PEP Review 2003-15
BASE LOAD LNG BY CASCADE REFRIGERATION
By
David Netzer And Richard Nielsen
(December 2003)
ABSTRACT
Movement of natural gas from remote locations with abundant supply of natural gas to the
consuming countries can be economically achieved only via the liquefaction route with shipping
by ocean tankers. About 160 million metric tons/y of new liquefied natural gas (LNG) capacity is
being implemented or in various planning stages in addition to the existing 100 million mt/y of
LNG capacity of about 20 global facilities.
Until recently, cascade refrigeration systems for LNG production accounted for about 3.5%
of the LNG global market with significant production at ConocoPhillips plant in Kenai, Alaska.
Almost all the balance of the LNG market, about 90%, is predominantly propane pre-cooled,
mixed refrigerant systems. Recent marketing efforts by a collaboration of ConocoPhillips and
Bechtel are increasing the market share of cascade refrigeration technology.
This Review evaluates the economics of a base loaded, generic cascade refrigeration LNG
plant nominally producing at least 600 million scf per stream day (4.375 million mt/y at 0.95 on
stream factor) of LNG using two 50% capacity refrigeration and liquefaction trains. The feed gas
is lean, containing less than 8 vol% (17 wt%) C2+, and also has low nitrogen and CO2 contents
(less than 1 vol% and 1.2 vol% respectively). A generic LPG (liquefied petroleum gas) recovery
process is used and the nitrogen stripping step is avoided. The refrigeration systems use a
combined cycle mode of propylene and ethylene cycles driven by gas turbines and methane
refrigeration driven by steam turbines. A closed loop methane refrigeration system is used for
this lean gas plant, however, an open loop methane cycle could be considered as an alternate.
The competing mixed refrigerant cycle technology is given a cursory review.
The proposed design represents a relatively low greenhouse gas emission plant (0.20 ton
CO2/ton of LNG, as opposed to a typically reported 0.25-0.35 ton CO2/ton LNG) with low NOx
emission.
INTRODUCTION
The production of Liquefied Natural Gas (LNG) has been commercially practiced since
1960. Since LNG is stored and delivered at atmospheric pressure and -160°C (-256°F), very
deep refrigeration is needed with the associated large energy consumption. Two types of LNG
plants have been built:
•
Peak shaving plants for seasonal adjustment and storage, mostly in the USA
•
Base load LNG plants for international trade with LNG shipped by dedicated LNG ocean
tankers.
The capacities of the peak shaving plants are rather small, more than an order of magnitude
smaller than base load LNG plants. Based load LNG is the focus of this report. Typically, the
liquefaction energy in a LNG plant is reported to be in the range of 9-12% of the heat energy in
the natural gas where 9-10% energy shrinkage is known to be a typical number for modern mega
tonnage capacity avoiding combined cycle systems. The capital investment in modern LNG
facilities is reported to be over $1.0 billion with 45-60% attributed to offsites and infrastructure
depending on one’s definition of offsites, particularly LNG storage, the marine system, and the
heat rejection method.
The thermal efficiency of LNG plant is determined by two major factors:
(a) The refrigeration cycle efficiency
(b) The power cycle efficiency.
Increasing the thermal efficiency of a LNG plant for a given turbine/driver configuration will
minimize on site gas consumption and will minimize greenhouse gas emissions of the
combustion CO2. The proposed design represents a relatively low greenhouse gas emission
plant (0.20 ton CO2 /ton of LNG, as opposed to a typically reported 0.25-0.35 ton CO2/ton LNG
[9]) also having low NOx emission (0.095 kg/ton). Further, the thermal inefficiency, of 9-12%, is
nearly proportional to the heat rejection and would almost proportionally affect the LNG
production rate for a given turbo machinery configuration and thus indirectly affect the cost of
production.
For evaluating the process economics, we use a base case of lean gas feed, 92 vol% (85
wt%) methane containing under 1.2 vol% CO2 and less than 1 vol% nitrogen supplied at 46.5
kg/cm2-a (650 psig). The nominal production rate of 525 mt/h (4.375 million mt/y at 0.95 on
stream factor or 600 million scf/sd) of LNG depends upon the average ambient air temperature
and the heat rejection temperature. The production concept is based upon two treating,
liquefaction and heat recovery trains each 50% of total plant capacity. Each train uses one
General Electric F7A frame, or equivalent gas turbines such as Rolls-Royce’s Trent model, in the
refrigeration system. Therefore, the performance of the gas turbine is a key factor in establishing
the LNG production rate. Waste heat is being recovered as 63 kg/cm2-g, 482°C (900 psig,
900°F) steam from the gas turbines exhaust gas at about 545°C, 300 mm water-g (1,020°F, 12
in. water-g). This steam drives two 50% methane refrigeration compressors and, on a separate
shaft, two power generation turbines. All turbines condense at summer conditions of 40.6°C, 57
mm Hg (105°F, 1.1 psia) using sea water at 27°C (81°F) maximum and discharged to 31-32°C
(88-90°F) maximum.
Unlike ConocoPhillips’ new design that uses an open loop methane refrigeration cycle, we
selected a closed loop cycle while recognizing that the open loop concept could avoid a separate
fuel gas compressor, an issue to be evaluated on a case specific basis [1,4,5]. Further, unlike
ConocoPhillips, we have decided to use propylene refrigerant rather than propane refrigerant,
thus reducing the volumetric refrigerant flow by 22% over propane and allowing more vendor
selection options for a single propylene compressor.
The recent design trend by both Bechtel and Kellogg Brown & Root seems to suggest the
use of air cooling as a method of heat rejection from the refrigeration cycle. From public
information released by Linde AG it is obvious that seawater cooling is their preferred method [2].
Needless to say this is a site specific issue. However, based on our preliminary analysis,
especially for fixed speed gas turbines such as the General Electric frame 7, unless heat
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PEP REVIEW 2003-15
rejection to sea water is prohibited by local law, water cooling would be the more economical
method of heat rejection in most cases: fresh water as first priority; once through sea water as a
second priority. A combination of these including rejecting the heat to seawater via a tempered
water cycle (loop) should be considered on a case by case basis.
The key capital cost elements in LNG facilities are in descending order:
1. Gas turbines steam turbine or motor drivers for refrigeration service.
2. Refrigeration compressors, typically over 100,000 kw
3. Steam and power generation including turbines waste heat recovery.
4. LNG storage, typically over 50,000 ton. (1,000,000 bbl)
5. LNG loading terminal including jetty or causeway.
6. Heat rejection system, in most cases, as suggested, seawater cooling.
7. Cold box and pre-chilling for gas liquefaction
8. Natural gas pre-treating for CO2 removal gas drying and mercury adsorption.
9. LPG and natural gasoline recovery as by products.
10. Fuel gas cold recovery and compression.
The above cost ranking illustrates that heat transfer, i.e. the cold box, is of secondary
importance compared with the capital investment associated with compression and several sitespecific factors. Nevertheless, the correct selection of the optimized refrigeration cycle and
associated drivers affects the compression and heat rejection systems. This choice becomes an
important item in the project evaluation especially that many of the site specific items are totally
independent of the liquefaction cycle technology.
The utilization of stranded gas as a source of energy to distant users, although expensive,
appears very economical as compared to gas to liquids. Recent announcement of GTL project
34,000 bpsd mixed liquids calls for investment of $675 million (or just under $20,000 per daily
barrel, which is just less than 80% of the benchmark $25,000 per daily barrel). The heat
equivalent of 34,000 bpsd is estimated at 200 million scfd of LNG, 1.46 million mtpy. On this
basis the GTL capital investment is 2.2 times higher per Btu and even after adjustment for
economy of scale costs about 1.8 times higher than modular construction of a LNG plant of 300
million scfd [12].
It is recognized that cost of transporting LNG say $0.75/MM Btu HHV ($35/ton) is double the
cost of transporting crude oil, about $17-21/ton. Further it is recognized the cost of re-gasification
(about $16 per ton) is avoided. Nevertheless even after considering all these factors LNG seems
by far the more economical fuel.
The thermal efficiency of GTL, based on 330 million scfd feed gas, is about 60% as opposed
to 94% efficiency for LNG. The additional gas consumption increases the cost of GTL by
additional $27 per ton, which cancels out the re-gasification, and half of the transportation cost
advantage of GTL.
The greenhouse gas emission from GTL production is about 2.0 ton of CO2 per ton of
synthetic crude, 10 times higher over the LNG case. The synthetic crude is yet to be fractionated
in a petroleum refinery to produce fuel oil to compete with LNG.
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CONCLUSIONS
From our review of the literature and economic evaluation, we conclude:
1. It is our judgment that the thermal efficiency of the cascade refrigeration design, as
suggested, is close or equal to the efficiency of the propane pre-cooled mixed
refrigerant system and somewhat lower than mixed fluid cascade (MFC) refrigeration.
However, at the end, no one process for gas liquefaction has a decisive advantage over
other processes that would make an LNG project economically viable in a given location
while a competing process technology would not. The calculated efficiency is 93.4%
plus 13.0 MW of electric power export, which is equivalent to 0.5% additional efficiency,
thus the total equivalent efficiency is 93.9%.
2. The production rate of LNG can be easily designed to increase from 526 mt/H (4.375
million mt/y) to 555 mt/H (4.619 million mt/y) by using the startup steam turbines on a
continuous basis as booster turbines and utilizing the auxiliary steam generator on a full
time basis. No electric power will be exported under this scenario.
3. Reducing the methane content of the raw natural gas feedstock from 92 vol% to 85
vol% increases production capacity by 0.2% while maintaining the same gas turbine
power. The LPG and natural gasoline production increases by 469%.
4. Increasing the feed gas pressure from 650 psig to 775 psig decreases LNG production
by 0.5 wt% at a constant raw gas plant feed rate but the total refrigeration compressor
horsepower decreases 2.0%.
5. Avoidance of superheating the suction to the methane compressor from about –153°C (243°F) to –101°C (-150°F) will reduce the total refrigeration power requirement by 1.7%
over the base case, which superheats the methane to –101°C (–150°F). However, this
option increases the capital investment by about $4.5 million.
6. The key factors in the overall relative economics of LNG are related to the correct
selection of the refrigeration compressors, the configurations of the compressor drivers,
the method of heat rejection, as well as some very site specific factors including climate
(especially ambient conditions), soil conditions, marine system design, and LNG
loading, storage and transportation logistics.
7. The total capital investment including 15% contingency is estimated to be $912 million.
The capital investment in LNG production is measurably lower than the investments
reported by others. However we believe our estimate, $181 per ton LNG based on US
Gulf Coast labor cost and productivity, that was corroborated by several sources of
expertise represents a realistic scenario. After adding 15% contingency the capital
investment is $209 per ton excluding the cost of land, owner’s cost (royalty fees), or any
unusual soil conditions.
8. Based upon Conclusion 7, using feed gas at $0.75/MM Btu higher heating value (HHV)
and by-products LPG at 9.62 ¢/lb (about 42 ¢/gal) and gasoline at $10.8 ¢/lb (about 65
¢/gal), the net production cost of LNG is 3.62 ¢/lb of LNG or $79.80 per ton LNG. At
23,550 Btu/lb (HHV), the net production cost at the LNG ocean tanker is US $1.54/MM
Btu HHV of LNG. For a return on investment of 25%, the product value is $2.54/MM Btu
HHV of LNG.
9. Based upon the $3.50/MM Btu HHV value of natural gas in the U.S., the product value
at the plant is estimated to be $2.60/MM Btu HHV or 6.12¢/lb of LNG after subtracting
transportation cost of $0.55/MM Btu HHV for about 2,000 miles shipping and $0.35/MM
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PEP REVIEW 2003-15
Btu HHV for re-gasification. At $2.60/MM Btu HHV, the plants return on investment
(ROI) before tax is estimated to be a favorable 26%.
LNG MARKET OVERVIEW
International LNG trade is expanding rapidly. In the 1980’s there were only two grass roots
LNG facilities built. In the 1990’s, six grass roots LNG plants were added to the list. Now, seven
new LNG plants are in advanced planning stage or in engineering and construction phases. As
shown below, over 90% of new LNG facilities are planned for tropical equatorial or sub-tropical
regions. Our selection of the generic ambient conditions reflects this reality. The median
capacity of the proposed new facilities is in the order of 4.5 million mt/y. This is a factor
considered in selecting our design capacity of 4.4 million mt/y for this evaluation.
Current LNG Production and Consumption
Total global LNG production is 110 million mt/y. Major consumption and production areas
are listed in Table 1. LNG Projects in advanced planning or the engineering and construction
phase total over 60 million mt/y as listed in Table 2. Twenty-two projects are reported to be
planed for the long term and total an additional 110 million mt/y (Table 3). Current and proposed
projects range in capacity from 3 to 11.2 million mt/y.
Table 1
MAJOR LNG PRODUCTION AND CONSUMPTION AREAS (2002)
Production
%
Consumption
%
Indonesia
23.4
Japan
48.0
Algeria
18.2
Korea
15.8
Malaysia
13.2
France
9.4
Qatar
12.2
Spain
9.1
Australia
6.5
Taiwan
4.8
Brunei
6.0
United States
4.7
Oman
5.6
Turkey
3.3
Nigeria
5.2
Belgium
2.4
Abu Dhabi
4.5
Italy
2.3
Trinidad
3.5
Greece
0.3
United States
(Alaska)
1.2
Libya
0.5
Source: [10, 11]
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Table 2
LNG PROJECTS -- ADVANCED PLANNING OR
ENGINEERING AND CONSTRUCTION STAGES
Project
Capacity, million mt/y
Trinidad, Atlantic LNG trains 3 & 4
8.5
Egypt, LNG train 1-2
7.2
Egypt, Damietta
5.0
Egypt, Idku
3.6
Nigeria LNG trains 4 & 5
11.2
Norway, Snohvit
4.0
Malaysia, Tiga
7.6
Qatar, Ras Laflan
4.7
Oman
3.0
Australia, North West Shelf, 4
4.2
Australia, Darwin
3.6
Total
62.6
Crude oil Equivalent
71 mm t/y, 1.5 mm BPSD
Estimated capital expenditure
$16,000 million
Sources: [7,11, 13, SRI Consulting]
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Table 3
LONG RANGE LNG PROJECTS
Project
Capacity,
million, mt/y
Project
Capacity,
million, mt/y
Equatorial Guinea
3.4
Tangguh, Indonesia
7.0
Brass Nigeria
5.0
North West Shelf, Australia
4.0
Trinidad, Atlantic LNG
5.2
Bontag Indonesia
3.0
Angola LNG
4.0
Sakhalin Island, Russia
9.6
Egypt LNG expansion
3.0
Bolivia
7.0
Egypt Damietta
5.0
Camisea, Peru
4.0
Nigeria LNG train 6
4.0
Gorgon, Australia
7.0
Algeria
4.0
Sunrise, Australia
5.0
Amazon
3.0
Yemen
3.0
Venezuela
4.0
Iran
4.8
Qatar, Ras Laffan 2 trains
15.6
_____
Total
110.6
Sources: [7, 11, 13, SRI Consulting]
LNG Terminals
In order to accommodate the projected production and consumption growth, new LNG regasification and unloading terminals need to be constructed. Forty terminals including 19 in
Japan and four in North America are in operation. Many terminals are in planning stage including
fourteen new LNG terminals for North America. The proposed North American terminals include
an off shore terminal near Oxnard, California where it is assumed LNG will be re-gasified on a
platform structure 21 miles from shore [3]. An additional terminal in Baja California Mexico is
mostly dedicated for Southern California. A LNG terminal in Freeport, Texas would become a
captive source for petrochemical feedstock for Dow Chemical’s ethylene production [3].
