chen4530urs2report

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YOTTA DESIGNS, INC.

Natural Gasoline Expansion

Natural Gasoline to LPG and Sales Gas

Curtis Edwards, Michael Polmear, Mark Colbenson CHEN 4530: Senior Design

Professor Clough

Mr. Sean Arendell – URS

5/5/2010

Wellhead

Gas

Outlet

Water

Outlet

Water

Oil

Oil

Inlet

Divertor

Mist

Extractor

Oil

Outlet

Baffle

Liquid

Level

Control

Figure 1 Three Phase Inlet Separator to Initiate the Magic (1).

Yotta Designs CHEN 4530 Senior Design Project May 5, 2010

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Table of Contents

Executive Summary .........................................................................................................................6

Project Description and Scope ........................................................................................................7

Problem Statement ......................................................................................................................7

Scope ...........................................................................................................................................8

Design Criteria..............................................................................................................................9

Wellhead Conditions .................................................................................................................9

Wellhead Flow Rate to Facility ............................................................................................... 11

Gas Re-Injection ..................................................................................................................... 11

Pipeline Gas Production Specifications ................................................................................. 12

Minimum Air Temperature Constraints .................................................................................. 12

Product Specifications ............................................................................................................ 12

Economic Considerations ....................................................................................................... 13

Background Information................................................................................................................. 13

The Yamal Megaproject ............................................................................................................. 13

Definitions ................................................................................................................................... 15

Natural Gas Processing ............................................................................................................. 16

Alternatives to the Proposed Process........................................................................................ 18

Dehydration............................................................................................................................. 19

Heat Integration ...................................................................................................................... 21

Column Optimization .............................................................................................................. 22

Refrigeration Cycle ................................................................................................................. 22

Recycle Operators .................................................................................................................. 23

Safety, Environmental, and Health Considerations ...................................................................... 23

Plant Safety (16)......................................................................................................................... 23

Environmental Concerns (16) .................................................................................................... 24

MSDS Summaries...................................................................................................................... 25

Natural Gasoline (17).............................................................................................................. 25

Liquefied Petroleum Gas (18) ................................................................................................ 26

Natural Gas (19) ..................................................................................................................... 26

Triethylene Glycol (20) ........................................................................................................... 27


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Project Premises and Corresponding Simulation Parameters...................................................... 27

Design Assumptions .................................................................................................................. 28

Economic Assumptions.............................................................................................................. 29

Overall Process Flow Diagrams .................................................................................................... 29

Natural Gasoline Process Flow Diagrams ................................................................................. 29

Natural Gasoline Expansion Process Flow Diagram ................................................................ 31

Process Descriptions ..................................................................................................................... 33

Inlet Separation and Natural Gasoline Recovery ...................................................................... 33

Natural Gasoline PFD ............................................................................................................. 33

Natural Gasoline Expansion PFD........................................................................................... 33

Approach................................................................................................................................. 34

Triethylene Glycol Dehydration.................................................................................................. 36

Triethylene Glycol Dehydration PFD ...................................................................................... 36

Approach................................................................................................................................. 37

Propane Refrigeration Cycle ...................................................................................................... 38

Propane Refrigeration Cycle PFD .......................................................................................... 38

Approach................................................................................................................................. 39

Sales Gas and LPG Recovery ................................................................................................... 40

PFD ......................................................................................................................................... 40

Approach................................................................................................................................. 41

Material and Energy Balances....................................................................................................... 43

Material and Energy Balances ................................................................................................... 43

Natural Gasoline Process Balances....................................................................................... 43

Expansion Process Balances ................................................................................................. 45

Process Description & Equipment Specifications ......................................................................... 50

Distillation Columns.................................................................................................................... 51

Estimating Column Pressure and Condenser Type............................................................... 51

Calculating Number of Trays .................................................................................................. 53

Determining the Dimensions of the Distillation Columns ....................................................... 54

Distillation Column Costing..................................................................................................... 58

Flash Drums ............................................................................................................................... 59

Three-Phase Separator .......................................................................................................... 59


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Flash Drums............................................................................................................................ 61

Heat Exchangers ........................................................................................................................ 62

Design of the Heat Exchangers.............................................................................................. 63

Pumps ........................................................................................................................................ 67

Compressors .............................................................................................................................. 67

Valves ......................................................................................................................................... 70

Storage Tank .............................................................................................................................. 71

Utility Summary .............................................................................................................................. 72

Estimation of Capital Investment and Total Product Cost ............................................................ 77

Economic Premises ................................................................................................................... 77

Venture Guidance Appraisal .................................................................................................. 77

Variable Costs......................................................................................................................... 78

Fixed Costs ............................................................................................................................. 79

Cash Flow ............................................................................................................................... 79

Capital Investment ..................................................................................................................... 80

Cost Indices ............................................................................................................................ 80

Commodity Chemicals............................................................................................................ 81

Total Permanent Investment (TPI) ......................................................................................... 81

Working Capital (WC) ............................................................................................................. 91

Operating Cost ........................................................................................................................... 91

Variable Cost .......................................................................................................................... 92

Fixed Cost ............................................................................................................................... 95

Profitability Analysis ....................................................................................................................... 99

Profitability ................................................................................................................................ 100

Cost of Capital ...................................................................................................................... 100

Net Present Value................................................................................................................. 100

Internal Rate of Return ......................................................................................................... 100

Return on Investment ........................................................................................................... 100

Break-Even Point .................................................................................................................. 101

Benefit-Cost Ratio................................................................................................................. 101

Depreciation .......................................................................................................................... 101

Salvage Percent ................................................................................................................... 102


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Accounts Receivable ............................................................................................................ 102

Corporate Income Tax .......................................................................................................... 102

Cash Flow Analyses ............................................................................................................. 102

Sensitivity Analysis................................................................................................................... 112

Present ROI and IRR for a +/- 100% Variation in TPI ......................................................... 112

Present ROI and IRR for a +/- 100% Variation in Fixed Operating Cost ............................ 113

Conclusion ................................................................................................................................... 113

Bibliography ................................................................................................................................. 115

Appendix A: Acronyms................................................................................................................. 118

Appendix B: Chemical Information .............................................................................................. 120

LPG MSDS (18) ............................................................................................................................ 120

Natural Gas MSDS (19) ................................................................................................................ 124

Natural Gasoline MSDS ............................................................................................................... 130

Propane MSDS............................................................................................................................ 130

TEG MSDS (20)............................................................................................................................ 130

Appendix C: Engineering Calculations ........................................................................................ 137

Design ...................................................................................................................................... 137

Costing ..................................................................................................................................... 160

Natural Gasoline Process ..................................................................................................... 160

Natural Gasoline Expansion Plant Process ......................................................................... 164

Appendix D: Computer Process Modeling .................................................................................. 173

Aspen HYSYS .......................................................................................................................... 173

Appendix E: Economic Spreadsheets ......................................................................................... 174

Total Capital Investment .......................................................................................................... 174

Natural Gasoline Process ........................................................................................................ 175

Natural Gasoline Expansion Plant ........................................................................................... 181


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Executive Summary

Natural gas processing represents an opportunity to exploit natural resources to provide

energy sources for a profit. The Yamal Peninsula contains a significant and valuable reserve for

the Russian economy. The objective of this project was to design two processes that separate

natural gasoline from a natural gas wellhead. Whereas the first process re-injects the overhead

streams back into the well, the second process expands the initial process to negate re-injection

and instead separate the overhead products into sales gas and liquefied petroleum gas (LPG)

product streams. These products are nearly equal in value to natural gasoline, with selling

prices of $50/bbl and $55/bbl, respectively, as compared with natural gasoline at $80/bbl. Both

processes yield approximately 10,000 bpd of natural gasoline, a scaled-up value from the 2,500

bpd that were initially being produced.

Product specifications are stringent for safety and energetic quality purposes. Design

specifications ensure that machinery functions in the extreme climate encountered above the

Arctic Circle. All design and product specifications were met, including the mitigation of hydrate

formation. Plant safety and environmental considerations were characterized and deemed

achievable through diligent planning and adherence to local and federal laws.

The economics of the two processes were estimated for a 15-year plant lifetime with one

design year and two years of construction. In the following summaries, economic parameters of

the natural gasoline and re-injection process will precede those of the expansion process to

produce the additional product streams. The total permanent investments were $170,600k and

$36,600k, respectively. The expansion process equipment was designed to use the reinjection

process equipment, thus realizing significant savings in investment. The internal rates of return

with the aforementioned selling prices were 52% and 164%, respectively. The break-even

points were during the first year of operation, the startup year for both processes. The benefits-

cost ratios were 14 and 67, respectively. For these reasons, the expansion process presents

favorable profitability, assuming that the well capacity remains fruitful and the selling prices

remain competitive.

Several improvements to the process merit further consideration. Dehydration

technology exists to minimize the product water content if the specifications change upon

integration with the approaching pipeline. Each of the distillation columns contains a reboiler

and condenser that could be integrated into a heat exchange network that contains process


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streams in an effort to reduce utility costs. The execution of the current design may be improved

by tightening convergence tolerances on columns and recycle functions.

Project Description and Scope

A plethora of natural gas reserves exist on and offshore of the Yamal Peninsula in

Northern Siberia, Russia. One of the clients of URS was initially recovering natural gasoline

from three wells on this peninsula, though due to the absence of a natural gas pipeline in the

area, they were flaring the excess natural gas. This client planned to expand their facility to

quadruple their production of natural gasoline, and initiate the recovery of liquefied petroleum

gas (LPG) and natural gas upon the imminent arrival of a natural gas pipeline. However, due to

the volatile nature of the Russian economy and the high cost of energy the client desired to

accomplish this expansion with minimal capital investment (2).

Problem Statement

Prior to the facility expansion project the client was producing natural gasoline from the

three remote wells. These wells are located far above the Arctic Circle on the Yamal Peninsula

in Northwestern Russia. There is little infrastructure in place in the area, and the natives sustain

themselves by hunting and fishing. The facility was producing approximately 2,500 BPD of

natural gasoline and burning all of the excess gas.

The expansions made to the facility were to consist of two phases. In the first, the

production of natural gasoline was to be increased to 10,000 BPD and the excess gas was to be

re-injected back into the reservoir instead of flared. As a natural gas pipeline was being routed

to the area and was to be in place within five years after the initiation of the expansion, the

second phase was to consist of the modification of the existing facility to include extra

processing equipment to separate and produce LPG and pipeline-quality natural “sales” gas

from the previously re-injected gas.

The remote location of the facility necessitated the consideration of several additional

factors that affected the design of the gas processing plant modifications. The most notable of

these was the extremely cold winter temperatures experienced at the location of the plant. The

wellheads are located in a region where permafrost exists, so the wellhead pipelines were

routed aboveground on piers to avoid degradation of the permafrost, and the equipment had to

be designed to a -60 °F design temperature. Also, due to the fact that electric power was


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unavailable at the site, natural gas engine-driven compressors were to be utilized for gas re-

injection. These were provided by Caterpillar, who has a presence in Siberia (2).

Scope

The first phase of the project was to involve the design of equipment to separate the

hydrocarbon liquids, water, and natural gas in the feed stream, to stabilize the hydrocarbon

liquids to shipping specifications, and to re-inject the residue gas back into the reservoir as

illustrated by the block flow diagram shown in Figure 2.

Inlet

Separation

C5+

RecoveryHydrocarbons

Water

Overheads

Wellhead

Compression Re-Injection

C5+ $

-$

-$

Figure 2. Block flow diagram for natural gasoline production facility

This process is hereafter to be referred to as the natural gasoline process.

Subsequent to the arrival of the natural gas pipeline, the plant modifications were to

include equipment to dehydrate the gas, reduce its hydrocarbon dew point, and compress it for

delivery to the pipeline. From the initial plant configuration the design was to be made easily

convertible to production of pipeline-quality gas. The block flow diagram for the facility following

the second plant modification, hereafter to be referred to as the natural gasoline expansion

process or simply the expansion process is shown in Figure 3.


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Inlet

Separation

C5+

Recovery

Hydrocarbons

Water

Overheads

Wellhead

Dehydration

C5+

Dry Gas

Water

LPG

Recovery

Sales Gas

LPG

C5+ Recycle

$

$

$

TEG

Compression

-$

-$

Refrigeration

HX

Cooled

Dry Gas

Propane

Figure 3. Block flow diagram for natural gasoline, sales gas, and LPG production (natural gasoline expansion

process)

The reduction in hydrocarbon dew point was to take place within the LPG recovery separation

train.

Additional requirements of the project included identifying and evaluating process

alternatives, identifying all the assumptions necessary for the design, and delineating the

requirements for the storage, shipping, and utility systems (2). Taken together, these factors

yielded a profitability analysis that favors expanding the process to produce sales gas and LPG.

Design Criteria

The following criteria and specifications were provided for the gas processing facility

expansion (2).

Wellhead Conditions

The conditions at the wellhead and inlet of the process are given in Table 1.


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Table 1. Process inlet conditions

Inlet Conditions

Wellhead Pressure (bar g)

150

Wellhead Temperature Range (°F)

20 - 50

Process Inlet Pressure (bar g)

103

The wellhead compositions are listed in Table 2.

Table 2. Composition of the wellhead stream (excluding water)

Composition

Component Mole %

N2 0.405

CO2 0.305

CH4 86.121

C2H6 6.637

C3H8 2.484

iC4H10 0.359

nC4H10 0.415

C5+ 3.274

COS/CS2 0.0

H2S 0.0

For the 3.274 mol % of the wellhead stream that is composed of C5+, ASTM D86 data

were given, which is shown in Table 3. ASTM D86 is a standard distillation-based assay used to

characterize petroleum, in which the temperature is recorded at which successive fractions of

the oil mixture have evaporated.


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Table 3. ASTM D86 oil characterization assay data of C5+ at wellhead

ASTM D86 Assay

Percent Evaporated (% Liquid Volume)

Temperature (°C)

Bubble Point 38.7

10 64.3

20 83.7

30 99.7

40 112

50 129

60 148.7

70 171.3

80 214.7

90 281

End Point 295

The data in Table 4 were also given for the C5+ content at the wellhead.

Table 4. Physical data for the C5+ at the wellhead

Natural Gasoline (C5+)

Density (g/cm3) 0.731

Average MW (Da) 101

Wellhead Flow Rate to Facility

The gas flow rate from the wellhead was to be determined based on a natural gasoline

(C5+) standard production rate of approximately 10,000 BPD. The flow rate of water in the

wellhead was to be found based on 1.5 bbl of water produced for every MMSCF of gas flow

from the inlet separator.

