based on the most restrictive limitation of

the variables analysed throughout the simulation.

5.2.1. Compositions of the cold box outlet streams

The compositions of the outlet streams should not be affected by variations of the cold box feed flow

rate. As previously mentioned, a sensitivity analysis of the Aspen simulation was performed with

respect to the flow rates of the cold box outlet streams by varying the feed flow rate of the cold box

from 0 up to 30 t/h. The results of the sensitivity analysis revealed a satisfactory response, i.e., they

confirmed the expected linear behaviour of the flow rate of each outlet stream (Figure 24).

Furthermore, the model showed to be valid from 2,8 up to 30 t/h. Below this range the simulation

yielded convergence errors. It is worth remarking that a split fraction of the main process stream,

was set to 1,1 t/h and it was impossible to get reliable results for feed gas flow rates inferior to the

split flow rate.

52

Table 12 lists the flow rates of each outlet stream for three distinct total feed flow rates (5,11 and 22

t/h). As shown in Table 12, the methane outlet stream had an excessive amount of ethane, meaning

that most ethane in the system ended up in the methane outlet stream. Moreover, the hydrogen

outlet stream had a H2 purity of 56,6% (w/w) (91,2% (v/v)) and contained a substantial amount of

methane. In fact, the composition of the hydrogen outlet stream was specified in the heat exchanger

deep cooler no.2, which was responsible for condensing 30% (w/w) of the feed stream regardless of

the feed flow rate, thus the hydrogen purity was fixed in the hydrogen outlet stream.

Table 12 Results of the flow rates of each outlet stream for three distinct total feed gas flow rates.

Hydrogen

outlet stream

Methane

outlet stream

Ethane

outlet stream

Feed gas flow rate in cold box

5

t/h

11

t/h

22

t/h

5

t/h

11

t/h

22

t/h

5

t/h

11

t/h

22

t/h

Qm H2 (t/h) 0,71 1,6 3,1 0,02 0,04 0,08 4,8x10-16 7,1x10-15 2,1x10-33

Qm CH4(t/h) 0,55 1,2 2,4 2,8 6,2 12,3 8,4x10-3 8,4x10-3 0,04

Qm C2H6 (t/h) 5,7x10-4 1,3x10-3 2,3x10-3 0,71 1,4 2,5 0,2 0,6 1,5

0

4

8

12

16

20

24

0 2 4 6 8 10 12 14 16 18 20 22 24 26 28 30

Tota

l flo

w ra

tes

of th

e ou

tlet s

tream

s (t/

h)

Feed gas flow rate of the cold box (t/h)

Hydrogen stream Methane stream Ethane stream

Figure 24 Model robustness test concerning the flow rates of each cold box outlet stream.

53


The results of the temperature distribution of each outlet stream showed that the model yielded a

stable response to the variations of the feed flow rate over a wide range, as depicted in Figure 25. As

the temperatures of the methane and ethane outlet streams were specified, it was not expected that

variations of the feed gas flow rate affected significantly those temperatures. Nevertheless, the

hydrogen outlet stream revealed a temperature variation of 15°C, somewhat steep for feed flow rates

up to 15 t/h. This variation might have occurred because the temperature of the hydrogen outlet

stream was the sole process variable in the simulation. Anyway, the model yielded a stable response

concerning the temperatures distributions, viz. the hydrogen stream temperature was practically

constant around the plant feed flow rate of 22,7 t/h.

5.2.3. Pressure distribution

As for the pressure distribution of each outlet stream, the results of the sensitivity analysis pointed

out that there were no significant pressure variations when the feed gas flow rate varied (Figure 26).

This behaviour was expected since the process streams pressures were specified and remained

constant throughout the cold box simulations.

-‐25 -‐22 -‐19 -‐16 -‐13 -‐10 -‐7 -‐4 -‐1 2 5

0 2 4 6 8 10 12 14 16 18 20 22 24 26 28 30

Stre

am te

mpe

ratu

re (°

C)

Feed gas flow rate of the cold box (t/h).


Figure 25 Model robustness test concerning the temperature distribution of each cold box outlet stream.

54

0

1000

2000

3000

4000

5000

6000

0 2 4 6 8 10 12 14 16 18 20 22 24 26 28 30

Eth

ylen

e fl

ow ra

te (k

g/h)

Feed gas flow rate of the cold box (t/h)

Ethylene


The increase of the feed gas flow rate of the cold box had a striking impact on the refrigerant flow

rate of the ethylene chiller. As a vapour fraction of 0,92 in the ethylene chiller was specified, the

increase of the cold box feed flow rate, led to a greater need of refrigerant in the chiller to ensure the

vapour fraction specification.

Figure 27 reveals a satisfactory response of the model, viz. the expected rise of the ethylene flow

rate when the feed gas flow rate increased.

Figure 12 Cold Box – Pressure Profile for the outlet streams of the C.B. 0

8

16

24

32

40

48

56

0 2 4 6 8 10 12 14 16 18 20 22 24 26 28 30

Stre

am p

ress

ure

(bar

)

Feed gas flow rate of the cold box (t/h).


Figure 26 Model robustness test concerning the pressure distribution of each cold box outlet stream.

Figure 27 Model robustness test concerning the flow rate of refrigerant supplied to the ethylene chiller.

55


The robustness analysis of the methane stripper revealed a reliable behaviour of the outlet streams

leaving the bottom and the top of the equipment. The increase of the cold box feed flow rate implied

a higher condensation rate in the ethylene chiller, and in turn higher flow rates of the methane

stripper outlets.

The stripper feed stream mainly consisted in ethane and, the bottom flow rate, rich in ethane, was

higher than the top flow rate, rich in methane. Moreover, some parameters were specified in this

block, for example the pressure distribution through the column, the number of trays, and the

methane concentration at the bottom of the stripper was set as 40% (w/w) of the total methane

present in the feed stream.

The results also showed a small range of the cold box feed flow rates (2,8-5,1 t/h), in which the

stripper top flow rates were higher than the bottom flow rates, because the heat duty of the stripper

reboiler was fixed. In fact, if the cold box feed flow rate was in between 2,8 and 5,1 t/h, the heat duty

of the reboiler was extremely high, vaporising most of the stripper feed stream.

Figure 28 Model robustness test concerning the flow rates of the methane stripper.

0

1

2

3

4

5

6

7

0 3 6 9 12 15 18 21 24 27 30

Tota

l flo

w ra

tes

of th

e m

etha

ne s

tripp

er (t

/h)

Feed gas flow rates of the cold box (t/h)

Bottom rich in ethane Top rich in methane Feed

56


57

6 Case studies

Contents

6.1. Case Study 1…………………………………………………………………………………59

6.2. Case Study 2…………………………………………………………………………………62 6.3. Case Study 3…………………………………………………………………………………67

58

Chapter 6 presents the analysis of three distinct plant scenarios by using the Aspen model that was

built and validated along Chapters 4 and 5, respectively. The first scenario was the Case Study 1,

corresponding to the current plant situation, in which the whole Aromatics 2 unit was working

normally, including the HDA reactor.

As previously introduced in Chapter 1, the main aim of this study was to confirm, whenever the HDA

reactor (Aro.2) was stopped, if it was viable to feed the hydroganation reactor (Aro.1) with high purity

hydrogen coming from the cold box unit (Aro.2). In such a case, the methanator outlet stream

(Aro.2), only containing CH4 and H2, was transported to the cold box unit through the existent piping

and equipments without any chemical reaction or physical process since these equipments were

stopped, as well. This situation is further described and analysed in Case Study 2 in this chapter.

Actually, no one knew if it was possible to forward the methanator outlet stream to the cold box unit

without operating the washing column, one of the equipments where the flow should pass through

right before reaching the cold box unit. For that reason this study also analysed the Case Study 3,

which was similar to the plant scenario of Case Study 2 except that in this new scenario the washing

column was normally working.

To simplify the simulations, since the washing column solvent only contained xylenes traces, the cold

box model never considered these compounds in any of the case studies.

It is worthwhile mentioning that all these case studies implied variations of the composition of the

cold box feed stream in the Aspen model implemented. As previously observed in Chapter 5, the

developed model was robust enough to undertake variations of the cold box feed flow rate and still

yield reliable responses. Before proceeding with the variations of the composition of the cold box

feed stream some parameters were set in the previous model to get reliable results. A new model

was used to conduct composition variations and analyse their impact on the cold box outlet streams.

Table 13 summarizes the parameters of the original model that were replaced leading to a new

model.

Table 13 Different parameters of the two simulation models.

Equipments outlet hot stream Original model New model

Ethylene chiller

Vapour fraction of 0,92 Temperature of -95°C

Deep cooler no.2 Vapour fraction of 0,7

DesignSpec Variable: T (-200 to 0°C)

Target: H2 purity of 90% (v/v)*

Deep cooler no.1 Vapour fraction of 0,95 Temperature of -91,1°C

*Minimum acceptable target, defined by the user.

59

6.1. Case Study 1

The Case Study 1 corresponded to the current plant situation, in which the HDA reactor was

normally working, as well as every equipment in Aromatics 2 facility (Figure 29).

So far, the developed Aspen model had just considered the main components, CH4/H2/C2H6.

However, to get a more realistic response from the model and predict how the plant currently works,

it was mandatory to consider a small flow rate of heavies4, benzene (C6H6) and toluene (C7H8), in the

cold box feed stream. It was known that practically these heavies should end up in the ethane outlet

stream, reinforcing the motivation for this case study, i.e., to verify if this fact was actually predicted

by the model. Moreover, a small flow rate of hydrogen sulfide (H2S) was also considered in the cold

box feed stream to prevent coking in the reactor system.

