DESIGN AND ECONOMIC FEASIBILITY OF A MODULAR
NATURAL GAS PROCESSING PLANT
BY
JAIYESIMI OLUSEGUN ADURAGBEMI
(CGRP/MSC/RPE/2014/020)
A THESIS SUBMITTED TO THE SCHOOL OF GRADUATE STUDIES,
UNIVERSITY OF PORT HARCOURT
IN PARTIAL FULFILLMENT OF THE REQUIREMENT FOR THE
AWARD OF THE MASTERS OF SCIENCE DEGREE IN GAS,
REFINING AND PETROCHEMICALS
CENTRE FOR GAS, REFINING AND PETROCHEMICALS
INSTITUTE OF PETROLEUM STUDIES
UNIVERSITY OF PORT-HARCOURT, CHOBA
RIVERS STATE, NIGERIA
(NOVEMBER, 2015)
a
CERTIFICATION
UNIVERSITY OF PORT HARCOURT
CENTER FOR GAS, REFINING AND PETROCHEMICALS
DESIGN AND ECONOMIC FEASIBILITY OF A MODULAR NATURAL
GAS PROCESSING PLANT
BY
JAIYESIMI OLUSEGUN ADURAGBEMI
(CGRP/MSC/RPE/2014/020)
THE BOARD OF EXAMINERS DECLARES AS FOLLOWS: THAT
THIS WORK IS THE ORIGINAL WORK OF THE CANDIDATE. THAT
IS ACCEPTED IN PARTIAL FULFILMENT OF THE REQUIREMENT
FOR THE AWARD OF THE MASTERS OF SCIENCE DEGREE IN
GAS, REFINING AND PETROCHEMICALS
Prof. Joel Ogbonna
(Supervisor)
…...……….
Signature
……...…….
Date
Prof. Igwe Godwin
(Co-supervisor)
……...…….
Signature
……...…….
Date
Dr. M. K Oduola
(Director CGRP)
……...…….
Signature
……...…….
Date
(External Examiner)
……...…….
Signature
……...…….
Date
Prof. O. M. O Etebu
(Chairman of Board)
……...…….
Signature
……...…….
Date
b
DEDICATION
I dedicate this masters thesis to the almighty God who was my source of strength
and inspiration through the course of my research.
i
ACKNOWLEDGEMENTS
I cannot afford not to acknowledge Professor Ogbonna Joel for his supervision and
guidance throughout the course of my research. Very important also are my father
and siblings who all supported me both financially and morally. Without them,
perhaps successful completion might not have been possible. Finally, i must not
forget to mention Professor Godwin Igwe for not just being my director, but also a
father and Dr. Koyejo Oduola for his selflessness when handling issues concerning
our welfare particularly during the internship.
ii
ABSTRACT
This work looked to design and carryout economic analysis of a modular natural gas
processing plant. So, an initial study into the processing stages involved in natural
gas processing was made. Four major stages are involved. These include removal of
acid gases from the raw natural gas, dehydration of the natural gas, recovery of the
natural gas liquids and fractionation of the recovered NGLs. The fourth stage is
optional, depending mostly on size of the plant facility and product prices. The
design of a 20.8 million standard cubic feet per annum natural gas processing plant
was done using aspen hysys simulation software. The acid gas removal unit utilized
diethanol amine solution while the dehydration unit utilized triethylene glycol. The
final products of the plant are methane, ethane, propane, butane and natural
gasoline. Finally, economic analysis was carried out with the use of hysys to
evaluate the capital and operating costs of the plant and Microsoft excel spreadsheet
to check for project feasibility and profitability. The estimated capital cost of the
project is about 14.5 million US Dollars.
iii
TABLE OF CONTENTS
DEDICATION ........................................................................................................................................ a
ACKNOWLEDGEMENTS ................................................................................................................... ii
ABSTRACT .......................................................................................................................................... iii
TABLE OF CONTENTS ..................................................................................................................... iv
LIST OF FIGURES.............................................................................................................................. vi
LIST OF TABLES ............................................................................................................................... vii
CHAPTER ONE.................................................................................................................................... 1
INTRODUCTION .................................................................................................................................. 1
1.1
Background to the study ..................................................................................................... 1
1.2
Statement of the problem .................................................................................................... 2
1.3
Objective of the study .......................................................................................................... 2
1.4
Significance of the research ............................................................................................... 3
CHAPTER TWO ................................................................................................................................... 4
LITERATURE REVIEW ....................................................................................................................... 4
2.1
Natural Gas Processing ...................................................................................................... 4
2.2
Acid Gas Treating................................................................................................................. 7
2.3
Natural Gas Dehydration................................................................................................... 12
2.4
Natural Gas Liquids Recovery.......................................................................................... 14
2.5
NGL Fractionation .............................................................................................................. 18
2.6
Modular Process Skid ........................................................................................................ 19
CHAPTER THREE ............................................................................................................................. 22
RESEARCH METHODOLOGY ........................................................................................................ 22
3.1
Overview of the Methodology ........................................................................................... 22
3.2
Design Approach ................................................................................................................ 23
3.3
Economic Approach ........................................................................................................... 29
CHAPTER FOUR ............................................................................................................................... 36
RESULTS AND DISCUSSION......................................................................................................... 36
4.1
Design Approach ................................................................................................................ 36
4.2
Economic Approach ........................................................................................................... 44
iv
CHAPTER FIVE ................................................................................................................................. 51
CONCLUSIONS AND RECOMMENDATIONS ............................................................................. 51
5.1
Conclusions ......................................................................................................................... 51
5.2
Recommendations ............................................................................................................. 52
APPENDIX A....................................................................................................................................... 55
APPENDIX B....................................................................................................................................... 61
APPENDIX C ...................................................................................................................................... 64
APPENDIX D ...................................................................................................................................... 65
APPENDIX E....................................................................................................................................... 66
v
LIST OF FIGURES
Figure
2.1: Simplified typical onshore treatment process............................................4
2.2: Schematic of amine gas-sweetening process flow diagram....................13
2.3: Simplified flow diagram for TEG dehydration..........................................15
2.4: Natural gas phase diagram.....................................................................16
2.5: Typical flow sheet of a cryogenic refrigeration plant................................17
2.6: Simplified flow diagram of a fractionation plant.......................................19
3.1: Overall methodology (design and economic approach)..........................22
3.2: Main flowsheet of the simulated natural gas processing plant................24
3.3: Simulated flowsheet of the acid removal unit..........................................25
3.4: Simulated flowsheet of the dehydration unit............................................26
3.5: Simulated flowsheet of the NGLs recovery unit.......................................27
3.6: Simulated flowsheet of the deethanizer...................................................28
3.7: Simulated flowsheet of the depropanizer................................................28
3.8: Simulated flowsheet of the debutanizer...................................................29
vi
LIST OF TABLES
Table
2.1: Natural Gas Specifications in the Sellable Gas Stream.............................5
3.1: List of the economic factors.....................................................................31
4.1: Thermodynamic properties of the key streams........................................36
4.2: Streams composition of the acid gas removal unit..................................38
4.3: Streams composition of the dehydration unit...........................................39
4.4: Streams composition of the NGLs recovery unit......................................40
4.5: Streams composition of the deethanizer..................................................41
4.6: Streams composition of the depropanizer...............................................42
4.7: Streams composition of the debutanizer..................................................43
4.8: Values of the economic factors................................................................44
4.9: Price ranges for raw gas and products....................................................45
4.10: Price feasibility analysis table................................................................46
4.11: Economic analysis based on raw material price....................................47
4.12: Economic analysis based on methane price..........................................47
4.13: Economic analysis based on NGLs composite price.............................48
4.14: Economic analysis based on interest rate.............................................48
4.15: Economic analysis based on escalation rates.......................................49
4.16: Economic analysis based on depreciation duration...............................49
4.17: Economic analysis based on tax rate....................................................50
A.1: List of Equipment and their Costs............................................................55
B.1: List of Utility Requirements by Equipment and Costs..............................61
vii
C.1: Total Utility Requirements and their Prices.............................................64
D.1: Project Capital Summary.........................................................................65
E.1: Cash Flow Analysis Results....................................................................66
viii
CHAPTER ONE
INTRODUCTION
1.1 Background to the study
Historically, natural gas was discovered as a consequence of prospecting for crude
oil. Natural gas was often an unwelcome by-product, as natural gas reservoirs were
tapped in the drilling process and workers were forced to stop drilling to let the gas
vent freely into the air. Now, and particularly after the crude oil shortages of the
seventies, natural gas has become an important source of energy in the world.
Natural gas produced from geological formations comes in a wide array of
compositions. The varieties of gas compositions can be broadly categorized into
three distinct groups: (1) non-associated gas that occurs in conventional gas fields,
(2) associated gas that occurs in conventional oil fields, and (3) continuous (or
unconventional) gas (Mokhatab et al., 2006).
Natural gas is a fossil fuel made primarily of methane, hydrogen and carbon.
However, it was not until recently that methods for obtaining this gas, bringing it to
the surface, and putting it to use were developed.
In Nigeria, Gas utilization is a primary goal of Nigeria’s petroleum and energy
policies. This is because, with a proven reserve of over 180 trillion cubic feet of
natural gas, Nigeria’s gas reserve is as much as the nation’s crude oil resources in
terms of value. Until recently, associated gas encountered during the normal course
of oil production has been largely flared. Starting from a very high baseline, gas
flaring has more than halved in Nigeria since 1996. Despite this, Nigeria remains the
second largest flaring country in the world and emits around $2.5 billion worth of gas
annually. But what is more surprising than the eye-popping economic loss is the fact
that so much gas is wasted despite the country’s rampant energy poverty.
In 2013, Nigerian gas production amounted to some 7.7 billion cubic feet per day
with some 15% being flared, 27% reinjection and the balance split between NLNG
feedstock, internal fuel usage, and a small percentage marketed as LPG.
1
There are different methods of developing and utilizing the natural gas resources.
The principal drivers for the development of natural gas are usually pressure to
reduce flaring, desire for economic growth and general enhancement of populace
quality of life, and desire for industrial development. On the other hand, the principal
barriers for the development are structure of investment (large investments in
pipelines and distribution systems are needed), inappropriate domestic pricing policy
(government policy may also heavily influence gas pricing, for example, through
social or sector policies).
Power generation, gas-to-liquids (GTL), LPG processing, fertilizer (ammonia/urea),
liquefied natural gas (LNG), methanol, petrochemicals are some of the development
options available for monetization of the natural gas.
This thesis looks into the economic feasibility of designing, constructing and
operating a modular gas processing plant as a link and opportunity in natural gas
utilization in Nigeria.
1.2 Statement of the problem
There are a number of general factors driving the need to reduce the gas volumes
being flared:

Flaring represents a significant economic loss.

Combustion products make a major contribution to environmental damage
through production of greenhouse gases.

Scarce information on economic feasibility of setting up a small scale natural
gas processing plant.

Energy is wasted by flaring of natural gas, but there is energy poverty in the
country.
1.3 Objective of the study
The key to succeed in any industry is “profitability.” This research is aimed at the
providing the following:

Detailed design of a natural gas processing plant that meets pipeline
specification.

Estimate of every equipment in the plant as well as overall cost of the project.

Estimate of every cost of operating the plant.
2

Economic analysis of the natural gas processing plant in order to determine
under which conditions it is economically feasible.

Detailed report that can serve as a guide for an entrepreneur that is willing to
invest in this project.
1.4 Significance of the research
There are so many benefits attached to this research if successful.

Reduction in environmental pollution due to natural gas flaring.

There will be information available to encourage smaller investors into setting
up modular gas processing plants.

Setting up modular gas processing plants will also create jobs.

Natural gas is an alternative source of energy.

Outputs of gas processing plant can be feedstock for the petrochemical
industry.

Sufficient supply of cooking gas to meet the growing awareness and demand.

