University of Miskolc

Faculty of Earth Science and Engineering

Petroleum and Natural Gas Institute

Feasibility analysis and process

simulation of high CO2 content gas

mixture dehydration using glycols for

pipeline transportation

Thesis

Author:

Dániel Kelemen

Petroleum Engineering MSc Student

Instructor:

László Kis

Assistant lecturer

Field Advisor:

Róbert Bíró

Chief Engineer (Production Engineering, MOL Hungary Plc.)

Miskolc, May 7, 2021

MISKOLCI EGYETEM

Műszaki Földtudományi Kar

Kőolaj és Földgáz Intézet

UNIVERSITY OF MISKOLC

Faculty of Earth Science & Engineering

Institute of Petroleum and Natural Gas H3515 Miskolc, Egyetemváros, HUNGARY

Tel: (36) 46 565 078

turzoz@kfgi.uni-miskolc.hu

www.kfgi.uni-miskolc.hu

Master Thesis

Project Assignment

For

Mr. Dániel Kelemen

MSc students in petroleum engineering

Title of Thesis:

Feasibility analysis and process simulation of high CO2 content gas mixture dehydration

using glycols for pipeline transportation

Tasks:

1. Create a comprehensive literature survey encompassing the investigation of the

relevant properties of CO2.

2. Detail and compare the different glycols used to remove water from hydrocarbon

gases and from CO2.

3. Investigate the different types of pipeline transport of CO2.

4. Create a model for every relevant investigated method in Aspen HYSYS.

5. Compare the different process models, carry out a case study to quantify the

differences between the different approaches.

6. Carry out an economical calculation to support the selection of the best water

removal technology.

Instructor: László KIS, assistant lecturer

Field Advisor: Róbert BÍRÓ, Chief Engineer (Production Engineering, MOL Hungary

Plc.)

Zoltán Turzó, PhD

Head of Institute

Miskolc, 2020. June 25.

Proof Sheet for thesis submission for Petroleum Engineering MSc students

Name of student: Daniel Kelemen Neptune code: EHOQSY

Title of Thesis: Feasibility analysis and process simulation of high CO2 content gas mixture dehydration using glycols for pipeline transportation

Declaration of Originality I hereby certify that I am the sole author of this thesis and that no part of this thesis has been published or submitted for publication. I certify that, to the best of my knowledge, my thesis does not infringe upon anyone’s copyright nor violate any proprietary rights and that any ideas, techniques, quotations, or any other material from the work of other people included in my thesis, published or otherwise, are fully acknowledged in accordance with standard referencing practices. 6th May 2021

Signature of the student

Statement of the Department Advisor1

Undersigned László KIS agree/ do not agree to the submission of this Thesis.

6th May 2021

Signature of

Department Advisor

Statement of the Industry Consultant2

Undersigned Róbert BÍRÓ agree/ do not agree to the submission of this Thesis.

6th May 2021

Signature of

Industry Consultant The thesis has been submitted: Date: 7th May 2021

Administrator of Petroleum and Natural Gas Institute

1 Thesis can be submitted regardless of the consultant’s consent. 2 If the student has no Industry Consultant delete paragraph as appropriate.

This Proof Sheet with the required signatures must made part of the Thesis on a page just after the Project

Assignment sheet.

ACKNOWLEDGEMENTS

First of all, I would like to thank my advisor, Róbert Bíró, for his professional

guidance, constant support through this work.

Equally I would like to express my appreciation to University of Miskolc and

the Petroleum and Natural Gas Institute for given me the opportunity of study

and preparing my thesis.

Special thanks to MOL Hungary Plc. for their help and supports that provided

to me, also for supporting this work with the most important data and software

usage.

Last but not least, I would like to thank you my friends and family. Without

their constant encouragement this work would not be possible.

Table of Content

List of Figures ....................................................................................................................... i

List of Tables ....................................................................................................................... iii

Nomenclature ...................................................................................................................... iv

Abbreviations ...................................................................................................................... vi

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

2. Theoretical Background .............................................................................................. 2

2.1. Fundamentals of Carbon Dioxide ....................................................................... 2

2.2. Thermophysical Properties of Carbon Dioxide ................................................. 4

2.3. Gas Hydrates ......................................................................................................... 9

2.4. Equations of State ............................................................................................... 16

2.5. Water Removal Processes .................................................................................. 19

2.6. Transportation .................................................................................................... 22

3. Case Study ................................................................................................................... 33

3.1. Model Building Principles .................................................................................. 34

3.2. Pressure Boosting and Cooling .......................................................................... 35

3.3. Case 1.1. – Dehydration with Glycol Technology ............................................ 38

3.4. Transportation of CO2 by pipeline .................................................................... 54

3.5. Case 1.2 – Inhibitor Injection to the Wet Stream ............................................ 58

3.6. Case 1.3 – No Dehydration nor Inhibition ........................................................ 60

3.7. Case 2. – Emitting the Carbon to the Atmosphere .......................................... 61

3.8. Economic Evaluation Sub-Cases ....................................................................... 63

4. Conclusion and Discussion ........................................................................................ 65

5. References ................................................................................................................... 67

Appendix A – Final Configuration of TEG Unit ........................................................ 70

Appendix B – Cost of Materials .................................................................................... 72

Appendix C – Additional Graphs ................................................................................. 73

i

List of Figures

Figure 1 Density of CO2 ....................................................................................................... 4

Figure 2 Viscosity of CO2 .................................................................................................... 5

Figure 3 Heat capacity of CO2 – tight interval ..................................................................... 6

Figure 4 Water content of CO2 ............................................................................................. 7

Figure 5 Pressure – Temperature diagram of CO2................................................................ 8

Figure 6 Effect of acid gases on hydrate formation temperature (CH4-CO2 and CH4-H2S

binary systems) .................................................................................................................... 10

Figure 7 Effect of acid gases on hydrate formation temperature of sweet natural gases ... 11

Figure 8 Typical dehydrator unit ........................................................................................ 20

Figure 9 Modified Beggs & Brill flow regime map ........................................................... 26

Figure 10 Simplified flowsheet .......................................................................................... 33

Figure 11 Pressure boosting (3 stages) ............................................................................... 35

Figure 12 Propane-based mechanical refrigeration unit ..................................................... 37

Figure 13 Cooling duty vs quantity of propane .................................................................. 37

Figure 14 Dehydrator unit .................................................................................................. 39

Figure 15 Water dew point vs contractor temperature ....................................................... 40

Figure 16 Absorber performance at lean TEG concentration of 98.50 wt% ...................... 41

Figure 17 Absorber performance at lean TEG concentration of 99.00 wt% ...................... 41

Figure 18 Remaining water content vs contractor temperature (TEG) .............................. 42

Figure 19 TEG loss vs contractor temperature ................................................................... 43

Figure 20 TEG mass fraction vs mass flow ........................................................................ 43

Figure 21 Remaining water content vs strip gas quantity (TEG) ....................................... 44

Figure 22 TEG mass fraction vs strip gas quantity ............................................................ 45

Figure 23 Reboiler duty vs TEG mass flow ....................................................................... 45

Figure 24 TEG mass fraction vs reboiler temperature with different strip gas quantities .. 46

Figure 25 Absorber performance at lean TEG concentration of 99.05 wt% ...................... 46

Figure 26 Dehydrator unit without heat exchanger ............................................................ 47

Figure 27 Cooling duty vs quantity of propane (relative) .................................................. 48

Figure 28 Effect of heat exchanger on reboiler duty (TEG) ............................................... 48

Figure 29 Remaining water content vs contractor temperature .......................................... 49

Figure 30 Glycol loss vs contractor temperature ................................................................ 49

Figure 31 Glycol mass fraction vs mass flow ..................................................................... 50

ii

Figure 32 Glycol mass fraction vs strip gas quantity ......................................................... 51

Figure 33 CO2 loss vs mass fraction of glycol ................................................................... 51

Figure 34 Number transfer unit conversion ........................................................................ 52

Figure 35 Height transfer unit conversion .......................................................................... 53

Figure 36 4th stage compression and transportation pipeline ............................................. 54

Figure 37 Process of pipeline sizing ................................................................................... 55

Figure 38 Hydrate formation conditions............................................................................. 58

Figure 39 Cost of sub-cases ................................................................................................ 64

Figure 40 Density of CO2 – high resolution ....................................................................... 73

Figure 41 Viscosity of CO2 – high resolution .................................................................... 73

Figure 42 Heat capacity of CO2 – wide interval ................................................................. 74

Figure 43 Mass enthalpy of CO2 ........................................................................................ 74

Figure 44 Mass entropy of CO2 .......................................................................................... 75

Figure 45 Pressure – Temperature diagram of gas mixture ................................................ 75

Figure 46 Water content of gas mixture (low temperature) ............................................... 76

Figure 47 Water content of gas mixture (high temperature) .............................................. 76

Figure 48 Reboiler duty vs strip gas quantity (relative) ..................................................... 77

Figure 49 Reboiler duty vs strip gas temperature (relative) ............................................... 77

Figure 50 Remaining water content vs contractor temperature (DEG) .............................. 78

Figure 51 DEG mass fraction vs strip gas quantity ............................................................ 78

Figure 52 Remaining water content vs strip gas quantity (DEG) ....................................... 79

Figure 53 DEG mass fraction vs mass flow ....................................................................... 79

Figure 54 Reboiler duty vs DEG mass flow ....................................................................... 80

Figure 55 DEG mass fraction vs reboiler temperature with different strip gas quantities . 80

Figure 56 Absorber performance at lean DEG concentration of 98.85 wt% ...................... 81

Figure 57 Remaining water content vs contractor temperature (TREG) ............................ 81

Figure 58 TREG mass fraction vs mass flow ..................................................................... 82

Figure 59 Remaining water content vs strip gas quantity (TREG) .................................... 82

Figure 60 TREG mass fraction vs strip gas quantity .......................................................... 83

Figure 61 TREG mass fraction vs reboiler temperature with different strip gas quantities 83

Figure 62 Absorber performance at lean TREG concentration of 97.85 wt% ................... 84

Figure 63 Glycol viscosity – wide interval ......................................................................... 84

Figure 64 Glycol viscosity – tight interval ......................................................................... 85

Figure 65 EU ETS allowance historical prices and forecast .............................................. 85

iii

List of Tables

Table 1 Fugacity relationships ............................................................................................ 12

Table 2 Physical parameters of glycols .............................................................................. 21

Table 3 Beggs & Brill flow regime boundary conditions ................................................... 27

Table 4 Horizontal liquid holdup calculation formulae ...................................................... 27

Table 5 β calculation formulae ........................................................................................... 28

Table 6 Composition of the input gas ................................................................................. 33

Table 7 Initial parameters of the input gas.......................................................................... 33

Table 8 Sub-cases ............................................................................................................... 34

Table 9 Pressure boosting parameters ................................................................................ 36

Table 10 MRU parameters .................................................................................................. 37

Table 11 ‘C’ values for maximum velocity calculation...................................................... 56

Table 12 Pipeline parameters – dehydrated gas .................................................................. 56

Table 13 Safety factors ....................................................................................................... 57

Table 14 Wall thickness parameters – dehydrated gas ....................................................... 57

Table 15 Inhibitor requirement for dense liquid phase ....................................................... 59

Table 16 Pipe data – not dehydrated gas ............................................................................ 60

Table 17 Determination of CO2 emission ........................................................................... 62

Table 18 CAPEX, OPEX, total costs and break-even allowance prices of the sub-cases .. 64

Table 19 TEG unit parameters ............................................................................................ 70

Table 20 Fixed costs ........................................................................................................... 72

Table 21 Specific costs ....................................................................................................... 72

Table 22 EU ETS allowance cost ....................................................................................... 72

iv

Nomenclature

A – area

a – attraction force between molecules

B(θ) – inclination factor

b – volume occupied by molecules

bara – absolute pressure

barg – gauge pressure

C – factor for contractor sizing

Cm – Langmuir constant

c0 – corrosion allowance (including

erosion)

c1 – negative tolerance

c2 – thinning allowance

D – diameter

d – depression of hydrate freezing point

e – wall thickness

F – design stress

Fr – Froude Number

f – friction factor

G – Gibbs energy

g – gravitational acceleration constant

(9.81 m/s2)

H – Henry’s constant

Hth – heat transfer coefficient

h – height

ID – inner diameter

K – vapor-liquid equilibrium constant

k – thermal conductivity

L – length

M – molar mass

m – mass

m – mass flow

Nvl – liquid velocity number

n – mole

p – pressure

OD – outer diameter

r – radius

q – heat transfer rate

R – universal gas constant (8.314J

molK)

Rth – thermal resistance

Re – Reynolds Number

ReH t – minimum specified value of upper

yield strength at calculation

temperature

Rp0.2 t – minimum 0.2% of proof strength

Rm – tensile strength at calculation

temperature

T – temperature

Ta – absolute temperature

V – volume

Vm – molar volume of the gas

v – velocity

vsl – superficial velocity

U – overall heat transfer coefficient

X – mass fraction in the liquid phase

x – mole fraction in the liquid phase

Y – mass fraction in the vapor phase

y – mole fraction in the vapor phase

Z – compressibility factor

Zb – depth of cover to centerline of pipe

z – joint coefficient

v

Greek symbols

α – binary interaction constant

β – coefficient in Beggs & Brill method

γ – activity coefficient

δ – correction factor

ϵ – relative roughness

εL – horizontal liquid holdup

εL(θ) – inclined liquid holdup

ϴ – fractional occupancy of guest

molecule

κ – heat capacity ratio

λ – volume factor

μ – dynamic viscosity

νm – number of cavities type m per water

molecule

ξ – chemical potential

ρ – density

σ – gas/oil interfacial tension

φ – fugacity

ω – acentric factor

Subscripts / Superscripts

a – analysis

amb – ambient

ap – additional packing

as – additional allow space

c – critical

calc – calculation

el – electrical

F – fluid

G – gas

H – hydrate

I – inhibitor

L – liquid

M – mixture

MT – empty hydrate lattice

ns – no-slip

ord – ordered

r – pseudo-reduced

th – thermal

tp – two-phase

V – vapor

W – water

vi

Abbreviations

CS – Carbon Steel

DEG – Diethylene Glycol

EG – Ethylene Glycol

EGR – Enhanced Gas Recovery

EOR – Enhanced Oil Recovery

EoS – Equation of State

GHG – Greenhouse Gas

GPA – Gas Processors Association

GSU – Gas Sweetening Unit

HC – Hydrocarbon

HTU – Height Transfer Unit

LNG – Liquified Natural Gas

MRU – Mechanical Refrigeration Unit

NGL – Natural Gas Liquid

NTU – Number Transfer Unit

ppmv – part per million in volume basis

PR – Peng-Robinson Equation of State

SRK – Soave-Redlich-Kwong Equation of State

STDm3 – GPA standard m3, volume of 1 mol gas at 101 325 Pa and 15 °C

TEG – Triethylene Glycol

thEUR – thousand euros

TREG – Tetraethylene Glycol

VLE – Vapor-liquid Equilibrium

1

1. Introduction

As the World’s energy demand is more and more increasing the production must follow

the tendency. In order to get energy (electricity and heat) today’s most common process is

to burn fossil fuels (coal, oil and gas). [1] [2] The products are energy and exhaust gases e.g.

water vapor, carbon dioxide (CO2), carbon monoxide (CO), nitrogen oxides (NOx) etc. The

physical and chemical properties make carbon dioxide a versatile substance, an important

raw material, on the other hand a contaminant in the atmosphere.

The sources – power plants, factories etc. – and the utilization points – chemical facilities,

oil and gas fields, agricultural facilities etc. – are rarely found close to each other, transport

may be required. For long distance transportation of gases, pipelines are the best practice.

With water, CO2 forms weak acid (carbonic acid, H2CO3), which harms the equipment.

In order to avoid the degradation, one can use different – active or passive – methods.

For high flowrates, usage of chemical inhibitors is not economical. Coating can protect the

pipelines, but not the pressure boosting and other equipment. Stainless steel can withstand

the corrosion, however using these materials can increase the cost significantly. By removing

the water from the mixture, thus eliminating one of the key components, one can get rid of

several problems: corrosion, chance of hydrate formation and liquid droplets in the

compressor. Then, the dry carbon dioxide can be transported and utilized safely. In addition,

total volume flow and transportation costs are reduced, storage efficiency is increased.

There are several methods in the industry for gas dehydration, but in most cases these

methods were engineered for natural gas mixtures, they are CO2 content sensitive. The most

fundamental method of gas processing – cold separation – cannot be used. It is not only

energy demanding, but the CO2 can freeze out, hydrate can form. Physical processes such as

molecule sieves, and some physical absorbents, like glycols or glycerin are not CO2 sensible.

