PEP Review 2003-15 BASE LOAD LNG BY CASCADE … LOAD LNG BY CASCADE REFRIGERATION By David Netzer...

PEP Review 2003-15BASE LOAD LNG BY CASCADE REFRIGERATION

ByDavid Netzer And Richard Nielsen

(December 2003)

ABSTRACT

Movement of natural gas from remote locations with abundant supply of natural gas to theconsuming countries can be economically achieved only via the liquefaction route with shippingby ocean tankers. About 160 million metric tons/y of new liquefied natural gas (LNG) capacity isbeing implemented or in various planning stages in addition to the existing 100 million mt/y ofLNG 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 andBechtel are increasing the market share of cascade refrigeration technology.

This Review evaluates the economics of a base loaded, generic cascade refrigeration LNGplant nominally producing at least 600 million scf per stream day (4.375 million mt/y at 0.95 onstream factor) of LNG using two 50% capacity refrigeration and liquefaction trains. The feed gasis 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) recoveryprocess is used and the nitrogen stripping step is avoided. The refrigeration systems use acombined cycle mode of propylene and ethylene cycles driven by gas turbines and methanerefrigeration driven by steam turbines. A closed loop methane refrigeration system is used forthis 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 tonCO2/ton of LNG, as opposed to a typically reported 0.25-0.35 ton CO2/ton LNG) with low NOxemission.

INTRODUCTION

The production of Liquefied Natural Gas (LNG) has been commercially practiced since1960. Since LNG is stored and delivered at atmospheric pressure and -160°C (-256°F), verydeep refrigeration is needed with the associated large energy consumption. Two types of LNGplants have been built:

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• 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 oceantankers.

The capacities of the peak shaving plants are rather small, more than an order of magnitudesmaller than base load LNG plants. Based load LNG is the focus of this report. Typically, theliquefaction energy in a LNG plant is reported to be in the range of 9-12% of the heat energy inthe natural gas where 9-10% energy shrinkage is known to be a typical number for modern megatonnage capacity avoiding combined cycle systems. The capital investment in modern LNGfacilities is reported to be over $1.0 billion with 45-60% attributed to offsites and infrastructuredepending on one’s definition of offsites, particularly LNG storage, the marine system, and theheat 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 willminimize on site gas consumption and will minimize greenhouse gas emissions of thecombustion CO2. The proposed design represents a relatively low greenhouse gas emissionplant (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%, isnearly proportional to the heat rejection and would almost proportionally affect the LNGproduction rate for a given turbo machinery configuration and thus indirectly affect the cost ofproduction.

For evaluating the process economics, we use a base case of lean gas feed, 92 vol% (85wt%) methane containing under 1.2 vol% CO2 and less than 1 vol% nitrogen supplied at 46.5kg/cm2-a (650 psig). The nominal production rate of 525 mt/h (4.375 million mt/y at 0.95 onstream factor or 600 million scf/sd) of LNG depends upon the average ambient air temperatureand 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 oneGeneral Electric F7A frame, or equivalent gas turbines such as Rolls-Royce’s Trent model, in therefrigeration system. Therefore, the performance of the gas turbine is a key factor in establishingthe 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, 12in. water-g). This steam drives two 50% methane refrigeration compressors and, on a separateshaft, two power generation turbines. All turbines condense at summer conditions of 40.6°C, 57mm 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, weselected a closed loop cycle while recognizing that the open loop concept could avoid a separatefuel gas compressor, an issue to be evaluated on a case specific basis [1,4,5]. Further, unlikeConocoPhillips, we have decided to use propylene refrigerant rather than propane refrigerant,thus reducing the volumetric refrigerant flow by 22% over propane and allowing more vendorselection options for a single propylene compressor.

The recent design trend by both Bechtel and Kellogg Brown & Root seems to suggest theuse of air cooling as a method of heat rejection from the refrigeration cycle. From publicinformation 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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rejection to sea water is prohibited by local law, water cooling would be the more economicalmethod of heat rejection in most cases: fresh water as first priority; once through sea water as asecond priority. A combination of these including rejecting the heat to seawater via a temperedwater 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 secondaryimportance compared with the capital investment associated with compression and several site-specific factors. Nevertheless, the correct selection of the optimized refrigeration cycle andassociated drivers affects the compression and heat rejection systems. This choice becomes animportant item in the project evaluation especially that many of the site specific items are totallyindependent 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 project34,000 bpsd mixed liquids calls for investment of $675 million (or just under $20,000 per dailybarrel, which is just less than 80% of the benchmark $25,000 per daily barrel). The heatequivalent of 34,000 bpsd is estimated at 200 million scfd of LNG, 1.46 million mtpy. On thisbasis the GTL capital investment is 2.2 times higher per Btu and even after adjustment foreconomy of scale costs about 1.8 times higher than modular construction of a LNG plant of 300million scfd [12].

It is recognized that cost of transporting LNG say $0.75/MM Btu HHV ($35/ton) is double thecost 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 seemsby far the more economical fuel.

The thermal efficiency of GTL, based on 330 million scfd feed gas, is about 60% as opposedto 94% efficiency for LNG. The additional gas consumption increases the cost of GTL byadditional $27 per ton, which cancels out the re-gasification, and half of the transportation costadvantage of GTL.

The greenhouse gas emission from GTL production is about 2.0 ton of CO2 per ton ofsynthetic crude, 10 times higher over the LNG case. The synthetic crude is yet to be fractionatedin 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, assuggested, is close or equal to the efficiency of the propane pre-cooled mixedrefrigerant system and somewhat lower than mixed fluid cascade (MFC) refrigeration.However, at the end, no one process for gas liquefaction has a decisive advantage overother processes that would make an LNG project economically viable in a given locationwhile 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 (4.375million mt/y) to 555 mt/H (4.619 million mt/y) by using the startup steam turbines on acontinuous basis as booster turbines and utilizing the auxiliary steam generator on a fulltime 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 85vol% increases production capacity by 0.2% while maintaining the same gas turbinepower. The LPG and natural gasoline production increases by 469%.

4. Increasing the feed gas pressure from 650 psig to 775 psig decreases LNG productionby 0.5 wt% at a constant raw gas plant feed rate but the total refrigeration compressorhorsepower 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, thisoption increases the capital investment by about $4.5 million.

6. The key factors in the overall relative economics of LNG are related to the correctselection 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 LNGloading, 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 investmentsreported by others. However we believe our estimate, $181 per ton LNG based on USGulf Coast labor cost and productivity, that was corroborated by several sources ofexpertise represents a realistic scenario. After adding 15% contingency the capitalinvestment is $209 per ton excluding the cost of land, owner’s cost (royalty fees), or anyunusual 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. At23,550 Btu/lb (HHV), the net production cost at the LNG ocean tanker is US $1.54/MMBtu HHV of LNG. For a return on investment of 25%, the product value is $2.54/MM BtuHHV of LNG.

9. Based upon the $3.50/MM Btu HHV value of natural gas in the U.S., the product valueat the plant is estimated to be $2.60/MM Btu HHV or 6.12¢/lb of LNG after subtractingtransportation cost of $0.55/MM Btu HHV for about 2,000 miles shipping and $0.35/MM

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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 rootsLNG facilities built. In the 1990’s, six grass roots LNG plants were added to the list. Now, sevennew LNG plants are in advanced planning stage or in engineering and construction phases. Asshown below, over 90% of new LNG facilities are planned for tropical equatorial or sub-tropicalregions. Our selection of the generic ambient conditions reflects this reality. The mediancapacity of the proposed new facilities is in the order of 4.5 million mt/y. This is a factorconsidered 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 areasare listed in Table 1. LNG Projects in advanced planning or the engineering and constructionphase total over 60 million mt/y as listed in Table 2. Twenty-two projects are reported to beplaned for the long term and total an additional 110 million mt/y (Table 3). Current and proposedprojects range in capacity from 3 to 11.2 million mt/y.

Table 1MAJOR 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 2LNG PROJECTS -- ADVANCED PLANNING ORENGINEERING 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 3LONG 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 re-gasification and unloading terminals need to be constructed. Forty terminals including 19 inJapan and four in North America are in operation. Many terminals are in planning stage includingfourteen new LNG terminals for North America. The proposed North American terminals includean off shore terminal near Oxnard, California where it is assumed LNG will be re-gasified on aplatform structure 21 miles from shore [3]. An additional terminal in Baja California Mexico ismostly dedicated for Southern California. A LNG terminal in Freeport, Texas would become acaptive 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 Mexiconear Tampico, Mexico.

LNG Specifications

The main specifications on LNG in Europe, Japan, Korea and the US Gulf Coast andNortheast are a maximum 0.5 vol% of C5 and heavier components and a higher heating value inthe range of 950 to 1,100 Btu/scf. Specifications in California are listed in Table 4.

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Table 4CALIFORNIA 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 systempreceded by propane refrigeration. Close to 90% of these plants are licensed by Air ProductsCorporation, Inc. (APCI). The mixed refrigerant liquefaction systems use a mixture of mostlymethane and ethane, about 1.2-2.0/1 mole ratio. Depending on the feed gas composition, up to3 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 toachieve maximum thermodynamic efficiency. The fundamentals of mixed refrigerant comparedto conventional refrigeration are discussed in PEP Report 29G Ethylene Plant Enhancement(2001).

One known significant exception to the propane pre-chilling, mixed refrigerant approach isthe Phillips Kenai Peninsula plant in Alaska that started up in 1969 [1]. This plant produces about1.5 million mt/y of LNG by a cascade refrigeration system using pure propane, ethylene andmethane refrigeration cycles. The efficiency of this cascade system has been reported to be onthe lower end of the scale, about 88% [US 5611216]. However, we understand the plant hasproven to be a reliable and profitable operation. Aside from refrigeration cycle as such, animportant 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 aremore likely to be on the order of 70% as opposed to more modern centrifugal compressors withthree dimensional blades that achieve adiabatic efficiencies above 80% and approach 85% aswe show later.

Over the past several years several patents issued to Phillips seem to suggest animprovement in cycle efficiency results from improved refrigeration load distribution, nitrogenstripping 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 refrigerantsystem is slightly higher on a consistent basis (ambient conditions, identical machinery, heatrecovery and heat rejection philosophy). Nevertheless, other design related factors, operationalflexibility and other considerations, suggest that an objective comparative evaluation could bemade only on a case by case basis and on a site specific basis and not on generic liquefactiontechnology as such nor on liquefaction efficiency or any other single factor [US 5669234, US5611216 and US 4680041].