Communities and societies in which the plants operate will also benefit.
3
CHAPTER TWO
LITERATURE REVIEW
2.1 Natural Gas Processing
Raw natural gas after transmission through the field-gathering network must be
processed before it can be moved into long-distance pipeline systems for use by
consumers. The objective of gas processing is to separate natural gas, condensate,
non-condensable, acid gases, and water from a gas-producing well and condition
these fluids for sale or disposal. The typical process operation modules are shown in
figure 2.1. Each module consists of a single piece or a group of equipment
performing a specific function. All the modules shown will not necessarily be present
in every gas plant. In some cases, little processing is needed; however, most natural
gas requires processing equipment at the gas processing plant to remove impurities,
water, and excess hydrocarbon liquid and to control delivery pressure. The unit
operations used in a given application may not be arranged in the sequence shown
in figure 2.1, although this sequence is typical.
Recovery Unit
Transported
Raw Gas
Phase
Separation
Solids
Water
Gas Treating
Condensate to Stabilization
Unit
Water
Inlet Gas
Compression
Recovery of Natural
Gas Liquids
Dehydration
Sales Gas
Compression
Gas to Pipeline,
Reinjection,
Flare
NGL to Fractionation
Figure 2.1: Simplified typical onshore treatment process (Mokhatab et al., 2006).
4
The choice of modules to be used and the arrangement of these modules are
determined during the design stage of each gas-field development project (Meyer
and Sharma, 1980).
The important factors that usually determine the extent of gas processing include the
processing objectives, the type or source of the gas, and the location and size of the
gas fields.
2.1.1 Processing Objectives
Listed below are the major objectives of natural gas processing (Mokhatab et al.,
2006):

To produce a sales gas stream that meets specifications of the type shown in
table 2.1. These specifications are mainly intended to meet pipeline
requirements and the needs of industrial and domestic consumers.

To maximize NGLs production by producing a lean gas stripped of most of the
hydrocarbons other than methane.

To deliver a commercial gas. Such gas must be distinguished by a certain
range of gross heating value lying.
Table 2.1: Natural Gas Specifications in the Sellable Gas Stream (Goar and
Arrington, 1978).
Characteristic
Specification
Water content
4–7 lb/MMscf (max)
Hydrogen sulfide content
1/4 grain/100 scf (max)
Gross heating value
950 Btu/scf (min)
Hydrocarbon dew point
15◦F at 800 psig (max)
Mercaptan content
0.2 grain/100 scf (max)
Total sulfur content
1–5 grain/100 scf (max)
Carbon dioxide content
1–3 mole percent (max)
Oxygen content
0–0.4 mole percent (max)
Sand, dust, gums, and free liquid
Commercially free
Typical delivery temperature
120oF
Typical delivery pressure
714.7 psia
5
2.1.2 Effect of Gas Type in Field Processing
The gas composition of the field is the most important issue in choosing a processing
scheme. In other words, depending on the type of reservoir and the composition of
the produced gas, the gas processing plant may contain extensive facilities for the
processing of the associated liquefiable hydrocarbons. Typically, associated gas is
very rich in liquefiable hydrocarbons and must undergo NGL and condensate
recovery to meet hydrocarbon dew point or minimum heating value requirements.
The gas processing scheme will also be dictated by the format of the sales contract
and its specifications. The contract may be totally different for each customer
depending on the composition and amount of gas, plant recoveries, and the
contractual preferences of the customer (Bullin and Hall, 2000).
2.1.3 Location of the Gas Field
The productivity of a gas reservoir can vary greatly and depend primarily on type,
location, and age. Because the location and output of the wells can vary widely, then
not surprisingly, the systems that have been designed to collect and process this
output also vary widely (Thorn et al., 1999). There are at least two aspects of
location
that
are
important:
remoteness
and
local
temperature
variation.
Temperature affects the tendency for hydrate formation in the gas gathering network.
Offshore platforms and “outback” are examples of remote locations. Even these
locales are not strictly comparable because one is sea based versus dry land based.
For the sea-based facility the produced fluid from each wellhead flows via a flow line
into a manifold and from there to the process facilities located on the platform deck.
Ship platforms are extremely limited with respect to size and allowable weight and
only those operations absolutely needed are performed. Facilities on the offshore
platform will generally process the gas to produce a low water content hydrocarbon
stream for export to shore through the subsea pipelines. This process ensures
minimal corrosion, as well as minimizing the potential for hydrate formation in the
raw gas pipeline. A dry-land outback facility has essentially unlimited area available
and can support operations not practical or desirable offshore, such as treating or
processing involving fire hazards (Manning and Thompson, 1991).
6
2.2 Acid Gas Treating
Natural gas, while ostensibly being hydrocarbon in nature, contains large amounts of
acid gases, such as hydrogen sulfide and carbon dioxide. Natural gas containing
hydrogen sulfide or carbon dioxide is referred to as sour, and natural gas free from
hydrogen sulfide is referred to as sweet. The corrosiveness nature of hydrogen
sulfide and carbon dioxide in the presence of water (giving rise to an acidic aqueous
solution) and because of the toxicity of hydrogen sulfide and the lack of heating
value of carbon dioxide, natural gas being prepared for sales is required to contain
no more than 5 ppm hydrogen sulfide and to have a heating value of no less than
920 to 980 Btu/scf. The actual specifications depend on the use, the country where
the gas is used, and the contract. However, because natural gas has a wide range of
composition, including the concentration of the two acid gases, processes for the
removal of acid gases vary and are subject to choice based on the desired end
product (Mokhatab et al., 2006).
There are many variables in treating natural gas. The precise area of application of a
given process is difficult to define. Several factors must be:

Types and concentrations of contaminants in the gas

The degree of contaminant removal desired

The selectivity of acid gas removal required

The temperature, pressure, volume, and composition of the gas to be
processed

The carbon dioxide–hydrogen sulfide ratio in the gas

The desirability of sulfur recovery due to process economics or environmental
issues.
In addition to hydrogen sulfide and carbon dioxide, gas may contain other
contaminants, such as mercaptans and carbonyl sulfide. The presence of these
impurities may eliminate some of the sweetening processes, as some processes
remove large amounts of acid gas but not to a sufficiently low concentration.
However, there are those processes that are not designed to remove (or are
incapable of removing) large amounts of acid gases. These processes are also
capable of removing the acid gas impurities to very low levels when the acid gases
are there in low to medium concentrations in the gas.
7
Process selectivity indicates the preference with which the process removes one
acid gas component relative to (or in preference to) another. For example, some
processes remove both hydrogen sulfide and carbon dioxide; other processes are
designed to remove hydrogen sulfide only. It is important to consider the process
selectivity for, say, hydrogen sulfide removal compared to carbon dioxide removal
that ensures minimal concentrations of these components in the product, thus the
need for consideration of the carbon dioxide to hydrogen sulfide in the gas stream
(Mokhatab et al., 2006).
2.2.1 Acid Gas Removal Processes
The processes that have been developed to accomplish gas purification vary from a
simple once-through wash operation to complex multistep recycling systems
(Speight, 1993). In many cases, process complexities arise because of the need for
recovery of the materials used to remove the contaminants or even recovery of the
contaminants in the original or altered form (Kohl and Riesenfeld, 1985; Newman,
1985).
There are two general processes used for acid gas removal: adsorption and
absorption (Speight, 1993). Adsorption is a physical–chemical phenomenon in which
the gas is concentrated on the surface of a solid or liquid to remove impurities.
Usually, carbon is the adsorbing medium, which can be regenerated upon desorption
[Speight (1993, 1999)]. The quantity of material adsorbed is proportional to the
surface area of the solid and, consequently, adsorbents are usually granular solids
with a large surface area per unit mass. Subsequently, the captured gas can be
desorbed with hot air or steam either for recovery or for thermal destruction.
Adsorbers are widely used to increase a low gas concentration prior to incineration
unless the gas concentration is very high in the inlet air stream. Adsorption is also
employed to reduce problem odours from gases. There are several limitations to the
use of adsorption systems, but it is generally felt that the major one is the
requirement for minimization of particulate matter and/or condensation of liquids
(e.g., water vapour) that could mask the adsorption surface and reduce its efficiency
drastically. Absorption differs from adsorption in that it is not a physical–chemical
surface phenomenon, but an approach in which the absorbed gas is ultimately
distributed throughout the absorbent (liquid). The process depends only on physical
8
solubility and may include chemical reactions in the liquid phase (chemisorption).
Common absorbing media used are water, aqueous amine solutions, caustic,
sodium carbonate, and non-volatile hydrocarbon oils, depending on the type of gas
to be absorbed. Usually, the gas–liquid contactor designs that are employed are
plate columns or packed beds.
Absorption is achieved by dissolution (a physical phenomenon) or by reaction (a
chemical phenomenon). Chemical adsorption processes adsorb sulfur dioxide onto a
carbon surface where it is oxidized (by oxygen in the flue gas) and absorbs moisture
to give sulfuric acid impregnated into and on the adsorbent.
As currently practiced, acid gas removal processes involve the chemical reaction of
the acid gases with a solid oxide (such as iron oxide) or selective absorption of the
contaminants into a liquid (such as ethanolamine) that is passed counter-current to
the gas. Then the absorbent is stripped of the gas components (regeneration) and
recycled to the absorber. The process design will vary and, in practice, may employ
multiple absorption columns and multiple regeneration columns (Mokhatab et al.,
2006).
Liquid absorption processes, which usually employ temperatures below 50 OC
(120OF), are classified either as physical solvent processes or as chemical solvent
processes. The former processes employ an organic solvent, low temperatures, or
high pressure. In chemical solvent processes, absorption of the acid gases is
achieved mainly by use of alkaline solutions such as amines or carbonates (Kohl and
Riesenfeld, 1985). Regeneration (desorption) can be brought about by the use of
reduced pressures and/or high temperatures, whereby the acid gases are stripped
from the solvent.
Amine washing of natural gas involves chemical reaction of the amine with any acid
gases with the liberation of an appreciable amount of heat and it is necessary to
compensate for the absorption of heat. Amine derivatives such as ethanolamine
(monoethanolamine),
diethanolamine,
triethanolamine,
methyldiethanolamine,
diisopropanolamine, and diglycolamine have been used in commercial applications
(Kohl and Riesenfeld, 1985; Speight, 1993; Polasek and Bullin, 1994).
9
2.2.1.1
Amine Processes
Chemical absorption processes with aqueous alkanolamine solutions are used for
treating gas streams containing hydrogen sulfide and carbon dioxide. However,
depending on the composition and operating conditions of the feed gas, different
amines can be selected to meet the product gas specification. Amines are
categorized as being primary, secondary, and tertiary depending on the degree of
substitution of the central nitrogen by organic groups. Primary amines react directly
with H2S, CO2, and carbonyl sulfide (COS). Examples of primary amines include
monoethanolamine (MEA) and the proprietary diglycolamine agent (DGA).
Secondary amines react directly with H2S and CO2 and react directly with some
COS. The most common secondary amine is diethanolamine (DEA), while
diisopropanolamine (DIPA) is another example of a secondary amine, which is not
as common anymore in amine-treating systems. Tertiary amines react directly with
H2S, react indirectly with CO2, and react indirectly with little COS. The most common
examples of tertiary amines are methyldiethanolamine (MDEA) and activated
methyldiethanolamine (Mokhatab et al., 2006).
Processes using ethanolamine and potassium phosphate are now widely used. The
ethanolamine process, known as the Girbotol process, removes acid gases
(hydrogen sulfide and carbon dioxide) from liquid hydrocarbons as well as from
natural and from refinery gases. The Girbotol treatment solution is an aqueous
solution of ethanolamine, which is an organic alkali that has the reversible property
of reacting with hydrogen sulfide under cool conditions and releasing hydrogen
sulfide at high temperatures. The ethanolamine solution fills a tower called an
absorber through which the sour gas is bubbled. Purified gas leaves the top of the
tower, and the ethanolamine solution leaves the bottom of the tower with the
absorbed acid gases. The ethanolamine solution enters a reactivator tower where
heat drives the acid gases from the solution. Ethanolamine solution, restored to its
original condition, leaves the bottom of the reactivator tower to go to the top of the
absorber tower, and acid gases are released from the top of the reactivator
(Mokhatab et al., 2006).
Depending on the application, special solutions such as mixtures of amines; amines
with physical solvents, such as sulfolane and piperazine; and amines that have been
10
partially neutralized with an acid such as phosphoric acid may also be used (Bullin,
2003).
The proper selection of the amine can have a major impact on the performance and
cost of a sweetening unit. However, many factors must be considered when
selecting an amine for a sweetening application (Polasek and Bullin, 1994).
Considerations for evaluating an amine type in gas treating systems are numerous. It
is important to consider all aspects of the amine chemistry and type, as the omission
of a single issue may lead to operational issues. While studying each issue, it is
important to understand the fundamentals of each amine solution.
While many of the recent published papers concerning amine selection or amine
conversions have focused on the utilization of MDEA over the older generic amines
(Bullin et al., 1990; Polasek et al., 1992), there are many recent cases where these
older generic amines have been the best and, even perhaps, the only choice for
recent new plant design (Jenkins and Haws, 2002). MEA and DEA have found the
most general application in the sweetening of natural gas streams. Even though a
DEA system may not be as efficient as some of the other chemical solvents are, it
may be less expensive to install because standard packaged systems are readily
available. In addition, it may be less expensive to operate and maintain (Arnold and
Stewart, 1999).
MEA is a stable compound and, in the absence of other chemicals, suffers no
degradation or decomposition at temperatures up to its normal boiling point. MEA
reacts with H2S and CO2 as follow:
2RNH2 + H2S
(RNH3)2S
2.1
(RNH3)2S + H2S
2(RNH3)HS
2.2
2RNH + CO2
RNHCOONH3R
2.3
These reactions are reversible by changing the system temperature. MEA also
reacts with carbonyl sulfide COS and carbon disulfide (CS2) to form heat-stable salts
that cannot be regenerated. DEA is a weaker base than MEA and therefore the DEA
system does not typically suffer the same corrosion problems but does react with
hydrogen sulfide and carbon dioxide:
11
2R2NH + H2S
(R2NH2)2S
2.4
(R2NH2)2S + H2S
2(R2NH2)HS
2.5
2R2NH + CO2
R2NCOONH2R2
2.6
DEA also removes carbonyl sulfide and carbon disulfide partially as its regenerable
compound without much solution losses (Mokhatab et al., 2006).
The general process flow diagram for an amine-sweetening plant varies little,
regardless of the aqueous amine solution used as the sweetening agent (figure 2.2).
2.3 Natural Gas Dehydration
Natural, associated, or tail gas usually contains water, in liquid and/or vapour form,
at source and/or as a result of sweetening with an aqueous solution. Operating
experience and thorough engineering have proved that it is necessary to reduce and
control the water content of gas to ensure safe processing and transmission.
The major reasons for removing the water from natural gas are as follow (Mokhatab
et al., 2006):