Gas Re-Injection

Before the plant modifications for LPG and natural gas recovery were implemented, in

the first phase of the facility expansions, gas was to be re-injected into the well at a pressure of

180 bar g.


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Pipeline Gas Production Specifications

After the facility expansion for the production of LPG and sales gas, in the second phase

of the project, the sales gas was specified to have a hydrocarbon dew point of not more than 0

°F at a pressure of 55 bar g.

Minimum Air Temperature Constraints

Owing to the frigid climate of the Siberian Yamal Peninsula the process and mechanical

design was to allow for a minimum air temperature of -60 °F. Furthermore, all equipment was to

be designed from low-temperature carbon steel impact tested to -65 °F. And finally, any air

coolers were to be designed for the use of air at 85 °F to account for the hottest ambient air

temperature that would likely be reached during the summer months.

Product Specifications

The criteria for the purity and production rate of natural gasoline were based on Reid

Vapor Pressure (RVP), a common measure of purity in the natural gas processing industry, and

standard flow rates, as illustrated in Table 5.

Table 5. Natural gasoline purity and flow rate product specifications

Natural Gasoline (C5+)

Max Reid Vapor Pressure (RVP in psia) 10

Min Standard Flow rate (BPD) 8,000

Design Standard Flow rate (BPD) 10,000

Max Standard Flow rate (BPD) 11,000

The LPG product was to be characterized as having a True Vapor Pressure (TVP) of 210 psia

at 100 °F, and a C5+ content of no more than 2.0% by volume.

Finally, the sales gas product specifications were to be based on Hydrocarbon (HC) dew

point, the temperature at which hydrocarbons begin condensing out of the gas, as well as the

CO2 and H2O content of the gas as shown in Table 6.


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Table 6. Sales gas purity product specifications

Sales Gas (Natural Gas)

Max Hydrocarbon Dew Point (°F)

0

Max CO2 Content (Mole %)

2.0

Max H2O Content (lbs/MMSCF)

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Economic Considerations

The sales prices of each of the three products were given and tabulated in Table 7.

Table 7. Sales prices of products

Product Sales Prices

Sales Gas $50/bbl $1.19/US gal

LPG $55/bbl $1.31/US gal

Natural Gasoline $80/bbl $1.90/US gal

Also, the operating cost of gas re-injection was given as $1.50/1000 SCF, and the future gas

sales price was given as $4/MMBTU.

Background Information

In order to better understand the expansions that were to be made to the operating

natural gas processing facility it was useful to situate this expansion within the Russian

economic and political climate in the area, to research natural gas processing and the unit

operations that are utilized in the industry and the proposed process, and several processing

alternatives.

The Yamal Megaproject

Owing to the fact that the gas reservoir to be modified is located on the Yamal Peninsula

in Northern Siberia, Russian Federation, information on the nature of Russian gas reserves and

the economic and political climate in which the project will take place are relevant concerns.


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Russia currently supplies one quarter of Europe’s natural gas, and plans to expand in this

market (3). However, over the last several years Russia has lost credibility with many of the

former Soviet Union countries and Europe as a reliable supplier of natural gas. In January of

2006 Russia cut off the gas supply to Ukraine and Moldova, and in late 2006 threatened to cut

supplies to Belarus and Georgia over pricing disputes (4). Then, again, in January of 2009,

fueled by ongoing political tensions between the two countries, Russia curtailed the flow of gas

through Ukraine during a particularly cold period of the winter, affecting in particular the Balkans

and Eastern Europe (3). These and other similar instances have encouraged some countries to

seek other sources of natural gas, and have incited criticism of Russia as using energy as a

political tool (4). However, petroleum and natural gas are vital to the Russian economy, and

Russia plans to further develop and expand the industry largely through the utilization of the

Yamal Peninsula. There are plans to increase production of natural gas from the peninsula by a

factor of almost 42 from 2011 to 2030, a reflection of the vast gas reserves present in the area

(5).

Russia owns approximately one third of the world’s gas reserves, which according to the

International Energy Agency consisted of 46.9 trillion cubic meters (tcm) of proven and probable

reserves at the beginning of 2001 (6). In the North Siberian Yamal Peninsula and adjacent

areas 11 gas and 15 oil, gas, and condensate fields have been discovered which contain

approximately 16 tcm of gas according to exploration and preliminary estimates. These fields

have even been projected to contain as much as 22 tcm of gas reserves (5). The majority of

these fields are owned and licensed to the corporation Gazprom, which is pursuing their

development under the ‘Yamal Megaproject.’

Gazprom is “one of the world’s largest energy companies,” and holds a monopoly in the

Russian gas market (5) (3). It is owned largely by the Russian government, though is a

privatized company which specializes in geological exploration, the production, transportation,

storage, processing, and marketing of hydrocarbons, and the marketing of heat and electric

power (5). In a 2007 initiative, Gazprom, in collaboration with the Yamal-Nenets Autonomous

Okrug (YaNAO) Administration, amended a 2002 draft program for the development of the

peninsula’s gas reserves with the aim of the expansion of the reserve fields and the construction

of gas pipelines on the Yamal Peninsula. The initiative plans for the launch of drilling at several

of the fields, the further development of the production capacities of the existing operational

fields, and the construction of a 2,500 km gas pipeline system. The company purports to be

taking into consideration the myriad environmental and social responsibility issues that the


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project warrants including pollution concerns and the welfare of the indigenous people and

wildlife. Due to the fact that the Yamal Peninsula is the most explored region for gas

production, is located near existing gas pipelines, and has significant reserves, Gazprom

regards the Yamal Megaproject as central to the development of the Russian economy (5).

Definitions

There are three product streams from the proposed process, these being natural

gasoline, liquefied petroleum gas, and natural gas. Each of these distinct products requires

definition.

Natural Gasoline (C5+) – Natural gasoline is a liquid product consisting of pentane, and

all of the hydrocarbons heavier than pentane. For the purposes of this project the purity

of this product stream is defined by a Reid Vapor Pressure of 10 psia at 100 °F, and its

temperature was to be as close to an upper limit of 400 °F as possible to meet common

shipping requirements (2).

Liquefied Petroleum Gas (LPG) – LPG is a liquid product consisting primarily of propane,

n-butane, and isobutane. The purity specifications for this product are a maximum True

Vapor Pressure of 210 psia at 100 °F and a C5+ content of no more than 2.0 % by

volume (2).

Natural Gas (Sales Gas) – Natural gas must meet certain quality specifications before

injection into a pipeline to ensure that the pipeline operates properly. Gas that does not

meet specification can lead to deleterious hydrate formation, operational problems in the

pipeline, pipeline deterioration, or even pipeline rupture (7). These quality measures

often include specifications on the energy content of the gas per volume, its hydrocarbon

dew point temperature, maximum levels of contaminants such as hydrogen sulfide,

carbon dioxide, nitrogen, water vapor, and oxygen, and maximum amounts of particulate

solids and liquid water, as these can damage the pipeline (7). For the expansion of the

relevant natural gas processing plant the quality measures required to be met included a

hydrocarbon dew point of no more than 0 °F, a CO2 content of no more than 2.0 mol %,

and a maximum water content of 4 lbs/MMSCF of gas (2).


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Natural Gas Processing

Natural gas exists in a variety of forms, its composition depending on the type, depth,

and location of the deposit as well as the geology of the area in which it is tapped. Oil and

natural gas are often found in the same reservoir, and natural gas is classified as associated-

dissolved when dissolved in crude oil, or non-associated when it occurs in the absence of oil.

The relevant raw gas being gathered for the proposed process is non-associated. The raw gas

most often consists of two to eight carbon hydrocarbons that are gaseous at underground

pressures, though condense to liquid at atmospheric pressure. These liquids are called

condensates or natural gas liquids (NGLs). The recovery of NGLs can involve any of several

initial processing steps depending on the particular composition of the well (7).

Due to the myriad possible compositions of the raw gas, these initial processing steps

can be quite complex. Producing areas can contain hundreds of wells, from which gas and

NGL is “gathered” via small-diameter pipes that connect the well to processing facilities. At the

wellhead the gas is often put through scrubbers to remove sand and any particulate matter

and/or heaters to ensure that the temperature does not drop low enough for hydrates to form in

the stream.

Hydrates are crystalline, ice-like solids that form with the water vapor in the stream, and

can pose serious risks to the process, as they have the potential to clog the valves and pipes

that the gas passes through during processing, thus leading to dead-heads. They form within a

certain temperature/pressure envelope, oftentimes above the freezing point of water, the limits

of which are dependent on the composition of the stream. Therefore, the avoidance of hydrate

formation was a concern in the design of the proposed process.

The various streams gathered at a given site can require differing initial processing steps

including heating, compression, scrubbing, carbon dioxide removal, and sulfur removal,

contributing to the complexity of the gathering process. After these steps are taken, the further

processing steps that are commonly performed include, but are not limited to the following (7).

Gas-Oil Separation

When natural gas is associated with crude oil it is first necessary to separate the gas

from the oil. Oftentimes pressure relief at the wellhead alone accomplishes this separation, and

just a simple closed tank is required. However, sometimes a multi-stage separation train is

required, in which a series of cylindrical shell, horizontal tanks are commonly utilized. These


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include an inlet at one end, and top and bottom outlets for gas and oil respectively. Separation

is effected by compressing and expanding the feed between tanks, causing heating and cooling

of the stream (7). As mentioned previously, the wellhead stream for the proposed process

consists of non-associated gas, eliminating the need for this step.

Condensate Separation

Condensate separation is most often accomplished through the use of mechanical

separators. It is at times preceded by a slug catcher to remove any free water from the

wellhead stream, and is usually employed when gas-oil separation is not required (7). In the

proposed process this step is carried out by the three phase separator unit.

Dehydration

It is necessary to remove any free water from the natural gas stream to avoid the

formation of hydrates in the process. The most common method of dehydration, and the one

employed in the proposed process, is absorption of water by glycol, though a variety of other

processes have been used, several of which will be discussed in the Alternatives to the

Proposed Process section. Triethylene glycol is the most common type of glycol used for this

purpose (7).

Contaminant Removal

Contaminants that must be removed during processing include hydrogen sulfide and

other sulfur-containing compounds, carbon dioxide, water vapor, helium, and oxygen. To

remove sulfurous compounds flow is often directed through a tower containing a solution of

amines. The amines absorb sulfur compounds from the gas stream, and have the advantage of

being able to be used repeatedly. Desulfurization can then be followed by a series of filter tubes

where gravity, centrifugal force, and flocculation of particulates elicit the removal of other stream

contaminants (7). As the wellhead feed stream in the proposed process does not contain

sulfurous compounds, helium, or oxygen, and carbon dioxide is present at acceptable levels,

this step was not required.

Nitrogen Extraction

Nitrogen, the excessive presence of which can lower the energy content of the gas, is

most often removed from natural gas streams via a nitrogen rejection unit (NRU), which also

works to further dehydrate the gas using molecular sieve beds. Separation can occur through


18

the use of cryogenic methods, a column, and a brazed aluminum plate fin heat exchanger, or by

an absorbing solvent (7). Once again, since nitrogen was present at acceptable levels in the

wellhead inlet stream to the proposed process, this step was not needed.

Methane Separation

Methane is the primary component of sales gas, and can be separated from natural gas

streams either as part of the NRU unit or in a separate unit operation. If done separately, there

are two primary methods that are utilized for this purpose, these being cryogenic methods and

absorption. The cryogenic approach, which is better at extracting the lighter liquids in the

stream, such as ethane, is accomplished by lowering the temperature of the gas stream to

around -120 °F. This is often done through the use of a turbo expander in combination with

external refrigerants, and results in the condensation of all stream components besides

methane. The absorption method can be carried out by using absorption oil to absorb the

majority of the NGLs, which are subsequently distilled from the absorbing liquid oil (7). In the

proposed process the use of absorption oils was unnecessary, and methane separation was

achieved by the Sales Gas refluxed absorption column.

Fractionation

Fractionation is the process of separating the various NGLs by virtue of the differing

boiling points of the hydrocarbons in the stream. This is generally done through successive

distillation of the NGL stream, though to produce LPG in the proposed process just a single

distillation column was required (7).

Which of these steps are performed depends on the composition of the raw gas, and

multiple steps can be performed in a single unit operation, or at different locations (7).

Alternatives to the Proposed Process

In modeling the natural gasoline with re-injection and natural gasoline expansion

processes, a number of process alternatives were identified and considered. These included

both alternative processes to those utilized to simulate these two natural gas processing plant

configurations, as well as methods for eliminating various assumptions and simplifications that

were utilized. A description of each process alternative is outlined below.


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Dehydration

Glycol Fluids Package

The Peng-Robinson fluids package was utilized to model both the natural gasoline with

re-injection and natural gasoline expansion processes. This fluids package is ideal for the

modeling of non-polar hydrocarbons, of which the vast majority of the process streams consist;

however, as triethylene glycol (TEG) is hydrophilic, it is not always modeled well by this

package (8). It was therefore considered to use the glycol fluids package in Aspen HYSYS for

the TEG dehydration cycle (9). This matter was discussed with Mr. Arendell, though for the

conditions and purposes of this simulation it was determined that the modeling of this cycle by

the Peng-Robinson fluids package would be sufficient. Mr. Arendell noted that the simplifying

assumption that the Peng-Robinson package accurately modeled TEG dehydration could result

in an underestimate of the amount of water absorbed by the TEG. Nevertheless, the amount of

water in the final sales gas stream came out to be about half of the maximum given specification

of 4 lbs/MMSCF, which was deemed an appropriate tolerance for any extra water that may have

been present in the stream due to an underestimate in its removal during dehydration.

Drizo®

The effectiveness with which a TEG dehydration cycle removes water from a natural gas

stream is dependent on the purity of the glycol upon regeneration. The GPSA section 20 on

dehydration quotes an achievable TEG purity of 98.6 wt% by reboiling TEG at 400°F at

atmospheric pressure (8). However, the proposed process achieves a purity of 99.0 wt% TEG

by reboiling at just below 400°F and just above atmospheric pressure. This very small

discrepancy may be due to the fact that TEG is modeled by the Peng-Robinson fluids package

rather than the glycol package, as discussed in the Glycol Fluids Package section; though as

the quoted value is nearly reproduced, the use of the Peng-Robinson package is further

justified. Various enhanced glycol recovery processes exist, each of which is based on the

principle of reducing the effective partial pressure of water in the vapor space of the lean (water-

deficient) glycol stream, allowing for higher glycol concentrations to be obtained at the same

temperature (8). This results in a greater water dew point depression than can generally be

achieved. The Drizo® process is among these enhanced dehydration processes.