Using the new model (Table 13), a single simulation was carried out by considering heavies in the

cold box feed stream, the total feed flow rate of which remained constant at 22,7 t/h. Appendix A1

presents the Aspen results of the single simulation relative to Case Study 1.

4 The word “heavies” stands for high molecular weight compounds, e.g., benzene and toluene.

Figure 29 Schematic overview of Case Study 2 (Aro.2 facility): green blocks working.

LHC I/II: ethylene plants; DC-MTH: methanator; GB1,2: compressors; DA-WASH: washing column; DC-

HDA: hydrodealkylation reactor; CB: cold box unit.

60


The composition of the cold box feed stream was slightly different from the one set in Chapter 4.

Table 14 lists the feed flow rates, weight and volume percentages compositions assumed in Case

Study 1.

Table 14 Case Study 1 – Feed flow rates and compositions of the cold box.

H2 CH4 C2H6 C6H6 C7H8 H2S Total

Qm (t/h) 3,3 15,2 4,12 0,023 0,023 2,27x10-3 22,7

C (wt%) 14,6 67,1 18,1 0,10 0,10 0,01 100

C (vol%) 60,2 34,8 5,0 0,01 0,01 0,002 100

The easiest way to check how the heavies evolved is by observing the composition of the cold box

outlet streams (Table 15). Assuming those feed conditions, it was observed that most heavies ended

up in the ethane outlet stream indeed, as expected.

H2 CH4 C2H6 C6H6 C7H8 H2S Total

Hydrogen outlet

stream

T = -2,8°C

P = 52,3 bar

C (vol%) 90 10 9,16x10-3 7,73 x10-15 5,39 x10-17 1,75x10-7 100

C (wt%) 53 46,8 0,08 1,76 x10-13 1,45 x10-15 1,74 x10-6 100

Qm (kg/h) 3228 2854 4,9 1,07 x10-11 8,85 x10-14 1,06 x10-4 6087

Qm (t/h) 3,2 2,8 4,9 x10-3 1,07 x10-14 8,85 x10-17 1,06 x10-7 6

Methane outlet

stream

T = -1,75°C

P = 5,12 bar

C (vol%) 4,65 84,4 10,9 2,31x10-5 2,88x10-6 2,88x10-3 100

C (wt%) 0,55 79,9 19,5 1,07x10-4 2,25x10-5 5,80x10-3 100

Qm (kg/h) 85,6 12368 3011 0,02 3,47x10-3 0,89 15466

Qm (t/h) 0,1 12 3 1,16x10-5 3,47x10-6 8,97x10-4 15

Ethane outlet

stream

T = 4,9°C

P = 6,8 bar

C (vol%) 8,42 x10-11 2,8 95,6 0,75 0,64 0,11 100

C (wt%) 5,58 x10-12 1,5 94,5 1,9 1,9 0,12 100

Qm (kg/h) 6,51 x10-11 17,5 1104 22,7 22,7 1,4 1168

Qm (t/h) 6,51 x10-14 1,75 x10-2 1,1 0,023 0,023 1,4 x10-3 1,16

Table 15 Case Study 1 – Aspen results for the flow rates, compositions and operating conditions of the cold box outlet streams.

61


The Aspen results concerning the operating conditions and the refrigerant flow rates in the ethylene

chiller are shown in Table 16.

Table 16 Case Study 1 – Aspen results for the operating conditions and flow rates of the refrigerant in the

ethylene chiller.


Table 17 presents the operating conditions, flow rates and compositions of the methane stripper

streams. The temperatures and pressures were in agreement with IP.21 database.

Table 17 Case Study 1 – Aspen results for the operating conditions, flow rates, and compositions of the

methane stripper streams.

Feed Bottom Top

Qm (kg/h) 3790 2851 939

CH2 (wt%) 0,27 7,2x10-10 1,08

CCH4 (wt%) 31,5 11,2 93,2

CC2H6 (wt%) 66,9 87,1 5,77

CH2S (wt%) 0,05 0,07 6,2x10-4

CC7H8 (wt%) 0,59 0,79 2,8x10-8

CC6H6 (wt%) 0,59 0,79 3,3x10-7

T (°C) -95 -94 -112

P (bar) 53,5 6,73 6,93

Vf ≈ 0 ≈0 ≈1

Column

No. trays 5

Condenser Absent

Reboiler (kW) 88

Inlet Outlet

Qm (kg/h) 1299 1299

T (°C) -73 -101

P (bar) 4,5 1,2

Vf ≈ 0 ≈ 1

62

6.2. Case Study 2

As already referred at the beginning of this chapter, the Case Study 2 represented the main goal of

this project. Unlike Case 1, in this plant scenario the HDA reactor was stopped. This case study is

sketched in Figure 30, displaying the methanator outlet stream, which was compressed and

transported to the cold box unit through the existent piping and equipments without suffering any

chemical reaction or physical process since these equipments were also stopped. The high purity

hydrogen that left the cold box unit was subsequently fed to the hydrogenation reactor located in the

Aromatics 1 facility. As sketched in Figure 30, the green blocks were working normally, whilst the

grey blocks were stopped.

Using the new Aspen model, the plant scenario of Case 2 technically meant to vary the composition

in the cold box feed stream. To observe the impact of the feed composition variation in the cold box

performance, several sensitivity analyses were performed, in which H2 purity in the cold box feed gas

stream varied from 0 up to 100% (v/v) whereas the total flow rate remained constant at 22,7 t/h. After

running the simulations, the sensitivity analyses revealed that the model representing the plant

scenario of Case 2, was only valid from 22,4 up to 90% (v/v) of H2 in the cold box feed gas.

Appendix A2 displays the Aspen results of the sensitivity analyses relative to Case Study 2.

Figure 30 Schematic overview of Case Study 2 (Aro.2 facility): grey blocks stopped and green blocks working.

LHC I/II: ethylene plants; DC-MTH: methanator; GB1,2: compressors; DA-WASH: washing column; DC-HDA:

hydrodealkylation reactor; CB: cold box unit.

63


The Aspen results revealed that from 0 up to 10% (v/v) of H2 in the cold box feed stream, the

separation between H2 and CH4 did not occur because the model was dealing with supercritical

fluids. Besides, the model faced convergence errors from 10 up to 22,4% (v/v), thus it did not yield

reliable results in this range. The results of the sensitivity analysis are plotted in Figure 31 that

identifies the validity range of the model and displays the composition distributions of the cold box

outlet streams.

Figure 31 shows that the ethane outlet stream was absent (null flow rate), as expected since C2H6

was not taken into account in this case study, because the HDA reactor was stopped. The methane

outlet stream revealed a realistic composition distribution, starting with a great flow rate of CH4,

which gradually reduced as the flow rate of H2 in the cold box feed stream increased. Last, but not

least, Figure 31 presents a predictable response of the total flow rate of the hydrogen outlet stream.

The latter gradually rose as H2 purity increased in the cold box feed stream up to 90% (v/v). Beyond

this value the total flow rate of the feed stream exceeded 22,7 t/h and led to convergence errors,

meaning that the results were unrealistic above 90% (v/v) of H2 in the cold box feed stream.

Table 18 lists the flow rates of the hydrogen and methane outlet streams along the range of H2 purity

in the cold box feed stream that might be applicable in the plant (60-75% (v/v) of H2). As previously

noticed in Chapter 5, the hydrogen outlet stream contained an significant amount of CH4 in Case

Study 2.

Figure 31 Case Study 2 - sensitivity analysis concerning the compositions of the cold box outlet

streams.

0

4

8

12

16

20

24

0 10 20 30 40 50 60 70 80 90 100

Tota

l flo

w ra

tes

of th

e ou

tlet s

tream

s (t/

h)

Purity of H2% (v/v) in the feed gas stream of the cold box


64

Table 18 Aspen results for the flow rates of the cold box outlet streams.

The equipment mainly responsible for the augment of H2 purity was the deep cooler no.2, in which a

DesignSpec varied the outlet temperature from -200 to 0°C to condense methane till a H2 purity of

90% (v/v) was achieved. Figure 32, shows a constant H2 purity of 90% (v/v) in the hydrogen outlet

stream over the model validity range. If the H2 purity in the cold box feed stream would exceed 90%

(v/v), the deep cooler no.2 would presumably condense H2 instead of CH4, but as this process would

not be feasible at the operating conditions specified the CH4/H2 separation would not occur after all,

i.e., everything that would enter into the cold box would end up in the hydrogen outlet stream.

Hydrogen outlet stream Methane outlet stream

CH2 feed gas (vol%) 60 65 70 75 60 65 70 75

Qm H2 (t/h) 3,6 4,2 5,1 6,1 0,1 0,1 0,1 0,1

Qm CH4 (t/h) 3,2 3,7 4,5 5,4 16 14,9 13,2 11,3

0

10

20

30

40

50

60

70

80

90

100

0 10 20 30 40 50 60 70 80 90 100

H2 p

urity

in th

e hy

drog

en o

utle

t stre

am

(%

(v/v

)


Figure 32 Case Study 2 – sensitivity analysis concerning the H2 composition of the hydrogen outlet stream.

65


Figure 33 displays a constant response for the temperature of the methane outlet stream, since this

variable was specified in the model. Conversely, the temperature of the hydrogen outlet stream

showed a rather steep variation of 100°C because it was the sole process variable in the system.

Anyway, in the range of H2 purity in the cold box feed stream practically applicable (60-75% (v/v)),

the temperature solely presented a smooth variation, meaning that it would be safe working within

this range without a significant impact on the temperature of the hydrogen outlet stream.

