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LARGE SCALE LNG
BACKGROUND PAPER
PR E P AR E D BY
T I T L E :
P R IN C IP A L E N G IN E E R ST R A T E G I C TE C H N O L O G
I E S
NA M E :
G .N.HU N T E R
RE V I S IO N H IS T O R Y
RE V IS IO N DE T A IL S : DA T E :
0 IS S U E D 25 T H O C T O B E R , 2006
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CONTENTS
1. INTRODUCTION 5
2. LNG SUPPLY CHAIN 5
2.1 Gas Pre-treatment 8
2.2 LNG Liquefaction 8
2.2.1 Gas Liquefaction Basics 8
2.2.2 Liquefaction Processes 10
2.2.3 Economy of Scale 10
2.2.4 Operating Performance/Availability 13
2.2.5 Process Selection 13
3. TECHNOLOGY VENDORS 14
3.1 APCI 15
3.1.1 Propane Pre-Cooled Mixed Refrigerant (PMR or C3-MR)
Process 15
3.1.2 AP-X Process 16
3.2 Phillips 17
3.3 Shell 21
3.3.1 SMR process. 21
3.3.2 DMR process. 21
3.3.3 PMR process 23
3.4 Linde 25
3.5 Axens 27
4. ASSOCIATED FACILITIES 28
4.1 LNG Storage Tanks 30
4.1.1 Total Storage Capacity 30
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4.1.2 Number of Tanks 30
4.1.3 Type of Containment 31
4.1.4 Pump Column for In-tank Pumps 35
4.1.5 Tank Pressure Control 36
4.1.6 Purging and Cooldown 36
4.1.7 Insulation 37
4.2 Jetty and Marine Facilities 37
4.2.1 Ship Size 37
4.2.2 Berth Occupancy 38
4.2.3 LNG Tanker Berth and Loading Dock 38
4.2.4 Safety in Port and Jetty Design 40
4.3 Shore-To-Ship Interface And Transfer Piping 41
4.3.1 LNG Loading Arms 41
4.3.2 Loading Line 43
4.4 Vapour Handling (Boil-Off Gas) 45
5. LIQUEFACTION EQUIPMENT SELECTION. 46
5.1 Main Cryogenic Heat Exchangers (MCHE) 47 5.1.1 Spiral Wound
(Coil Wound) Heat Exchangers 47
5.1.2 Plate Fin (Brazed Aluminium) Heat Exchangers 49
5.1.3 Core-in-kettle 51
5.1.4 Cold Boxes 51
5.2 Compressors and Drivers 54
5.2.1 Combustion Gas Turbines 54
5.2.2 Electric Motors 55
6. ATTACHMENTS 55
6.1 LNG Trains Operating or being executed 55
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6.2 LNG Train Current Maximum Capacity 55
6.3 LNG EPC Contractors Experience 55
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1. INTRODUCTION This Background Paper presents an overview of
Large Scale Baseload LNG export plants, provides an introduction to
the processes available for the liquefaction section of a Baseload
LNG plant, and discusses some of the technology selection issues
that affect LNG plant configuration.
Since the first LNG liquefaction train came into operation in
1964, some 110 trains have been brought on line (or are currently
in design or construction phases). Tables listing trains, sorted by
process licensor and EPC contractor, are included as Attachments
6.1 and 6.3.
Starting in the 1960s, with single train capacity less than
1Mtpa (million tonne per annum), train sizes have increased
eight-fold, with 7.8Mtpa trains under construction in Qatar. Refer
to Attachment 6.2 for a table of current maximum train capacities
for each technology.
For the purposes of this paper I have used the term Mini for
plants less than 300,000tpa, Mid-scale for plants between 0.3 and
2Mtpa and large for plants over 2.0Mtpa. In this document I have
concentrated on providing background regarding Large Scale LNG
plants. Refer to the separate October 2nd, 2006 Background Paper
which covered Mini & Mid-scale LNG.
No attempt has been made herein to cover floating LNG.
2. LNG Supply Chain The typical LNG supply chain is comprised of
facilities similar to those below, of which only those within the
Base Load Liquefaction Plant box will be discussed in this
memorandum.
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The production of LNG from natural gas involves two distinct
process steps, being:
Gas Pre-treatment - the removal of impurities and contaminants
such as acid gas, mercaptans, water and mercury; and
LNG Liquefaction - Pre-cooling and the removal of heavy
hydrocarbons to prevent hydrocarbon
freeze-up and plugging of the cryogenic equipment and to ensure
that the calorific value of the LNG meets specification; and
liquefaction and sub-cooling.
The typical process scheme for an LNG plant (using the Phillips
Process):
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The Darwin LNG plant (which uses the Phillips LNG Process):
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2.1 Gas Pre-treatment
In a typical LNG plant the feed gas will be delivered at high
pressure (e.g. up to 90 bar, 1300psi) from upstream gas fields. The
gas is metered and its pressure controlled to the design pressure
of the plant. The gas is first pre-treated to remove any impurities
that interfere with processing or are not desired in the final
products.
Pre-treatment upstream of a liquefaction unit traditionally
consists of an acid gas removal step, in which CO2 and sulphur
compounds (H2S, COS and mercaptans) are removed, a dehydration step
and a mercury removal step. Treating unit requirements are
determined by the liquefaction unit requirements (water, CO2),
specifications of the LNG product (H2S, COS, organic sulphur
compounds), material protection (mercury) and environmental
restrictions (SO2 and hydrocarbon emissions). In addition waste
streams have also to fulfil minimum specifications.
Where there are high levels of H2S and limitations on the SO2
emissions, the removed sulphur components are recovered as
elemental sulphur. Environmental limitations to hydrocarbon
emissions can require incineration of CO2 acid gas even in the
absence of sulphur compounds. The mercury removal step can be
positioned upstream of the acid gas removal or downstream of the
dehydration step.
Most of the operational base load LNG plants process feed gases
with only low concentrations of CO2, mercury and water. This type
of gas requires the minimum of treating, often comprising a CO2
removal unit, molecular sieves for drying and a carbon bed for
mercury removal. The relative capital investment for acid gas
removal in a LNG plant increases significantly with increasing CO2
content. At 2 mol% CO2 the acid gas unit represents 6% of the
processing equipment cost but at 14 mol% CO2 it represents 15% of
the processing equipment cost. New developments such as membrane
technologies are starting to be considered as an option for bulk
removal of CO2 but solvent absorption remains the most cost
effective treatment process for meeting LNG specifications.
The LNG product specification (e.g. heating value/Wobbe number
etc) for the end market for the LNG will also determine the
pre-treatment (and liquefaction) processing requirements. Most LNG
contracts specify a range of acceptable heating values for the LNG
sold into a particular market. In most cases, this requires that a
certain fraction of the heavier hydrocarbon components found in
natural gas be removed prior to liquefaction, so that the LNG does
not exceed the upper heating value limit. Some natural gases also
require removal of the heavy ends to prevent operating problems in
the liquefaction cycle, such as freezing of aromatic hydrocarbons
at low temperatures.
The remaining gas is made up mainly of methane and typically
contains less than 0.1 mol% of pentane and heavier
hydrocarbons.
2.2 LNG Liquefaction
Studies of the different liquefaction processes by independent
consultants suggest there is not one of them, on its own, that is
substantially more efficient than the others in all situations.
Rather, each technology can be competitive within a certain range
of train sizes. The ultimate choice of which process to select will
remain dependent on project-specific variables and the potential
development state of novel processes.
2.2.1 Gas Liquefaction Basics
The liquefaction section is the key LNG plant element.
Liquefaction processes mainly use mechanical refrigeration, in
which heat is transferred from the natural gas, through exchanger
surfaces, to a separate closed loop refrigerant fluid. The
refrigerant loop uses the cooling effect of fluid expansion,
requiring work input via a compressor.
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LNG plants consist of parallel units, called trains, which treat
and liquefy natural gas and send the LNG to storage tanks.
Liquefaction train capacity is primarily determined by the
liquefaction process, refrigerant used, and largest available size
of the compressor/ driver combination that drives the cycle and the
heat exchangers that cool the natural gas.
Basic principles for cooling and liquefying gas using
refrigerants involve matching, as closely as possible, the cooling/
heating curves of the process gas and the refrigerant. This results
in a more efficient thermodynamic process requiring less power per
unit of LNG produced, and it applies to all liquefaction processes.
Typical natural gas/refrigerant cooling curves are shown in this
figure:
Observing the cooling curve of a typical gas liquefaction
process, three zones can be noted in the process of the gas being
liquefied. These include a pre-cooling zone, followed by a
liquefaction zone, and completed by a sub-cooling zone. All of
these zones are characterized by having different curve slopes, or
specific heats, along the process. All of the LNG processes are
designed to closely approach the cooling curve of the gas being
liquefied. This is done by using specially mixed multi-component
refrigerants that will match the cooling curve at the different
zones/ stages of the liquefaction process to achieve high
refrigeration efficiency and reduce energy consumption.