The process regenerates glycol by solvent stripping as opposed to the conventional gas

stripping that is ordinarily employed (10). The solvent is obtained from the natural gas itself,

and is composed of paraffinic and aromatic hydrocarbons (BTEX) that exhibit a C5+ boiling point


20

range (10) (8). In the regeneration column for the process, heavy hydrocarbons and water are

condensed from the overhead while non-condensable species are vented to the atmosphere

nearly free of BTEX. The condensed hydrocarbons are separated from the water, vaporized,

and superheated before being routed to the lean-glycol stripping column where they serve as

the stripping gas (8). This results in glycol purities of up to 99.998 wt% according to the

manufacturer, yielding water dew point depressions of upwards of 100 °C (10). The process

can even be supplemented with drying of the solvent by a solid desiccant, which can yield glycol

purities of as high as 99.999 wt% and water dew point depressions of 121 °C (8).

Aside from the fact that this process is exceedingly complicated to model, the extent of

glycol regeneration achieved and the water dew point depressions reached are unnecessary to

meet the sales gas purity specifications of 4 lbs/MMSCF of water and a hydrocarbon dew point

of 0 °F for the proposed process (11). While the Drizo® dehydration system seems well-suited

to applications with very stringent water removal criteria, its implementation in the proposed

process was deemed superfluous.

Coldfinger®

Another proprietary process for achieving enhanced glycol purities upon regeneration is

the Coldfinger® process. In this process, a bundle of condensing tubes (the cold finger), in

which rich TEG is commonly utilized as the coolant, is inserted into the vapor space of a surge

tank half full of lean TEG. The cold finger continuously condenses equilibrium water vapor,

which is discharged from the unit via a collecting trough placed beneath the finger. This

continuous condensation maintains the partial pressure of water in the vapor below its

equilibrium vapor pressure, which works to further draw water out of the lean TEG liquid phase.

The process results in glycol regeneration of upwards of 99.7 wt% TEG in the lean glycol

stream (8). Due to the fact that this process is not in equilibrium, while Apsen HYSYS models

all unit operations as if they were in equilibrium, the Coldfinger® process could not be easily

modeled using this software (11). And further, the extent of glycol regeneration achieved was

again deemed unnecessary to reach the water removal specifications required by the natural

gasoline expansion process.

Advanced Prism® Membranes

A fairly novel process for natural gas dehydration is membrane separation technology.

Advanced Prism® Membranes utilize the principle of selective gas permeation, in which the


21

driving force for separation is differing gas partial pressures on either side of a membrane, to

separate water from a natural gas stream (12). These units consist of bundles of hundreds of

thousands of hollow-fiber membranes enclosed in a pressure-rated casing. The gas to be

dehydrated is passed through this casing on the outside of the hollow fibers. Due to the faster

permeation rate of water through these membranes as compared with the hydrocarbons in the

stream, water diffuses through to the inside of the fibers, in which a lower pressure is

maintained. The many fibers provide a large area for membrane separation, resulting in

significant water removal from the stream. These separators can be arranged in parallel, in

series, or in a cascade fashion, and can yield gas streams of up to 98% purity (13). These units

have the potential to incur lower maintenance costs and operate with less downtime than

comparable dehydration units as there are no moving parts involved in the separation. In

addition, raw material costs can be lowered through the use of membrane separation as no

chemical inventory is required for their operation (12).

While membrane-based dehydration systems show much potential for the economic

dehydration of natural gas, this avenue was not pursued due to the impossibility of modeling

these units in Aspen HYSYS. However, this technology could provide a very viable option for

dehydration in similar natural gas recovery processes to the proposed processes.

Heat Integration

The implementation of a heat exchange network to minimize the utilization of process

utilities was investigated in designing the proposed processes. However, heat integration

among utility streams was determined to be infeasible given the thermal properties of these

streams. For instance, the chilled water that is used in the condensers of each column, after

being heated to 90 °F, no longer possesses enough of a cooling capacity to be used for any

other heat exchange processes. Similarly, the low and high pressure steam utilized in the

reboilers of each of the columns, after being condensed, no longer possesses enough of a

heating capacity to be used for any other heat exchange processes.

Any feasible heat integration that could be performed on the proposed processes would

have to involve process streams as opposed to solely utility streams. Though, due to

inexperience in the natural gas processing industry, this option was not considered. A possible

route for heat exchange would be to use the overhead from the sales gas column to cool the

inlet to this column as mentioned by Mr. Arendell. Another option brought up by Mr. Arendell

would be to use the hot C5+ product stream to run the reboiler of one of the other distillation


22

columns. While configurations like these have the potential to reduce the utility costs of the

plant, process stream heat integration greatly complicates the startup and control of the given

processes (14). For this reason a cost/benefit analysis would need to be performed in order to

determine if this type of heat integration would be economical without overcomplicating the

control systems for the processes.

A final process alternative relating to heat exchange that was considered was the

expansion of the propane refrigeration cycle to include cooling streams to each of the column

condensers. As the propane, which is recycled, would replace non-recycled cooling water

utilities in each of these condensers, this option may have provided an economic advantage to

the proposed natural gasoline expansion process, however, without a full economic analysis on

this expanded refrigeration cycle, its economic feasibility cannot be determined. Unfortunately,

time constraints did not permit proper investigation of this alternative.

Column Optimization

Due to the numerous variables involved in distillation and absorber column design there

are a plethora of alternate configurations that each of these could assume. For example,

columns could have different numbers of trays, pressures, inlet temperatures, etc. Nonetheless,

by adjusting column parameters such that the design specifications were met, the column

designs were optimized towards the designs that would actually be implemented in industry.

Further optimization was performed by changing the number of trays and the feed tray location

such that reboiler duties were minimized, thus minimizing the amount of heating utility required

to run the column. Finally, the feed tray was chosen based on matching the temperature of the

inlet stream to the inlet tray temperature as closely as possible. This provides for a smoother

temperature profile up the length of the column and allows for better control of column dynamics

upon disturbances (14). The design of each of the distillation and absorption columns was

honed throughout the project by conversations and parameters suggested by Mr. Arendell and

Professor Clough.

Refrigeration Cycle

The propane in the propane refrigeration cycle was modeled as pure propane; however,

this is not entirely accurate. In actuality, refrigeration-grade propane consists of 98% propane

and 2% ethane by weight (15). While this simplifying assumption may have resulted in the

modeling of the processes in this cycle slightly differently due to the different composition of this

exchange fluid, it was determined upon the discovery of the actual composition of refrigeration


23

propane that modeling it as pure propane would be sufficient for the proposed process.

However, to model the heat exchange more accurately in this aspect of the simulation, the

correct composition of refrigeration-grade propane would need to be used.

Recycle Operators

Two recycle operators were used in the design of the natural gasoline expansion

process, one in the TEG dehydration cycle and one that recycles C5+ back from the LPG

recovery column to the C5+ column. During the implementation of these operators into the

design, the performance of intermittent mass and energy balances was overlooked, and these

were performed only when the natural gas expansion process was complete. It was then

discovered that the sensitivities for mass and energy flows aligning with one another on either

side of the recycle operators were not tight enough, resulting in a 5.6 % discrepancy in the

energy balance on the process. However, in attempting to tighten the tolerances of the recycle

operators it was found that this was impossible with a complete process including two of these

operators, and Aspen HYSYS was unable to converge with tightened tolerances. While

unfortunate, this discrepancy does not invalidate the proposed process as the mass balance

was very nearly closed. Though in modeling similar processes, it is advisable to tighten the

tolerances for mass and energy flows on any recycle operators while the process is being

modeled, rather than at the end, when it is already complete.

Safety, Environmental, and Health Considerations

Natural gas processing presents a wide variety of safety, environmental, and health

considerations. These are reviewed herein.

Plant Safety (16)

In order to safely operate a natural gas processing plant it is of the utmost concern that

all industry safety standards and protocols are strictly adhered to. Basic safety measures such

as extensive personnel training for the operation of equipment containing flammable and

explosive hydrocarbons under high pressures and at high temperatures must be implemented.

All equipment should be preventatively maintained on a regular schedule. The control systems

of the plant should be optimized for safety, and backup and emergency shutdown systems

should be included for all major unit operations. Systems should be in place to monitor all

equipment for leaks, fluid levels, pressure and temperature, such that any irregularities will

quickly become evident before any situations become critically dangerous. Risk assessment


24

protocols should be in place to identify and evaluate any and all potential risks associated with

new or modified process equipment. Also, emergency response procedures should be in place

for any emergency that might occur.

Furthermore, the operation of a natural gas processing plant in particular requires that

concern be paid to various extra safety matters relating to flammability, fire, and explosions.

Again, all industry standard protocols regarding these risks should be followed with care.

Processes should be appropriately segregated from flammable product storage areas, or if this

cannot be achieved, blast walls should be implemented where they are necessary. The plant

should be designed such that potential ignition sources are avoided, such as the elimination of

fixtures that could leak flammable material onto or near heated piping or equipment. Finally the

specific dangers associated with each flammable material in the process should be known and

accounted for. For instance, pressurized flammable gasses can result in jet fires, while

flammable liquid spills can lead to pool fires. The specific dangers of the hydrocarbons and

chemicals employed in the proposed processes will further be outlined in the MSDS summary

section below.

Environmental Concerns (16)

There are several environmental concerns associated with natural gas processing

including fugitive emissions, gas flaring, and wastewater treatment, each of which will be

implemented in the proposed processes.

Gas release to the environment is common in natural gas processing. Fugitive gas

emissions to the environment can occur from leaks in piping, valves, flanges, or other process

connections. In addition, emissions can occur during the loading and unloading of any

hydrocarbon streams or products. These emissions, which can include greenhouse gasses,

can be minimized through the installation of monitoring systems as well as by the maintenance

of stable tank pressures and vapor spaces. Oftentimes flammable gasses are flared from

natural gas processes either for byproduct disposal or as a safety measure for emergencies.

The proposed natural gasoline expansion process includes a Flare Gas stream in the TEG

regeneration cycle. For the sake of safety as well as the prevention of the release of

greenhouse gasses to the atmosphere, this process should be carried out in the most controlled

manner possible.


25

Wastewater treatment is also a concern for streams that are contaminated with

hydrocarbons or other chemicals. In the proposed processes the water separated from the inlet

separator, Water 1, as well as the overhead to the TEG Regeneration column, Water 2, are

wastewater streams. These would be treated with an onsite wastewater treatment unit, such

that no contaminated liquids were released into the environment. In addition, all plants that deal

with wastewater treatment should include secondary containment basins with impervious

surfaces to further prevent the release of deleterious compounds into groundwater or soil.

MSDS Summaries

The pertinent points of the MSDS for each of the products and chemicals in the process

are given below. While this information provides many of the key relevant safety issues

involved with working with these chemicals, it DOES NOT substitute for the actual MSDSs.

These should be reviewed and kept in an accessible location at the plant.

It is to be noted that the hazards of propane will not be summarized, as propane is

present in significant quantity in LPG, and the hazards can therefore be assumed to be the

same as for LPG.

Natural Gasoline (17)

May contain benzene, cyclohexane, xylene, and/or toluene

o Can be carcinogenic due to presence of benzene

Clear, colorless liquid with a distinct hydrocarbon odor

Flash point: -45 °F

Extinguishing media: dry chemical, foam, carbon dioxide

Unusual fire and explosion hazards:

o Flames impinging on a product storage vessel above the liquid level can cause

vessel failure within nine minutes, resulting in a boiling liquid expanding vapor

explosion.

o Liquid product will change to vapor quickly at temperatures well below ambient

and form flammable mixtures with air.


26

o Vapors are heavier than air, and can travel long distances to an ignition source.

Inhalation risk to respiratory and central nervous systems potentially resulting in death

Frostbite can occur

Liquefied Petroleum Gas (18)

May contain propane, propylene, and/or butane

Clear, colorless gas

Flash point: -156 °F; Autoignition temperature: 842 °F

Extinguishing media: dry chemical, foam, carbon dioxide, water spray


o Containers of product may rupture upon exposure to heat or flame.

o Approach a flame-enveloped container only from the sides, and never from the

head ends.



Freeze burns can occur

Natural Gas (19)

May contain natural gas, benzene, and/or n-hexane

o Can be carcinogenic due to presence of benzene

Clear, colorless gas

Flash point: <100 °F

Extinguishing media: dry chemical, foam, carbon dioxide, water fog

o Do not use a direct stream of water to extinguish, as natural gas will float, and

can reignite on the surface of water.


27


o Do not weld, heat, or drill on or near containers of the product.

o Do not enter confined-space fire without full bunker gear including a face shield,

bunker coat, gloves, rubber boots, and a positive-pressure breathing apparatus.



Triethylene Glycol (20)

Clear, colorless liquid with no odor

Slightly flammable

Skin irritant

Flash point: 351 °F; Autoignition temperature: 700 °F

Can form explosive mixture with air above flash point

Extinguishing media: dry chemical, alcohol foam, carbon dioxide

o Water or foam may cause frothing

Project Premises and Corresponding Simulation Parameters

The design specifications were given by Mr. Arendell and URS Corporation. Table 8

outlines these specifications and the corresponding simulation parameters.


28

Table 8. Project specifications and corresponding simulation and design parameters.

Design Specification Design Value Current Design Value

Wellhead Flow Rates to Facility

Adjusted by design group to produce approximately 10,000 BPD of natural gasoline.

Adjust 1

Natural gasoline and re-injection process dry wellhead flow rate (bpd): 1.068E5 Natural gasoline expansion dry wellhead flow rate (bpd): 1.068E5

Produced Water Rate

1.5 bbl water per MMSCF of gas flow out of the inlet separator. Adjust 2

Natural gasoline and re-injection process wet wellhead and water production flow rate from inlet separator (bpd):

360.4 344.7

Natural gasoline expansion wet wellhead and water production flow rates from inlet separator (bpd):

360.4 344.7

Gas Re-Injection

Normal gas injection pressure (bar g): 180 180 Future Pipeline Gas Product Specification

Maximum hydrocarbon dew point (°F) at 55 barg: 0 -41.2

Future pipeline gas delivery pressure (bar g): 55 55 Product Specifications

Natural Gasoline (C5+ product)

Maximum Reid Vapor Pressure (RVP) (psia): 10 10 Design standard flow rate (bpd): 8,000-11,000

Re-Injection standard flow rate (bpd): 9,202 Expansion standard flow rate (bpd): 9,869

Liquefied Petroleum Gas Maximum true vapor pressure (TVP) (psia) at 100°F: 210 208.2

C5+ content (% volume maximum): 2.0 0.98 Future Residue (Sales) Gas

Maximum CO2 content (mole %): 2.0 0.0032 H2O content (lbs/MMSCF): 4 2.0

Maximum hydrocarbon dew point (°F): 0 -41.2

These results demonstrate that every design and product specification was met.