6.2.3. Pressure distributions

The results for the pressure distribution of each outlet stream (Figure 34) revealed that there were no

substancial pressure variations when the purity of H2 in the feed gas stream of the cold box varied

since the outlet streams pressures were specified throughout the cold box unit. The ethane outlet

stream was not plotted because its flow rate was null, as previously mentioned in section 6.2.1.

-‐110

-‐95

-‐80

-‐65

-‐50

-‐35

-‐20

-‐5

10

0 10 20 30 40 50 60 70 80 90 100

Stre

am te

mpe

ratu

re (°

C )


Hydrogen Stream Methane Stream

Figure 33 Case Study 2 – sensitivity analysis concerning the temperature of the cold box outlet streams.

66


As Case Study 2 did not consider ethane neither heavies in the cold box feed gas, the flow rate of

ethylene refrigerant to be fed to the cold box was lower than in Case Study 1. The results depicted

in, (Figure 35) showed that the ethylene chiller partially condensed the feed gas, only containing H2

and CH4, in a small range from 22,4 to 24,6% (v/v) of H2 in the cold box feed stream. On the other

hand, for hydrogen purities beyond 24,6% (v/v), the ethylene chiller was unable to proceed with the

condensation of CH4 and just a cooling process took place.

0

10

20

30

40

50

60

0 10 20 30 40 50 60 70 80 90 100

Stream

pressure (bar)


Hydrogen Stream Methane Stream

Figure 34 Case Study 2 - sensitivity analysis concerning the pressure of the cold box outlet streams.

Figure 35 Case Study 2 - sensitivity analysis concerning the refrigerant flow rates in the cold box unit.

Partial

condensation

Cooling

0

0,5

1

1,5

2

2,5

3

0 10 20 30 40 50 60 70 80 90 100

Eth

ylen

e flo

w ra

te (

t/h)


Partial

condensation

Cooling

67


The results of the methane stripper were in close agreement with the ethylene chiller results, i.e.,

since there was no ethane in the cold box feed gas, the stripper feed flow rate was null (Figure 36).

Yet, small flow rates through the methane stripper were noticed in a small range from 22,4 up to

24,6% (v/v) of H2 in the cold box feed stream, corresponding to a minor condensation of CH4 in the

ethylene chiller, as previously observed in section 6.2.4.

6.3. Case Study 3

The Case Study 3, which was similar to Case Study 2, was finally studied. Actually, no one knew if it

was possible to forward the methanator outlet stream to the cold box unit without operating normally

the washing column. Thus, the Case Study 3 considered this equipment in the flowsheet as shown in

Figure 37. Obviously, the compositions of the cold box feed stream were different in the Case

Studies 2 and 3, because in the latter the washing column operated with a BTX solvent, whereas in

the former this equipment was stopped.

0

1

2

3

0 10 20 30 40 50 60 70 80 90 100

Methane

strip

per fl

ow ra

tes (t/h)


Top BoWom Feed

Figure 36 Case Study 2 - sensitivity analysis concerning the flow rates of the methane stripper.

68

In the Case Study 3 a single simulation was run. Case Study 3 was analogous to Case Study 1 since

both aimed to ascertain in which outlet stream the heavies from the washing solvent ended up, but it

did not consider ethane neither hydrogen sulfide in the cold box feed stream.


The composition of the cold box feed stream in Case Study 3 was distinct from the one in Case

Study 2, as the washing column operated with a BTX solvent. Table 19 presents the composition of

the cold box feed stream, considering identical flow rates of heavies to the ones assumed in the

Case Study 1.

Table 19 Case Study 3 – Feed flow rates and compositions of the cold box.

H2 CH4 C6H6 C7H8 Total

Qm (t/h) 4 18,6 0,027 0,027 22,7

C (wt%) 17,8 81,9 0,125 0,125 100

C (vol%) 63,3 36,7 0,01 9,5x10-3 100

Figure 37 Schematic overview of Case Study 3 (Aro.2 facility): grey blocks stopped and green blocks working.

LHC I/II: ethylene plants; DC-MTH: methanator; GB1,2: compressors; DA-WASH: washing column; DC-HDA:

hydrodealkylation reactor; CB: cold box unit.

69

Table 20 displays the Aspen results for the flow rates and compositions of the cold box outlet

streams. Unlike Case Study 1, the heavies ended up in the methane outlet stream, meaning that

they passed through the heat exchanger deep cooler no.2. This equipment worked at the lowest

temperature of the cold box unit, hence plugging was likely to occur under those hrash conditions.

Table 20 Case Study 3 – Aspen results for the flow rates, compositions and operating conditions of the cold box

outlet streams.

H2 CH4 C6H6 C7H8 Total

Hydrogen outlet stream

T = 0,8°C

P = 52,2 bar

C (vol%) 90 10 3,1x10-13 1,7x10-15 100

C (wt%) 53,1 46,9 7,1x10-12 4,8x10-14 100

Qm (kg/h) 3942 3481 5,2x10-10 3,5x10-12 7423

Qm (t/h) 3,9 3,5 5,2x10-13 3,5x10-15 7

Methane outlet stream

T = -1,7°C

P = 5,1 bar

C (vol%) 5,28 95 0,03 0,03 100

C (wt%) 0,69 98,9 0,18 0,18 100

Qm (kg/h) 106,2 15134 27,7 27,7 15298

Qm (t/h) 0,1 15 0,03 0,03 15

Ethane outlet stream

T = 4,9°C

P = 6,8 bar

C (vol%) 3,5x10-14 4,5x10-10 4,29 95,7 100

C (wt%) 7,7x10-16 7,9x10-11 3,66 96,3 100

Qm (kg/h) 2,0x10-21 2,1x10-16 9,7x10-6 2,5x10-4 2,6x10-4

Qm (t/h) 2,0x10-24 2,1x10-19 9,7x10-9 2,5x10-7 2,6x10-7

70


The Aspen results for the operating conditions of the ethylene chiller and the refrigerant flow rates

are shown in Table 21.

Table 21 Case Study 3 – Aspen results for the operating conditions and flow rates of the refrigerant in the

ethylene chiller.


The Aspen results for the operating conditions, flow rates and compositions of the methane stripper

streamsare shown in Table 22. It is worth to highlight the huge unrealistic temperatures of the

stripper, maybe due to severe restrictions in the simulation.

Table 22 Case Study 3 – Aspen results for the operating conditions, flow rates and compositions of the methane

stripper streams.

Feed Bottom Top

Qm (kg/h) 57,3 7,1x10-4 57,3

CH2 (wt%) 9,3x10-3 7,6x10-4 9,4x10-3

CCH4 (wt%) 3,3 3,1x10-9 3,3

CC7H8 (wt%) 48,4 94,7 48,4

CC6H6 (wt%) 48,2 5,3 48,2

T (°C) -95 191 169

P (bar) 53,5 6,7 6,9

Vf ≈ 0 ≈0 ≈1

Column

No. trays 5

Condenser (kw) Absent

Reboiler (kw) Error

Inlet Outlet

Qm (kg/h) 661 661

T (°C) -73,41 -101,19

P (bar) 4,51 1,18

Vf ≈ 0 ≈ 1

71

7 Conclusions and future work

Contents

7.1. Conclusions………………………………………………………………………………….72

7.2. Future work…………………………………………………………………………………..73

72

7.1. Conclusions

This work investigated the hydrogen purification in the cold box (Aromatics-2 facility) at The Dow

Chemical Company’s plant in The Netherlands, with the major aim of enhancing the benzene

production in the hydrogenation reactor (Aromatics-1 facility), whenever the hydrodealkylation

reactor was shut down.

The development of a cold box model, by using the Aspen plus 8.0 software, that predicted well the

current plant situation succeeded, yielding mass balances deviations below 10% with respect to the

IP.21 process data. The model robustness was tested by varying the cold box feed flow rates. The

sensitivity analyses revealed reliable responses of the model for the compositions of the outlet

streams, which did not significantly vary. Another conclusion drawn was that the hydrogen purity was

raised from 60 to 91,2% (v/v).

Although the developed model was already a good representation of the cold box unit, it was not

robust to vary the feed gas composition and still yield a stable response. Therefore, a similar model

was implemented to analyse three distinct plant scenarios. For all these scenarios, the H2 purity was

set as 90% (v/v) in the hydrogen outlet stream for a cold box feed flow rate of 22,7 t/h.

The results of the Case Study 1, which considered the current plant scenario, confirmed that the

heavies ended up in the ethane outlet stream. Nevertheless, the results also revealed an excessive

flow rate of ethane in the methane outlet stream.

In the Case Study 2, the simulation showed a reliable performance of the cold box unit, i.e.,

whenever the HDA reactor was stopped, the methanator outlet stream (60% (v/v)), only containing

H2 and CH4, was forwarded to the cold box unit to achieve the desired H2 purity of 90% (v/v), and the

rise of benzene production in 35%. The methane stripper was dispensable in this scenario.

The Case Study 3 had the same purpose as the previous one, but it considered the washing column

working normally. The heavies ended up in the methane outlet stream, instead of ending up in the

ethane outlet stream as in the current plant (Case Study 1). This situation might lead to a plugging

drawback. Moreover, the methane stripper presented unreliable temperatures of its outlet streams.

Briefly, this project bears a high potential for benzene production of The Dow Chemical Company.

73

7.2. Future work

The developed models revealed distinct validity ranges due to their hrash restrictions. To study the

cold box performance beyond the models validity ranges, it is suggested to test the model with

different approaches, such as: varying the split fraction that flows into the reboiler; varying the H2

purity goal; inserting design/geometry parameters in the equipments; testing distinct methods for the

components physical properties.