The natural gas, being a mixture of compounds, liquefies over a
wide temperature range. Matching of heat curves by minimising the
temperature difference between the cooling process gas and
refrigerant streams can be achieved by using more than one
refrigerant to cover the temperature range and using the
refrigerant at different pressure levels to further split the
temperature ranges to closely matching ones. The process gas side
is normally operated at high pressure (e.g. 40 50 bara) to reduce
equipment size and provide more efficient refrigeration.
The liquefaction cooling curve performance is a benchmark that
is reviewed in LNG technology comparisons and is often
misunderstood or incorrectly applied when considering energy
performance relative to lifecycle cost. Caution should be used with
this type of comparison. Detailed knowledge of each liquefaction
process design, the options they can achieve at different
performance levels along this curve, and these options' cost
impact, is required for a valid comparison.
Because LNG liquefaction requires a significant amount of
refrigeration energy, the refrigeration system represents a large
portion of an LNG facility. A number of liquefaction processes have
been developed, with the differences mainly confined to the type of
refrigeration cycles employed. The most commonly utilized LNG
technologies are described below. These processes are used in
current plants or are applied in projects in progress. However,
there are other processes developed for baseload LNG applications,
which are being considered for future projects (and of these only
the Axens process is discussed here).
The liquefaction section typically accounts for 30% to 40% of
the capital cost of the overall liquefaction plant, which in turns
accounts for 25% to 35% of total LNG export plant costs. Key
equipment items include compressors used to circulate the
refrigerants, compressor drivers, and heat exchangers used to cool
and liquefy the gas and exchange heat between refrigerants. For
recent baseload LNG plants, this equipment is among the largest of
its type and at the leading edge of technology.
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The composition of the refrigerant gives an added control
parameter as it can be made either from pure or mixed components.
With a mixed refrigerant the composition can be adjusted to suit
the process conditions. The aluminium heat exchangers used [e.g.
the spiral wound heat exchangers (SWHE) or the plate fin heat
exchangers (PFHE)] have very large surface areas and a large number
of passes, to enable close temperature approaches.
2.2.2 Liquefaction Processes
Variations of three liquefaction processes have used for large
scale base-load plants. These are:
The classical cascade in which a three-stage pre-cooling cycle
is followed by a three-stage ethylene liquefaction cycle and a
three-stage methane sub-cooling cycle;
The single flow mixed refrigerant process in which a mixture of
nitrogen, methane, ethane, propane and normal butane is used as the
refrigerant.
The propane pre-cooled mixed refrigerant (PMR) process in which
pre-cooling is undertaken in a three-stage propane cycle compressor
and pre-cooling heat exchangers. Liquefaction and sub-cooling are
undertaken using a two-stage mixed refrigerant compressor,
separator, liquefier and sub-cooler.
As a general statement, LNG technology licensors have focused on
three aspects of LNG production; these are:
The compression required in the refrigeration cycles;
The power to drive the refrigeration cycles with the exception
of the Kenai plant, all of the earliest plants used steam turbine
driven compressors. Now, combustion gas turbines are the norm.
Certain licensors are looking to utilize the largest frame
turbines, whilst others are considering aero-derivative GTs;
and
The Main Cryogenic Heat Exchanger (MCHE) that is used to chill
the incoming gas. Until recently, APCI with its spiral-wound heat
exchanger (SWHE) dominated the design and manufacture of this
component. Linde also manufactures such a unit. Brazed aluminium
plate-fin heat exchangers (PFHE) are now challenging the dominance
of the SWHE. These high efficiency units used for MCHE service in
both the Phillips and Axens processes, and are also used for lesser
services in the APCI AP-X and the Statoil/Linde MFCP processes.
2.2.3 Economy of Scale
The trend towards larger train size seems inevitable as the
still relatively young LNG industry seeks to lower unit capital
costs through economies of scale. Moreover, there is a growing
consensus that the hyper trains will prove technically feasible by
introducing new line-ups and/or paralleling known equipment. But as
train size increases, an array of factors including overall capital
requirements, added complexity during construction, specific
venture conditions such as resource availability and market
considerations make it unlikely that hyper trains will be a "one
size fits all" solution for the industry.
To meet the demand for larger LNG trains, the providers of LNG
technologies (including APCI, Phillips, Linde, Shell and Axens)
have been engineering ever larger throughput trains, with nominal
capacities of around 5million tonne per annum (Mtpa) now in
operation and up to 7.8Mtpa being constructed (refer to Attachment
6.2).
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The technical limitations involved in scaling up train size vary
between the different liquefaction processes. The size of the
combustion gas turbines used to drive the compressors is a
limitation in most of the processes and in the case of the APCI PMR
process, the size of the spiral-wound exchangers is also a
limitation. Other challenges include the size of the large diameter
cryogenic piping headers and control valves. Economies of scale as
a result of bigger individual trains have significantly reduced
liquefaction plant capital costs and enhanced the competitiveness
of LNG in international energy markets. In the 1960s 1Mtpa was
regarded as a large train and train capacity increased slowly
throughout the 1970s and 1980s such that a decade ago, the largest
LNG production capacity per train was around 2.5Mtpa. By the late
1990s, this had risen to 3Mtpa and the most recent trains have been
constructed with capacities of 4 to 5Mtpa. Now the industry is
setting its sights on super sized trains that are capable of
producing 7.8Mtpa. The following chart compares the growth in train
size over time (cerise curve) with the associated increase in LNG
Carrier (LNGC) capacity (blue diamonds and green curve):
One limitation to train size has been the feasibility of
manufacturing larger heat exchangers and transporting them from the
manufacturing site to the field. For spiral wound heat exchangers
(SWHE) employed by APCI, winding the tubing and manufacturing and
transporting the heat exchanger has been strongly related to the
shell size, which is limited to 18-20 feet in diameter. This
manufacture and transport limit for SWHEs capped the APCI C3-MR and
Shell DMR processes to around 5Mtpa. The technology developers have
overcome this (in APCIs case, refer to their AP-X process, and in
Shells case refer to their PMR process) so that the same size of
spiral wound heat exchanger being manufactured today can be used to
produce much higher output. As a result, heat exchangers compatible
with a 7.8Mtpa train are within the industry's capability and the
ultimate true single LNG Main Cryogenic Heat Exchanger (MCHE) could
perhaps support 8-10Mtpa trains from a technical feasibility
standpoint. As train sizes have increased, operators have employed
ever-larger gas turbine drivers to power refrigerant compressors.
Earlier gas turbine driven LNG plants used smaller, dual shaft gas
turbines as compressor drivers. The 28 MW ISO rating of the GE
Frame 5 gas turbine limited maximum possible train capacity to
about 2.7Mtpa without resorting to multiple compressor-drivers in
parallel (such as used in the Phillips process). LNG train
production of 3.3Mtpa or more for the C3-MCR liquefaction process
was made possible with the use of GE Frame 6 gas turbine (ISO
rating of 38.5 MW) as driver for the propane refrigeration cycle
compressor and a GE
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Frame 7 gas turbine (ISO rating of 80 MW) for the driver for the
mixed refrigerant compressor. The use of the GE Frame 7 gas turbine
drivers for both C3 and mixed refrigerant cycles compressors
enabled individual train capacity up to 4 to 5Mtpa. Now APCI is
employing GE Frame 9 gas turbine drivers in their AP-X process.
Such increases come at a price, including more expensive drivers,
more difficult start up and greatly reduced flexibility of
operation. Further optimization and technical innovation is
required for gas turbine drivers and compressors to meet these high
planned production rates and to cope with the increasing
requirements for availability and operability.
Another potential bottleneck in the ultimate realizable train
size is the ability to extract impurities from the feed gas. If the
feed gas contains a concentration of CO2 and/or H2S of 20-50% or
greater, treating plant could easily limit train size. Removal of
these two most common impurities in natural gas is a costly and
complex step in LNG manufacture. Such treating often requires high
pressure, thick walled, large diameter vessels. While these vessels
are limited in size to around 18 to 20 feet, this bottleneck can be
removed if parallel design is employed. While technical feasibility
appears within the industry's grasp, LNG export ventures must
consider an array of other factors when considering whether a super
sized train fits their needs, including:
The overall capital cost of the project, rather than unit
costs;
Whether the plant is to be onshore, on a gravity-based structure
or floating;
Adequacy of gas resource base;
Complexity of scheduling, and additional manpower and other
resources during construction phase;
The robustness of process and design (requires 20-40 years of
operation);
Safety over the life of the project;
Compatibility of operation and maintenance with existing trains
for an expansion project, as well as similarity of equipment for
spare parts;
General ease of maintenance and accessibility of major
components, including supplemental refrigerant;
Ease of start-up (and restart);
Minimum emissions; and
Experience of the technology partner with the process.