Natural gas in the right conditions can combine with liquid or free water to
form solid hydrates that can plug valves fittings or even pipelines

Water can condense in the pipeline, causing slug flow and possible erosion
and corrosion

Water vapour increases the volume and decreases the heating value of the
gas

Sales gas contracts and/or pipeline specifications often have to meet the
maximum water content of 7 lb/MMscf
There are several methods of dehydrating natural gas. The most common of these
are liquid desiccant (glycol) dehydration, solid desiccant dehydration, and
refrigeration (i.e., cooling the gas). The first two methods utilize mass transfer of the
water molecule into a liquid solvent (glycol solution) or a crystalline structure (dry
desiccant). The third method employs cooling to condense the water molecule to the
liquid phase with the subsequent injection of inhibitor to prevent hydrate formation.
However, the choice of dehydration method is usually between glycol and solid
desiccants (Mokhatab et al., 2006).
12
Figure 2.2: Schematic of amine gas-sweetening process flow diagram (Mokhatab et
al., 2006).
2.3.1 Glycol Dehydration
Among the different gas drying processes, absorption is the most common
technique, where the water vapour in the gas stream becomes absorbed in a liquid
solvent stream. Glycols are the most widely used absorption liquids as they
approximate the properties that meet commercial application criteria. Several glycols
have been found suitable for commercial application.
The commonly available glycols and their uses are described as follows (Katz et al.,
1959):

Monoethylene glycol (MEG); high vapour equilibrium with gas so tend to lose
to gas phase in contactor. Use as hydrate inhibitor where it can be recovered
from gas by separation at temperatures below 50oF

Diethylene glycol (DEG); high vapour pressure leads to high losses in
contactor. Low decomposition temperature requires low reconcentrator
13
temperature (315 to 340oF) and thus cannot get pure enough for most
applications

Triethylene glycol (TEG); most common. Reconcentrate at 340–400oF, for
high purity. At contactor temperatures in excess of 120 oF, there is a tendency
to high vapour losses. Dew point depressions up to 150 oF are possible with
stripping gas

Tetraethylene glycol (TREG); more expensive than TEG but less loss at high
gas contact temperatures. Reconcentrate at 400 to 430oF
TEG is by far the most common liquid desiccant used in natural gas dehydration. It
exhibits most of the desirable criteria of commercial suitability listed here (Manning
and Thompson, 1991; Hubbard, 1993):

TEG is regenerated more easily to a concentration of 98–99% in an
atmospheric stripper because of its high boiling point and decomposition
temperature

TEG has an initial theoretical decomposition temperature of 404 oF, whereas
that of diethylene glycol is only 328oF

Vaporization losses are lower than monoethylene glycol or diethylene glycol.
Therefore, the TEG can be regenerated easily to the high concentrations
needed to meet pipeline water dew point specifications

Capital and operating costs are lower
2.4 Natural Gas Liquids Recovery
Most natural gas is processed to remove the heavier hydrocarbon liquids from the
natural gas stream. These heavier hydrocarbon liquids, commonly referred to as
natural gas liquids (NGLs), include ethane, propane, butanes, and natural gasoline
(condensate). Recovery of NGL components in gas not only may be required for
hydrocarbon dew point control in a natural gas stream (to avoid the unsafe formation
of a liquid phase during transport), but also yields a source of revenue, as NGLs
normally have significantly greater value as separate marketable products than as
part of the natural gas stream. Lighter NGL fractions, such as ethane, propane, and
butanes, can be sold as fuel or feedstock to refineries and petrochemical plants,
while the heavier portion can be used as gasoline-blending stock.
14
WET GAS
Figure 2.3: Simplified flow diagram for TEG dehydration (Manning and Thompson,
1991).
The price difference between selling NGL as a liquid and as fuel, commonly referred
to as the “shrinkage value,” often dictates the recovery level desired by the gas
processors. Regardless of the economic incentive, however, gas usually must be
processed to meet the specification for safe delivery and combustion. Hence, NGL
recovery profitability is not the only factor in determining the degree of NGL
extraction. The removal of natural gas liquids usually takes place in a relatively
centralized processing plant, where the recovered NGLs are then treated to meet
commercial specifications before moving into the NGL transportation infrastructure.
Figure 2.4 shows the phase behaviour of a natural gas as a function of pressure and
temperature. Some plants operate at inlet pressures above the critical point and thus
re-vaporize NGLs when the temperature drops below the retrograde temperature. It
is therefore important to know where we are on the phase envelope.
15
Figure 2.4: Natural gas phase diagram.
The basic NGL recovery processes are related to figure 2.4, where possible. Such
processes include (Mokhatab et al., 2006):

Refrigeration Processes
 Mechanical Refrigeration
 Self-Refrigeration
 Cryogenic Refrigeration

Lean Oil Absorption

Solid Bed Adsorption

Membrane Separation Process
2.4.1 Cryogenic Refrigeration
Cryogenic refrigeration processes traditionally have been used for NGL recovery.
These plants have a higher capital cost but a lower operational cost. Moreover, they
contain numerous moving parts and are complicated to operate (Ewan et al., 1975).
16
As the natural gas enters a turboexpander or expansion valve, the gas expands, it
supplies work to the turbine shaft, thus reducing the gas enthalpy. The turbine can
be connected to a compressor, which recompresses the gas with only a small loss in
overall pressure. This results in a higher treated gas pressure, which can be
increased to the pipeline specification by a second compression.
Figure 2.5: Typical flow sheet of a cryogenic refrigeration plant (Ewan et al., 1975).
Although there are variations in the design of expander plants, most expander plants
have the same basic process flow as shown in figure 2.5. The inlet gas is first cooled
in the high-temperature, gas-to-gas heat exchanger and then in the propane chiller.
The partially condensed feed gas is sent to a separator. The liquid from the
separator is fed to the demethanizer, and the gas is cooled further in the lowtemperature gas to gas exchanger and fed into a second cold separator. Gas from
the cold separator expands through the expansion turbine to the demethanizer
pressure, which varies from 100 to 450 psia. The turbo expander simultaneously
produces cooling/condensing of the gas and useful work, which may be used to
recompress the sales gas (Mokhatab et al., 2006).
17
Cryogenic refrigeration process is generally the most technically advanced type of
NGL recovery used today. This combines high recovery levels (typically allowing full
recovery of all of the propane and heavier NGLs and recovery of 50% to more than
90% of the ethane) with low capital cost and easy operation (Lee et al., 1999).
2.5 NGL Fractionation
The bottom liquid from the NGL recovery plant may be sold as a mixed product. This
is common for small, isolated plants where there is insufficient local demand. The
mixed product is transported by truck, rail, barge, or pipeline to a central location for
further processing. Often it is more economical to fractionate the liquid into its
various components, which have a market value as purity products. However, as the
relative prices of natural gas and NGLs fluctuate, the relative incentive to extract the
NGLs from the gas changes. The process of separating a stream of NGLs into its
components is called fractionation. At the fractionation plant, liquids will be separated
into commercial quality products and then delivered to the market by tankers
(exports) and tank trucks (domestic consumption). NGLs are fractionated by heating
mixed NGL streams and passing them through a series of distillation towers
(Mokhatab et al., 2006).
Fractionators are usually named for the overhead or top product, as depicted in the
fractionation flow schematic of figure 2.6. Therefore, a deethanizer implies that the
top product is ethane; a depropanizer indicates that the top product is propane, etc.
Natural gas liquids are normally fractionated by boiling the lighter products from the
heavier products in the following order (Tuttle and Allen, 1976):

Deethanizer: The first step in the fractionating sequence is to separate the
ethane and propane, with the ethane going overhead and the propane and
heavier components passing from the bottom of the fractionator.

Depropanizer: The next step in the processing sequence is to separate the
propane and the isobutane, with the propane going overhead and the
isobutane and heavier components passing from the bottom of the
depropanizer.