In order to proceed with this study and test this project aim in the field, it is recommended to find out

the cold box start-up operating conditions, only feeding H2 and CH4 to ascertain if it is possible

running the cold box unit while the washing column is stopped. Furthermore, it is essential to

investigate the likelihood of plugging in Case Study 3. Afterwards, it is crucial to perform a simulation

before forwarding the hydrogen outlet stream to the hydrogenation reactor (Aro.1).

At last, for the situation when the HDA reactor is working normally, it is suggested to reduce the

ethane purity in the methane outlet stream of the cold box unit.

74


75

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79

Appendixes Appendix A1 – Case Study 1

S1 S2 S3 S4 S5 S6 S7 S8 S9 Substream: MIXED Mole Flow (kmol/h) H2 1,64E+03 1,64E+03 1,64E+03 1,60E+03 0,00E+00 4,25E+01 3,23E-11 1,56E+03 8,22E+01 CH4 9,50E+02 9,50E+02 9,50E+02 1,78E+02 0,00E+00 7,71E+02 1,09E+00 9,02E+02 4,75E+01 C2H4 0,00 0,00 0,00 0,00 46,34 0,00 0,00 0,00 0,00 C2H6 1,37E+02 1,37E+02 1,37E+02 1,63E-01 0,00E+00 1,00E+02 3,67E+01 1,30E+02 6,85E+00 H2S 6,67E-02 6,67E-02 6,67E-02 3,11E-06 0,00E+00 2,63E-02 4,03E-02 6,33E-02 3,33E-03 C6H6 2,91E-01 2,91E-01 2,91E-01 1,37E-13 0,00E+00 2,11E-04 2,91E-01 2,76E-01 1,45E-02 C7H8 2,47E-01 2,47E-01 2,47E-01 9,60E-16 0,00E+00 3,77E-05 2,47E-01 2,34E-01 1,23E-02

Mole Frac H2 6,02E-01 6,02E-01 6,02E-01 9,00E-01 0,00E+00 4,65E-02 8,42E-13 6,02E-01 6,02E-01

CH4 3,48E-01 3,48E-01 3,48E-01 1,00E-01 0,00E+00 8,44E-01 2,84E-02 3,48E-01 3,48E-01 C2H4 0,00 0,00 0,00 0,00 1,00 0,00 0,00 0,00 0,00 C2H6 5,02E-02 5,02E-02 5,02E-02 9,16E-05 0,00E+00 1,10E-01 9,57E-01 5,02E-02 5,02E-02 H2S 2,44E-05 2,44E-05 2,44E-05 1,75E-09 0,00E+00 2,88E-05 1,05E-03 2,44E-05 2,44E-05 C6H6 1,06E-04 1,06E-04 1,06E-04 7,73E-17 0,00E+00 2,31E-07 7,57E-03 1,06E-04 1,06E-04 C7H8 9,03E-05 9,03E-05 9,03E-05 5,39E-19 0,00E+00 4,13E-08 6,42E-03 9,03E-05 9,03E-05

Mass Flow (kg/h) H2 3,31E+03 3,31E+03 3,31E+03 3,23E+03 0,00E+00 8,56E+01 6,51E-11 3,15E+03 1,66E+02

CH4 1,52E+04 1,52E+04 1,52E+04 2,85E+03 0,00E+00 1,24E+04 1,75E+01 1,45E+04 7,62E+02 C2H4 0,00 0,00 0,00 0,00 1299,90 0,00 0,00 0,00 0,00 C2H6 4,12E+03 4,12E+03 4,12E+03 4,90E+00 0,00E+00 3,01E+03 1,10E+03 3,91E+03 2,06E+02 H2S 2,27E+00 2,27E+00 2,27E+00 1,06E-04 0,00E+00 8,97E-01 1,37E+00 2,16E+00 1,14E-01 C6H6 2,27E+01 2,27E+01 2,27E+01 1,07E-11 0,00E+00 1,65E-02 2,27E+01 2,16E+01 1,14E+00 C7H8 2,27E+01 2,27E+01 2,27E+01 8,85E-14 0,00E+00 3,48E-03 2,27E+01 2,16E+01 1,14E+00

Mass Frac H2 1,46E-01 1,46E-01 1,46E-01 5,30E-01 0,00E+00 5,53E-03 5,58E-14 1,46E-01 1,46E-01

80

CH4 6,71E-01 6,71E-01 6,71E-01 4,69E-01 0,00E+00 8,00E-01 1,50E-02 6,71E-01 6,71E-01 C2H4 0,00 0,00 0,00 0,00 1,00 0,00 0,00 0,00 0,00 C2H6 1,81E-01 1,81E-01 1,81E-01 8,05E-04 0,00E+00 1,95E-01 9,45E-01 1,81E-01 1,81E-01 H2S 1,00E-04 1,00E-04 1,00E-04 1,74E-08 0,00E+00 5,80E-05 1,18E-03 1,00E-04 1,00E-04 C6H6 1,00E-03 1,00E-03 1,00E-03 1,76E-15 0,00E+00 1,07E-06 1,94E-02 1,00E-03 1,00E-03 C7H8 1,00E-03 1,00E-03 1,00E-03 1,45E-17 0,00E+00 2,25E-07 1,94E-02 1,00E-03 1,00E-03 Total Flow (kmol/h) 2731,60 2731,60 2731,60 1779,62 46,34 913,60 38,38 2595,02 136,58 Total Flow (kg/h) 22722,00 22722,00 22722,00 6087,41 1299,90 15466,38 1168,21 21585,90 1136,10 Total Flow (cum/h) 1185,93 853,55 683,83 788,65 2,64 3969,71 119,88 810,87 42,68 Temperature (°C) 10,05 -60,46 -91,75 -2,83 -73,42 -1,75 4,98 -60,46 -60,46 Pressure (bar) 54,51 54,16 53,82 52,26 4,52 5,12 6,84 54,16 54,16 Vapor Frac 1,00 1,00 0,95 1,00 0,00 1,00 0,99 1,00 1,00 Liquid Frac 0,00 0,00 0,05 0,00 1,00 0,00 0,01 0,00 0,00 Solid Frac 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 Enthalpy (J/kmol) -3,10E+07 -3,35E+07 -3,50E+07 -8,33E+06 3,59E+07 -7,36E+07 -8,41E+07 -3,35E+07 -3,35E+07 Enthalpy (J/kg) -3,73E+06 -4,02E+06 -4,21E+06 -2,43E+06 1,28E+06 -4,35E+06 -2,76E+06 -4,02E+06 -4,02E+06 Enthalpy (kW) -23538,17 -25399,98 -26541,49 -4115,74 462,49 -18673,05 -896,34 -24129,98 -1270,00 Entropy (J/kmol-K) -6,57E+04 -7,56E+04 -8,33E+04 -4,13E+04 -1,43E+05 -1,00E+05 -1,91E+05 -7,56E+04 -7,56E+04 Entropy (J/kg-K) -7,90E+03 -9,09E+03 -1,00E+04 -1,21E+04 -5,10E+03 -5,90E+03 -6,28E+03 -9,09E+03 -9,09E+03 Density (kmol/cum) 2,30 3,20 3,99 2,26 17,56 0,23 0,32 3,20 3,20 Density (kg/cum) 19,16 26,62 33,23 7,72 492,57 3,90 9,75 26,62 26,62 Average (MW) 8,32 8,32 8,32 3,42 28,05 16,93 30,44 8,32 8,32 Liq Vol 60F (cum/h) 150,60 150,60 150,60 95,32 2,48 52,06 3,22 143,07 7,53 S10 S11 S12 S13 S14 S15 S16 S17 S18 Substream: MIXED Mole Flow (kmol/h)

H2 8,22E+01 1,60E+03 3,23E-11 4,25E+01 3,74E+01 1,02E-08 1,56E+03 1,01E-08 3,23E-11 CH4 4,75E+01 1,78E+02 1,09E+00 7,52E+02 6,98E+02 2,00E+01 9,02E+02 1,89E+01 1,09E+00 C2H4 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 C2H6 6,85E+00 1,63E-01 3,67E+01 5,43E+01 5,25E+01 8,26E+01 1,30E+02 4,59E+01 3,67E+01 H2S 3,33E-03 3,11E-06 4,03E-02 9,69E-03 9,52E-03 5,70E-02 6,33E-02 1,66E-02 4,03E-02 C6H6 1,45E-02 1,37E-13 2,91E-01 3,16E-05 3,16E-05 2,91E-01 2,76E-01 1,80E-04 2,91E-01 C7H8 1,23E-02 9,60E-16 2,47E-01 4,38E-06 4,38E-06 2,47E-01 2,34E-01 3,34E-05 2,47E-01

81

Mole Frac H2 6,02E-01 9,00E-01 8,42E-13 5,00E-02 4,75E-02 9,87E-11 6,02E-01 1,57E-10 8,42E-13 CH4 3,48E-01 1,00E-01 2,84E-02 8,86E-01 8,86E-01 1,94E-01 3,48E-01 2,92E-01 2,84E-02 C2H4 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 C2H6 5,02E-02 9,16E-05 9,57E-01 6,39E-02 6,66E-02 8,01E-01 5,02E-02 7,08E-01 9,57E-01 H2S 2,44E-05 1,75E-09 1,05E-03 1,14E-05 1,21E-05 5,52E-04 2,44E-05 2,57E-04 1,05E-03 C6H6 1,06E-04 7,73E-17 7,57E-03 3,72E-08 4,01E-08 2,82E-03 1,06E-04 2,77E-06 7,57E-03 C7H8 9,03E-05 5,39E-19 6,42E-03 5,16E-09 5,56E-09 2,39E-03 9,03E-05 5,15E-07 6,42E-03