Therefore, choosing a liquefaction process is not a matter of
simply comparing one technology against another. Rather, it is part
of a whole range of factors that must be considered in order to
achieve maximum value in project implementation. Furthermore, as
engineering and construction firms come up with innovative
combinations of equipment, the concept of an LNG "train" could
blur. Finding sales outlets for the large incremental LNG
production from a hyper train remains a significant commercial
challenge. There are only a few buyers willing to make large
purchase commitments from single trains. In general, customers are
seeking small tranches of LNG to match the requirements of newly
liberalizing and competitive markets. Moreover, many buyers are
insisting on long ramp-up periods in their purchase contracts.
Delays in production build-up could quickly erode any specific cost
benefits from the new hyper trains. At the same time, potential
economies of scale do little to help the producer maintain a steady
cash flow stream, particularly if it is a relatively new seller
with most of his "eggs" in a single large train "basket."
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2.2.4 Operating Performance/Availability
The ability of a facility to achieve the required annual
production and export volumes is dependent on the overall system
availability which is in turn a function of equipment reliability,
sparing philosophy, planned shutdowns and unscheduled
shutdowns.
Normal strategy for a baseload facility is that a single
component failure or malfunction must not cause a shutdown of the
plant through a cascade of events and may only affect the operation
of a specific facility or unit. Some projects adopt a policy
whereby in general, sparing of equipment is minimized and only
selected equipment has installed spares. In establishing the
economics for a project, it is important to establish the overall
system availability as soon as practicable so that technical
limitations and commercial expectations are properly aligned.
Optimized sparing is determined by reliability modelling and
cost-benefit analysis.
The design availability for LNG plants ranges from 340 days per
year to 350 days per year corresponding to 93 to 95+ percent
availability. Downtime for schedule maintenance of refrigeration
gas turbines is generally set at an average of 7 days per year. In
practice the compressors become the critical path for maintenance
and downtime. Best in class performers achieve around 10 days per
year for unplanned downtime.
2.2.5 Process Selection
Each process has its merits and, depending on plant capacity,
more than one process may be economical. The
choice of optimal process can vary based on site location, feed
gas price and ambient conditions, and
evaluation of a number of processes may be necessary to
determine the best economically over a
developments full life cycle.
Choosing the optimum process is crucial to reducing plant
capital cost as reduction in liquefier costs also reduces utilities
and offsites costs. The choice of liquefaction cycle depends on
many factors of which the major ones are:
Machinery configuration and available drivers;
Specific power requirement (affecting machinery capital cost and
operating cost);
NGL recovery or nitrogen rejection requirement;
Heat exchanger type and surface area;
Required flexibility; and
Ease of operation/start-up/shutdown. All of these issues should
be considered in process technology selection. Contacts will be
made with the LNG liquefaction licensors, LNG experienced EPC
contractors and main equipment vendors to obtain data and develop
designs to enable valid comparisons and optimum selections to be
made. Technology selection of process and equipment will be based
on technical and economic considerations:
Depending on the stage of project development, sufficient
process details must be developed to define main equipment and
operating parameters to evaluate options using relevant
criteria;
Technical considerations include process and equipment
experience, reliability, process efficiency, site
conditions and environmental impact among others;
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Economic issues include capital cost, operating cost and life
cycle costing. All of these aspects will need to be evaluated to
arrive at the optimum solution;
Technical risks associated with a process relate to the track
record of the process in operation and any
developments required for the project e.g. capacity
increase;
Process efficiency, for example energy required to produce LNG,
is not solely related to the thermodynamic efficiency of the
liquefaction process but also to the efficiency of the main
equipment, such as the main refrigerant compressors and
drivers;
Site conditions may favour one type of process over other. For
example, with very cold ambient
temperatures multi-mixed refrigerant processes may offer the
optimum solution;
Process requirements and configuration will have an influence on
selection. A requirement for greater LPG recovery may suit
processes with lower pre-cooling temperatures;
Wider feed gas range will require better process adaptability
and may favour mixed refrigerant
processes with the added flexibility of changing refrigerant
composition; and
Refrigerants made up from components that can be produced in the
process (in the fractionation unit) will obviate the need for
external supply to make up refrigerant losses.
3. Technology Vendors
Some 80% of total liquefaction capacity uses the Air Products
and Chemicals, Inc. (APCI) propane/mixed-refrigerant system. APCIs
strategy was so effective that they were the only successful
liquefaction process supplier for about 25 years. Their pre-cooled
propane mixed refrigerant (C3-MR) system became the standard in
which project investors could rest assured it would operate as
advertised.
Early patents for APCI's precooled propane mixed refrigerant
(C3/MR) process have expired. So, not only are new competitors
vying for market share, but also former customers (such as Shell)
are devising systems in-house.
APCI dominance has recently been successfully challenged by
Phillips, Shell, Linde-Statoil and IFP-Axens (although the latter
has not yet found a commercial application).
The following technologies are potential candidates for a Large
Scale LNG Project (listed in order of experience refer to
Attachment 6.1):
Air Products and Chemicals, Inc. (APCI) Propane Pre-Cooled Mixed
Refrigerant (PMR or C3-MR) Process and the AP-X Process
Technologies;
ConocoPhillips Optimized Cascade (OCP) Process Technology;
Shell Dual Mixed Refrigerant (DMR) Technology;
Statoil-Linde Mixed Fluid Cascade (MFC) Process Technology;
IFP-Axens Liquefin Technology.
Selection of an EPC Contractor will be run in parallel with LNG
liquefaction technology selection. Attachment 6.3 provides
background regarding contractors experience with designing and
constructing LNG trains since the first in 1964.
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3.1 APCI
APCI is based in Allentown, Pennsylvania.
APCI markets both their Propane Pre-Cooled Mixed Refrigerant
(PMR or C3-MR) and the AP-X processes:
3.1.1 Propane Pre-Cooled Mixed Refrigerant (PMR or C3-MR)
Process
The Propane Pre-cooled Mixed Refrigerant (C3-MR) process,
developed by APCI, began to dominate the industry from the late
1970s on. This process accounts for a very significant proportion
of the worlds baseload LNG production capacity and train capacities
of up to 5.0Mtpa have been built.
There are two main refrigerant cycles. The pre-cooling cycle
uses a pure component, propane. The liquefaction and sub-cooling
cycle uses a mixed refrigerant (MR) made up of nitrogen, methane,
ethane and propane. The pre-cooling cycle uses propane at three or
four pressure levels and can cool the process gas down to 40 C. It
is also used to cool and partially liquefy the MR. The pre-cooling
is achieved in kettle-type exchangers with propane refrigerant
boiling and evaporating in a pool on the shell side, and with the
process streams flowing in immersed tube passes. A centrifugal
compressor with side streams recovers the evaporated C3 streams and
compresses the vapour to 15 25 bara to be condensed against water
or air and recycled to the propane kettles. In the MR cycle the
partially liquefied refrigerant is separated into vapour and liquid
streams, which are used to liquefy and sub-cool the process stream
by cooling from typically -35C to between -150C to -160C. This is
carried out in a proprietary spiral wound heat exchanger (SWHE),
the main cryogenic heat exchanger (MCHE). The MCHE consists of two
or three tube bundles arranged in a vertical shell, with the
process gas and refrigerants entering the tubes at the bottom and
flowing upward under pressure.. The process gas passes through all
the bundles to emerge liquefied at the top. The liquid MR stream is
extracted after the warm or middle bundle and is flashed across a
Joule Thomson valve or hydraulic expander
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on to the shell side. It flows downwards and evaporates to
provide the bulk of cooling for the lower bundles. The vapour MR
stream passes to the top (cold bundle) and is liquefied and
sub-cooled, and is flashed across a JT valve or expander into the
shell side over the top of the cold bundle. It flows downwards to
provide the cooling duty for the top bundle and, after mixing with
liquid MR, part of the duty for the lower bundles. The overall
vaporised MR stream from the bottom of the MCHE is recovered and
compressed by the MR compressor to 45 48 bara. It is cooled and
partially liquefied first by water or air and then by the propane
refrigerant, and recycled to the MCHE. In earlier plants all stages
of the MR compression were normally been centrifugal, however, in
some recent plants axial compressors have been used for the LP
stage and centrifugal for the HP stage. The SEGAS Damietta LNG
plant in Egypt (at 5Mtpa, the largest APCI C3-MR plant):
3.1.2 AP-X Process
A limitation of the C3-MR process, which caps the capacity of
this process to around 5Mtpa for a single train, is the physical
size of the SWHE required, and the ability to manufacture and
transport the largest examples of these. While the SWHE could be
paralleled to increase single train capacity (refer to the Shell
PMR process below), APCI introduced a new technology called AP-X
Liquefaction Process Technology, which allows LNG production trains
to produce approximately 8Mtpa, over a 50 percent increase from
today's 5Mtpa standard. This new design marries APCIs standard LNG
technology with the company's air separation technology.