The aim of this work is to investigate the possibilities of a glycol-based dehydration and

transportation system of a high CO2 content gas mixture. To compare the different available

technologies by modelling the operation of these units, using Aspen HYSYS® and methods

presented in open literature. Finally, to present – expenditure sensitive – suggestion for the

described system, treating a sour gas mixture leaving an existing – amine based – gas

sweetening unit (GSU) and transporting it to the utilization point, a hydrocarbon field for

enhanced oil recovery, enhanced gas recovery (EOR/EGR). The source of the gas mixture

and the reservoir are not in the scope of this work, the investigation of the GSU and the

reinjection are excluded from this work.

2

2. Theoretical Background

2.1. Fundamentals of Carbon Dioxide

Carbon dioxide (CO2) is the fully oxidized form of carbon, it is a colorless, non-

flammable gas. In low concentrations the gas is odorless; however, at sufficiently high

concentrations, it has sharp sour odor. Carbon dioxide is non-toxic, however can cause

asphyxia. It is heavier than dry air, it’s molecular weight is 44.01 g/mol, it has a specific

gravity of ~1.52. [3] At normal temperature, gaseous carbon dioxide is not reactive. Its

molecule is relatively stable and does not easily break down into simpler compounds.

It occurs naturally as a trace gas in Earth's atmosphere, currently the most abundant

greenhouse gas. Its role in the Earth’s heat balance is outstanding, but not fully understood

today (~20% of the greenhouse effect). [4]

Carbon dioxide was in the scope of the scientific interest for a long time, its physical

properties are extensively researched. In the XIX. century the carbon dioxide was mainly

used as a working fluid in refrigerators. Faraday successfully liquefied CO2 in 1823, thus

began the industrial development. Thilorier followed his work on large scale, in 1835

succeeded in carrying out extensive experiments on the expansion, vapor pressure, density,

and enthalpy changes of the liquid CO2 during evaporation. He was the first who could

produce solid carbon dioxide called “dry ice” as well. In 1884 Raydt established a facility to

produce liquid CO2, in 1889 the production of CO2 from flue gasses began in Berlin. [5]

In the XX. century the research turned to the supercritical and to the dense phase regions.

The ambition behind this change was the intended application of supercritical CO2 as a

cooling medium in nuclear power plants in the early 50’s. Furthermore, supercritical carbon

dioxide is proved to be a suitable solvent and an outstanding tool in enhanced oil recovery.

Today’s scientific interest focuses on the greenhouse effect of CO2 and the influence of

anthropogenic carbon on the carbon cycle. These topics are not fully understood yet. [7] [8]

Natural sources of CO2 include volcanoes, hot springs and geysers, and it is freed from

carbonate rocks by dissolution in water and acids. Carbon dioxide is water-soluble, it

naturally occurs in groundwater, rivers and lakes, ice caps, glaciers and brine. It is present

in deposits of petroleum and natural gas.

There are two main sources of carbon dioxide in the atmosphere, the carbon cycle and

human activity. The first one is a slow biogeochemical cycle which is a complex carbon

exchange process between the atmosphere, oceans, soil, rocks and the biosphere.

3

Atmospheric carbon can be reduced by living creatures. Plants and other photoautotrophs

(e.g. plankton or algae) use the energy of the Sun to produce carbohydrates from atmospheric

carbon dioxide and water through photosynthesis. With water, CO2 forms weak acid

(carbonic acid, H2CO3), that can dissolve rocks. Metal ions (mainly calcium, magnesium

and potassium) transported to the oceans, where combined with bicarbonate ions form

calcium carbonate, one of the most important ingredients of shell (through diagenesis it

becomes limestone). CO2 is emitted to the atmosphere by volcanos, or by living creatures

burning sugar to get energy. [9]

The anthropogenic carbon dioxide is the result of the human activity. Burning fossil fuel

– coal, oil and natural gas – to get electricity or heat; transportation; agriculture; animal

husbandry and industrial activities are all emitters. Deforestation has double effect on

atmospheric CO2 concentration, not only by reducing the CO2 removal potential, but

reducing the carbon containing biomass at the same time.

Several international conventions and protocols have been formulated to try to reduce

CO2 emission such as, the Intergovernmental Panel on Climate Change in 1988,

the UN Framework Convention on Climate Change in 1992, Kyoto Protocol in 1997,

the EU Emission Trading System in 2005 and the Paris Agreement in 2016. To date, there

is no worldwide agreement on these laws, and many countries and industries do not comply

with these conventions. In general, the anthropogenic CO2 emission can be controlled by

reducing energy intensity, limiting carbon intensity, or improving CO2 sequestration. [4]

Carbon dioxide is a versatile industrial material, used for example, as an inert gas in

welding and fire extinguishers; as a pressurizing gas in air guns; in petroleum recovery; as a

supercritical fluid solvent to replace organic solvents in decaffeination of coffee and

supercritical drying. It is added to drinking water and carbonated beverages including soda

or beer. Dry ice is used as a refrigerant and as an abrasive in dry-ice blasting. It is a feedstock

for fuel and chemical (fertilizers, plastics, and polymers) synthesis. The electronics industry

uses carbon dioxide in manufacturing applications, including semiconductor device

manufacturing, surface cleaning, and circuit board assembly. It is also used to purge,

pressurize and cool equipment. In water and soil treatment, carbon dioxide is safer to handle

than mineral acids for pH control. Carbon dioxide is used in healthcare as well. [4] [10]

Transport of any substance is more economical in liquid form if the required conditions

are reasonable from engineering point of view (e.g. LNG). For long range transport on land

the pipeline is the most cost-efficient way. CO2 is transported in liquid form or more often

in its dense phase, however in some cases gaseous form is more economical. [6]

4

2.2. Thermophysical Properties of Carbon Dioxide

Thermophysical properties are the governing factors of every reaction of the substance.

Knowledge of these properties is essential to understand its physical and chemical processes.

The principles of thermodynamics are widely applicable in correlating and predicting the

properties of mixtures. The properties of greatest interest are density, viscosity and heat

capacity of gases and liquids; heat of vaporization; and the effects of pressure on the

enthalpies of compressible fluids. Enthalpy and entropy charts of pure substances and

mixtures enable prediction of temperature change when gases are expanding or being

compressed. [6]

Heat capacity, enthalpy, and entropy are the important thermodynamic parameters of CO2

mixtures, because they affect the heat transfer and energy consumption of compression and

purification processes. Density and viscosity are important for transportation calculations,

as they have significant effect on pressure drop. [11]

2.2.1. Density, Viscosity and Heat Capacity

Figure 1 – 3 show the density, viscosity and heat capacity of CO2, enthalpy and entropy

graphs can be found in Appendix C. These graphs were generated with the property table

tool of Aspen HYSYS®.

Figure 1 Density of CO2 Own edit

5

Density and viscosity show similar behavior with pressure and temperature. Below the

critical point, more significant change at lower pressures are observable, as the CO2 changes

from gas to liquid phase. This phenomenon presented by the 0 °C isothermal lines.

Figure 2 Viscosity of CO2 Own edit

In the vicinity of the critical point (20 °C and 40 °C isothermals) a small change in

pressure have high impact on the given property. Figure 40 and 41 show the density and

viscosity with higher resolution in the vicinity of the critical point. On these graphs the two-

phase region can be seen more accurately as well.

The temperature increment reduces the effect of pressure on density and viscosity.

Above the critical temperature, as the distance between the system parameters and the

critical point increases, the effect of pressure even less significant. Furthermore, as the

temperature increases the difference between the parameters belonging to the isothermal

lines reduces. It is because the fact, that the substance change phase from gas to supercritical

phase. This trend is true in case of heat capacity, entropy and enthalpy as well.

Heat capacity on the other hand does not follow this behavior. It has a maximum point at

the critical point, but similarly to density and viscosity, as the temperature increases the

effect of pressure become less significant. This can be explained with the phase change from

dense liquid to supercritical phase. If the temperature is above the critical, this change will

not occur.

6

Figure 3 Heat capacity of CO2 – tight interval Own edit

2.2.2. Water Content

Correct determination of water content of mixtures is essential when one designs

dehydration systems (particularly TEG systems). To meet extremely low water dewpoint

specifications, it is necessary to determine the water content in equilibrium with a hydrate.

There are several methods to determine the water content of a pure gas or gas mixture.

One should keep in mind that these correlations are valid only in certain ranges of

temperature, pressure and composition. Correlations such as McKetta & Wehe and Olds,

et.al are applicable only for lean, sweet natural gasses. Even the Wichert & Wichert method

is limited to 50 H2S equivalent mol% (~71 mol% of CO2, if there is no H2S in the mixture).

The H2S equivalent mol% can be obtained by multiplying the CO2 mol% by 0.70 and adding

the initial H2S mol% to it.

mol%H2S(equiv) = 0.7mol%CO2 +mol%H2S(initial) (1)

The saturated water content of a gas depends on pressure, temperature, and composition.

The effect of composition increases with pressure. Acid gas components – carbon dioxide

and hydrogen sulfide – increase the solubility of water in natural gas due to the attraction of

water for these molecules. The equilibrium water content of an acid gas mixture varies

significantly with pressure, temperature and mixture composition. [11]

7

Figure 4 Water content of CO2 Source: [11]

Figure 4 shows the water content of pure CO2, compared to methane at different

temperatures. At low pressure, the water content of CO2 decreases. At higher pressures the

water content increases with increasing pressure, as density and attraction of water of CO2

increases as well. At the vicinity of the critical point (31.04 °C, 73.83 barg) the

thermophysical properties of carbon dioxide changes significantly with small change in

pressure or temperature. Thus, the water content shows significant changes as well. 31.1 °C

and 50 °C isothermals represent this effect. [11]

From Figure 4 one can observe that, the saturated water content of pure CO2 significantly

higher, than the water content of the sweet natural gas. Furthermore, a small amount of acid

gas does not have a strong effect on the water content. However, low concentration of

methane can significantly reduce the water content of CO2. [11]

2.2.3. CO2 in Different Phases

As mentioned above, the research on carbon dioxide has a long history, most of its

physical properties are well known. In order to understand these changes,

the pressure-temperature (p-T) diagram (Figure 5), the temperature-entropy (T-s) diagram

and the pressure-enthalpy (p-H) diagram (Figure 43) are the best tools.

Phase diagrams, especially the p-T diagram have different form in case of a pure

substance, a binary system or a multicomponent system. A mixture’s p-T diagram is called

the phase envelope diagram. In the envelope there is multiphase system, outside the envelope

only a single phase can exist.

8

The lines in the p-T diagram represent conditions where two phases coexist and are the

boundaries for single phase regions. All three lines meet at the triple point. The vaporization

curve ends at the critical point, which corresponds to the highest pressure and highest

temperature at which pure chemical species can exist in vapor-liquid equilibrium. [6]

The pressure – temperature diagram is the simplest phase diagram of a pure substance.

This diagram shows the pressure and temperature boundaries of the different phases. The

lines (equilibrium curves) represent the phase boundaries, the conditions where two phases

(solid-gas, gas-liquid and solid-liquid) are coexist, the phase transitions occur along these

lines. It does not provide any information about volume properties. On a pressure-volume

(p-V) diagram these boundaries become areas, i.e. regions where two phases coexist.

Figure 5 Pressure – Temperature diagram of CO2 Based on: [12]

Below the triple point (-56.57 °C, 5.16 barg) liquid carbon dioxide cannot exist. The gas

freezes, the solid CO2 evaporates without liquid phase (sublimates). At the triple point the

three phases are in equilibrium. Even a little change in temperature or pressure disturbs the

balance; freezing, melting, solidification, sublimation, evaporation and condensation can be

observed simultaneously.

The critical point is the end point of an equilibrium curve. Substances have more critical

points. From these, the gas-liquid critical point (31.04 °C, 73.83 barg) was first discovered;

this point is the most investigated. In the vicinity of the critical point, the physical properties

of the phases change dramatically, with both phases becoming ever more similar.

At the critical point the phase boundary disappears. Above the critical pressure and

9

temperature, a state of matter exists that is continuously connected with (can be transformed

without phase transition into) both the liquid and the gaseous state, it is called supercritical

fluid. The region above the critical pressure, but below the critical temperature is called

dense or dense liquid phase.

In general, most Equation of State display different behavior around the critical point.

It is understandable, since the formation of the supercritical phase is alike to a phase

transition, with the generated phase bearing mixed characteristics of liquid and gaseous

phases, e.g. the material has similar density to the liquid phase, but similar viscosity to the

gas phase. The essence of the phenomena is that the difference in the density of the liquid

and gas phases decreases to that extent that it becomes equal. High above the critical point

the functions are steady, as is the supercritical phase itself. However, in the vicinity of the

critical point the phase bears unsteady characteristic: most of the state equations displays

sharp changes as a response for a relatively small change in the temperature or pressure.

In dense or dense liquid region there is a continuous progression from gas to liquid,

without a distinct phase change. In this region the density of carbon dioxide increases with

decreasing temperature. Today’s pipelines are operated in this region. [13]

2.3. Gas Hydrates

Clathrate hydrate or gas hydrate is a solid, “ice-like”, white, flammable material.

A physical combination of water and other small molecules, there is no chemical bond

between the particles. Their formation in gas and/ or NGL systems can plug pipelines,

equipment, and instruments, restricting or interrupting flow.

There are three recognized crystalline structures for such hydrates. In all cases water

molecules build the lattice and hydrocarbons, N2, CO2 and H2S occupy the cavities.

Smaller molecules (CH4, C2H6, CO2, H2S) stabilize a body-centered cubic called

Structure I. Larger molecules (C3H8, i-C4H10, n-C4H10) form a diamond-lattice called

Structure II. Normal paraffin molecules larger than n-C4H10 do not form Structure I and II

hydrates as they are too large to stabilize the lattice. However, some iso-paraffins and cyclo-

alkanes larger than pentane, are known to form Structure H hydrates.

Gas composition determines structure type. Mixed gases will typically form Structure II.

Limiting hydrate numbers (ratio of water molecules to molecules of included gaseous

component) are calculated using the size of the gas molecules and the size of the cavities in

H2O lattice. [14]

10

From a practical point of view, the structure type does not affect the appearance,

properties, or problems caused by the hydrate. It does, however, have a significant effect on

the pressure and temperature at which hydrates form. Structure II hydrates are more stable

than Structure I. This is the reasons why gases containing propane and isobutane will form

hydrates at higher temperatures than similar gas mixtures without these components. [11]

2.3.1. Formation of Hydrates

Gas hydrates form at relatively high pressure and low temperature. In the nature these

conditions can be found in environments such as deep-water continental margins and

under the permafrost of arctic regions. Gas hydrates can form anywhere in a pipeline or

process stream, but they are particularly likely to form downstream of orifices or valves due

to Joule-Thomson expansion effects. [15]

At a given pressure, pure methane forms hydrates at the lowest temperature. Adding

ethane through butane, hydrogen sulfide, or carbon dioxide raises the formation temperature

significantly. However, nitrogen will lower the hydrate formation temperature of the

mixture. Figure 6 and 7 show the effect of CO2 and H2S on hydrate formation temperature

in binary and multicomponent systems.

Figure 6 Effect of acid gases on hydrate formation temperature (CH4-CO2 and CH4-H2S

binary systems) Own edit

The presence of H2S in multicomponent mixture or in a CH4-H2S binary system results

in a substantially higher hydrate formation temperature at a given pressure. CO2 in general,

has smaller impact. In a methane - carbon dioxide binary system with the increasing CO2

content the hydrate formation temperature increases as well, however in a natural gas

mixture it slightly reduces the hydrate formation temperature. [15]

11

Figure 7 Effect of acid gases on hydrate formation temperature of sweet natural gases Own edit

2.3.2. Hydrate Prediction Methods

Prediction of hydrate formation conditions considered significant; the literature of this

topic is extensive. A common point of prediction models that, they are based on the ratio of

hydrate formation components. If the conditions – pressure, temperature and composition –

are favorable, and there is any water in the system, hydrate will form. The quantity of water

only effects the quantity of hydrate formed. There are several methods to predict hydrate

formation e.g. p-T or hydrate phase diagrams, three-phase equilibrium calculations,

distribution coefficient method etc.