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Our understanding of Phillips’s patents is that their technology is driven by the desire to havean equal load distribution among the propane cycle, ethylene cycle and methane cycle. This loadequalization is achieved in part by superheating the methane and ethylene refrigerant vapors fedto the compressors probably to about -46°C (-50°F). This assumed design methodology allowsconventional carbon steel metallurgy in these compressors. Design and constructionconsiderations then allow six identical gas turbines, such as frame 5D (nominal 30,000 kw), to beused. Based on other information published by ConocoPhillips, we understand that the moremodern design uses open loop methane refrigeration where the methane refrigerationcompressor is used also as the fuel gas compressor [4]. Methane is fed as fuel to the gasturbines after cold recovery, thus helping to equalize the loads while avoiding a fuel gascompressor. Based on Phillips U.S. Patent 5,611,216, the estimated power consumption for38°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 gascompression. Based on the above, it is our judgment the drawbacks of load equalization couldbe up to 5% additional refrigeration power and, in case of fixed speed gas turbines, an actualreduction in production capacity. Further, based on a patent and other information published byConocoPhillips, we believe that this could result in more complex refrigeration cycle where thepropane 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 byPhillips in Kenai could have an advantage by allowing higher suction pressure to the methanecompressor, about 1.7 kg/cm2-a (24 psia) instead of an estimated 1.1 kg/cm2-a (16 psia) for anassumed open loop compressor. Only one electric motor driven fuel gas compressor is used forboth trains. In case of outage of the fuel gas compressor, back up is provided by a draw from thefeed gas.

Our base case design includes superheating of the refrigerant gas to the methanecompressor. 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 atsuction temperatures under about –130°C (-200°F) and industrial experience by otherscompressing 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 firststage of the methane compressor from about -152°C (-240°F) to about –101°C (-150°F). Thiswill 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 compressormetallurgy at low temperatures. Based on our proposed approach of using a combined cycle anddriving the methane compressors with steam turbines, many of the above issues as applied byConocoPhillips for the gas turbine power cycle are becoming academic for the combined cycleapproach suggested.

Further, it is our understanding that the ethylene refrigeration cycle in the ConocoPhillipsdesign, aside from superheating the suction to the first stage to about -46°C (-50°F), comprisesonly a single side load with a probable goal of obtaining the compression in a single casing. Inour approach, no superheating of the suction to the ethylene compressors is employed. Twoside loads are used, reducing refrigeration load by 2% and increasing production capacity by 2-3%. 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. Theloads of the propylene compressor are about 47.3%, the refrigeration ethylene compressionloads are about 36.4% and the methane compressor loads are 16.3% of the total refrigerationload. In reality the combined load on the steam turbines may be about 27% of the total motivepower 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 ambienttemperature times, the start-up steam turbine attached to the gas turbine increases its relativesteam consumption and electric power export drops to near zero.

The actual production capacity is determined by the performance of the industrial equivalentof 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 platecapacity to show realistic mid life performance. Derating of 4% is possible and typically servesas a performance guarantee by the manufacturer. The generic 4.375 million t/y at 0.95 onstream factor (600 million scf/d) of LNG is simply a result of turbine performance and is related toambient conditions and perhaps to the mechanical condition of the turbo machinery. In additionto climatic assumptions, several generic assumptions are made as to infrastructure, feed gascomposition and LNG specifications. The cost estimate and subsequent economics are highlydependant upon the specific site location but could be adjusted based on variance ininfrastructure. The site assumptions, offsites definitions and turbine performances are included inthe review. Appropriate adjustments could be made for different site locations and ambientconditions. As said, the total capacity could probably be raised from 4.375 million mt/y to 4.60million t/y at 0.95 on stream factor by shifting the steam turbine loads from excess powergeneration to continuous boosting of the gas turbines.

In addition to the ConocoPhillips cascade refrigeration processes, Linde AG has developedthe Mixed Fluid Cascade (MFC) process where each of the three refrigerants cycles operate incascade using mixed refrigerant [US 6253574]. This concept is being engineered for theNorwegian North Sea project offshore near Hammerfest and is claimed to be of extremely highthermal efficiency, about 250 kwh refrigeration load per ton of LNG. However, in this project heatis rejected to 5°C (41°F) seawater as opposed to the more conventional 27°C (81°F) in moretypical LNG projects in tropical zones. In this temperature range the claimed refrigeration load ismore like 330 kwh per ton [2] of LNG as opposed to 333 kwh refrigeration per ton LNG calculatedfor 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 Review2002-7 (2003).

A recent interesting patent by Exxon suggests a new approach for producing pressurizedLNG (PLNG) [US 6,016,665]. LNG at about 30 kg/cm2-a and -95°C (410 psig, and -140°F) wouldbe produced by two cascade cycles using propane and ethylene as opposed to the current threecascade cycles comprising propylene, ethylene and methane refrigerants. Producing PLNG isestimated to save 50% or more of refrigeration power compared with LNG, along with avoidingmethane refrigeration, which substantially reduces the capital investment for the liquefactioncycle. However, this PLNG concept brings yet unknown issues related to transportation andstorage 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 onstream factor, 602 MM scf/d) of LNG depending on average ambient air temperature and heatrejection temperature. The feedstock is a lean raw natural gas containing about 92 vol% ofmethane. The levels of nitrogen and carbon dioxide are also relatively low. The generic processis a conventional cascade refrigeration process. The plant contains two equal capacity trains forgas processing including treating to remove carbon dioxide and water. Liquid petroleum gasesare separated into byproduct natural gasoline, mixed C3/C4 LPG and fuel gas in a single trainunit. The fuel gas is consumed within the plant. The plant produces all its utilities; about 13,000kwh of electricity can be sold outside the plant. The process design and utility rates used in thisreport are based on computations, published information, and nonconfidential information from


licensors and vendors. The design may or may not be similar to processes licensed or otherwisein use today.

The production concept uses two 50% liquefaction trains with General Electric F7A frame orequivalent gas turbines such as the Rolls-Royce Trent model. The performance of the gasturbine is a key factor in establishing the LNG production rate. Waste heat is being recovered as63 kg/cm2-g, 482°C (900 psig, 900°F) steam from the turbine exhaust gas at about 545°C-300mm water-g (1,020°F, 12 in. water-g) driving two 50% methane refrigeration compressors and, ona 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 LNGplants. This led to using two 50% GE F7A gas turbines, or equivalent, along with two GE, orequivalent, 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 4.375million mt/y of LNG at nominal atmospheric pressure, -160°C (-260°F). No liquid expanders areused, but in general a liquid expander can reduce energy consumption by 2-2.5% but was judgedto offer only a marginal economic advantage.

The process flow diagram of the LNG plant, Figure 1 (attached), is shown with the insideboundary 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 ischaracterized in Table 6. The process stream flows are summarized in Table 7 and refrigerantflows are summarized in Table 8. These flows are for the total plant consisting of two trains. Theprocess 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-Kwongequation of state method. Major equipment is listed with size and materials of construction inTable 9. On the equipment list, the two trains are identified in the equipment numbers by A andB. The letters C, D and higher indicate multiple pieces of equipment are used per train. Thefacility is self-contained in terms of electric power, fuel gas and all other utilities. Internal utilitiesaverage consumptions are summarized in Table 10. An average of 13,000 KW of interruptibleelectricity is sold outside the plant. The steam used to generate this power could be shifted torun booster turbines and increase LNG production by 6%.


Table 5DESIGN 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:

Speed, rpm Stage No. Efficiency

Propylene 2,600 1 0.837

2 0.842

3 0.838

Ethylene 3,600 1 0.824

2 0.847

3 0.844

Methane 8,500 1 0.775

2 0.788

3 0.814

Source: GE, The Elliott Company


Table 6FEED 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


Table 7LNG BY CASCADE REFRIGERATION

STREAM FLOWS

CAPACITY: 525 MT/H (9,646 MILLION LB/Y OR 602 MILLION SCF/SD OR 17.056 M3/SD)LNG

AT 0.95 STREAM FACTOR

1 2 3 4 5 6

Mol Wt

Raw Natural Gas Fresh

Feed

Molecular Sieve Dryer

Feed Treated GasAmine Unit Fuel Gas

CO2 Rich Gas

Prechilled Gas

Water 18.02 659 941 2 47 628 1

Carbon Dioxide 44.01 16,214 64 64 146 16,004 2Nitrogen 28.01 5,833 5,833 5,833 0 0 5,824Methane 16.04 1,062,337 1,061,082 1,061,082 1,151 105 1,056,040Ethane 30.07 101,144 101,039 101,039 94 10 97,754Propane 44.1 56,725 56,673 56,673 47 5 49,825Isobutane 58.12 14,199 14,189 14,189 10 0 10,377n-Butane 58.12 17,121 17,110 17,110 10 0 10,980Isopentane 72.15 7,775 7,770 7,770 5 0 3,062n-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

Total, lb/H 1,294,516 1,277,205 1,276,266 1,517 16,752 1,236,415kg/H 587,183 579,330 578,905 688 7,599 560,828lb-mole/H 72,339 71,855 71,802 82 405 70,859

Stream Flows, lb/Hr


Table 7 (Continued)LNG BY CASCADE REFRIGERATION

STREAM FLOWS



7 8 9 10 11 12

Mol Wt Raw LPGLNG to Storage

LNG Product

LNG Tank Flash

N2/Fuel Flash

Cold Fuel Gas

Water 18.02 1 1 1 0 0 0

Carbon Dioxide 44.01 62 2 2 0 0 0Nitrogen 28.01 9 2,839 2,147 692 2,985 3,677Methane 16.04 5,042 998,451 982,403 16,048 57,589 73,637Ethane 30.07 3,285 97,743 97,740 3 12 14Propane 44.1 6,848 49,825 49,825 0 0 0Isobutane 58.12 3,812 10,377 10,377 0 0 0n-Butane 58.12 6,130 10,980 10,980 0 0 0Isopentane 72.15 4,709 3,062 3,062 0 0 0n-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 77,328kg/H 18,076 533,347 525,753 7,594 27,481 35,075lb-mole/H 941 67,162 66,137 1,025 3,697 4,722

Stream Flows, lb/Hr


Table 7 (Concluded)LNG BY CASCADE REFRIGERATION

STREAM FLOWS



13 14 15 16 17

Mol Wt C3+Natural

GasolineLPG

ProductC2 Rich Fuel Gas

Total Fuel Gas

Water 18.02 1 1 0 0 47

Carbon Dioxide 44.01 61 62 0 0 146Nitrogen 28.01 0 0 0 9 3,686Methane 16.04 2 0 2 5,040 79,828Ethane 30.07 119 0 118 3,167 3,276Propane 44.1 6,737 72 6,628 148 195Isobutane 58.12 3,858 179 3,630 3 14n-Butane 58.12 6,179 420 5,709 1 12Isopentane 72.15 4,733 4,605 104 0 5n-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 87,214kg/H 14,365 6,913 7,368 3,796 39,559lb-mole/H 521 201 317 423 5,227

Stream Flows, lb/Hr



REFRIGERANT STREAM FLOWS



Propylene Refrigerant Ethylene Refrigerant Methane Refrigerant

StreamNo. lb/H 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


Table 9BASE LOAD LNG BY CASCADE REFRIGERATION

MAJOR EQUIPMENT

CAPACITY: 9,646 MILLION LB/YR (4,375,300 mt/y)LIQUIFIED NATURAL GASAT 0.95 STREAM FACTOR

EQUIPMENTNUMBER NAME SIZE MATERIAL OF CONSTRUCTION REMARKS

------------------- ---------------------------------------------- ---------------------------------------- ---------------------------------------------------------- ----------------------------------------------------------------------------------

COLUMNS

C-101A,B CO2 ABSORBER 9 FT DIA SHELL: C.S. 22 VALVE TRAYS, 24 INCH SPACING44 FT T-T TRAYS:

C-102A,B CO2 STRIPPER 11 FT DIA SHELL: C.S. 25 VALVE TRAYS, 24 INCH SPACING50 FT T-T TRAYS:

C-201A-F MOLE SIEVE DRYER 10 FT DIA SHELL: C.S.15 FT T-T PACKING: MOLECULAR SIEVES

C-202A-D MERCURY ADSORBER 10 FT DIA SHELL: C.S.10 FT T-T PACKING: HG ADSORBENT

C-401 DEETHANIZER, TOP SECTION 6.5 FT DIA SHELL: C.S. 10 VALVE TRAYS , 24 INCH SPACING20 FT T-T TRAYS: C.S.