A modification of the APCI process, the APX process is a hybrid
C3-MR cycle adding a third refrigerant cycle (nitrogen expander) to
conduct LNG sub-cooling duties outside the MCHE. The new process
employs three refrigeration stages, propane (C3), mixed refrigerant
(MR) followed by a nitrogen (N2) expander cycle, and there are
several options of a suite of refrigerants. It is expected to be
similar to C3-MCR in that it is highly efficient. The addition of
the nitrogen expander cycle to sub-cool LNG can reduce the
refrigerant flow requirements of C3 and MR per unit of LNG
production. Thus, by using proven compression equipment without
duplicating or paralleling requirement, the individual train size
can be greatly increased.
With the new AP-X Hybrid LNG Process, train capacities up to
eight million metric tons per year are feasible in tropical
climates, in existing compressor frame sizes without
duplicate-parallel compression equipment, and
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using a spool-wound main cryogenic heat exchanger (MCHE). The
AP-X process is being utilised for the 6 trains, each of 7.8Mtpa
capacity, that are currently under development for RasGas and
Qatargas in Qatar.
The process cycle is an improvement over the propane-precooled,
mixed refrigerant (C3-MR) process in that the LNG is sub-cooled
using a simple, efficient nitrogen expander loop instead of mixed
refrigerant. In addition to improving the efficiency, the use of
the nitrogen expander loop makes greatly increased capacity
feasible. It does this by reducing the flow of both propane and
mixed refrigerant. The nitrogen-expander loop is a simplified
version of the cycle employed in hundreds of air separation plans
and nitrogen liquefiers worldwide.
Volumetric flow of mixed refrigerant at the low-pressure,
compressor suction is about 60 percent of that required by the
C3-MR process while mass flow of propane is about 80 percent of
that needed by the latter process. Propane is used to provide
cooling to a temperature of about -30C. The feed is then cooled and
liquefied by mixed refrigerant, exiting the MCHE at about -120C.
Final sub-cooling of the LNG is done using cold, gaseous nitrogen
from the nitrogen expander.
The AP-X train can be operated at a reduced production rate of
about 65 percent (5.2Mtpa) without the nitrogen expander loop by
adjusting the composition of the mixed-refrigerant inventory. The
producer can expand capacity later by adding the nitrogen-expander
cycles.
The power split between the propane, mixed-refrigerant and
nitrogen loops is flexible and can be manipulated by changing the
temperature range of the three refrigerant loops.
3.2 Phillips
ConocoPhillips is based in Houston, Texas.
Phillips Petroleum Company developed the original Optimized
Cascade Process (OCP) in the 1960s. The objective was to develop a
liquefaction technology that permitted easy start-up and smooth
operation for a wide range of feed gas conditions. This process was
first used in 1969 at their Kenai, Alaska LNG facility. That
facility was constructed by Bechtel and was the first plant to ship
LNG to Japan, and it has achieved more than 37 years of
uninterrupted supply to its Japanese customers.
The first new use of this process was for Train 1 of the
Atlantic LNG plant in Trinidad which came on line in 1999. The
process has emerged as a serious competitor to APCI's dominant
liquefaction technology and other
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LNG trains using this process have since been installed in
Trinidad (4), Darwin (1), Idku, Egypt (2) and Equatorial Guinea
(1). Train capacities of up to 5.2Mtpa are now operating. Bechtel,
which is the engineering and construction firm that has exclusive
rights to the Phillips process, is working with ConocoPhillips to
design even larger trains, with an 8Mtpa train having been
evaluated. Refrigeration and liquefaction of the process gas is
achieved in a cascade process using three pure component
refrigerants; propane, ethylene and methane, each at two or three
pressure levels. This is carried out in a series of brazed
aluminium plate fin heat exchangers (PFHE) arranged in cold boxes.
Pre-cooling can be carried out in a corein-kettle type exchanger.
The refrigerants are circulated using centrifugal compressors.
This process uses two pure refrigerants - propane and ethylene
circuits, and a methane flash circuit - cascaded to provide maximum
LNG production by utilizing the horsepower available from 6 gas
turbines. Each refrigerant circuit has parallel compression trains
using two 50% compressors with common process equipment. Frame 5
gas turbine drivers have been used for most plants, although Darwin
LNG uses GE LM25000+ aero-derivative gas turbines. Brazed aluminium
heat exchangers and core-in-kettle exchangers are used for the feed
gas, propane, ethylene and methane circuits. All of these heat
exchangers, with the exception of the propane chillers, are housed
in two "cold boxes." All compressor inter-cooling, after-cooling
and propane refrigerant condensing is provided by fin-fan heat
exchangers.
The Phillips' "two-in-one" design provides an added advantage as
train size increases, especially when production availability is
important. This design incorporates two drivers per refrigeration
service. As a consequence, if one driver goes down, the entire
train's production capacity is not lost. According to Bechtel,
production capacity using the Cascade process has an inherent
availability advantage of about 4 to 5% over a typical single
driver-compressor train arrangement. Despite negative perceptions
that the 2-in-1 compressor concept means more drivers and
compressors to maintain, those OCP plants operating achieve greater
than 95% availability. For example, Train 1 at Atlantic LNG has
operated with an availability of over 96% and the original Kenai
plant has achieved some 97% availability over its 37 year life. The
stated benefits of the 2-in-1 OCP concept include:
The ability to operate with one compressor down and still
produce at approx. 60% throughput. The ability to turndown
throughput even lower with only 3 of the 6 compressors operating (I
recall
from published papers that this can be as low as some 40%).
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High availability as maintenance can be undertaken on individual
compressors (shut down) while the
remainder of the plant continues to produce (albeit at reduced
capacity as noted above). In fact Atlantic LNG debottlenecked their
Train 1 by changing out all of their compressor drivers (from GE
Frame 5Cs to 5Ds) one by one while still producing from the plant -
no shutdowns involved! Because of this, availability of these OCP
plants is high. Bechtel believe that because APCI plants do not
have this flexibility, that the best that an APCI plant can achieve
is circa 93%.
They can take advantage of the smaller proven gas turbines (such
as GE Frame 5 and LM2500+ ) that
are available, rather than having to take the risk with larger
KW drivers (note that Darwin LNG is the first use of the more
efficient aeroderivative gas turbines in an LNG liquefaction
plant).
Unlike the APCI plants which have little turn down capability,
the OCPs 2 x 50% compressor design means that the following
operating ranges that can be utilised (in addition to running in
full recycle mode):
Operating Range % of design capacity
Full plant 80 to 105 One compressor offline 60 to 80 Three
compressors offline 30 to 60
Another feature of the 2-in-1 OCP process is that should a
project be initially constrained to gas feedstock volumes or
reserves that can only support an initial 1.5Mtpa plant, then there
is a way that a Phillips OCP process can manage this as a phased
1.5+1.5 plant. In discussion with Bechtel they recounted work that
they did on a study some 8 years ago for a client who wanted to
start at 1.5Mtpa and grow to 3.0Mpta. This project did not proceed,
however the idea is one worthy of consideration. Due to the
'standard' 2-in-1 design's 2 x 50% compression trains, the
compression can be installed in 2 stages with 3 compressors
initially (for the 1.5Mtpa capacity) and with the other 3
compressors (and drivers) being installed when the plant has gas
reserves and/or market to support a 3.0Mtpa plant. For this, the
gas pre-treatment and utilities would be designed and installed in
2 x trains (and if their were two distinct gas feedstocks, such as
lean CSM gas and richer/hotter pipeline sales gas, there is the
possibility of dedicating each of the pre-treatment trains to a
different gas feedstock). Of course this 1.5+1.5 facilities
arrangement and phased execution costs more than going ahead with a
Greenfield 3.0Mtpa plant in a single go. Even if this resulted in
the same total costs of two separate 1.5 trains, from
efficiency/operability/maintainability points of view the
integrated 1.5+1.5 would be preferable.
The OCP can provide a facility with high thermal efficiency. The
process utilizes proven technology and equipment, and has a wide
range of operational flexibility. Turndown rates to 10% are
achievable for long-term operation. Due to the pure component
systems, the plant has easy start-up and operation. The plant
boasts low utility and reduced flaring requirements, because
refrigerants are not flared on typical upset conditions. This leads
to reduced requirements for maintenance and operational
staffing.
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The four train Atlantic LNG plant, with the 5.2Mtpa train
(currently the worlds largest) in the foreground:
A schematic of the two train BG Idku plant in Egypt:
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3.3 Shell
Shell is based in the Hague.