Since the main drive factor was the natural gas industry, most of the researchers focused

on sweet natural gases – like Katz or McLeod & Campbell method. The Baille & Wichert

method deals with the effect of H2S only. These methods are not applicable for high CO2

content gases. [11]

2.3.2.1. Ng & Robinson Method

Ng & Robinson made their experiment with different mixture including hydrocarbons

– e.g. methane, isobutane etc. – and inert components – e.g. CO2 and N2. They modified the

Parrish-Prausnitz method, determined new Kihara parameters and changed the applied EoS

to Peng-Robinson. Their results can be classified as 2-Phase – where no aqueous water

present – and 3-Phase – where aqueous water is present – models. [16] [17] [18]

In the 2-Phase model there is no aqueous water in the system – vapor only, HC liquid

only, vapor-HC liquid, HC liquid-HC liquid and vapor-HC liquid-HC liquid systems.

12

The general fugacity equation of the 2-Phase model can be written as:

φWMT = φW,0

MT + (dφW

MT

dp)T

p (2)

Combining Eq. 2 with the linear regressed plots presented in [16], the fugacity of water

over the unfilled hydrate lattice as a function of temperature and pressure can be obtained.

Table 1 summarizes the relations. These relationships depend on hydrate structure but

independent of the composition. [16]

Table 1 Fugacity relationships

Hydrate type φW,0MT

dφWMT

dp

Type I 14.269 −5 393

Ta 0.00036Ta − 0.1025

Type II 18.062 −6 512

Ta 0.0001109Ta − 0.03192

After the fugacities are calculated the chemical potential difference can be determined as:

ΔξWH = ln(

φWRTa

fWMT

) (3)

The 3-Phase model can be applied if a free aqueous phase is present in the system

– aqueous only, vapor-aqueous, HC liquid-aqueous, and vapor-HC liquid-aqueous systems.

This model based on the van der Waals-Platteuw equation, which relates the thermodynamic

properties of the hydrate to their components as:

ΔξWH = ξW

MT − ξW = RTa∑νm ln (1 +∑Cmjφjj

)

m

(4)

Eq. 4 was not applicable if liquid water was in the system. Ng & Robinson modified the

expression as: [17]

ΔξWL = RTa [∏{1 + 3(αj − 1)yj

2 − 2(αj − 1)yj3}

j

] ×

[∑ νmm ln(1 + ∑ Cmjφjj ) + ln(xW)]

(5)

13

2.3.2.2. CSM Method

The Colorado School of Mines method is based on the SRK EoS and statistical

thermodynamics. This model is applicable for all possible scenarios – vapor-hydrate,

HC liquid-hydrate, aqueous-hydrate, vapor-HC liquid-hydrate, vapor-aqueous-hydrate, and

vapor-HC liquid-aqueous-hydrate systems.

It differs from most of the published method, as they use the van der Waals-Platteeuw

hydrate equation (Eq. 4) coupled with an EoS model. The standard state used in these

equations is a hypothetical empty hydrate lattice. The CSM model uses an alternative

derivation of these equations and a different standard state. It also uses a new aqueous phase

model tailored for the presence of water and mixed inhibitors. This model relies on the

equilibrium equivalence of the fugacity of water in the hydrate phase to that of water in the

fluid (aqueous, liquid hydrocarbon, or vapor) phase: [19] [20] [21]

φWH = φW

F (6)

For fugacity and chemical potential calculation they derived a different relation:

φWH = φW0e

ξWH −GW0RTa

(7)

and

ξWH = GW

MT + RTa∑νm ln (1 −∑Θimi

) + RTa ln(γWH )

m

(8)

2.3.3. Prevention of Hydrate Formation

The main hazard of hydrates, that they can reduce the effective cross-section area, thus

restricting the flow. This is a virtuous cycle, as the hydrate forms on the wall of the pipe the

cross-section area reduces, the velocity increases, the temperature reduces. This temperature

reduction enhances the hydrate formation, finally it can plug the pipeline. To make a plugged

pipeline operable again it takes time, heating and/ or chemicals. Even if the pipeline is not

plugged, the solid particles can harm the equipment – e.g. compressors, separators, etc.

Avoiding hydrate formation is a top priority in the hydrocarbon industry. Prevention has

several ways:

- Maintaining the system temperature above the hydrate formation temperature, by

heating or insulation

- Removing the water from the system, which solves the corrosion problems as well

- Injection of chemical inhibitors [11]

14

Dehydration and inhibition can be achieved by several ways, Chapter 2.3.3.1 and 2.5

detailing these topics. The most feasible and economical solution is different in every case,

the decision-making process should include an extensive investigation.

2.3.3.1. Inhibition

Heating, isolation or dehydration are not always feasible. In these cases, chemical

inhibition can be an effective method of preventing hydrate formation. This solution requires

minimal investment, the expenditures depend on the volumetric flow rate of the gas, on the

composition of gas mixture and the time interval of application. Chemical inhibition utilizes

injection of thermodynamic inhibitors (equilibrium inhibitors) or low dosage hydrate

inhibitors (LDHIs).

Thermodynamic inhibitors are the traditional inhibitors (e.g. glycols or methanol), they

lower the hydrate formation temperature. LDHIs are either kinetic hydrate inhibitors (KHIs)

or antiagglomerants (AAs). They do not lower the temperature of hydrate formation but do

diminish its effect. KHIs lower the rate of hydrate formation, which inhibits its development

for a defined duration. AAs allow the formation of hydrate crystals but restrict them to sub-

millimeter size. In the following only thermodynamic inhibitors are detailed.

2.3.3.1.1. Thermodynamic Inhibitors

Thermodynamic inhibitors should be injected to the process stream, where they can

combine with the condensed aqueous phase to lower the hydrate formation temperature at a

given pressure. With simple separation, the water with the inhibitor can be removed from

the process stream and after regeneration the inhibitor can be reused.

Beside methanol, ethylene glycol, diethylene glycol, and triethylene glycol have been

used for hydrate inhibition. The most popular is ethylene glycol because of its lower cost,

lower viscosity, and lower solubility in liquid hydrocarbons. DEG and TEG used in

dehydration rather than inhibition. Hydrate inhibition with methanol or EG is widely used.

The choice of the inhibitor used is influenced by several factors. The advantages of

methanol are the followings: it costs less than EG; requires lower concentrations in the

aqueous phase; applicable at very low temperatures; has lower viscosity than EG; during

regeneration, contaminants in the water phase leave with the water, not with the methanol;

can be transported in vapor phase. The disadvantages are: higher inhibitor losses; more

difficult to recover methanol from the aqueous phase; more toxic than EG; more flammable

than EG (has a lower flash point).

15

Advantages of EG are: very low solubility losses to the hydrocarbon phases; easier to

regenerate from water phase; less toxic and less flammable than methanol; can also provide

corrosion inhibition for “top of the line corrosion” in pipelines. The disadvantages are:

higher concentrations required in the aqueous phase; higher viscosity makes physical

separation from hydrocarbon liquid phase more difficult; transported in liquid phase, not

very effective in removing hydrates from downstream of the injection point; during

regeneration, contaminants in the water phase accumulate in the EG phase, requiring special

regeneration designs.

2.3.3.1.2. Determination of Required Inhibitor Injection

The required depression of freezing point can be calculated by Eq. 9, where KH empirical

constant is 1 297 for ethylene glycol and methanol.

d =KHXI

MI(1 − XI) (9)

Rearranging Eq. 9 to the minimum inhibitor concentration in the free water phase can be

approximated as:

XI =dMI

KH + dMI (10)

Eq. 9 and 10 should not be used beyond 20 wt% for methanol and 50 wt% for glycols.

For methanol concentrations up to about 50 wt%, Eq. 11 provides better accuracy.

d = −72 ln(xW) (11)

Eq. 12 and 13 can be used to convert between mass and mole percent.

mol%I =

wt%I

MI

(wt%I

MI +100 − wt%I

Mw)100 (12)

wt%I =mol%IMI

mol%IMI +MW(100 − mol%I)100 (13)

Once the required inhibitor concentration has been calculated, the mass of lean inhibitor

solution required in the water phase may be calculated from Eq. 14.

mI =XrichmW

Xlean − Xrich (14)

The amount of inhibitor injected must be sufficient to prevent freezing of the water phase,

and has to cover the vaporization and solubility losses of the inhibitor as well. For methanol,

the vapor pressure is sufficiently high, a significant quantity of inhibitor will be lost to the

vapor phase. [11]

16

2.4. Equations of State

Equation of State (EoS) is a thermodynamic equation relating state variables which

describe the state of matter under a given set of physical conditions, such as pressure, volume

and temperature (PVT), or internal energy. Equations of State can also describe solids,

including the transition of solids from one crystalline state to another.

In a practical context, Equations of State are instrumental for PVT calculations in process

engineering problems, such as petroleum gas/ liquid equilibrium calculations. A successful

PVT model based on a fitted EoS can be helpful to determine the state of the flow regime,

the parameters for handling the reservoir fluids, and pipe sizing.

The most famous EoS is the Ideal Gas Law or General Gas Equation:

pV = nRTa (15)

It was first stated by Benoit Paul Émile Clapeyron in 1834. It assumes that the gas

molecules are perfect spheres, they are not taking up space, the collisions are perfectly

inelastic – there is no change in kinetic energy –, there is no attraction or repulsion between

the molecules. This model is roughly accurate for weakly polar gases at low pressures and

moderate temperatures but becomes increasingly inaccurate at higher pressures and lower

temperatures and fails to predict condensation from a gas to a liquid.

In the engineering practice, for short calculations with gases a modified gas law

(Engineering Gas Law or Real Gas Law) is generally used.

pV = ZnRTa (16)

The ‘Z’ compressibility or gas deviation factor is an empirical, pressure, temperature and

quality dependent constant. It describes the deviation of a real gas from ideal gas behavior.

It is simply defined as the ratio of the molar volume of a gas to the molar volume of an ideal

gas at the same conditions (temperature and pressure).

There are several new EoS models from different authors in the literature. These models

can describe reality more accurately, but most of them require extensive and complex

calculations or measurements. However, at present there is no single Equation of State, that

accurately predicts the properties of all substances under all conditions.

In 1873 van der Waals was the first who modified the Ideal Gas Law. He introduced two

constants ‘a’ and ‘b’. The subsequent models are based on his work. The van der Waals

Equation can be written as:

(p + a1

Vm2) (Vm − b) = nRTa (17)

17

2.4.1. Soave-Redlich-Kwong Equation of State

Redlich and Kwong introduced an empirical equation in 1949. This model is more

accurate than the van der Waals at temperature above critical temperature and not as complex

as the Benedict-Webb-Rubin model. Since that time, numerous modified Redlich-Kwong

equations have been proposed: Redlich and Dunlop in 1963; Chueh and Prausnitz in 1967;

Wilson in 1969; Zudkvitch and Joffe in 1970; and others. Some have introduced deviation

functions to fit pure substance PVT data while others have improved the equation's capability

for vapor-liquid equilibrium (VLE) predictions. [22] [23]

Modification made on the Redlich-Kwong model by Soave, in 1972 resulted in the

Soave-Redlich-Kwong (SRK) equation. A relatively simple equation, but it is capable of

generating reasonably accurate equilibrium ratios in VLE calculations. On the other hand, it

fails to generate satisfactory density values for the liquid even though the calculated vapor

densities are generally acceptable. While the Peng-Robinson formula focuses on

hydrocarbon system, the SRK model can generate more accurate results in case of sour, or

acidic gas mixtures.

The original Redlich-Kwong equation can be written as:

p =

RT

Vm − b−

a

√TaVm(Vm + b)

(18)

Soave replaced the a

√T term by a(T), which is more temperature dependent. The Soave-

Redlich-Kwong equation can be written as: [24]

p =RTa

Vm − b(T)−

a(T)

Vm(Vm + b(T)) (19)

Where

a(T) = 0.42747δR2Tc

2

pc (20)

b(T) = 0.08664RTcpc

(21)

δ = (1 + (0.48508 + 1.55171ω − 0.15613ω2)(1 − √Tr))2

(22)

2.4.2. Cubic-Plus-Association Equation of State

The CPA Equation of State combines the Soave-Redlich-Kwong equation with

association terms to describe the polar/ association effect. The CPA EoS, proposed by

Kontogeorgis et al. can be expressed as: [25] [26]

18

p =RTa

Vm − b(T)−

a(T)

Vm(Vm − b(T))−RT

2 Vm(1+

1

Vm

∂ln g

∂ (1Vm))∑xi

i

∑(1− xAi)

Ai

(23)

xAi represents the fraction of association sites that do not form bonds with other active

sites:

xAi =

1

1 + (1Vm)∑ xjj ∑ XBj∆

AiBjBj

(24)

∆AiBj= g(Vm) [e𝜁AiBj

RT − 1] bijβAiBj (25)

Where ∆AiBj describes association strength between site ‘A’ on molecule i and site ‘B’ on

molecule j. In the association strength ∆AiBj term ζAiBj and βAiBj represent cross-association

energy and effective cross-association volume, respectively. g(Vm) is the radial distribution

function for the reference fluid defined as:

g(Vm) =1

1 − 1.9η (26)

Where:

η =b

4Vm (27)

2.4.3. Peng- Robinson Equation of State

Peng and Robinson introduced their semi-empirical model in 1976. It was developed to

handle both vapor and liquid properties near equilibrium conditions. The development of

this equation was focused on natural gas systems. This model is simple to use and accurately

represent the vapor pressures of pure substances and the relationships temperature, pressure,

and phase compositions in binary and multicomponent systems, therefore it is widely used

to the present day in the industry. The Peng-Robinson equation can be written as: [22]

p =RTa

Vm − b(T)−

a(T)

Vm(Vm + b(T)) + b(Vm − b(T)) (28)

Where

a(T) = 0.45724δR2Tc

2

pc (29)

b(T) = 0.07780RTcpc

(30)

δ = (1 + (0.37464 + 1.54266ω − 0.26992ω2)(1 − √Tr))2

(31)

19

2.4.4. Wilson Sour Water Equilibrium Model

Wilson’s model is very similar to the Van Krevelen model, but he removed some of the

limitations and expanded the temperature range within the model is accurate. The

assumption that H2S and CO2 only exist in aqueous solutions as ionized species made by

Van Krevelen is not true when acid gases are present in the absence of NH3 or other basic

components. This method considers the chemical equilibrium between ionic species of H2S

or CO2 and undissociated H2S or CO2 in the liquid. However, it leaves out the consideration

of the equilibrium between dissolved CO2 and carbonic acid (H2CO3), because the presence

of other acidic or basic component does not affect this equilibrium. By this method, the

partial pressure of H2S or CO2 in the vapor phase above a solution can be calculated from

the concentrations of the undissociated species as: [27]

ppartialH2S= HH2SCH2S (32)

ppartialCO2= HCO2CH2CO3 (33)

2.5. Water Removal Processes

Removing the water from the gas mixture is a key process in order to prevent corrosion,

erosion and hydrate formation. It has not only safety benefits, but additional cost reduction

can be achieved in the transportation, and in case of underground storage the total effective

storage capacity is increased. The most basic method is cold separation, which consist of

cooling, water condensation and separation. The other methods are adsorption – the adsorber

is solid –, absorption – the absorber is fluid – or application of membranes, molecular sieves.

In case of sour gases cryogenic separation is not recommended because of hydrate formation

or CO2 freezing.

2.5.1. Dehydration by Liquid Desiccant

In natural gas absorption dehydration systems, the solvent should be hygroscopic,

noncorrosive, non-volatile, easily regenerated to high purity, insoluble in liquid

hydrocarbons and unreactive with hydrocarbons, CO2 and sulfur compounds.

The glycols are close to meet these criteria. EG, DEG, TEG and TREG all possess suitable

traits. TEG is the most commonly applied glycol in dehydration as it has low vapor pressure

(low carry-over losses) and regenerable to high concentrations. TREG has lower vapor

pressure and withstands higher regeneration temperatures, but it is more viscous and

expensive, DEG is preferred in colder climates due to its lower viscosity. EG is more

commonly used in hydrate inhibition than dehydration. [15] [28]

20

Figure 8 shows a typical absorption-based dehydrator unit. This standard configuration

is applied regardless of the quality of the absorber.

Figure 8 Typical dehydrator unit Own edit

The wet gas first goes to an inlet scrubber to remove all the remaining liquid water.

Then, the gas enters the bottom of the absorber tower (contractor) and flows countercurrent

to the lean absorber. The dry gas leaves the tower on the top, while the rich absorber

(with the removed water) on the bottom. The rich absorber enters the flash tank, where most

of the volatile components vaporize, while the liquid leaves on the bottom and goes to the

regenerator unit (regenerator tower and reboiler), where the absorber reaches the required

concentration by distillation. The distillation temperature cannot be higher, than the

decomposition temperature of the absorber, and it is practical to be as low as possible in

order to reduce the carry-over losses. The water vapor leaves on top from the condenser. The

lean absorber leaving the reboiler, goes to the pump and enters the dehydrator tower on top,

thus closing the loop.