Interesting enough, a LNG terminal is also being contemplated by Repsol for the Gulf of Mexico
near Tampico, Mexico.
LNG Specifications
The main specifications on LNG in Europe, Japan, Korea and the US Gulf Coast and
Northeast are a maximum 0.5 vol% of C5 and heavier components and a higher heating value in
the range of 950 to 1,100 Btu/scf. Specifications in California are listed in Table 4.
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Table 4
CALIFORNIA LNG SPECIFICATIONS
Specification
Value
Methane, vol% min.
88
Ethane, vol% max.
6
Propane-Pentane, vol% max.
3
Hexane and Heavier, vol% max.
0.2
Nitrogen plus Carbon Dioxide, vol% max.
1.4-3.5
High Heating Value, Btu/scf
970-1150
Source: [8]
PROCESS REVIEW
The most common refrigeration system in prior LNG projects is the mixed refrigerant system
preceded by propane refrigeration. Close to 90% of these plants are licensed by Air Products
Corporation, Inc. (APCI). The mixed refrigerant liquefaction systems use a mixture of mostly
methane and ethane, about 1.2-2.0/1 mole ratio. Depending on the feed gas composition, up to
3 mol% nitrogen and 6-12 mol% propane may be added to optimize the refrigerant boiling curve.
The refrigerant cooling curve is adjusted to follow closely the feed gas cooling curve in order to
achieve maximum thermodynamic efficiency. The fundamentals of mixed refrigerant compared
to conventional refrigeration are discussed in PEP Report 29G Ethylene Plant Enhancement
(2001).
One known significant exception to the propane pre-chilling, mixed refrigerant approach is
the Phillips Kenai Peninsula plant in Alaska that started up in 1969 [1]. This plant produces about
1.5 million mt/y of LNG by a cascade refrigeration system using pure propane, ethylene and
methane refrigeration cycles. The efficiency of this cascade system has been reported to be on
the lower end of the scale, about 88% [US 5611216]. However, we understand the plant has
proven to be a reliable and profitable operation. Aside from refrigeration cycle as such, an
important factor in refrigeration power is the adiabatic efficiency of the refrigeration compressors.
It is reasonable to assume that the compressors operated by ConocoPhillips in Kenai, Alaska are
more likely to be on the order of 70% as opposed to more modern centrifugal compressors with
three dimensional blades that achieve adiabatic efficiencies above 80% and approach 85% as
we show later.
Over the past several years several patents issued to Phillips seem to suggest an
improvement in cycle efficiency results from improved refrigeration load distribution, nitrogen
stripping for nitrogen rich gas, open loop methane refrigeration and LPG recovery. Our opinion,
on a purely thermodynamic concept, is the overall thermal efficiency of the mixed refrigerant
system is slightly higher on a consistent basis (ambient conditions, identical machinery, heat
recovery and heat rejection philosophy). Nevertheless, other design related factors, operational
flexibility and other considerations, suggest that an objective comparative evaluation could be
made only on a case by case basis and on a site specific basis and not on generic liquefaction
technology as such nor on liquefaction efficiency or any other single factor [US 5669234, US
5611216 and US 4680041].
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PEP REVIEW 2003-15
Our understanding of Phillips’s patents is that their technology is driven by the desire to have
an equal load distribution among the propane cycle, ethylene cycle and methane cycle. This load
equalization is achieved in part by superheating the methane and ethylene refrigerant vapors fed
to the compressors probably to about -46°C (-50°F). This assumed design methodology allows
conventional carbon steel metallurgy in these compressors.
Design and construction
considerations then allow six identical gas turbines, such as frame 5D (nominal 30,000 kw), to be
used. Based on other information published by ConocoPhillips, we understand that the more
modern design uses open loop methane refrigeration where the methane refrigeration
compressor is used also as the fuel gas compressor [4]. Methane is fed as fuel to the gas
turbines after cold recovery, thus helping to equalize the loads while avoiding a fuel gas
compressor. Based on Phillips U.S. Patent 5,611,216, the estimated power consumption for
38°C (100°F) feed gas at 650 psig, is 371 kwh/ton, assuming very lean gas, say 96 vol%
methane, at an unreported heat rejection temperature but based upon Phillips brochures,
speculated to be 38°C (100°F) [4]. This estimate probably includes 15 kwh/ton of fuel gas
compression. Based on the above, it is our judgment the drawbacks of load equalization could
be up to 5% additional refrigeration power and, in case of fixed speed gas turbines, an actual
reduction in production capacity. Further, based on a patent and other information published by
ConocoPhillips, we believe that this could result in more complex refrigeration cycle where the
propane cycle, ethylene cycle, and methane cycle are more heat integrated [US 5611216, 4].
Nevertheless, at the end, we believe the concept of cascade refrigeration is very sound.
In our approach we judged that the older closed loop methane refrigeration as used by
Phillips in Kenai could have an advantage by allowing higher suction pressure to the methane
compressor, about 1.7 kg/cm2-a (24 psia) instead of an estimated 1.1 kg/cm2-a (16 psia) for an
assumed open loop compressor. Only one electric motor driven fuel gas compressor is used for
both trains. In case of outage of the fuel gas compressor, back up is provided by a draw from the
feed gas.
Our base case design includes superheating of the refrigerant gas to the methane
compressor. We also examine an alternate design where the refrigerant gas is not superheated.
Some reported experience by The Elliott Company compressing LNG tanker boil off gas at
suction temperatures under about –130°C (-200°F) and industrial experience by others
compressing nitrogen at –180°C (-292°F) seems to support our suggested alternate approach.
Nevertheless, a more conservative approach will call for superheating the suction of the first
stage of the methane compressor from about -152°C (-240°F) to about –101°C (-150°F). This
will slightly move toward equalizing the loads while increasing the total refrigeration load by 2%.
The very conservative potential operator is brought into a “comfort zone” in terms of compressor
metallurgy at low temperatures. Based on our proposed approach of using a combined cycle and
driving the methane compressors with steam turbines, many of the above issues as applied by
ConocoPhillips for the gas turbine power cycle are becoming academic for the combined cycle
approach suggested.
Further, it is our understanding that the ethylene refrigeration cycle in the ConocoPhillips
design, aside from superheating the suction to the first stage to about -46°C (-50°F), comprises
only a single side load with a probable goal of obtaining the compression in a single casing. In
our approach, no superheating of the suction to the ethylene compressors is employed. Two
side loads are used, reducing refrigeration load by 2% and increasing production capacity by 23%. However using two casings somewhat increases our capital investment, possibly by about
$5 million. Again, in the proposed approach, no attempt is made to equalize the loads. The
loads of the propylene compressor are about 47.3%, the refrigeration ethylene compression
loads are about 36.4% and the methane compressor loads are 16.3% of the total refrigeration
load. In reality the combined load on the steam turbines may be about 27% of the total motive
power in the facility. However 20% of the steam motive power, about 6.2% of the total power, is
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exported on the average as electric power outside the boundary limits. During high ambient
temperature times, the start-up steam turbine attached to the gas turbine increases its relative
steam consumption and electric power export drops to near zero.
The actual production capacity is determined by the performance of the industrial equivalent
of GE 7FA turbines installed on the same shaft as the propylene and ethylene compressors.
Turbine output as provided by GE is conservatively derated by 2-2.5% from the name plate
capacity to show realistic mid life performance. Derating of 4% is possible and typically serves
as a performance guarantee by the manufacturer. The generic 4.375 million t/y at 0.95 on
stream factor (600 million scf/d) of LNG is simply a result of turbine performance and is related to
ambient conditions and perhaps to the mechanical condition of the turbo machinery. In addition
to climatic assumptions, several generic assumptions are made as to infrastructure, feed gas
composition and LNG specifications. The cost estimate and subsequent economics are highly
dependant upon the specific site location but could be adjusted based on variance in
infrastructure. The site assumptions, offsites definitions and turbine performances are included in
the review. Appropriate adjustments could be made for different site locations and ambient
conditions. As said, the total capacity could probably be raised from 4.375 million mt/y to 4.60
million t/y at 0.95 on stream factor by shifting the steam turbine loads from excess power
generation to continuous boosting of the gas turbines.
In addition to the ConocoPhillips cascade refrigeration processes, Linde AG has developed
the Mixed Fluid Cascade (MFC) process where each of the three refrigerants cycles operate in
cascade using mixed refrigerant [US 6253574]. This concept is being engineered for the
Norwegian North Sea project offshore near Hammerfest and is claimed to be of extremely high
thermal efficiency, about 250 kwh refrigeration load per ton of LNG. However, in this project heat
is rejected to 5°C (41°F) seawater as opposed to the more conventional 27°C (81°F) in more
typical LNG projects in tropical zones. In this temperature range the claimed refrigeration load is
more like 330 kwh per ton [2] of LNG as opposed to 333 kwh refrigeration per ton LNG calculated
for the cascade refrigeration as proposed in this report with heat rejection at 35°C (95°F) to 27°C
(81°F) sea water [1-2, US 5611216, US 6253574]. The MFC process is reviewed in PEP Review
2002-7 (2003).
A recent interesting patent by Exxon suggests a new approach for producing pressurized
LNG (PLNG) [US 6,016,665]. LNG at about 30 kg/cm2-a and -95°C (410 psig, and -140°F) would
be produced by two cascade cycles using propane and ethylene as opposed to the current three
cascade cycles comprising propylene, ethylene and methane refrigerants. Producing PLNG is
estimated to save 50% or more of refrigeration power compared with LNG, along with avoiding
methane refrigeration, which substantially reduces the capital investment for the liquefaction
cycle. However, this PLNG concept brings yet unknown issues related to transportation and
storage of the PLNG.
PROCESS DESCRIPTION
In this section, we describe the nominal production of 525 mt/h (4.375 million mt/y at 0.95 on
stream factor, 602 MM scf/d) of LNG depending on average ambient air temperature and heat
rejection temperature. The feedstock is a lean raw natural gas containing about 92 vol% of
methane. The levels of nitrogen and carbon dioxide are also relatively low. The generic process
is a conventional cascade refrigeration process. The plant contains two equal capacity trains for
gas processing including treating to remove carbon dioxide and water. Liquid petroleum gases
are separated into byproduct natural gasoline, mixed C3/C4 LPG and fuel gas in a single train
unit. The fuel gas is consumed within the plant. The plant produces all its utilities; about 13,000
kwh of electricity can be sold outside the plant. The process design and utility rates used in this
report are based on computations, published information, and nonconfidential information from
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PEP REVIEW 2003-15
licensors and vendors. The design may or may not be similar to processes licensed or otherwise
in use today.
The production concept uses two 50% liquefaction trains with General Electric F7A frame or
equivalent gas turbines such as the Rolls-Royce Trent model. The performance of the gas
turbine is a key factor in establishing the LNG production rate. Waste heat is being recovered as
63 kg/cm2-g, 482°C (900 psig, 900°F) steam from the turbine exhaust gas at about 545°C-300
mm water-g (1,020°F, 12 in. water-g) driving two 50% methane refrigeration compressors and, on
a separate shaft, two power generation turbines. All condense at summer conditions of 40.6°C,
57 mm Hg (105°F, 1.1 psia) using seawater at 27°C (81°F) maximum and discharged at 32°C
(90°F) maximum.
The nominal design capacity is based upon the median production capacities of future LNG
plants. This led to using two 50% GE F7A gas turbines, or equivalent, along with two GE, or
equivalent, steam turbines operating at full capacity 347 days per year (0.95 on stream factor)
with a reasonable expectation to achieve 350 days per year. The result is a nominal 4.375
million mt/y of LNG at nominal atmospheric pressure, -160°C (-260°F). No liquid expanders are
used, but in general a liquid expander can reduce energy consumption by 2-2.5% but was judged
to offer only a marginal economic advantage.
The process flow diagram of the LNG plant, Figure 1 (attached), is shown with the inside
boundary limits equipment divided into five sections:
•
Amine Unit (Section 100), Sheet 1
•
Dryers and Mercury Removal (Section 200), Sheet 1
•
Liquefaction (Section 300), Sheets 2 and 3
•
LPG Fractionation (Section 400), Sheet 4
•
Heat Recovery System (Section 500), Sheet 5.
The offsite units, not shown in Figure 1, are divided into an additional four sections:
•
Storage (Section 600)
•
Marine System (Section 700)
•
Relief System (Section 800)
•
General Offsites (Section 900).
The design basis and assumptions are summarized in Table 5. The feed gas is
characterized in Table 6. The process stream flows are summarized in Table 7 and refrigerant
flows are summarized in Table 8. These flows are for the total plant consisting of two trains. The
process was modeled using Aspen Technology’s Aspen Plus process simulator, version 11.1-0,
built September 20, 2001. The physical properties method was the Soave Redlich-Kwong
equation of state method. Major equipment is listed with size and materials of construction in
Table 9. On the equipment list, the two trains are identified in the equipment numbers by A and
B. The letters C, D and higher indicate multiple pieces of equipment are used per train. The
facility is self-contained in terms of electric power, fuel gas and all other utilities. Internal utilities
average consumptions are summarized in Table 10. An average of 13,000 KW of interruptible
electricity is sold outside the plant. The steam used to generate this power could be shifted to
run booster turbines and increase LNG production by 6%.
11
PEP REVIEW 2003-15
Table 5
DESIGN BASIS AND ASSUMPTIONS
Plant capacity, scf/sd
Nominal 600 million
mt/yr
Nominal 4.4
On stream factor
0.95
Number of trains
2 (one LPG unit)
Site
Eastern Mediterranean
Elevation, ft
0 (sea level)
Barometric pressure, psia
14.7
Relative humidity, %
60
Ambient air temperature, °C (°F)
21 (70) average, 38 (100) maximum
Average January temperature, °C (°F)
14.5 (58); low 11 (52), high 18 (64)
Average April temperature, °C (°F)
19 (66); low 15 (59), high 23 (73)
Average August temperature, °C (°F)
27 (81); low 23 (73), high 31 (88)
Average October temperature, °C (°F)
24 (75); low 20 (68), high 28 (82)
Cooling water
Once through sea water
Cooling water temperature, °C (°F)
19 (66) average; 27 (80.6) maximum
Sea water maximum temperature rise, °C (°F)
5 (9) With built in capability to 4°C rise
Sea water circulation rate, ft/sec
12 maximum
Feed gas
Lean, sweet natural gas, see Table 6
Acid gas treatment
MEA absorption of CO2
Compressor drivers
Propylene/Ethylene
Gas turbine (GE frame FA7 or equivalent)
Methane
Steam turbine
Fuel gas
Electric motor
Compressor efficiency:
Propylene
Ethylene
Methane
Speed, rpm
Stage No.