Mass Flow (kg/h) H2 1,66E+02 3,23E+03 6,51E-11 8,56E+01 7,54E+01 2,05E-08 3,15E+03 2,05E-08 6,51E-11 CH4 7,62E+02 2,85E+03 1,75E+01 1,21E+04 1,12E+04 3,20E+02 1,45E+04 3,03E+02 1,75E+01 C2H4 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 C2H6 2,06E+02 4,90E+00 1,10E+03 1,63E+03 1,58E+03 2,48E+03 3,91E+03 1,38E+03 1,10E+03 H2S 1,14E-01 1,06E-04 1,37E+00 3,30E-01 3,24E-01 1,94E+00 2,16E+00 5,67E-01 1,37E+00 C6H6 1,14E+00 1,07E-11 2,27E+01 2,47E-03 2,46E-03 2,27E+01 2,16E+01 1,40E-02 2,27E+01 C7H8 1,14E+00 8,85E-14 2,27E+01 4,04E-04 4,04E-04 2,27E+01 2,16E+01 3,07E-03 2,27E+01

Mass Frac H2 1,46E-01 5,30E-01 5,58E-14 6,21E-03 5,87E-03 7,20E-12 1,46E-01 1,22E-11 5,58E-14 CH4 6,71E-01 4,69E-01 1,50E-02 8,75E-01 8,71E-01 1,12E-01 6,71E-01 1,80E-01 1,50E-02 C2H4 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 C2H6 1,81E-01 8,05E-04 9,45E-01 1,18E-01 1,23E-01 8,71E-01 1,81E-01 8,20E-01 9,45E-01 H2S 1,00E-04 1,74E-08 1,18E-03 2,40E-05 2,53E-05 6,81E-04 1,00E-04 3,37E-04 1,18E-03 C6H6 1,00E-03 1,76E-15 1,94E-02 1,79E-07 1,92E-07 7,97E-03 1,00E-03 8,34E-06 1,94E-02 C7H8 1,00E-03 1,45E-17 1,94E-02 2,93E-08 3,14E-08 7,97E-03 1,00E-03 1,83E-06 1,94E-02 Total Flow (kmol/h) 136,58 1779,62 38,38 848,81 787,45 103,17 2595,02 64,79 38,38 Total Flow (kg/h) 1136,10 6087,41 1168,21 13783,03 12843,81 2851,57 21585,90 1683,36 1168,21 Total Flow (cum/h) 30,02 560,90 2,31 2691,88 138,87 164,38 653,63 162,07 2,31 Temperature (°C) -104,86 -78,32 -54,46 -65,00 -149,33 -54,56 -91,05 -54,56 -54,56 Pressure (bar) 54,16 52,61 7,00 5,29 5,61 6,57 53,82 6,57 6,57 Vapor Frac 0,90 1,00 0,00 1,00 0,08 0,63 0,95 1,00 0,00 Liquid Frac 0,10 0,00 1,00 0,00 0,92 0,37 0,05 0,00 1,00 Solid Frac 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00

82

Enthalpy (J/kmol) -3,58E+07 -1,06E+07 -9,98E+07 -7,51E+07 -8,59E+07 -9,10E+07 -3,49E+07 -8,58E+07 -9,98E+07 Enthalpy (J/kg) -4,30E+06 -3,09E+06 -3,28E+06 -4,63E+06 -5,27E+06 -3,29E+06 -4,20E+06 -3,30E+06 -3,28E+06 Enthalpy (kW) -1358,21 -5224,68 -1064,28 -17713,37 -18793,19 -2609,00 -25183,28 -1544,62 -1064,37 Entropy (J/kmol-K) -8,80E+04 -5,11E+04 -2,59E+05 -1,06E+05 -1,79E+05 -2,04E+05 -8,30E+04 -1,72E+05 -2,59E+05 Entropy (J/kg-K) -1,06E+04 -1,49E+04 -8,51E+03 -6,53E+03 -1,10E+04 -7,40E+03 -9,98E+03 -6,63E+03 -8,51E+03 Density (kmol/cum) 4,55 3,17 16,59 0,32 5,67 0,63 3,97 0,40 16,60 Density (kg/cum) 37,84 10,85 505,01 5,12 92,49 17,35 33,02 10,39 505,17 Average (MW) 8,32 3,42 30,44 16,24 16,31 27,64 8,32 25,98 30,44 Liq Vol 60F (cum/h) 7,53 95,32 3,22 47,15 43,81 8,13 143,07 4,90 3,22

S19 S20 S21 S22 S23 S24 S25 S26 S27

Substream: MIXED Mole Flow (kmol/h) H2 4,25E+01 1,64E+03 5,03E+00 1,64E+03 1,64E+03 1,60E+03 0,00E+00 3,74E+01 1,60E+03 CH4 7,71E+02 8,75E+02 7,45E+01 8,75E+02 9,50E+02 1,78E+02 0,00E+00 6,98E+02 1,78E+02 C2H4 0,00 0,00 0,00 0,00 0,00 0,00 3,56 0,00 0,00 C2H6 1,00E+02 5,26E+01 8,44E+01 5,26E+01 1,37E+02 1,63E-01 0,00E+00 5,25E+01 1,63E-01 H2S 2,63E-02 9,52E-03 5,71E-02 9,52E-03 6,67E-02 3,11E-06 0,00E+00 9,52E-03 3,11E-06 C6H6 2,11E-04 3,16E-05 2,91E-01 3,16E-05 2,91E-01 1,37E-13 0,00E+00 3,16E-05 1,37E-13 C7H8 3,77E-05 4,38E-06 2,47E-01 4,38E-06 2,47E-01 9,60E-16 0,00E+00 4,38E-06 9,60E-16

Mole Frac H2 4,65E-02 6,38E-01 3,06E-02 6,38E-01 6,02E-01 9,00E-01 0,00E+00 4,75E-02 9,00E-01 CH4 8,44E-01 3,41E-01 4,53E-01 3,41E-01 3,48E-01 1,00E-01 0,00E+00 8,86E-01 1,00E-01 C2H4 0,00 0,00 0,00 0,00 0,00 0,00 1,00 0,00 0,00 C2H6 1,10E-01 2,05E-02 5,13E-01 2,05E-02 5,02E-02 9,16E-05 0,00E+00 6,66E-02 9,16E-05 H2S 2,88E-05 3,71E-06 3,47E-04 3,71E-06 2,44E-05 1,75E-09 0,00E+00 1,21E-05 1,75E-09 C6H6 2,31E-07 1,23E-08 1,77E-03 1,23E-08 1,06E-04 7,73E-17 0,00E+00 4,01E-08 7,73E-17 C7H8 4,13E-08 1,71E-09 1,50E-03 1,71E-09 9,03E-05 5,39E-19 0,00E+00 5,56E-09 5,39E-19

Mass Flow (kg/h) H2 8,56E+01 3,30E+03 1,01E+01 3,30E+03 3,31E+03 3,23E+03 0,00E+00 7,54E+01 3,23E+03 CH4 1,24E+04 1,40E+04 1,20E+03 1,40E+04 1,52E+04 2,85E+03 0,00E+00 1,12E+04 2,85E+03 C2H4 0,00 0,00 0,00 0,00 0,00 0,00 100,00 0,00 0,00 C2H6 3,01E+03 1,58E+03 2,54E+03 1,58E+03 4,12E+03 4,90E+00 0,00E+00 1,58E+03 4,90E+00

83

H2S 8,97E-01 3,24E-01 1,95E+00 3,24E-01 2,27E+00 1,06E-04 0,00E+00 3,24E-01 1,06E-04 C6H6 1,65E-02 2,46E-03 2,27E+01 2,46E-03 2,27E+01 1,07E-11 0,00E+00 2,46E-03 1,07E-11 C7H8 3,48E-03 4,04E-04 2,27E+01 4,04E-04 2,27E+01 8,85E-14 0,00E+00 4,04E-04 8,85E-14

Mass Frac

H2 5,53E-03 1,75E-01 2,67E-03 1,75E-01 1,46E-01 5,30E-01 0,00E+00 5,87E-03 5,30E-01 CH4 8,00E-01 7,42E-01 3,15E-01 7,42E-01 6,71E-01 4,69E-01 0,00E+00 8,71E-01 4,69E-01 C2H4 0,00 0,00 0,00 0,00 0,00 0,00 1,00 0,00 0,00 C2H6 1,95E-01 8,36E-02 6,69E-01 8,36E-02 1,81E-01 8,05E-04 0,00E+00 1,23E-01 8,05E-04 H2S 5,80E-05 1,71E-05 5,14E-04 1,71E-05 1,00E-04 1,74E-08 0,00E+00 2,53E-05 1,74E-08 C6H6 1,07E-06 1,30E-07 5,99E-03 1,30E-07 1,00E-03 1,76E-15 0,00E+00 1,92E-07 1,76E-15 C7H8 2,25E-07 2,13E-08 5,99E-03 2,13E-08 1,00E-03 1,45E-17 0,00E+00 3,14E-08 1,45E-17 Total Flow (kmol/h) 913,60 2567,07 164,54 2567,07 2731,60 1779,62 3,56 787,45 1779,62 Total Flow (kg/h) 15466,38 18931,22 3790,78 18931,22 22722,00 6087,41 100,00 12843,81 6087,41 Total Flow (cum/h) 2897,37 660,69 8,37 373,62 669,06 472,63 0,20 1863,78 341,63 Temperature (°C) -64,41 -95,00 -95,00 -147,28 -95,00 -106,29 -73,42 -106,29 -147,28 Pressure (bar) 5,29 53,47 53,47 53,12 53,47 52,96 4,52 5,45 53,12 Vapor Frac 1,00 1,00 0,00 0,69 0,94 1,00 0,00 0,99 1,00 Liquid Frac 0,00 0,00 1,00 0,31 0,06 0,00 1,00 0,01 0,00 Solid Frac 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 Enthalpy (J/kmol) -7,59E+07 -3,15E+07 -9,27E+07 -3,51E+07 -3,52E+07 -1,14E+07 3,59E+07 -7,69E+07 -1,26E+07 Enthalpy (J/kg) -4,48E+06 -4,27E+06 -4,02E+06 -4,76E+06 -4,23E+06 -3,33E+06 1,28E+06 -4,72E+06 -3,67E+06 Enthalpy (kW) -19257,99 -22453,84 -4234,54 -25004,73 -26688,37 -5627,51 35,58 -16826,30 -6211,54 Entropy (J/kmol-K) -1,10E+05 -7,60E+04 -2,14E+05 -9,99E+04 -8,43E+04 -5,57E+04 -1,43E+05 -1,15E+05 -6,38E+04 Entropy (J/kg-K) -6,49E+03 -1,03E+04 -9,27E+03 -1,35E+04 -1,01E+04 -1,63E+04 -5,10E+03 -7,06E+03 -1,87E+04 Density (kmol/cum) 0,32 3,89 19,66 6,87 4,08 3,77 17,56 0,42 5,21 Density (kg/cum) 5,34 28,65 452,97 50,67 33,96 12,88 492,57 6,89 17,82 Average (MW) 16,93 7,37 23,04 7,37 8,32 3,42 28,05 16,31 3,42 Liq Vol 60F (cum/h) 52,06 139,13 11,47 139,13 150,60 95,32 0,19 43,81 95,32