21

Application of the lean/rich heat exchanger and the dry gas/ lean absorber heat exchanger

is not essential, but significantly improves the energy efficiency of the gas treating process.

Without these apparatus additional heating and cooling is needed. Stripping gas is also

optional, but if high purity absorber is needed, application of stripping gas is necessary. The

absorber tank has no effect on the process but an important part of the whole system.

Another common configuration is, when instead of reflux, the rich absorber after leaving

the dehydrator tower goes to the top of the regenerator tower, acts as a cooling medium and

only then enters the flash tank.

Operating conditions for glycol units are governed principally by the degree of

dehydration required, the physical properties of the glycol solutions, and the inlet pressure

of the gases. Lower temperatures and high pressure enhance absorption capacity but can lead

to hydrate formation. Typical operation conditions are: 30 - 70 barg and 16 - 40 °C. [29]

2.5.1.1. Physical Properties of Glycols

As it was mentioned above, glycols are favorable substances for absorption, DEG, TEG

and TREG are widely applied. In case of natural gas dehydration TEG is the most common

choice. EG is more common in low temperature technologies – because its low viscosity –

and on offshore platforms – it can hold more salt than the other glycols. [11]

Glycols have some common properties e.g. they are highly hydrophile, react with oxygen

and form corrosive acidic compounds; the products increase the potential of foaming and

glycol carryover. Low pH accelerates glycol decomposition. They do not freeze solid, but

form a dense, highly viscous solution. Their decomposition temperatures are lower, than

their boiling points. However, the difference in their physical properties decide their

application in a dehydration. The most important parameters are listed in Table 2. [15] [29]

Table 2 Physical parameters of glycols Based on: [11], [15]

EG DEG TEG TREG

Decomposition temperature at 1 atm [°C] 165 164 208 238

Maximum suggested regeneration temperature [°C] 160 160 204 234

Density [kg/m3] at 25 °C

at 60 °C

1 110

1 085

1 113

1 088

1 119

1 092

1 120

1 092

Viscosity [cP] at 25 °C

at 60 °C

16.5

4.68

28.2

6.99

37.3

8.77

44.6

10.2

Heat Capacity at 25 °C [kJ/(kgK)] 2.43 2.30 2.22 2.18

22

EG is a colorless, practically odorless, low volatility, low-viscosity, hygroscopic liquid.

It is completely miscible with water and many organic liquids. It is favorable substance as

antifreezer, as heat transfer fluid, as humectant in paper, leader and glue manufacturing. It is

common in natural gas dehydration and hydrate inhibition. [30]

DEG is a colorless, low-volatility, low viscosity, hygroscopic liquid. Under normal

conditions, diethylene glycol has no detectable odor; however, under high vapor

concentrations, a slightly sweet odor may be detected. It is completely miscible with water

and many organic liquids. The reactivity and solubility of diethylene glycol provide the basis

for many applications. The uses for diethylene glycol are numerous: gas dehydration,

plasticizer for paper, antifreezer for paints, heat transfer fluid, lubricant etc. [30]

TEG is a colorless, odorless liquid, denser than water. It has high viscosity and boiling

point. Miscible with water and insoluble in most hydrocarbons – except ethanol, acetone,

acetic acid, glycerin, pyridine and aldehydes. Heavier paraffin hydrocarbons are essentially

insoluble in TEG. Aromatic hydrocarbons, however, are very soluble, and significant

amounts of aromatic hydrocarbons may be absorbed in the TEG at contactor conditions.

It has higher boiling point and lower volatility than DEG. TEG is known for its hygroscopic

quality and its ability to dehumidify fluids, one of the most common used physical

absorbents in dehydration. The viscosity of TEG and its aqueous solutions increases

significantly as temperature decreases. Sulfur compounds and CO2 are soluble in glycol.

TEG is more soluble in dense CO2 than in natural gas, high acidic content increases the

carryover losses. [11] [28]

TREG is a colorless to straw-colored, low-volatility, low viscosity, hygroscopic liquid

with mild odor. It is completely miscible with water and many organic liquids. It is more

stable and less volatile than DEG or TEG, it has higher decomposition temperature, but it is

more viscous. It is used for dehydration, solvent for dyes and ink, plasticizer. [15]

2.6. Transportation

Significant emitters – power plants, refineries, factories etc. – and the utilization points

– specialized chemical plants, hydrocarbon fields – are rarely close to each other,

transportation of carbon dioxide is usually necessary. It can be terrain – road, rail or pipeline

– or water transportation – fluvial or marine – with ships. In case of tank transportation (road,

rail or ship) special thermal isolated, high pressure containers are needed. These containers

increase the capital costs and limit the maximum mass flow.

23

CO2 can be transported on road, rail and water as a sub-cooled liquid, through pipelines

as a gas, as supercritical fluid, as dense liquid or as subcooled liquid, depending on the

pressure and temperature conditions. CO2 utilization systems require safe, reliable and cost-

efficient solutions for transmission of CO2 from the capturing facility to the end point.

If smaller amounts, for shorter inland distances are transported, road is the ideal solution.

For longer distances rail or pipeline transport are more feasible. A pipeline is CAPEX

intensive, but the OPEX (depending on the distance) is significantly below the rail

transportation. Additionally, the main advantage of the pipeline is, that it is a continuous

form of transit, while the others are periodic. If the source and the end point are connected

by river or sea, ships are recommended. [31]

2.6.1. Transportation by Road

“A 'hazardous substance' is any substance that has one or more of the following intrinsic

'hazardous properties': Explosiveness; Flammability; Ability to oxidise (accelerate a fire);

Human toxicity (acute or chronic); Corrosiveness (to human tissue or metal); Ecotoxicity

(with or without bioaccumulation); Capacity, on contact with air or water, to develop one

or more of the above properties.” [32] According to this classification CO2 (irrespectively

of phase) is a hazardous substance and restricted by the Agreement Concerning the

International Carriage of Dangerous Goods by Road (ADR for short). The ADR regulates

the transportation conditions (pressure, temperature) and all the safety measures.

In most cases greater amount of gas is not feasible to transport on road. Liquification

improves the transportable quantity per truck but specially isolated, high pressure containers

are needed.

2.6.2. Pipeline Transportation

In case of terrain transportation, pipelines are the most economical solution. [13] Unlike

hydrocarbons, the performance parameters of CO2 – phase, density, viscosity, specific heat,

diffuse coefficient, entropy, etc. – can change easily. The impurities can influence the CO2

vapor pressure and phase boundaries. Phase change can cause pressure drop, even can choke

the pipeline, avoidance of phase change is necessary. [33]

Compared with hydrocarbon pipeline transportation, the properties of CO2 result in

different processes, safety measures. Density of gaseous CO2 is higher than air, it is easy to

gather in low shallow places. CO2 can affect the surrounding environment, even causing

asphyxia. Material selection for pipeline, even blow-down pipeline and its valves need to be

specially considered. [13]

24

Carbon dioxide can be transported in gaseous, in liquid, in dense or in supercritical phase.

Each phase has its upsides and downsides. Gas phase is flexible solution in term of quantity,

on the other hand it requires greater pipe diameter, increasing the investment costs.

Gas flow in a pipe causes great pressure loss due to friction. Great pressure drop should be

avoided, because as the pressure decreases the density decreases, the actual flowrate and the

velocity increases, this cause the cooling of the substance. If the CO2 is cooled enough it can

condense to liquid phase. Greater velocity increases erosion as well. Gaseous form is suitable

for lower amounts, shorter-distances and for gas phase sources. It is more favorable for

transport in a densely populated area, compared to other phases. [13]

In case of liquid transportation, not the pressure drop, but the temperature change cause

complications. Liquification of CO2 involves cooling, which is energy demanding.

This can be achieved by heat exchangers, air coolers or mechanical refrigeration units

(MRU). In order to protect the booster pump, CO2 have to be liquified before entering the

pump. After pressure boosting, it is still necessary to be cooled. Liquid CO2 pipelines have

lower operation pressure, and usually need a thermal insulation layer. It fits for lower

amount, and shorter distances or CO2 source in liquid phase. If CO2 not cooled enough and

the ambient temperature is high, it can evaporate and choke the pipeline. [13] [33]

Transportation in supercritical or dense phase requires higher operating pressures, as the

pressure have to be greater, than the critical pressure. If the temperature is greater than the

critical, the pipeline operates in supercritical phase. If the pressure greater than the critical,

but the temperature is below it, the pipeline is in dense liquid phase operation. Installing an

insulation layer or even heating the substance are essential of supercritical transportation.

This kind of additional energy investment is not required for dense phase. Unlike gas

transportation, supercritical and dense transportation needs to have the minimum pressure

above the critical pressure to keep its high density. These forms of transportation are suitable

for larger quantities and longer distances. It is suitable in some areas with lower population

density. Most of the long-distance CO2 pipelines operated in the dense liquid region. [13]

In liquid and supercritical forms, the necessity of insulation layer needs to be determined

by thermodynamic calculation. For longer distances, it is required to avoid phase change.

During transportation the pressure drops. If CO2 enters as supercritical fluid, at some

point it evaporates. This phase transition causes an even bigger pressure drop.

It chokes or may block the pipeline. This means that, there is a maximum safe transport

distance. If there is a need to transport CO2 farther than over this maximum distance,

installation of one or more pressure boosting stations is required.

25

The phase transition occurs when the pressure and temperature drop below the critical

pressure or temperature, it follows that the safe distance depends on the pressure drop

– pipe diameter, roughness, velocity, elevation etc. – and on the heat transfer – insulation,

ambient temperature etc. – either isothermal or adiabatic conditions are assumed. In [33] an

extensive investigation on the long-distance CO2 pipeline transportation can be found. The

final decision of phase of transport depends on economic considerations.

2.6.3. Pipeline Sizing

The minimum basic parameter requirement to design a pipeline are the followings:

- characteristics and physical properties of the fluid,

- desired mass-flow or volume rate of the fluid to be transported,

- pressures, temperatures, and elevations at starting and end point,

- distance of the two points and equivalent length introduced by valves and fittings.

These basic parameters are needed to design a piping system. Assuming steady-state flow,

there are several equations, which are based on the general energy equation

– Bernoulli equation – that can be employed to design the piping system. [34]

2.6.3.1. Pressure Drop Determination

Determination of the pressure drop is one of the most important calculations in any

process that has connection with pipes. For a new pipeline it defines the inner diameter; for

existing pipeline the flow capacity; if the entry and exit point pressures are known.

There are several methods in the literature for multiphase flow pressure drop calculation.

Most of them applicable either for horizontal or vertical sections. The Beggs & Brill method

is one of the few correlations that can deal with any direction of inclination. It is popular in

the petroleum industry, as the track of a long-distance pipeline going through the landscape

meets these conditions.

It was developed using 1" and 1½” sections of pipe – filed with water and air – that could

be inclined at any angle from the horizontal and it deals with all the three components of the

basic pressure gradient equation; the frictional pressure loss, the elevation and the kinetic

energy loss. Beggs and Brill constructed their method based on Field Units, the equations

are valid in this unit system. It is based on the Fanning correlation, as if only a single-phase

fluid is flowing, it devolves to the Fanning gas or Fanning liquid correlation. The two-phase

friction factor is calculated based on the Fanning friction factor and the input gas-liquid ratio.

Since its original publication it has been modified, here the modified method is presented.

[35] [36] [37]

26

The basic pressure drop equation can be written as:

dp

dl=

(dpdl)elevation

+ (dpdl)friction

1 − Ekinetic

(34)

Where

(dp

dl)elevation

=1

144ρM sin(θ) (35)

(dp

dl)friction

= 1.294 ∙ 10−3ftpvM2 ρnsp

(36)

Ekinetic = 2.16 ∙ 10−4vMvsGρns

p (37)

The first step in the calculation of elevation term is to determine the appropriate flow

regime – segregated, intermittent, transition or disturbed – based on the no-slip liquid

holdup. One can use either a flow pattern map, like Figure 9 or the boundary conditions

summarized in Table 3. Eq. 38 to 41 show the boundaries of the flow regimes.

L1 = 316λL0.302 (38)

L2 = 9.252 ∙ 10−4λL

−2.2484 (39)

L3 = 0.1λL−1.4516 (40)

L4 = 0.5λL−6.738 (41)

Figure 9 Modified Beggs & Brill flow regime map Based on: [37]

27

In order to determine the flow regime, calculation of a dimensionless parameter, called

Froude Number, is required. Froude number is an important parameter of comparison, where

the weight of the fluid is a significant. It represents the ratio of the inertial forces on an

element of fluid to its weight. It can be calculated using Eq. 42.

Fr = √vMgID

(42)

Table 3 summarizes the conditions of the flow regimes. There is an AND logical

connection between liquid volume fraction and Froude number.

Table 3 Beggs & Brill flow regime boundary conditions

Liquid volume fraction (λL) Froude Number (Fr) Flow regime

<0.01 <L1 Segregated

≥0.01 <L2

0.01≤ λL <0.04 L3< Fr ≤L1 Intermittent

≥0.4 L3< Fr ≤L4

<0.4 ≥L4 Disturbed

≥0.4 >L4

– L2< Fr <L3 Transition

After the flow regime is determined, the liquid holdup should be calculated.

It has two parts, first the liquid holdup is calculated for horizontal flow, then it is modified

for inclined flow. The liquid holdup should be always greater or equal than λL, if this

condition not fulfilled, λL should be used. Liquid holdup is calculated with different formula

for each flow regime, Table 4 summarizes the equations.

Table 4 Horizontal liquid holdup calculation formulae

Flow regime Formula

Segregated εL(0) =0.98λL

0.4846

Fr0.0868 (43)

Intermittent εL(0) =0.845λL

0.5351

Fr0.0173 (44)

Disturbed εL(0) =1.065λL

0.5824

Fr0.0609 (45)

Transition εL(0) = AεL(0)segregated + BεL(0)intermittent (46)

28

In the transition region as its name suggests, the fluid flow pattern is between the

segregated and intermittent regime, fluid properties are the mixture of the two regimes.

‘A’ and ‘B’ parameters in Eq. 47 and 48 represent this mixing. They can be calculated as:

A =L3 − Fr

L3 − L2 (47)

B = 1 − A (48)

Once the horizontal liquid holdup is determined, the actual liquid holdup should be

calculated as:

εL(θ) = B(θ)εL(0) (49)

Where B(θ) is the inclination factor. It can be calculated as:

B(θ) = 1 + β [sin(1.8θ) −1

3sin3(1.8θ)] (50)

β is a function of flow regime, direction of inclination (uphill or downhill), the Froude

number and the liquid velocity number. It is always greater than 0, if a negative value is

calculated, β=0 must be used. The liquid velocity number can be expressed as:

Nvl = 1.938vsL√ρLgσ

4

(51)

Table 5 β calculation formulae

Direction Flow regime β

Uphill

Segregated β = (1 − λL) ln (0.011Nvl

3.539

λL3.768Fr1.614

) (52)

Intermittent β = (1 − λL) ln (2.96λL

0.305Fr0.0978

Nvl0.4473 ) (53)

Disturbed β = 0 (54)

Downhill All β = (1 − λL) ln (4.7Nvl

0.1244

λL0.3692Fr0.5056

) (55)

The actual liquid holdup determines the mixture density, which effects the elevation term.

For the frictional losses one should calculate the ‘S’ empirical parameter and the no-slip

friction factor. For the no-slip friction factor one needs the Reynolds number and the

assumption that the pipe wall is smooth. The Reynolds number should be expressed as:

Re =ρnsvMID

μns (56)

29

This method uses the Fanning friction factor, which can be determined from the Fanning

friction factor chart or can be calculated with the Chen equation:

1

√fns= −4 log [0.2698ϵ −

5.0452

Relog {0.3539ϵ1.1098 +

5.8506

Re0.8981}] (57)

This friction factor should be normalized for actual conditions as:

ftp = fnseS (58)

The empirical parameter ‘S’ can be calculated as:

S =y

−0.0523 + 3.18y − 0.872y2 + 0.01853y4 (59)

Where

y =λL

εL2 (60)

Eq. 59 is discontinuous if 1 < y <1.2, in this interval Eq. 61 should be used.