Debutanizer: The next fractionating step is separation of butanes from
pentanes plus (C5+) stream. The butanes (both iso and normal) pass
overhead and the pentanes plus from the bottom of the fractionator.
18
Figure 2.6: Simplified flow diagram of a fractionation plant (Mokhatab et al., 2006).
2.6 Modular Process Skid
A modular process skid is a process system contained within a frame that allows the
process system to be easily transported. Individual skids can contain complete
process systems and multiple process skids can be combined to create larger
process systems or entire portable plants. They are sometimes called “a system in a
box.” An example of a multi-skid process system might include a raw materials skid,
a utilities skid and a processing unit which work in tandem.
Process skids are considered an alternative to traditional stick-built construction
where process system parts are shipped individually and installed incrementally at
the manufacturing site. They provide the advantage of parallel construction, where
process systems are built off-site in a fabrication facility while civil site upgrades are
completed at the plant site simultaneously. Skids are not always appropriate. If
individual process parts are large and cannot reasonably be contained within the
frame of a modular process skid, traditional construction methods are preferred.
19
2.6.1 Skid design and layout
Process skids are designed to contain a complete process system, a complete unit
of operations or to organize a manufacturing process into logical units. All skids
have the following characteristics in common:

Portable design

Small footprint

Gathered process connections

Controlled assembly

Factory acceptance testing (FAT) before installation

Accessible layout
2.6.2 Modular process skid components
Modular process skids typically contain the following equipment:

Controls

Electrical wiring

Flanges

Flow meters

Heat exchangers

Fractionators

Instrumentation

Insulation

Piping

Pumps

Tanks

Tubing

Valves
2.6.3 Skid applications

Batch processing

Bio waste deactivation systems

Centrifuge systems

Chemical processing

Chemical reactors
20

Clean-in-place systems

Coating systems

Continuous production systems

Demonstration plants

Distillation

Flavour mixing

Food and beverage processing

Fuel delivery systems

In-line blending systems

Mixing systems

Perfume mixing

Petroleum processing

Pilot plants

Processing plants

Pump carts

Raw materials processing

Refining

Wastewater treatment systems
(Reference: https://en.m.wikipedia.org/wiki/Modular_process_skid.)
21
CHAPTER THREE
RESEARCH METHODOLOGY
3.1 Overview of the Methodology
The methodology consists of both the design approach and the economic approach.
Input data (i.e. flowrate, composition, T, P)
Process
Synthesis
Step
Simulation and
Optimization
Step

Synthesize a typical NG processing
flowsheet from literature

Run Aspen HYSYS steady state
simulation
Perform design specifications &
sensitivity analysis
Optimize fractionation precooling, reflux
ratios, utilities, and column sizes


Economic
Evaluation
Step




Economic
Analysis
Step


Specify raw material and products costs
Specify utilities cost (optional)
Use ICARUS cost evaluator for
estimation of capital and operating costs
and product sales
Export evaluator results to Microsoft
Excel
Manipulate economic factors
Perform cash flow analysis
Check for project feasibility
Stop and Make Decision
Figure 3.1: Overall methodology (design and economic approach). Source: Self developed.
22
The design approach is aimed at constructing a natural gas processing plant for a
173,437 metric tonnes per annum raw natural gas which has the ability to remove
the acid gases, dehydrate the gas, recover the natural gas liquids and fractionate the
NGLs into ethane, propane, butane and natural gasoline.
The economic approach aims at analyzing conditions under which this project can be
economically feasible by making reasonable best case and worst case assumptions
on raw material and products pricing, interest rate, taxation, e.t.c.
The following is a description of the proposed methodology in sequential steps:

First step is to construct typical flowsheets which are based on widely used
natural gas processing plant technology by using the literature.

Second step is the simulation of the synthesized flowsheet using Aspen
HYSYS simulation package. The appropriate thermodynamic property method
should be selected before entering the simulation environment.

The third step is to evaluate the costs of constructing and operating the
natural gas processing plant by mapping, sizing and evaluating the costs.

The fourth step is to export the evaluation results to Microsoft Excel
environment where further economic analysis is carried out by manipulating
economic factors.
A schematic representation of these steps is shown in Figure 3.1.
3.2 Design Approach
To design a natural gas processing plant, three phases are very important. These
include:

Acid gas removal unit

Dehydration unit

NGLs recovery unit
The flowsheet for each of these units have been synthesized from the literature.
After understanding the processes and the operating ranges, Aspen HYSYS
simulator was then used to design and simulate the natural gas processing plant.
The initial procedure for the design is as follows:
23

Open the Aspen HYSYS software package and select a new case. This
automatically takes you to the properties environment.

Create two different component lists. The two component lists will both have
the following components: hydrogen sulfide, carbon dioxide, nitrogen,
methane, ethane, propane, iso-butane, n-butane, iso-pentane, n-pentane, nhexane, a hypothetical component for heptane plus fractions and water.

In addition to these components, triethylene glycol (TEG) and diethanol amine
(DEA) will be added to the first and second component lists respectively.

For the first component list, select the peng-robinson fluid property package
for its calculations while the acid gas fluid property package is selected for the
second component list.

Proceed to the simulation environment to design the natural gas processing
plant.
Figure 3.2: Main flowsheet of the simulated natural gas processing plant. Source:
Self developed.
Shown in figure 3.2 below is the feed process description for the simulated natural
gas processing plant. Description of each stage of the entire process will be stated in
detail below.
3.2.1 Feed Conditioning
The initial feed (raw sour gas) is obtained from an associated natural gas feed from
the Obigbo oil field in rivers state, Nigeria. The feed contains less than 0.45 mol% of
hydrogensulfide which classifies it as a sweet gas, it is still important to remove this
sulphur compound before further processing.
24
3.2.2 Acid Gas Removal Unit
In this process, raw gas is sweetened by reaction with DEA solution. The raw gas
enters a separator vessel at 40.81oC and 6895kPa. The gaseous overhead product
of the separator is then sent to the absorber where it enters the bottom stage of the
absorber column and reacts with lean DEA solution which enters the top stage of the
column at 38.60oC and 6895kPa. The absorber has a total of 20 stages. Wet gas
leaves the top of the absorber at 38.34oC and 5000kPa and fed to the dehydration
unit. The rich DEA solution leaves the bottom of the absorber at 45.53 oC and
6895kPa to be regenerated at the DEA stripper. The regenerated DEA solution is
recycled back to the absorber. Figure 3.3 below shows the flowsheet of the acid
removal unit.
Figure 3.3: Simulated flowsheet of the acid removal unit. Source: Self developed.
3.2.3 Dehydration Unit
As a result of removing the acid gases aqueous DEA, water was carried over by the
sweetened gas and must be removed. The wet gas enters a separator vessel and
25
the overhead vapour enters the bottom of an absorber while dry TEG enters the
absorber from the top stage at 27.97oC and 5000kPa. Dry gas exits the top of the
absorber at 39.13oC and 4900kPa while the wet TEG leaves the bottom at 38.77 oC
and 4900kPa. The absorber column has 4 stages. The wet TEG is sent to a stripper
to be regenerated and recycled while the dry gas is sent to the NGLs recovery unit.
Figure 3.4 below shows the flowsheet of the dehydration unit.
Figure 3.4: Simulated flowsheet of the dehydration unit. Source: Self developed.
3.2.4 NGLs Recovery and Fractionation Unit
This unit consists of the NGLs recovery system and an NGL fractionation train which
all together comprises of four columns namely the demethanizer, deethanizer,
depropanizer and debutanizer.
3.2.4.1
NGLs Recovery Unit
The dry gas first passes through a heat exchanger and chiller which drop its
temperature from 39.13oC to -60oC. The precooled gas then flashes into a separator
vessel. The liquid portion from the separator then enters a J-T valve which drops its
pressure from 4900kPa to 2145kPa thereby causing a sharp drop in temperature to 82.91oC.
26
Figure 3.5: Simulated flowsheet of the NGLs recovery unit. Source: Self developed.
The separator vapour enters a turboexpander that also reduces its pressure from
4900kPa to 2145kPa thereby dropping its temperature to -92.62oC. The vapour and
liquid then recombine at the top stage of the demethanizer. The demethanizer is a
reboiled absorber column with 15 stages. The residual gas (methane) exits the top
stage at -102.8oC and 2000kPa while the NGLs exit the reboiler at 23.78 oC and
2050kPa. In order to meet pipeline requirement, the residual gas is heated by the dry
gas, undergoes compression to 7000kPa and cooled to 50Oc by air cooler. The
NGLs then pass through a J-T valve which forces the pressure down to 1300kPa
with temperature of 7.43oC before entering the fractionation train. Figure 3.5 below
shows the flowsheet of the NGLs recovery unit.
3.2.4.2
Deethanizer
The deethanizer is a partial condenser distillation column which has 20 stages.
Ethane vapour exits the condenser reflux drum at -42.24oC and 700kPa while the
deethanizer bottoms exit the reboiler at 41.44oC and 800kPa into a pump that boosts
the pressure to 1500kPa and 42.1oC before it enters the depropanizer. Figure 3.6
below shows the flowsheet of the deethanizer.
27
Figure 3.6: Simulated flowsheet of the deethanizer. Source: Self developed.
3.2.4.3
Depropanizer
The depropanizer is a total condenser distillation column which has 25 stages.
Propane liquid exits the condenser reflux drum at -25.48oC and 200kPa while the
depropanizer bottoms exit the reboiler at 22.46oC and 205kPa into a pump that
boosts the pressure to 1000kPa and 23oC before it enters the debutanizer. Figure
3.7 below shows the flowsheet of the depropanizer.
Figure 3.7: Simulated flowsheet of the depropanizer. Source: Self developed.
3.2.4.4
Debutanizer
The debutanizer is a total condenser distillation column which has 17 stages. Butane
liquid exits the condenser reflux drum at -6.48oC and 105kPa while natural gasoline
28
is drawn from the reboiler at 42.6oC and 105kPa. Figure 3.8 below shows the
flowsheet of the debutanizer.
Figure 3.8: Simulated flowsheet of the debutanizer. Source: Self developed.
3.3 Economic Approach
The economic approach has two sections which are economic evaluation and
economic analysis.
3.3.1 Economic Evaluation
This is carried out to generate estimates for the capital and operating costs of the
natural gas processing plant. The economics section of the simulation environment
on HYSYS has the capabilities to help achieve this task. The procedure for achieving
this is as follows:

Activating the Costing Engine

Mapping Unit Operations to Equipment

Sizing Equipment

Economic Evaluation
When the costing engine has been activated, input raw and products material stream
prices and time based factors such as the project’s economic life, the start of
project’s engineering and the length of plant’s start up.
Then, the equipment required for every unit operation is determined by mapping
each operation. All equipment are then sized accordingly. The material of
29
construction by default is carbon steel but can be altered if desired. The capital and
operating costs are now determined by economic evaluation.
In other to perform economic/investment analysis, the economic evaluation results
are then exported to the Microsoft excel environment.
3.3.2 Economic Analysis
At this point, it is necessary to determine under what conditions the natural gas
processing plant project will be feasible before embarking or investing in it. The
report of the economic evaluation has the following sections:

Run Summary

Executive Summary

Cash Flow

Project Summary

Equipment Summary

Utility Summary

Utility Resource Summary

Raw Material Summary

Products Summary
The detailed components of the capital costs and operating costs are presented
along with cash flow analysis.
To perform economic analysis, the cash flow tab as well as the raw material and
products summary tabs will be extracted onto another Microsoft Excel document for
economic analysis. On the economic analysis document, some of the economic
factors will be defined and fixed while some will be varied with discretion. Table 3.1
below is the list of the economic factors that must be provided.
30
Table 3.1: List of the economic factors. Source: Self developed.
Economic Factors
Units
Duration of EPC Phase (DTEPC)
Year
Duration of EPC Phase & Start-up (DT)
Year
Working Capital Percentage (WORKP)
per cent /year
Operating Charges (OPCHG)
per cent /year
Plant Overhead (PLANTOVH)
per cent /year
Total Project Cost (CAPT)
Cost
Total Raw Material Cost (RAWT)
cost /year
Total Product Sales (PRODT)
cost /year
Total Utilities Cost (UTILT)
cost/year
Annual Interest Rate (ROR)
per cent
Economic Life of Project (ECONLIFE)
Year
Percentage Salvage value (SAVAL)
per cent
Depreciation Method (DEPMETH)
Depreciation Duration (DEPDUR)
Year
Project Capital Escalation (ESCAP)
per cent /year
Products Escalation (ESPROD)
per cent /year
Raw Material Escalation (ESRAW)
per cent /year
Operating and Maintenance Labour Escalation (ESLAB)
per cent /year
Utilities Escalation (ESUT)
per cent /year
Tax Rate (TAXR)
per cent /year
31
G and A Expenses (GA)
per cent /year
Total Operating Labour Cost (OPL)
cost /year
Total Maintenance Cost (MT)
cost /year
Also, the following equations will be used to set up the cash flow analysis worksheet.
𝑆 = 𝑃𝑅𝑂𝐷𝑇 × (1 +
𝑊𝐶 = 𝐶 ×
𝐸𝑆𝑃𝑅𝑂𝐷 𝑁
)
100
𝑊𝑂𝑅𝐾𝑃
3.2
100
𝐶 = 𝐶𝐴𝑃𝑇 × (1 +
𝐸𝑆𝐶𝐴𝑃 𝑃
100
)
3.3
𝑇𝐶 = 𝑊𝐶 + 𝐶
3.4
𝑅𝑀𝐶 = 𝑅𝐴𝑊𝑇 × (1 +
𝑂𝐿𝐶 = 𝑂𝑃𝐿 × (1 +
𝑀𝐶 = 𝑀𝑇 × (1 +
𝐸𝑆𝑅𝐴𝑊 𝑁
100
)
𝐸𝑆𝐿𝐴𝐵 𝑁
100
)
𝐸𝑆𝐿𝐴𝐵 𝑁
100
𝑈𝐶 = 𝑈𝑇𝐼𝐿𝑇 × (1 +
𝑂𝐶 = 𝑂𝐿𝐶 ×
3.1
)
𝐸𝑆𝑈𝑇 𝑁
100
)
𝑂𝑃𝐶𝐻𝐺
𝑃𝐿𝐴𝑁𝑇𝑂𝑉𝐻
100
𝑆𝑂𝐶 = 𝑅𝑀𝐶 + 𝑂𝐿𝐶 + 𝑀𝐶 + 𝑈𝐶 + 𝑂𝐶 + 𝑃𝑂
𝐺𝐴
3.6
3.7
3.8
3.9
100
𝑃𝑂 = (𝑂𝐿𝐶 + 𝑀𝐶) ×
3.5
3.10
3.11
𝐺&𝐴 = 𝑆𝑂𝐶 × 100
3.12
𝑂𝑃 = 𝑆𝑂𝐶 + 𝐺&𝐴
3.13
𝑅 = 𝑆 − 𝑂𝑃 − 𝑇𝐶
3.14
32
𝐷𝐸𝑃 = (𝐶𝐴𝑃𝑇 − 𝐶𝐴𝑃𝑇 ×
𝑆𝐴𝐿𝑉𝐴𝐿
100
𝐼 = 𝑅 − 𝐷𝐸𝑃
𝑇𝐴𝑋 = 𝐼 ×
)⁄𝐷𝐸𝑃𝐷𝑈𝑅
3.15
3.16
𝑇𝐴𝑋𝑅
3.17
100
𝑁𝐼 = 𝐼 − 𝑇𝐴𝑋
3.18
𝐶𝐹 = 𝑁𝐼 + 𝐷𝐸𝑃
3.19
𝐹𝑉 = 𝑃𝑉 × (1 +
𝑁𝑃𝑉 = ∑𝑁
𝑖=0
∑𝑁
𝑖=0
∑𝑃𝑌
𝑖=0
𝐶𝐹𝑖
(1+
𝐼𝑅𝑅 𝑖
)
100
𝐶𝐹𝑖
(1+
𝑅𝑂𝑅 𝑖
)
100
∑𝑁
𝑖=0
𝑃𝐼 =
𝑅𝑂𝑅 𝑁
100
)
𝐶𝐹𝑖
(1+
𝑅𝑂𝑅 𝑖
)
100
3.20
3.21
=0
3.22
=0
3.23
𝐶𝐹𝑖
𝑇𝐶
+
𝑅𝑂𝑅 𝑃
𝑅𝑂𝑅 𝑖
(1+
) (1+ 100 )
100
𝑇𝐶
𝑅𝑂𝑅 𝑃
(1+
)
100
3.24
S=Product Sales
WC=Working Capital
C=Escalated Capital Cost
TC=Total Capital Cost
RMC=Raw Material Cost
OLC=Operating Labour Cost
MC=Maintenance Cost
UC=Utility Cost
OC=Operating Charges
33
PO=Plant Overhead Cost
G&A=General and Administrative Costs
SOC=Subtotal Operating Cost
OP=Total Operating Cost
R=Gross Income
DEP=Depreciation
I=Income before Tax
TAX=Taxes
NI=Net Income
CF=Cash Flows
FV=Future Value of Cash
PV=Present Value of Cash
NPV=Net Present Value of Cash Flows
IRR=Internal Rate of Return
PY=Pay-out Time
P=Year of Plant Start-up
N=Number of Years
PI=Profitability Index
Economic feasibility means profitability and the minimum requirement for a project’s
profitability is a positive difference between the revenue and the total operating cost.
To establish the feasible region, reasonable upper and lower limit of prices are set
for the raw material and the products. The raw material in this case is the raw sour
gas while the products are residual gas (methane) and NGLs composite. Therefore,
setting two of these three variables at a combination of either upper limit prices,
lower limit prices or upper and lower limit prices, the feasible price range of the third
34
variable will be determined by iterations. The solver tool of the Microsoft excel will be
used to determine the optimum price that reduces the objective function, total
products sales minus total operating costs, to zero.
After the feasible region has been determined, further manipulations of some of the
economic factors will be done to perform economic analysis. The factors which will
be varied include:

Raw Material Price

Products Prices

Interest Rate

Escalation Rates

Tax Rate

Depreciation Duration
35
CHAPTER FOUR
RESULTS AND DISCUSSION
4.1 Design Approach
4.1.1 Natural Gas Processing Plant
Table 4.1 shows some of the thermodynamic properties of all the key material
streams of the overall natural gas processing plant.
Table 4.1: Thermodynamic properties of the key streams.
Material Streams
Vapour
Raw Sour Gas
Wet Gas
Dry Gas
1.0000
0.9980
1.0000
Fraction
Temperature
oC
40.92
38.35
39.13
Pressure
kPa
6895
5000
4900
Molar Flow
kgmole/h
1000
988.0
984.4
Mass Flow
kg/h
1.979e+04
1.923e+04
1.909e+04
Liquid Volume
m3/h
58.28
57.56
57.35
kJ/h
-8.287e+07
-7.842e+07
-7.756e+07
Methane
Ethane
Propane
1.0000
0.9996
0.0000
Flow
Heat Flow
Vapour
Fraction
Temperature
oC
50.00
-41.82
-25.48
Pressure
kPa
6995
700.0
200.0
Molar Flow
kgmole/h
865.5
60.77
26.36
Mass Flow
kg/h
1.400e+04
1832
1162
Liquid Volume
m3/h
46.44
5.132
2.294
kJ/h
-6.473e+07
-5.417e+06
-3.314e+06
Flow
Heat Flow
36
Table 4.1: Thermodynamic properties of the key streams (cont.).
Vapour Fraction
Butane
Gasoline
0.0000
0.0000
Temperature
oC
-6.483
42.60
Pressure
kPa
105.0
105.0
Molar Flow
kgmole/h
21.52
10.19
Mass Flow
kg/h
1251
837.6
Liquid Volume
m3/h
2.192
1.287
kJ/h
-3.355e+06
-1.943e+06
Flow
Heat Flow
As observed, 1000kgmole/h of raw sour gas was fed into the entire system but the
value dropped along the acid gas removal and dehydration units. These are due to
reduction in the level of impurity in the system and minor losses that can be
recovered and used as fuel in the plant. Also only propane, butane and natural
gasoline are recovered in liquefied phase while every other process stream is in
vapour phase.
37
4.1.2 Acid Gas Removal Unit
Table 4.2 shows the compositions of the main inlet and outlet streams of the acid
gas removal unit. 800kgmole/h of lean DEA solution (30% by weight) was used to
strip off the acid gases from the natural gas. The acid gas content has been
sufficiently reduced. As anticipated water was carried over from the acid gas removal
unit, hence, the natural gas must be dehydrated before fractionation. In trying to
regenerate the DEA solution, minor losses there were minor losses of the DEA. The
regenerated DEA solution is made up water and DEA and recycled as lean DEA
solution.
Table 4.2: Streams composition of the acid gas removal unit.
Compositions
Raw Sour Gas
Wet Gas
Comp Mole Fraction (H2S)
0.0015
0.0000
Comp Mole Fraction (Nitrogen)
0.0041
0.0041
Comp Mole Fraction (CO2)
0.0116
0.0001
Comp Mole Fraction (Methane)
0.8582
0.8677
Comp Mole Fraction (Ethane)
0.0646
0.0653
Comp Mole Fraction (Propane)
0.0271
0.0274
Comp Mole Fraction (i-Butane)
0.0125
0.0126
Comp Mole Fraction (n-Butane)
0.0092
0.0093
Comp Mole Fraction (i-Pentane)
0.0042
0.0042
Comp Mole Fraction (n-Pentane)
0.0028
0.0028
Comp Mole Fraction (n-Hexane)
0.0016
0.0016
Comp Mole Fraction (H2O)
0.0000
0.0022
Comp Mole Fraction (C7+*)
0.0026
0.0025
38
4.1.3 Dehydration Unit
Table 4.3 shows the compositions of the main inlet and outlet streams of the
dehydration unit. 2.399kgmole/h of triethylene glycol was used to dehydrate the wet
natural gas. TEG losses occur both at the absorber and the stripper, therefore, it is
made up with fresh TEG and recycled back. The water content has been reduced
sufficiently and the natural gas is ready to be sent to the fractionation train.
Table 4.3: Streams composition of the dehydration unit.
Compositions
Wet Gas
Dry Gas
Comp Mole Fraction (H2S)
0.0000
0.0000
Comp Mole Fraction (Nitrogen)
0.0041
0.0042
Comp Mole Fraction (CO2)
0.0001
0.0001
Comp Mole Fraction (Methane)
0.8677
0.8706
Comp Mole Fraction (Ethane)
0.0653
0.0655
Comp Mole Fraction (Propane)
0.0274
0.0274
Comp Mole Fraction (i-Butane)
0.0126
0.0126
Comp Mole Fraction (n-Butane)
0.0093
0.0093
Comp Mole Fraction (i-Pentane)
0.0042
0.0042
Comp Mole Fraction (n-Pentane)
0.0028
0.0028
Comp Mole Fraction (n-Hexane)
0.0016
0.0015
Comp Mole Fraction (H2O)
0.0022
0.0000
Comp Mole Fraction (C7+*)
0.0025
0.0019
39
4.1.4 NGLs recovery unit.
Table 4.4 shows the compositions of the main inlet and outlet streams of the NGLs
recovery unit. 99 percent methane purity was achieved by this unit. This represents a
high degree of effectiveness of the system. Nitrogen, being lighter than methane, is
present in the residual gas (methane). The fraction of methane in the bottoms
product of this column is only 0.08 percent which shows that the column is also
efficient in terms of methane recovery.
Table 4.4: Streams composition of the NGLs recovery unit.
Compositions
Dry Gas
Methane
Demeth. Btms
Comp Mole Fraction (H2S)
0.0000
0.0000
0.0000
Comp Mole Fraction (Nitrogen)
0.0042
0.0047
0.0000
Comp Mole Fraction (CO2)
0.0001
0.0000
0.0004
Comp Mole Fraction (Methane)
0.8706
0.9900
0.0008
Comp Mole Fraction (Ethane)
0.0655
0.0049
0.5063
Comp Mole Fraction (Propane)
0.0274
0.0002
0.2253
Comp Mole Fraction (i-Butane)
0.0126
0.0000
0.1043
Comp Mole Fraction (n-Butane)
0.0093
0.0000
0.0766
Comp Mole Fraction (i-Pentane)
0.0042
0.0000
0.0345
Comp Mole Fraction (n-Pentane)
0.0028
0.0000
0.0229
Comp Mole Fraction (n-Hexane)
0.0015
0.0000
0.0126
Comp Mole Fraction (H2O)
0.0000
0.0000
0.0002
Comp Mole Fraction (C7+*)
0.0019
0.0000
0.0159
40
4.1.5 Deethanizer
Table 4.5 shows the compositions of the main inlet and outlet streams of the
deethanizer. 99 percent ethane purity was also achieved by this column. This also
represents a high degree of effectiveness of the system. Much of the residual carbon
dioxide will be present in the ethane product. Traces of water vapour still present in
the system also come out with the ethane product. The fraction of ethane in the
bottoms product of this column is only 0.02 percent which represents high column
efficiency with respect to ethane recovery.
Table 4.5: Streams composition of the deethanizer.
Compositions
Demeth.
Ethane
Deeth. Btms
Btms*
Comp Mole Fraction (H2S)
0.0000
0.0000
0.0000
Comp Mole Fraction (Nitrogen)
0.0000
0.0000
0.0000
Comp Mole Fraction (CO2)
0.0004
0.0008
0.0000
Comp Mole Fraction (Methane)
0.0008
0.0017
0.0000
Comp Mole Fraction (Ethane)
0.5063
0.9900
0.0002
Comp Mole Fraction (Propane)
0.2253
0.0071
0.4537
Comp Mole Fraction (i-Butane)
0.1043
0.0000
0.2134
Comp Mole Fraction (n-Butane)
0.0766
0.0000
0.1568
Comp Mole Fraction (i-Pentane)
0.0345
0.0000
0.0707
Comp Mole Fraction (n-Pentane)
0.0229
0.0000
0.0469
Comp Mole Fraction (n-Hexane)
0.0126
0.0000
0.0258
Comp Mole Fraction (H2O)
0.0002
0.0004
0.0000
Comp Mole Fraction (C7+*)
0.0159
0.0000
0.0326
41
4.1.6 Depropanizer
Table 4.6 shows the compositions of the main inlet and outlet streams of the
depropanizer. By level of product purity, this appears to be the most effective column
in the fractionation unit with 99.95 percent propane purity. The fraction of propane in
the bottoms product of this column is only 0.00 percent which represents 100
percent column efficiency with respect to propane recovery.
Table 4.6: Streams composition of the depropanizer
Compositions
Deeth.
Propane
Deprop. Btms
Btms*
Comp Mole Fraction (H2S)
0.0000
0.0000
0.0000
Comp Mole Fraction (Nitrogen)
0.0000
0.0000
0.0000
Comp Mole Fraction (CO2)
0.0000
0.0000
0.0000
Comp Mole Fraction (Methane)
0.0000
0.0000
0.0000
Comp Mole Fraction (Ethane)
0.0002
0.0004
0.0000
Comp Mole Fraction (Propane)
0.4537
0.9995
0.0000
Comp Mole Fraction (i-Butane)
0.2134
0.0001
0.3907
Comp Mole Fraction (n-Butane)
0.1568
0.0000
0.2872
Comp Mole Fraction (i-Pentane)
0.0707
0.0000
0.1295
Comp Mole Fraction (n-Pentane)
0.0469
0.0000
0.0858
Comp Mole Fraction (n-Hexane)
0.0258
0.0000
0.0472
Comp Mole Fraction (H2O)
0.0000
0.0000
0.0000
Comp Mole Fraction (C7+*)
0.0326
0.0000
0.0596
42
4.1.7 Debutanizer
Table 4.7 shows the compositions of the main inlet and outlet streams of the
debutanizer. Two products are obtained from this column. Butane product has 99.9
percent purity while the pentane plus fraction (gasoline) exit with almost 100 percent.
This also represents a high degree of effectiveness of the system. The fraction of
butane in the bottoms product of this column is only 0.0 percent which represents
100 percent column efficiency with respect to butane recovery.
Table 4.7: Streams composition of the debutanizer.
Compositions
Deprop.
Butane
Gasoline
Btms*
Comp Mole Fraction (H2S)
0.0000
0.0000
0.0000
Comp Mole Fraction (Nitrogen)
0.0000
0.0000
0.0000
Comp Mole Fraction (CO2)
0.0000
0.0000
0.0000
Comp Mole Fraction (Methane)
0.0000
0.0000
0.0000
Comp Mole Fraction (Ethane)
0.0000
0.0000
0.0000
Comp Mole Fraction (Propane)
0.0000
0.0000
0.0000
Comp Mole Fraction (i-Butane)
0.3907
0.5758
0.0000
Comp Mole Fraction (n-Butane)
0.2872
0.4232
0.0000
Comp Mole Fraction (i-Pentane)
0.1295
0.0010
0.4007
Comp Mole Fraction (n-Pentane)
0.0858
0.0000
0.2670
Comp Mole Fraction (n-Hexane)
0.0472
0.0000
0.1468
Comp Mole Fraction (H2O)
0.0000
0.0000
0.0000
Comp Mole Fraction (C7+*)
0.0596
0.0000
0.1855
43
4.2 Economic Approach
4.2.1 Economic Factors
Table 4.8 shows the base case input values required for economic analysis as well
as estimates from Aspen HYSYS economic evaluation. The prices of the raw sour
gas, methane and natural gas liquids are set at $2.5/MCF, $3/MMBTU and
$10/MMBTU respectively.
Table 4.8: Values of the economic factors.
Economic Factors
Units
Duration of EPC Phase (DTEPC)
year
Duration of EPC Phase & Start-up (DT)
year
Working Capital Percentage (WORKP)
per cent /year
Operating Charges (OPCHG)
per cent /year
Plant Overhead (PLANTOVH)
per cent /year
Total Project Cost (CAPT)
US Dollars
Value
0.653846
0.845627
5
25
50
1.45E+07
1.90E+07
Total Raw Material Cost (RAWT)
US Dollars /year
Total Product Sales (PRODT)
US Dollars /year
Total Utilities Cost (UTILT)
US Dollars /year
3.91E+06
per cent
20
Annual Interest Rate (ROR)
Economic Life of Project (ECONLIFE)
Percentage Salvage value (SAVAL)
year
per cent
20
20
Straight Line
Depreciation Method (DEPMETH)
Depreciation Duration (DEPDUR)
3.69E+07
year
10
Project Capital Escalation (ESCAP)
per cent /year
3
Products Escalation (ESPROD)
per cent /year
3
Raw Material Escalation (ESRAW)
per cent /year
2.5
Operating and Maintenance Labour
per cent /year
3
44
Escalation (ESLAB)
Utilities Escalation (ESUT)
per cent /year
2.5
Tax Rate (TAXR)
per cent /year
40
G and A Expenses (GA)
per cent /year
8
Total Operating Labour Cost (OPL)
US Dollars /year
Total Maintenance Cost (MT)
US Dollars /year
1.18E+06
166554
4.2.2 Determination of Region of Economic Feasibility.
Following the past trend in the prices of natural gas from wellhead, residual
gas(methane) and NGLs composite, a set of low and high prices, avoiding being
overly optimistic, were set as shown in the table 4.9 below. As initially stated, these
values are discretionary and therefore are not rigid.
Table 4.9: Price ranges for raw gas and products.
Price Units
Low
High
Raw Sour Gas
$US/MCF
1
2.9
Methane
$US/MMBTU
1.5
4
NGLs Composite
$US/MMBTU
3
12
Table 4.10 below now shows the results of different scenarios for price
combinations. The table was obtained with absolute difference between the total
product sales and total operating cost set as the objective function. The objective
function is to approach zero as a feasible price range is determined for combination
of two extreme values. The table reveals that only 2 extremes are not feasible within
the specified price limits.
45
Table 4.10: Price feasibility analysis table.
Scenario 1 (fixed prices for raw sour gas and methane)
Raw Sour Gas
$US/MCF
1
1
2.9
2.9
Methane
$US/MMBTU
1.5
4
1.5
4
NGLs Composite
$US/MMBTU
3.22
3
11.18
3.82
Scenario 2 (fixed prices for raw sour gas and NGLs Composite)
Raw Sour Gas
$US/MCF
1
1
2.9
2.9
Methane
$US/MMBTU
1.58
1.5
N/F
1.5
NGLs Composite
$US/MMBTU
3
12
3
12
Scenario 3 (fixed prices for NGLs Composite and methane)
Raw Sour Gas
$US/MCF
N/F
2.9
2.7
2.9
Methane
$US/MMBTU
1.5
1.5
4
4
NGLs Composite
$US/MMBTU
3
12
3
12
* N/F- Not feasible
4.2.3 Economic Analysis
By adjusting one of raw gas price, methane price, NGLs composite price, interest
rate, escalation rates, depreciation duration and tax rate with every other economic
factor kept at the defined base values, their effects on the net present value (NPV),
internal rate of return (IRR), pay-out and profitability index (PI) of a 20 year cash flow
were documented. The different scenarios are tabulated below.
4.2.3.1
Scenario 1-Analysis Based On Raw Gas Price
It is evident from table 4.11 that profitability increases with decreasing cost of the raw
natural gas which ensures shorter pay-out period. Also, as long as there is
profitability, the IRR increases with decreasing cost. From this, the incentive to
attract investors into this business will be to make the wellhead price of natural gas
be very low to help guarantee profitability in the face of other uncertainties inherent
in other economic factors. This is reasonable because the gas which would have
been flared will eventually be monetized.
46
Table 4.11: Economic analysis based on raw material price
$0/MCF
$1/MCF
$2.5/MCF
NPV ($)
8.08E+07
5.5E+07
1.62E+07
IRR (%)
164.73
99.34
38.33
Pay-out (years)
1.74
2.29
5.35
PI
7.194
5.212
2.238
4.2.3.2
Scenario 2-Analysis Based On Methane Price
It is also evident from table 4.12 that profitability increases with increasing price of
methane. As long as there is profitability, the IRR increases with increasing price.
The price of methane is controlled by market forces of demand and supply. In as
much as it is mouth watering to keep the price as high as possible, revenue is a
product of price and quantity. It is therefore important to think of patronage when
fixing the price of methane. Potential patronage will be from process plants that burn
natural gas a source of energy or use it as feed to their processes like in fertilizers
production.
Table 4.12: Economic analysis based on methane price
$1.5/MMBTU
$3/MMBTU
$4/MMBTU
NPV ($)
-1.05E+07
1.62E+07
3.39E+07
IRR (%)
6.82%
38.33
59.73
0
5.35
3.37
1.946E-01
2.238
3.601
Pay-out (years)
PI
4.2.3.3
Scenario 3-Analysis Based On NGLs Composite Price
Table 4.13 shows that profitability increases with increasing price of the NGLs. As
long as there is profitability, the IRR increases with increasing price.
Similar to
methane, the prices of the NGL components are market driven. NGLs sold as
composite to petrochemical plants is less valuable than each of its components.
Ethane being the most valuable component is primarily used in petrochemicals
manufacture, but propane and butane are also used domestically for cooking.
47
Table 4.13: Economic analysis based on NGLs composite price
$7/MMBTU
$10/MMBTU
$12/MMBTU
NPV ($)
-1.97E+06
1.62E+07
2.82E+07
IRR (%)
17.72
38.33
52.65
0
5.35
3.78
8.492E-01
2.238
3.164
Pay-out (years)
PI
4.2.3.4
Scenario 4-Analysis Based On Interest Rate
Profitability increases and pay-out period shortens as the interest rate decreases.
The value of money is not constant with time, but to assume 20 percent interest rate
is reasonably high enough to provide assurances or reliability of the estimates from
the economic analysis despite having an internal rate of return at 38.33%. So, it can
be guaranteed that all the initial investments will be recovered in less than six years
while the business remains profitable under given conditions.
Table 4.14: Economic analysis based on interest rate
5%
10%
20%
NPV ($)
8.53E+07
4.78E+07
1.62E+07
IRR (%)
38.33
38.33
38.33
Pay-out (years)
4.06
4.40
5.35
PI
6.718
4.358
2.238
4.2.3.5
Scenario 5-Analysis Based On Escalation Rates
Escalations are expected to occur in the cost of goods and services over time. The
rates at which these escalations occur are quite unpredictable as well, but the
assumed base rates provide a good comparison with a case of zero escalation.
Assuming no escalation in the future costs and sales represents lower profitability
relative to the base values. So, one can be rest assured that the profitability of the
business with respect to escalations cannot be worse than if there were no
escalations.
48
Table 4.15: Economic analysis based on escalation rates
4.2.3.6
0%
Base values (%)
NPV ($)
9.10E+06
1.62E+07
IRR (%)
32.62
38.33
Pay-out (years)
6.13
5.35
PI
1.718
2.238
Scenario 6-Analysis Based On Depreciation Duration
Depreciation is a form of incentive to an investor to reduce the amount of tax that is
to be paid by recovering parts of the initial investment from net income before
taxation. The earlier the capital investment is recovered the better because of the
time value of money and tax being negative cash flow. This will result in higher NPV
and internal rate of return as well as lower payout as evident from table 4.16.
Table 4.16: Economic analysis based on depreciation duration
5 years
10 years
20 years
NPV ($)
1.66E+07
1.62E+07
1.55E+07
IRR (%)
39.47
38.33
37.35
Pay-out (years)
4.97
5.35
5.55
PI
2.272
2.238
2.191
4.2.3.7
Scenario 7-Analysis Based On Tax Rate
Taxation is one of the means by which a country’s government generate revenue.
Therefore, it is also to the advantage of the government that investors come into this
business. In view of this fact, it is necessary for the government to set the tax rate at
a low value to attract investors.Table 4.17 shows that the higher the taxation the
lower the profitability and longer the pay-out.
49
Table 4.17: Economic analysis based on tax rate
10%
30%
40%
NPV ($)
3.04E+07
2.09E+07
1.62E+07
IRR (%)
53.13
43.31
38.33
Pay-out (years)
3.77
4.65
5.35
PI
3.330
2.602
2.238
50
CHAPTER FIVE
CONCLUSIONS AND RECOMMENDATIONS
5.1 Conclusions
In an effort to design and carry out the economic feasibility of a modular natural gas
processing plant, the following tasks were carried out:

Literature on the processing of natural gas was first reviewed in order to
understand the sequence of operations and select processes to adopt for the
plant design.