BOTTOM SECTION 13 FT DIA TRAYS: C.S. 16 VALVE TRAYS , 24 INCH SPACING32 FT T-T TRAYS: C.S.

C-402 DEBUTANIZER 12 FT DIA SHELL: C.S. 26VALVE TRAYS, 18 INCH SPACING52 FT T-T TRAYS: C.S.

COMPRESSORS

K-301A,B START-UP/PEAKING TURBINE 9,387 BHP 2.5% CrK-302A,B GAS TURBINE 108,621 BHP 2.5% CrK-303A,B ETHYLENE COMPRESSOR 41,571 BHP 3.5% Ni ALLOY; STAGE 3, C.S. 3 STAGESK-304A,B PROPYLENE COMPRESSOR 56,322 BHP C.S. 3 STAGESK-305A,B ELECTRICITY GENERATOR 26,820 BHP C.S.K-306A,B METHANE COMPRESSOR STEAM 21,000 BHP 2.5% Cr

TURBINEK-307A,B METHANE COMPRESSOR 19,200 BHP C.S. 3 STAGESK-308A,B FUEL GAS COLD BLOWER 1,006 BHP 304 SSK-309A,B FUEL GAS COMPRESSOR 10,058 BHP C.S.K-501A-D AIR BLOWER 350 BHP C.S. ELECTRIC MOTOR

HEAT EXCHANGERS

E-101A-H FEEED GAS/AMINE OVERHEAD 20,000 SQ FT SHELL: C.S.EXCHANGER 7.5 MMBTU/HR TUBES: C.S.

E-102A,B AMINE TRIM COOLER 7,500 SQ FT SHELL: C.S. AIR COOLER11 MMBTU/HR TUBES: C.S.

E-103A,B LEAN/RICH AMINE EXCHANGER 1 3,,000 SQ FT SHELL: C.S.11 MMBTU/HR TUBES: C.S.

E-104A,B LEAN/RICH AMINE EXCHANGER 2 3,000 SQ FT SHELL: C.S.11 MMBTU/HR TUBES: C.S.

E-105A,B LEAN/RICH AMINE EXCHANGER 3 3,000 SQ FT SHELL: C.S.11 MMBTU/HR TUBES: C.S.

E-106A,B AMINE RECLAIMER 200 SQ FT SHELL: C.S.1 MMBTU/HR TUBES: C.S.

E-107A,B STRIPPER REBOILER 8,200 SQ FT SHELL: C.S.52 MMBTU/HR TUBES: C.S.

E-108A,B STRIPPER CONDENSER 3,800 SQ FT SHELL: C.S. AIR COOLERMMBTU/HR TUBES: C.S.

E-201A,B REGENERATOR GAS COOLER 22 SQ FT SHELL: C.S. AIR COOLER7 MMBTU/HR TUBES: C.S.

E-301A,B START-UP TURBINE STEAM 21,000 SQ FT SHELL: C.S.CONDENSER 133 MMBTU/HR TUBES: TITANIUM


Table 9 (Continued)BASE LOAD LNG BY CASCADE REFRIGERATION

MAJOR EQUIPMENT



------------------- ---------------------------------------------- ---------------------------------------- ---------------------------------------------------------- ----------------------------------------------------------------------------------

HEAT EXCHANGERS (CONT.)

E-302A-X PROPYLENE CONDENSER 32,000 SQ FT SHELL: C.S.42 MMBTU/HR TUBES: TITANIUM ALLOY

E-303A,B POWER GENERATOR STEAM SQ FT SHELL: C.S.CONDENSER 79 MMBTU/HR TUBES: TITANIUM ALLOY

E-304A,B FEED GAS FIRST CHILLER 17,000 SQ FT SHELL: C.S.16 MMBTU/HR TUBES: C.S.

E-305A,B FEED GAS SECOND CHILLER 30,000 SQ FT SHELL: C.S.25 MMBTU/HR TUBES: C.S.

E-306A,B LNG/FUEL GAS EXCHANGER SQ FT SHELL: C.S.0.6 MMBTU/HR TUBES: C.S.

E-307A,B FEED GAS PRE-COOLER NO. 1 SQ FT SHELL: ALUMINUM1 MMBTU/HR PLATES: ALUMINUM



E-310A,B FUEL GAS COLD RECOVERY NO. 3 SQ FT SHELL: ALUMINUM1.2 MMBTU/HR PLATES: ALUMINUM

E-311A,B FEED GAS LIQUEFIER SQ FT SHELL: ALUMINUM69 MMBTU/HR PLATES: ALUMINUM

E-312A,B FUEL GAS COLD RECOVERY NO. 2 SQ FT SHELL: ALUMINUM1 MMBTU/HR PLATES: ALUMINUM

E-313A,B LNG SUBCOOLER NO. 1 SQ FT SHELL: ALUMINUM20 MMBTU/HR PLATES: ALUMINUM

E-314A,B FUEL GAS COLD RECOVERY NO. 1 SQ FT SHELL: ALUMINUM1 MMBTU/HR PLATES: ALUMINUM

E-315A,B LNG SUSCOOLER NO. 2 SQ FT SHELL: ALUMINUM19 MMBTU/HR PLATES: ALUMINUM

E-316A,B LNG SUBCOOLER NO. 3 SQ FT SHELL: ALUMINUM20 MMBTU/HR PLATES:ALUMINUM

E-317A,B FUEL GAS COMPRESSOR INTERSTAGE SQ FT SHELL: NONE AIR COOLERCOOLER 1.7 MMBTU/HR TUBES: C.S.

E-318A,B METHANE REFRIGERANT CONDENSER SQ FT SHELL: ALUMINUM74 MMBTU/HR PLATES: ALUMINUM

E-319A,B METHANE REFRIGERANT DESUPERHEATER SQ FT SHELL: ALUMINUM18 MMBTU/HR PLATES: ALUMINUM

E-320A,B METHANE REFRIGERANT COOLER SQ FT SHELL: ALUMINUM22 MMBTU/HR PLATES:ALUMINUM

E-321A,B METHANE COMPRESSOR INLET SQ FT SHELL: C.S.COOLER 6 MMBTU/HR TUBES: C.S.

E-323A,B ETHYLENE ECONOMIZER FIRST 25,000 SQ FT SHELL: C.S.DESUPERHEATER 12 MMBTU/HR TUBES: C.S.

E-324A,B ETHYLENE ECONOMIZER SECOND 20,000 SQ FT SHELL: C.S.DESUPERHEATER 17 MMBTU/HR TUBES: C.S.

E-325A,B ETHYLENE DESUPERHEATER 32,000 SQ FT SHELL: C.S.36 MMBTU/HR TUBES: C.S.

E-326A-P ETHYLENE CONDENSER 30,000 SQ FT SHELL: C.S.32 MMBTU/HR TUBES: C.S.

E-327A,B TURBINE STEAM CONDENSER 100 MMBTU/HR TUBES: TITANIUM


Table 9 (Continued)BASE LOAD LNG BY CASCADE REFRIGERATION

MAJOR EQUIPMENT



------------------- ---------------------------------------------- ---------------------------------------- ---------------------------------------------------------- ----------------------------------------------------------------------------------

HEAT EXCHANGERS (CONCLUDED)

E-401 DEETHANIZER OVERHEAD 2,800 SQ FT SHELL: C.S. KETTLE TYPECONDENSER MMBTU/HR TUBES: C.S.

E -402 DEETHANIZER REBOILER 7,200 SQ FT SHELL: C.S.MMBTU/HR TUBES: C.S.

E-403 DEBUTANIZER REBOILER 4,800 SQ FT SHELL: C.S.MMBTU/HR TUBES: C.S.

E-404 NATURAL GASOLINE COOLER 1,600 SQ FT SHELL: C.S. AIR COOLER, FINED TUBESMMBTU/HR TUBES: C.S.

E-405 DEBUTANIZER CONDENSER 24,000 SQ FT SHELL: C.S. AIR COOLERMMBTU/HR TUBES: C.S.

E-406 FUEL GAS/LPG EXCHANGER 1,000 SQ FT SHELL: C.S.MMBTU/HR TUBES: C.S.

E-501A,B HRSG SUPSERHEATER TUBES 31,000 SQ FT SHELL: NONE FINNED TUBES IN G-50164 MMBTU/HR TUBES: 5% Cr ALLOY

E-502A,B HRSG STEAM GENERATION TUBES 110,000 SQ FT SHELL: NONE FINNED TUBES IN G-501175 MMBTU/HR TUBES: C.S.

E-503A,B HRSG ECONOMIZER TUBES NO. 1 32,000 SQ FT SHELL: NONE FINNED TUBES IN G-50152 MMBTU/HR TUBES: C.S.

E-504A,B HRSG ECONOMIZER TUBES NO. 2 32,000 SQ FT SHELL: NONE FINNED TUBES IN G-501108 MMBTU/HR TUBES: C.S.

FURNANCESF-201A,B DRYER REGENERATION FURNACE 18 MMBTU/HR TUBES: C.S.F-501 AUXILIARY BOILER 375 MMBTU/HR TUBES: 5% Cr ALLOY

TANKS

T-101 AMINE SURGE DRUM 42,000 GAL C.S.T-601A,B LNG STORAGE 37,800,000 GAL 9% Ni ALLOY 8 DAYS PRODUCTION, DOUBLE WALL WITH PERLITE

INSULATION T-602 LPG 840,000 GAL C.S. 14 DAYS PRODUCTION

T-603 NATURAL GASOLINE 840,000 GAL C.S. 14 DAYS PRODUCTIONT-604A-F PROPYLENE REFRIGERANT 50,000 GAL C.S.T-605A-D ETHYLENE REFRIGERANT 50,000 GAL C.S.T-606A,B FRESH WATER 5,000 GAL C.S.T-607A,B DEMINERALIZED BOILER WATER 5,000 GAL C.S.T-608 DIESEL FUEL 3,000 GAL C.S. FOR EMERGENCY GENERATORS

PRESSURE VESSELS

V-101A,B WATER KNOCK OUT DRUM 1,000 GAL C.S.V-102A,B CO2 ABSORBER OVERHEAD DRUM 800 GAL C.S.V-103A,B FUEL GAS SEPARATOR 600 GAL C.S.V-201A,B CONDENSATE DRUM 900 GAL C.S.V-301A,B HIGH PRESSURE PROPYLENE 42,000 GAL C.S.