The APCI C3/MR designs were traditionally the industry and
Shells standard. Shell has developed its own processes based on its
experience with operating APCI process LNG trains in Brunei,
Malaysia, Australia, Oman and Nigeria.
Shell has designed Single-Mixed-Refrigerant (SMR),
Double-Mixed-Refrigerant (DMR) and Parallel-Mixed-Refrigerant (PMR)
processes. With three different mixed-refrigerant cycles, Shell can
design LNG trains with capacities ranging from 0.5 to 7.0Mtpa. The
process options in the Shell portfolio are all based on two
refrigeration cycles in series. Depending on the required capacity
and local circumstances, different choices are made with respect to
the line-up of rotating equipment, type of refrigerant, ambient
cooling medium and types of cryogenic heat exchangers.
The Shell Double-Mixed-Refrigerant (DMR) process overcomes the
inherent limitations of using a single component refrigerant in
pre-cooling in the C3/MR design; the additional degree of freedom
resulting from the use of two MR cycles allows full utilisation of
power in a design with two mechanically driven compressors.
Furthermore, it allows keeping the compressors at their best
efficiency points over a very wide range (up to 50C) of ambient
temperature variations and changes in feed gas composition.
3.3.1 SMR process.
The traditional C3/MR design loses its advantage in economy of
scale for small capacity liquefaction trains. For this purpose, the
more cost-effective Shell single-mixed-refrigerant (SMR) process
has been developed. This type of design is most suitable for the
lower capacity ranges of 0.5 to 1.5Mtpa of LNG.
The process uses one refrigerant loop, which is used both for
pre-cooling the natural gas circuit as well as for pre-cooling the
refrigerant. A dedicated, spool-wound heat exchanger provides for
pre-cooling the mixed refrigerant. A bundle break in the cryogenic
heat exchanger allows overhead cooling of the scrub column for
removal of heavy hydrocarbons.
The unit operates at two mixed-refrigerant pressure levels.
Three-stage compression drives the mixed-refrigerant loop. The
driver could be electric drive or mechanical-drive gas
turbines.
3.3.2 DMR process.
The Shell DMR process has two separate mixed-refrigerant loops,
hence the name. The DMR process is designed for the mid and high
capacity ranges of 1.5 to 4.5Mtpa of LNG. Shell is also developing
a promising air-cooled DMR version with capacities from 2.0 to
4.0Mtpa for tropical conditions.
The pre-cool mixed refrigerant circuit is used just like the
propane pre-cool loop in the C3/MR process. This circuit pre-cools
both the natural gas circuit and the main mixed-refrigerant loop.
The main difference between this process and the C3/MR design is in
the use of two spool-wound heat exchangers, rather than a multiple
of kettle exchangers for extracting heat from the circuits. Also, a
less complicated two-stage centrifugal compressor is included in
this design.
The choice of an end flash system integrated with the
liquefaction design is dependent on the nitrogen content in the
feed gas and the requirement of increased design capacities. For
the large-scale Shell DMR process, the capacity of the system can
be boosted by some 5.0 percent to 10 percent by applying an
endflash system.
The DMR design was selected for the Sakhalin LNG Project
currently under construction, with an annual design capacity of
9.6Mtpa (two trains @4.8Mtpa each). It is also being used for the
Woodside NWS Train 5 and Pluto Projects.
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Process configuration is similar to the APCI propane pre-cooled
mixed refrigerant process, with the pre-cooling conducted by a
mixed refrigerant (made up mainly of ethane and propane) rather
than pure propane. Another main difference is that the pre-cooling
is carried out in SWHE rather than kettles. The pre-cooling and
liquefaction SWHEs for Sakhalin were supplied by Linde.
The DMR process has two separate mixed refrigerant cooling
cycles. One is for pre-cooling gas to about - 50C (PMR cycle), and
the other is for final cooling and liquefaction (MR cycle). This
concept allows the designer to choose the load on each cycle. It
also uses proven equipment, e.g. spiral-wound heat exchangers
(SWHEs), throughout the process.
PMR vapour from the pre-cool exchangers is routed via knock-out
vessels to a two-stage centrifugal PMR compressor. De-superheating,
condensation and sub-cooling of the PMR is achieved by using
induced-draft air coolers.
The PMR compressor is driven by a single gas turbine, equipped
with an electric starter motor/ generator. The refrigerant
compressors are driven by two Frame 7 gas turbines, equipped with a
separate variable speed starter/ helper motor. An axial compressor
is also used as part of the cold refrigerant compression
stages.
The cooling for liquefaction of the natural gas is provided by a
second mixed refrigerant cooling cycle (MR cycle). This cycle's
refrigerant consists of a mixture of nitrogen, methane, ethane and
propane. Mixed refrigerant vapour from the shell side of the main
cryogenic heat exchanger is compressed in an axial compressor,
followed by a two-stage centrifugal compressor. Inter-cooling and
initial de-superheating is achieved by air cooling. Further
de-superheating and partial condensation is achieved by the PMR
pre-cooling cycle. The mixed refrigerant vapour and liquid are
separated, and further cooled in the main cryogenic heat exchanger,
except for a small slipstream of vapour MR, which is routed to the
end flash exchanger.
Shell has also developed technology to further push the propane
cycle capacity, by employing double casing instead of single casing
equipment. This reliable method brings the propane-MR process
closer to a capacity of 5Mtpa. Another possibility for the
propane-MR process is to transfer power from the propane cycle to
the mixed refrigerant cycle. The closer coupling between the two
cycles by mechanical interlinking of compressors is an operational
challenge.
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A further development of the DMR process is the
electrically-driven DMR (LNG GameChanger) design. The Gamechanger
concept is based on parallel line-up of electrically driven
refrigerant compressors around a common set of cryogenic spool
wound exchangers. Electric motors of 65MW have already been
constructed for LNG service. Motors up to 80MW are considered
feasible. The current electrically driven DMR design is
particularly attractive in the 58Mtpa capacity range. Electrically
driven LNG trains can compete with mechanically driven trains
because the increase in cost is compensated for by the increased
availability. Other benefits of the electric option are the
variable size and speed of the driver, the increased vendor base
and the potential to make a step change reduction in overall plant
carbon dioxide (CO2) emissions, by using combined cycle electric
power generation.
3.3.3 PMR process
Shell's portfolio also includes designs for ultra large trains
based on the company's own liquefaction technology.
The company's Parallel-Mixed-Refrigerant (PMR) process employs a
common pre-cool cycle serving two parallel liquefaction cycles
using three large industrial-type gas turbines, which enable a
production of 6 to 7Mtpa. Much of the equipment has been used in
previous designs. The train can run at 65% capacity when one of the
liquefaction cycles is down.
Shell has developed the PMR process to meet the current
challenge of the industry for larger train sizes in tropical
conditions. With a single pre-cooling cycle and two PMR cycles, the
capacity can be boosted up to 8Mtpa with three GE Frame 7
compressors in a tropical climate. The process can either use C3 or
MR in pre-cooling. Proven refrigerant cycles can be used and the
design can currently be applied, without step changes in
technology. The capacity can be increased further with different
(larger) drivers.
Gas receipt and natural gas treating is followed by a single
propane pre-cooling cycle. After pre-cooling, the flow is
distributed over two parallel strings, each having a scrub column
for NGL extraction and an MR cycle for liquefaction and sub-cooling
of the natural gas. The scrub column overheads are cooled by the MR
to create reflux and ensure the required extraction level. Each
liquefaction cycle has its own MR circuit, driven by a gas turbine.
The sub-cooled liquid from the two liquefaction cycles is combined
in an end-flash system, where fuel gas is flashed off and LNG is
sent to storage at atmospheric pressure. The split-propane
technology is applied to limit the propane suction flows to
acceptable levels. In this arrangement, the four-stage propane
compressor is split over two casings the first machine compresses
the low-pressure (LP) and high pressure (HP) propane to discharge
pressure, whereas the second machine handles the medium pressure
(MP) and high-high-pressure (HHP) flows. This split arrangement
results in a lower volumetric flow per stage, for the same
compression duty. In order to achieve the targeted 8Mtpa production
capacity with the driver configuration chosen, an extended
end-flash system is required for the conditions prevalent in this
study. The application of such a system allows an increase in the
run-down temperature from the main cryogenic heat exchanger (MCHE)
in the liquefaction cycle. The additional flash gas that is
generated as a result of this, which cannot be accommodated in the
fuel gas system, is compressed, condensed and recycled back into
the end-flash gas system. In this way, power in the MR circuit is
freed up and replaced by power in the end-flash gas compressor.
This relieves the power constraint in the MR circuit and enables
higher LNG production capacity. The extended line-up as chosen here
results in some 4% to 5% additional LNG capacity. As an alternative
for the extended end-flash system, larger compressor drivers like
GE-Frame 9 or Siemens V84.2 can be used to achieve a production
capacity of 8Mtpa or higher. Two of the key parameters in this
process design are the scrub column overheads temperature and the
cut-point temperature between the C3 and MR cooling cycles.