Typical flowsheets of natural gas processing plant units based on widely used
technology were synthesized.

Aspen HYSYS simulation package was used to design a 20.8 million standard
cubic feet per annum (173 thousand metric tonnes per annum) natural gas
processing plant.

The costs of setting up and operating the natural gas processing plant were
estimated using Aspen HYSYS economic evaluator.

Economic analysis for a period of 20 years was carried out with the help of
Microsoft excel to determine the net present value, internal rate of return, the
payout period and profitability index of the project.

Also, the economic analysis was done to determine under which conditions
the natural gas processing plant will be feasible.
From the results of the design and economic analysis, the following observations
were made:

The design results were in agreement with the expected outcomes based on
the literature.

The estimated capital cost of the project is about 14.5 million US Dollars.

For a 20 year cash flow analysis at $2.5/MCF, $3/MMBTU, $10/MMBTU,
20%, 10 year and 40% raw natural gas, methane, NGL composite, discount
rate, depreciation period and tax rate respectively:
 The NPV is about 16.2 million US Dollars.
51
 The IRR is about 38.33%.
 The payout period is about 5.35 years.
 The profitability index is about 2.238.

But assuming zero escalations in the costs of materials and services under
the same conditions:
 The NPV is about 91 million US Dollars.
 The IRR is about 32.62%.
 The payout period is about 6.13 years.
 The profitability index is about 1.718.
5.2 Recommendations
From the outcomes of the study, the following recommendations were brought up:

It is recommended that the natural gas liquids fractionation train be included in
the plant’s configuration in order to add more value the NGLs.

More awareness should be created on the prospects in building modular
natural gas processing plants.

Entrepreneurs with the financial capabilities should invest in this project as the
return on investment is high and the payout period is quite early.

Government should create incentives such as allowing high depreciation rate
and imposing small tax rate.