Efficiency
2,600
1
0.837
2
0.842
3
0.838
1
0.824
2
0.847
3
0.844
1
0.775
2
0.788
3
0.814
3,600
8,500
Source: GE, The Elliott Company
12
PEP REVIEW 2003-15
Table 6
FEED GAS CHARACTERIZATION
Pressure, kg/cm2-a (psig)
46.6 (650)
Temperature, °C (°F)
17 (62.6)
Higher Heating Value, Btu/scf
1,095
Composition:
Volume %
Variance, vol%
Nitrogen
0.29
0.1-1.0
Carbon Dioxide
0.47
0.3-1.2
Hydrogen Sulfide
10 ppm
0-50 ppm
Methane
91.69
88.0-96.0
Ethane
4.66
2.5-5.5
Propane
1.78
0.8-3.0
Isobutane
0.34
0.2-0.6
n-Butane
0.41
0.2-0.6
Isopentane
0.15
0.1-0.4
n-Pentane
0.10
0.05-0.3
C6+
0.11
Total
100.00
13
PEP REVIEW 2003-15
Table 7
LNG BY CASCADE REFRIGERATION
STREAM FLOWS
3
CAPACITY: 525 MT/H (9,646 MILLION LB/Y OR 602 MILLION SCF/SD OR 17.056 M /SD)
LNG
AT 0.95 STREAM FACTOR
Stream Flows, lb/Hr
1
Mol Wt
2
3
4
Raw Natural Molecular
Gas Fresh Sieve Dryer
Amine Unit
Feed
Feed
Treated Gas Fuel Gas
Water
18.02
659
Carbon Dioxide
44.01
16,214
Nitrogen
28.01
5,833
Methane
16.04
1,062,337
Ethane
30.07
101,144
Propane
44.1
Isobutane
6
CO2 Rich
Gas
Prechilled
Gas
2
47
628
64
64
146
16,004
2
5,833
5,833
0
0
5,824
1,061,082
1,061,082
1,151
105
1,056,040
101,039
101,039
94
10
97,754
56,725
56,673
56,673
47
5
49,825
58.12
14,199
14,189
14,189
10
0
10,377
n-Butane
58.12
17,121
17,110
17,110
10
0
10,980
Isopentane
72.15
7,775
7,770
7,770
5
0
3,062
n-Pentane
72.15
5,703
5,698
5,698
5
0
1,803
Hexane + Heavier
86.18
6,806
6,806
6,806
0
0
747
1,294,516
1,277,205
1,276,266
1,517
16,752
1,236,415
587,183
579,330
578,905
688
7,599
560,828
72,339
71,855
71,802
82
405
70,859
Total, lb/H
kg/H
lb-mole/H
941
5
14
1
PEP REVIEW 2003-15
Table 7 (Continued)
LNG BY CASCADE REFRIGERATION
STREAM FLOWS
3
CAPACITY: 525 MT/H (9,646 MILLION LB/Y OR 602 MILLION SCF/SD OR 17.056 M /SD)
LNG
AT 0.95 STREAM FACTOR
Stream Flows, lb/Hr
Mol Wt
7
8
9
10
11
12
Raw LPG
LNG to
Storage
LNG
Product
LNG Tank
Flash
N2/Fuel
Flash
Cold Fuel
Gas
Water
18.02
1
Carbon Dioxide
44.01
62
2
2
0
0
0
Nitrogen
28.01
9
2,839
2,147
692
2,985
3,677
Methane
16.04
5,042
998,451
982,403
16,048
57,589
73,637
Ethane
30.07
3,285
97,743
97,740
3
12
14
Propane
44.1
6,848
49,825
49,825
0
0
0
Isobutane
58.12
3,812
10,377
10,377
0
0
0
n-Butane
58.12
6,130
10,980
10,980
0
0
0
Isopentane
72.15
4,709
3,062
3,062
0
0
0
n-Pentane
72.15
3,895
1,803
1,803
0
0
0
Hexane + Heavier
86.18
6,059
747
747
0
0
0
Total, lb/H
39,851
1,175,829
1,159,087
16,742
60,586
77,328
kg/H
18,076
533,347
525,753
7,594
27,481
35,075
941
67,162
66,137
1,025
3,697
4,722
lb-mole/H
1
15
1
0
0
0
PEP REVIEW 2003-15
Table 7 (Concluded)
LNG BY CASCADE REFRIGERATION
STREAM FLOWS
3
CAPACITY: 525 MT/H (9,646 MILLION LB/Y OR 602 MILLION SCF/SD OR 17.056 M /SD)
LNG
AT 0.95 STREAM FACTOR
Stream Flows, lb/Hr
Mol Wt
13
14
15
16
17
C3+
Natural
Gasoline
LPG
Product
C2 Rich
Fuel Gas
Total Fuel
Gas
Water
18.02
1
1
0
0
47
Carbon Dioxide
44.01
61
62
0
0
146
Nitrogen
28.01
0
0
0
9
3,686
Methane
16.04
2
0
2
5,040
79,828
Ethane
30.07
119
0
118
3,167
3,276
Propane
44.1
6,737
72
6,628
148
195
Isobutane
58.12
3,858
179
3,630
3
14
n-Butane
58.12
6,179
420
5,709
1
12
Isopentane
72.15
4,733
4,605
104
0
5
n-Pentane
72.15
3,911
3,853
42
0
5
Hexane + Heavier
86.18
6,067
6,049
10
0
0
Total, lb/H
31,669
15,241
16,243
8,369
87,214
kg/H
14,365
6,913
7,368
3,796
39,559
521
201
317
423
5,227
lb-mole/H
16
PEP REVIEW 2003-15
Table 8
LNG BY CASCADE REFRIGERATION
REFRIGERANT STREAM FLOWS
3
CAPACITY: 525 MT/H (9,646 MILLION LB/Y OR 602 MILLION SCF/SD OR 17.056 M /SD)
LNG
AT 0.95 STREAM FACTOR
Propylene Refrigerant
Ethylene Refrigerant
Methane Refrigerant
Stream
lb/H
No.
kg/H
Stream No.
lb/H
kg/H
Stream No.
lb/H
kg/H
20
3,338,600
1,514,363
30
1,730,000
784,715
40
615,000
278,959
21
2,850,752
1,293,079
31
625,396
283,675
41
341,036
154,691
22
2,850,752
1,293,079
32
1,351,198
612,893
42
215,437
97,721
23
2,446,704
1,109,806
33
1,104,604
501,040
43
273,964
124,268
24
2,013,637
913,370
34
1,204,604
546,399
44
274
124
25
2,013,637
913,370
35
994,372
451,039
45
215,437
97,721
26
1,706,387
774,004
36
779,588
353,615
46
113,966
51,694
37
779,588
353,615
47
113,966
51,694
38
702,537
318,665
48
92,911
42,144
17
PEP REVIEW 2003-15
Table 9
BASE LOAD LNG BY CASCADE REFRIGERATION
MAJOR EQUIPMENT
CAPACITY: 9,646 MILLION LB/YR (4,375,300 mt/y)
LIQUIFIED NATURAL GAS
AT 0.95 STREAM FACTOR
EQUIPMENT
NUMBER
-------------------
NAME
----------------------------------------------
SIZE
----------------------------------------
MATERIAL OF CONSTRUCTION
REMARKS
---------------------------------------------------------- ----------------------------------------------------------------------------------
COLUMNS
C-101A,B
CO2 ABSORBER
C-102A,B
CO2 STRIPPER
C-201A-F
MOLE SIEVE DRYER
C-202A-D
MERCURY ADSORBER
C-401
DEETHANIZER, TOP SECTION
BOTTOM SECTION
C-402
DEBUTANIZER
9
44
11
50
10
15
10
10
6.5
20
13
32
12
52
FT DIA
FT T-T
FT DIA
FT T-T
FT DIA
FT T-T
FT DIA
FT T-T
FT DIA
FT T-T
FT DIA
FT T-T
FT DIA
FT T-T
SHELL: C.S.
TRAYS:
SHELL: C.S.
TRAYS:
SHELL: C.S.
PACKING: MOLECULAR SIEVES
SHELL: C.S.
PACKING: HG ADSORBENT
SHELL: C.S.
TRAYS: C.S.
TRAYS: C.S.
TRAYS: C.S.
SHELL: C.S.
TRAYS: C.S.
9,387
108,621
41,571
56,322
26,820
21,000
BHP
BHP
BHP
BHP
BHP
BHP
2.5% Cr
2.5% Cr
3.5% Ni ALLOY; STAGE 3, C.S.
C.S.
C.S.
2.5% Cr
19,200
1,006
10,058
350
BHP
BHP
BHP
BHP
C.S.
304 SS
C.S.
C.S.
20,000
7.5
7,500
11
3,,000
11
3,000
11
3,000
11
200
1
8,200
52
3,800
SQ FT
MMBTU/HR
SQ FT
MMBTU/HR
SQ FT
MMBTU/HR
SQ FT
MMBTU/HR
SQ FT
MMBTU/HR
SQ FT
MMBTU/HR
SQ FT
MMBTU/HR
SQ FT
MMBTU/HR
SQ FT
MMBTU/HR
SQ FT
MMBTU/HR
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: TITANIUM
22 VALVE TRAYS, 24 INCH SPACING
25 VALVE TRAYS, 24 INCH SPACING
10 VALVE TRAYS , 24 INCH SPACING
16 VALVE TRAYS , 24 INCH SPACING
26VALVE TRAYS, 18 INCH SPACING
COMPRESSORS
K-301A,B
K-302A,B
K-303A,B
K-304A,B
K-305A,B
K-306A,B
K-307A,B
K-308A,B
K-309A,B
K-501A-D
START-UP/PEAKING TURBINE
GAS TURBINE
ETHYLENE COMPRESSOR
PROPYLENE COMPRESSOR
ELECTRICITY GENERATOR
METHANE COMPRESSOR STEAM
TURBINE
METHANE COMPRESSOR
FUEL GAS COLD BLOWER
FUEL GAS COMPRESSOR
AIR BLOWER
3 STAGES
3 STAGES
3 STAGES
ELECTRIC MOTOR
HEAT EXCHANGERS
E-101A-H
E-102A,B
FEEED GAS/AMINE OVERHEAD
EXCHANGER
AMINE TRIM COOLER
E-103A,B
LEAN/RICH AMINE EXCHANGER 1
E-104A,B
LEAN/RICH AMINE EXCHANGER 2
E-105A,B
LEAN/RICH AMINE EXCHANGER 3
E-106A,B
AMINE RECLAIMER
E-107A,B
STRIPPER REBOILER
E-108A,B
STRIPPER CONDENSER
E-201A,B
REGENERATOR GAS COOLER
E-301A,B
START-UP TURBINE STEAM
CONDENSER
22
7
21,000
133
18
AIR COOLER
AIR COOLER
AIR COOLER
PEP REVIEW 2003-15
Table 9 (Continued)
BASE LOAD LNG BY CASCADE REFRIGERATION
MAJOR EQUIPMENT
CAPACITY: 9,646 MILLION LB/YR (4,375,300 mt/y)
LIQUIFIED NATURAL GAS
AT 0.95 STREAM FACTOR
EQUIPMENT
NUMBER
-------------------
NAME
----------------------------------------------
SIZE
----------------------------------------
MATERIAL OF CONSTRUCTION
REMARKS
---------------------------------------------------------- ----------------------------------------------------------------------------------
HEAT EXCHANGERS (CONT.)
E-302A-X
E-303A,B
E-304A,B
E-305A,B
E-306A,B
E-307A,B
E-308A,B
E-309A,B
E-310A,B
E-311A,B
E-312A,B
E-313A,B
E-314A,B
E-315A,B
E-316A,B
E-317A,B
E-318A,B
E-319A,B
E-320A,B
E-321A,B
E-323A,B
E-324A,B
E-325A,B
E-326A-P
E-327A,B
PROPYLENE CONDENSER
32,000 SQ FT
42 MMBTU/HR
POWER GENERATOR STEAM
SQ FT
CONDENSER
79 MMBTU/HR
FEED GAS FIRST CHILLER
17,000 SQ FT
16 MMBTU/HR
FEED GAS SECOND CHILLER
30,000 SQ FT
25 MMBTU/HR
LNG/FUEL GAS EXCHANGER
SQ FT
0.6 MMBTU/HR
FEED GAS PRE-COOLER NO. 1
SQ FT
1 MMBTU/HR
FEED GAS PRE-COOLER NO. 2
SQ FT
20 MMBTU/HR
FEED GAS PRE-COOLER NO. 3
SQ FT
23 MMBTU/HR
FUEL GAS COLD RECOVERY NO. 3
SQ FT
1.2 MMBTU/HR
FEED GAS LIQUEFIER
SQ FT
69 MMBTU/HR
FUEL GAS COLD RECOVERY NO. 2
SQ FT
1 MMBTU/HR
LNG SUBCOOLER NO. 1
SQ FT
20 MMBTU/HR
FUEL GAS COLD RECOVERY NO. 1
SQ FT
1 MMBTU/HR
LNG SUSCOOLER NO. 2
SQ FT
19 MMBTU/HR
LNG SUBCOOLER NO. 3
SQ FT
20 MMBTU/HR
FUEL GAS COMPRESSOR INTERSTAGE
SQ FT
COOLER
1.7 MMBTU/HR
METHANE REFRIGERANT CONDENSER
SQ FT
74 MMBTU/HR
METHANE REFRIGERANT DESUPERHEATER
SQ FT
18 MMBTU/HR
METHANE REFRIGERANT COOLER
SQ FT
22 MMBTU/HR
METHANE COMPRESSOR INLET
SQ FT
COOLER
6 MMBTU/HR
ETHYLENE ECONOMIZER FIRST
25,000 SQ FT
DESUPERHEATER
12 MMBTU/HR
ETHYLENE ECONOMIZER SECOND
20,000 SQ FT
DESUPERHEATER
17 MMBTU/HR
ETHYLENE DESUPERHEATER
32,000 SQ FT
36 MMBTU/HR
ETHYLENE CONDENSER
30,000 SQ FT
32 MMBTU/HR
TURBINE STEAM CONDENSER
100 MMBTU/HR
SHELL: C.S.
TUBES: TITANIUM ALLOY
SHELL: C.S.
TUBES: TITANIUM ALLOY
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: ALUMINUM
PLATES: ALUMINUM
SHELL: ALUMINUM
PLATES: ALUMINUM
SHELL: ALUMINUM
PLATES: ALUMINUM
SHELL: ALUMINUM
PLATES: ALUMINUM
SHELL: ALUMINUM
PLATES: ALUMINUM
SHELL: ALUMINUM
PLATES: ALUMINUM
SHELL: ALUMINUM
PLATES: ALUMINUM
SHELL: ALUMINUM
PLATES: ALUMINUM
SHELL: ALUMINUM
PLATES: ALUMINUM
SHELL: ALUMINUM
PLATES:ALUMINUM
SHELL: NONE
TUBES: C.S.