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The scrub column overheads temperature sets the level of propane
recovery. In order to achieve a specification of an LNG product
quality of say 1,110 British thermal units (BTU) per standard cubic
foot (scf), a propane recovery level of approximately 40% to 45% is
required. At the pressures prevalent in the scrub column, this is
accomplished by using MR cooling to approximately -45C. The lowest
propane cooling level is the cut-point temperature between
pre-cooling and liquefaction. This temperature is optimised so that
the utilised power is balanced to a 1/2 ratio between the
pre-cooling and the liquefaction cycles. The ability to tune the
power balance exactly to the installed mechanical refrigeration
capacity in the PMR process is an advantage over the conventional
C3/MR process where, due to other constraints, the pre-cooling
cycle cannot be fully loaded.
The Shell PMR technology for large LNG trains has a number of
advantages. The processes are robust through the application of
well-proven equipment (e.g. spool wound heat exchangers, proven
rotating equipment) without requiring further scale-up. In fact,
the main equipment in the PMR process is already in operation in
plants such as Nigeria LNG, North West Shelf LNG and Malaysia LNG.
The parallel line-up of the liquefaction cycles improves the
reliability of the train since the LNG production can be designed
to continue at 60% of the train capacity when one of the
liquefaction cycles trips. Moreover, it allows high production
capacity with only two refrigeration cycles in series, compared
with three in most other large train concepts. Due to the PMR
line-up and the application of split-propane technology, the
installed power of three GE-Frame 7 drivers can be fully utilised.
Since these machines form a large part of the liquefaction unit
cost, this makes PMR designs very cost effective. The Shell PMR
process has a high efficiency through the use of two very efficient
refrigeration cycles. The parallel line-up reduces the pressure
drop in the system, which also helps to improve efficiency. Another
advantage is the absence of a third cycle with associated
efficiency losses due to temperature approaches in the cooling of
the third cycle refrigerant. Comparison by Shell of different
large-train processes for similar conditions and design premises
has shown that the Shell PMR process has an efficiency up to 10%
better than alternative processes.
In a tropical climate, where air-cooling and GE-Frame 7 drivers
are used, an LNG production capacity of 8Mtpa can be achieved with
the Shell PMR process. This production capacity can be achieved
without stopouts in main equipment or process technology. The Shell
PMR process can be designed to produce different grades of LNG. The
installed driver power can be utilised fully by changing the
cut-point temperature between the pre-cooling and liquefaction
refrigerant cycles.
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3.4 Linde
Linde Engineering is based in Munich.
Their business model is to market their patented technologies
from early concept through to implementation. Therefore the
Statoil/Linde MFCP is exclusively available through Linde.
Linde is highly experienced with cryogenic processing
technologies (to be expected for a company whose founder Karl Linde
in 1873 invented the first mechanical refrigeration), with many
decades of ethylene, gas processing, air separation and LNG
experience.
Linde is well established in providing small scale LNG peak
shaving plants (with 5 plants, starting with their first LNG peak
shaver in 1972), base load plants (with 4 very small plants plus
Xinjiang and Snohvit) and satellite plants (with 7 small
plants).
Their flagship LNG project is the Statoil-Linde process 4.3Mtpa
Snohvit LNG project due for commissioning mid-2007. Their next
biggest plant constructed is the 0.43Mtpa Xinjiang Phase 1 plant in
China, and like Black & Veatch they have submitted a LSTK bid
for the ~0.8Mtpa Phase II expansion of this.
They are co-developer with Statoil of the LNG liquefaction
technology being utilised for the Snohvit LNG development in
Norway. The Statoil/ Linde LNG Technology Alliance was established
to develop alternative LNG baseload plants for the North Sea and
this work resulted in a new LNG baseload process, the Mixed Fluid
Cascade Process (MFCP).
Additionally, Linde manufacture cryogenic exchangers (both plate
fin heat exchangers and spiral wound heat exchangers) and have done
so for decades. Linde pioneered spiral wound heat exchangers in
early 1900s. As well as being used within their own process, Linde
spiral wound heat exchangers (SWHE) are being installed within APCI
and Shell liquefaction processes, in new projects and as
replacement for old APCI cryogenic exchangers, on many world-scale
LNG trains. The MFCP uses three mixed refrigerants to provide the
cooling and liquefaction duty. Pre-cooling is carried out in a
plate fin heat exchanger (PFHE) by the first mixed refrigerant, and
the liquefaction and sub-cooling are carried out in spiral wound
heat exchanger (SWHE) by the other two refrigerants. The SWHE may
also be used for the pre-cooling stage. The refrigerants are made
up of components selected from methane, ethane, propane and
nitrogen. The 3 refrigerant compression systems can have separate
drivers or integrated to have 2 strings of compression. Frame 6 and
Frame 7 gas turbine drivers have been proposed for large LNG trains
(> 4Mtpa). A novel feature of the Snhvit project is that all
electric motor drivers will be used for the main refrigerant
compressors, with sizes up to 60 MW.
Within this proprietary process, purified natural gas is
pre-cooled, liquefied and sub-cooled by three separate mixed
refrigerant cycles. The pre-cooling cycle's cold is transferred to
natural gas via two PFHEs, whereas the cold of the liquefaction and
sub-cooling cycle is transferred via two SWHEs by the other two
refrigerants. The three refrigerant compression systems can have
separate drivers or be integrated to have two strings of
compression.
The MFCP is a classic cascade process with one important
difference - mixed component refrigerant cycles replace single
component refrigerant cycles, and thereby improve thermodynamic
efficiency and operational flexibility. The MFCP concept is built
up by well-known elements. The size and complexity of the separate
SWHE applied in the MFCP are considerably less when compared with
today's single unit used in dual-flow LNG plants. Last, but not
least, MFCP allows larger, single compressors to handle refrigerant
over a larger temperature scale.
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For Snohvit the process facilities are barge-mounted for ease of
fabrication, to save costs since the site is remote. The barge
forms the permanent foundation for the process equipment. The
barge, with process plant facilities installed, was fabricated in
Spain and dry towed to Hammerfest on a heavy lift vessel.
The Snohvit process barge being floated into its prepared space
at site:
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The completed Snohvit LNG plant on Melkoya Island,
Hammerfest:
3.5 Axens
Axens is based in Paris.
The Institute Francais de Petrole (IFP) developed the Liquefin
process that is licensed through Axens. While this process has not
yet been accepted for a commercial application, BP invested in its
evaluation for commercial application and it is being considered
for several projects in Iran (where companies such as APCI and
ConocoPhillips are excluded due to USAs trade embargo).
Claims made by Axens for the Liquefin process are that it
produces LNG cheaper than with any other process and that very high
capacities can be reached with a simple scheme and standard
compressors. No plants have been built so these claims are yet not
proven.
The Liquefin process is a two-mixed refrigerant process designed
for base load projects of train sizes up to 6Mtpa. All cooling and
liquefaction is conducted in plate-fin heat exchangers (PFHE),
arranged in cold boxes. The PFHE arrangement is at the heart of the
liquefaction technology. The PFHEs are non-proprietary and can be
supplied by a number of independent vendors.
The refrigerants are made up of components from methane, ethane,
propane, butane and nitrogen. The first mixed refrigerant is used
at three different pressure levels to pre-cool the process gas and
pre-cool and liquefy the second mixed refrigerant. The second mixed
refrigerant is used to liquefy and sub-cool the process gas.
Using a mixed refrigerant for the pre-cooling stage, the
temperature is decreased down to a range of - 50C to - 80C
depending on refrigerant composition. At these temperatures, the
cryogenic mixed refrigerant can be completely condensed, no phase
separation is necessary and, moreover, the quantity of cryogenic
refrigerant is substantially reduced.
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Overall necessary power is decreased, as the quantity of
cryogenic mixed refrigerant is lower, and a good part of the energy
necessary to condense it is shifted from the cryogenic cycle to the
pre-refrigeration cycle. Moreover, this shifting of energy allows a
better repartition of exchange loads. The same number of cores in
parallel can be used all along between the ambient and the
cryogenic temperature, allowing a very compact design for the heat
exchange line. A very significant advantage of this new scheme is
the possibility to adjust the power balance between the two cycles,
making it possible to use the full power provided by two identical
gas drivers.
The Liquefin process is very flexible. It offers more than one
possibility to reach large, highly competitive capacities, either
by using very large gas turbines (combined cycle) to produce
electricity, and large electrical motors (up to 70 MW) in parallel
on each cycle, or by using larger gas turbines. Frame 7 gas
turbines are proposed for large LNG trains. The Frame 9 has very
recently been qualified for mechanical drive. With Liquefin, this
would allow capacities of 7 to 8 Mtpa with only two main
drivers.