Government should also invest in pipeline infrastructures that modular natural
gas processing plants can easily connect with.
52
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Speight, J. ((1993).). Gas Processing: Environmental Aspects and Methods. Oxford,
England: Butterworth Heinemann.
Thorn, R., Johansen, G., & Hammer, E. (April 14–17, 1999). Three-Phase Flow
Measurement in the Offshore Oil Industry: Is There a Place for Process
53
Tomography? 1st World Congress on Industrial Process Tomography. Buxton,
Greater Manchester.
54
APPENDIX A
Table A.1: List of Equipment and their Costs.
Total
Component
Direct
Equipment
Equipment
Installed
Name
Cost
Cost
Weight
Weight
(USD)
(USD)
KG
KG
75600
28300
3600
6377
251000
80400
6900
16829
56000
11200
750
2731
242900
92200
12100
18563
Knock
109100
30200
4800
7991
Separator
100400
17700
1600
5086
Pressure Valve
0
0
0
0
Recycle
0
0
0
0
DEA Pump
110600
64800
2000
4656
Gas Treating
Heat
74200
22900
2800
5660
Unit
Exchanger
Gas Treating
Cooler
62600
14900
1400
3663
Mixer
0
0
0
0
Area Name
Gas Treating
Stripper
Unit
Reboiler
Gas Treating
Stripper Main
Unit
Tower
Gas Treating
Stripper
Unit
Condenser
Gas Treating
Acid Gas
Unit
Absorber Main
Tower
Gas Treating
Unit
Gas Treating
Unit
Gas Treating
Unit
Gas Treating
Unit
Gas Treating
Unit
Unit
Gas Treating
Unit
55
Table A.1: List of Equipment and their Costs (cont.).
Gas
Absorber Main
163300
39100
3800
8133
Dehydration
Tower
0
0
0
0
Recycle
0
0
0
0
MIXER
0
0
0
0
Flash Tank
89600
16900
1400
4213
Surge
98400
27200
4300
7118
Cooler
45200
8900
310
1690
TEG Pump
75600
48600
1300
2265
Unit
Gas
Heater
Dehydration
Unit
Gas
Pressure Valve
Dehydration
Unit
Gas
Heat
Dehydration
Exchanger2
Unit
Gas
Dehydration
Unit
Gas
Dehydration
Unit
Gas
Dehydration
Unit
Gas
Dehydration
Unit
Gas
Dehydration
Unit
Gas
Dehydration
Unit
56
Table A.1: List of Equipment and their Costs (cont.).
Gas
Knockout Tank
106800
21100
1800
5161
161500
23800
1500
8209
102300
43200
6400
9514
313700
122200
12600
23863
55400
12900
1100
3173
43300
8900
310
1603
100300
31600
4200
6317
Dehydration
Unit
Gas
Heat
Dehydration
Exchanger
Unit
Gas
Stripper Main
Dehydration
Tower
Unit
NGLs
Debutanizer
Recovery/
Condenser
Fractionation
Unit
NGLs
Debutanizer
Recovery/
Main Tower
Fractionation
Unit
NGLs
Debutanizer
Recovery/
Reboiler
Fractionation
Unit
NGLs
Deethanizer
Recovery/
Reboiler
Fractionation
Unit
NGLs
Deethanizer
Recovery/
Condenser
Fractionation
Unit
57
Table A.1: List of Equipment and their Costs (cont.).
NGLs
Deethanizer
Recovery/
Main Tower
198700
60600
4600
10273
225700
83200
8700
15481
43500
8800
270
1591
45800
10100
520
1948
263600
102900
8800
16108
166200
80500
12300
15101
58200
22000
2500
3809
Fractionation
Unit
NGLs
Demethanizer
Recovery/
Main Tower
Fractionation
Unit
NGLs
Demethanizer
Recovery/
Reboiler
Fractionation
Unit
NGLs
Depropanizer
Recovery/
Reboiler
Fractionation
Unit
NGLs
Depropanizer
Recovery/
Main Tower
Fractionation
Unit
NGLs
Depropanizer
Recovery/
Condenser
Fractionation
Unit
NGLs
Heat
Recovery/
Exchanger
Fractionation
Unit
58
Table A.1: List of Equipment and their Costs (cont.).
NGLs
Pump1
43200
14800
500
1919
Separator
136900
30400
5000
9946
1.78E+06
1.63E+06
13500
23645
J-T Valve
0
0
0
0
Air Cooler
103200
54600
5300
8917
Chiller
107100
21000
2200
5138
Compressor1
969800
826100
5200
13977
Recovery/
Fractionation
Unit
NGLs
Recovery/
Fractionation
Unit
NGLs
Compressor2
Recovery/
Fractionation
Unit
NGLs
Recovery/
Fractionation
Unit
NGLs
Recovery/
Fractionation
Unit
NGLs
Recovery/
Fractionation
Unit
NGLs
Recovery/
Fractionation
Unit
59
Table A.1: List of Equipment and their Costs (cont.).
NGLs
Pump2
43100
14800
490
1864
0
0
0
0
226800
126500
4400
6520
Recovery/
Fractionation
Unit
NGLs
J-T Valve 2
Recovery/
Fractionation
Unit
NGLs
Turboexpander
Recovery/Fra
ctionation
Unit
60
APPENDIX B
Table B.1: List of Utility Requirements by Equipment and Costs.
Cost per
Description
Fluid
Item Description
Rate
Hour
(USD)
TEG Pump @ Gas
Electricity
Dehydration Unit
DEA Pump @ Gas
Electricity
Treating Unit
0.75
KW
0.058125
55.02
KW
4.26405
2.22
KW
0.17205
1000.4
KW
77.531
3
KW
0.2325
140.06
KW
10.85465
1.5
KW
0.11625
4.0128
M3/H
0.127206
46.0036
M3/H
1.458316
48.7301
M3/H
1.544745
Pump1 @ NGLs
Electricity
Recovery/ Fractionation
Unit
Compressor2 @ NGLs
Electricity
Recovery/ Fractionation
Unit
Air Cooler @ NGLs
Electricity
Recovery/ Fractionation
Unit
Compressor1 @ NGLs
Electricity
Recovery/ Fractionation
Unit
Pump2 @ NGLs
Electricity
Recovery/ Fractionation
Unit
Cooling
Water
Cooling
Water
Cooling
Water
Water
Water
Refrigerant
Cooler @ Gas
Dehydration Unit
Condenser @ T-100 @
Gas Treating Unit
Cooler @ Gas Treating
Unit
61
Table B.1: List of Utility Requirements by Equipment and Costs (cont.).
Condenser @
Refrigerant
- Propylene
Refrigerant
Deethanizer @ NGLs
Recovery/ Fractionation
1.62126
TON/H
0.220492
5.83218
TON/H
0.209959
44.4268
TON/H
7.552566
16.7899
TON/H
2.854287
0.57582
TON/H
10.31300
1.19382
TON/H
0.30253
TON/H
5.41842
12.4062
TON/H
222.1965
Unit
Refrigerant
- Ethane
Chiller @ NGLs
Refrigerant
Recovery/ Fractionation
Unit
Condenser @
Refrigerant
- Freon 12
Refrigerant
Debutanizer @ NGLs
Recovery/ Fractionation
Unit
Condenser @
Refrigerant
- Freon 12
Refrigerant
Depropanizer @ NGLs
Recovery/ Fractionation
Unit
Steam
@690KPA
Reboiler @ Deethanizer
Steam
@ NGLs Recovery/
Fractionation Unit
Reboiler @
Steam
@690KPA
Steam
Depropanizer @ NGLs
Recovery/ Fractionation
21.38142
4
Unit
Reboiler @
Steam
@690KPA
Steam
Demethanizer @ NGLs
Recovery/ Fractionation
Unit
Steam
@690KPA
Steam
Reboiler @ T-100 @
Gas Treating Unit
62
Table B.1: List of Utility Requirements by Equipment and Costs (cont.).
Steam
@690KPA
Steam
@690KPA
Reboiler @ Debutanizer
Steam
@ NGLs Recovery/
3.51992
TON/H
63.04176
0.00014
TON/H
0.002525
Fractionation Unit
Steam
Heater @ Gas
Dehydration Unit
63
APPENDIX C
Table C.1: Total Utility Requirements and their Prices.
Description
Fluid
Rate
Units
1419.96
KW
Cost Rate
Units
KW
0.0775
USD/KWH
Water
98.7466
M3
M3/H
0.0317
USD/M3
Refrigerant
1.62126
TON
TON/H
0.13600022
USD/TON
Refrigerant
5.83219
TON
TON/H
0.03600004
USD/TON
Refrigerant
61.2168
TON
TON/H
0.16999999
USD/TON
Steam
17.9985
TON
TON/H
17.9099999
USD/TON
Electricity
Cooling
Water
Refrigerant Propylene
Refrigerant Ethane
Refrigerant Freon 12
Steam
@690KPA
Rate
Units
64
APPENDIX D
Table D.1: Project Capital Summary.
Design,
Eng,
Constr.
Proc.
Matrl. Cost
Cost
(USD)
(USD)
Purchased
Constr.
Manhrs.
Cost
(USD)
Constr.
Manpower
Cost (USD)
5.04E+06
Equipment
Equipment
Total Cost
(USD)
5.04E+06
2413
25035.3873
2.50E+04
492703.029
11945.5
107952.39
6.01E+05
Civil
0
0
0
0.00E+00
Steel
142404.21
768
6508.23
1.49E+05
Instrumentation
1657737.9
13199
121188.9
1.78E+06
Electrical
1080349.92
6773
58755.6
1.14E+06
Insulation
121233.06
3872
26264.04
1.47E+05
Paint
37438.245
3098
20768.52
5.82E+04
917163
0
7.78E+06
0
0
0.00E+00
0
295950.33
13298.16
3.67E+05
269623.2
203219.64
37438.8
6.74E+05
0
0
0
0.00E+00
100014.12
186555.33
8892.15
3.34E+05
Setting
Piping
Other
4.93E+06
Subcontracts
G and A
Overheads
Contract Fee
Escalation
Contingencies
Total Project
Cost
Adjusted Total
Project Cost
0
1.81E+07
1.45E+07
65
APPENDIX E
Table E.1: Cash Flow Analysis Results.
CAP
Unescalated
(Capital
Cumulative
Costs)
Capital Cost
USD/year
0
Cumulative
Working
Capital Cost
Capital Cost
Capital
USD/year
USD/year
USD/year
USD/year
0
0
0
0
1
15658153.11
14478181.33
14912526.77
2
0
14478181.33
0
14912526.77
0
3
0
14478181.33
0
14912526.77
0
4
0
14478181.33
0
14912526.77
0
5
0
14478181.33
0
14912526.77
0
6
0
14478181.33
0
14912526.77
0
7
0
14478181.33
0
14912526.77
0
8
0
14478181.33
0
14912526.77
0
9
0
14478181.33
0
14912526.77
0
10
0
14478181.33
0
14912526.77
0
11
0
14478181.33
0
14912526.77
0
12
0
14478181.33
0
14912526.77
0
13
0
14478181.33
0
14912526.77
0
14
0
14478181.33
0
14912526.77
0
15
0
14478181.33
0
14912526.77
0
16
0
14478181.33
0
14912526.77
0
17
0
14478181.33
0
14912526.77
0
18
0
14478181.33
0
14912526.77
0
19
0
14478181.33
0
14912526.77
0
20
0
14478181.33
0
14912526.77
0
year
14912526.77 745626.3384
66
Table E.1: Cash Flow Analysis Results (cont.).
Operating Labor
Maintenance
Raw Materials
Cost
Cost
Utilities
USD/year
USD/year
USD/year
USD/year
0
0
0
0
0
1
6750182.36
421931.3683
59382.93332
1388313.444
2
19988031.1
1255479.669
176697.1386
4110948.538
3
20487731.88
1293144.059
181998.0528
4213722.251
4
20999925.18
1331938.381
187457.9943
4319065.307
5
21524923.31
1371896.532
193081.7342
4427041.94
6
22063046.39
1413053.428
198874.1862
4537717.988
7
22614622.55
1455445.031
204840.4118
4651160.938
8
23179988.12
1499108.382
210985.6241
4767439.962
9
23759487.82
1544081.633
217315.1929
4886625.961
10
24353475.01
1590404.082
223834.6486
5008791.61
11
24962311.89
1638116.205
230549.6881
5134011.4
12
25586369.69
1687259.691
237466.1787
5262361.685
13
26226028.93
1737877.482
244590.1641
5393920.727
14
26881679.65
1790013.806
251927.869
5528768.745
15
27553721.64
1843714.22
259485.7051
5666987.964
16
28242564.68
1899025.647
267270.2763
5808662.663
17
28948628.8
1955996.417
275288.3845
5953879.229
18
29672344.52
2014676.309
283547.0361
6102726.21
19
30414153.13
2075116.598
292053.4472
6255294.365
20
31174506.96
2137370.096
300815.0506
6411676.725
year
67
Table E.1: Cash Flow Analysis Results (cont.).
Operating
Plant
Subtotal Operating
G and A
Charges
Overhead
Costs
Costs
USD/year
USD/year
USD/year
USD/year
0
0
0
0
0
1
105482.8421
240657.1508
8965950.099
717276.0079
2
313869.9173
716088.4038
26561114.77
2124889.182
3
323286.0148
737571.0559
27237453.32
2178996.265
4
332984.5952
759698.1876
27931069.64
2234485.572
5
342974.1331
782489.1332
28642406.78
2291392.543
6
353263.3571
805963.8072
29371919.16
2349753.533
7
363861.2578
830142.7214
30120072.91
2409605.833
8
374777.0955
855047.0031
30887346.18
2470987.695
9
386020.4084
880698.4132
31674229.43
2533938.354
10
397601.0206
907119.3656
32481225.74
2598498.059
11
409529.0512
934332.9465
33308851.18
2664708.094
12
421814.9228
962362.9349
34157635.1
2732610.808
13
434469.3705
991233.823
35028120.49
2802249.64
14
447503.4516
1020970.838
35920864.36
2873669.149
15
460928.5551
1051599.963
36836438.05
2946915.044
16
474756.4118
1083147.962
37775427.64
3022034.212
17
488999.1041
1115642.401
38738434.34
3099074.747
18
503669.0773
1149111.673
39726074.83
3178085.986
19
518779.1496
1183585.023
40738981.72
3259118.537
20
534342.5241
1219092.573
41777803.93
3342224.315
year
68
Table E.1: Cash Flow Analysis Results (cont.).
OP (Operating
year
Costs)
SP (Products Sales)
R (Revenue)
USD/year
USD/year
USD/year
0
0
0
0
1
9683226.106
5869484.144
-19471895.07
2
28686003.95
39162085.78
10476081.83
3
29416449.58
40336948.35
10920498.77
4
30165555.22
41547056.8
11381501.59
5
30933799.32
42793468.51
11859669.18
6
31721672.69
44077272.56
12355599.87
7
32529678.74
45399590.74
12869911.99
8
33358333.88
46761578.46
13403244.58
9
34208167.78
48164425.81
13956258.03
10
35079723.8
49609358.59
14529634.79
11
35973559.27
51097639.35
15124080.07
12
36890245.91
52630568.53
15740322.62
13
37830370.13
54209485.58
16379115.45
14
38794533.51
55835770.15
17041236.64
15
39783353.09
57510843.25
17727490.16
16
40797461.86
59236168.55
18438706.7
17
41837509.08
61013253.61
19175744.53
18
42904160.81
62843651.22
19939490.4
19
43998100.25
64728960.75
20730860.5
20
45120028.25
66670829.58
21550801.33
69
Table E.1: Cash Flow Analysis Results (cont.).
DEP
(Depreciation
E (Earnings
Expense)
Before Taxes)
TAX (Taxes)
Earnings)
USD/year
USD/year
USD/year
USD/year
0
0
0
0
0
1
1158254.506
-20630149.58
0
-20630149.58
2
1158254.506
9317827.319
3727130.928
5590696.392
3
1158254.506
9762244.264
3904897.706
5857346.559
4
1158254.506
10223247.08
4089298.832
6133948.248
5
1158254.506
10701414.68
4280565.87
6420848.806
6
1158254.506
11197345.36
4478938.145
6718407.218
7
1158254.506
11711657.49
4684662.995
7026994.492
8
1158254.506
12244990.08
4897996.031
7346994.047
9
1158254.506
12798003.53
5119201.411
7678802.116
10
1158254.506
13371380.28
5348552.113
8022828.17
11
0
15124080.07
6049632.029
9074448.043
12
0
15740322.62
6296129.048
9444193.572
13
0
16379115.45
6551646.179
9827469.269
14
0
17041236.64
6816494.656
10224741.98
15
0
17727490.16
7090996.064
10636494.1
16
0
18438706.7
7375482.679
11063224.02
17
0
19175744.53
7670297.81
11505446.72
18
0
19939490.4
7975796.162
11963694.24
19
0
20730860.5
8292344.2
12438516.3
20
0
21550801.33
8620320.532
12930480.8
year
NE (Net
70
Table E.1: Cash Flow Analysis Results (cont.).
TEX (Total Expenses
TED (Total
(Excludes Taxes and
CF (CashFlow for
Earnings)
Depreciation))
Project)
USD/year
USD/year
USD/year
0
0
0
0
1
-19471895.07
25341379.21
-19471895.07
2
6748950.898
28686003.95
6748950.898
3
7015601.065
29416449.58
7015601.065
4
7292202.754
30165555.22
7292202.754
5
7579103.312
30933799.32
7579103.312
6
7876661.725
31721672.69
7876661.725
7
8185248.999
32529678.74
8185248.999
8
8505248.553
33358333.88
8505248.553
9
8837056.623
34208167.78
8837056.623
10
9181082.676
35079723.8
9181082.676
11
9074448.043
35973559.27
9074448.043
12
9444193.572
36890245.91
9444193.572
13
9827469.269
37830370.13
9827469.269
14
10224741.98
38794533.51
10224741.98
15
10636494.1
39783353.09
10636494.1
16
11063224.02
40797461.86
11063224.02
17
11505446.72
41837509.08
11505446.72
18
11963694.24
42904160.81
11963694.24
19
12438516.3
43998100.25
12438516.3
20
12930480.8
45120028.25
12930480.8
year
71
Table E.1: Cash Flow Analysis Results (cont.).
PVCI
FVCI (Future
(Present
PVI (Present
PVO (Present
Value of
Value of
Value of Cash
Value of Cash
Cumulative Cash
Cumulative
Inflows)
Outflows)
Inflows)
Cash Inflows)
USD/year
USD/year
USD/year
USD/year
0
0
0
0
0
1
4891236.786
21117816.01
5869484.144
4891236.786
2
27195892.9
22509121.44
46205466.75
32087129.69
3
23343141.41
19283187.09
95783508.45
55430271.1
4
20036196.37
16519509.09
156487266.9
75466467.47
5
17197735.22
14151863.58
230578188.8
92664202.69
6
14761389.4
12123511.32
320771099.2
107425592.1
7
12670192.57
10385839.79
430324909.7
120095784.7
8
10875248.62
8897199.64
563151470.2
130971033.3
9
9334588.399
7621907.622
723946190
140305621.7
10
8012188.376
6529392.267
918344786.6
148317810.1
11
6877128.356
5655816.672
1153111383
155194938.4
12
5902868.505
4843639.347
1436364228
161097806.9
13
5066628.8
4148115.402
1777846560
166164435.7
14
4348856.387
3552486.348
2189251642
170513292.1
15
3732768.399
3042401.734
2684612813
174246060.5
16
3203959.542
2605573.066
3280771545
177450020
17
2750065.274
2231477.465
3997939107
180200085.3
18
2360472.693
1911103.934
4860370580
182560558
19
2026072.395
1636736.089
5897173657
184586630.4
20
1739045.473
1401766.122
7143279217
186325675.9
year
72
Table E.1: Cash Flow Analysis Results (cont.).
PVCOP
(Present Value
of Cumulative
PVCO (Present
Cash
Value of
PV (Present
Outflows,
Cumulative Cash
Value of Cash
NPV (Net
Products)
Outflows)
Flows)
Present Value)
USD/year
USD/year
USD/year
USD/year
0
0
0
0
0
1
21117816.01
21117816.01
-16226579.23
-16226579.23
2
43626937.46
43626937.46
4686771.457
-11539807.77
3
62910124.54
62910124.54
4059954.32
-7479853.448
4
79429633.63
79429633.63
3516687.285
-3963166.163
5
93581497.22
93581497.22
3045871.637
-917294.5263
6
105705008.5
105705008.5
2637878.075
1720583.548
7
116090848.3
116090848.3
2284352.774
4004936.322
8
124988048
124988048
1978048.98
5982985.302
9
132609955.6
132609955.6
1712680.777
7695666.079
10
139139347.9
139139347.9
1482796.109
9178462.188
11
144795164.5
144795164.5
1221311.684
10399773.87
12
149638803.9
149638803.9
1059229.158
11459003.03
13
153786919.3
153786919.3
918513.3985
12377516.43
14
157339405.6
157339405.6
796370.0396
13173886.47
15
160381807.4
160381807.4
690366.6646
13864253.13
16
162987380.4
162987380.4
598386.4762
14462639.61
17
165218857.9
165218857.9
518587.8084
14981227.42
18
167129961.8
167129961.8
449368.7592
15430596.18
19
168766697.9
168766697.9
389336.3066
15819932.48
20
170168464
170168464
337279.3504
16157211.83
year
73
Table E.1: Cash Flow Analysis Results (cont.).
IRR (Internal Rate of
Return)
PO (Payout Period)
Percent
year
0.38325792
5.347739547
PI (Profitability Index)
2.238246559
74