SHELL: ALUMINUM
PLATES: ALUMINUM
SHELL: ALUMINUM
PLATES: ALUMINUM
SHELL: ALUMINUM
PLATES:ALUMINUM
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
TUBES: TITANIUM
19
AIR COOLER
PEP REVIEW 2003-15
Table 9 (Continued)
BASE LOAD LNG BY CASCADE REFRIGERATION
MAJOR EQUIPMENT
CAPACITY: 9,646 MILLION LB/YR (4,375,300 mt/y)
LIQUIFIED NATURAL GAS
AT 0.95 STREAM FACTOR
EQUIPMENT
NUMBER
-------------------
NAME
----------------------------------------------
SIZE
----------------------------------------
MATERIAL OF CONSTRUCTION
REMARKS
---------------------------------------------------------- ----------------------------------------------------------------------------------
HEAT EXCHANGERS (CONCLUDED)
E-401
E -402
DEETHANIZER OVERHEAD
CONDENSER
DEETHANIZER REBOILER
E-403
DEBUTANIZER REBOILER
E-404
NATURAL GASOLINE COOLER
E-405
DEBUTANIZER CONDENSER
E-406
FUEL GAS/LPG EXCHANGER
E-501A,B
HRSG SUPSERHEATER TUBES
E-502A,B
HRSG STEAM GENERATION TUBES
E-503A,B
HRSG ECONOMIZER TUBES NO. 1
E-504A,B
HRSG ECONOMIZER TUBES NO. 2
FURNANCES
F-201A,B
F-501
DRYER REGENERATION FURNACE
AUXILIARY BOILER
2,800 SQ FT
MMBTU/HR
7,200 SQ FT
MMBTU/HR
4,800 SQ FT
MMBTU/HR
1,600 SQ FT
MMBTU/HR
24,000 SQ FT
MMBTU/HR
1,000 SQ FT
MMBTU/HR
31,000 SQ FT
64 MMBTU/HR
110,000 SQ FT
175 MMBTU/HR
32,000 SQ FT
52 MMBTU/HR
32,000 SQ FT
108 MMBTU/HR
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: C.S.
TUBES: C.S.
SHELL: NONE
TUBES: 5% Cr ALLOY
SHELL: NONE
TUBES: C.S.
SHELL: NONE
TUBES: C.S.
SHELL: NONE
TUBES: C.S.
18 MMBTU/HR
375 MMBTU/HR
TUBES: C.S.
TUBES: 5% Cr ALLOY
KETTLE TYPE
AIR COOLER, FINED TUBES
AIR COOLER
FINNED TUBES IN G-501
FINNED TUBES IN G-501
FINNED TUBES IN G-501
FINNED TUBES IN G-501
TANKS
T-101
T-601A,B
AMINE SURGE DRUM
LNG STORAGE
T-602
T-603
T-604A-F
T-605A-D
T-606A,B
T-607A,B
T-608
LPG
NATURAL GASOLINE
PROPYLENE REFRIGERANT
ETHYLENE REFRIGERANT
FRESH WATER
DEMINERALIZED BOILER WATER
DIESEL FUEL
42,000 GAL
37,800,000 GAL
C.S.
9% Ni ALLOY
840,000
840,000
50,000
50,000
5,000
5,000
3,000
GAL
GAL
GAL
GAL
GAL
GAL
GAL
C.S.
C.S.
C.S.
C.S.
C.S.
C.S.
C.S.
1,000
800
600
900
42,000
GAL
GAL
GAL
GAL
GAL
C.S.
C.S.
C.S.
C.S.
C.S.
42,000 GAL
C.S.
25,000 GAL
C.S.
15,000 GAL
25,000 GAL
C.S.
3.5% Ni ALLOY
8 DAYS PRODUCTION, DOUBLE WALL WITH PERLITE
INSULATION
14 DAYS PRODUCTION
14 DAYS PRODUCTION
FOR EMERGENCY GENERATORS
PRESSURE VESSELS
V-101A,B
V-102A,B
V-103A,B
V-201A,B
V-301A,B
V-302A,B
V-303A,B
V-304A,B
V-305A,B
WATER KNOCK OUT DRUM
CO2 ABSORBER OVERHEAD DRUM
FUEL GAS SEPARATOR
CONDENSATE DRUM
HIGH PRESSURE PROPYLENE
FLASH DRUM
MID PRESSURE PROPYLENE
FLASH DRUM
LOW PRESSURE PROPYLENE
FLASH DRUM
RAW LPG SEPARATOR
HIGH PRESSURE ETHYLENE
FLASH DRUM
20
PEP REVIEW 2003-15
Table 9 (Concluded)
BASE LOAD LNG BY CASCADE REFRIGERATION
MAJOR EQUIPMENT
CAPACITY: 9,646 MILLION LB/YR (4,375,300 mt/y)
LIQUIFIED NATURAL GAS
AT 0.95 STREAM FACTOR
EQUIPMENT
NUMBER
-------------------
NAME
----------------------------------------------
SIZE
----------------------------------------
MATERIAL OF CONSTRUCTION
REMARKS
---------------------------------------------------------- ----------------------------------------------------------------------------------
PRESSURE VESSELS (CONCLUDED)
V-306A,B
MID PRESSURE ETHYLENE
FLASH DRUM
LOW PRESSURE ETHYLENE
FLASH DRUM
HIGH PRESSURE METHANE
FLASH DRUM
MID PRESSURE METHANE
FLASH DRUM
LOW PRESSURE METHANE
FLASH DRUM
LNG FLASH DRUM
RAW LPG FLASH DRUM
DEETHANIZER REFLUX DRUM
DEBUTANIZER REFLUX DRUM
CONDENSATE DRUM
HRSG STEAM DRUM
DEAERATOR
AUXILIARY BOILER STEAM DRUM
V-307A,B
V-308A,B
V-309A,B
V-310A,B
V-311A,B
V-401
V-402
V-403
V-501A,B
V-502A,B
V-503A,B
V-504A,B
20,000 GAL
3.5% Ni ALLOY
15,000 GAL
3.5% Ni ALLOY
30,000 GAL
3.5% Ni ALLOY
15,000 GAL
304 SS
7,000 GAL
304 SS
30,000
800
800
2,200
7,000
10,000
12,000
10,000
GAL
GAL
GAL
GAL
GAL
GAL
GAL
GAL
304 SS
C.S.
C.S.
C.S.
C.S.
C.S.
C.S.
C.S.
MISCELLANEOUS EQUIPMENT
M-301A,B
M-302A,B
M-501
SCREEN
AIR FILTER
STACK
15 FT DIA, BASE
50 TALL
C.S.
HEAT RECOVERY STEAM GENERATOR
DEMINERALIZER (OR
DESALTATION UNIT)
100 GPM
C.S.
SECTION
---------------
OPERATING BHP
-------------------------
SPECIAL EQUIPMENT
S-301
COLD BOX
PACKAGED UNITS
G-301
G-501A,B
G-502A,B
PUMPS
100
200
300
400
500
600
700
OPERATING
------------------4
0
4
6
4
20
8
SPARES
-------------8
1,620
4
6
2
10
2
200
60
2,000
9,770
17,600
21
PEP REVIEW 2003-15
Table 10
LNG BY CASCADE REFRIGERATION
UTILITIES SUMMARY
3
CAPACITY: 525 MT/H (9,646 MILLION LB/Y OR 602 MILLION SCF/SD OR 17.056 M /SD)
LNG
AT 0.95 STREAM FACTOR
Superheated Steam,
Fuel Gas,
a
Section
MM Btu/H
(900°F), lb/H
Electricity, kw
c
76,000
1,200
d
200
36
300
1,640
x
e
400
Battery Limits Total
900 psig, 482°C
saturated, lb/H
b
100
500
Steam, 100 psig
f
500,000
7,600
12,000
x
x (start up)
______
_______
1,500
1676
88,000
500,000
10,300
g
600
595
700
11,300
h
800
x
900
x
x
74
350
All Other
Total Offsite plus Other
74
______
_______
12,230
Export
_____
______
_______
13,000
Total Average Use
1,750
88,000
500,000
35,530
a
Lower heating value.
b
Average; 194,000 lb/H when processing maximum CO2 content feed gas.
c
At average CO2 content raw feed gas; 2,000 kw at maximum CO2 feed gas.
d
x indicates consumption is included in “all other” category.
e
At 27°C (81°F).
f
Maximum.
g
LNG and other product pumps average running 10% of the time. Peak load is 5,950 kw.
h
Spare sea water pumps off; 16,400 kw with spare pumps running.
22
PEP REVIEW 2003-15
Section 100 – Amine Unit
Raw natural gas feedstock at an average 17°C (63°F) and 650 psia enters one of the two
trains at its amine unit (Figure 1, sheet 1) where acid gas (CO2 and H2S) is removed following
free water separation (V-101). The feed gas is then heated to 36°C (97°F) by exchange in E-101
with treated gas prior to entering the CO2 absorber (C-101) and being contacted with solvent, in
our case, monoethanol amine (MEA). (For reasons of consistency, a generic MEA for CO2
removal system is shown; similar configurations could apply to diethanol amine (DEA), diglycol
amine (DGA), methyldiethanol amine (MDEA) and other solvents.) Using 20 wt% MEA (0.35 mol
CO2/mol MEA), CO2 is removed to under 50 ppm. On leaving the absorber at 50°C (122°F), the
CO2 free natural gas heats the feed gas while being chilled to 23°C (73°F). Entrained liquid MEA
is separated from the treated gas (V-102). The gas will next be dried.
Meanwhile, MEA solution is charged to the absorber at about 46°C (115°F). The CO2 rich
solution leaves the bottom of the absorber at about 38°C (100°F) and is reduce in pressure to
about 100 psia. Light hydrocarbon gases dissolved in the MEA solution separate in V-103. The
gases are scrubbed with lean MEA solution and sent to the fuel gas system. The rich solution is
heated by a series of exchangers (E-103 to E-105) with 122°C (252°F) lean solution from the
bottom of the CO2 stripper (C-102). CO2 is stripped from the MEA solution with vaporized lean
solution either from storage (E-106) or from a reboiler (E-107). The stripper is refluxed with a
small amount of make up water (about 18 lb/H) added at the cooler (E-108). CO2 rich gas is
vented at 54°C (129°F).
Section 200 – Dryers and Mercury Removal
The natural gas feed is next dried by molecular sieve beds (C-201). Of the three beds, one
is in service, one is held ready for service and one is being regenerated by stripping with hot (F201) natural gas. Water condensed (E-201) from the wet regeneration gas flows to the amine
unit (Section 100) to supply make up water.
Many natural gases around the globe are known to contain traces of mercury that is harmful
to aluminum alloy in the downstream cryogenic section. Mercury removal is necessary for safety
due to corrosion of cryogenic aluminum heat exchangers. Adsorption beds are used for trace
mercury removal. The “mercury free” gas then proceeds to the liquefaction section.
Section 300 -- Liquefaction
Liquefaction is conducted in four steps: (1) prechilling by heat exchange with pure liquid
propylene refrigerant, (2) separating raw liquid petroleum gas, (3) cryogenic cooling and
condensing using successively colder pure liquid ethylene and methane refrigerants and (4)
flashing methane vapor from the liquid. For economy, the refrigeration loops are cascaded, i.e.,
seawater cools the high pressure propylene, propylene cools the high pressure ethylene and
ethylene cools the high pressure methane. The prechilling occurs in large kettle type
exchangers. The cryogenic exchange occurs in aluminum plate-fin exchangers located in an
insulated cold box.
Dry treated feed gas enters the chilling train of the gas liquefaction section to be cooled by
two propylene refrigeration levels in kettle type heat exchangers. The feed gas at 23°C (73°F)
and 610 psia is first cooled by exchange with propylene refrigerant at –4°C (24°F) in E-304 to –
1°C (30°F) and again to –31°C (-24°F) with –32°C (-25°F) propylene (E-305) (Figure 1, sheet 2).
Much of the C3+ raw liquid petroleum gas is condensed and separated in knock out drum V-304.
The liquid is separated into saleable byproducts in Section 400 while the vapor, now meeting the
23
PEP REVIEW 2003-15
natural gas C5+ heavies specification, proceeds to the cold box for deep refrigeration via ethylene
and methane refrigeration.
The three-stage propylene compressor does the propylene refrigeration work. Propylene is
flashed to three successively lower pressures (V-301 to V-303) to provide three refrigerant levels.
Both the propylene and ethylene compressors (K-304 and K-303 respectively) are driven by a
large gas turbine (K-302). Heat is recovered from the turbine exhaust (about 545°C, 1020°F) in
Section 500. The ethylene compressor stages 1 and 2 share the same case; the third stage is
separate. The train is started up with a steam turbine (K-301). The driving steam is condensed
by exchange with seawater (E-301) that is chlorinated (G-301). Seawater also cools the high
pressure propylene refrigeration loop (E-302) and also condenses (E-303) steam from the
electricity generator (K-305).
The bulk of the liquefaction occurs in the ethylene refrigeration cycle while the methane
refrigeration cycle provides the sub cooling duty. Gas enters the cold box (Figure 1, sheet 3) at 32°C (-26°F) and sub-cooled liquid leaves the box at -152°C (-241°F) and 36.5 kg/cm2-a, (520
psia). In the cold box, the gas is cooled by a series of heat exchanges in aluminum plate-fin
exchangers (E-307 thru E-316) with successively colder refrigerant loops and cold fuel gas.
As mentioned, three-stage compressor K-303 drives the ethylene refrigeration system. Most
of the heat of compression is removed by cooling with propylene between stages 2 and 3 (E-323
and E-324) and at high pressure (E-325 and E-326). Ethylene is flashed at three successive
pressures (V-305 to V-306) to obtain the three temperature levels. The methane compressor (K307) is steam driven (K-306). Seawater condenses the steam (E-327). Stages 2 and 3 are in
one case. The inlet temperature is heated to –101C (-150F) by exchange with a slip stream of
the natural gas (E-321) to allow the compressor to be made of less expensive metal. Most of the
heat of compression is removed by exchange with ethylene (E-318 to E-320). Again the
methane is flashed to three successively lower pressures (V-308 to E-310) to obtain three
temperature levels.
On leaving the cold box, the sub-cooled liquid flashes at 1.3 kg/cm2-a (19 psia) in flash drum
V-311 prior to entering to LNG storage tank at 1.1 kg/cm2-a (16 psia). The flash tank provides
protection to the storage tank and also pre-flashes most of the nitrogen prior to storage. For high
nitrogen content feedstocks, perhaps pre-flashing at 4-5 kg/cm2-a (55-70 psia) would be
preferable prior to flashing at 1.3 kg/cm2-a (19 psia).
The flashed gas and vapor boiled off from the tanker during loading are combined and
compressed to 1.9 kg/cm2-a (27 psia) prior to recovering the “cold” by exchange with the gas in
the cold box (E-314, E-312, E-310, E-307). Flashed gas after cold recovery at about –31°F (35°C) is being compressed by K-309, using electric motor drivers to about 21-23 kg/cm2-a (299330 psia) and used as turbine fuel. Interstage (E-317) to about 38-49°C (100-120°F) is provided
depending upon the ambient air temperature. The flash temperature (first stage of methane
refrigeration) is continuously adjusted to provide the correct amount of fuel gas needed for the
gas turbines (K-302).
Section 400 – LPG Fractionation
The raw LPG at –32°C (-25°F) 594 psia is preflashed to remove light gases (V-401, Figure
1, sheet 4). Ethane rich gas and occasionally distillate are separated from the flashed liquid by
distillation (C-401). Overhead vapor is partially condensed (E-401) and vapors separated from
the liquid (V-402), which is normally all returned to the column as reflux. Steam (nominal 50 psig)
to reboiler E-402 provides energy for the separation.
24
PEP REVIEW 2003-15
The C3 and heavier bottoms are fractionated in debutanizer C-402 to produce C3/C4 LPG
distillate and natural gasoline bottoms products. Overhead vapor is usually totally condensed (E403) with reflux drum V-403 handling surges. The LPC distillate product is cooled to about 35°C
(95°F) by exchange with the fuel gas product (E-406) and stored offsite (Section 600). The
natural gasoline product is cooled by air cooler E-404 to about 49°C (120°F) depending upon the
ambient temperature and stored offsite (Section 600).