The process was developed using the following assumptions and specifications with

regard to design and economics.

Design Assumptions

Peng-Robinson equation of state is valid for the entire process

Ignore light ends in Oil Manager


29

Mole fractions of C5+ hypothetical components are distributed equally in the Dry Well

inlet material stream

The dehydration step effectively removes enough water to negate hydrate formation

Economic Assumptions

The plant is an expansion on an existing process that has produced 2,500 bpd of natural

gasoline for an unknown duration of time

Conservatively, the salvage value of the current process is unknown and assumed to be

negligible

The purchase of land is not required

The plant is to be constructed in the Yamal Peninsula, Siberia, Russia

Plant is operated 90% of the year for 7,884 operating hours

No royalties

Overall Process Flow Diagrams

Process flow diagrams are linear representations of the process. Both processes were

modeled in Aspen HYSYS V7.0 (9).

Natural Gasoline Process Flow Diagrams

The current plant in the Yamal Peninsula produces 2,500 bpd of natural gasoline. The

overhead products from inlet separation and C5+recovery are reinjected into the well,

representing a significant cost for compression power and significant profit loss for the

contained sales gas and LPG products. Figure 4 is the final Aspen HSYSYS process flow

diagram (PFD) diagram.

30

Figure 4. Final simulation of “current” natural gasoline production.

Material streams are blue, energy streams are red, and special controls are illustrated in neon green. Large blue arrows represent

products.

31

Natural Gasoline Expansion Process Flow Diagram

The proposed design curtails the necessity for re-injection by expanding the separation

train to include sales gas and LPG recovery processes. The process builds off of the overhead

streams that were re-injected. The additional sales gas and LPG recovery streams required

dehydration, refrigeration, and heat exchange prior to the final separation steps.

32

Figure 5. Final simulation of “expanded” natural gasoline production.

Material streams are blue, energy streams are red, and special controls are illustrated in neon green. Large blue arrows represent

products.

33

Process Descriptions

Inlet Separation and Natural Gasoline Recovery

In the inlet separation and natural gasoline recovery portion of the process water in the

wellhead stream is separated from the liquid hydrocarbons and natural gas prior to natural

gasoline recovery in the C5+ distillation column.

Natural Gasoline PFD

Figure 6 depicts the process of inlet separation of the wellhead stream, natural gasoline

recovery, and natural gas compression for re-injection into the well:

1Wellhead

50 °F

2190 psia

2 3

Water 1

Hydrocarbons

C5+ Column

Overhead 1

Qheat

3-Phase Inlet

Separator

Overhead 2

4Qcomp1

Overhead 3

Qc1

Qr1

C5+

$C5+ Storage

C5+

Qcomp2

Overhead 2

Compressor

Re-Injection

Compressor 1

Qcomp3

Re-Injection

Compressor 2

5

Qcomp4

Re-Injection

Compressor 3

6 Re-Injection

105 °F

1505 psia

50 °F

155 psia

400 °F

160 psia

322 °F

2625 psia

69 °F

605 psia66 °F

605 psia

-$

-$

Figure 6. “Current” natural gasoline production process flow diagram with select stream conditions.

Natural Gasoline Expansion PFD

In the natural gasoline expansion process, inlet separation proceeds in a similar fashion

to the natural gasoline process; however, the natural gasoline recovery column includes a

recycled feed from the LPG recovery process. In this process the natural gas was not

compressed for re-injection, but rather simply piped to the TEG dehydration cycle as illustrated

in Figure 7.


34

1Wellhead

50 °F

2190 psia

2 3

Water 1

Hydrocarbons

C5+ Column

Overhead 1

Qheat

3-Phase Inlet

Separator

Overhead 2

4

Qcomp1

Overhead 3 to TEG Dehydration

Qc1

Qr1

C5+

$C5+ Storage

C5+

Overhead 2

Compressor

105 °F

1505 psia

50 °F

165 psia

393 °F

175 psia

69 °F

605 psia

C5+ Recycled

278 °F

200 psia

C5+ from Sales/LPG Recovery

300 °F

255 psia

-$

66 °F

605 psia

Figure 7. Natural gasoline expansion process flow diagram showing overhead to TEG dehydration, and

recycle from LPG recovery process.

Approach

Separation of the water, liquid hydrocarbons, and natural gas in wellhead natural gas

streams is often accomplished via mechanical three phase separation units (7). This is the

approach utilized in the both of the proposed processes. Consultation with Mr. Arendell

confirmed this method of inlet separation as being well-suited to the processes being modeled.

To assist in the modeling of C5+ recovery the GPSA section on fractionation as well as

specifications provided by Mr. Arendell were utilized (21). In addition, the GPSA section on

separation equipment was used to determine a residence time for vessel sizing of the inlet

separator (22). The specifications used to converge this phase of the process are as follows:

Hypothetical C5+ components in dry well stream of equal composition, adding to the total

mole fraction of C5+ in the stream as given by the problem statement

Dry Well and Water Well streams at 50 °F and 150 bar g as per the problem statement

Pressure drop across inlet heater and valve to 103 bar g as given in problem statement

Temperature reached after heating by inlet heater and expansion by inlet valve 105 °F in

stream 3 to avoid hydrate formation


35

Pressure drop of 62.05 psi across inlet separator as specified by Mr. Arendell

Pressure of 175 psia at the reboiler of the C5+ column, with a 10 psi pressure drop up

the column

Reid Vapor Pressure (RVP) column specification of 10 psia at reboiler stage of C5+

column as per the problem statement

Temperature column specification at condenser stage of 50 °F of C5+ column to avoid

hydrate formation

Feed stage 1 for natural gasoline process and stages 1 and 10 for the hydrocarbons and

C5+ Recycled streams, respectively, for the natural gasoline expansion process

Compression of C5+ overhead to 41.71 bar to match the pressure of the overhead from

the inlet separator before their combination

Gas pressure after Re-injection compressors 1, 2, and 3 of 77.35 bar, 129.4 bar, and

181 bar respectively, meeting the re-injection pressure specification of 180 bar g given in

the problem statement

For the inlet separation and C5+ recovery stage of both processes the flow rates of the

Dry Well and Well Water streams were determined by given downstream parameters. Adjust

operator 1 was implemented to set the flow rate of the Dry Well stream such that the flow rate of

the C5+ stream was 10,000 standard BPD, the design specification given in the problem

statement. Similarly, Adjust operator 2 and the Water Spreadsheet were introduced to set the

flow rate of the Water Well stream such that 1.5 bbl of water was produced from the inlet

separator for every MMSCF of gas flow from the inlet separator, another given specification.

Converging distillation columns in Aspen HYSYS requires the specification of two

process variables to account for the two degrees of freedom in the column. As the C5+ product

stream purity specification was a RVP of 10 psia, one of the column specifications for the C5+

column was that the reboiler stage has an RVP of 10 psia. Originally, the other specification to

account for the final degree of freedom in the column was a reflux ratio of 0.5 suggested by Mr.

Arendell. However, using this configuration hydrates were found in Overhead 2 of this column.

The second column specification was therefore changed to a 50 °F temperature at the

condenser stage of the column. This was found to eliminate hydrate formation in the overhead.


36

A storage tank was also modeled to hold one day’s worth of natural gasoline product, or 10,000

bbl.

The only difference in the inlet separation phase of the processes is that in the natural

gasoline process the combined overhead gasses are compressed for re-injection into the well,

whereas in the natural gasoline expansion process the combined overhead streams are simply

routed to the dehydration system. For the re-injection process three compressors were

modeled to bring the gas up to re-injection pressure so that the cost of these could be

determined using an available costing equation. Therefore, to size the compressors within the

maximum horsepower constraint of the costing equation Adjust functions 3 and 4 were used to

adjust the pressures of the respective outlet gas streams such that the a horsepower of 5990

was achieved in the first two re-injection compressors. The final re-injection compressor works

to bring the natural gas pressure up to re-injection specification. In reality, only one large

compressor would be used for re-injection; however, Mr. Arendell agreed that this alternate

configuration employed for economic convenience was adequate to model this portion of the

process.

Triethylene Glycol Dehydration

To produce sales-quality natural gas, excess water must be removed from the gas to

meet the common standards for sales gas as well as to protect the pipeline from damage.

Triethylene Glycol Dehydration PFD

In an effort to mitigate deleterious hydrate formation, dehydration is a crucial step to

remove water from the process. Here, triethylene glycol was used to absorb water in the TEG

contactor as shown in Figure 8.


37

Overhead 3 from

Inlet Separation

Dry Gas

Rich TEG

Lean TEG

TEG Flash Drum

5

Flare Gas

TEG Air Cooler

TEG

TEG Pump

3

TEG HX

69 °F

605 psia

95 °F

690 psia

75 °F

100 psia

Regenerated TEG

6

178 °F

15 psia

7

300 °F

90 psia

400 °F

16 psia

Qc2

Qr2

TEG

Regeneration

Column

300 °F

30 psia

Water 2

Propane HX

Propane Out to

Refrigeration

Cycle

Propane In from

Refrigeration

Cycle69 °F

593 psia

253 °F

15 psia

12 to Sales/LPG Recovery

-36 °F

18 psia

-38 °F

17 psia

-33 °F

590 psia

11

-40 °F

510 psia

Qpump

-$

-$

TEG Contactor

Figure 8. TEG dehydration to remove water from the overhead gas stream prior to sales gas and LPG

recovery.

Approach

Liquid desiccant dehydration equipment, more specifically triethylene dehydration, can

be easily automated for use in remote areas. The primary source for glycol dehydration was the

corresponding GPSA section (8). It was recommended by Mr. Arendell to model the dehydration

cycle with triethylene glycol (TEG) and to generate a flow diagram with the assistance of the

outlined GPSA unit operations.

TEG is the most common liquid desiccant used for natural gas dehydration. It was

recommended within the GPSA document that the design employ a 3 gal. TEG/lb water

absorbed ratio (8). The dehydration cycle was converged with the following specifications:

Circulation rate of TEG: 3 gal. TEG/ lb water absorbed

The TEG contactor was specified to have a 5 psia pressure drop.

The TEG regeneration column is run at atmospheric pressure with a 10 kPa pressure

drop up the column.

The reboiler on the TEG regeneration column is specified at a temperature of 400 °F.


38

Reflux ratio of the TEG regeneration column is specified as 0.50

Process stream pressure drop of 10 psi, air intake temperature of 85°F, and air intake

pressure of 14.7 psia for the air cooler

The pressure drop for the pump was specified to 684.5 psia

Commonly, an inlet scrubber is installed to prevent accidental dumping of large

quantities of water, hydrocarbons, or corrosion inhibitors into the TEG absorber (8). However, it

was decided to exclude the scrubber in the proposed process because the feed stream was

completely vapor. The first unit operation in the dehydration system is the TEG contactor in

which glycol enters on the top stage and absorbs the water from the counter-current vapor

stream. The water-rich TEG is then subjected to a flash drum which flashes off most of the

soluble gas and flares it. The water is then removed from the water-rich TEG stream within the

TEG regeneration column. The column removes the water from the TEG at atmospheric

pressure with heat (400 °F). The lean-TEG stream is then cooled with the bottoms of the flash

drum, which brings the TEG closer to the feed conditions for the TEG absorber. The lean-TEG

is then pumped and further cooled with an air cooler to return the stream to the absorber feed

conditions. The glycol dehydration is a crucial step in ensuring pipeline quality LPG and sales

gas.

Propane Refrigeration Cycle

The purpose of the propane refrigeration cycle is to cool the dehydrated gas stream prior

to entering the sales gas recovery column.

Propane Refrigeration Cycle PFD

Figure 9 illustrates the two-stage refrigeration cycle that was used to cool the dehydrated

gas stream prior to sales gas and LPG recovery:


39

Dry Gas from

TEG

Dehydration

69 °F

593 psia

Propane HX

Propane Out

Propane In

12 to Sales/LPG Recovery

-33 °F

590 psia

11

-40 °F

510 psia

-36 °F

18 psia

Liquid

-38 °F

17 psia

Liquid

Propane

Recycled Liquid

Propane

Flash Drum 1

Suction Drum

Propane Vapor

Qcomp3

Propane

Compressor 1

15

Propane Recycled Vapor

16

Qcomp4

17

Propane

Compressor 2Propane Air Cooler

67 °F

60 psia

24 °F

58 psia

Propane

Flash Drum 3

Economizer

24 °F

60 psia

55 °F

58 psia

155 °F

187 psia

18

Propane

Flash Drum 2

Accumulator

Propane

Liquid

19

95 °F

177 psia

26 °F

62 psia

Liquid Vent

Vapor Vent

Figure 9. Two-stage propane refrigeration cycle to reduce the temperature of the gas stream prior to sales

gas recovery in the reboiled absorber.

Approach

The cooling of the dehydrated gas stream to -33 °F partially condenses the stream to

enhance separation in the reboiled absorber (21). The primary guiding document for the

refrigeration cycle was the corresponding GPSA section (23). In concordance with this

document, it was recommended by Mr. Arendell to model a two-stage refrigeration cycle with an

economizer. This system saves on refrigeration costs by reducing compressor duty while not

investing in the additional equipment required for a three-stage system.

The cycle was simulated with pure propane; however, refrigeration-grade propane

contains 98% propane and 2% w/w ethane (15). Therefore, the refrigeration cycle is idealized

and adaptable once the exact composition of the on-site refrigerant is determined. The cycle

was built with minimal stream specifications. The specifications were as follow:

Vapor/Phase Fraction of 1.0 and temperature of -38 °F in Propane Out to allow for a 5 °F

approach temperature with process Stream 11

Inlet pressure drop of 1.5 psi in Suction Drum


40

Pressure of 60 psia in Stream 15

Process stream pressure drop of 10 psi, air intake temperature of 85°F, and air intake

pressure of 14.7 psia for the air cooler

Vapor/Phase Fraction of 0 and temperature of 95°F in Stream 18

Pressure of 62 psia in Stream 19

Inlet and vapor outlet pressure drops of 2 psi in the Economizer

In brief, the propane refrigerant undergoes four steps with the intent to evaporate in the

process heat exchanger, thereby cooling the process stream from 70 °F to -33 °F.