S = ln(2.2y − 1.2) (61)

2.6.3.2. Heat Loss During Transportation

The pressure drop during transportation highly depends on the density and viscosity of

the fluid, which are temperature dependent properties. In order to accurately calculate the

pressure drop, the heat transfer processes – heat generated by friction, heat convention inside

the pipe, heat transfer to the environment – should be known as well. During heat transfer,

part of the internal energy of the fluid migrates to the environment through the pipe wall as

heat. This can be written with the Fourier’s law of heat conduction.

qx = −kA∂T

∂x (62)

or in radial system as:

qradial = −kAradial∂T

∂r= −k2πrL

∂T

∂r (63)

In Eq. 62 and 63 the ‘k’ thermal conductivity, a quality dependent parameter. It shows

how much energy can the given substance transfer in a unit of time for 1 °C temperature

difference. Its unit most commonly is W/(m°C). [38]

Pipelines can be considered as cylinders, where the radius is negligible compared to the

length. It can be assumed that heat is transferred only in the radial direction. Solving Eq. 63

for a given temperature difference it can be written as:

q =2πkL(Tinner − Touter)

ln (routerrinner

) (64)

30

For comparison of insulation materials, the thermal resistance is commonly used, its

common unit is °C/W. By definition the thermal resistance is:

Thermal resistance =Thermal potential difference

Heat flow (65)

For one layered radial system it is:

Rth =

ln (routerrinner

)

2πkL

(66)

In case of a buried pipeline the thermal resistance of ambience can be calculated as:

Rth amb =routerkamb

ln

(

2Zb +√4Zb

2 − 2router2

router)

(67)

The reciprocal value of thermal resistance called heat transfer coefficient; it can be

expressed as:

Hth amb =1

Rth amp (68)

If the total heat transfer should be calculated it is characterized by the overall heat transfer

coefficient as:

U =1

A∑Rth=

1

A Rth overall (69)

2.6.3.3. Wall Thickness Determination

Once the inner diameter of the piping segment has been determined, the pipe wall

thickness must be calculated. There are many factors, that affect the pipe wall thickness

requirement. These parameters can be divided into three groups: 1.) operation conditions:

maximum and working pressures and temperatures; 2.) physical and chemical properties of

the transported fluid: density, viscosity, pH, water content, heat capacity etc.; and

3.) properties of pipe material: grade, yield strength, tensile strength. Value of safety factor

and exact calculation method are defined by the relevant documents of the country. [34]

In Hungary the MSZ EN 13480 Metallic industrial piping standard is authoritative, from

this group Part 3: Design and calculation details the wall thickness determination of piping

components. Through the case study the process presented by this standard was applied.

The minimal required wall thickness depends on the piping material, the diameter, the

joint coefficient, the calculation pressure and temperature. The diameter is determined by

the maximum allowable pressure drop and velocity, the other parameters are detailed by the

document.

31

As the standard states: “The joint coefficient z shall be used in the calculation of the

thicknesses of components which include one or several butt welds, other than

circumferential, and shall not exceed the following values:

- for equipment subject to destructive and non-destructive testing which confirms that

the whole series of joints show no significant imperfections: 1;

- for equipment subject to random non-destructive testing: 0.85;

- for equipment not subject to non-destructive testing other than visual inspection: 0.7.”

[39]

The pressure and temperature in the calculation procedure must be considered as the most

severe conditions of coincident pressure and temperature, which may prevail in the piping

system over a long time.

The design stress must be the lowest of the time-independent stress values – ReH t, Rp0.2 t

and Rm – as Eq. 70 presents.

F = min {ReH t1.5

;Rp0.2 t

1.5;Rm2.4} (70)

In case of straight pipe with an OD/ID ≤ 1.7 the minimal required wall thickness can be

determined using Eq. 71 or 72.

e =pcOD

2Fz + pcalc (71)

or

e =pcID

2Fz − pcalc (72)

The final wall thickness should be greater than ‘e’ and the added allowances.

These allowances are: ‘c0‘ is the corrosion or erosion allowance; ‘c1’ is the absolute value of

the negative tolerance taken from the material standards or as provided by the pipe

manufacturer; ‘c2’ is the thinning allowance for possible thinning during manufacturing

process.

eord ≥ e + c0 + c1 + c2 (73)

‘eord’ is the next greater standard wall thickness that fulfills the above condition. After the

‘eord’ is determined the maximum pressure should be checked. For this calculation the ‘ea’

analysis wall thickness should be used.

ea = eord − c0 − c1 − c2 (74)

If ‘ea’ fulfills the stress requirements the determination of straight section is complete.

32

2.6.3.4. Effect of Corrosion

Pipelines have important role in transport of fluids for long distances. Because carbon

steel is a commonly used as material, corrosion is one of the most serious problems in the

petroleum industry. Application of stainless steel could eliminate the corrosion, but due to

the cost of this special material and the required long distances, no project could be feasible.

Corrosion is defined by the National Association of Corrosion Engineers (NACE) as “the

degradation of a material, usually a metal, because of a reaction with its environment”. [40]

Corrosion could lead to failures, breakdowns and leaks, which are followed by large losses

of product, environmental pollution and ecological disasters. These cause significant

economic loss. [41]

CO2 and H2S are frequent components in the produced hydrocarbon. From the several

forms of corrosion, sweet corrosion (caused by carbon dioxide mixed with water) and sour

corrosion (caused by hydrogen sulfide) are the crucial problems in the petroleum industry.

Removal of these components from the produced oil and gas is critical from operation safety

point of view. In case of oil production, the oil could act as a natural inhibitor, however in

gas production such natural inhibition is not available, outer source of inhibition is needed.

The use of chemical inhibitors has been acknowledged practical and economical method of

combating CO2 corrosion. The inhibiting molecule retards the rate of corrosion by acting at

the metal-corrosive medium interface. [41]

The mechanism of carbon dioxide corrosion is a complicated process, that is influenced

by many factors and conditions – i.e. temperature, pH, partial pressure of CO2, etc. –, but

due to its importance, this type of corrosion is well investigated.

There are three main ways to avoid corrosion, removal of water (dehydration) or

application of passive or active inhibition. Passive methods are special coatings, insulation

materials, they separate the metal from the environment. Active corrosion inhibition includes

cathodic protection – when another metal is sacrificed to protect the pipeline – and adding

inhibitors to the wet gas stream – reducing the hazard of hydrate formation as well.

Application of inhibitors are common in case of pre-processed stream (e.g. flowlines) and

during gas processing (e.g. dehydration, sweetening etc.). In order to avoid hydrate

formation in wet gas pipelines, ethylene glycol and/ or methanol is often added.

33

3. Case Study

The aim of the study is to investigate the different possibilities of a CO2 dehydration and

transport system. The source of the gas mixture is an operating, amine-based gas sweetening

unit. The gas of the GSU is pressure boosted, dehydrated and transported to reinjection to

the target reservoir. The initial transportation pressure is the sum of the pressure needed to

reinject the gas to the formation (100 barg) and the pressure loss during transportation.

The GSU and reservoir are the boundary points, their operation is excluded from this

work. Table 6 and 7 summarizes the initial parameters of the gas mixture leaving the

sweetening unit. Figure 10 shows a simplified flowsheet of the process.

Table 6 Composition of the input gas

Component mol% g/m3

C1 4.021 27.430

C2 0.235 3.010

C3 0.092 1.720

i-C4 0.019 0.460

n-C4 0.035 0.860

i-C5 0.012 0.370

n-C5 0.012 0.370

C6 0.131 4.810

C7 0.030 1.290

C8 0.007 0.350

CO2 95.271 1 783.050

N2 0.135 1.610

Total 100.00 1 825.330

Table 7 Initial parameters of the input gas

Parameter Value Unit

Temperature 20 °C

Pressure 0.5 barg

Volume Flow 10 000 STDm3/h

Molecular

Weight 42.92 g/mol

Specific

gravity 1.47 –

Water Content 11 778.57 mg/STDm3

Specific Heat 0.907 kJ/(kgK)

Heat Capacity

Ratio 1.29 –

Figure 10 Simplified flowsheet Own edit

GSU Pressure

Boosting

Inhibitor Injection

Target Reservoir

Pressure

BoostingTransport by PipelineWater Removal

34

Table 8 Sub-cases

Case Number Short Description

Case 1. Transport the carbon dioxide by pipeline to utilization

Case 1.1. Dehydration with glycol technology

Case 1.1.1. Dehydration with DEG

Case 1.1.2. Dehydration with TEG

Case 1.1.3. Dehydration with TREG

Case 1.2. No dehydration but application of thermodynamic inhibition

Case 1.2.1. Inhibition with EG

Case 1.2.2. Inhibition with methanol

Case 1.3. No dehydration, nor inhibition

Case 2. Emitting the CO2 to the atmosphere and paying the emission allowances

Through the Case Study the sub-cases summarized in Table 8 are investigated, evaluated

and compared to each other. The final evaluation and recommendation are based on

estimated expenditures – CAPEX and OPEX – on operation interval of 10 years.

3.1. Model Building Principles

The modelling carried out by Aspen HYSYS® v12 (or HYSYS for short) software.

The software solves mass and energy balances, calculates vapor/ liquid equilibria, with

mathematical models – Equations of State – to predict the physical and chemical processes

taking place in the investigated system. Originally, it was developed as a chemical

engineering software, but today it is one of the most popular tools in the industry and

academia for steady-state and dynamic simulation, process design, performance modelling,

and optimization. In the petroleum industry it is used in gas processing, process engineering

and in refineries as well.

Through the simulation the Sour-Soave-Redlich-Kwong (Sour-SRK) fluid package is

used as default. This package is a mixture of the Soave-Redlich-Kwong Equation of State

and the Wilson sour water model. Through dehydration the CPA fluid package is used, based

on the software’s recommendation. This model uses the Cubic-Plus Association term. For

the propane-based mechanical refrigeration unit the Peng-Robinson fluid package is used.

This package uses the Peng-Robinson EoS.

35

3.2. Pressure Boosting and Cooling

3.2.1. Compressors and Air Coolers

As mentioned above, for dehydration higher pressure is required. In order to achieve this

pressure, a 3-stage pressure boosting system – with reciprocating compressors – was

modelled. Figure 11 shows the pressure boosting model made in HYSYS and Table 9

summarizes the most important parameters. The final (4th) stage of pressure boosting

detailed in Chapter 3.4.

Figure 11 Pressure boosting (3 stages) Own edit

As recommended in API Standard 618 Reciprocating Compressors for Petroleum,

Chemical, and Gas Industry Services, the compressor discharge temperature should not be

higher than 150 °C (300 °F) because of the danger involving the auto-combustion of

lubricating cylinder oil in the presence of hot, compressed air. The pressure ratio was

selected considering this regulation. In case of pure CO2, this ratio, as a rule of thumb can

be approximated with 2.6. The impurities have effect on the heat capacity of the mixture,

thus increasing the allowable pressure ratio. The selected pressure ratio was assumed equal

for all stages. After each compression stage, the gas is cooled back to 50 °C by air coolers

and the free liquid is separated before entering the next stage. The pressure losses of the

piping system are neglected.

36

Table 9 Pressure boosting parameters

Compressors

Parameter 1st Stage 2nd Stage 3rd Stage

Adiabatic Efficiency [%] 92.00 92.00 92.00

Polytropic Efficiency [%] 92.875 92.875 92.875

Pressure Ration [-] 3.16 3.16 3.16

Inlet Pressure [barg] 0.50 3.42 12.64

Outlet Pressure [barg] 3.77 12.99 42.14

Outlet Temperature [°C] 108.0 145.3 149.0

Air coolers

Air inlet temperature [°C] 40 40 40

Process stream pressure drop [bar] 0.35 0.35 0.35

Process stream outlet temperature [°C] 50 50 50

Removed heat [kWth] 277.0 510.6 590.8

Energy requirement [kWel] 3.05 6.06 6.04

3.2.2. Mechanical Refrigeration Unit

As mentioned above, after each stage the gas is cooled back to 50 °C. This high

temperature is reducing the efficiency of dehydration, additional cooling is required.

A propane-based mechanical refrigeration unit (MRU) was designed, its cooling duty is used

to reduce the temperature of processed gas. Figure 12 shows the model of the unit. This loop

process is based on the Joule-Thomson expansion. As the gas or liquid is flowing from a

higher pressure into a lower pressure region without significant kinetic energy change, work

is done, causing a change in internal energy but the enthalpy is constant. The internal energy

can increase, if work is done on the fluid and decreased, if the work is done by the fluid.

The loop starts at the condensed (liquid) propane. The liquid flows through a J-T valve,

its pressure and temperature drop, the vapor phase appears. In the heat exchanger the propane

evaporates, removes energy from the ambience, thus reducing the temperature of the other

medium, in this case the gas mixture. The vapor goes to a tank or separator, and after pressure

boosting, to the condenser, where it become liquid and the process starts over. In the practice

a special shape of heat exchanger called chiller – which allows the propane to expand – is

needed to deliver the cold energy. In the simulation a pair of ideal heater and cooler was

used with a connected duty. The result is the same from modeling point of view, but with

this solution it is easier to follow the model, and the different units are more portable.

37

Figure 12 Propane-based mechanical refrigeration unit Own edit

As Figure 13 shows, if no other parameter (pressure and temperature) is modified the

cooling duty is changing linearly with the quantity of propane.

Figure 13 Cooling duty vs quantity of propane Own edit

Table 10 MRU parameters

Parameter Value Unit

Propane mass flow 2 666 kg/h

Cooling duty 161.1 kWth

Compressor energy requirement 107.9 kWel

Condensation temperature 50 °C

Chiller temperature -15 °C

38

Condenser

Number of bays 1 –

Number of fans per bay 1 –

Removed heat 256.9 kWth

Energy requirement 12.88 kWel

With a propane-based MRU the process gas mixture is cooled back to 26 °C. After a J-T

valve the pressure drops to 40.5 barg and the temperature to 25 °C. These are the inlet

parameters of the dehydrator unit.

3.3. Case 1.1. – Dehydration with Glycol Technology

In order to prevent corrosion and hydrate formation the water content of the gas has to be

reduced. Furthermore, by removing the water the total volume flow is reduced, thus reducing

the transportation cost, and increasing CO2 injection efficiency. On the other hand, complete

water removal is not essential nor economical in this situation. An industrial best practice

presented in [42], 50 ppmv water content is generally accepted for pipeline transport. This

is the goal of dehydration.

The basic buildup of an absorption-based dehydration unit is the same as described in

Chapter 2.5.1. Flash tank was not built in the system, because of low operation pressure.

This configuration was used through the case study with different substances, resulting

different operational parameters (e.g. absorber mass flow and concentration, regeneration

temperature, strip gas quantity etc.). As mentioned above EG is common practice in cold

technologies – natural gas conditioning – and hydrate inhibition, however in case of water

removal the first substance of consideration is TEG. Beside TEG, DEG and TREG were

investigated.

The basic steps of design is based on the method presented in [11] and [28], the final

parameters are the result of HYSYS modelling. For the optimization, the results presented

in the following chapter are used. These investigations are made with the Case Study tool of

the software. The figures show simple relationships, every other parameter – not shown on

the graphs – considered constant during the investigations. In order to present the result in a

more representative form, some graphs show relative results. This always means the ratio

compared to the initial – presented in tables – values. This chapter presents the design

process of TEG unit as an example; the graphs belonging to DEG and TREG can be found

in Appendix C.

39

Figure 14 shows the dehydrator unit built in HYSYS. The “Sour gas to absorber” and the

“Sour gas to dehydration” material streams on Figure 11 and 14 are connected by virtual

stream (op-100). All parameters – composition, pressure, temperature, volume flow, etc. –

are the same values.

The gas enters the T-100 contactor tower on the bottom, while the absorber on the upper

section. The dried gas leaves on the head (stream Ovhd). The inlet stream of 4th stage

pressure boosting compressor is connected to this stream by another virtual stream (op-101).

The model of last pressure boosting stage, and transportation is seen on Figure 36.

In order to prevent foaming, the solvent’s temperature should be greater than the inlet

gas’ temperature. This limitation was considered during the simulation.

Figure 14 Dehydrator unit Own edit

If one wants to check the water dew point of the processed gas, one can observe falsely

high values. Unfortunately, HYSYS cannot calculate the water dew point accurately, if there

is glycol in the system. Following the software’s suggestion, a splitter (X-100) is used to

remove the glycol. This solution may not seem ideal, but it is a common practice [42].

This stream is not used in the following processes, only needed to determine the correct

water dew point.

40

3.3.1. Design of a Glycol-based Dehydration Unit

As it was mentioned above, the gas enters the unit at 25 °C and 40.50 barg. In order to

achieve the 50 ppmv water content – ~-13 °C water dew point at transportation pressure –

98.00 wt% or higher concentration of TEG is required. Figure 15 is based on the equilibrium

between vapor and liquid water phase, and applicable up to 100 bara. The inlet gas water

content is 571.03 mg/STDm3; the required is 38.10 mg/STDm3; 532.93 mg/STDm3 is

removed, the water removal is ~93.32 %. This high removal – with 3 stages – is not possible

with TEG concentration below 99.00 wt%. (see Figure 16) Based on Figure 17 a circulation

ratio of 17 is sufficient, the lean TEG mass flow is 17 times of the removed water mass flow.