Section 500 – Heat Recovery System
Heat in the turbine exhaust (From K-302, sheet 2) is recovered by steam generation (Figure
1, sheet 5). Steam at 63 bar-g and 480°C (900 psig and 900°F) would represent a class break
point and thus was selected. Steam at 63 kg/cm2-g (900 psig) is fed to four turbines driving the
two 14,300 kw methane compressors and two nominal 20,000 kw power generators. Steam at
about 8.0 kg/cm2-g (114 psig) is uncontrollably extracted as a heat source for the MEA
regeneration (Figure 1, sheet 1) and LPB fractionation (Figure 1, sheet 4).
Turbine exhaust at about 545°C (1020°F) and 12 in water-g flows to heat recovery steam
generator (HRSG) G-501. The exhaust gas leaves the HRSG at 182°C (360°F) and flows to
stack M-501. Condensate at about 40°C (104°F) from condensate drum V-501 is pumped at
about 115 psia through HRSG coils E-504 and preheated to about 145°C (290°F). The
condensate is then deaerated (V-502) along with demineralized or desalted make up water from
G-502. The deaerated boiler feed water is pumped at 1115 psia through the HRSG at coils E503 into steam drum V-502 at 965 psia. Liquid water from the steam drum circulates through coil
E-502. Steam from V-502 is superheated to 482°C (900°F) and flows to users at 915 psia.
For start up and peak loads, auxiliary boiler F-501 with steam drum V-504 generates 915
psia, 482°C (900°F) superheated steam. The auxiliary boiler is a packaged unit.
Section 600 – Storage
Storage is provided for LNG product, LPG and natural gasoline byproducts, amine and the
propylene and ethylene refrigerants. Two LNG storage tanks, each 145,000 M3 (900,000 bbl) are
double walled, insulated and further contained by concrete walls. Each tank is installed with four
3,000 M3/hr send out pumps. LPG is stored in two 20,000 bbl tanks with a refrigerated vapor
recovery system. Natural gasoline is stored in two fixed roof 20,000 bbl tanks also having vapor
recovery.
Make up amine is stored in two diked 1,000 bbl tanks. Make up propylene is stored in six
bullets each holding 50,000 gal under about 275 psia pressure. Make up ethylene is stored in
four refrigerated bullets each holding 50,000 gal at about 265 psia and –34°C (–30°F). Also,
fresh water and demineralized boiler feed water are each stored in two 5,000 bbl tanks. A 3,000
bbl diesel fuel tank is provided to fuel the two 1,600 kw emergency power generators.
Section 700 – Marine System
A 1,500 meter long jetty with appropriate support structures is provided for loading LNG onto
ships (not included). The jetty pipe rack has two LNG lines each 30 in. in diameter. Two vapor
recovery lines are 20 in. diameter each. LPG and natural gasoline are loaded through two 10 in.
and two 8 in. diameter lines respectively. Two refrigerant make up lines of 3-1/2 nickel alloy are
4 in. diameter. Two 4,000 volt electric cables for the sea water pumps each carry 8,000 kw.
The cost of the seawater coolant (shown in Section 300, Figure 1,sheet 3) is also accounted
as part of the marine system. Two seawater intake suction lines buried 15 ft below seawater
25
PEP REVIEW 2003-15
surface are 7 ft I.D. coated carbon steel pipes. Two pumping structures have 5 pumps on each
structure. Driven by 2,500 Hp motors, each can pump 45,000 gpm. Also included are a pump
house and a pump control center. The seawater is screened, chlorinated, filtered and transferred
via two 7 ft diameter concrete pipes laid under the water. The cooling seawater is returned to the
sea through two 1000 meter water discharge lines, 7 ft each, made of coated concrete laid under
water.
Section 800 – Relief System
Two flare lines are provided. The cold flare line is 24 in. in diameter, 500 m (1,500 ft) long,
and insulated. The larger warm flare line is 30 in. diameter and the same length. The two flare
stacks are 50 m (150 ft) tall.
Section 900 – General Offsites
General offsites includes roads and general infrastructure and such items as process and
sanitary drain systems, telecommunications, control building including a digital control system
and auxiliary, craft shops, control laboratory. Fire protection and ballast water systems are also
general offsite items. Emergency power is provided by two 1,600 kw diesel generators.
A nitrogen generation package supplies 40 tpd (800 scfm) of 99.5% purity nitrogen. Back up
is provided by 800 ton (1,000 m3) of liquid nitrogen. The evaporator uses sea water. Instrument
air is compressed at 800 scfm by two 250 kw (335 Hp) compressors.
PROCESS DISCUSSION
Our grass-roots LNG plant is conceptually designed using ambient and cooling air and water
temperatures for a subtropical region (Table 5).
Our general turbo machinery configuration is two half plant capacity turbines in parallel, one
in each train. Thus a forced or planed shut down of a gas turbine will result in a 50% reduction in
LNG production. Whereas in a propane pre-cooled, mixed refrigerant system, the drivers are
arranged in series. For example, in a typical 4.5 million mt/y LNG plant, a single frame 7 turbine
is dedicated to low pressure and mid pressure mixed refrigerant service and a single frame 7 is
used for high pressure mixed refrigerant and propane compression. Then a forced or planed
shut down of any turbine results in 100% production loss. On this basis the common on stream
factor is elevated from 340 days per year (0.93) to 347 days per year (0.95). The ConocoPhillips
process utilizing 6 turbines will have some advantage in this respect since an outage of a single
turbine will still allow production in excess of 60-65% of the name plate capacity.
Plant Capacity
The capacity range is determined by market outlook as mentioned in the LNG Market
Overview section. The ultimate plant capacity is determined by matching the performance of
available turbine driver configurations. On a generic basis we decided upon two 7FA GE turbines
or equivalent, one in each train. A gas turbine is used for propylene and ethylene refrigeration
while steam turbines were used for methane refrigeration and power generation. In our design, a
forced or elective shut down of a given gas turbine will reduce production by 50% unlike the
ConocoPhillips six half capacity frame 5D concept which could allow perhaps over 60-65%
production in the event of a forced shut down of a gas turbine. Nevertheless, we believe that
gains from reducing the number of turbines out weigh the reduction in flexibility.
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PEP REVIEW 2003-15
The gas turbines are equipped with start up turbines running on steam from the auxiliary
boiler. The start up turbine could be used to boost capacity when the gas turbine is de-rated due
to high ambient temperatures. Furthermore, operation of the auxiliary boiler (250,000 lb/hr 900
psig/482°C (900°F) steam) will allow one to maintain higher production at all times.
As shown in the design basis (Table 5), the average ambient temperature is assumed to be
21°C (70°F). Since the speed of the gas turbine is fixed, the maximum sea water temperature,
27°C (81°F) in our case, sets the discharge pressure of the propylene refrigerant. No credit
could be taken for lower seawater temperatures. Since the average ambient air temperature in
August is 27°C (81°F), gas turbine performance declines by 4.5% with respect to the average.
Running the startup booster steam turbine could make this deficiency. Conversely, the average
January temperature is 14.5°C (58°F); thus turbine performance increases by 5%. The average
daily variance of 7°C (12.6°F) represents about 5% of turbine performance.
Feed Gas
The feed gas composition (Table 6) does not necessarily represent the gas composition in
any subtropical location. This is a lean gas (92.1 vol% methane) also low in CO2 and nitrogen
that would be a good representative composition for a majority of global locations. Our
conceptual design will not be affected by C2+ contents under 15 vol% nor by CO2 or nitrogen
contents under 3.0 vol% each. This design would be applicable to the great majority (over about
80% of world’s stranded gas) of the potential market. High nitrogen content, say above 3.0 vol%,
may require a second, higher pressure flash stage. Any significant H2S content will require
incineration of the CO2 vent gas and, at some locations such as Qatar, the installation of sulfur
recovery units.
Over time as a gas field ages the gas supply pressure decreases and the CO2 content
increases. In our case the pressure is initially assumed to be about 800 psia. Our 650 psig
supply pressure and 0.47 vol% CO2 is the mid-life pressure after 10-15 years and average CO2
content.
Acid Gas Removal
MEA is assumed to be more economical solvent than MDEA for CO2 contents under 1.5
vol%. Each train has an amine unit. A single amine unit serving both trains could be a viable
choice as well. However we judged that the savings in capital investment versus the extra
reliability and flexibility of two units favored selection of the dual train system. Since the H2S
content is very minimal, acid gas could be safely incinerated after stripping from the MEA
solution. The CO2 content is reduced to under 50 ppm. Thus DEA, MDEA or another solvent
could be selected depending upon the operational preference and general business
considerations of the LNG producer. On a global basis, about 35% of LNG plants are using
MEA.
In our design, the amine unit is sized for 1.2 vol% CO2 in the raw feed gas even though the
average is 0.47 mol% CO2. The maximum amine circulation rate is 725 gpm per train based on
1.2 mol% CO2. Circulation based on the average CO2 content is 283 gpm per train.
No CO2 is being recovered from amine unit or from the flue gases. However, a high CO2
greenhouse gas emission tax could make CO2 recovery from the amine unit an economically
viable option, especially if no H2S is present. The CO2 vent gas is contaminated with
hydrocarbons, and incineration of the vent gas in the auxiliary boiler could at time become a
viable option especially when traces of H2S are present.
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PEP REVIEW 2003-15
Nitrogen Removal
Since the nitrogen content of the gas is less than 1.0 vol% (0.3 vol%), no particular nitrogen
removal is needed. Flashing fuel gas at 1.3 kg/cm2-a (19 psia) in the process and flashing from
storage at 16 psia diverts the bulk of the nitrogen into the fuel gas system. In our case, 67% of
the nitrogen is flashed to the fuel gas. This however would have no impact on the operation of
the gas turbines.
Gas drying and mercury removal
Molecular sieve gas drying at 23°C (73°F) is used to protect the cryogenic exchangers and
piping from any potential freezing.
Many natural gases around the globe are known to contain traces of mercury that is harmful
to aluminum alloy in the downstream cryogenic section. Mercury in light hydrocarbons is very
briefly reviewed in PEP Report 29G Ethylene Plant Enhancement (April 2001) and in PEP
Review 91-1-4 Removal of mercury from ethylene plant feedstock and cracked gas streams (July
1992). Mercury removal is necessary for safety due to corrosion of cryogenic aluminum heat
exchangers. To remove mercury, the dry natural gas is passed through mercury adsorption
beds. Carbon impregnated with sulfur (Calgon Carbon) or silver on alumina is used for vaporous
hydrocarbons. Experienced contractors should be contacted about adsorbents and process
configurations suitable for specific applications.
Refrigeration Cycles
The refrigeration cycle was simulated on the basis of the flow diagram shown in Figure 1,
sheets 2 and 3, using efficiencies for the compressors (Table 5) and turbines based upon
preliminary flow data. The final flow rates are only very slightly different. Because the flow
limitation on the assumed propane (used the ConocoPhillips process) compressor becomes a
bottleneck, propylene refrigeration is selected in order to shrink the suction volume by 22%.
Features of the refrigeration system include:
•
Two, one for each train, GE Frame FA or equivalent, propylene and ethylene
refrigeration gas turbines
•
Two GE or equivalent steam turbines in a closed methane loop
•
Two 50% HRSG units generating steam at 63 kg/cm2- g, 480°C (900 psig/900°F)
•
Two 50% plant capacity GE steam turbines operating on a single steam cycle with a
single uncontrolled extraction of steam at a nominal 8 kg/cm2-g (115 psig). Turbine
steam is condensed at 41°C (106°F), 1.1 psia by seawater at 27°C to 32°C (to 81°F
90°F).
•
All the refrigeration loads, totaling about 1,000 MM Btu/H, are calculated on the basis of
seawater at 27°C (81°F) and rejecting heat at 35°C (95°F).
Selection of the gas turbine is not necessarily easy for any location and situation. Steam
turbine drivers, electric motor drivers, combined cycle gas turbine–steam turbine could all be
considered. Our design basis is a gas turbine coupled to an exhaust gas heat recovery steam
generator to drive both the propylene and the ethylene compressors. A start up booster steam
turbine is on the same shaft. As a result of the design basis and the preliminary recommendation
of GE-Novo Pingone, the methane compressor is divided into two cases on the same shaft driven
at 8,500 rpm by a steam turbine. The ethylene compressor is also split.
28
PEP REVIEW 2003-15
In each train, a General Electric FA7 or equivalent gas turbine drives two refrigeration
compressors in series: about 31,900 kw ethylene refrigeration and 41,500 kw propylene
refrigeration. A 7,000 kw (9,400 Hp) steam turbine is used for start up as well for load boosting
during periods of high ambient temperatures. The estimated gas turbine performance is
summarized in Table 11. The compressors have 6 wheels and two side loads. The gas turbine
name plate capacity at 21°C (70°F) is 75,990 kw (101,900 Hp). About 660 kw (890 Hp) or 1.6%
of power to the propylene compressor is lost through speed reduction thus the real net power
available is 75,900 kw (101,000 Hp). De-rating due to refrigeration losses leaks about 230 kw
(300 Hp). Mechanical deterioration of 2.3% would suggest a conservative net load of 73,400 kw
(98,430 Hp) at 21°C (70°F).
The gas turbine fuel consumption is based on 10,730 Btu/h per kw LHV (8,000 Btu/h per Hp)
as suggested by GE. This is the basis for the fuel gas consumption and the set point of the flash
gas, which controlled by the terminal temperature of the sub-cooling section in the cold box,
calculated at –152°C (-242°F). The gas turbine is at the core of the design philosophy, and
clearly demonstrates that availability of turbo machinery is a key factor in the establishing the
design basis.
A Frame 5 variable speed turbine has higher NOx emission (about 100 ppm) that in some
locations would require a de-NOx process. The 25 ppmv NOx emission with the Frame 7 turbine
does not require NOx removal.
Based on our cost estimate, we are considering that the combined cycle mode, as
suggested, represents a lower capital investment per ton of LNG compared with using exclusively
gas turbine drivers commonly practiced in the industry. Furthermore, the combined cycle mode
presents the higher efficiency, lower CO2 greenhouse gas emissions and probably all together
the most economical configuration, at least in our case. Some locations may need to obtain fresh
water for the combined cycle’s steam production and for employees from a packaged
desalination unit instead of demineralizing water. Integration of desalination and power
generation is very commonly practiced.
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PEP REVIEW 2003-15
Table 11
ESTIMATED PERFORMANCE OF GAS TURBINE (FRAME 7)
Load Condition
Base
Base
Base
Base
4.
4.
4.
Inlet Loss
in H2O
4.
Exhaust Pressure Loss
in H2O
12.0
Ambient Temperature
°C (°F)
15 (59)
Fuel Type
Methane
12.0
12.0
20 (68)
Methane
12.0
28 (82.4)
Methane
40 (104)
Methane
Fuel LHV
Btu/lb
21,515
21,515
21,515
21,515
Fuel Temperature
°C (°F)
26.7 (80)
26.7 (80)
26.7 (80)
26.7 (80)
Output
Hp
106,630
102,690
96,430
87,130
Heat Rate (LHV)
Btu/Hp-H
7,925
8,000
8,140
8,410
Heat Cons. (LHV)
MM Btu/H
845.0
821.5
784.9
732.8
2234
2179
2090
1959
549 (1020)
553 (1028)
560 (1040)
569 (1057)
3
Exhaust Flow
x10 lb/H
Exhaust Temperature
°C (°F)
Exhaust Loss
in H2O
Application
12.0 @ ISO conditions
Mechanical Drive
Combustion System
DLN Combustor
a
Emissions
NOx at 15% O2
ppmv
25.