FLASH DRUMV-302A,B MID PRESSURE PROPYLENE 42,000 GAL C.S.

FLASH DRUMV-303A,B LOW PRESSURE PROPYLENE 25,000 GAL C.S.

FLASH DRUMV-304A,B RAW LPG SEPARATOR 15,000 GAL C.S.V-305A,B HIGH PRESSURE ETHYLENE 25,000 GAL 3.5% Ni ALLOY

FLASH DRUM


Table 9 (Concluded)BASE LOAD LNG BY CASCADE REFRIGERATION

MAJOR EQUIPMENT



------------------- ---------------------------------------------- ---------------------------------------- ---------------------------------------------------------- ----------------------------------------------------------------------------------

PRESSURE VESSELS (CONCLUDED)

V-306A,B MID PRESSURE ETHYLENE 20,000 GAL 3.5% Ni ALLOYFLASH DRUM

V-307A,B LOW PRESSURE ETHYLENE 15,000 GAL 3.5% Ni ALLOYFLASH DRUM

V-308A,B HIGH PRESSURE METHANE 30,000 GAL 3.5% Ni ALLOYFLASH DRUM

V-309A,B MID PRESSURE METHANE 15,000 GAL 304 SS FLASH DRUM

V-310A,B LOW PRESSURE METHANE 7,000 GAL 304 SSFLASH DRUM

V-311A,B LNG FLASH DRUM 30,000 GAL 304 SSV-401 RAW LPG FLASH DRUM 800 GAL C.S.V-402 DEETHANIZER REFLUX DRUM 800 GAL C.S.V-403 DEBUTANIZER REFLUX DRUM 2,200 GAL C.S.V-501A,B CONDENSATE DRUM 7,000 GAL C.S.V-502A,B HRSG STEAM DRUM 10,000 GAL C.S.V-503A,B DEAERATOR 12,000 GAL C.S.V-504A,B AUXILIARY BOILER STEAM DRUM 10,000 GAL C.S.

MISCELLANEOUS EQUIPMENT

M-301A,B SCREENM-302A,B AIR FILTERM-501 STACK 15 FT DIA, BASE C.S.

50 TALL

SPECIAL EQUIPMENT

S-301 COLD BOX

PACKAGED UNITS

G-301G-501A,B HEAT RECOVERY STEAM GENERATORG-502A,B DEMINERALIZER (OR

DESALTATION UNIT) 100 GPM C.S.

PUMPS

SECTION OPERATING SPARES OPERATING BHP--------------- ------------------- -------------- -------------------------

100 4 8 1,620200 0300 4 4 200400 6 6 60500 4 2 2,000600 20 10 9,770700 8 2 17,600


Table 10 LNG BY CASCADE REFRIGERATION

UTILITIES SUMMARY



SectionFuel Gas,MM Btu/Ha

Steam, 100 psigsaturated, lb/H

Superheated Steam,900 psig, 482°C

(900°F), lb/H Electricity, kw

100 76,000b 1,200c

200 36 xd

300 1,640e 500,000f 7,600

400 12,000 x

500 x (start up) ______ _______ 1,500

Battery Limits Total 1676 88,000 500,000 10,300

600 595g

700 11,300h

800 x

900 x x

All Other 74 350

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.


Section 100 – Amine Unit

Raw natural gas feedstock at an average 17°C (63°F) and 650 psia enters one of the twotrains at its amine unit (Figure 1, sheet 1) where acid gas (CO2 and H2S) is removed followingfree water separation (V-101). The feed gas is then heated to 36°C (97°F) by exchange in E-101with treated gas prior to entering the CO2 absorber (C-101) and being contacted with solvent, inour case, monoethanol amine (MEA). (For reasons of consistency, a generic MEA for CO2removal system is shown; similar configurations could apply to diethanol amine (DEA), diglycolamine (DGA), methyldiethanol amine (MDEA) and other solvents.) Using 20 wt% MEA (0.35 molCO2/mol MEA), CO2 is removed to under 50 ppm. On leaving the absorber at 50°C (122°F), theCO2 free natural gas heats the feed gas while being chilled to 23°C (73°F). Entrained liquid MEAis 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 richsolution leaves the bottom of the absorber at about 38°C (100°F) and is reduce in pressure toabout 100 psia. Light hydrocarbon gases dissolved in the MEA solution separate in V-103. Thegases are scrubbed with lean MEA solution and sent to the fuel gas system. The rich solution isheated by a series of exchangers (E-103 to E-105) with 122°C (252°F) lean solution from thebottom of the CO2 stripper (C-102). CO2 is stripped from the MEA solution with vaporized leansolution either from storage (E-106) or from a reboiler (E-107). The stripper is refluxed with asmall amount of make up water (about 18 lb/H) added at the cooler (E-108). CO2 rich gas isvented 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, oneis in service, one is held ready for service and one is being regenerated by stripping with hot (F-201) natural gas. Water condensed (E-201) from the wet regeneration gas flows to the amineunit (Section 100) to supply make up water.

Many natural gases around the globe are known to contain traces of mercury that is harmfulto aluminum alloy in the downstream cryogenic section. Mercury removal is necessary for safetydue to corrosion of cryogenic aluminum heat exchangers. Adsorption beds are used for tracemercury 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 liquidpropylene refrigerant, (2) separating raw liquid petroleum gas, (3) cryogenic cooling andcondensing 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 andethylene cools the high pressure methane. The prechilling occurs in large kettle typeexchangers. The cryogenic exchange occurs in aluminum plate-fin exchangers located in aninsulated cold box.

Dry treated feed gas enters the chilling train of the gas liquefaction section to be cooled bytwo 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


natural gas C5+ heavies specification, proceeds to the cold box for deep refrigeration via ethyleneand methane refrigeration.

The three-stage propylene compressor does the propylene refrigeration work. Propylene isflashed 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 alarge gas turbine (K-302). Heat is recovered from the turbine exhaust (about 545°C, 1020°F) inSection 500. The ethylene compressor stages 1 and 2 share the same case; the third stage isseparate. The train is started up with a steam turbine (K-301). The driving steam is condensedby exchange with seawater (E-301) that is chlorinated (G-301). Seawater also cools the highpressure propylene refrigeration loop (E-302) and also condenses (E-303) steam from theelectricity generator (K-305).

The bulk of the liquefaction occurs in the ethylene refrigeration cycle while the methanerefrigeration 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, (520psia). In the cold box, the gas is cooled by a series of heat exchanges in aluminum plate-finexchangers (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. Mostof the heat of compression is removed by cooling with propylene between stages 2 and 3 (E-323and E-324) and at high pressure (E-325 and E-326). Ethylene is flashed at three successivepressures (V-305 to V-306) to obtain the three temperature levels. The methane compressor (K-307) is steam driven (K-306). Seawater condenses the steam (E-327). Stages 2 and 3 are inone case. The inlet temperature is heated to –101C (-150F) by exchange with a slip stream ofthe natural gas (E-321) to allow the compressor to be made of less expensive metal. Most of theheat of compression is removed by exchange with ethylene (E-318 to E-320). Again themethane is flashed to three successively lower pressures (V-308 to E-310) to obtain threetemperature levels.

On leaving the cold box, the sub-cooled liquid flashes at 1.3 kg/cm2-a (19 psia) in flash drumV-311 prior to entering to LNG storage tank at 1.1 kg/cm2-a (16 psia). The flash tank providesprotection to the storage tank and also pre-flashes most of the nitrogen prior to storage. For highnitrogen content feedstocks, perhaps pre-flashing at 4-5 kg/cm2-a (55-70 psia) would bepreferable 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 andcompressed to 1.9 kg/cm2-a (27 psia) prior to recovering the “cold” by exchange with the gas inthe 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 (299-330 psia) and used as turbine fuel. Interstage (E-317) to about 38-49°C (100-120°F) is provideddepending upon the ambient air temperature. The flash temperature (first stage of methanerefrigeration) is continuously adjusted to provide the correct amount of fuel gas needed for thegas 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, Figure1, sheet 4). Ethane rich gas and occasionally distillate are separated from the flashed liquid bydistillation (C-401). Overhead vapor is partially condensed (E-401) and vapors separated fromthe 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.


The C3 and heavier bottoms are fractionated in debutanizer C-402 to produce C3/C4 LPGdistillate and natural gasoline bottoms products. Overhead vapor is usually totally condensed (E-403) 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). Thenatural gasoline product is cooled by air cooler E-404 to about 49°C (120°F) depending upon theambient 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 (Figure1, sheet 5). Steam at 63 bar-g and 480°C (900 psig and 900°F) would represent a class breakpoint and thus was selected. Steam at 63 kg/cm2-g (900 psig) is fed to four turbines driving thetwo 14,300 kw methane compressors and two nominal 20,000 kw power generators. Steam atabout 8.0 kg/cm2-g (114 psig) is uncontrollably extracted as a heat source for the MEAregeneration (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 steamgenerator (HRSG) G-501. The exhaust gas leaves the HRSG at 182°C (360°F) and flows tostack M-501. Condensate at about 40°C (104°F) from condensate drum V-501 is pumped atabout 115 psia through HRSG coils E-504 and preheated to about 145°C (290°F). Thecondensate is then deaerated (V-502) along with demineralized or desalted make up water fromG-502. The deaerated boiler feed water is pumped at 1115 psia through the HRSG at coils E-503 into steam drum V-502 at 965 psia. Liquid water from the steam drum circulates through coilE-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 915psia, 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 thepropylene and ethylene refrigerants. Two LNG storage tanks, each 145,000 M3 (900,000 bbl) aredouble walled, insulated and further contained by concrete walls. Each tank is installed with four3,000 M3/hr send out pumps. LPG is stored in two 20,000 bbl tanks with a refrigerated vaporrecovery system. Natural gasoline is stored in two fixed roof 20,000 bbl tanks also having vaporrecovery.