84

S28 S29 S30 S31 S32 S33 Substream: MIXED Heat stream Mole Flow (kmol/h) H2 3,74E+01 4,25E+01 5,03E+00 1,02E-08 0,00E+00 CH4 6,98E+02 7,52E+02 5,45E+01 2,00E+01 0,00E+00 C2H4 0,00 0,00 0,00 0,00 46,34 C2H6 5,25E+01 5,43E+01 1,80E+00 8,26E+01 0,00E+00 H2S 9,52E-03 9,69E-03 1,70E-04 5,70E-02 0,00E+00 C6H6 3,16E-05 3,16E-05 3,99E-08 2,91E-01 0,00E+00 C7H8 4,38E-06 4,38E-06 2,86E-09 2,47E-01 0,00E+00

Mole Frac H2 4,75E-02 5,00E-02 8,20E-02 9,87E-11 0,00E+00 CH4 8,86E-01 8,86E-01 8,89E-01 1,94E-01 0,00E+00 C2H4 0,00 0,00 0,00 0,00 1,00 C2H6 6,66E-02 6,39E-02 2,94E-02 8,01E-01 0,00E+00 H2S 1,21E-05 1,14E-05 2,76E-06 5,52E-04 0,00E+00 C6H6 4,01E-08 3,72E-08 6,50E-10 2,82E-03 0,00E+00 C7H8 5,56E-09 5,16E-09 4,65E-11 2,39E-03 0,00E+00

Mass Flow (kg/h) H2 7,54E+01 8,56E+01 1,01E+01 2,05E-08 0,00E+00 CH4 1,12E+04 1,21E+04 8,75E+02 3,20E+02 0,00E+00 C2H4 0,00 0,00 0,00 0,00 1299,90 C2H6 1,58E+03 1,63E+03 5,42E+01 2,48E+03 0,00E+00 H2S 3,24E-01 3,30E-01 5,78E-03 1,94E+00 0,00E+00 C6H6 2,46E-03 2,47E-03 3,12E-06 2,27E+01 0,00E+00 C7H8 4,04E-04 4,04E-04 2,63E-07 2,27E+01 0,00E+00

Mass Frac H2 5,87E-03 6,21E-03 1,08E-02 7,20E-12 0,00E+00

CH4 8,71E-01 8,75E-01 9,32E-01 1,12E-01 0,00E+00 C2H4 0,00 0,00 0,00 0,00 1,00 C2H6 1,23E-01 1,18E-01 5,77E-02 8,71E-01 0,00E+00 H2S 2,53E-05 2,40E-05 6,16E-06 6,81E-04 0,00E+00 C6H6 1,92E-07 1,79E-07 3,32E-09 7,97E-03 0,00E+00

85

C7H8 3,14E-08 2,93E-08 2,80E-10 7,97E-03 0,00E+00 Total Flow (kmol/h) 787,45 848,81 61,37 103,17 46,34 Total Flow (kg/h) 12843,81 13783,03 939,21 2851,57 1299,90 Total Flow (cum/h) 31,98 2003,63 109,16 5,45 538,22 Temperature (°C) -147,28 -106,82 -112,01 -93,93 -101,19 Pressure (bar) 53,12 5,45 6,93 6,73 1,19 Vapor Frac 0,00 0,99 1,00 0,00 1,00 Liquid Frac 1,00 0,01 0,00 1,00 0,00 Solid Frac 0,00 0,00 0,00 0,00 0,00 Enthalpy (J/kmol) -8,59E+07 -7,67E+07 -7,40E+07 -1,01E+08 4,73E+07 Enthalpy (J/kg) -5,27E+06 -4,72E+06 -4,84E+06 -3,64E+06 1,69E+06 Enthalpy (kW) -18793,19 -18088,33 -1262,04 -2884,49 609,38 88 Entropy (J/kmol-K) -1,81E+05 -1,15E+05 -1,12E+05 -2,52E+05 -7,58E+04 Entropy (J/kg-K) -1,11E+04 -7,07E+03 -7,29E+03 -9,10E+03 -2,70E+03 Density (kmol/cum) 24,62 0,42 0,56 18,94 0,09 Density (kg/cum) 401,60 6,88 8,60 523,41 2,42 Average (MW) 16,31 16,24 15,31 27,64 28,05 Liq Vol 60F (cum/h) 43,81 47,15 3,34 8,13 2,48

86

Appendix A2 – Case Study 2

VARY 1 VARY 2 S1 S1 S1 S1 S4 S4 S4 S4 S4

MIXED MIXED MIXED MIXED MIXED MIXED MIXED MIXED MIXED Total mass

flow H2 mass

flow H2 mol frac CH4 mol frac H2 mass flow CH4 mass flow Total flow H2 mol frac Temperature t/h t/h t/h t/h t/h C

22,7 0,80 0,224 0,78 0,65 0,57 1,22 0,90 -93,23 22,7 0,90 0,246 0,75 0,75 0,66 1,41 0,90 -68,13 22,7 1,00 0,267 0,73 0,85 0,75 1,60 0,90 -55,45 22,7 1,10 0,287 0,71 0,95 0,84 1,79 0,90 -46,08 22,7 1,20 0,306 0,69 1,05 0,93 1,98 0,90 -38,87 22,7 1,30 0,325 0,68 1,16 1,02 2,17 0,90 -33,19 22,7 1,40 0,342 0,66 1,26 1,11 2,36 0,90 -28,61 22,7 1,50 0,359 0,64 1,36 1,20 2,55 0,90 -24,84 22,7 1,60 0,375 0,63 1,46 1,29 2,75 0,90 -21,70 22,7 1,70 0,390 0,61 1,56 1,37 2,94 0,90 -19,05 22,7 1,80 0,405 0,59 1,66 1,46 3,13 0,90 -16,78 22,7 1,90 0,419 0,58 1,76 1,55 3,32 0,90 -14,82 22,7 2,00 0,433 0,57 1,87 1,64 3,51 0,90 -13,13 22,7 2,10 0,446 0,55 1,97 1,73 3,70 0,90 -11,64 22,7 2,20 0,459 0,54 2,07 1,82 3,89 0,90 -10,33 22,7 2,30 0,471 0,53 2,17 1,91 4,08 0,90 -9,13 22,7 2,40 0,483 0,52 2,27 2,00 4,27 0,90 -8,10 22,7 2,50 0,495 0,51 2,37 2,09 4,46 0,90 -7,18 22,7 2,60 0,506 0,49 2,47 2,18 4,65 0,90 -6,34 22,7 2,70 0,516 0,48 2,57 2,27 4,84 0,90 -5,59 22,7 2,80 0,527 0,47 2,68 2,36 5,03 0,90 -4,91 22,7 2,90 0,537 0,46 2,78 2,44 5,22 0,90 -4,29 22,7 3,00 0,546 0,45 2,88 2,53 5,41 0,90 -3,73 22,7 3,10 0,556 0,44 2,98 2,62 5,60 0,90 -3,21