25.
25.
CO
ppmv
15.
15.
15.
15.
Source: GE
a
Emission information is based on GE’s recommended measurement methods. NOx emissions are corrected to 15%
O2 without heat rate correction and are not corrected to ISO reference condition per 40CFR 60.335(a)(1)(i). NOx levels
shown will be controlled by algorithms within the SPEEDTRONIC control system.
The side loads of the propylene refrigeration, -4°C (+25°F) and +17°C (63°F), were
determined by turbo machinery analysis and wheels configuration provided by Novo Pingone. A
different machinery design could result in different side loads conditions.
The ethylene and methane refrigeration cycles use plate-fin core aluminum exchangers
packed into a cold box with total weight of not to exceed 1,000 ton and no less than 2 cold boxes.
Budgetary quotations and technical recommendations by Chart Industries, La Cross, Wisconsin,
and their agent, SME Associates, Houston, are incorporated in the design of the cold box. The
aluminum exchangers dictate mercury removal beds prior to the gas chilling train as discussed
previously. Refinery grade propylene, typically 95 wt% propylene and 5 wt% propane, could be
acceptable as the refrigerant.
Heat Rejection
The seawater circulation piping was sized to a maximum velocity of 12 ft/second due to
concerns of erosion. At the current design, the maximum temperature rise is set at 5°C (9°F)
30
PEP REVIEW 2003-15
however it is recognized that in some locations a maximum rise of 3.5-4.0°C (6-7°F) is allowed. If
the two spare sea water pumps would be run, the total circulation will increase by 25%, the
temperature rise will be controlled at 4°C (7°F) and the velocity at the pipes will reach 11.6 ft/sec.
All heat is rejected to sea water. The constant speed gas turbine drive calls for a propylene
refrigerant pressure of 220 psia based on a maximum sea water temperature of 27°C (81°F).
Propylene is condensed at 35°C (95°F) and turbine steam condenses at 41°C (105°F). The
maximum ambient air temperature along the Eastern Mediterranean coast, for example, is about
38°C (100°F) and this sets the discharge pressure of the propylene compressor if air cooling is
applied. For an air cooler condenser, this would require a minimum propylene condensing
temperature of about 50-52°C (122-125°F) and a discharge pressure of about 300-310 psia.
Compression power increases by 7,000-7,500 kw per gas turbine and booster steam turbine train
of 73,000 kw thus reducing LNG production by 10%. The avoided sea water circulation per
refrigeration train would be 110,000 gpm, about 4,000 kw of circulation power. On the other
hand, the air cooler fans for 500 MM Btu/hr heat rejection per production train would consume
4,000 kw as well. The economics of this design decision are briefly discussed below in the
section “Capital Cost”.
Steam and power generation
One HRSG (Heat Recovery Steam Generation) system is used per train (Figure 1, sheet 5).
Each HRSG is dedicated to one gas turbine. About 250,000 lb/hr of steam at 900 psig, 482°C
(900°F) is generated. Although a single 500,000 lb/hr steam HRSG is possible, we believe the
extra duct work, hot by pass stack and on stream time considerations could not be justified for
the marginal savings. The auxiliary start up boiler, 250,000 lb/hr capacity, will provide the back
up steam needed for methane compression in the event the HRSG is by passed. The auxiliary
boiler also can provide boosting capacity for the gas turbine during high ambient temperatures.
Two nominal 20,000 kw, 22,000 kw design capacity, power generators are assumed, thus
dependence on the local electricity grid is avoided in the event one generator is forced to shut
down (Figure 1, sheet 2). Under the normal scenario, the facility is self sufficient in electric
power. The generator size allows for 13,000 kw of power to be exported to nearby users,
probably under an interrupting power purchase contract. If sale of excess power becomes a very
valuable source of revenue, a steam pressure of 1,500 psig should be given consideration.
The power generator steam turbine also produces the low pressure steam consumed in the
amine and LPG Fraction sections (Figure 1, sheets 1 and 4). Uncontrolled extraction steam at
about 100 psig is higher than the 50 psig steam the process requires. As the flow rate slows
down, the extracted steam pressure decreases. Controlled extraction is very expensive and not
required for this application.
LPG Fractionation
Liquefied petroleum gas (LPG) removal is required to meet the LNG specifications for C5+
heavies and heating value. Since the feed gas is very low in nitrogen and CO2 while lean in C2+,
a generic LPG recovery is selected and the nitrogen stripping step is avoided.
Ethane separation from the raw LPG is not incorporated in our design. Instead, ethane rich
gas from LPG fractionation is routed to the turbine fuel system. This is due to the fact that only a
small quantity of ethane must be recovered to meet the LNG specifications. Our higher heating
value of the LNG is 1,091 Btu/scf, just below the specification maximum of 1,100 Btu/scf, and
essentially unchanged from the raw natural gas feedstock (1,095 Btu/scf). Our recovery of
31
PEP REVIEW 2003-15
ethane and propane is only 3.4 wt% and 12.2 wt% respectively. Recovery of ethane and
propane for sale as petrochemical feedstock is easier to do during regasification of the LNG at
the receiving terminal.
The pressure of the deethanizer column is set to allow at least sufficient fuel gas pressure to
the gas turbines. In the event that the LPG unit is forced to shut down, crude LPG after flashing
light ends, can be routed to storage or to flare.
For rich gas of 85 vol% methane, the LPG production rate may increase by a factor of four,
which would then control the design of the LPG fractionation section.
LNG storage and loading capacity
LNG storage and loading capacity is a key investment item that will change drastically
depending on location and LNG delivery contract. For an location where the bulk of the product
is delivered to a market a relatively short distance, about 2,000 miles, two LNG tanks (145,000
M3, 900,000 bbl) seem appropriate considering the maximum capacity of LNG tankers and the
average tanker capacity of 750,000 bbl. We assume the double containment tanks with an outer
shell of concrete would be 25-30% more expensive than single containment tanks. However
double containment reduces plot area if dikes can be avoided and also increases safety. The
basis of storage capacity is nominal 8 days flexible capacity. If more marketing flexibility is
required, to reach farther markets for instance, a third LNG tank could become reasonable. The
bottom line is simply stated: just like with turbine selection, storage capacity, flexibility, safety and
on stream factor are inter-related to capital investment.
LNG tanker loading is assumed to take under 16 hours. On this basis, the evacuation rate is
9,000 M3/hr per tank, using 7 bars pressure drop (40,000 gpm, 100 psi). Each tank is equipped
with four (3+ spare) 3,000 M3/hr, 750 kw pumps. All vapors displaced from LNG tankers during
loading, 12-16 hours per send-out cycle, is recycled to LNG storage. Heat leakage into the
tankers is accounted as general refrigeration losses.
Design Cases
The gas turbine limits the plant’s feedstock rate. Allowing for losses due to gears,
refrigeration losses and mechanical deterioration discussed previously, the maximum power at
21°C (70°F) ambient temperature is 98,430 Hp.
The propylene compressors and stage 3 of the ethylene are made of conventional carbon
steel. Stages 1 and 2 of the ethylene compressors are made of fine grain carbon steel (3.5% Ni).
In the base case design, the feed to the first stage of the methane compressor is superheated to
–101°C (-150°F) in order to use less expensive metallurgy. Without superheating, the methane
compressor is made of 304 SS at double the cost of a conventional carbon steel compressor.
With superheating, fine grain carbon steel (3.5% Ni) is used at about 15% more cost than
conventional carbon steel for the same suction volume compressor.
At constant plant fresh gas conditions, superheating reduces the work of the propylene
compressor while shifting more of the load to the ethylene and methane compressors (Table 12).
The net result is a reduction of 230 Hp (0.29%) in the gas turbine load from 98,660 Hp without
superheating to 98,430 Hp. The methane compressor load increases by 2,200 Hp (12.9%) from
17,000 Hp to 19,200 Hp. The LNG production rate decreases slightly with superheating (0.3
mt/H) while the LPG recovery is unchanged (fuel gas is higher).
The sensitivity of the process to feedstock inlet pressure and to feedstock methane content
are also shown in Table 12. For the fresh feed gas supply pressure at 775 psig, as occurs when
32
PEP REVIEW 2003-15
the plant is new instead of our average design pressure of 650 psig, reduces the gas turbine
power usage by 2,040 Hp (1,520 kw) or 2%. The fresh feed rate could be increased slightly.
Assume the raw natural gas feed composition were richer, methane content dropping from
92 vol% to 85 vol% with constant distribution of the C2 and heavier hydrocarbons, CO2 and
nitrogen. Loading the gas turbine drive, the fresh feed rate increases to 644.1 mt/H. The LNG
production rate is essentially unchanged, increasing by 1.1 mt/H to 526.9 mt/H. Raw LPG
increases by 66.5 mt/H (370%) to 84.6 mt/H from 18.0 mt/H. The C3/C4 LPG product is estimated
to increase to 34.5 mt/H from 7.4 mt/H. Natural gasoline production similarly increases to 32.3
mt/H from 6.9 mt/H. The methane compressor load is reduced by 660 Hp (3.5%) from 19,200 Hp
to 18,535 Hp.
If the rate of fresh feed is held constant at 5,789 mt/H, the gas turbine is unloaded and the
power drops 7,305 Hp to 91,125 Hp from 98,430 Hp. In this case, the rates of the C3/C4 LPG and
natural gas by-products are estimated to be 30.98 mt/H and 29.07 mt/H respectively.
33
PEP REVIEW 2003-15
Table 12
SUMMARY OF REFRIGERANT COMPRESSOR POWER REQUIREMENTS
Lean Gas________________________
Cold C1
High Feed
Gas
Compressor
Pressure
Inlet
Base Case
Fresh Feed,
vol% C1
Fresh Feed
Supply, psig
Plant Fresh
Feed, mt/H
Methane
Compressor
Inlet, °C (°F)
Rich Gas_________________
Loaded
Compressor
Unloaded
Compressor
92.1
92.1
92.1
85.0
85.0
650
650
775
650
650
578.9
578.9
578.9
644.1
578.9
-101 (-150)
-153 (-243)
-101 (-150)
-101 (-150)
LNG, mt/H
525.8
526.1
-101 (-150)
523.2
526.9
473.2
LPGa. mt/H
18.05
18.05
19.37
84.57
76.01
Refrigerant
Stage
Methane
1
3,876
2,849
3,819
3,871
3,867
2
5,226
4,559
5,144
5,043
4,653
3
10,097
9,594
9,871
9,621
8,769
Total
19,199
17,002
18,834
18,535
17,289
1
8,642
8,977
7,256
8,083
7,739
2
12,022
12,466
10,548
11,752
11,095
3
22,095
21,889
22,830
21,906
19,706
Total
42,759
43,332
40,634
41,741
38,540
1
19,605
19,606
20,377
19,987
19,034
2
19,003
18,814
18,647
19,334
17,646
3
17,063
16,907
16,732
17,368
15,905
Total
55,671
55,327
55,756
56,689
52,585
98,430
98,659
96,390
98,430
91,125
117,629
115,661
115,224
116,965
108,414
328
328
331
342
Ethylene
Propylene
____________________Power, Hp per train_______________________
Total, C2 + C3
Total , per train
Plant kwH/mt
LNG
a
Raw LPG stream separated.
333
34
PEP REVIEW 2003-15
COST ESTIMATE
In the following subsections, we discuss the investment and operating costs for a grassroots
LNG plant located on the U.S. Gulf Coast. Capital and operational cost factors must be applied
to estimate the capital costs and operating costs for realistic locations. The plant treats and
liquefies raw natural gas feedstock supplied by pipeline downstream of a slug catcher. The plant
uses a closed loop cascade refrigeration circuit with pure propylene, ethylene and methane
refrigerants.
The base case plant capacity is a nominal 525 mt/h (4.375 million mt/y at 0.95 on stream
factor, 600 MM scf/d) of LNG. The feedstock is a lean raw natural gas containing about 92 vol%
of methane and relatively low nitrogen and carbon dioxide contents. The plant contains two
equal capacity trains for gas processing including treating to remove carbon dioxide and water.
The plant produces all its utilities; about 13,000 kwh of interruptible electricity is sold outside the
plant. An U.S. Gulf Coast plant location is assumed for capital cost estimation; additional
shipping charges and a construction cost factor can be applied for other locations. Grassroots
construction on a cleared, level site with utilities access is assumed. Overnight construction is
also assumed, that is, there is no allowance for finance interest or price escalation before or
during construction. Licensee fees are also excluded.
On the basis of published data the capital investment for 600 million scf/d cascade
refrigeration plant built with two trains of 300 million scf/d is very competitive with a propane
precooled mixed refrigerant plant built in a single train of 600 million scf/d. The proposed
cascade plant can be expanded by adding a train or reduced in half by building only one train.
Because of the parallel train configuration, one achieving a modular capacity of 300 MM Scfd
LNG the cost per production unit is almost linear. This is certainly not the case with the
precooled mixed refrigerant process where the cost of a 600 million scfd plant is considerably
lower than the cost of two 300 million scfd plants. Therefore one advantage of the evaluated
cascade scheme is an additional modular production of 300 million scfd could be added at a very
competitive investment cost, compared with the mixed refrigerant process, when the market
develops.
Production cost and some equipment costs are estimated using the PEP Cost database,
version 3.1.5, developed by PEP. However equipment costs are mainly based on venders’
estimates since much very large or special equipment is involved. In the course of developing
our design, we had discussions with individuals at several contractors and equipment vendors
whom we cordially thank for their advice, information and assistance. Foster Wheeler of Canada,
St Katherine, Ontario, provided technical data and cost information related to the steam cycle.
Chart Industries in LaCross, Wisconsin and their agent, SME Associates, LLC in Houston, Texas
provided technical information and cost data for the “cold box” cryogenic section and heat
rejection system. Since turbine and compressor efficiencies are critical to the conceptual design,
the compression machinery sizing, performances and pricing were provided verbally by General
Electric and their joint venture Novo Pignone of Italy. The Elliott Company confirmed the
performance and provided additional data and the pricing information.
Capital Cost
Table 13 summarizes the capital investment for the plant. Since the plant is designed in
modules of 300 million scf/d, cost exponents are not applicable. Table 14 itemizes the capital
cost by section. The estimates are in mid-2002 U.S. dollars (PEP Cost Index of 620) and are
based on construction on the U.S. Gulf Coast. The base case battery limits investment estimate
totals $542 million. With a 15% contingency, the total fixed capital investment (TFC) for the plant
35
PEP REVIEW 2003-15
is about $912 million including off-site utilities and storage. (With a 25% contingency factor, the
TFC is $992 million.)
The total capital investment of $912 million in LNG production is measurably lower than the
investments reported by others, which range from $205/mt to $380/mt of LNG [9]. The basis of
these literature values is not reported (whether ships are included in the highest values for
instance). However we believe our estimate, $181 per ton LNG based on US Gulf Coast labor
cost and productivity, that was corroborated by several sources of expertise represents a realistic
scenario. After adding 15% contingency the capital investment is $209 per ton excluding the cost
of land, owner’s cost (royalty fees), or any unusual soil conditions. As discussed in the sections
above, the process differences in our plant from the conventional processes can account for at
least a portion of the savings.
The Liquefaction section contains the largest and most expensive individual pieces of
equipment, costing $353 million; compression and refrigeration account for $261 million of this.