Make up amine is stored in two diked 1,000 bbl tanks. Make up propylene is stored in sixbullets each holding 50,000 gal under about 275 psia pressure. Make up ethylene is stored infour 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 3,000bbl 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 ontoships (not included). The jetty pipe rack has two LNG lines each 30 in. in diameter. Two vaporrecovery 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 are4 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 accountedas part of the marine system. Two seawater intake suction lines buried 15 ft below seawater


surface are 7 ft I.D. coated carbon steel pipes. Two pumping structures have 5 pumps on eachstructure. Driven by 2,500 Hp motors, each can pump 45,000 gpm. Also included are a pumphouse and a pump control center. The seawater is screened, chlorinated, filtered and transferredvia two 7 ft diameter concrete pipes laid under the water. The cooling seawater is returned to thesea through two 1000 meter water discharge lines, 7 ft each, made of coated concrete laid underwater.

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 flarestacks are 50 m (150 ft) tall.

Section 900 – General Offsites

General offsites includes roads and general infrastructure and such items as process andsanitary drain systems, telecommunications, control building including a digital control systemand auxiliary, craft shops, control laboratory. Fire protection and ballast water systems are alsogeneral 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 upis provided by 800 ton (1,000 m3) of liquid nitrogen. The evaporator uses sea water. Instrumentair 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 watertemperatures for a subtropical region (Table 5).

Our general turbo machinery configuration is two half plant capacity turbines in parallel, onein each train. Thus a forced or planed shut down of a gas turbine will result in a 50% reduction inLNG production. Whereas in a propane pre-cooled, mixed refrigerant system, the drivers arearranged in series. For example, in a typical 4.5 million mt/y LNG plant, a single frame 7 turbineis dedicated to low pressure and mid pressure mixed refrigerant service and a single frame 7 isused for high pressure mixed refrigerant and propane compression. Then a forced or planedshut down of any turbine results in 100% production loss. On this basis the common on streamfactor is elevated from 340 days per year (0.93) to 347 days per year (0.95). The ConocoPhillipsprocess utilizing 6 turbines will have some advantage in this respect since an outage of a singleturbine 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 MarketOverview section. The ultimate plant capacity is determined by matching the performance ofavailable turbine driver configurations. On a generic basis we decided upon two 7FA GE turbinesor equivalent, one in each train. A gas turbine is used for propylene and ethylene refrigerationwhile steam turbines were used for methane refrigeration and power generation. In our design, aforced or elective shut down of a given gas turbine will reduce production by 50% unlike theConocoPhillips 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 thatgains from reducing the number of turbines out weigh the reduction in flexibility.


The gas turbines are equipped with start up turbines running on steam from the auxiliaryboiler. The start up turbine could be used to boost capacity when the gas turbine is de-rated dueto high ambient temperatures. Furthermore, operation of the auxiliary boiler (250,000 lb/hr 900psig/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 be21°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 creditcould be taken for lower seawater temperatures. Since the average ambient air temperature inAugust 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 averageJanuary temperature is 14.5°C (58°F); thus turbine performance increases by 5%. The averagedaily 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 inany subtropical location. This is a lean gas (92.1 vol% methane) also low in CO2 and nitrogenthat would be a good representative composition for a majority of global locations. Ourconceptual design will not be affected by C2+ contents under 15 vol% nor by CO2 or nitrogencontents under 3.0 vol% each. This design would be applicable to the great majority (over about80% 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 requireincineration of the CO2 vent gas and, at some locations such as Qatar, the installation of sulfurrecovery units.

Over time as a gas field ages the gas supply pressure decreases and the CO2 contentincreases. In our case the pressure is initially assumed to be about 800 psia. Our 650 psigsupply pressure and 0.47 vol% CO2 is the mid-life pressure after 10-15 years and average CO2content.

Acid Gas Removal

MEA is assumed to be more economical solvent than MDEA for CO2 contents under 1.5vol%. Each train has an amine unit. A single amine unit serving both trains could be a viablechoice as well. However we judged that the savings in capital investment versus the extrareliability and flexibility of two units favored selection of the dual train system. Since the H2Scontent is very minimal, acid gas could be safely incinerated after stripping from the MEAsolution. The CO2 content is reduced to under 50 ppm. Thus DEA, MDEA or another solventcould be selected depending upon the operational preference and general businessconsiderations of the LNG producer. On a global basis, about 35% of LNG plants are usingMEA.

In our design, the amine unit is sized for 1.2 vol% CO2 in the raw feed gas even though theaverage is 0.47 mol% CO2. The maximum amine circulation rate is 725 gpm per train based on1.2 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 CO2greenhouse gas emission tax could make CO2 recovery from the amine unit an economicallyviable option, especially if no H2S is present. The CO2 vent gas is contaminated withhydrocarbons, and incineration of the vent gas in the auxiliary boiler could at time become aviable option especially when traces of H2S are present.


Nitrogen Removal

Since the nitrogen content of the gas is less than 1.0 vol% (0.3 vol%), no particular nitrogenremoval is needed. Flashing fuel gas at 1.3 kg/cm2-a (19 psia) in the process and flashing fromstorage at 16 psia diverts the bulk of the nitrogen into the fuel gas system. In our case, 67% ofthe nitrogen is flashed to the fuel gas. This however would have no impact on the operation ofthe gas turbines.

Gas drying and mercury removal

Molecular sieve gas drying at 23°C (73°F) is used to protect the cryogenic exchangers andpiping from any potential freezing.

Many natural gases around the globe are known to contain traces of mercury that is harmfulto aluminum alloy in the downstream cryogenic section. Mercury in light hydrocarbons is verybriefly reviewed in PEP Report 29G Ethylene Plant Enhancement (April 2001) and in PEPReview 91-1-4 Removal of mercury from ethylene plant feedstock and cracked gas streams (July1992). Mercury removal is necessary for safety due to corrosion of cryogenic aluminum heatexchangers. To remove mercury, the dry natural gas is passed through mercury adsorptionbeds. Carbon impregnated with sulfur (Calgon Carbon) or silver on alumina is used for vaporoushydrocarbons. Experienced contractors should be contacted about adsorbents and processconfigurations 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 uponpreliminary flow data. The final flow rates are only very slightly different. Because the flowlimitation on the assumed propane (used the ConocoPhillips process) compressor becomes abottleneck, 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 ethylenerefrigeration 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 asingle uncontrolled extraction of steam at a nominal 8 kg/cm2-g (115 psig). Turbinesteam is condensed at 41°C (106°F), 1.1 psia by seawater at 27°C to 32°C (to 81°F90°F).

• All the refrigeration loads, totaling about 1,000 MM Btu/H, are calculated on the basis ofseawater 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. Steamturbine drivers, electric motor drivers, combined cycle gas turbine–steam turbine could all beconsidered. Our design basis is a gas turbine coupled to an exhaust gas heat recovery steamgenerator to drive both the propylene and the ethylene compressors. A start up booster steamturbine is on the same shaft. As a result of the design basis and the preliminary recommendationof GE-Novo Pingone, the methane compressor is divided into two cases on the same shaft drivenat 8,500 rpm by a steam turbine. The ethylene compressor is also split.


In each train, a General Electric FA7 or equivalent gas turbine drives two refrigerationcompressors in series: about 31,900 kw ethylene refrigeration and 41,500 kw propylenerefrigeration. A 7,000 kw (9,400 Hp) steam turbine is used for start up as well for load boostingduring periods of high ambient temperatures. The estimated gas turbine performance issummarized in Table 11. The compressors have 6 wheels and two side loads. The gas turbinename 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 poweravailable 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 flashgas, 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, andclearly demonstrates that availability of turbo machinery is a key factor in the establishing thedesign basis.

A Frame 5 variable speed turbine has higher NOx emission (about 100 ppm) that in somelocations would require a de-NOx process. The 25 ppmv NOx emission with the Frame 7 turbinedoes not require NOx removal.

Based on our cost estimate, we are considering that the combined cycle mode, assuggested, represents a lower capital investment per ton of LNG compared with using exclusivelygas turbine drivers commonly practiced in the industry. Furthermore, the combined cycle modepresents the higher efficiency, lower CO2 greenhouse gas emissions and probably all togetherthe most economical configuration, at least in our case. Some locations may need to obtain freshwater for the combined cycle’s steam production and for employees from a packageddesalination unit instead of demineralizing water. Integration of desalination and powergeneration is very commonly practiced.


Table 11ESTIMATED PERFORMANCE OF GAS TURBINE (FRAME 7)

Load Condition Base Base Base Base

Inlet Loss in H2O 4. 4. 4. 4.

Exhaust Pressure Loss in H2O 12.0 12.0 12.0 12.0

Ambient Temperature °C (°F) 15 (59) 20 (68) 28 (82.4) 40 (104)

Fuel Type Methane Methane Methane 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

Exhaust Flow x103 lb/H 2234 2179 2090 1959

Exhaust Temperature °C (°F) 549 (1020) 553 (1028) 560 (1040) 569 (1057)

Exhaust Loss in H2O 12.0 @ ISO conditions

Application Mechanical Drive

Combustion System DLN Combustor

Emissionsa

NOx at 15% O2 ppmv 25. 25. 25.

CO ppmv 15. 15. 15. 15.

Source: GEa 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 levelsshown 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), weredetermined by turbo machinery analysis and wheels configuration provided by Novo Pingone. Adifferent machinery design could result in different side loads conditions.

The ethylene and methane refrigeration cycles use plate-fin core aluminum exchangerspacked 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. Thealuminum exchangers dictate mercury removal beds prior to the gas chilling train as discussedpreviously. Refinery grade propylene, typically 95 wt% propylene and 5 wt% propane, could beacceptable as the refrigerant.

Heat Rejection

The seawater circulation piping was sized to a maximum velocity of 12 ft/second due toconcerns of erosion. At the current design, the maximum temperature rise is set at 5°C (9°F)


however it is recognized that in some locations a maximum rise of 3.5-4.0°C (6-7°F) is allowed. Ifthe two spare sea water pumps would be run, the total circulation will increase by 25%, thetemperature 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 propylenerefrigerant 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). Themaximum ambient air temperature along the Eastern Mediterranean coast, for example, is about38°C (100°F) and this sets the discharge pressure of the propylene compressor if air cooling isapplied. For an air cooler condenser, this would require a minimum propylene condensingtemperature 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 trainof 73,000 kw thus reducing LNG production by 10%. The avoided sea water circulation perrefrigeration train would be 110,000 gpm, about 4,000 kw of circulation power. On the otherhand, the air cooler fans for 500 MM Btu/hr heat rejection per production train would consume4,000 kw as well. The economics of this design decision are briefly discussed below in thesection “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 theextra duct work, hot by pass stack and on stream time considerations could not be justified forthe marginal savings. The auxiliary start up boiler, 250,000 lb/hr capacity, will provide the backup steam needed for methane compression in the event the HRSG is by passed. The auxiliaryboiler 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, thusdependence on the local electricity grid is avoided in the event one generator is forced to shutdown (Figure 1, sheet 2). Under the normal scenario, the facility is self sufficient in electricpower. 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 veryvaluable 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 theamine and LPG Fraction sections (Figure 1, sheets 1 and 4). Uncontrolled extraction steam atabout 100 psig is higher than the 50 psig steam the process requires. As the flow rate slowsdown, the extracted steam pressure decreases. Controlled extraction is very expensive and notrequired 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 richgas from LPG fractionation is routed to the turbine fuel system. This is due to the fact that only asmall quantity of ethane must be recovered to meet the LNG specifications. Our higher heatingvalue of the LNG is 1,091 Btu/scf, just below the specification maximum of 1,100 Btu/scf, andessentially unchanged from the raw natural gas feedstock (1,095 Btu/scf). Our recovery of


ethane and propane is only 3.4 wt% and 12.2 wt% respectively. Recovery of ethane andpropane for sale as petrochemical feedstock is easier to do during regasification of the LNG atthe receiving terminal.