87

22,7 3,20 0,565 0,44 3,08 2,71 5,79 0,90 -2,73 22,7 3,30 0,573 0,43 3,18 2,80 5,98 0,90 -2,30 22,7 3,40 0,582 0,42 3,28 2,89 6,17 0,90 -1,90 22,7 3,50 0,590 0,41 3,39 2,98 6,37 0,90 -1,52 22,7 3,60 0,598 0,40 3,49 3,07 6,56 0,90 -1,18 22,7 3,70 0,606 0,39 3,59 3,16 6,75 0,90 -0,86 22,7 3,80 0,614 0,39 3,69 3,25 6,94 0,90 -0,56 22,7 3,90 0,621 0,38 3,79 3,34 7,13 0,90 -0,29 22,7 4,00 0,628 0,37 3,89 3,43 7,32 0,90 -0,03 22,7 4,10 0,635 0,36 3,99 3,52 7,51 0,90 0,21 22,7 4,20 0,642 0,36 4,09 3,60 7,70 0,90 0,44 22,7 4,30 0,649 0,35 4,20 3,69 7,89 0,90 0,65 22,7 4,40 0,655 0,35 4,30 3,78 8,08 0,90 0,85 22,7 4,50 0,661 0,34 4,40 3,87 8,27 0,90 1,04 22,7 4,60 0,667 0,33 4,50 3,96 8,46 0,90 1,21 22,7 4,70 0,673 0,33 4,60 4,05 8,65 0,90 1,38 22,7 4,80 0,679 0,32 4,70 4,14 8,84 0,90 1,53 22,7 4,90 0,685 0,32 4,80 4,23 9,03 0,90 1,68 22,7 5,00 0,690 0,31 4,91 4,32 9,22 0,90 1,82 22,7 5,10 0,696 0,30 5,01 4,41 9,41 0,90 1,95 22,7 5,20 0,701 0,30 5,11 4,50 9,60 0,90 2,07 22,7 5,30 0,706 0,29 5,21 4,59 9,79 0,90 2,19 22,7 5,40 0,711 0,29 5,31 4,67 9,99 0,90 2,30 22,7 5,50 0,716 0,28 5,41 4,76 10,18 0,90 2,41 22,7 5,60 0,721 0,28 5,51 4,85 10,37 0,90 2,51 22,7 5,70 0,726 0,27 5,61 4,94 10,56 0,90 2,61 22,7 5,80 0,730 0,27 5,72 5,03 10,75 0,90 2,70 22,7 5,90 0,735 0,27 5,82 5,12 10,94 0,90 2,79 22,7 6,00 0,739 0,26 5,92 5,21 11,13 0,90 2,87 22,7 6,10 0,743 0,26 6,02 5,30 11,32 0,90 2,95 22,7 6,20 0,748 0,25 6,12 5,39 11,51 0,90 3,03 22,7 6,30 0,752 0,25 6,22 5,48 11,70 0,90 3,10 22,7 6,40 0,756 0,24 6,32 5,57 11,89 0,90 3,17

88

22,7 6,50 0,760 0,24 6,42 5,66 12,08 0,90 3,23 22,7 6,60 0,764 0,24 6,53 5,74 12,27 0,90 3,30 22,7 6,70 0,767 0,23 6,63 5,83 12,46 0,90 3,36 22,7 6,80 0,771 0,23 6,73 5,92 12,65 0,90 3,41 22,7 6,90 0,775 0,23 6,83 6,01 12,84 0,90 3,47 22,7 7,00 0,778 0,22 6,93 6,10 13,03 0,90 3,52 22,7 7,10 0,782 0,22 7,03 6,19 13,22 0,90 3,57 22,7 7,20 0,785 0,21 7,13 6,28 13,41 0,90 3,62 22,7 7,30 0,789 0,21 7,24 6,37 13,60 0,90 3,67 22,7 7,40 0,792 0,21 7,34 6,46 13,80 0,90 3,72 22,7 7,50 0,795 0,20 7,44 6,55 13,99 0,90 3,76 22,7 7,60 0,798 0,20 7,54 6,64 14,18 0,90 3,80 22,7 7,70 0,802 0,20 7,64 6,73 14,37 0,90 3,84 22,7 7,80 0,805 0,20 7,74 6,82 14,56 0,90 3,88 22,7 7,90 0,808 0,19 7,84 6,90 14,75 0,90 3,92 22,7 8,00 0,811 0,19 7,94 6,99 14,94 0,90 3,96 22,7 8,10 0,814 0,19 8,05 7,08 15,13 0,90 3,99 22,7 8,20 0,816 0,18 8,15 7,17 15,32 0,90 4,02 22,7 8,30 0,819 0,18 8,25 7,26 15,51 0,90 4,06 22,7 8,40 0,822 0,18 8,35 7,35 15,70 0,90 4,09 22,7 8,50 0,825 0,18 8,45 7,44 15,89 0,90 4,12 22,7 8,60 0,827 0,17 8,55 7,53 16,08 0,90 4,15 22,7 8,70 0,830 0,17 8,65 7,62 16,27 0,90 4,18 22,7 8,80 0,833 0,17 8,76 7,71 16,46 0,90 4,20 22,7 8,90 0,835 0,16 8,86 7,80 16,65 0,90 4,23 22,7 9,00 0,838 0,16 8,96 7,89 16,84 0,90 4,26 22,7 9,10 0,840 0,16 9,06 7,97 17,03 0,90 4,28 22,7 9,20 0,842 0,16 9,16 8,06 17,22 0,90 4,30 22,7 9,30 0,845 0,16 9,26 8,15 17,41 0,90 4,33 22,7 9,40 0,847 0,15 9,36 8,24 17,61 0,90 4,35 22,7 9,50 0,850 0,15 9,46 8,33 17,80 0,90 4,37 22,7 9,60 0,852 0,15 9,57 8,42 17,99 0,90 4,39 22,7 9,70 0,854 0,15 9,67 8,51 18,18 0,90 4,41

89

22,7 9,80 0,856 0,14 9,77 8,60 18,37 0,90 4,43 22,7 9,90 0,858 0,14 9,87 8,69 18,56 0,90 4,45 22,7 10,00 0,861 0,14 9,97 8,78 18,75 0,90 4,47 22,7 10,10 0,863 0,14 10,07 8,87 18,94 0,90 4,49 22,7 10,20 0,865 0,14 10,17 8,96 19,13 0,90 4,51 22,7 10,30 0,867 0,13 10,28 9,04 19,32 0,90 4,52 22,7 10,40 0,869 0,13 10,38 9,13 19,51 0,90 4,54 22,7 10,50 0,871 0,13 10,48 9,22 19,70 0,90 4,56 22,7 10,60 0,873 0,13 10,58 9,31 19,89 0,90 4,57 22,7 10,70 0,875 0,13 10,68 9,40 20,08 0,90 4,59 22,7 10,80 0,877 0,12 10,78 9,49 20,27 0,90 4,60 22,7 10,90 0,878 0,12 10,88 9,58 20,46 0,90 4,62 22,7 11,00 0,880 0,12 10,98 9,67 20,65 0,90 4,63 22,7 11,10 0,882 0,12 11,09 9,76 20,84 0,90 4,64 22,7 11,20 0,884 0,12 11,19 9,85 21,03 0,90 4,66 22,7 11,30 0,886 0,11 11,29 9,94 21,23 0,90 4,67 22,7 11,40 0,887 0,11 11,39 10,03 21,42 0,90 4,68 22,7 11,50 0,889 0,11 11,49 10,12 21,61 0,90 4,69 22,7 11,60 0,891 0,11 11,59 10,20 21,80 0,90 4,71 22,7 11,70 0,892 0,11 11,69 10,29 21,99 0,90 4,72 22,7 11,80 0,894 0,11 11,79 10,38 22,18 0,90 4,73 22,7 11,90 0,896 0,10 11,90 10,47 22,37 0,90 4,74 22,7 12,00 0,897 0,10 12,00 10,56 22,56 0,90 4,75 22,7 12,10 0,899 0,10 12,10 10,65 22,75 0,90 4,76

90



flow H2 mass

flow Pressure H2 mass flow CH4 mass flow Total flow Temperature Pressure H2 mass flow t/h t/h bar t/h t/h t/h C bar t/h

22,7 0,00 52,26 0,15 21,19 21,34 -1,75 5,12 0,00 22,7 0,00 52,26 0,15 21,26 21,41 -1,75 5,12 0,00 22,7 0,00 52,26 0,15 21,08 21,22 -1,75 5,12 0,00 22,7 0,00 52,26 0,15 20,89 21,03 -1,75 5,12 0,00 22,7 0,00 52,26 0,15 20,70 20,84 -1,75 5,12 0,00 22,7 0,00 52,26 0,14 20,51 20,65 -1,75 5,12 0,00 22,7 0,00 52,26 0,14 20,32 20,46 -1,75 5,12 0,00 22,7 0,00 52,26 0,14 20,13 20,27 -1,75 5,12 0,00 22,7 0,00 52,26 0,14 19,94 20,08 -1,75 5,12 0,00 22,7 0,00 52,26 0,14 19,75 19,89 -1,75 5,12 0,00 22,7 0,00 52,26 0,14 19,57 19,70 -1,75 5,12 0,00 22,7 0,00 52,26 0,14 19,38 19,51 -1,75 5,12 0,00 22,7 0,00 52,26 0,13 19,19 19,32 -1,75 5,12 0,00 22,7 0,00 52,26 0,13 19,00 19,13 -1,75 5,12 0,00 22,7 0,00 52,26 0,13 18,81 18,94 -1,75 5,12 0,00 22,7 0,00 52,26 0,13 18,62 18,75 -1,75 5,12 0,00 22,7 0,00 52,26 0,13 18,43 18,56 -1,75 5,12 0,00 22,7 0,00 52,26 0,13 18,25 18,37 -1,75 5,12 0,00 22,7 0,00 52,26 0,13 18,06 18,18 -1,75 5,12 0,00 22,7 0,00 52,26 0,13 17,87 17,99 -1,75 5,12 0,00 22,7 0,00 52,26 0,12 17,68 17,80 -1,75 5,12 0,00 22,7 0,00 52,26 0,12 17,49 17,61 -1,75 5,12 0,00 22,7 0,00 52,26 0,12 17,30 17,42 -1,75 5,12 0,00 22,7 0,00 52,26 0,12 17,11 17,23 -1,75 5,12 0,00 22,7 0,00 52,26 0,12 16,93 17,04 -1,75 5,12 0,00 22,7 0,00 52,26 0,12 16,74 16,85 -1,75 5,12 0,00 22,7 0,00 52,26 0,12 16,55 16,66 -1,75 5,12 0,00 22,7 0,00 52,26 0,11 16,36 16,47 -1,75 5,12 0,00