There is total 5.62 kg/h water in the 9 851 STDm3/h inlet gas and 0.38 kg/h should remain

in the gas stream. Based on the circulation ratio, 89.19 kg/h 99.00 wt% lean TEG is needed.

These literature-based values are the first iteration step of the design. The final parameters

are the product of HYSYS modelling.

Figure 15 Water dew point vs contractor temperature Source: [11]

Based on Figure 15 98.00 wt% TEG is able to remove the water, on the other hand

Figure 16 and 17 prove that 99.00 wt% TEG is needed, if 3 stages are used. On Figure 16

the red line shows the theoretically possibly maximum water removal and the green line is

the required water removal.

41

Through the simulation TEG concentration higher than 98.14 wt% was unachievable

without strip gas. As an industrial practice, to achieve 97.00 wt% or higher purity, strip gas

is required. In case of DEG, because of its lower decomposition temperature (164 °C), strip

gas is essential, in contrast the higher decomposition temperature of TREG (238 °C) reduces

or even eliminates the strip gas requirement. As strip gas pure, dry N2 was selected and used.

As calculated above 89.19 kg/h 99.00 wt% TEG is needed. Running the simulation

resulted 43.27 ppmv water content, which satisfies the goal 50 ppmv.

Figure 16 Absorber performance at lean TEG concentration of 98.50 wt% Source: [11]

Figure 17 Absorber performance at lean TEG concentration of 99.00 wt% Source: [11]

42

3.3.1.1. Main Influencing Parameters

As one can see from the discussion above, there are several factors influencing the

operation of a dehydrator unit. The result is infinite number of possible operation mode.

From these, should be one, most preferred operation. The goal of optimization is to find the

parameters belonging to this optimal operation. The most preferred operation could mean

several conditions depending on the criteria of optimization. It could mean the highest

achievable water removal, the least energy consumption, the least operational cost or least

total expenditures including CAPEX and OPEX. This study aims for the least operational

cost, least total chemical (glycol, propane, nitrogen) and least energy (heat and electricity)

consumption. The following section provides a short overview of the effect of different

factors affecting the operation of a glycol-based dehydrator unit.

As it was mentioned above, glycols have favorable physical properties for dehydration.

However, they have different basic physical qualities (density, viscosity, decomposition

temperature, etc.) which are the main indicators how a dehydrator unit will operate. Another

important parameter is the cost of substances, that can decide which absorber should be used.

3.3.1.1.1. Effect of Inlet Temperature

Contractor temperature is one of the most important parameters determining the operation

of the dehydrator unit. As one can see on Figure 15 in order to achieve lower water dewpoint

at constant contact temperature, higher purity of TEG is required. Or with a constant TEG

concentration by reducing the contact temperature, the water dewpoint will be reduced.

However, reducing the temperature increases the heat requirement of regeneration and

viscosity of the glycol which drag along the increment of pressure drop and pumping duty.

Figure 18 Remaining water content vs contractor temperature (TEG) Own edit

43

As one can see on Figure 18 and 19 lower contractor temperatures are favorable, as not

only higher water removal efficiency is achievable but at lower contractor temperatures the

loss of glycol is less. The mass fraction of circulated TEG has negligible effect, compared

to the contractor temperature.

Figure 19 TEG loss vs contractor temperature Own edit

3.3.1.1.2. Effect of Glycol Concentration

If no other parameters changed, higher water removal can be achieved by greater glycol

quantity. This increment can be achieved by higher mass flow (with lower glycol purity) or

higher purity (with lower mass flow). Higher purity means higher strip gas requirement, on

the other hand, higher mass flow requires higher reboiler duty.

Figure 20 TEG mass fraction vs mass flow Own edit

44

3.3.1.1.3. Effect of Strip Gas

The main aim of adding stripping gas, is to reduce the partial pressure of gaseous water,

to enhance water evaporation and thus reduction the water content of the glycol. The

stripping gas could come from the flash gas, from the dehydrated gas or from a third source.

As it was mentioned above, for the case study nitrogen was selected, as strip gas. As one can

see on Figure 21 the increased quantity of strip gas results lower remaining water content.

However, the relation between the removed water and the quantity of strip gas used is not

linear, more like an exponential relation. The same behavior is observable on Figure 22

between the quantity of strip gas and purity of TEG. This relation can be approached with a

logarithmic function.

Figure 21 Remaining water content vs strip gas quantity (TEG) Own edit

In conclusion one can say, that the increment of stripping gas quantity enhances the glycol

regeneration, thus purer glycol can be circulated, which leads to higher water removal

efficiency. The quantity of strip gas has greater effect on the lower volume flow regions; a

small amount of strip gas can enhance the regeneration process compared to a no strip gas

scenario. On the other hand, there is a limit where greater amount of strip gas will not result

significant change in dehydration efficiency.

One should bear in mind, that the strip gas is injected to the reboiler. If greater amount of

strip gas is needed, it can influence the reboiler duty. As one can see on Figure 23 after a

minimal value, the duty shows linear relation with TEG mass flow. The temperature of strip

gas does not affect this trend however, it shifts the curve. This trend is true in case DEG as

well. (see Figure 54) The high decomposition temperature of TREG allows high grade

regeneration without strip gas. The effect of strip gas temperature was not investigated.

45

Figure 22 TEG mass fraction vs strip gas quantity Own edit

In most cases the quantity and temperature of strip gas does not have significant effect on

the energy requirement. Figure 23 and 54 are only valid in case of nitrogen gas. As one can

see, the temperature of strip gas and required duty have linear relation, while the quantity of

strip gas and reboiler duty have logarithmic relation. (see Figure 48 and 49)

Figure 23 Reboiler duty vs TEG mass flow Source: Own edit

3.3.1.1.4. Effect of Regeneration Temperature

The regeneration temperature should be lower than the decomposition temperature. For

TEG 200 - 204 °C common choice, achieving high water removal, but staying below the

decomposition temperature of TEG (208 °C). The higher the regeneration temperature the

less strip gas required. Figure 24 shows not only the importance of high regeneration

temperature but one can notice the decreasing effect of strip gas with increasing quantity.

46

Figure 24 TEG mass fraction vs reboiler temperature with different strip gas quantities Own edit

Decomposition temperature of DEG is 164 °C, of TREG is 238 °C, resulting higher strip

gas requirement for DEG, while TREG is regeneratable to required purity with low amount

or even without any strip gas. (see Figure 55 and 61)

3.3.1.1.5. Effect of Number of Stages

Inside the contactor tower each stage has its own pressure, temperature and belonging

vapor/ liquid ratio. As seen on Figure 16, 17 and 25 a given water removal efficiency,

requires lower and lower amount of glycol with the increment of stages. Or for a given

circulation ratio, with more stages higher water removal is achievable. This trend is valid in

case of DEG and TREG as well. (see Figure 56 and 62) The horizontal red line shows the

desired water removal, the vertical lines show the glycol requirements.

Figure 25 Absorber performance at lean TEG concentration of 99.05 wt% Own edit

47

3.3.1.1.6. Effect of Heat Exchangers

As it was described before, during dehydration lower temperatures are advantageous, on

the other hand the regeneration requires higher temperatures. This temperature swing

unequivocally involves a heat exchanger, where the rich glycol is pre-heated while the lean

glycol – leaving the regenerator – is cooled back. This heat exchanger is not crucial for the

process but increases the efficiency. Without it, additional heating and cooling is required.

A basic design consideration is, that using a static apparatus (i.e. heat exchanger) is more

economical, than operating an additional cooler unit (e.g. air cooler or MRU). The low

temperature of contractor tower cannot be achieved with air coolers, MRU is needed.

Removing the exchanger has double effect, the lean glycol has to be cooled back, and the

rich glycol enters the regenerator tower on lower temperature, increasing the reboiler duty

or additional pre-heating is needed.

In order to quantify the increased cooling duty, the heat exchanger was removed from the

system and an ideal cooler was inserted to cool down the lean TEG to the original contractor

temperature. (see Figure 26)

Figure 26 Dehydrator unit without heat exchanger Own edit

48

It is practical to use the same propane-based MRU unit – that cools the gas mixture before

dehydration – for cooling down the glycol as well. This means greater propane mass flows,

greater vessels, duties etc.

Figure 13 showed the linear relationship between cooling duty and propane requirement,

Figure 27 quantifies the propane increment in relative form. One can notice that the trend

not only linear but the gradient equals unity.

Figure 27 Cooling duty vs quantity of propane (relative) Own edit

As mentioned above, removing the heat exchanger effects the regeneration as well.

Figure 28 represents the effect of inlet temperature on reboiler duty. As the inlet temperature

decreases, the reboiler duty increases. This relationship is not linear, as seen on the deviation

of datapoints from the reference (grey, dotted) line.

Figure 28 Effect of heat exchanger on reboiler duty (TEG) Own edit

49

3.3.2. Comparison of Glycols and Selection of Favorable Substance

The purpose of this section is to graphically demonstrate the differences of investigated

glycols and to present the selection process. Most of the curves presented here, can be found

in the previous chapter or in Appendix C, however comparison is more representative, if

the given parameter of different substances is in one graph. In this section DEG, TEG and

TREG presented with the same color – blue, orange and grey respectively. There were some

calculations where the software could not calculate the TREG parameters properly – based

on values presented in open literature. These values were not accepted, nor presented here.

Figure 29 Remaining water content vs contractor temperature Own edit

Figure 29 shows the remaining water content as function of contractor temperature.

In the lower temperature zone – below 30 °C – the values are close to each other, DEG

results higher remaining water content.

Figure 30 Glycol loss vs contractor temperature Own edit

50

DEG has the highest absorber loss, which leads to greater make up requirement, thus

higher OPEX. Loss of TREG is negligible even at higher temperatures, TEG loss also stays

relatively low. (see Figure 30)

In order to achieve the required water content, a fixed amount of pure glycol is needed.

The pure amount of glycol is the product of mass fraction and mass flow, Figure 31

represents this relation. DEG has the lowest mass flow requirement at certain mass fraction,

or at a certain mass flow the lowest mass fraction is required. This statement only valid above

the critical mass fraction – where the curve becomes flatter –, below that value the substance

cannot attract more water even with greater mass flows. TEG has higher mass fraction

requirement than DEG.

Figure 31 Glycol mass fraction vs mass flow Own edit

Greater purity means reduced mass flow, however achieving the required mass fraction

has different conditions – regeneration temperature, quantity of strip gas – depending on the

substance. Figure 32 shows the strip gas requirement of each glycol. This graph is an

exception. All the other graphs show the different substances at same conditions. On this

graph the lines belonging to DEG, TEG and TREG are valid for regeneration temperatures

of 160 °C, 204 °C and 230 °C respectively.

As one can see on Figure 32 the statement, which indicates the great effect of low amount

of strip gas compared to no strip gas scenario and the reducing trend of effect, is true in all

cases.

51

Figure 32 Glycol mass fraction vs strip gas quantity Own edit

Another important parameter of the system is the loss of CO2 during dehydration.

Removing carbon dioxide from the gas with the water is favorable in case of natural gas

conditioning, as it increases the ratio of valuable components. In this case the aim of the

system is to purify the CO2, reducing its quantity should be avoided. Still these losses are

negligible compared to the total quantity.

Figure 33 CO2 loss vs mass fraction of glycol Own edit

Based on the graphs presented in this section and the material cost in Appendix B, the

proper glycol can be selected. With DEG the lowest mass flow can be obtained – with

constant mass fraction –, it requires the lowest regeneration duty and has the lowest cost

from the three, but then again it has the greatest glycol loss and the greatest strip gas

requirement. The loss of TREG is negligible, can be regenerated to high purity without or

52

with little amount of strip gas and can achieve the required water content with the lowest

purity, but it has the greatest regeneration duty requirement, and the greatest cost.

Furthermore, the high uncertainty of TREG calculation made the results unreliable.

Selection of TREG as representative glycol for this scenario not recommended.

In most aspects TEG is in between DEG and TREG. After several iteration TEG proved

to have the lowest total OPEX, thus the economic evaluation of this substance will represent

the cost of dehydration.

3.3.3. Contactor Tower Specification

The contractor tower and the regenerator tower are the biggest static components of the

system, they have to be sized properly. The applied process in details can be found in [11],

as an example the contractor tower sizing is presented here.

In the simulation environment the contractor tower was designed with theoretical stages.

These stages, as the name suggests only exists in theory, a conversion to real packing is

needed. For this, transfer units – number (NTU) and height (HTU) – are used. Figure 34

provides an approximate conversion from theoretical stages to transfer units, as a function

of circulation ratio. The contractor has 4 theoretical stages and circulation ratio of 17.74.

From Figure 34 NTU is 7.1.

Figure 34 Number transfer unit conversion Source: [11]

53

Height transfer unit depends on the mass transfer rate, they are inversely proportional.

Increment in HTU represents a decrease in mass transfer. Increasing specific packing area

and glycol circulation rate decreases HTU, and gas rate and gas density increase it.

Figure 35 Height transfer unit conversion Source: [11]

The total height of the contactor column is based on the packing required, plus an

additional 3 m to allow space for vapor disengagement on top, and inlet gas distribution, rich

glycol surge volume at the bottom of the column and for the liquid distributor. Figure 35

was used to estimate the HTU of structured packing with gas density of 92.52 kg/m3

(40.5 bar, 25 °C) and 300 m2/m3 specific packing area. HTU equals 0.7. The total height

includes two 200 mm thick additional packing, the final height is calculated with Eq. 75.

Total Height = NTU ∙ HTU + hap + has = 7.1 ∙ 0.7 + 0.4 + 3 = 8.37 m (75)

Diameter of the contactor tower is set by the gas velocity; it is identical to sizing a

separator. Here only the equations and results are presented, the value of parameters and

constants can be found in Appendix A.

vmass = Cbubble tray ∙ √ρL(ρL − ρV) = 52 618 kg/m2h (76)

A =mmassvmass

= 0.3397m2 (77)

Dbubble tray = √4A

π= 0.6577 m (78)

Dstructured packing = √Cbubble tray

Cstructured packing∙ Dbubble tray = 0.6770 m (79)

54

The height calculated with Eq. 75 (8.37 m) includes additional space, but Eq. 79

determined only the width of the packing (0.677 m). These values do not include the wall

thickness of the tower. In practice, an 8.50 m tall, 0.70 m wide tower can be used.

Low quantity as presented in the case study results in a small dehydrator unit. It is possible

that periodic operation of regeneration is more feasible. In that case greater apparatus and

storage capacity are required. This option is not in the scope of the case study.

As one can see from Table 19 the final parameters do not significantly deviate from the

literature-based values presented in the previous chapter, justifying the usage of the

presented graphs for rough calculations. One also should bear in mind that Figure 15 - 17

are made for natural gas, not for CO2. For more accurate design EoS methods are essential.

3.4. Transportation of CO2 by pipeline

The injection pressure of CO2 was specified in the beginning of the case study,

as 100 barg. This pressure determines the injection phase of CO2, it can be either supercritical

or dense liquid phase, depending on the temperature. During transportation it can be dense,

supercritical or gas phase, but phase change during transportation should be avoided.

The advantages of transportation in dense or supercritical phase are smaller pipe diameter

and smaller pressure drop, on the other hand in case of dense phase additional cooling, in

case of supercritical phase additional thermal insulation is required. The third option – gas

phase – requires greater diameter and pressure boosting on-site. Without any calculation,

one can realize, this solution has greater CAPEX – greater pipe diameter – and

OPEX – greater pressure boosting – demand, this scenario is therefore not investigated.

The high-pressure mixture cannot enter the pipe after compression, as the high

temperature decreases the mechanical strength of the steel. The mixture was cooled back to

50 °C with air coolers. In case of supercritical transportation, this is the inlet temperature of

the pipe. In order to achieve dense phase, the substance should be cooled below the critical

temperature. This additional cooling provided by same the propane-based MRU, used

before. In the model for simplicity ideal cooler was used.

Figure 36 4th stage compression and transportation pipeline Own edit

55

3.4.1. Pipeline sizing

Line sizing limited only to pipeline diameter and wall thickness determination, it excludes

the design of fittings, valves and other components. The following methods and regulations

were applied: API RP14E recommended practice for maximum allowable velocity;

MSZ EN 13480-3 standard and 2/2010. (I. 14.) KHEM edict for wall thickness;

and Beggs & Brill method for pressure drop calculations. The final inner and outer diameter

of the pipe is the result of an iterative process. Figure 37 shows the steps of the iteration.