The pressure of the deethanizer column is set to allow at least sufficient fuel gas pressure tothe gas turbines. In the event that the LPG unit is forced to shut down, crude LPG after flashinglight 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 drasticallydepending on location and LNG delivery contract. For an location where the bulk of the productis delivered to a market a relatively short distance, about 2,000 miles, two LNG tanks (145,000M3, 900,000 bbl) seem appropriate considering the maximum capacity of LNG tankers and theaverage tanker capacity of 750,000 bbl. We assume the double containment tanks with an outershell of concrete would be 25-30% more expensive than single containment tanks. Howeverdouble containment reduces plot area if dikes can be avoided and also increases safety. Thebasis of storage capacity is nominal 8 days flexible capacity. If more marketing flexibility isrequired, to reach farther markets for instance, a third LNG tank could become reasonable. Thebottom line is simply stated: just like with turbine selection, storage capacity, flexibility, safety andon stream factor are inter-related to capital investment.

LNG tanker loading is assumed to take under 16 hours. On this basis, the evacuation rate is9,000 M3/hr per tank, using 7 bars pressure drop (40,000 gpm, 100 psi). Each tank is equippedwith four (3+ spare) 3,000 M3/hr, 750 kw pumps. All vapors displaced from LNG tankers duringloading, 12-16 hours per send-out cycle, is recycled to LNG storage. Heat leakage into thetankers 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 at21°C (70°F) ambient temperature is 98,430 Hp.

The propylene compressors and stage 3 of the ethylene are made of conventional carbonsteel. 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 methanecompressor 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 thanconventional carbon steel for the same suction volume compressor.

At constant plant fresh gas conditions, superheating reduces the work of the propylenecompressor 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 withoutsuperheating to 98,430 Hp. The methane compressor load increases by 2,200 Hp (12.9%) from17,000 Hp to 19,200 Hp. The LNG production rate decreases slightly with superheating (0.3mt/H) while the LPG recovery is unchanged (fuel gas is higher).

The sensitivity of the process to feedstock inlet pressure and to feedstock methane contentare also shown in Table 12. For the fresh feed gas supply pressure at 775 psig, as occurs when


the plant is new instead of our average design pressure of 650 psig, reduces the gas turbinepower 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 from92 vol% to 85 vol% with constant distribution of the C2 and heavier hydrocarbons, CO2 andnitrogen. Loading the gas turbine drive, the fresh feed rate increases to 644.1 mt/H. The LNGproduction rate is essentially unchanged, increasing by 1.1 mt/H to 526.9 mt/H. Raw LPGincreases by 66.5 mt/H (370%) to 84.6 mt/H from 18.0 mt/H. The C3/C4 LPG product is estimatedto increase to 34.5 mt/H from 7.4 mt/H. Natural gasoline production similarly increases to 32.3mt/H from 6.9 mt/H. The methane compressor load is reduced by 660 Hp (3.5%) from 19,200 Hpto 18,535 Hp.

If the rate of fresh feed is held constant at 5,789 mt/H, the gas turbine is unloaded and thepower drops 7,305 Hp to 91,125 Hp from 98,430 Hp. In this case, the rates of the C3/C4 LPG andnatural gas by-products are estimated to be 30.98 mt/H and 29.07 mt/H respectively.


Table 12SUMMARY OF REFRIGERANT COMPRESSOR POWER REQUIREMENTS

Lean Gas________________________ Rich Gas_________________

Base Case

Cold C1

CompressorInlet

High FeedGas

PressureLoaded

CompressorUnloaded

Compressor

Fresh Feed,vol% C1 92.1 92.1 92.1 85.0 85.0Fresh FeedSupply, psig 650 650 775 650 650Plant FreshFeed, mt/H 578.9 578.9 578.9 644.1 578.9MethaneCompressorInlet, °C (°F) -101 (-150) -153 (-243) -101 (-150) -101 (-150) -101 (-150)

LNG, mt/H 525.8 526.1 523.2 526.9 473.2

LPGa. mt/H 18.05 18.05 19.37 84.57 76.01

Refrigerant Stage ____________________Power, Hp per train_______________________

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

Ethylene 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

Propylene 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

Total, C2 + C3 98,430 98,659 96,390 98,430 91,125

Total , per train 117,629 115,661 115,224 116,965 108,414

Plant kwH/mtLNG 333 328 328 331 342a Raw LPG stream separated.


COST ESTIMATE

In the following subsections, we discuss the investment and operating costs for a grassrootsLNG plant located on the U.S. Gulf Coast. Capital and operational cost factors must be appliedto estimate the capital costs and operating costs for realistic locations. The plant treats andliquefies raw natural gas feedstock supplied by pipeline downstream of a slug catcher. The plantuses a closed loop cascade refrigeration circuit with pure propylene, ethylene and methanerefrigerants.

The base case plant capacity is a nominal 525 mt/h (4.375 million mt/y at 0.95 on streamfactor, 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 twoequal 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 theplant. An U.S. Gulf Coast plant location is assumed for capital cost estimation; additionalshipping charges and a construction cost factor can be applied for other locations. Grassrootsconstruction on a cleared, level site with utilities access is assumed. Overnight construction isalso assumed, that is, there is no allowance for finance interest or price escalation before orduring construction. Licensee fees are also excluded.

On the basis of published data the capital investment for 600 million scf/d cascaderefrigeration plant built with two trains of 300 million scf/d is very competitive with a propaneprecooled mixed refrigerant plant built in a single train of 600 million scf/d. The proposedcascade 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 ScfdLNG the cost per production unit is almost linear. This is certainly not the case with theprecooled mixed refrigerant process where the cost of a 600 million scfd plant is considerablylower than the cost of two 300 million scfd plants. Therefore one advantage of the evaluatedcascade scheme is an additional modular production of 300 million scfd could be added at a verycompetitive investment cost, compared with the mixed refrigerant process, when the marketdevelops.

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 developingour design, we had discussions with individuals at several contractors and equipment vendorswhom 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, Texasprovided technical information and cost data for the “cold box” cryogenic section and heatrejection system. Since turbine and compressor efficiencies are critical to the conceptual design,the compression machinery sizing, performances and pricing were provided verbally by GeneralElectric and their joint venture Novo Pignone of Italy. The Elliott Company confirmed theperformance and provided additional data and the pricing information.

Capital Cost

Table 13 summarizes the capital investment for the plant. Since the plant is designed inmodules of 300 million scf/d, cost exponents are not applicable. Table 14 itemizes the capitalcost by section. The estimates are in mid-2002 U.S. dollars (PEP Cost Index of 620) and arebased on construction on the U.S. Gulf Coast. The base case battery limits investment estimatetotals $542 million. With a 15% contingency, the total fixed capital investment (TFC) for the plant


is about $912 million including off-site utilities and storage. (With a 25% contingency factor, theTFC is $992 million.)

The total capital investment of $912 million in LNG production is measurably lower than theinvestments reported by others, which range from $205/mt to $380/mt of LNG [9]. The basis ofthese literature values is not reported (whether ships are included in the highest values forinstance). However we believe our estimate, $181 per ton LNG based on US Gulf Coast laborcost and productivity, that was corroborated by several sources of expertise represents a realisticscenario. After adding 15% contingency the capital investment is $209 per ton excluding the costof land, owner’s cost (royalty fees), or any unusual soil conditions. As discussed in the sectionsabove, the process differences in our plant from the conventional processes can account for atleast a portion of the savings.

The Liquefaction section contains the largest and most expensive individual pieces ofequipment, 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 $142million are the largest costs. The $32 million allotment for General Service Facilities assumes astand alone grassroots unit. The other category of general service items includes a fire stationand equipment, cafeteria, fences and all other support facilities. At an existing plant site most ofthese services would already be available at a modest incremental cost, if any, and beconsiderably less than at a totally new location.

For cost estimation, the amine unit cost is treated as a packaged unit using a curve typeestimate. Nevertheless, for high CO2 content (say over 4% as opposed to 0.2-1.2 vol % in thegiven design basis and more so for significant sulfur content, which is not the case here), the costof the gas treating could reach a significant proportion and factor equipment estimate item byitem would be appropriate for achieving a credible cost estimate. Further, the feed gas isassumed to enter to ISBL at 45 bar-g. This is not necessarily the case in all situations andexpensive pre-compression could be required depending on location and depletion level of thegas reservoir. The gas inlet system may include slug catcher equipment that we consider to bean upstream unit within the gas production facility.

We chose sea water cooling over air cooling for our process design. The capital investmentassociated with sea water circulation for two trains is about $40 MM. About $27 MM could beattributed to two refrigeration trains rejecting 1,000 MM Btu/hr and the balance for steam turbinescondensers rejecting 500 MM Btu/H. The investment associated with refrigeration cycle heatrejection is an additional $18 MM a total $45 in heat rejection associated with refrigeration. Theappropriate investment in air cooling rejecting 1,000 MM Btu/hr at a temperature difference of12°C (22°F) is estimated to be $34 MM. On this basis, the potential additional 10% extra LNGproduction capacity by far out weighs the reduction of $11 MM in capital investment. Given thesame situation for variable speed gas turbines such as frame 5D this issue of air cooling versussea water cooling could turn more towards a break even situation.

Regarding heat rejection from the steam/power cycle, about 450-500 MM Btu/hr, theestimated cost of steam condensers is relatively smaller than propylene condensers because of ahigher 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 addedto the operating 24 propylene condensing exchangers, thus allowing scale removal whileoperating at maximum summer sea water temperatures. As to the steam condensers of themethane refrigeration compressors, one common spare is added to the two existing condensers.


The same is the case with the power generation condensers where the common spare alsoserves 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 concretecould be 25-30% more expensive than single containment tanks. However this reduces plot areaand 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 inlower investment, and indeed if power back up from a local grid exists this could be a viableapproach. However, at this point the facility is designed to be fully self sustained and no singleequipment failure will cause a total shut down.

Making the gas treating, methane refrigeration and power generation units into single trainunits is estimated to reduce the capital cost by $22-25 million or about 2.5%. Our assessment isthis cost reduction does not justify the potential reduction in the plant’s on stream factor.