Total offsites cost $370 million. Storage of $143 million and marine system costs of $142
million are the largest costs. The $32 million allotment for General Service Facilities assumes a
stand alone grassroots unit. The other category of general service items includes a fire station
and equipment, cafeteria, fences and all other support facilities. At an existing plant site most of
these services would already be available at a modest incremental cost, if any, and be
considerably less than at a totally new location.
For cost estimation, the amine unit cost is treated as a packaged unit using a curve type
estimate. Nevertheless, for high CO2 content (say over 4% as opposed to 0.2-1.2 vol % in the
given design basis and more so for significant sulfur content, which is not the case here), the cost
of the gas treating could reach a significant proportion and factor equipment estimate item by
item would be appropriate for achieving a credible cost estimate. Further, the feed gas is
assumed to enter to ISBL at 45 bar-g. This is not necessarily the case in all situations and
expensive pre-compression could be required depending on location and depletion level of the
gas reservoir. The gas inlet system may include slug catcher equipment that we consider to be
an upstream unit within the gas production facility.
We chose sea water cooling over air cooling for our process design. The capital investment
associated with sea water circulation for two trains is about $40 MM. About $27 MM could be
attributed to two refrigeration trains rejecting 1,000 MM Btu/hr and the balance for steam turbines
condensers rejecting 500 MM Btu/H. The investment associated with refrigeration cycle heat
rejection is an additional $18 MM a total $45 in heat rejection associated with refrigeration. The
appropriate investment in air cooling rejecting 1,000 MM Btu/hr at a temperature difference of
12°C (22°F) is estimated to be $34 MM. On this basis, the potential additional 10% extra LNG
production capacity by far out weighs the reduction of $11 MM in capital investment. Given the
same situation for variable speed gas turbines such as frame 5D this issue of air cooling versus
sea water cooling could turn more towards a break even situation.
Regarding heat rejection from the steam/power cycle, about 450-500 MM Btu/hr, the
estimated cost of steam condensers is relatively smaller than propylene condensers because of a
higher heat transfer coefficient. Once sea water cooling is established for the refrigeration cycle,
the incremental cost of sea water circulation for the steam cycle is relatively small. On this basis,
total sea water cooling system is used.
Because sea water cooling is a high maintenance service, two spare exchangers are added
to the operating 24 propylene condensing exchangers, thus allowing scale removal while
operating at maximum summer sea water temperatures. As to the steam condensers of the
methane refrigeration compressors, one common spare is added to the two existing condensers.
36
PEP REVIEW 2003-15
The same is the case with the power generation condensers where the common spare also
serves as a condenser to the startup turbines.
The cost is based on demineralizing water instead of desalination of water.
For LNG storage, we assume the double containment tanks with outer shell of concrete
could be 25-30% more expensive than single containment tanks. However this reduces plot area
and increases safety. The basis of LNG storage capacity is nominal 8 days flexible capacity.
Two 50% power generators are chosen. It is recognized that a single unit would result in
lower investment, and indeed if power back up from a local grid exists this could be a viable
approach. However, at this point the facility is designed to be fully self sustained and no single
equipment failure will cause a total shut down.
Making the gas treating, methane refrigeration and power generation units into single train
units is estimated to reduce the capital cost by $22-25 million or about 2.5%. Our assessment is
this cost reduction does not justify the potential reduction in the plant’s on stream factor.
37
PEP REVIEW 2003-15
Table 13
LNG BY CONVENTIONAL CASCADE REFRIGERATION
TOTAL CAPITAL INVESTMENT
CAPACITY: 9,646 MILLION LB/YR (4,380,000 T/YR)
LNG
AT 0.95 STREAM FACTOR
PEP COST INDEX: 620
COST
($1,000)
------------
PACKAGE UNITS:
AMINE UNIT (SECTION 100)
DRYING/MERCURY REMOVAL (SECTION 200)
LIQUEFACTION (SECTION 300)
LPG FRACTIONATION (SECTION 400)
HEAT RECOVERY/STEAM GENERATION (SECTION 500)
BATTERY LIMITS, INSTALLED
CONTINGENCY, 15%
CAPACITY
EXPONENT
-------------------------UP
DOWN
---------- ----------
22,000
15,000
354.16
15,000
65,000
--------471,160
70,670
--------541,830
BATTERY LIMITS INVESTMENT
OFF-SITES, INSTALLED
STORAGE (SECTION 600)
MARINE SYSTEM (SECTION 700)
143,000
142,140
---------
UTILITIES & STORAGE
WASTE TREATMENT (SECTION 800)
285,140
5,000
GENERAL SERVICES FACILITIES (SECTION 900)
32,000
--------322,140
TOTAL
CONTINGENCY, 15%
OFF-SITES INVESTMENT
48,320
--------370,460
TOTAL FIXED CAPITAL
912,290
38
PEP REVIEW 2003-15
Table 14
LNG BY CONVENTIONAL CASCADE REFRIGERATION
CAPITAL INVESTMENT BY SECTION
CAPACITY: 9,646 MILLION LB/YR (4,380,000 T/YR)
LNG
AT 0.95 STREAM FACTOR
PEP COST INDEX: 620
COST
($1,000)
-----------PACKAGE UNITS:
AMINE UNITS (2) (SECTION 100)
DRYING/MERCURY REMOVAL UNITS (2) (SECTION 200)
LIQUEFACTION SECTION (2) (SECTION 300)
PROPYLENE PRE-CHILLING, INCLUDING ETHYLENE CONDENSING
PROPYLENE COMPRESSORS (2) INCLUDING GEAR BOXES
GAS TURBINES (2)
START UP/BOOSTER STEAM TURBINE
ETHYLENE COMPRESSORS (2)
METHANE COMPRESSORS (2) AND STEAM TURBINES (2)
COLD BOXES (2) (1,200,000 LB) INCLUDING EXCHANGERS
FUEL GAS COMPRESSOR (1), ELECTRIC MOTOR
COLD FUEL GAS BLOWER (1)
REFRIGERANT PRESSURE VESSELS (18)
TOTAL, LIQUEFACTION SECTION
22,000
15,000
33,030
31,850
115,150
17,150
38,250
53,900
45,920
5,040
3,600
10,000
--------289,010
15,000
LPG FRACTIONATION (SECTION 400)
HEAT RECOVERY/STEAM GENERATION (SECTION 500)
HEAT RECOVERY/STEAM GENERATORS (HRSG) (2)
HOT BY-PASS STACK (1)
AUXILIARY BOILER, PACKAGE UNIT (1)
WATER TREATING DEAERATORS, STACK, BOILER FEED WATER PUMPS
POWER GENERATORS (2), INCLUDING CONDENSORS
HEAT REJECTION REFRIGERATION
TOTAL, HEAT RECOVERY/STEAM GENERATION SECTION
BATTERY LIMITS, INSTALLED
CONTINGENCY, 15%
20,000
1.50
5,000
5,500
15,000
18,000
--------65,000
471,160
70,670
--------541,830
BATTERY LIMITS INVESTMENT
39
PEP REVIEW 2003-15
Table 14 (Continued)
LNG BY CONVENTIONAL CASCADE REFRIGERATION
CAPITAL INVESTMENT BY SECTION
CAPACITY: 9,646 MILLION LB/YR (4,380,000 T/YR)
LNG
AT 0.95 STREAM FACTOR
PEP COST INDEX: 620
COST
($1,000)
-----------OFF-SITES, INSTALLED
STORAGE (SECTION 600)
LNG TANKS (2), DOUBLE CONTAINMENT
LPG TANK (1)
NATURAL GASOLINE TANK (1)
ETHYLENE REFRIGERANT TANK (4)
110,000
5,000
3,000
4,000
PROPYLENE REFRIGERANT TANK (6)
AMINE TANK (2)
WATER STORAGE TANKS (2)
LNG PRODUCT PUMPS (8)
6,000
1,000
2,000
12,000
--------143,000
TOTAL, STORAGE
MARINE SYSTEM (SECTION 700)
JETTY/CAUSEWAY (4,500 FT)
SEA WATER INTAKE SYSTEM INCLUDING CHLORINATION (2)
CONCRETE PIPES, ONE SUCTION LINE, ONE DISCHARGE LINE (4,500 FT EACH)
SEA WATER PUMPS, INSTALLED (10)
LNG PIPES, 36 IN. I.D. (2) 4,500 FT EACH
VAPOR RETURN LINE, 24 IN. I.D. (1) 4,500 FT
POWER CABLES (2)
LNG TANKER OFF GAS COMPRESSOR (1)
70,000
9,000
20,000
27,000
6,000
4,000
2,000
4,140
--------142,140
TOTAL, MARINE SYSTEM
UTILITIES & STORAGE--
285,140
WASTE TREATMENT (SECTION 800)
COLD FLARE LINE, 24 IN. I.D. (2)
WARM FLARE LINE, 36 IN. I.D. (2)
FLARE STACKS, 150 FT (2)
2,000
2,000
1,000
--------5,000
TOTAL WASTE TREATMENT
GENERAL SERVICES FACILITIES (SECTION 900)
BALLAST WATER TANK AND TREATING (1)
FIRE FIGHTING SYSTEM (FIRE RINGS, PUMPS AND WATER TANK)
SEWAGE & DRAINAGE
SECURITY & LIGHTING
INSTRUMENT AIR
40
3,000
5,000
2,000
2,000
1,000
PEP REVIEW 2003-15
Table 14 (Concluded)
LNG BY CONVENTIONAL CASCADE REFRIGERATION
CAPITAL INVESTMENT BY SECTION
CAPACITY: 9,646 MILLION LB/YR (4,380,000 T/YR)
LNG
AT 0.95 STREAM FACTOR
PEP COST INDEX: 620
COST
($1,000)
------------
NITROGEN AND LIQUID NITROGEN SYSTEM (1)
DIGITAL CONTROL SYSTEM, INCLUDING COMPUTER AND AUXILIARY
CONTROL ROOM, BLAST PROOF (1)
SUBSTATION (1)
GENERAL BUILDINGS, SHOPS, ADMINISTRATION BUILDING, LAB
ROADS AND PARKING
TELECOMMUNICATIONS
GENERAL ALLOWANCE FOR OTHER ITEMS
TOTAL GENERAL SERVICES FACILITIES
1,500
1,500
2,000
1,500
3,000
2,000
2,000
5,500
--------32,000
--------322,140
TOTAL OFFSITES
41
PEP REVIEW 2003-15
Production Cost
Table 15 shows the estimated production cost of LNG product including feedstock and byproduct prices and utilities. The feedstock and product values and utility costs are based upon
PEP Yearbook 2002 prices for the U.S. Gulf Coast. All production costs are expressed as ¢/lb of
isooctane product. By-products and utilities generated are shown as credits (negative values).
Credit is shown for 13,000 kw of interruptible electricity sold outside the plant and valued at 75%
of non-interruptible power value. The value of raw natural gas feedstock is for distressed gas
where the price is set by negotiation. We used a value of $0.75/MM Btu HHV or 1.735 ¢/lb.
The gross raw material cost 1.94 ¢/lb of LNG product. Of the gross raw material cost, 0.27
¢/lb is recovered as by-product credits for LPG and natural gasoline. A utilities credit for
electricity exported of –0.03 ¢/lb of LNG brings the net variable cost to 1.64 ¢/lb.
Labor costs are based upon 9 operators per shift (two marine operators and 7 plant
operators). Total U.S. operating labor compensation, including benefits, is $37.27/hr each. In
likely locations for a LNG plant, the prevailing cost of labor is considerably lower. Maintenance
labor is estimated as 2.5%/yr of the battery limits investment.
Since online control
instrumentation is assumed, control laboratory labor is estimated as 20% of the direct operating
labor cost. Total labor cost is estimated to be 0.28 ¢/lb. In the U.S., the staff is estimated to be
92 employees (Table 16). The combined cycle mode as suggested may have increased the
staffing by about 20-25% over the common practice when using exclusively gas turbines.
42
PEP REVIEW 2003-15
Table 15
LNG BY CASCADE REFRIGERATION
PRODUCTION COSTS
PEP COST INDEX: 620
VARIABLE COSTS
UNIT COST
-------------------------
RAW MATERIALS
RAW NATURAL GAS
CONSUMPTION
PER LB
--------------------------
¢/LB
----------
1.735 ¢/LB
1.11684 LB
1.94
--------1.94
9.62 ¢/LB
10.8 ¢/LB
-0.014014 LB
-0.013149 LB
-0.13
-0.14
---------0.27
GROSS RAW MATERIALS
BY-PRODUCTS
LPG
NATURAL GASOLINE
TOTAL BY-PRODUCTS
UNIT COST
-------------------------
CONSUMPTION
CONSUMPTION
PER LB
PER KG
-------------------------- --------------------------
UTILITIES
ELECTRICITY
-0.03
---------0.03
TOTAL UTILITIES
43
PEP REVIEW 2003-15
Table 15 (Concluded)
LNG BY CONVENTIONAL CASCADE REFRIGERATION
PRODUCTION COSTS
PEP COST INDEX: 620
RAW GAS COST ($/MM BTU HHV)
CAPACITY (MILLION LB/YR)*
INVESTMENT ($ MILLIONS)
BATTERY LIMITS (BLI)
OFFSITES
TOTAL FIXED CAPITAL (TFC)
0.50
0.75#
1.00
9,646
------------
9,646
------------
9,646
------------
912.3
NEGL
--------912.3
912.3
NEGL
--------912.3
912.3
NEGL
--------912.3
1.29
-0.27
-0.03
--------0.99
1.94
-0.27
-0.03
--------1.64
2.58
-0.27
-0.03
--------2.28
0.03
0.24
0.01
--------0.28
0.03
0.24
0.01
--------0.28
0.03
0.24
0.01
--------0.28
0.28
NEGL
--------1.55
0.28
NEGL
--------2.20
0.28
NEGL
--------2.84
0.22
0.19
--------1.96
0.22
0.19
--------2.61
0.22
0.19
--------3.25
0.95
--------2.91
0.95
--------3.56
0.95
--------4.20
0.06
--------2.97
0.06
--------3.62
0.06
--------4.26
3.15
--------6.12
2.50
--------6.12
1.86
--------6.12
PRODUCTION COSTS (¢/LB)
RAW MATERIALS
BY-PRODUCTS
UTILITIES
VARIABLE COSTS
OPERATING LABOR, 9/SHIFT, $37.27/HR
MAINTENANCE LABOR, 2.5%/YR OF BLI
CONTROL LAB LABOR, 20% OF OPER LABOR
LABOR COSTS
MAINTENANCE MATERIALS, 3%/YR OF BLI
OPERATING SUPPLIES, 10% OF OPER LABOR
TOTAL DIRECT COSTS
PLANT OVERHEAD, 80% OF LABOR COSTS
TAXES AND INSURANCE, 2%/YR OF TFC
PLANT CASH COSTS
DEPRECIATION, 10%/YR OF TFC
PLANT GATE COSTS
G&A, SALES, RESEARCH
NET PRODUCTION COST
ROI BEFORE TAXES, 33.5, 26.6, or 19.8^%/YR OF TFC
PRODUCT VALUE
----------------------------------* OF LNG
# BASE CASE
^ RESPECTIVELY
44
PEP REVIEW 2003-15
Table 16
U.S. STAFFING ESTIMATE
Operations
28, 2 per shift
Maintenance
24, 16 day time plus 2 on shift
Engineering
8, 4 day time plus 1 on shift
Marine Operators
8, 2 per shift
Administrative
6, 2 day time plus 1 on shift
Purchasing
2, day time
Laboratory
4, 1 per shift
Security
8, 2 per shift
Site Manager, Head of Operations, Head of
Engineering, Head of Administration
4
Total
92
Maintenance materials are estimated at 3%/yr of the battery limits investment; operating
supplies at 10% of operating labor costs. Adding maintenance materials and supplies to the
variable materials and labor costs produces a total direct operating cost of 2.20 ¢/lb of LNG
product.