Figure 37 Process of pipeline sizing Own edit

Diameter of the pipe depends on the maximum allowable velocity – erosion velocity –

and pressure drop of the fluid. Flow velocity is restricted in order to avoid erosion of the

pipe wall, reduce noise and allow corrosion inhibition; minimalization of pressure drop

decreases the pressure boosting requirements. In Eq. 80 ‘C’ is empirical constant, its value

depends on the material of pipe, operation continuity and solid content of the fluid.

vmax =C

√ρ (80)

In case of carbon steel and continuous operation ‘C’ equals 122. [43] As one can see in

Table 12, the fluid velocities are lower than vmax in each case, the selected diameter fulfills

the velocity condition.

Determination of vmax

based on density

Selection of nominal

diameter

Pressure drop

calculation

Pressure drop

requirement

fullfilled?

No

Yes

Wall thickness

calculation

Inner and outer

diameter determination

Pressure drop and

maximum velocity

requirements fullfilled?

Yes

End of process

No

Maximum velocity

requirement fullfilled?

Yes

NoSelection of different

nominal diameter

56

Table 11 ‘C’ values for maximum velocity calculation Source: [43]

Condition Value of C

Carbon steel, continuous operation 122

Carbon steel, periodic operation 150

Stainless steel or duplex, continuous operation 250

Stainless steel or duplex, periodic operation 400 – 450

Wall thickness determination follows the process presented in [39]; applied safety factor

determined by [44]. As CO2 is not stable liquid at atmospheric conditions and the pipeline

is not in built-up area the applied safety factor is 1.4.

Table 12 Pipeline parameters – dehydrated gas

Parameter Dense phase Supercritical phase

Volume flow [STDm3/h] 9 838 9 838

Length of pipe [m] 10 000 10 000

Inner diameter [mm] 101.7 101.7

Inlet pressure [barg] 109.3 111.7

Outlet pressure [barg] 105.3 104.7

Pressure drop [bar] 4.0 7.0

Inlet temperature [°C] 30 50

Outlet temperature [°C] 7.8 38.58

Inlet density [kg/m3] 611.9 376.5

Outlet density [kg/m3] 879.9 475.9

Inlet vmax [m/s] 4.932 6.288

Outlet vmax [m/s] 5.981 5.592

Inlet velocity [m/s] 0.998 1.622

Outlet velocity [m/s] 0.694 1.283

Weight of pipe [kg] 16 673 16 673

Insulation required No Yes

Insulation thickness [mm] – 25

Insulation thermal conductivity [W/(mK)] – 0.04

MRU required Yes No

MRU cooling duty [kWth] 457.5 –

57

As one can see, the same pipe dimensions are applicable both for dense and supercritical

phases, the difference between compression duties are negligible. The final decision is based

on economic evaluation. The results are detailed in Chapter 3.8.

Table 13 Safety factors Source: [44]

Area If the substance is stable liquid at

atmospheric condition

If the substance is unstable liquid

and gas at atmospheric condition

Built-up area 1.7 2.0

Object of virtu 1.7 2.0

Other area 1.3 1.4

Table 14 Wall thickness parameters – dehydrated gas

Parameter Dense phase Supercritical phase

Material grade [–] P355N P355N

Yield stress [MPa or N/mm2] 355 355

Safety factor [–] 1.4 1.4

Design stress [MPa or N/mm2] 253.6 253.6

Calculation pressure [barg] 120 120

Calculation temperature [°C] 30 50

Outer diameter [mm] 114.3 114.3

Inner diameter [mm] 101.7 101.7

Joint efficiency [–] 1 1

e [mm] 2.70 2.70

c0 [mm] 2 2

c1 [mm] 1 1

c2 [mm] 0 0

e+c0+c1+c2 [mm] 5.70 5.70

eord [mm] 6.3 6.3

58

3.5. Case 1.2 – Inhibitor Injection to the Wet Stream

The previous chapters discussed the scenario, when the water is removed from the

mixture, thus solving the corrosion and hydrate formation hazard. Application of chemicals

as inhibitor proved to be economical solution for preventing hydrate formation. This chapter

investigates this alternative solution.

Hydrate formation was investigated both with Ng & Robinson and CSM methods, the

results presented by Figure 38. CSM failed to accurately predict the hydrate formation.

The hydrate formation depends on the ratio of the hydrate forming components, not on the

water content. Water content influences the quantity of hydrate formed. The formation

temperature difference between the dehydrated and wet gas (green continuous and dashed

lines) is clearly overestimated. The method also failed in the vicinity of critical point. Further

calculations were not made with this method.

Based on Ng & Robinson method the hydrate formation temperature at the operation

pressure – 110-100 barg – is ~16-17 °C. During normal operation this temperature should

not be reached. This only can occur in case of dense liquid phase operation, as the in order

to transport the substance in supercritical phase, the temperature have to be greater than the

critical temperature (~28 °C).

Figure 38 Hydrate formation conditions Own edit

Hydrate prevention calculations were made for winter – ambient temperature = 5 °C –

and summer – ambient temperature = 15 °C – conditions. At winter the heat loss is greater,

the substance is cooled more effectively by the environment. From hydrate formation point

59

of view, winter operation is the worst scenario. This is the opposite as it would be for natural

gas. Natural gas conditioning involves the cooling of the gas thus, the pipe inlet temperature

is low. At summer the greater ambient temperature causes faster temperature increment, the

gas will reach the hydrate formation temperature sooner.

Accounting in the accuracy of the methods, 5 °C difference between the operational

minimum temperature and the hydrate formation temperature was determined as safety.

The required inhibitor – methanol and EG – quantity was determined for summer and winter

operation, with and without dehydration. The values presented in Table 15 were calculated

with Eq. 11 – 13 from Chapter 2.3.3.1.2.

From Table 15 several conclusions become clear. Winter operation has greater inhibitor

requirement. For inhibition, methanol is more economical solution as it has lower cost

(see Table 21) and lower required quantity.

If the gas is not dehydrated, application of inhibitor is essential in order to avoid hydrate

plugs. If the water is removed from the system one can state with high certainty, that this

low water mass flow will not cause any restriction in the system during normal operation.

Inhibitor may be needed in case of pressurizing or depressurizing the pipeline.

If inhibitor is added to the stream, the same pipe dimensions are applicable as it was for

dehydrated gas. Special polymer coating is suggested to protect the inner wall against

corrosion.

Table 15 Inhibitor requirement for dense liquid phase

Dehydrated Not dehydrated

Water mass flow [kg/h] 0.375 21.43

Hydrate formation temperature [°C] 17 17

Summer (Tamb=15 °C)

Minimum system temperature [°C] 16.89 16.89

Required formation temperature [°C] 11.89 11.89

Methanol required [kg/h] 0.05 2.70

Ethylene glycol required [kg/h] 0.09 5.24

Winter (Tamb=5 °C)

Minimum system temperature [°C] 7.75 7.75

Required formation temperature [°C] 2.75 2.75

Methanol required [kg/h] 0.13 7.54

Ethylene glycol required [kg/h] 0.26 14.62

60

3.6. Case 1.3 – No Dehydration nor Inhibition

As it was discussed in Chapter 3.5, operation without dehydration or inhibition only

possible if the gas is transported in supercritical phase. For this scenario the same pipeline

sizing calculation were made, as before. For additional safety, greater corrosion allowance

was determined, and the inner wall of the pipe is coated with special polymer.

With the greater corrosion allowance the minimal required wall thickness was greater

than the used before (6.3 mm), a next standard thickness (8 mm) had to be used. The greater

wall thickness reduced the inner diameter, causing greater pressure drop. The outlet pressure

was not sufficient with the DN100 pipe, instead DN150 was used. The greater diameter

reduced the pressure drop, on the hand the capital cost increased.

Table 16 Pipe data – not dehydrated gas

Parameter Supercritical phase Unit

Volume flow 9 854 STDm3/h

Inner diameter 152.3 mm

Outer diameter 168.3 mm

Inlet pressure 112.6 barg

Outlet pressure 111.6 barg

Pressure drop 1.0 bar

Inlet temperature 50 °C

Outlet temperature 37.05 °C

Inlet density 399.7 kg/m3

Outlet density 546.0 kg/m3

Inlet maximum velocity 6.102 m/s

Outlet maximum velocity 5.221 m/s

Inlet velocity 0.682 m/s

Outlet velocity 0.4992 m/s

Weight of pipe 31 424 kg

Insulation thickness 25 mm

Insulation thermal conductivity 0.04 W/(mK)

Corrosion allowance 3 mm

eord 8.0 mm

61

3.7. Case 2. – Emitting the Carbon to the Atmosphere

The chapters before dealt with the problems of dehydration and transportation. As an

alternative, emission of CO2 to the atmosphere was investigated.

3.7.1. The EU Emissions Trading System

The EU Emissions Trading System (EU ETS) is a ‘cap and trade’ system. The EU ETS

legislation creates allowances, which are essentially rights to emit greenhouse gases

equivalent to the global warming potential of 1 tonne of CO2 (tCO2e). The system allows

trading of allowances, so that the total emissions of the installations and aircraft operators

stays within the cap and the least-cost measures can be taken up to reduce emissions. The

level of the cap determines the number of allowances available in the whole system.

It is a major tool of the European Union in its efforts to meet emissions reductions targets

now and into the future. As the first and largest emissions trading system for reducing GHG

emissions, the EU ETS covers more than 11 000 power stations and industrial plants

in 31 countries, and flights between airports of participating countries.

Each year, a proportion of the allowances are given to certain participants for free, while

the rest are sold, mostly through auctions. At the end of a year the participants must return

an allowance for every ton of CO2e they emit during that year. If a participant has insufficient

allowances, then it must either take measures to reduce its emissions or buy more allowances

on the market. Participants can acquire allowances at auction, or from each other.

The System caps the total volume of GHG emissions from installations and aircraft

operators responsible for around 50% of EU GHG emissions. Heavy energy-using

installations consisting power generation – power stations and other combustion plants with

≥20 MW thermal rated input (except hazardous or municipal waste installations) –;

oil refineries and gas plants; carbon capture, transport in pipelines and geological storage;

manufacturing – coke ovens, iron and steel, cement clinker, glass, lime, bricks, ceramics,

pulp, paper and board –; aviation sector; aluminium sector; petrochemicals and other

chemicals – ammonia, nitric, adipic and glyoxylic acid production – are all under the

restriction of the EU ETS.

3.7.2. Quantity of Emitted CO2

Let’s assume that the GSU is under the restriction of the EU ETS and must pay the

allowance price. By the 2/2010 KHEM edict, hydrocarbon as a flammable hazardous gas,

cannot be emitted to the atmosphere, the gas mixture must be flared, and the allowance is

payed after the total emitted CO2. The flaring system is already available.

62

First, one has to calculate the total emitted CO2. Assume chemically perfect burning, thus

stochiometric ratios can be used. Based on the general equation of the perfect combustion of

paraffins the total quantity of CO2 can be calculated.

CnH2n+2 + (3n + 1

2)O2 → n CO2 + (n + 1)H2O (81)

Total 9 845 STDm3/h waterless gas leaves the GSU. As the molar and volume percent of

gases are equal, based on the composition in Table 6 and on the molar volume the gas

mixture – calculated from Eq. 16 Vm@15°C, 101325Pa, Z=0.9946 = 23.64 m3/kmol – the molar flow

of each component can be calculated. From Eq. 81 one can see that the mole of formed

carbon dioxide is equal to the moles of carbon can be found in the original hydrocarbon

molecule. The formed carbon dioxide is the product of the molar weight and mole of CO2.

The total emission is the sum of the inherent and the formed carbon dioxide. Table 17

summarizes the values and steps of the calculation.

Table 17 Determination of CO2 emission

Mole/ Volume

fraction [-]

Volume flow

[m3/h]

Molar flow

[kmol/h]

Molar flow of

CO2 [kmol/h]

Mass flow of

CO2 [kg/h]

C1 0.04021 396 16.84 16.84 740.93

C2 0.00235 23 0.98 1.97 86.60

C3 0.00092 9 0.39 1.16 50.86

C4 0.00054 5 0.23 0.90 39.80

C5 0.00024 2 0.10 0.50 22.11

C6 0.00131 13 0.55 3.29 144.83

C7 0.00030 3 0.13 0.88 38.70

C8 0.00007 1 0.03 0.23 10.32

CO2 0.95271 9 380 398.89 398.89 17 555.03

N2 0.00135 13 0.57 0.00 0.00

Total 1.0000 9 845 419 425 18 689

The 18 689 kg/h multiplied by 8 400 – total working hours in a year – gives the total

annular emission as 156 989 116 kg/year or 156 989 ton/year. This value can be assumed as

constant, the cost of ETS allowance was calculated based on this emission.

63

3.8. Economic Evaluation Sub-Cases

Since the GSU is the source of the gas mixture, its expenditures are excluded in all cases.

Total cost can be divided into two groups, OPEX – including energy (electricity and gas),

chemicals and maintenance – and CAPEX – capital cost of apparatus – of air coolers,

compressors, heat exchangers, separators, towers etc. For OPEX calculation, the specific

costs are presented in Table 21, the CAPEX cost of different technological units can be

found in Table 20. Cost of propane for MRU was neglected, as a petroleum company do not

have to buy this substance.

Personal costs are neglected, as the dehydrator technology will be part of the existing

plant. Maintenance is considered as 1.5% of total CAPEX. The price of consumables

– chemicals and energy– was inflated by 1.5% annually. Total 10 years of operation was

investigated.

With air coolers the lowest achievable process stream temperature is 10 °C greater, than

the inlet air temperature. This means, when the average ambient temperature is below 15 °C,

the required 25 °C – inlet temperature of TEG unit – is achievable without the MRU. The

OPEX of the MRU was accounted in only for 6 months annually.

The considered injected methanol quantity was determined as the values presented in

Table 15 multiplied by 1.3 to cover the vaporization and additional losses.

Based on the phase of CO2 during transportation, Case 1.1 – 1.3 can be divided into a)

and b) options, where a) means carbon dioxide transported in dense liquid phase, b) mean it

transported in supercritical phase. In option a) additional cooling is required to reduce the

temperature below the critical, in b) the pipeline covered with heat insulation to keep the

temperature above the critical temperature.

Case 1.1 a) and b) – water removal – technically feasible, there is financial difference.

In Case 1.2 – water is not removed, but inhibitor is added to the stream – in option a) hydrate

will form, methanol is required, in option b) the temperature is high enough to prevent the

hydrate formation, on the other hand corrosion could damage the pipe, inner polymer coating

and greater corrosion allowance were used. Case 1.3 – when the water was not removed;

inhibitor was not added – only feasible in option b), as in option a) hydrate could plug the

pipeline. Polymer inner coating and greater corrosion allowance were used. Case 1.2 b) and

Case 1.3 b) essentially the same, thus reducing the technically feasible scenarios to five –

1.1a, 1.1b, 1.2a, 1.2b and 2.

64

For Case 2. the EU ETS allowance price used, is the average value of the prices published

in [45] and [46]. Based on historical data (see Figure 65) the forecasts did not prove to be

accurate. More up to date long-term forecast was not accessible. To forecast accurate

allowance prices, is not in the scope of this study. These estimates are always reevaluated

and published – like in [47] – by professional analysts, who follow the recent events.

Figure 39 Cost of sub-cases Own edit

Table 18 CAPEX, OPEX, total costs and break-even allowance prices of the sub-cases

Case

Number

CAPEX

[thEUR]

1st year OPEX

[thEUR]

Total Cost

[thEUR]

Break-even allowance price

[EUR/tCO2e]

Case 1.1a 9 829 1 241 23 116 14.80

Case 1.1b 9 602 1 177 22 200 14.21

Case 1.2a 8 913 1 363 23 519 15.06

Case 1.2b 8 887 1 253 22 297 14.28

Case 2 – 4 160 61 304 –

The financial results of the scenarios are close to each other, the greatest difference is

~1.3 million EUR. This difference is within the error range of estimation, considering the

depths of design. It is also obvious, that every solution has greater result, than emitting the

CO2 to the atmosphere and pay the allowance. This is mainly due to the size of the plant and

the fact the most cost intensive technological unit – carbon capture – was not part of the

study. As the total cost of scenarios, the break-even prices are close to each other as well.

65

4. Conclusion and Discussion

The aim of this work was to investigate and compare the options of a water removal and

transportation system for a water saturated CO2 rich gas mixture. For this purpose, an

extensive literature review was carried out. It included the relevant properties of carbon

dioxide; conditions and circumstances of gas hydrate formation and prediction methods;

comparison of different glycols used in dehydration; and different solutions of CO2

transportation, emphasizing the types of pipeline transportation.