Table 13LNG BY CONVENTIONAL CASCADE REFRIGERATION

TOTAL CAPITAL INVESTMENT

CAPACITY: 9,646 MILLION LB/YR (4,380,000 T/YR)LNG


PEP COST INDEX: 620

CAPACITYEXPONENT

COST --------------------------($1,000) UP DOWN------------ ---------- ----------

PACKAGE UNITS:AMINE UNIT (SECTION 100) 22,000DRYING/MERCURY REMOVAL (SECTION 200) 15,000LIQUEFACTION (SECTION 300) 354.16LPG FRACTIONATION (SECTION 400) 15,000HEAT RECOVERY/STEAM GENERATION (SECTION 500) 65,000

---------BATTERY LIMITS, INSTALLED 471,160

CONTINGENCY, 15% 70,670---------

BATTERY LIMITS INVESTMENT 541,830

OFF-SITES, INSTALLEDSTORAGE (SECTION 600) 143,000MARINE SYSTEM (SECTION 700) 142,140

---------

UTILITIES & STORAGE 285,140WASTE TREATMENT (SECTION 800) 5,000

GENERAL SERVICES FACILITIES (SECTION 900) 32,000---------

TOTAL 322,140

CONTINGENCY, 15% 48,320---------

OFF-SITES INVESTMENT 370,460

TOTAL FIXED CAPITAL 912,290


Table 14LNG BY CONVENTIONAL CASCADE REFRIGERATION

CAPITAL INVESTMENT BY SECTION



PEP COST INDEX: 620

COST($1,000)------------

PACKAGE UNITS:AMINE UNITS (2) (SECTION 100) 22,000DRYING/MERCURY REMOVAL UNITS (2) (SECTION 200) 15,000

LIQUEFACTION SECTION (2) (SECTION 300)PROPYLENE PRE-CHILLING, INCLUDING ETHYLENE CONDENSING 33,030PROPYLENE COMPRESSORS (2) INCLUDING GEAR BOXES 31,850GAS TURBINES (2) 115,150START UP/BOOSTER STEAM TURBINE 17,150ETHYLENE COMPRESSORS (2) 38,250METHANE COMPRESSORS (2) AND STEAM TURBINES (2) 53,900COLD BOXES (2) (1,200,000 LB) INCLUDING EXCHANGERS 45,920FUEL GAS COMPRESSOR (1), ELECTRIC MOTOR 5,040COLD FUEL GAS BLOWER (1) 3,600REFRIGERANT PRESSURE VESSELS (18) 10,000

---------TOTAL, LIQUEFACTION SECTION 289,010

LPG FRACTIONATION (SECTION 400) 15,000

HEAT RECOVERY/STEAM GENERATION (SECTION 500)HEAT RECOVERY/STEAM GENERATORS (HRSG) (2) 20,000HOT BY-PASS STACK (1) 1.50AUXILIARY BOILER, PACKAGE UNIT (1) 5,000WATER TREATING DEAERATORS, STACK, BOILER FEED WATER PUMPS 5,500POWER GENERATORS (2), INCLUDING CONDENSORS 15,000HEAT REJECTION REFRIGERATION 18,000

---------TOTAL, HEAT RECOVERY/STEAM GENERATION SECTION 65,000

BATTERY LIMITS, INSTALLED 471,160CONTINGENCY, 15% 70,670

---------BATTERY LIMITS INVESTMENT 541,830


Table 14 (Continued)LNG BY CONVENTIONAL CASCADE REFRIGERATION




PEP COST INDEX: 620

COST($1,000)------------

OFF-SITES, INSTALLEDSTORAGE (SECTION 600)LNG TANKS (2), DOUBLE CONTAINMENT 110,000LPG TANK (1) 5,000NATURAL GASOLINE TANK (1) 3,000ETHYLENE REFRIGERANT TANK (4) 4,000

PROPYLENE REFRIGERANT TANK (6) 6,000AMINE TANK (2) 1,000WATER STORAGE TANKS (2) 2,000LNG PRODUCT PUMPS (8) 12,000

---------TOTAL, STORAGE 143,000

MARINE SYSTEM (SECTION 700)JETTY/CAUSEWAY (4,500 FT) 70,000SEA WATER INTAKE SYSTEM INCLUDING CHLORINATION (2) 9,000CONCRETE PIPES, ONE SUCTION LINE, ONE DISCHARGE LINE (4,500 FT EACH) 20,000SEA WATER PUMPS, INSTALLED (10) 27,000LNG PIPES, 36 IN. I.D. (2) 4,500 FT EACH 6,000VAPOR RETURN LINE, 24 IN. I.D. (1) 4,500 FT 4,000POWER CABLES (2) 2,000LNG TANKER OFF GAS COMPRESSOR (1) 4,140

---------TOTAL, MARINE SYSTEM 142,140

UTILITIES & STORAGE-- 285,140

WASTE TREATMENT (SECTION 800)COLD FLARE LINE, 24 IN. I.D. (2) 2,000WARM FLARE LINE, 36 IN. I.D. (2) 2,000FLARE STACKS, 150 FT (2) 1,000

---------TOTAL WASTE TREATMENT 5,000

GENERAL SERVICES FACILITIES (SECTION 900)BALLAST WATER TANK AND TREATING (1) 3,000FIRE FIGHTING SYSTEM (FIRE RINGS, PUMPS AND WATER TANK) 5,000SEWAGE & DRAINAGE 2,000SECURITY & LIGHTING 2,000INSTRUMENT AIR 1,000


Table 14 (Concluded)LNG BY CONVENTIONAL CASCADE REFRIGERATION




PEP COST INDEX: 620

COST($1,000)------------

NITROGEN AND LIQUID NITROGEN SYSTEM (1) 1,500DIGITAL CONTROL SYSTEM, INCLUDING COMPUTER AND AUXILIARY 1,500CONTROL ROOM, BLAST PROOF (1) 2,000SUBSTATION (1) 1,500GENERAL BUILDINGS, SHOPS, ADMINISTRATION BUILDING, LAB 3,000ROADS AND PARKING 2,000TELECOMMUNICATIONS 2,000GENERAL ALLOWANCE FOR OTHER ITEMS 5,500

---------TOTAL GENERAL SERVICES FACILITIES 32,000

---------TOTAL OFFSITES 322,140


Production Cost

Table 15 shows the estimated production cost of LNG product including feedstock and by-product prices and utilities. The feedstock and product values and utility costs are based uponPEP Yearbook 2002 prices for the U.S. Gulf Coast. All production costs are expressed as ¢/lb ofisooctane 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 gaswhere 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 forelectricity 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 plantoperators). Total U.S. operating labor compensation, including benefits, is $37.27/hr each. Inlikely locations for a LNG plant, the prevailing cost of labor is considerably lower. Maintenancelabor is estimated as 2.5%/yr of the battery limits investment. Since online controlinstrumentation is assumed, control laboratory labor is estimated as 20% of the direct operatinglabor cost. Total labor cost is estimated to be 0.28 ¢/lb. In the U.S., the staff is estimated to be92 employees (Table 16). The combined cycle mode as suggested may have increased thestaffing by about 20-25% over the common practice when using exclusively gas turbines.



PRODUCTION COSTS

PEP COST INDEX: 620

VARIABLE COSTSCONSUMPTION

UNIT COST PER LB ¢/LB------------------------- -------------------------- ----------

RAW MATERIALSRAW NATURAL GAS 1.735 ¢/LB 1.11684 LB 1.94

---------GROSS RAW MATERIALS 1.94

BY-PRODUCTSLPG 9.62 ¢/LB -0.014014 LB -0.13NATURAL GASOLINE 10.8 ¢/LB -0.013149 LB -0.14

---------TOTAL BY-PRODUCTS -0.27

CONSUMPTION CONSUMPTIONUNIT COST PER LB PER KG

------------------------- -------------------------- --------------------------UTILITIES

ELECTRICITY -0.03---------

TOTAL UTILITIES -0.03


Table 15 (Concluded)LNG BY CONVENTIONAL CASCADE REFRIGERATION

PRODUCTION COSTSPEP COST INDEX: 620

RAW GAS COST ($/MM BTU HHV) 0.50 0.75# 1.00

CAPACITY (MILLION LB/YR)* 9,646 9,646 9,646------------ ------------ ------------

INVESTMENT ($ MILLIONS)BATTERY LIMITS (BLI) 912.3 912.3 912.3OFFSITES NEGL NEGL NEGL

--------- --------- ---------TOTAL FIXED CAPITAL (TFC) 912.3 912.3 912.3

PRODUCTION COSTS (¢/LB)

RAW MATERIALS 1.29 1.94 2.58BY-PRODUCTS -0.27 -0.27 -0.27UTILITIES -0.03 -0.03 -0.03

--------- --------- ---------VARIABLE COSTS 0.99 1.64 2.28

OPERATING LABOR, 9/SHIFT, $37.27/HR 0.03 0.03 0.03MAINTENANCE LABOR, 2.5%/YR OF BLI 0.24 0.24 0.24CONTROL LAB LABOR, 20% OF OPER LABOR 0.01 0.01 0.01

--------- --------- ---------LABOR COSTS 0.28 0.28 0.28

MAINTENANCE MATERIALS, 3%/YR OF BLI 0.28 0.28 0.28OPERATING SUPPLIES, 10% OF OPER LABOR NEGL NEGL NEGL

--------- --------- ---------TOTAL DIRECT COSTS 1.55 2.20 2.84

PLANT OVERHEAD, 80% OF LABOR COSTS 0.22 0.22 0.22TAXES AND INSURANCE, 2%/YR OF TFC 0.19 0.19 0.19

--------- --------- ---------PLANT CASH COSTS 1.96 2.61 3.25

DEPRECIATION, 10%/YR OF TFC 0.95 0.95 0.95--------- --------- ---------

PLANT GATE COSTS 2.91 3.56 4.20

G&A, SALES, RESEARCH 0.06 0.06 0.06--------- --------- ---------

NET PRODUCTION COST 2.97 3.62 4.26

ROI BEFORE TAXES, 33.5, 26.6, or 19.8^%/YR OF TFC 3.15 2.50 1.86--------- --------- ---------

PRODUCT VALUE 6.12 6.12 6.12

-----------------------------------* OF LNG# BASE CASE^ RESPECTIVELY


Table 16U.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 ofEngineering, Head of Administration 4

Total 92

Maintenance materials are estimated at 3%/yr of the battery limits investment; operatingsupplies at 10% of operating labor costs. Adding maintenance materials and supplies to thevariable materials and labor costs produces a total direct operating cost of 2.20 ¢/lb of LNGproduct.

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 aplant cash cost of 2.61 ¢/lb. Depreciation further adds 0.95 ¢/lb to bring the plant gate cost to3.56 ¢/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 investmentcosts. LNG or natural gas is a mature commodity sold through well established networks incompetitive markets. With G&S, sales and R&D expenses, the net production cost becomes3.62 ¢/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 valueis lower than the $2.60/MM Btu HHV calculated at $0.75/MM Btu HHV raw gas cost from reportedinformation because we have credit for electricity and our capital investment is lower than otherdesigns [6].