Coincidentally, the propane stream reduces from -35.6 °F to -38 °F during vaporization, allowing

for a five degree Fahrenheit approach. This represents the first of the four steps. The stream

then passes through a suction drum to knock out any liquids prior to compression. This was

initially modeled as a flash drum but was corrected to a tank to avoid background equilibrium

calculations. Secondly, the vapor is compressed in two different compressors. The second

compressor combines the vapor product from the economizer with the one-time compressed

vapor that is once-removed from the process heat exchanger. This is the energy saving step

that characterizes this system as a two-stage cycle. The propane stream is still in the

superheated vapor form following compression, thus giving way to the third step of

condensation in the air cooler. The stream completely condenses via heat exchange with air.

This step necessarily cools the stream to prepare for a two-step expansion via passing through

the economizer. There is a vapor vent potential in the accumulator to isolate liquid refrigerant

prior to expansion. Thus, the fourth and final step is expansion to reduce the pressure and

temperature of the refrigerant prior to heat exchange with the process stream.

Sales Gas and LPG Recovery

The sales gas and LPG portion of the natural gasoline expansion process purifies and

recovers natural sales gas and LPG, as well as removing excess C5+ from the natural gas and

recycling it back to the C5+ recovery column.

PFD

Figure 10 illustrates the configuration of the unit operations required to perform the

functions described above:


41

Dry Gas from

TEG Dehydration

Propane HX

Propane Out to

Refrigeration

Cycle

Propane In from

Refrigeration

Cycle

69 °F

593 psia

12

-36 °F

18 psia

-38 °F

17 psia

-33 °F

590 psia

11

-40 °F

510 psia

Sales Gas Column

Sales Gas

Qcomp4

Propane

Compressor 2

-40 °F

500 psia

LPG Column

Qc4

Qr4

Heavy LPG 13

Qr3

C5+ Recycle to Inlet Separation

237 °F

510 psia

184 °F

255 psia

LPG to Pipeline

115 °F

245 psia

300 °F

255 psia

Sales Gas Compressed to Pipeline

27.69 °F

812.4 psia

$

$

Figure 10. Sales gas recovery in the reboiled absorbed and LPG recovery in the distillation column. The

bottoms product of the LPG column recycles to the C5+ recovery column to enhance yield.

Approach

Sales gas recovery is often accomplished by an absorption tower in industry (7). Upon

suggestion by Mr. Arendell, a reboiled absorber specifically was implemented for this purpose.

In order to model LPG recovery, the GPSA section on fractionation as well as a patent by

Mealey were utilized to determine various operating parameters including relevant temperatures

to the process (21) (24). Using these sources as well as input from Mr. Arendell, sales gas and

LPG of appropriate qualities were recovered from the natural gasoline expansion process.

Following are the specifications that were used to achieve convergence of this portion of the

process:

Temperature and pressure of gas stream lowered to -40.01 °F and 510 psia,

respectively, with Propane Heat Exchanger and Sales Valve before entrance to Sales

Gas Column to meet ideal absorption conditions, as suggested by Mr. Arendell

Pressure of 510 psia at the reboiler of the sales gas absorption column, with a 10 psi

pressure drop up the column

Temperature column specification of 237 °F at the Sales Gas column reboiler

Sales gas pipeline injection pressure of 55 psig as specified in the problem statement

Pressure of 255 psia before entrance to LPG column


42

Pressure of 255 psia at the reboiler of the LPG column, with a 10 psi pressure drop up

the column

Temperature column specification in the LPG column of 300°F at the reboiler

Composition column specification on the LPG product stream of the LPG column of

1.420 X 10-2 fraction by volume of the lowest molecular weight hypothetical C5+

component

The convergence of the absorption and distillation columns in the sales gas and LPG

recovery portion of the process in such a way that all of the purity specifications on the two

product streams were met was a difficult process. As the specifications for sales gas and LPG

purity were given in terms of a hydrocarbon (HC) dew point and maximum vapor pressure

specification, respectively, these were first used as column convergence parameters. The

refluxed absorber, which has only one degree of freedom, was converged based on an a HC

dew point specification such that the HC dew point of the sales gas was below the given

maximum value of 0 °F. The LPG distillation column was then converged using specifications

for True Vapor Pressure (TVP) at 100 °F and reflux ratio. However, with this set of parameters,

it was impossible to meet the other purity specification of the LPG product stream, that it should

contain less than 2 % of C5+ species by volume.

Eventually, upon the suggestion of a reboiler temperature of 300 °F for the LPG column

by Mr. Arendell, all of the product purity specifications were met. First of all, one of the column

convergence parameters of the LPG column was set to meet the given reboiler temperature.

Then, in order to meet the C5+ content specification for the LPG product, a column parameter

controlling the composition of the highest boiling hypothetical C5+ component in the LPG stream

was created. By adjusting this parameter downwards, the stream was purified of C5+ to

acceptable levels. However, at this point the TVP specification of the stream was not being

met. This was accomplished by changing the convergence parameter of the Sales Gas column

to reboiler temperature, then adjusting this value upwards. This resulted in more heavy

hydrocarbons being reboiled into the Sales Gas stream, which increased the HC dew point of

this stream (within acceptable allowances), while decreasing the TVP of the LPG product. By

this method, and the fact that the water content specification of the Sales Gas stream was easily

met by previous TEG dehydration, all of the product specifications for both the Sales Gas and

the LPG were met.


43

Material and Energy Balances

The performance of material and energy balances was necessary to validate the accuracy with

which Aspen HYSYS was able to model the natural gas processes.

Material and Energy Balances

Material and energy balances for both processes were valid, except the overall energy

balance on the expansion process. This imbalance may be attributed to relaxed recycle

sensitivities. A suggestion from Professor Clough and Mr. Arendell to tighten the sensitivities

from 10 to less than unity and increase the number of iterations from 10 to over 100 arose

during the final presentation. This approach was explored and proved to be constrained by time.

At this point in the design, single iterations of one sensitivity unit were allowed to run for several

hours without completing an iteration. Therefore, it is recommended to investigate this approach

upon first reaching convergence of the recycle function. Overall and unit operation balances

were performed in an effort to pinpoint the imbalances.

Balances were completed about the entire processes to validate conservation of mass

and energy. The imbalance of the process, given by Equation 1, demonstrated the validity of

PFD convergence.

Equation 1. Equation to calculate imbalance for material and energy streams.

𝐼𝑚𝑏𝑎𝑙𝑎𝑛𝑐𝑒 = (𝑇𝑜𝑡𝑎𝑙 𝐹𝑙𝑜𝑤 𝑜𝑓 𝑂𝑢𝑡𝑙𝑒𝑡 𝑆𝑡𝑟𝑒𝑎𝑚𝑠) − (𝑇𝑜𝑡𝑎𝑙 𝐹𝑙𝑜𝑤 𝑜𝑓 𝐼𝑛𝑙𝑒𝑡 𝑆𝑡𝑟𝑒𝑎𝑚𝑠)

Furthermore, the relative imbalance, illustrated in Equation 2, normalizes the imbalance to the

total flow of inlet streams. The expected value of this figure is zero; however, HYSYS is

accurate to 0.02 % (9).

Equation 2. Equation to calculate relative imbalance for material and energy streams.

𝑅𝑒𝑙𝑎𝑡𝑖𝑣𝑒 𝐼𝑚𝑏𝑎𝑙𝑎𝑛𝑐𝑒 (%) =𝐼𝑚𝑏𝑎𝑙𝑎𝑛𝑐𝑒

𝑇𝑜𝑡𝑎𝑙 𝐹𝑙𝑜𝑤 𝑜𝑓 𝐼𝑛𝑙𝑒𝑡 𝑆𝑡𝑟𝑒𝑎𝑚𝑠× 100

A relative mass imbalance of zero percent demonstrates that the PDF is fully converged.

Natural Gasoline Process Balances

The mass balance for the natural gasoline process is shown in Table 9:


44

Table 9. Natural gasoline material balance

C5+ Material Streams

Inlet lb/hr Outlet lb/hr

Dry Well 5.58E+05 Water 1 5.10E+03

Water Well 5.38E+03 LNG 1.01E+05

C5+ Vapor Product 0

Re-Injection Gas 4.57E+05

Total (lb/hr) 5.64E+05 5.64E+05

Imbalance (lb/hr) 0

Relative Imbalance 0%

The relative imbalance is acceptable.

The energy balance for the natural gasoline process is shown in Table 10.

Table 10. Natural gasoline energy balance

C5+ Energy Streams

Inlet Btu/hr Outlet Btu/hr

Dry Well -9.75E+08 Water 1 -3.49E+07

Water Well -3.69E+07 Qc1 1.00E+05

Qheat 3.32E+07 LNG -7.73E+07

Qr1 2.08E+07 C5+ Vapor Product 0.00E+00

Qcomp1 6.04E+05 Re-Injection Gas -7.89E+08

Qcomp2 2.04E+07

Qcomp3 2.04E+07

Qcomp4 1.55E+07

Total -9.01E+08 -9.01E+08

Imbalance (Btu/hr) 4000




45

The unit operation balance exposed an imbalance about the C5+ column, as seen in

Table 11.

Table 11. Unit operation balances with C5+ column detail for natural gasoline process. Remarkable (>1)

imbalances denoted in red.

C5+ Process Imbalance

Unit Op Name Mass Flow (lb/hr) Energy Flow (Btu/hr) Volume Flow (bpd)

3-Phase Inlet Separator 1.59E-06 4.84E-07 1.09E-04

C5+ Column 3.97E-06 4.00E+03 2.72E-04

Inlet Heater 3.17E-06 1.36E-06 2.17E-04

Inlet Valve 2.38E-06 1.22E-06 1.63E-04

LNG Storage Tank 7.94E-06 3.41E-06 5.43E-04

Overhead 2 Compressor 4.76E-06 2.05E-06 3.26E-04

Overhead Mixer 7.94E-07 -3.54E-07 5.43E-05

Re-Injection Compressor 1 5.56E-06 2.39E-06 3.80E-04



Wellhead Mixer -1.13E-10 1.99E-07 -3.02E-11

Total 0 4000 0

C5+ Process C5+ Column Imbalance


Condenser -4.41E-12 -3.39E-01 7.07E-13

Main TS 1.59E-06 3.92E+03 1.09E-04

Reboiler 7.94E-07 9.43E00 5.43E-05

Total 0 4000 0

The energy imbalance observed in the C5+ column existed but was negated when normalized,

as seen in the overall energy balance.

Expansion Process Balances

Similarly, a material balance about the expansion process is shown in Table 12.


46

Table 12. Expansion process material balance

Expansion Process Material Streams

Inlet lb/hr Outlet lb/hr

Dry Well 5.58E+05 Water 1 5.10E+03

Water Well 5.38E+03 Flare Gas 1.45E+01

Water 2 3.02E+02

Sales Gas Compressed 4.39E+05

LPG 1.14E+04

Propane Liquid 0 0.00E+00

Propane Vapor 0 0.00E+00

LNG 1.08E+05

C5+ Vapor Product 6.54E+00

Total (lb/hr) 5.64E+05 5.64E+05

Imbalance (lb/hr) 23



The problematic energy balance for the natural gasoline process is shown in Table 13:


47

Table 13. Expansion energy balance

Expansion Process Energy Streams

Inlet Btu/hr Outlet Btu/hr

Dry Well -9.75E+08 Water 1 -3.49E+07

Water Well -3.69E+07 Qc1 1.35E+05

Qheat 3.32E+07 Flare Gas -2.48E+04

Qr1 2.08E+07 Qc2 1.35E+05

Qcomp1 5.88E+05 Water 2 -1.59E+06

Qr2 8.42E+05 Sales Gas Compressed -8.42E+08

Qpump 1.75E+04 Qc4 4.27E+07

Qcomp2 1.12E+07 LPG -1.30E+07

Qr4 4.19E+07 Propane Liquid 0 0

Qr3 4.62E+06 Propane Vapor 0 0

Qcomp3 7.65E+06 LNG -8.28E+07

Qcomp4 1.02E+07 C5+ Vapor Product -4.56E+03

Total (Btu/hr) -8.81E+08 -9.31E+08

Imbalance (Btu/hr) -4.97E+07

Relative Imbalance 5.6%

The relative imbalance of 5.6% is the point of discrepancy for the expansion process.

The unit operation balance exposed significant and numerous energy imbalances, as

seen in Table 14:


48

Table 14. Unit operation balances for the expansion process. Remarkable (>1) imbalances denoted in red.

Expansion Process Imbalance


3-Phase Inlet Separator 2.38E-06 8.25E-07 1.63E-04

C5+ Column 1.03E-05 -8.75E+03 7.06E-04

C5+ Recycle -4.68E+00 3.14E+03 -4.38E-01

C5+ Recycle Valve 7.14E-06 3.07E-06 4.89E-04

Inlet Heater 9.52E-06 4.09E-06 6.52E-04

Inlet Valve 3.17E-06 1.56E-06 2.17E-04

LNG Storage Tank 2.62E-05 1.12E-05 1.79E-03

LPG Column 1.19E-05 -2.40E+04 8.15E-04

LPG Valve 5.56E-06 2.39E-06 3.80E-04

Overhead 2 Compressor 1.27E-05 5.46E-06 8.70E-04

Overhead Mixer 7.94E-07 -3.54E-07 5.43E-05

Propane Air Cooler 2.54E-05 -4.93E+07 1.74E-03

Propane Compressor 1 1.43E-05 6.14E-06 9.78E-04

Propane Compressor 2 1.51E-05 6.48E-06 1.03E-03

Propane Flash Drum 1 1.75E-05 7.51E-06 1.20E-03



Propane HX 2.30E-05 -3.09E-01 1.58E-03

Propane Mixer 1.59E-06 6.82E-07 1.09E-04

Propane Valve 1 8.73E-06 3.75E-06 5.98E-04

Propane Valve 2 7.94E-06 3.41E-06 5.43E-04

Sales Compressor 1.35E-05 5.80E-06 9.24E-04

Sales Gas Column 2.38E-05 -1.14E+02 1.63E-03

Sales Valve 4.76E-06 2.05E-06 3.26E-04

TEG Air Cooler 2.46E-05 -3.05E+05 1.68E-03

TEG Contactor 1.59E-05 6.21E+01 1.09E-03

TEG Flash Drum 1.67E-05 7.17E-06 1.14E-03

TEG HX 2.22E-05 9.55E-06 1.52E-03

TEG Pump 1.98E-05 8.53E-06 1.36E-03

TEG Recycle 2.77E+01 -6.44E+04 1.69E+00

TEG Regeneration Column 1.11E-05 1.01E+00 7.61E-04

TEG Valve 1 3.97E-06 1.71E-06 2.72E-04

TEG Valve 2 6.35E-06 2.73E-06 4.35E-04

Wellhead Mixer -1.13E-10 1.99E-07 -3.02E-11

Total 0 -4.97E+07 0


49

The significant negative energy imbalance was caused by a combination of columns, recycle

functions, and an air cooler imbalance. The final value with a seventh-order magnitude exposes

the source of error.