The process has been reported by Axens as representing a real
breakthrough with a total cost reduction per ton of LNG of 20% when
compared to the APCI C3-MR process. The cost reductions arrive
from: 1) increasing the plant capacity; 2) reducing the heat
exchanger costs; 3) all-over plate-fin heat exchangers; 4) a
compact plot area; and 5) multi-sourcing of all equipment,
including heat exchangers.
It is particularly well-adapted to the range of 4 to 8Mtpa per
train.
4. Associated Facilities As well as the LNG liquefaction process
plan described above, other specialised facilities are required for
an LNG development including those systems for; storage, loading,
LNG pumpout and boil-off handling.
Storage
The system consists of one or more specially designed tanks.
Ships to transport the LNG arrive at the terminal at specific
intervals. The minimum required storage capacity is the volume of
LNG discharged to the largest ship expected at the terminal. In
practice, the installed storage is larger than this minimum. The
extra storage provides a cushion to account for scheduled and
unscheduled delays in ship arrival. The storage tanks represent a
substantial capital cost. The volume of LNG stored in these tanks
is large and a failure of one or more tanks could have disastrous
consequences. Because of the exacting design and operational
techniques used, the modern LNG industry has had an excellent
safety record.
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Loading
The LNG loading system consists of all the facilities,
infrastructure and equipment required to safely dock the LNG ship,
to establish the necessary ship-to-shore interfaces, and for
transferring the cargo from the onshore tanks to the ship's tanks.
The system also includes facilities for disconnecting the
ship-to-shore interface at the end of the loading operation, and
for undocking the ship prior to its return voyage.
Specifically, the loading system consists of:
Breasting and mooring dolphins for securing the LNG ships to the
loading berth;
The loading platform which supports the loading arms and the
control building;
The control and emergency safety systems housed within the
control building;
The liquid loading arms for transferring LNG from the shore
piping to the ship;
The vapour return arm for returning vapour from the ship to
shore (created from the ship's cargo tanks because of the LNG being
pumped in);
Connections for transfer of utilities (e.g. nitrogen) from the
shore to the ship;
The piping, valves and vessels required for transferring the
LNG;
The return vapour and the utilities between the loading berth
and the main terminal facilities; and
If the loading berth or jetty is some distance from shore, a
pier connecting the jetty to the shore, will be required to provide
both access to the jetty and support the ship-to-shore piping.
LNG Pumpout
The LNG tanks operate at very low pressure, just slightly above
atmospheric pressure. Pumping of cryogenic liquids, especially at
the high rates required in LNG facilities, is specialized
technology. In modem terminals the first stage pumps are almost
always installed inside the storage tanks, and referred to as
in-tank pumps. The second stage pumps, when required (if for
instance there is a long loading jetty), are located outside the
tanks. These second-stage pumps discharge at a pressure
sufficiently high to satisfy the battery limit pressure at the
terminal fence.
Boil-Off Handling
LNG is a cryogenic liquid having a temperature, at atmospheric
pressure, of about -1620C. Heat entering the LNG (often referred to
as "heat in-leak") causes the LNG to warm up. However, in the
storage tanks the LNG needs to be maintained at a sufficiently low
temperature, consistent with the low operating pressure. Hence,
heat absorbed by the LNG has to be released by "flashing" (or
boiling-off) some of the liquid to gas.
Handling of boil-off gas requires compression equipment that is
costly to install and operate. Every effort is made to reduce the
amount of boil-off gas produced. Three main factors cause LNG
boil-off;
The LNG loaded into the ship may be slightly warmer than the
temperature in the storage tanks;
The energy used by the loading pumps is ultimately transferred
to the LNG as heat; and
Ambient heat transferred into the LNG through the cryogenic
insulation in pipes, equipment and storage tanks.
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Boil-off gas is essentially gasified LNG at atmospheric
pressure. It has substantial fuel value and, except in emergencies,
should not be vented or flared. Design and operation of the
boil-off gas handling system involves setting priorities for its
efficient disposition, including its utilization as fuel gas and
reliquefaction. Boil-off gas generated in the loading mode can be
many times greater than the gas generated in the period between
loadings (the period between loadings is referred to as the
"holding mode"). Hence, larger compression equipment may be needed
for the loading mode.
4.1 LNG Storage Tanks
LNG storage tanks account for a large portion of the cost of an
LNG plant. LNG is stored in double-walled tanks at atmospheric
pressure. The storage tank is a tank within a tank, with insulation
between the walls of the tanks.
Important factors to consider while specifying the LNG storage
system include:
Total storage capacity required;
Number of tanks required;
Type of containment preferred;
Applicable codes to be used; and
Other considerations like tank internals, commissioning and
insulation.
4.1.1 Total Storage Capacity
Determination of total storage capacity is seldom a simple and
straightforward exercise. Clearly, the minimum required capacity
would be the volume of the largest LNG tanker expected at the
terminal, plus a small margin above this (buffer). Another way to
look at the storage requirement is in terms of number of days of
LNG production.
The number of days of buffer (spare volume) storage capacity
varies. For instance, the Darwin LNG storage tank at 188,000m3
capacity has only some 2.5days of buffer when loading 145,000m3
ships. 5 to 10 days is more typical however.
Computer simulations are helpful in fine-tuning decisions
regarding the LNG storage capacity, and also the number of LNG
ships, their size and speed, and their utilization among different
facilities. It is important to note that the primary determinant of
storage capacity is the philosophy adopted by the owners. Computer
simulations can be used as a tool for fine tuning the capacity,
after the basic philosophy has been established.
The theoretical volume of storage required, assuming there are
no delays in LNG ship arrivals and no variations in LNG production
rate, is easy to calculate. In practice there will be events-both
scheduled and unexpected-that will cause deviations from this
theoretically ideal situation. These could include, for example,
predictable events like maintenance turnaround at the liquefaction
plant, scheduled maintenance for the LNG ships, seasonal variations
in LNG delivery, maintenance at downstream power plants, or
seasonal variations in sendout requirements. Other disruptions that
are anticipated, but whose timing cannot be predicted, might
include unscheduled downtime at the liquefaction plant,
weather-related ship delays, or unexpected downtime at the power
plant.
4.1.2 Number of Tanks
Once the total storage capacity is established the number of
tanks should be decided.
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The minimum number of tanks can be determined based on the total
storage volume required and the maximum capacity of a single tank.
The latter number however is not fixed and will depend on the type
of containment, type of construction and applicable codes.
For Greenfield LNG export plants the norm has been to install
two 50% tanks as part of the initial development. This was to
provide assurance that at least one tank would be available should
one tank require scheduled or unscheduled maintenance. For Darwin
LNG however, the Operator (Phillips) weighed up the risks and made
the decision to install a single tank (at 188,000m3, the largest
above-ground tank installed at that time). Despite the excellent
record of reliability with LNG tanks many owners prefer two smaller
sized tanks instead of a single large tank. The baseload nature of
the facility and the implications of a long-term take or pay
contract often favour multiple tanks.
Most of the LNG tanks in service have capacities of 100,000
cubic meters or less. Aboveground tanks with an inner metal wall
have now been built for capacity as high as 200,000 cubic meters.
However, the maximum capacity is limited by the availability of the
required wall thickness 9% nickel plate. Below-ground tanks using
the membrane type design with reinforcing concrete have been built
for capacities as high as 200,000 cubic metres, and above-ground
tanks with concrete inner and outer walls have been proposed for
250,000 cubic metres capacity.
In specifying storage tank capacity it is important to remember
that the "usable" volume in the tank is less than the total tank
volume. The minimum level to which the LNG in the tank can be
lowered, will be limited by the LNG pumps' ability. Similarly, to
avoid tank overfill it will be necessary to limit the maximum fill
level to less than the full height of liquid container. The ratio
of usable volume to built-up volume will depend on the tank height,
the pumpout arrangement, the LNG pump characteristics, and the
instrumentation/control philosophy Typically only about 95% of the
volume is usable (for example the Darwin LNG of nominal 188,00m3
capacity has actually some 200,000m3 of built capacity, or 94%
usable). The LNG at the bottom of the tank that isnt used is called
the heel.
4.1.3 Type of Containment
A variety of storage tank designs have been employed in LNG
service, all of which are characterized by their heavy insulation
and special material requirements.
LNG storage tanks are classified in terms of containment type
(single, double or full), and erection method (in-ground,
semi-buried and above-ground). All three containment types are
designed to store LNG safely and contain any spills in the unlikely
event of a leak in the primary liquid container, and are designed
with an inner and outer wall separated by insulation materials. The
inner wall must be designed for LNGs low cryogenic temperature and
the material used most extensively is 9% nickel steel, as this
remains ductile at cryogenic temperatures.