At $0.50/MM Btu HHV as the cost of the raw natural gas fresh feedstock, net production costdrops 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 tojump 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 BtuHHV.


Profitability

As mentioned in the Product Cost section above, the net production cost of the LNG in theU.S. is estimated to be 3.62 ¢/lb or $1.54/MM Btu HHV. To determine a value at the productionplant, transportation and the cost of re-gasification are subtracted from the $3.50/MM Btu HHVU.S. value of natural gas. An allowance for about 2,000 miles of transportation with large tankersis assumed at $0.55/MM Btu HHV. A re-gasification cost of $0.35/MM Btu HHV is assumed. Theprice 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 basecapacity plant, just greater than the 25% ROI value frequently used to screen potential projects inthe petroleum or chemicals industries. However, the LNG business has some characteristics of autility, for instance, long term sales contracts and a dedicated supply of raw gas. Utility returnson investment are lower since these market risks are deemed to be lower. A lower ROI may thusbe acceptable. A lower return may happen due to shipping longer distances or higher re-gasification costs than assumed.

At $0.50/MM Btu HHV as the cost of the raw natural gas fresh feedstock, ROI jumps to avery 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.


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 SpringNational Meeting, paper 54e (Feb. 2001)

7. Cambridge Energy Research Associates, Inc., “CERA’s LNG Quarterly Review, Betweenthe 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 (June30, 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 liquefactionplants,” Oil Gas J., 101, 32 (Aug. 18, 2003) 50-54

Patents

US 4680041 DeLong, B. W., (to Phillips Petroleum Company), “Method for cooling normallygaseous material,” US Patent 4,680,041 (Jul. 14, 1987)

US 5611216 Low, W. R., (not assigned), “Method of load distribution in a cascaded refrigerationprocess,” US Patent 5,611,216 (Mar. 18, 1997)

US 5669234 Houser, C. G., et al., (to Phillips Petroleum Company), “Efficiency improvement ofopen-cycle cascaded refrigeration process,” US Patent 5,669,234 (Sep. 23, 1997)

US 6016665 Cole, E. T., et al., (to Exxon Production Research Company), “Cascaderefrigeration 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 streamrich in hydrocarbons,” US Patent 6,253,574 (Jul. 3, 2001)


PEP Publications

PEP Report29G

Nielsen, Richard. “Ethylene Plant Enhancement,” PEP Report 29G, SRIConsulting, Menlo Park, CA (April 2001)

PEP Review91-1-4

Ma, James J. L. “Removal of Mercury from Ethylene Plant Feedstock andCracked Gas Streams,” PEP Review 91-1-4, SRI Consulting, Menlo Park, CA(July 1992)

PEP Review2002-7

Cesar, Marcos A. “Liquefied Natural Gas by the Mixed Fluid CascadeProcess, PEP Review 2002-7, SRI Consulting, Menlo Park, CA (November2003)

PEPYearbook

Wang, Shao-Hwa (Sean), ed. “PEP Yearbook International 2002”, Vol. 1E,SRI Consulting, Menlo Park, CA (2002)

V-101Water Knock

Out Drum

T-101Amine

Surge Drum

C-101CO2

Absorber

V-102CO Absorber2

Overhead Drum

V-103Fuel GasSeparator

C-102CO Stripper2

F-201Dryer Regeneration

Furnace

V-201Condensate

Drum

C-201 A,B,CMole Sieve

Dryer

C-202 A,B MercuryAdsorber

4

5

2

1

3

Natural GasO17 C

650 psig

O60 C

O46 C

O23 C

O50 C

V-101

V-102

C-101

F-201

Condensate to Amine Unit

Mole Sieve Drying

O54 CCO (g)2

To Vent

Natural Gasto E-304

Water

SteamSteam

To Fuel Gas

O38 C100 psia

From AmineSump

T-101

E-108

Water

E-101

E-102

E-201

V-201

C-201 A,B,CC-202 A,B

V-103

E-103 E-104 E-105

C-102

E-107

E-106

200349

PEP Report 2003-15

O36 C

O122 C

SECTION 100: AMINE UNIT SECTION 200: DRYERS & MERCURY REMOVAL

LNG BY CASCADE REFRIGERATION PROCESSTWO TRAINS (One Shown)

Figure 1 (Sheet 1 of 5)

M-301Screen

G-301Chlorination

Unit

M-302Air

Filter

K-301Start-up/PeakingTurbine

K-302Gas

Turbine

K-303Ethylene

Compressor

K-304Propylene

Compressor

V-301High Pressure

Propylene FlashDrum

K-305ElectricityGenerator

V-302Mid Pressure

Propylene FlashDrum

V-303Low Pressure

Propylene FlashDrum

V-304Raw LPGSeparator

3

Steam 915 psig, 482OC

Fuel Gas 322 psia

OAir 2,155,000 lb/H 21 C

OSteam, 915 psia, 382 C

Combustor

4 in. Water-g

Air Compressor

Condensate to V-501, Sh. 5

From E-327

Exhaust Gas to Heat Recovery System, Sh. 5

K- 305

Uncontrolled Extraction

Steam to Amine & LPG Units

To Condensate/SteamDrum V-501, Sh. 5

E-302 A-L

O17 C137 psia

-4OC74 psia

O-32 C28 psia

Natural Gas toCold Box, Sh. 3

V-304

O-32 C594 psia

Raw LPG toLPG Unit, Sh. 4

O23 C610 psia

O-1 CTreated Dry Natural Gas

E-301

3 2 12

M-301

G-301

M-302

K-301

K-302

K-303

K-304

V-301

V-302

V-303

20

22

25

6

7

Speed Reducer

HotExpander

E-323Sh. 3

E-325Sh. 3

E-304 E-305

Fuel Gas fromCold Box, Sh. 3

To E-327,Sh. 3

35OC137 psia 6OC

72 psia-32OC26 psia

E-303

SeaE-324Sh. 3

E-326Sh. 3

Fuel Gas to K-309, Sh. 3

O1482 C

60OC220 psia

35OC

Speed Reducer

23

24

23

21

O-31 C

E-306

200350

PEP Report 2003-15

SECTION 300: LIQUEFACTION

13



K-303Ethylene

Compressor(See Sh. 2)

V-305HighPressure

EthyleneFlash Drum

V-306Mid Pressure

EthyleneFlash Drum

V-307Low Pressure

EthyleneFlash Drum

K-306Methane

CompressorSteam Turbine

K-307Methane

Compressor

V-308High PressureMethane Flash

Drum

V-309Mid Pressure

Methane FlashDrum

V-310Low PressureMethane Flash

Drum

K-308Fuel Gas

Cold Blower

K-309Fuel Gas

Compressor

V-311LNG Flash

Drum

T-601LNG StorageTank (Offsite)

K-309

6

18

19

12

11

10

8 9

30 34

31

37

32

33

35

36

38

41

44

47

4243

46

45

25

O-4 C

C =Refrigerant3

300 psia

O-32 CC =Refrigerant3

O-43 C

E-325

O-57 C120 psia

O-28 C295 psia

V-305

E-323

26 psia

V-106O-74 C

64 psia

V-307O-92 C

27 psia

O-90 C

530 psia

O-21 C

540 psia

Stage 3 Stage 2

Stage 1K-107

O482 C, Steam

900 psia

Sea Water

E-327

To V-501, Sh. 5

E-318

E-309 E-311

E-310

E-314

594 psia

O-32 C

From V-304Sh. 2

E-307

E-315

O-128 C

V-308

O-109 C

273 psia

V-309

V-310

O-153 C28 psia

E-308

C =Refrigerant3 K-303, Sh.2

E-317

Fuel Gas

From Tankers

O-134 C

28 psia

LNG to Loading

K-308

V-311

O-157 C

19 psia

O24 C

O-3 C

58 psia

O-92 C24 psia

O-37 C61 psia

O-73 C O-51 C

O-51 C

O-34 C

To E-306Sh. 2

From E-306Sh. 2

O-33 C

O-76 C O-93 C

E-312O-91 C

O-90 CO-72 C

O-113 C

O-108 C

E-313

E-319 E-320

E-321

E-316

O-152 C

O-109 C

299 psia

O-129 C

113 psia

O-3 C

105 psiaO-23 C

245 psia

O-75 CE-326

E-324

C =Refrigerant3 O-4 C

T-601

M

40

NNF

16 psia

Cold Box

O24 C127 psia

O-41 C117 psia

O-3 C

O-101 C24 psia

200351

PEP Report 2003-15

SECTION 300: (Concluded)

O-55 C O-72 C

O38 C



O-81 C

V-401Raw LPG

Flash Drum

C-401Deethanizer

V-402DeethanizerReflux Drum

C-402Debutanizer

V-403DebutanizerReflux Drum

7

13

15

V-401

C3

O-32 C544 psia

SteamO149 C

14

O149 C

Fuel Gas

Natural Gasoline

C C43

LPG

V-403

E-402

Steam

To VaporRecovery, NNF

C Rich Distillate to2

Fuel Gas, NNF

O-32 CRef.

E-403

E-406

C-401 C-402

Raw LPGfrom V-304

Sh. 2

O-19 C440 psia

E-401

V-402

450 psia

O133 C 50 psig

O49 C

16

O35 C

E-404

50 psig

E-405

O-30 C435 psia

200352

PEP Report 2003-15

SECTION 400: LPG FRACTIONATION

O49 C

154 psia

LNG BY CASCADE REFRIGERATION PROCESSONE TRAIN


V-501Condensate

Drum

V-502HRSG Steam

Drum

G-501Heat Recovery

Steam Generator

G-502Demineralizer

or Desaltation Unit

F-501Auxiliary

Boiler

V-504Auxiliary

Boiler SteamDrum

K-501Air

Blower

V-503Deaerator

M-501Stack

K-501

Steam Condensate

Exhaust Gas fromGas Turbine K-302, Sh. 2

O40 C

12 in Water-g

Make up Water49,726 lb/H

O313 C

Steam 900 psig, 482OC250,000 lb/H

Vent

100 psig

Flue Gas

O182 C Flue Gas

Steam 900 psig, 482OC250,000 lb/H., Max.

Fuel Gas

Air

V-502

O549 C

V-501

G-501

F-501

V-504

V-503

M-501

E-501 E-502 E-503 E-504

1115 psia

Boiler FeedWater

G-502 O149 C

200353

PEP Report 2003-15

SECTION 500: HEAT RECOVERY SYSTEM

O498 C

O282 C



Steam

PEP Review 2003-15 BASE LOAD LNG BY CASCADE … LOAD LNG BY CASCADE REFRIGERATION By David Netzer...

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Transcript of PEP Review 2003-15 BASE LOAD LNG BY CASCADE … LOAD LNG BY CASCADE REFRIGERATION By David Netzer...