Detailed column imbalances are shown in Table15.

Table 15. Column balance details for expansion process. Remarkable (>1) imbalances denoted in red.

C5+ Column Imbalance


Condenser -1.76E-12 7.48E-01 0

Main TS 1.59E-06 -8.68E+03 1.09E-04

Reboiler 7.94E-07 -6.89E+01 5.43E-05

Total 0 -9.00E+03 0

TEG Contactor


TS-1 -2.26E-10 62 0

TEG Regeneration Column


Condenser 1.59E-06 5.81E-07 1.09E-04

Main TS 7.94E-07 1.01E+00 5.43E-05

Reboiler 4.41E-12 1.71E-07 5.89E-14

Total 0 0 0

LPG Column


Condenser 1.59E-06 3.33E-01 1.09E-04

Main TS 7.94E-07 -2.40E+04 5.43E-05

Reboiler -1.13E-10 4.00E-01 -7.54E-12

Total 0 -24000 0

Sales Gas Column


Main TS -1.13E-10 -1.13E+02 -3.02E-11

Reboiler 7.94E-07 -1.51E+00 5.43E-05

Total 0 0 0


50

The additive effect of these imbalances was corroborated by the overall energy imbalance. This

analysis expands the imbalance beyond the recycle functions. However, both recycle balances

are shown in Table 16 to provide additional clarity to the imbalance:

Table 16. Recycle balances for the expansion process. Remarkable (>1) imbalances denoted in red.

C5+ Recycled Stream 14 C5+ Recycled Imbalance Relative Imbalance

Vapour 1.58E-01 1.58E-01 -2.62E-05 -0.02%

Temperature (°F) 2.78E+02 2.78E+02 -2.61E-02 -0.01%

Pressure (psia) 2.00E+02 2.00E+02 0.00E+00 0.00%

Molar Flow (lbmole/hr) 1.13E+02 1.13E+02 -8.42E-02 -0.07%

Mass Flow (lb/hr) 7.37E+03 7.36E+03 -4.68E+00 -0.06%

Std Ideal Liq Vol Flow (bpd) 7.69E+02 7.69E+02 -4.40E-01 -0.06%

Molar Enthalpy (Btu/lbmole) -5.48E+04 -5.48E+04 -1.30E+01 0.02%

Molar Entropy (Btu/lbmole-F) 2.94E+01 2.94E+01 1.02E-02 0.03%

Heat Flow (Btu/hr) -6.22E+06 -6.22E+06 3.14E+03 -0.05%

Total -0.21%

TEG Recycle

Stream 10 Lean TEG Imbalance Relative Imbalance

Vapour 0.00E+00 0.00E+00 0.00E+00 0.00%

Temperature (°F) 9.50E+01 9.50E+01 0.00E+00 0.00%

Pressure (psia) 6.90E+02 6.90E+02 0.00E+00 0.00%

Molar Flow (lbmole/hr) 4.96E+01 4.98E+01 1.99E-01 0.40%

Mass Flow (lb/hr) 6.92E+03 6.95E+03 2.77E+01 0.40%

Std Ideal Liq Vol Flow (bpd) 4.21E+02 4.22E+02 1.68E+00 0.40%

Molar Enthalpy (Btu/lbmole) -3.24E+05 -3.24E+05 1.31E-01 0.00%

Molar Entropy (Btu/lbmole-F) 3.52E+01 3.52E+01 -1.18E-05 0.00%

Heat Flow (Btu/hr) -1.61E+07 -1.62E+07 -6.44E+04 0.40%

Total 1.60%

The imbalances observed in the recycle functions were primarily energy parameters, except for

the mass flow in the TEG Recycle function. This value was apparently negligible; an imbalance

was not observed in the overall mass balance.

Process Description & Equipment Specifications

Equipment was designed to accommodate the greater demand of 10,000 bpd, for both

the natural gasoline and re-injection and expansion project.


51

Distillation Columns

In the re-injection process there is only one column, which is responsible for the

recovery of natural gasoline at a standard flow rate of 9182 bpd. In the expansion process,

which is capable of recovering natural gasoline, LPG, and sales gas there are a total of five

columns. The purpose of the first distillation column in the separation train (the C5+ Column) is

to separate the heavier hydrocarbons (C5 and above) from the lighter hydrocarbons. An

absorption column and a distillation column are pertinent unit operations within the glycol

dehydration step of the process. The first column in the dehydration step is an absorption

column which acts to remove the remaining water from the process stream by contacting the

stream with a TEG stream. The vapor outlet of the TEG contactor goes on to undergo further

separation to produce LPG and sales gas. The bottoms product from the TEG contactor is the

feed for the TEG regeneration distillation column. The purpose of the TEG regeneration column

is to remove the water from the glycol restoring it to a purity of 99.0 wt%. A refluxed absorber

column is used to separate the heavier hydrocarbons from the sales gas product. The bottoms

stream from the refluxed absorber column is the feed to the LPG recovery distillation column.

The LPG distillation column separates the LPG product from the heavier hydrocarbons (C5+).

The heavier hydrocarbons are then recycled back to the C5+Column.

Estimating Column Pressure and Condenser Type

The column operating conditions are important for obtaining the desired product

specifications. In conjunction with recommendations from Mr. Arendell and Professor Clough,

the diagram in Figure 11 was followed to determine an appropriate column pressure and

condenser type.


52

Calculate bubble-

point pressure (PD)

of distillate at

120°F

Start

Distillate and bottoms

are known or estimated

Calculate bubble-

point pressure (PD)

of distillate at

120°F

PD > 215 psia

Choose a

refrigerant so as to

operate partial

condenser at

415 psia

PD > 365 psia

Estimate bottoms

Pressure

(PB)

PD < 365 psia

Use partial condenser

PD < 215 psia

Use total condenser

(reset PD to 30 psia

If PD <30 psia)

Calculate bubble-

point temperature

(TB) of bottoms at

PB

Lower pressure

PD appropriately

TB > bottoms

decomposition or critical temperature

TB < bottoms

decomposition or

critical temperature

Figure 11. Decision tree to determine column pressures and condenser types.

The final column operating parameters are shown in Table 17 where the number of trays and

feed tray locations were iterated and optimized to reduce the reboiler duty.


53

Table 17. Final column operating temperatures and pressures.

Column Inlet

Temperature (oC)

Inlet Pressure

(kPa)

Distillate Temperature

(oC)

Distillate Pressure

(kPa)

Bottoms Temperature

(oC)

Bottoms Pressure

(kPa)

C5+

Recovery 19.09 136.6 4171 1379 10.09 1138 213 1207

TEG Contactor Absorber

35 20.55 4757 4171 20.90 4089 20.77 4123

TEG

Regeneration 148.6 206.8 122.9 101.3 204 110.0

Sales Gas Refluxed

Absorber -40 3516 -39.77 3447 113.9 3516

LPG

Recovery 84.46 1758 45.90 1689 148.9 1758

Calculating Number of Trays

In order to determine the number of trays that each distillation column needs, the

distillation columns were first attempted as shortcut distillation columns in Aspen HYSYS.

However, after being unable to make the columns converge another method was utilized. A

trial-and-error method which included changing the number of trays until the duty of the

condensers and reboilers were minimized was utilized instead. The number of trays that yielded

the minimum duty was chosen as the actual number of trays. The resulting number of trays and

reflux ratios, if applicable, can be seen below in Table 18.


54

Table 18. Final column operating tray and reflux ratio specifications.

Column Number of

Trays

Number of

Distillate Trays

Number of

Bottoms Trays

Reflux

ratio

C5+ Recovery 10 4 6 1.43x10-3

TEG

Absorber 5 2 3 N/A

TEG Regeneration 5 2 3 0.50

Sales Gas

Refluxed Absorber 12 5 7 N/A

LPG Recovery 12 5 7 25.8

Determining the Dimensions of the Distillation Columns

The determination of the diameter for a distillation column is a relatively straightforward process

requiring only vapor flow rate, G, liquid flow rate, L, pressure, liquid density, ρL, and vapor

density, ρG. Using the aforementioned properties, the flow ratio parameter is calculated using

Equation 3, seen below.

Equation 3.Flow ratio parameter of liquid and vapor flow rates

𝐹𝐿𝐺 = (𝐿

𝐺) ∗ (

𝜌𝐺

𝜌𝐿

)0.5

The parameter CSB is estimated using the obtained value for the flow ratio parameter,

18-in. tray spacing, and a correlation established by Fair in 1961 (25).

Using the parameter CSB, the empirical capacity parameter, C, is calculated for use in

the determination of the flooding velocity. The appropriate equation includes a surface tension

factor, FST, a foaming factor, FF, and a hole-area factor, FHA. Equation 4 demonstrates the

calculation of this parameter:

Equation 4.Empirical capacity parameter calculation

𝐶 = 𝐶𝑆𝐵 ∗ 𝐹𝑆𝑇 ∗ 𝐹𝐹 ∗ 𝐹𝐻𝐴

The flooding velocity, Uf , is computed from the empirical capacity parameter, based on a force

balance on a suspended liquid droplet which can be referenced below in Equation 5 (25).


55

Equation 5. Flooding velocity

𝑈𝑓 = 𝐶 ∗ (𝜌𝐿 − 𝜌𝐺

𝜌𝐺

)0.5

The inner diameter of the tower, DT, is computed using tower cross-sectional diameter, At,

downcomer area, AD, and flooding fraction, f. The equation can be referenced in Equation 6:

Equation 6. Inner tower diameter.

𝐷𝑇 = [4𝐺

(𝑓𝑈𝑓)𝜋 (1 − (𝐴𝐷

𝐴𝑡) 𝜌𝐺 )

]

0.5

Following the calculation for the diameter of the column, it is necessary to calculate the

length and thickness of the column, which are pertinent components in the costing of distillation

columns.

The thicknesses of each of the distillation columns in the process were calculated to

ensure column rigidity and strength to stand up to potential earthquake hazards and wind loads.

While the Yamal Peninsula is not in a particularly earthquake-prone location, it was determined

prudent to overdesign the columns for this situation, which does not require an over-excess of

additional carbon steel for construction.

To calculate the column thicknesses, the thickness of steel needed for the top of the

columns was first determined. This requires the determination of the inner diameter of the

column, the design pressure, the maximum allowable stress of the steel, and the weld efficiency

for construction. The inner diameter was found using Equation 6, as shown in the section

above, and the design pressure was calculated using Equation 7:

Equation 7. Design pressure calculation

𝑃𝑑 = exp (0.60608 + 0.91615[𝐿𝑛(𝑃𝑜)] + 0.0015655[𝐿𝑛(𝑃𝑜)]2 )

Since none of the components being distilled are corrosive, it was deemed sufficient to

construct the columns from Grade C carbon steel, for which the maximum allowable stress is

12650 psig (26) (27) (28). For carbon steel up to 1.25 in. thick, only a 10% spot X-ray check of

the welds are necessary and Seider et al. cite a weld efficiency of 0.85 to be sufficient (25).


56

With these parameters defined it was then possible to calculate the thickness of carbon steel at

the top of the columns necessary to withstand the internal pressure, tp, as shown in Equation 8:

Equation 8.Thickness at top of column

𝑡𝑝 =𝑃𝑑𝐷𝑖

2𝑆𝐸 − 1.2𝑃𝑑

Here Pd is the design pressure, Di is the internal diameter of the column, S is the

maximum allowable stress of Grade C carbon steel, and E is the weld efficiency. When the

calculated thicknesses were not sufficient for rigidity, an increased thickness was used to

sustain the columns.

The next step in the process was to calculate tw, the excess necessary thickness of the

column at the bottom to withstand earthquake and wind load. For this, the approximate outer

diameter of the column, the column length, and the maximum allowable stress are required.

The lengths of the columns were calculated as the sum of the space between the trays and that

required for a sump below the trays and a disengagement space above the trays. Seider et al.

cite 10 ft. and 4 ft. for sump and disengagement spaces respectively for a column with 2 ft. tray

spacing (25). These were scaled for the 18 in. tray spacing used for each of the columns in the

process according to Equation 9:

Equation 9.Column length

𝐿 = (𝑁 − 1)𝐻𝑡,2 +10𝐻𝑡,2

𝐻𝑡,1

+4𝐻𝑡,2

𝐻𝑡,1

In this equation L is the column length, N is the number of trays, Ht,1 is the tray spacing

requiring 10 ft. and 4 ft. sump and disengagement spaces respectively, and Ht,2 is the tray

spacing of the columns in the process. To calculate the approximate outer diameter of the

column, Seider et al. recommend utilizing a conservative estimate for the total thickness of the

column wall, twall, to ensure that the column can stand up to any earthquake or high winds that

may occur (25). The approximate outer diameters of the columns, Do, were calculated using

Equation 10:

Equation 10. Outer column diameter calculation

𝐷𝑜 = 𝐷𝑖 + 2𝑡𝑤𝑎𝑙𝑙


57

The excess thickness required for earthquake and wind hazards, tw, was calculated using

Equation 11:

Equation 11. Excess thickness calculation

𝑡𝑤 =0.22(𝐷𝑜 + 18)𝐿2

𝑆𝐷𝑜2

The variable S in Equation 11 is again the maximum allowable stress of Grade C carbon

steel. From this excess thickness required at the bottom of the column, the total wall thickness

at the bottom of the column, tw allbottom, was calculated using Equation 12:

Equation 12. Thickness of bottom of column

𝑡𝑤𝑎𝑙𝑙𝑏𝑜𝑡𝑡𝑜𝑚 = 𝑡𝑝 + 𝑡𝑤

To find the thickness of the columns to use for costing purposes it was then necessary to

find the average thickness of the columns over their lengths, tv , for which Equation 13 was used:

Equation 13. Average thickness of column throughout length

𝑡𝑣 =𝑡𝑝 + 𝑡𝑤𝑎𝑙𝑙𝑏𝑜𝑡𝑡𝑜𝑚

2

Despite the chemicals being distilled not being cited as being corrosive, for the sake of

prudence in design an allowance for corrosion was nevertheless added to the average wall

thickness to account for any corrosion that might occur over the planned 18 year lifespan of the

plant. The final calculated thickness was determined using Equation 14:

Equation 14.Corrosion insurance to find column thickness

𝑡𝑓 = 𝑡𝑣 + 𝑡𝑐

In Equation 14 tf is the final calculated column wall thickness, tv is the average wall

thickness, and tc is the excess thickness allowed for corrosion. Taking into account the fact that

steel is manufactured in set increments, the thickness of each of the columns was adjusted

accordingly.