91

22,7 0,00 52,26 0,11 16,17 16,28 -1,75 5,12 0,00 22,7 0,00 52,26 0,11 15,98 16,09 -1,75 5,12 0,00 22,7 0,00 52,26 0,11 15,79 15,90 -1,75 5,12 0,00 22,7 0,00 52,26 0,11 15,61 15,71 -1,75 5,12 0,00 22,7 0,00 52,26 0,11 15,42 15,52 -1,75 5,12 0,00 22,7 0,00 52,26 0,11 15,23 15,33 -1,75 5,12 0,00 22,7 0,00 52,26 0,11 15,04 15,14 -1,75 5,12 0,00 22,7 0,00 52,26 0,10 14,85 14,96 -1,75 5,12 0,00 22,7 0,00 52,26 0,10 14,66 14,77 -1,75 5,12 0,00 22,7 0,00 52,26 0,10 14,47 14,58 -1,75 5,12 0,00 22,7 0,00 52,26 0,10 14,29 14,39 -1,75 5,12 0,00 22,7 0,00 52,26 0,10 14,10 14,20 -1,75 5,12 0,00 22,7 0,00 52,26 0,10 13,91 14,01 -1,75 5,12 0,00 22,7 0,00 52,26 0,10 13,72 13,82 -1,75 5,12 0,00 22,7 0,00 52,26 0,09 13,53 13,63 -1,75 5,12 0,00 22,7 0,00 52,26 0,09 13,34 13,44 -1,75 5,12 0,00 22,7 0,00 52,26 0,09 13,15 13,25 -1,75 5,12 0,00 22,7 0,00 52,26 0,09 12,97 13,06 -1,75 5,12 0,00 22,7 0,00 52,26 0,09 12,78 12,87 -1,75 5,12 0,00 22,7 0,00 52,26 0,09 12,59 12,68 -1,75 5,12 0,00 22,7 0,00 52,26 0,09 12,40 12,49 -1,75 5,12 0,00 22,7 0,00 52,26 0,09 12,21 12,30 -1,75 5,12 0,00 22,7 0,00 52,26 0,08 12,02 12,11 -1,75 5,12 0,00 22,7 0,00 52,26 0,08 11,84 11,92 -1,75 5,12 0,00 22,7 0,00 52,26 0,08 11,65 11,73 -1,75 5,12 0,00 22,7 0,00 52,26 0,08 11,46 11,54 -1,75 5,12 0,00 22,7 0,00 52,26 0,08 11,27 11,35 -1,75 5,12 0,00 22,7 0,00 52,26 0,08 11,08 11,16 -1,75 5,12 0,00 22,7 0,00 52,26 0,08 10,89 10,97 -1,75 5,12 0,00 22,7 0,00 52,26 0,08 10,70 10,78 -1,75 5,12 0,00 22,7 0,00 52,26 0,07 10,52 10,59 -1,75 5,12 0,00 22,7 0,00 52,26 0,07 10,33 10,40 -1,75 5,12 0,00 22,7 0,00 52,26 0,07 10,14 10,21 -1,75 5,12 0,00

92

22,7 0,00 52,26 0,07 9,95 10,02 -1,75 5,12 0,00 22,7 0,00 52,26 0,07 9,76 9,83 -1,75 5,12 0,00 22,7 0,00 52,26 0,07 9,57 9,64 -1,75 5,12 0,00 22,7 0,00 52,26 0,07 9,39 9,45 -1,75 5,12 0,00 22,7 0,00 52,26 0,06 9,20 9,26 -1,75 5,12 0,00 22,7 0,00 52,26 0,06 9,01 9,07 -1,75 5,12 0,00 22,7 0,00 52,26 0,06 8,82 8,88 -1,75 5,12 0,00 22,7 0,00 52,26 0,06 8,63 8,69 -1,75 5,12 0,00 22,7 0,00 52,26 0,06 8,45 8,50 -1,75 5,12 0,00 22,7 0,00 52,26 0,06 8,26 8,31 -1,75 5,12 0,00 22,7 0,00 52,26 0,06 8,07 8,13 -1,75 5,12 0,00 22,7 0,00 52,26 0,06 7,88 7,94 -1,75 5,12 0,00 22,7 0,00 52,26 0,05 7,69 7,75 -1,75 5,12 0,00 22,7 0,00 52,26 0,05 7,50 7,56 -1,75 5,12 0,00 22,7 0,00 52,26 0,05 7,32 7,37 -1,75 5,12 0,00 22,7 0,00 52,26 0,05 7,13 7,18 -1,75 5,12 0,00 22,7 0,00 52,26 0,05 6,94 6,99 -1,75 5,12 0,00 22,7 0,00 52,26 0,05 6,75 6,80 -1,75 5,12 0,00 22,7 0,00 52,26 0,05 6,56 6,61 -1,75 5,12 0,00 22,7 0,00 52,26 0,04 6,38 6,42 -1,75 5,12 0,00 22,7 0,00 52,26 0,04 6,19 6,23 -1,75 5,12 0,00 22,7 0,00 52,26 0,04 6,00 6,04 -1,75 5,12 0,00 22,7 0,00 52,26 0,04 5,81 5,85 -1,75 5,12 0,00 22,7 0,00 52,26 0,04 5,62 5,66 -1,75 5,12 0,00 22,7 0,00 52,26 0,04 5,44 5,47 -1,75 5,12 0,00 22,7 0,00 52,26 0,04 5,25 5,28 -1,75 5,12 0,00 22,7 0,00 52,26 0,04 5,06 5,09 -1,75 5,12 0,00 22,7 0,00 52,26 0,03 4,87 4,91 -1,75 5,12 0,00 22,7 0,00 52,26 0,03 4,68 4,72 -1,75 5,12 0,00 22,7 0,00 52,26 0,03 4,50 4,53 -1,75 5,12 0,00 22,7 0,00 52,26 0,03 4,31 4,34 -1,75 5,12 0,00 22,7 0,00 52,26 0,03 4,12 4,15 -1,75 5,12 0,00 22,7 0,00 52,26 0,03 3,93 3,96 -1,75 5,12 0,00

93

22,7 0,00 52,26 0,03 3,74 3,77 -1,75 5,12 0,00 22,7 0,00 52,26 0,02 3,56 3,58 -1,75 5,12 0,00 22,7 0,00 52,26 0,02 3,37 3,39 -1,75 5,12 0,00 22,7 0,00 52,26 0,02 3,18 3,20 -1,75 5,12 0,00 22,7 0,00 52,26 0,02 2,99 3,01 -1,75 5,12 0,00 22,7 0,00 52,26 0,02 2,81 2,82 -1,75 5,12 0,00 22,7 0,00 52,26 0,02 2,62 2,64 -1,75 5,12 0,00 22,7 0,00 52,26 0,02 2,43 2,45 -1,75 5,12 0,00 22,7 0,00 52,26 0,02 2,24 2,26 -1,75 5,12 0,00 22,7 0,00 52,26 0,01 2,05 2,07 -1,75 5,12 0,00 22,7 0,00 52,26 0,01 1,87 1,88 -1,75 5,12 0,00 22,7 0,00 52,26 0,01 1,68 1,69 -1,75 5,12 0,00 22,7 0,00 52,26 0,01 1,49 1,50 -1,75 5,12 0,00 22,7 0,00 52,26 0,01 1,30 1,31 -1,75 5,12 0,00 22,7 0,00 52,26 0,01 1,12 1,12 -1,75 5,12 0,00 22,7 0,00 52,26 0,01 0,93 0,94 -1,75 5,12 0,00 22,7 0,00 52,26 0,01 0,74 0,75 -1,75 5,12 0,00 22,7 0,00 52,26 0,00 0,55 0,56 -1,75 5,12 0,00 22,7 0,00 52,26 0,00 0,37 0,37 -1,75 5,12 0,00 22,7 0,00 52,26 0,00 0,18 0,18 -1,75 5,12 0,00

94



flow H2 mass

flow Total flow Temperature Pressure C2H4 mass flow Top stripper flow

Bottom stripper flow

Feed stripper flow

t/h t/h t/h C bar t/h t/h t/h t/h 22,7 -0,15 0,26 4,98 6,84 1,64 1,84 0,71 2,55 22,7 -0,15 0,00 - - 0,60 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,55 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,51 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,49 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,47 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,46 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,45 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,45 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,44 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,45 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,45 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,45 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,46 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,46 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,47 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,48 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,49 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,50 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,51 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,52 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,53 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,54 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,55 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,56 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,57 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,58 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,60 0,00 0,00 0,00

95

22,7 -0,15 0,00 - - 0,61 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,62 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,63 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,65 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,66 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,67 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,68 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,70 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,71 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,72 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,74 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,75 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,76 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,78 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,79 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,81 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,82 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,83 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,85 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,86 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,88 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,89 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,90 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,92 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,93 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,95 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,96 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,97 0,00 0,00 0,00 22,7 -0,15 0,00 - - 0,99 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,00 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,02 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,03 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,05 0,00 0,00 0,00

96

22,7 -0,15 0,00 - - 1,06 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,08 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,09 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,10 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,12 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,13 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,15 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,16 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,18 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,19 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,21 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,22 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,24 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,25 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,26 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,28 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,29 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,31 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,32 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,34 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,35 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,37 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,38 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,40 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,41 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,43 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,44 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,46 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,47 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,49 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,50 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,52 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,53 0,00 0,00 0,00

97

22,7 -0,15 0,00 - - 1,55 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,56 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,58 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,59 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,60 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,62 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,63 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,65 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,66 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,68 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,69 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,71 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,72 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,74 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,75 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,77 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,78 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,80 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,81 0,00 0,00 0,00 22,7 -0,15 0,00 - - 1,83 0,00 0,00 0,00

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