Our production cost estimate includes charges for plant overhead (80% of labor costs),
property taxes and insurance (2%/yr of total fixed cost) and depreciation (10%/yr, straight line).
Adding the overhead and taxes and insurance costs to the total direct operating cost gives a
plant cash cost of 2.61 ¢/lb. Depreciation further adds 0.95 ¢/lb to bring the plant gate cost to
3.56 ¢/lb.
General and administrative expenses (G&A), sales, and research and development (R&D)
expenses for this process are assumed to be 1% of the battery limits plus off-site investment
costs. LNG or natural gas is a mature commodity sold through well established networks in
competitive markets. With G&S, sales and R&D expenses, the net production cost becomes
3.62 ¢/lb of LNG product or $1.54/MM Btu HHV.
At 25% ROI before taxes, the product value is 5.98 ¢/lb or $ 2.54/MM Btu HHV. This value
is lower than the $2.60/MM Btu HHV calculated at $0.75/MM Btu HHV raw gas cost from reported
information because we have credit for electricity and our capital investment is lower than other
designs [6].
At $0.50/MM Btu HHV as the cost of the raw natural gas fresh feedstock, net production cost
drops to 2.97 ¢/lb ($1.26/MM Btu HHV) from 3.62 ¢/lb ($1.54/MM Btu HHV) of the base case at
$0.75/MM Btu HHV. Raising the cost to $1.00/MM Btu HHV, causes the net production cost to
jump to 4.26 ¢/lb or $1.81/MM Btu HHV.
Using a 25% contingency increased the net production cost to 3.79 ¢/lb or $1.69/MM Btu
HHV.
45
PEP REVIEW 2003-15
Profitability
As mentioned in the Product Cost section above, the net production cost of the LNG in the
U.S. is estimated to be 3.62 ¢/lb or $1.54/MM Btu HHV. To determine a value at the production
plant, transportation and the cost of re-gasification are subtracted from the $3.50/MM Btu HHV
U.S. value of natural gas. An allowance for about 2,000 miles of transportation with large tankers
is assumed at $0.55/MM Btu HHV. A re-gasification cost of $0.35/MM Btu HHV is assumed. The
price of LNG at the plant then is $2.60/MM Btu HHV or 6.12 ¢/lb of LNG product.
With these assumptions, the before tax return on investment (ROI) is 26.4% for the base
capacity plant, just greater than the 25% ROI value frequently used to screen potential projects in
the petroleum or chemicals industries. However, the LNG business has some characteristics of a
utility, for instance, long term sales contracts and a dedicated supply of raw gas. Utility returns
on investment are lower since these market risks are deemed to be lower. A lower ROI may thus
be acceptable. A lower return may happen due to shipping longer distances or higher regasification costs than assumed.
At $0.50/MM Btu HHV as the cost of the raw natural gas fresh feedstock, ROI jumps to a
very attractive 33.3% from 26.6% of the base case at $0.75/MM Btu HHV. Raising the cost to
$1.00/MM Btu HHV, causes the ROI to decline to a marginally acceptable 19.7%.
Using a 25% contingency decreased the ROI to 22.7% from 26.4% with 15% contingency.
46
PEP REVIEW 2003-15
REFERENCES
Literature
1.
ConocoPhillips, “Kenai Liquefied Natural Gas Operation,” sales brochure
2.
Linde Science and Technology (Jan. 2003)
3.
Zeus Development LNG report (Sept. 2003)
4.
Houser, C. G., et al., “Phillips Optimized Cascade Process,” GASTECH 96, Vienna, Austria.
(December 1996)
5.
Herandez, Rick, “ConocoPhillips – Bechtel Global LNG Collaboration,” (ca. 2003)
6.
Kotzot, H. J. “LNG Plant Size versus LNG Transportation Distance,” 2001 AICHE Spring
National Meeting, paper 54e (Feb. 2001)
7.
Cambridge Energy Research Associates, Inc., “CERA’s LNG Quarterly Review, Between
the Reality and the Hype,” Special Report®, (c2003)
8.
Mak, J., et al. “LNG Flexibility,” Hydrocarbon Engineering, 8, 10 (Oct. 2003), 26-31
9.
Yost, C., et al. “Benchmarking study compares LNG plant costs,” Oil Gas J., 101, 15 (Apr.
14, 2003) 56-59
10. Sen, C. T., “LNG poised to consolidate its place in global gas trade,” Oil Gas J., 101, 24
(June 23, 2003) 72-74, 76-81
11. Sen, C. T., “New supply projects to push LNG into major markets,” Oil Gas J., 101, 25 (June
30, 2003) 64-68, 71
12. Tullo, A. H. “Catalyzing GTL,” C&E News 81, 29 (July 21, 2003) 18-19
13. Avidan, A., et al. “Study evaluates design considerations of larger, more efficient liquefaction
plants,” Oil Gas J., 101, 32 (Aug. 18, 2003) 50-54
Patents
US 4680041 DeLong, B. W., (to Phillips Petroleum Company), “Method for cooling normally
gaseous material,” US Patent 4,680,041 (Jul. 14, 1987)
US 5611216 Low, W. R., (not assigned), “Method of load distribution in a cascaded refrigeration
process,” US Patent 5,611,216 (Mar. 18, 1997)
US 5669234 Houser, C. G., et al., (to Phillips Petroleum Company), “Efficiency improvement of
open-cycle cascaded refrigeration process,” US Patent 5,669,234 (Sep. 23, 1997)
US 6016665 Cole, E. T., et al., (to Exxon Production Research Company), “Cascade
refrigeration process for liquefaction of natural gas,” US Patent 6,016,665 (Jan. 25,
2000)
US 6253574 Stockmann, R., et al., (to Linde Aktiengesellschaft), “Method for liquefying a stream
rich in hydrocarbons,” US Patent 6,253,574 (Jul. 3, 2001)
47
PEP REVIEW 2003-15
PEP Publications
PEP Report
29G
Nielsen, Richard. “Ethylene Plant Enhancement,” PEP Report 29G, SRI
Consulting, Menlo Park, CA (April 2001)
PEP Review
91-1-4
Ma, James J. L. “Removal of Mercury from Ethylene Plant Feedstock and
Cracked Gas Streams,” PEP Review 91-1-4, SRI Consulting, Menlo Park, CA
(July 1992)
PEP Review
2002-7
Cesar, Marcos A. “Liquefied Natural Gas by the Mixed Fluid Cascade
Process, PEP Review 2002-7, SRI Consulting, Menlo Park, CA (November
2003)
PEP
Yearbook
Wang, Shao-Hwa (Sean), ed. “PEP Yearbook International 2002”, Vol. 1E,
SRI Consulting, Menlo Park, CA (2002)
48
PEP REVIEW 2003-15
Figure 1 (Sheet 1 of 5)
LNG BY CASCADE REFRIGERATION PROCESS
TWO TRAINS (One Shown)
SECTION 200: DRYERS & MERCURY REMOVAL
SECTION 100: AMINE UNIT
F-201
2
C-202 A,B
C-201 A,B,C
Natural Gas
to E-304
3
23OC
V-201
V-102
Natural Gas
E-201
17OC
650 psig
Mole Sieve Drying
E-101
1
Condensate to
Amine Unit
50OC
V-101
CO2 (g)
To Vent
5
54OC
46OC
Water
C-101
60OC
C-102
E-108
T-101
E-102
Water
36OC
4
To Fuel Gas
Steam
Steam
38OC
100 psia
E-106
V-103
122OC
E-107
From Amine
Sump
E-103
T-101
Amine
Surge Drum
PEP Report 2003-15
V-101
Water Knock
Out Drum
C-101
CO2
Absorber
V-102
CO2 Absorber
Overhead Drum
V-103
V-201
Fuel Gas Condensate
Separator
Drum
E-104
E-105
F-201
Dryer Regeneration
Furnace
C-102
CO2 Stripper
C-201 A,B,C
Mole Sieve
Dryer
C-202 A,B
Mercury
Adsorber
2003
49
Figure 1 (Sheet 2 of 5)
LNG BY CASCADE REFRIGERATION PROCESS
TWO TRAINS (One Shown)
SECTION 300: LIQUEFACTION
K-302
Fuel Gas
322 psia
Exhaust Gas to Heat
Recovery System, Sh. 5
M-302 4 in. Water-g
Air 2,155,000 lb/H 21OC
Combustor
O
Steam, 915 psia, 382 C
G-301
K-303
1482OC
K-301
Speed
Reducer
To E-327,
Sh. 3
3
2
1
2
3
35OC
60OC 137 psia
20
220 psia
Hot
Air
Compressor Expander
From E-327
K-304
Speed
Reducer
1
6OC
72 psia
E-301
-32OC
26 psia
22
Condensate to V-501, Sh. 5
E-302 A-L
M-301
V-301
17OC
137 psia
O
35 C
To Condensate/Steam
Drum V-501, Sh. 5
21
E-303
25
V-302
-4OC
74 psia
K- 305
V-303
23
-32OC
28 psia
24
Sea
Steam 915 psig,
482OC
Uncontrolled
Extraction
Steam to Amine
& LPG Units
Fuel Gas from
Cold Box, Sh. 3
E-324
Sh. 3
6
E-323
Sh. 3
V-304
E-306
-32OC
594 psia
E-326
Sh. 3
E-325
Sh. 3
7
E-304
Natural Gas to
Cold Box, Sh. 3
E-305
Raw LPG to
LPG Unit, Sh. 4
Fuel Gas to K-309, Sh. 3
O
Treated Dry Natural Gas
M-301
Screen
PEP Report 2003-15
G-301
Chlorination
Unit
M-302
Air
Filter
K-301
Start-up/
Peaking
Turbine
K-302
Gas
Turbine
3
23 C
610 psia
K-303
K-304
Ethylene
Propylene
Compressor Compressor
V-301
High Pressure
Propylene Flash
Drum
-1OC
K-305
Electricity
Generator
V-302
Mid Pressure
Propylene Flash
Drum
-31OC
V-303
Low Pressure
Propylene Flash
Drum
V-304
Raw LPG
Separator
2003
50
Figure 1 (Sheet 3 of 5)
LNG BY CASCADE REFRIGERATION PROCESS
TWO TRAINS (One Shown)
SECTION 300: (Concluded)
24OC
-41OC
117 psia
300 psia
30
K-303, Sh.2
C3 =Refrigerant
26 psia
-43OC
58 psia
-37OC
61 psia
34
E-324
C3 =Refrigerant
-4OC
24OC
127 psia
-3OC
31
E-325
482OC, Steam
900 psia
K-107
-92OC
24 psia
E-323
O
-3 C
Stage 3
C3 =Refrigerant
-4OC
C3 =Refrigerant
-32OC
E-326
Stage 2
O
-21 C
540 psia
-57OC
120 psia
V-305
Stage 1
44
-23OC
245 psia
40
37
-3OC
105 psia
-81OC
-75OC
E-327
47
-28OC
295 psia
Sea Water
To V-501, Sh. 5
41
-101OC
24 psia
32
33
35
-90OC
V-106
-74OC
64 psia
V-308
530 psia
V-309
-129OC
113 psia
O
-109 C
273 psia
42
43
36
Cold Box
45
V-307
-92OC
27 psia
V-310
46
-153OC
28 psia
E-318
38
E-321
O
11
25
O
-73 C
-134OC
28 psia
16 psia
-51 C
K-308
O
E-319
O
-55 C
-72 C
E-320
-128OC
-90OC
-72OC
E-308
From V-304
Sh. 2
To E-306
Sh. 2
-32OC
594 psia
E-307
E-311
E-309
E-310
-33OC
E-315
E-313
E-312
-108 C
-76OC
-93OC
-157OC
19 psia
E-316
-109 C
O
-51OC
8
9
E-314
-113OC
LNG to
Loading
T-601
18
E-317
38OC
Fuel Gas
12
From E-306
Sh. 2
From
Tankers
O
-91OC
6
NNF
10
V-311
-152OC
299 psia
O
-34 C
19
M
K-309
K-303
Ethylene
Compressor
(See Sh. 2)
PEP Report 2003-15
V-305
HighPressure
Ethylene
Flash Drum
V-306
Mid Pressure
Ethylene
Flash Drum
V-307
Low Pressure
Ethylene
Flash Drum
K-306
Methane
Compressor
Steam Turbine
K-307
Methane
Compressor
V-308
High Pressure
Methane Flash
Drum
V-309
Mid Pressure
Methane Flash
Drum
V-310
K-308
Low Pressure
Fuel Gas
Methane Flash Cold Blower
Drum
K-309
Fuel Gas
Compressor
V-311
T-601
LNG Flash LNG Storage
Drum
Tank (Offsite)
2003
51
Figure 1 (Sheet 4 of 5)
LNG BY CASCADE REFRIGERATION PROCESS
ONE TRAIN
SECTION 400: LPG FRACTIONATION
-19OC
440 psia
Raw LPG
from V-304
Sh. 2
-32OC
544 psia
-32OC
Ref.
C3
To Vapor
Recovery, NNF
154 psia
-30OC
435 psia
E-401
V-401
V-403
V-402
7
16
E-405
C2 Rich Distillate to
Fuel Gas, NNF
15
E-406
C-401
Fuel Gas
49OC
C3C4
LPG
35OC
C-402
13
133OC
450 psia
50 psig
50 psig
Steam
149OC
O
149 C
E-402
Steam
E-403
14
E-404
V-401
Raw LPG
Flash Drum
PEP Report 2003-15
C-401
Deethanizer
V-402
Deethanizer
Reflux Drum
C-402
Debutanizer
Natural Gasoline
49OC
V-403
Debutanizer
Reflux Drum
2003
52
Figure 1 (Sheet 5 of 5)
LNG BY CASCADE REFRIGERATION PROCESS
TWO TRAINS (One Shown)
SECTION 500: HEAT RECOVERY SYSTEM
V-501
Steam
Condensate
Steam 900 psig, 482OC
250,000 lb/H
O
40 C
V-502
M-501
G-501
498OC
313OC
182OC Flue Gas
O
Exhaust Gas from
Gas Turbine K-302, Sh. 2
549 C
12 in Water-g
E-501
E-502
E-503
E-504
Vent
1115 psia
G-502
149OC
100 psig
Make up Water
49,726 lb/H
Steam
V-503
V-504
Boiler Feed
Water
282OC
Flue Gas
Steam 900 psig, 482OC
250,000 lb/H., Max.
Fuel Gas
F-501
Air
K-501
V-501
Condensate
Drum
G-501
Heat Recovery
Steam Generator
K-501
Air
Blower
PEP Report 2003-15
G-502
Demineralizer
or Desaltation Unit
V-504
Auxiliary
Boiler Steam
Drum
V-502
HRSG Steam
Drum
V-503
Deaerator
M-501
Stack
F-501
Auxiliary
Boiler
2003
53