58

The calculated column dimensions can be referenced below in Table 19. Worth noting

is the fact that a design diameter was used for calculation of the purchase cost of the distillation

columns. The actual diameter was rounded up to the next whole meter diameter.

Table 19. Final column design parameters.

Distillation Column Diameter

(m)

Design Diameter

(m)

Number of Trays

Length (m)

Thickness (m)

C5+ Recovery 1.87 2.00 10 7.32 0.022

TEG Contactor

Absorber 2.25 3.00 5 5.03 0.069

TEG Regeneration 0.26 0.50 5 5.03 0.0095

Sales Gas

Refluxed Absorber 3.31 4.00 12 8.23 0.083

LPG Recovery 3.92 4.00 12 8.23 0.051

Distillation Column Costing

All of the distillation columns were modeled with a carbon steel shell, carbon steel sieve

trays, and 68 kg steel couplings, flanged manholes, and flanged nozzles. The cost of the

columns (distillation and absorption) is based off a design diameter and the length of the

column, whereas the cost of the trays is only based upon the column diameter.

The costs of the connections were calculated using the thickness of the respective

distillation column. It was estimated that each column needed five couplings, three flanged

manholes, and five flanged nozzles. All of the estimated costs can be clearly seen in Table 20.


59

Table 20. Final column costs

Distillation Column

Purchase Cost of Column

Installation Cost of Trays

Couplings Cost

Flanged Manholes

Cost

Flanged Nozzles

Cost

Total Cost of

Distillation Column

C5+ Recovery $ 132,000 $ 17,000 $ 5,000 $ 28,000 $ 21,000 $ 203,000

TEG Contactor Absorber

$ 158,000 $21,000 $ 3,000 $ 40,000 $ 27,000 $ 249,000

TEG

Regeneration $ 32,000 $5,000 $ 5,000 $ 19,000 $ 14,000 $ 75,000

Sales Gas Refluxed

Absorber $ 285,000 $39,000 $ 3,000 $ 50,000 $ 35,000 $ 412,000

LPG Recovery $ 285,000 $39,000 $ 5,000 $32,000 $31,000 $ 392,000

Total Purchase Cost of Columns $ 1,330,000

Expansion Purchase Cost $ 1,130,000

Flash Drums

Three-Phase Separator

The three-phase separator is the first unit operation in natural gasoline recovery

process. Intuitively, it has three product streams, a vapor stream and two liquid streams. The

primary purpose of the three-phase separator is to remove most of the water from the feed

stream. The vapor stream is primarily composed of methane, the hydrocarbon stream is

composed of mostly larger hydrocarbons with some water and smaller hydrocarbons, and the

water stream contains 99.99% water. The amount of water in the feed was determined using an

Adjust function in Aspen HYSYS, such that there was 1.5 bbl of water for every MMSCF of

vapor out of the three-phase separator. The feed conditions can be seen below in Table 21.

Table 21. Feed conditions to the three-phase inlet separator in Stream 3.

Stream Temperature

(oC) Pressure

(kPa) Mass Flow Rate (kg/hr)

3 (Feed to 3-Phase Separator)

40.56 1.038x104 2.556x105


60

The feed compositions can be seen below in Table 22.

Table 22. Composition of Stream 3, the inlet material stream to the three-phase inlet separator.

Stream Mole

%

N2

Mole %

CO2

Mole %

CH4

Mole %

C2H6

Mole %

C3H8

Mole %

iC4H10

Mole %

nC4H10

Mole %

C5+

Mole %

H20

3 0.4 0.3 85.13 6.56 2.46 0.35 0.41 3.24 1.15

After the three-phase separator the liquid hydrocarbon stream is subjected to a

distillation column in order to separate natural gasoline (C5+) from lighter hydrocarbons and

water. The natural gasoline is the first product in the expansion plant and the sole product in the

re-injection plant. The vapor stream from the three-phase separator was combined with the

overhead of the natural gasoline recovery column. The water stream from the three-phase

separator was subject to treatment to remove all organic components. The product stream

conditions can be seen below in Table 23.

Table 23. Product stream conditions for the three phase inlet separator.

Stream Temperature

(oC) Pressure

(kPa) Mass Flow Rate

(kg/hr)

Overhead 1 ( Vapor Stream) 19.09 4171 2.039x105

Water 1 (Water Stream) 19.09 4171 4.938x104 Hydrocarbons (Hydrocarbon Stream) 19.09 4171 2313

The compositions of each of the resulting streams can be seen in Table 24.

Table 24. Product stream compositions from the three-phase inlet separator.

Stream Mole

%

N2

Mole %

CO2

Mole %

CH4

Mole %

C2H6

Mole %

C3H8

Mole %

iC4H10

Mole %

nC4H10

Mole %

C5+

Mole %

H20

Overhead 1 0.42 0.31 89.17 6.66 2.30 0.29 0.31 0.47 0.06

Water 1 0.00 0.01 0.00 0.00 0.00 0.00 0.00 0.00 99.99

Hydrocarbons 0.03 0.16 17.20 5.95 6.44 1.86 2.72 65.59 0.03


61

Flash Drums

Flash drums were only necessary in the expansion plant process. Three flash drums

were used in the propane refrigeration cycle and one was used in the TEG dehydration cycle.

The TEG flash drum was used to remove the hydrocarbons which contaminated the water-rich

glycol stream. Two of the flash drums in the propane refrigeration cycle are primarily safety

precautions these being the suction drum and accumulator (23 p. 13). The other flash drum in

the propane refrigeration cycle is an economizer, which is responsible for separating vapor and

liquid propane before the liquid propane is again used as a utility fluid for the overhead gas heat

exchanger.

Calculating the Diameter and Height of the Flash Drums

In order to calculate the diameter and height of the flash drums, a 3 to 1 length to

diameter ratio was used. The volume was calculated using the volumetric flow rate and a

residence time of 5 minutes. The calculation of the volume can be seen in Equation 15.

Equation 15. Calculation of Flash Drum Volume

𝑉 = �̇� (𝑚3

𝑚𝑖𝑛)∗ 5 𝑚𝑖𝑛

Using the 3 to 1 length to diameter ratio, the diameter was solved for using the volume of

a cylinder equation, which can be seen below in Equation 16.

Equation 16. Calculation of Flash Drum Diameter

𝐷 = 2 ∗ (𝑉

6𝜋)

13

The length of the flash drums was calculated by simply multiplying the diameter by a

factor of three. The thickness of the flash drum was calculated using the same method utilized

for the distillation columns.

Calculating the Weight of the Flash Drums using the Diameter, Height, and Thickness

In order to calculate the costs of the flash drums it was necessary to find the weight of

the flash drums which is a function of the diameter, height, and thickness. The weight of the

shell and the two elliptical heads was calculated using an equation from Seider et al., which can

be seen below in Equation 17. The diameter, length, and thickness must all be in meters.


62

Equation 17. Weight of shell and two elliptical heads

𝑊 = 𝜋(𝐷 + 𝑡𝑠) ∗ (𝐿 + 0.8𝐷) ∗ 𝑡𝑠 ∗ 17304.17𝑙𝑏

𝑚3

Flash Drum Costing

Since most of the cost of flash drums were unable to be estimated using the costing

worksheet, the cost of the vessels was estimated using the Matches website (29). The three-

phase separator was modeled as a carbon steel air sweep, dry, without a motor separator. The

remaining flash drums were modeled as either large or medium carbon steel vertical pressure

vessels. In addition, three 68 kg couplings were used for each flash drums. The specifications

and costs of each of the flash drums can be seen in Table 25.

Table 25. Final flash drum design parameters and costs.

Flash Separator

Diameter (m)

Length (m)

Thickness (m)

Weight of Vessel (lb)

Couplings Cost

Total Cost of Flash

Separator

TEG Flash Drum 0.52 1.57 0.013 753 $ 3,000 $ 17,000

Propane Flash Drum 1

10.8 32.4 0.013 3.13x105 $ 3,000 $ 712,000

Propane Flash Drum 2 2.16 6.47 0.013 1.26x104 $ 3,000 $ 73,000

Propane Flash Drum 3 5.28 15.8 0.013 7.49x104 $ 3,000 $ 250,000


2.47 7.42 0.013 1.65x104 $ 3,000 $ 121,000

Total Purchase Cost of Flash Drums $ 1,170,000

Expansion Purchase Cost $ 1,050,000

Heat Exchangers

Heat exchangers are vital unit operations in most industrial processes. Choosing the

materials of construction and dimensions are important in ensuring safe operation and proper

heat exchange. In the process, double pipe heat exchangers were used when there was a small


63

heat transfer area and fixed-tube sheet and floating head shell-and-tube heat exchangers were

used when the heat transfer area was large. In addition to the aforementioned heat exchangers,

air-cooled heat exchangers were used for some process streams. For all of the heat exchangers

used in the process carbon steel was used as the material of choice due to each of the utilized

chemicals non-corrosive properties as well as to design for the frigid temperatures of Siberia.

Design of the Heat Exchangers

The log mean temperature difference method was used to determine the area necessary

for proper heat exchange. For the log mean temperature difference method, the duty and log

mean temperature difference are taken directly from Aspen HYSYS. The log mean temperature

difference was checked using Equation 18.

Equation 18. Log mean temperature difference for shell-and-tube heat exchanger

∆𝑇𝐿𝑀 =∆𝑇1 − ∆𝑇2

ln (∆𝑇1

∆𝑇2)

Where ∆𝑇1 is 𝑇ℎ,𝑖𝑛 − 𝑇𝑐 ,𝑜𝑢𝑡 and ∆𝑇2 = 𝑇ℎ,𝑜𝑢𝑡 − 𝑇𝑐,𝑖𝑛.

In order to solve for the heat transfer area, an overall heat transfer coefficient, U, was

estimated using the typical range of overall heat transfer coefficients that are relevant to the

fluids flowing through the heat exchanger, which was obtained from literature (25). Once the U

value was selected, the heat transfer area was calculated using Equation 19.

Equation 19. Heat exchange area calculation

𝐴 =𝑄

𝑈𝐹𝑇∆𝑇𝐿𝑀

In the Equation 19, A is the heat transfer area, Q is the required duty, and FT is the correction

factor, which is also found within the HYSYS interface. The correction factor was either less

than or equal to 1 for all of the heat exchangers in the process.

Heat Exchanger Costing

The material of construction for all of the exchangers in the system is carbon steel

because there is no risk of corrosion. Chilled water is used as the utility fluid for each of the


64

columns with condensers. Both low and high pressure steam are used as the utility fluid for

each of the columns with reboilers. Low pressure steam was used as the utility fluid for the inlet

heater. Both the propane and TEG heat exchangers utilize utility fluid that is the same as the

process stream. In each of the heat exchangers found within the process, the process stream is

run through the tubes of the heat exchanger. The flow rates of both the process stream and the

utility fluid dictate the size of the heat exchangers. Each of the heat exchangers were priced

using the costing equations provided in the costing spreadsheet.

The observed operating conditions for each of the heat exchangers can be seen in Table

26:

Table 26. Heat exchanger operating conditions and corresponding utility requirements

Heat Exchanger

Mass Flow Rate

(kg/hr)

Inlet Temperature

(oC)

Outlet Temperature

(oC)

Inlet Pressure

(kPa)

Outlet Pressure

(kPa) Fluid

Inlet Heater

2.556x105 9.912 113.9 1.510x104 1.509x104 Process

4.104x104 147.6 146.7 446.1 436.1 50 psig

Steam

TEG Heat

Exchanger

3277 24.03 148.9 689.5 620.5 TEG

3140 204.0 81.26 110 106.6 TEG

Propane Heat

Exchanger

2.076x105 20.90 -36.17 4089 4068 Process

9.759x104 -37.53 -38.89 123.7 116.8 Propane

TEG Air Cooler

3140 79.84 35.00 4826 4757 TEG

4.179x105 29.44 30.21 101.4 101.4 Air

Propane Air Cooler

1.354x105 68.28 35 1289 1220 Propane

1.589x106 29.44 61.75 101.4 101.4 Air

The observed operating conditions for each of the condensers can be seen in Table 27:


65

Table 27. Condenser operating conditions

Condenser Mass

Flow Rate (kg/hr)

Inlet Temperature

(oC)

Outlet Temperature

(oC)

Inlet Pressure

(kPa)

Outlet Pressure

(kPa) Fluid

C5+

Condenser

3815 28.10 9.99 1138 1128 Process

3945 7.22 15.56 1138 1090 Chilled Water

TEG Regeneration

Condenser

199.2 99.33 99.00 101.3 101.3 Process

837.8 7.22 32.22 101.3 101.3 Chilled Water

LPG Condenser

1.384x105 62.97 45.86 1689 1679 Process

4.127x105 7.22 32.22 1689 1641 Chilled Water

The observed operating conditions for each of the reboilers can be seen in Table 28.

Table 28. Reboilers operating conditions.

Reboiler Mass

Flow Rate (kg/hr)

Inlet Temperature

(oC)

Outlet Temperature

(oC)

Inlet Pressure

(kPa)

Outlet Pressure

(kPa) Fluid

C5+

Reboiler

8.032x104 137.9 200.7 1207 1193 Process

1.136x104 207.3 207.0 1793 1783 245.3 psig

Steam

TEG Regeneration

Reboiler

3380 144.1 204.0 110.0 100 Process

471.8 214.6 214.3 2068 2058

285.3 psig

Steam

LPG Reboiler

1.555x105 133.5 148.9 1758 1748 Process

2.211x104 185.7 185.3 1136 1126 150 psig

Steam

Sales Gas Reboiler

2.810x104 95.67 113.9 3516 3506 Process

2286 147.6 146.7 446.1 436.1 50 psig

Steam


66

The equipment costs for the heat exchangers were determined using the calculated heat

transfer area and the appropriate costing equation. The heat transfer area and the

corresponding cost for the heat exchangers can be found in Table 29.

Table 29. Heat exchanger cost.

Heat Exchanger ∆T lm

(oF)

U (Btu/ft2-hr-

oF)

Duty (Btu/hr)

Heat Transfer Area (m2)

Purchased Cost

Inlet Heater 132.10 250.00 8.35x107 234.92 $ 37,000

TEG Heat Exchanger 91.24 40 1.00x106 25.49 $2,000

Propane Heat Exchanger

28.11 200.00 3.15x107 519.95 $ 110,000

TEG Air Cooler 36.23 20.00 3.05x105 39.16 $ 24,000

Propane Air Cooler 12.82 10.00 5.01x107 26313.35 $ 494,000

Total Purchase Cost of Heat Exchangers $ 667,000

Expansion Purchase Cost $ 630,000

The heat transfer area and the corresponding cost for the condensers can be found in

Table 30.