In order to quantify the different scenarios a case study was made, including a wide-

spread process simulation model, built with ASPEN HYSYS® v12. Several alternative

scenarios – different chemicals, different processes and solutions – were investigated,

the total costs of 10 years of operation – capital and operational – were compared.

Based on the financial results presented in Chapter 3.8, the suggested method for this

specific problem is to dehydrate the gas with TEG and transport it in supercritical phase to

the end point. This conclusion can be changed with more detailed design process. One should

also bear in mind the following factors have great influence:

- The most cost intensive technology, the carbon capture was not part of the study.

Based on the size of the topic, it could be an independent study.

- For the study, preliminary design was made. The accuracy of this depth cannot be

compared to the accuracy of a detailed design.

- The capital and operational costs are low, due to the size of the plant. A plant this

small could be a pilot plant, in which case positive financial results are not targeted.

- The pipeline is relatively short, cannot be compared to long range pipelines.

- Chemical inhibitors could be economical alternative if the end point allows their

application. However, if the end point is a utilization facility, chemical inhibitors

could pollute the substance, their application should be avoided. Feasibility of

inhibitors also highly depending on the quantity. As the CAPEX and OPEX of

dehydration is not linear with the capacity, the required inhibitor quantity, thus the

cost is linear.

From the study the following conclusion can be drawn:

- Dehydration has great influence on the total cost.

- Chemical inhibitors could be economical alternative for lower quantities, but it may

be not the case for greater amount of gas.

66

The following recommendations are made for further studies:

- Investigation of different chemicals e.g. glycerol and different water removal

processes e.g. molecule sieves.

- Investigation of the effect of the gas quantity, thus the size of the plant.

- Investigation of the effect of transportation distance.

The case study justified the usage of the following methods:

- Figure 15 - 17 for rough determination of TEG requirement, as the results predicted

by them are close to the simulation.

- Results of DEG and TEG calculations – quantity, remaining water content, strip gas

requirement etc. – are consistent with values given in the open literature – e.g. [11]

[29] and [42] –; CPA EoS can predict the reactions of these materials accurately.

On the other hand, it revealed some limitations and errors of published methods.

These are:

- CPA EoS failed to predict the remaining water content of the processed gas stream,

if the TREG quantity was reduced to 60 kg/h or below. The absorber and regenerator

towers could not converge without significantly reducing the error tolerance. This

made the results unreliable.

- The CSM method failed to accurately predict the hydrate formation conditions

– pressure and temperature – of the high CO2 content gas.

- Both CPA and Sour-SRK EoS failed to accurately predict the inhibitor requirement

both for dehydrated and not dehydrated gas.

67

5. References

[1] Enerdata, Global Energy Statistical Yearbook 2020, 24 June 2020. [Online]. Available:

https://yearbook.enerdata.net/.

[2] International Energy Agency, Electricity Information: Overview (2020 edition), 2020.

[3] The Engineering ToolBox, „Carbon Dioxide - Thermophysical Properties,” [Online]. Available:

https://www.engineeringtoolbox.com/CO2-carbon-dioxide-properties-d_2017.html. [12. October

2020].

[4] I. Karamé, H. Srour and J. Shaya, Carbon Dioxide Chemistry, Capture and Oil Recovery, 2018.

[5] Wiley Online Library, Ullmann's Encyclopedia of Industrial Chemistry, Weinheim, Germany , 2014.

[6] J. E. García, Fluid Dynamics of Carbon Dioxide Disposal into Saline Aquifers, Berkeley, California,

2003.

[7] Earth Science Communications Team, Global Climate Change: Vital Signs of the Planet, National

Aeronautics and Space Administration, [Online]. Available: https://climate.nasa.gov/causes/. [12.

October 2020.].

[8] S. R. Weart, The Discovery of Global Warming, Harvard University Press, 2003.

[9] H. Riebeek, The Carbon Cycle, 2016. [Online]. Available:

https://earthobservatory.nasa.gov/features/CarbonCycle.

[10] Air Products and Chemicals, Inc., Carbon Dioxide - Product Stewardship Summary, 2009.

[11] Gas Processors Suppliers Association, Engineering Data Book, Tulsa, Oklahoma, 2012.

[12] DDBST, „Dortmund Data Bank GmbH,” [Online]. Available:

http://www.ddbst.de/en/EED/PCP/VAP_C1050.php. [15 august 2020].

[13] H. Wang, J. Chen and Q. Li, A Review of Pipeline Transportation Technology of Carbon Dioxide,

2019.

[14] J. F. Gabitto and C. Tsouris, Physical Properties of Gas Hydrates: A Review, Journal of

Thermodynamics, p. 12, 2010.

[15] A. J. Kidnay and W. R. Parrish, Fundamentals of Natural Gas Processing, Boca Raton, 2006.

[16] H. Ng and D. Robinson, A method for predicting the equilibrium gas phase water content in gas-

hydrate equilibrium, Industrial & Engineering Chemistry Fundamentals, %1. Vol. 19., 1980.

[17] H. Ng and D. Robinson, The measurement and prediction of hydrate formation in liquid hydrocarbon-

water system, Industrial & Engineering Chemistry Fundamentals, %1. Vol. 15., %1. Num.4., 1976.

[18] H. Ng and D. Robinson, The prediction of hydrate formation in condensed systems, AIChE Journal,

%1. Vol. 23., %1. Num. 4., 1977.

[19] E. D. Sloan and C. A. Koh, Clathrate Hydrates of Natural Gases, New York: Taylor & Francis Group,

2007.

68

[20] A. Ballard and E. S. Jr, The next generation of hydrate prediction I. Hydrate standard states and

incorporation of spectroscopy, Fluid Phase Equilibria, pp. 371-383, 2002.

[21] M. Jager, A. Ballard and E. S. J. , The next generation of hydrate prediction II. Dedicated aqueous

phase fugacity model for hydrate prediction, Fluid Phase Equilibria, %1. Vol. 211, pp. 85-107, 2003.

[22] D. Peng and D. B. Robinson, A New Two-Constant Equation of State, 1976.

[23] R. Sashay, M. Edison and B. Mohamed, A Review of the Equations of State and their Applicability in

Phase Equilibrium Modeling, Johannesburg, 2013.

[24] G. Soave, Equilibrium constants from a modified Redlich-Kwong equation of state, 1972.

[25] G. M. Kontogeorgis, E. Voutsas, I. Yakoumis and D. Tassios, An Equation of State for Associating

Fluids, Ind. Eng. Chem. Res, 1996.

[26] G. Kontogeorgis and G. Folas, Thermodynamic Models for Industrial Applications: From Classical

and Advanced Mixing Rules to Association Theories, Chichester, West Sussex, UK: John Wiley &

Sons, 2010.

[27] G. M. Wilson, A New Correlation of NH3, CO2, and FLS Volatility Data from Aqueous Sour Water

Systems, 1980.

[28] J. M. Campbell, Gas Conditioning and Processing Volume 2: The Equipment Modules, 2012.

[29] J. M. Campbell, Gas Conditioning and Processing Volume 1: The Basic Principles, 2012.

[30] MEGlobal, „MEGlobal An EQUATE Company,” [Online]. Available:

https://www.meglobal.biz/products-and-applications/product-literature/. [17 August 2020].

[31] F. K., CO2 Interim Storage as a Tool for CO2 Market Development: a Comprehensive Technical

Assessment. Master Thesis, Stanford University, Stanford, California, 2011.

[32] Economic Commission for Europe Inland Transport Comittee, Agreement Concerninig the

International Carriage of Dangerous Goods by Road, New York and Geneva: United Nations, 2020.

[33] A. Witkowski, M. Majkut and S. Rulik, Analysis of pipeline transportation systems for carbon dioxide

sequestration, Archives of Thermodynamics, pp. 117-140, 2014.

[34] Society of Petroleum Engineers, Petroleum Engineering Handbook Volume III Facilities and

Construction Engineering, Richardson, Texas: Society of Petroleum Engineers, 2006.

[35] Fekete Associates Inc., IHS Markit Energy, [Online]. Available:

http://www.fekete.com/san/webhelp/piper/webhelp/c-te-pressure.htm. [18 February 2021].

[36] G. Takács, Production Engineering Fundamentals Vol. 1, Miskolc: Miskolci Egyetemi Kiadó, 2012.

[37] J. P. B. H. Dale Beggs, A Study of Two-Phase Flow in Inclined Pipes, Journal of Petroleum

Technology, pp. 607-617, May 1973.

[38] J. P. Holman, Heat Transfer, New York, New York: McGraw-Hill, 2010.

[39] Magyar Szabványügyi Testület, MSZ EN 13480-3 Metallic industrial piping Part 3: Design and

calclation, Brussels: Magyar Szabványügyi Testület, 2002.

[40] National Association of Corrosion Engineers, Corrosion Basics, An Introduction, Houston, Texas:

National Association of Corrosion Engineers, 1984.

69

[41] H. M. A. El-Lateef, V. M. Abbasov, L. I. Aliyeva and T. A. Ismayilov, Corrosion Protection of Steel

Pipelines Against CO2 Corrosion-A Review, Chemistry Journal, pp. 52-63, 2012.

[42] L. E. Øi and M. Fazlagic, Glycol Dehydration of Captured Carbon Dioxide Using Aspen HYSYS

Simulation, 2014.

[43] American Petroleum Institute, Recommended Practice for Design and Installation of Offshore

Production Platform Piping Systems, Dallas, Texas: American Petroleum Institute, 2019.

[44] Közlekedési, Hírközlési és Energiaügyi Minisztérium , 2/2010. (I. 14.) KHEM rendelet a Kőolaj- és

Földgázbányászati Biztonsági Szabályzatról, Budapest: Közlekedési, Hírközlési és Energiaügyi

Minisztérium , 2010.

[45] S. Schjølset, „Reuters: Business & Financial News, US & Internatinal,” [Online]. Available:

https://ec.europa.eu/clima/sites/clima/files/docs/0094/thomson_reuters_point_carbon_en.pdf. [30

March 2021].

[46] M. Mazzoni és P. Ruff, „ICIS: Independent Commodity Intelligence Services,” [Online]. Available:

https://www.icis.com/explore/resources/higher-carbon-prices-on-utilities-and-industries/. [30 March

2021].

[47] S. Twidale, „Reuters: Business & Financial News, US & Internatinal,” [Online]. Available:

https://www.reuters.com/article/idUSKBN29N0ZJ. [30 March 2021].

[48] Ember NPO, „Daily EU ETS carbon market price (Euros),” [Online]. Available: https://ember-

climate.org/data/carbon-price-viewer/. [31 March 2021].

[49] European Union, EU ETS Handbook, 2015.

[50] World Resources Institute, Guidelines for Carbon Dioxide Capture, Transport, and Storage,

Washington, DC, 2002.

[51] J. A. Dean, Lange's Handbook of Chemistry, McGraw-Hill Professional, 1998.

[52] The Engineering ToolBox, „Carbon Dioxide Gas - Specific Heat,” [Online]. Available:

https://www.engineeringtoolbox.com/carbon-dioxide-d_974.html. [12. October 2020].

[53] The Engineering ToolBox, „Carbon dioxide - Thermal Conductivity,” [Online]. Available:

https://www.engineeringtoolbox.com/carbon-dioxide-thermal-conductivity-temperature-pressure-

d_2019.html. [12. October 2020].

[54] Sandbag, „Smarter Climate Policy,” [Online]. Available: https://sandbag.be/index.php/carbon-price-

viewer/. [22 January 2021].

[55] E. E, Process simulation of natural gas dehydration by absorption in triethylene glycol. Master Thesis,

Telemark University College, Porsgrunn, Norway, 2012.

[56] B. V., Conditioning of CO2 coming from a CO2 capture process for transport and storage purposes.

Master Thesis, NTNU, Trondheim, Norway, 2009.

[57] H. Li, Thermodynamic Properties of CO2 Mixtures and Their Applications in Advanced Power Cycles

with CO2 Capture Processes, Royal Institute of Technology, Stockholm, Sweden, 2008.

70

Appendix A – Final Configuration of TEG Unit

Table 19 TEG unit parameters

Parameter Value Unit

Inlet Gas

Pressure 40.5 barg

Temperature 25 °C

Water content 602.4 mg/STDm3

Volume Flow 9 851 STDm3

Mass Density 92.52 kg/m3

Outlet Gas

Pressure 40.35 barg

Temperature 25.65 °C

Water content 40.17 mg/STDm3

49.98 ppmv

Volume Flow 9 839 STDm3

Lean TEG

Pressure 41.5 barg

Temperature 24.80 °C

Concentration 98.85 wt%

Mass Flow 93.15 kg/h

Rich TEG

Pressure 40.5 barg

Temperature 25.52 °C

Concentration 84.43 wt%

Mass Flow 108.9 kg/h

T-100 contactor tower

Number of theoretical stages 4 –

Head pressure 40.35 barg

Head temperature 25.80 °C

Bottom pressure 40.5 barg

Bottom temperature 26.56 °C

Type of contractor Structured Packing –

71

Thickness of one packing unit 200 mm

Additional allow space 3 m

Cbubble tray 176 –

Cstructured packing 384 –

Specific area of the packing (As) 300 m2/m3

Diameter 700.0 mm

Height 8 500 mm

T-101 regenerator tower

Reboiler temperature 204 °C

Reboiler Duty 41.91 kWth

Head temperature 62.39 °C

Strip gas mass flow 35 kg/h

Other equipment

TEG pump duty 0.13 kWe

R/L TEG heat exchanger duty 12.61 kWth

72

Appendix B – Cost of Materials

All values are based on previous informational bids and existing technologies. The values

have been modified with an unpublished multiplier in order to respect secrecy.

Table 20 Fixed costs

Apparatus Cost Currency

Pressure Boosting (1st – 3rd stage) 4 356 thEUR

Pressure Boosting (4th stage) 1 425 thEUR

Mechanical Refrigeration Unit 1 375 thEUR

Mechanical Refrigeration Unit (with dense phase condenser) 1 709 thEUR

Glycol Unit 910 thEUR

Table 21 Specific costs

Cost Element Cost Currency

Electricity 82 EUR/kWh

Methanol 700 EUR/ton

EG 910 EUR/ton

DEG 1 050 EUR/ton

TEG 1 400 EUR/ton

TREG 4 550 EUR/ton

Propane 570 EUR/ton

Table 22 EU ETS allowance cost

Year Average Annular Price Currency

2021 26.5 EUR/ton

2022 30 EUR/ton

2023 34.5 EUR/ton

2024 37 EUR/ton

2025 37.5 EUR/ton

2026 38 EUR/ton

2027 40 EUR/ton

2028 45 EUR/ton

2029 50 EUR/ton

2030 52 EUR/ton

73

Appendix C – Additional Graphs

Figure 40 Density of CO2 – high resolution Own edit

Figure 41 Viscosity of CO2 – high resolution Own edit

74

Figure 42 Heat capacity of CO2 – wide interval Own edit

Figure 43 Mass enthalpy of CO2 Own edit

75

Figure 44 Mass entropy of CO2 Own edit

Figure 45 Pressure – Temperature diagram of gas mixture Own edit

76

Figure 46 Water content of gas mixture (low temperature) Own edit

Figure 47 Water content of gas mixture (high temperature) Own edit

77

Figure 48 Reboiler duty vs strip gas quantity (relative) Own edit

Figure 49 Reboiler duty vs strip gas temperature (relative) Own edit

78

Figure 50 Remaining water content vs contractor temperature (DEG) Own edit

Figure 51 DEG mass fraction vs strip gas quantity Own edit

79

Figure 52 Remaining water content vs strip gas quantity (DEG) Own edit

Figure 53 DEG mass fraction vs mass flow Own edit

80

Figure 54 Reboiler duty vs DEG mass flow Own edit

Figure 55 DEG mass fraction vs reboiler temperature with different strip gas quantities Own edit

81

Figure 56 Absorber performance at lean DEG concentration of 98.85 wt% Own edit

Figure 57 Remaining water content vs contractor temperature (TREG) Own edit

82

Figure 58 TREG mass fraction vs mass flow Own edit

Figure 59 Remaining water content vs strip gas quantity (TREG) Own edit

83

Figure 60 TREG mass fraction vs strip gas quantity Own edit

Figure 61 TREG mass fraction vs reboiler temperature with different strip gas quantities Own edit

84

Figure 62 Absorber performance at lean TREG concentration of 97.85 wt% Own edit

Figure 63 Glycol viscosity – wide interval Own edit

85

Figure 64 Glycol viscosity – tight interval Own edit

Figure 65 EU ETS allowance historical prices and forecast Based on: [48]