Safe use of LNG, or any cryogenic substance, requires an
understanding of how materials behave at cryogenic temperatures. At
extremely low temperatures, carbon steel loses its ductility and
becomes brittle. Therefore, the material selected for tanks,
piping, and other equipment that comes in contact with LNG is
critical. The use of high nickel content steels, aluminium, and
stainless steels is costly but necessary to prevent embrittlement
and material failures. High alloy steels composed of 9% nickel and
stainless steel will be used for the inner tank of LNG storage
tanks and for other LNG applications. In much of the discussion
above the primary liquid containment is assumed to be constructed
of 9% Ni steel. In addition to 9% Ni, other materials that are
suitable for cryogenic service include aluminium and stainless
steel. Aluminium is no longer considered economical for large LNG
tanks. However, stainless steel is a viable material and is
routinely used in the membrane-type design. The membrane technology
for LNG tanks relies on a post-tensioned concrete outer tank for
structural strength and a steel-corrugated membrane for liquid and
gas tightness. Membrane type tanks have been used extensively in
Japan where in-ground tanks have been built with capacity as high
as 200,000 cubic meters. Membrane technology has also been used
successfully for above ground LNG tanks.
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The main features of the typical single, double and full
containment tanks are summarised below:
1. Single Containment Tank
Double-walled with an interior tank is made of 9% nickel, while
the outer tank is made of carbon steel;
Only the inner primary container is required to meet the low
temperature ductility requirement for
storage of the product;
The outer container serves primarily to contain insulation and
vapour and to provide a weather shield. In the event of leakage
from the primary container the outer tank is not designed to
contain the refrigerated liquid;
The tank is surrounded by a bund wall or dike to contain any
leakage;
Single containment tanks are less expensive and rely on a
separate impoundment to contain the design
spill;
The required distance between the earthen type bund wall and the
tank adds significantly to the total land area. This type of
impoundment system has a large footprint, resulting in a large heat
flux exclusion zone;
The cost of a single containment tank is about 65% that of a
corresponding full-containment tank. If land is scarce this cost
advantage might be reduced; and
The construction time for a single containment tank will be
about four months less compared to a full or double containment
tank.
2. Double Containment Tank
Both the inner self-supporting primary container and the
secondary container are capable of
independently containing the refrigerated liquid;
The secondary tank, typically a concrete wall, is located
outside the primary tank. In the event of a leak, the secondary
tank contains the cryogenic liquid and limits the surface area and
vaporization of an LNG liquid pool. However, it is not intended to
contain any vapour resulting from such a leakage;
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The primary liquid container and the shell to contain the
insulation are similar to a single-containment tank. However, in
addition to it there is a surrounding concrete wall that is capable
of containing the cryogenic liquid in the event of a leakage from
the primary container. Unlike the bund wall surrounding a
single-containment tank, this wall is located close to the primary
container. This ensures that the liquid pool, in the event of LNG
leakage, has a smaller surface area compared to the
single-containment system.
The 2 x 140,000m3 LNG double containment storage tanks at the BG
Idku, Egypt LNG facility:
3. Full Containment
A full containment tank typically consists of a 9% Ni inner tank
with a prestressed concrete outer tank. The reinforced concrete
roof is lined with carbon steel, with the liner also functioning as
formwork for the concrete.
Both the self-supporting primary container and the secondary
container are capable of independently containing the refrigerated
liquid;
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The inner tank contains the LNG under standard operating
conditions. The outer shell, bottom and roof are made out of
pre-stressed concrete;
The outer tank supports the outer roof and is also intended to
contain the LNG;
Full containment tanks offer the highest level of safety;
The outer tank or wall composed of approximately 1 metre of
concrete is one to two meters away from
the inner tank;
The outer tank is capable both of containing the refrigerated
liquid and of controlled venting of the vapour resulting from
product leakage after a credible event;
The outer tank, which includes a reinforced concrete roof lined
with carbon steel, can be designed to
withstand realistic impacts from missiles or flying objects;
Concrete provides good resistance to heat radiation from nearby
LNG fires. There will be a significant time delay before structural
weakening of the reinforcement occurs;
Concrete also provides good protection against possible LNG
spills on the tank roof. The effects of cold-shock, if any, will
most likely be restricted to a small area, and generally should not
affect the vapour-tight integrity of the tank;
The cost significantly more and require about six months longer
to construct than the equivalent single containment tanks; and
Typically, full containment type tanks are used in sites where
the public is nearby or where security
issues exist. Outside the U.S., virtually all new above-ground
LNG storage tanks have been either full-containment or
double-containment designs.
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4.1.4 Pump Column for In-tank Pumps
Modern storage tanks have no side or bottom penetrations. All
penetrations, including those for LNG sendout, are through the roof
of the tank to avoid siphoning of the full content of the tank in
case of piping failures. This design substantially reduces the
amount of LNG spilled in the unlikely event of a rupture or leakage
in the sendout piping.
In-tank pumps are provided to transfer the LNG out of the tanks
and into the sendout system. In older facilities the LNG pumps were
usually located external to the tank, and a cryogenic line from the
bottom of the tank conveyed the LNG to the pump suction. In modern
facilities, for safety reasons, LNG tanks are designed with no
bottom or side penetrations. Instead, in-tank pumps, located at the
bottom of the tank and inside a pump column, are used. The
fabrication and installation of the pump columns requires
coordination with the pump supplier.
The LNG transfer pumps comprise in-tank pumps submerged within
pipe deepwells extending from the top of the domed roof to the
bottom of the inner tank. Typically, four pump deepwells are
installed, but only three wells will contain pumps, leaving one
well as a spare. Crane/winch facilities above the roof are used to
extract and re-install pumps undergoing maintenance.
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The 188,000m3 Darwin LNG storage tank showing the external riser
and pump access platform:
4.1.5 Tank Pressure Control
The tank pressure must be controlled within a narrow range.
During normal operation, the vapour handling system will increase
or decrease the boil-off' gas removal rate to maintain the required
pressure. However, properly designed over-pressure protection and
vacuum protection systems must be installed to handle upset
conditions and unusual circumstances. Typically, metal roof tanks
are restricted to a design pressure of less than 150 millibar
gauge. Concrete roof tanks can be designed to withstand a much
higher internal pressure, perhaps as high as 300 millibar gauge. A
higher design pressure allows a greater range of operating
pressures, and may also permit direct return of vapour to the ship,
without the need for compression. LNG tanks are usually designed
for vacuum conditions between 0 and -10 millibar gauge. Under
normal operation a vacuum condition is not expected, but a vacuum
protection system is required to safeguard against upset
conditions.
4.1.6 Purging and Cooldown
When an LNG tank is put into service, such as during initial
commissioning, the atmosphere in the tank has to be changed from
air to natural gas. Natural gas vapour is primarily methane which,
in certain concentration ranges, can form a flammable mixture with
oxygen. To avoid the possibility of forming a flammable mixture the
oxygen content in the tank must be reduced to less than 12%. In
practice a margin of safety is included and the oxygen content
should be reduced to around 8% or 9%. This is accomplished by
purging the tank with nitrogen, which is an inert gas. The annular
space between the inner and outer tanks contains the insulation,
usually loose perlite. Effective purging of the perlite is also a
requirement and means to accomplish this must be provided.
Cooldown of the tank is a sensitive operation, and must be
completed prior to filling it with LNG. Cooldown is accomplished in
a slow and gradual manner with cooldown rates (degrees per hour)
limited by the tank vendor specifications. Cooldown must be not
only gradual but also uniform, so that temperature gradients within
the tank are within the limits specified by the tank vendor.
Cooldown is accomplished by spraying liquid nitrogen or LNG into
the tank. A spray ring, located below the suspended deck of the
tank, ensures uniform spraying and cooldown. Sufficient number of
thermocouples, located at suitable intervals, are provided to
monitor the
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cooldown progress. If a source of liquid nitrogen is
conveniently available it would be preferable to have the first
tank cooled and ready to receive LNG when the first ship arrives.
Subsequent tanks can then be cooled using LNG from the first
tank.
4.1.7 Insulation
Insulation is necessary to limit heat leak into the LNG tanks.
Heat leak typically averages around 0.05% to 0.06% of full tank
contents per day. Different types of insulation are used in
different parts of the tank. Typically, the annular space between
the inner and outer tanks is filled with loose perlite (expanded
mica). In addition, a resilient blanket, such as fibreglass
material, is installed on the outside of the inner tank. This
blanket provides resiliency for the perlite as the tank contracts
due to temperature changes, and prevents settling of the perlite.
The blanket also facilitates flow of the purge gas during the tank
inerting process. In membrane type tanks an internal insulation
such as rigid PVC foam is used to transmit liquid pressure from the
membrane to the concrete tank.
Heat leak from the roof of the LNG tank is limited by installing
insulation on the suspended deck (which is suspended from tile
roof). There is no insulation immediately beneath the roof, and the
vapour space between the suspended