Table 30. Condenser cost.

Condenser ∆T lm

(oF)

U (Btu/ft2-hr-

oF)

Duty (Btu/hr)

Heat Transfer Area (m2)

Purchased Cost

C5+ Condenser 9.55 140 1.34x105 9.33 $ 2,000

TEG Regeneration Condenser

141.80 140 8.57x104 0.40 $ 1,000

LPG Condenser 58.25 140 4.22x107 480.39 $ 72,000

Total Purchase Cost of Condensers $ 75,000

The heat transfer area and the corresponding cost for the reboilers can be found in

Table 31.


67

Table 31. Reboiler costs

Reboiler ∆T lm

(oF) U

(Btu/ft2-hr-oF) Duty

(Btu/hr) Heat Transfer

Area (m2) Purchased

Cost

C5+

Reboiler 47.99 250 2.09x107 161.76 $ 32,000

TEG Regeneration Reboiler

56.69 300 8.55x105 4.67 $ 2,000

LPG Reboiler

78.92 250 4.24x107 199.45 $ 41,000

Sales Gas Reboiler 75.09 250 4.65x106 23.02 $ 2,000

Total Purchase Cost of Reboilers $ 77,000

Pumps

The process does not heavily rely on the use of pumps. The pump used for the process

was a cast steel centrifugal pump with an explosion-proof alternating current electric motor. The

volumetric flow rate (Q) and the power requirement were obtained from the Aspen HYSYS

simulation. The pump operates at 75% efficiency. The pump operating conditions and 2010 cost

adjusted purchase costs can be referenced in Table 32.

Table 32. Pump operating conditions and cost

Pump Q

(gal/min) ∆P

(psi) Head (ft)

Power Requirement

(kW)

Purchase Cost of Pump

Purchase Cost of Motor

Total Purchase Cost of Pump

1 12.27 684.5 1458.1 5.116 $ 6,000 $ 2,000 $ 8,000

Compressors

Compressors are a pivotal unit operation in both the re-injection process and the

expansion plant process. Both processes use a total of four compressors. The compressors in

the process are used to increase the pressure of a vapor streams, essentially acting as a pump

for vapor streams. The compressors used in both processes were modeled as carbon steel gas

engine reciprocating compressors with a pressure rating of 7000kPa. The operating conditions

for the compressors in the re-injection process can be seen in Table 33:


68

Table 33. Compressor operating conditions (re-injection process)

Compressor Mass Flow

Rate

(kg/hr)

Inlet Temperature

(oC)

Inlet Pressure

(kPa)

Outlet Temperature

(oC)

Outlet Pressure

(kPa)

Inlet Vapor

Fraction

Outlet Vapor

Fraction

Overhead 2 Compressor

3525 10.09 1138 106.4 4171 0.999 1.000

Re-Injection Compressor

1

2.075x105 20.49 4171 75.95 7735 1.000 1.000


2

2.075x105 75.95 7735 126.4 1.294x104 1.000 1.000


3

2.075x105 126.4 1.294x104 161.2 1.810x104 1.000 1.000

The operating conditions for the compressors in the expansion plant process can be

seen in Table 34:

Table 34. Compressor operating conditions (expansion plant process)


Rate

(kg/hr)

Inlet Temperature

(oC)

Inlet Pressure

(kPa)

Outlet Temperature

(oC)

Outlet Pressure

(kPa)

Inlet Vapor

Fraction

Outlet Vapor

Fraction


3801 10.04 1138 104.3 4171 0.999 1.000

Propane Compressor

1

9.759x104 -39.16 106.5 19.18 413.7 1.000 1.000

Propane Compressor

2

1.354x105 12.49 399.9 68.28 1289 1.000 1.000

Sales Compressor

1.991x105 -39.77 3447 -2.39 5601 1.000 1.000

The cost of the compressors for each of the processes was estimated using the power

requirement and the costing equations. The adiabatic efficiency of each of the compressors is

75% and the polytropic efficiency calculated with the Shultz polytropic method is between 76-


69

77%. The power requirement, as well as the cost for each of the gas engine reciprocating

compressors for the re-injection process can be seen in Table 35:

Table 35. Compressor operating conditions and cost (re-injection process)


Rate (kg/hr)

Adiabatic Efficiency

Power Requirement

(kW)

Purchase Cost

Overhead 2 Compressor 3525 75.0% 167 $ 272,000

Re-Injection

Compressor 1 2.075x105 75.0% 5989 $ 8,050,000

Re-Injection

Compressor 2 2.075x105 75.0% 5988 $ 8,040,000

Re-Injection

Compressor 3 2.075x105 75.0% 4537 $ 6,190,000

Total Purchase Cost of Compressors $ 22,550,000

The power requirement, as well as the cost for each of the gas engine reciprocating

compressors for the expansion plant process can be seen in Table 36:

Table 36. Compressor operating conditions and cost (expansion process)


Rate (kg/hr)

Adiabatic Efficiency

Power Requirement

(kW)

Purchase Cost

Overhead 2 Compressor 3801 75.0% 172 $ 280,000

Propane

Compressor 1 9.759x104 75.0% 2241 $ 3,172,000

Propane

Compressor 2 1.354x105 75.0% 2993 $ 4,172,000

Sales Compressor 1.991x105 75.0% 3279 $ 4,548,000

Total Purchase Cost of Compressors $ 12,171,000


70

Valves

Valves are necessary for the reduction of pressure and temperature of the streams for which

they are implemented. Valves were used on the following streams in the expansion process:

Three-phase separator feed (1.508x104 kPa 1.038x104 kPa)

Feed stream to the TEG flash drum (4123 kPa 689.5 kPa)

Feed stream to the TEG regeneration column (620.5 kPa 206.8 kPa)

Feed stream to the Sales Gas column (4068 kPa 3516 kPa)

Feed stream to the LPG column (3516 kPa 1758 kPa)

C5+ recycle stream (1758 kPa 1379 kPa)

Propane liquid stream (1220 kPa 427.5 kPa)

Propane recycled liquid stream (413.7 kPa 123.7 kPa)

The only valve utilized in the re-injection process is the valve on the three-phase separator

feed. Most of the valves were modeled as carbon steel butterfly construction diaphragm valves,

the cost of which is based on their nominal diameter. In order to find the nominal diameter, the

cross–sectional area was calculated using the velocity, ν, and volumetric flow rate, Q, obtained

from HYSYS. The calculation for cross-sectional area, Ac, can be seen in Equation 20.

Equation 20. Cross-sectional area calculation

𝐴𝑐 =𝑄

𝑣

Using the equation for cross-sectional area the nominal diameters were calculated. The

two propane valves were modeled as carbon steel gated flanged valves, the cost of which is

also based on their nominal diameter. The calculation of the nominal diameter was performed in

the same way as for diaphragm valves. The valve specifications and purchase costs can be

seen in Table 37.


71

Table 37. Valve operating conditions and costs

Valve Velocity

(m/s)

Volumetric Flow Rate (m3/s)

Cross Sectional Area (m2)

Diameter (m)

Purchase Cost

Inlet Valve 232.80 0.64 0.0028 0.059 $ 2,000

TEG Valve 1 0.41 0.0011 0.0027 0.059 $ 2,000

TEG Valve 2 0.46 0.0020 0.0043 0.074 $ 2,000

Sales Valve 581.8 1.34 0.0023 0.054 $ 2,000

LPG Valve 3.01 0.028 0.0094 0.109 $ 2,000

C5+ Recycle Valve

0.97 0.0058 0.0060 0.087 $ 2,000

Propane Valve 1

40.21 1.15 0.029 0.191 $ 3,000

Propane Valve 2

25.82 1.78 0.069 0.256 $ 2,000

Total Purchase Cost of Valves $ 17,000

Expansion Purchase Cost $15,000

Storage Tank

A storage tank was necessary in both processes for the storage of natural gasoline. The

storage tank was modeled to hold the daily production volume of natural gasoline. In order to

calculate the daily production volume the volumetric flow rate (gallons/min) was multiplied by the

amount of minutes in a day (1440 min/day), which can be seen in Equation 21.

Equation 21. Calculation of the Daily Production Volume

𝑉𝐷𝑃 = �̇� (𝑔𝑎𝑙𝑙𝑜𝑛𝑠

𝑚𝑖𝑛) ∗ 1440

𝑚𝑖𝑛

𝑑𝑎𝑦

The daily production volume was calculated to be 508,511 gallons, the tank was

assumed to have a volume equal to the daily production value. The storage tank was sized with

the assistance of an API standard tank sizing sheet provided by Sean Arendell. The dimensions

were specified using a design volume of 540,000 gallons, which was the daily production

volume for the expansion plant. Utilizing the larger tank initially saves from having to re-

purchase a new storage tank following the expansion of the plant. The cost of the storage tank

was estimated utilizing the online cost estimate website Matches (29). The storage tank was


72

modeled as a carbon steel API floating roof storage tank. The storage tank dimensions and

costs for the expansion plant can be seen in Table 38.

Table 38. Storage tank operating conditions and cost

Storage Tank

Volumetric Flow Rate (gallons/min)

Volume of Tank

(gallons)

Diameter (m)

Length (m)

Total Cost of Storage Tank

LNG Storage

Tank 374.73 5.396x105 14.63 12.19 $ 435,000

Utility Summary

Utility requirements were calculated for both the natural gasoline re-injection and the

expansion plant processes based on annual consumption rates. One exception in the expansion

plant process was the propane refrigeration cycle which is responsible for cooling the overhead

from the TEG dehydration cycle. For the refrigeration cycle, the cost of propane was a one-time

capital expense. The residence time of a unit of propane within the refrigeration determined the

required amount of propane. The residence time was estimated to be one hour (McKetta, 1994).

This scaling factor was multiplied by the hourly flow rate to yield a total consumption term. The

cost of propane is $50/ bbl. The initial glycol necessary for the TEG dehydration cycle was

calculated using a residence time of one hour, which was estimated to be 6949 lb. Unlike the

propane refrigeration cycle, the TEG dehydration loses approximately 36.5 lb/hr; therefore it

was necessary to purchase additional glycol as a utility to compensate for the hourly losses of

glycol.

Table 39 defines pertinent conversion factors. Importantly, the plant operated for 90 % of

the year or 7,884 hours of the 8,760 hours per year.

Table 39. Conversion table for utility summary

Conversion Factors

Factor Value Units

Operating Time 0.90

Volumetric Conversion 264.17 US gallons/m3

Density Water, 20oC 998.2 kg/m3

Hours in a Year 8,760 h/yr

Operating Time 7,884 h/yr


73

Each of the condensers used in both processes utilized chilled water as the utility fluid.

The reboilers were serviced by low and high pressure steam. The inlet heater present in both

processes used low pressure steam as the utility fluid. In addition to heat exchangers, the

expansion plant process required electricity to run the air coolers and the pump in the process.

On-site sales gas was used to power the compressors in both of the proposed processes.

In order to properly dispose of the organic impurities found within the two water product

streams, it was necessary to subject the streams to waste water treatment. This is done in

accordance with the U.S. Clean Water Act of 1977 (25). The treatment costs were determined

to be $0.15 per pound of organic impurities.

Table 40 shows the utility summary for the natural gasoline re-injection process, which

excluded the use of propane refrigeration and TEG dehydration.


74

Table 40. Utility summary for the natural gasoline re-injection process

Condensers

Condenser Utility Type Mass (ton) Yearly Mass Requirement

(ton-day)

Unit Cost

($/ton-day)

Annualized Utility Cost

($)

C5+ Condenser Chilled Water 11.2 3680 $ 1.20 $ 4,000

Reboilers

Reboiler Utility Type Mass flow

Rate (kg/hr)

Yearly Mass Requirement

(kg)

Unit Cost ($/1000 kg)


($)

C5+ Reboiler 245.3 psig

Steam

1.14x104 8.96x107 $ 0.012 $ 1,050,000

Heat Exchangers

Heat Exchanger

Utility Type Mass flow

Rate (kg/hr)


(kg)



($)

Inlet Heater 50 psig

Steam

4.10x104 3.24x108 $ 0.0066 $ 2,140,000

Compressors

Compressor Utility Type

Heat Flow

Rate (MMBTU/hr)

Yearly Heat

Requirement (MMBTU)

Unit Cost ($/MMBTU)

Annualized

Utility Cost ($)


Sales Gas 1.79 1.41x104 $ 4.00 $ 57,000

Re-Injection Compressor 1

Sales Gas 64.3 5.07x105 $ 4.00 $ 2,030,000

Re-Injection

Compressor 2 Sales Gas 64.3 5.07x105 $ 4.00 $ 2,030,000

Re-Injection

Compressor 3 Sales Gas 48.7 3.84x105 $ 4.00 $ 1,540,000

Waste Water Treatment

Waste Stream Wastewater

Treatment

Mass Flow

Rate (lb/hr)


(lb)

Unit Cost ($/lb organic

removed)

Annualized

Utility Cost ($)


Organic Impurities

0.72 5.71x103 $ 0.15 $ 856

Total Utilities Cost $ 8,850,000


75

Table 41 shows the utility summary for the expansion plant process, which includes the

use of propane refrigeration and TEG dehydration.

Table 41. Utility summary for the expansion plant process

Condensers

Condenser Utility

Type Mass (ton)


(ton-day)

Unit Cost

($/ton-day)


($)

C5+Condenser Chilled Water

11.2 3680 $ 1.20 $ 4,000

TEG Condenser

Chilled Water

7.14 2340 $ 1.20 $ 3,000

LPG Condenser

Chilled Water

3510 1.15x106 $ 1.20 $1,390,000

Reboilers

Reboiler Utility Type

Mass flow Rate (kg/hr)

Yearly Mass Requirement (kg)



($)

C5+ Reboiler 245.3 psig

Steam

1.14x104 8.96x107 $ 0.012 $ 1,050,000

TEG Reboiler 285.3 psig

Steam

4.72x102 3.72x106 $ 0.012 $ 46,000

LPG Reboiler 150 psig

Steam

2.21x104 1.74x108 $ 0.011 $ 1,830,000

Sales Gas

Reboiler

50 psig

Steam

2.29x103 1.80x107 $ 0.0066 $ 119,000

Heat Exchangers

Heat Exchanger

Utility Type

Mass flow Rate (kg/hr)

Yearly Mass Requirement (kg)



($)

Inlet Heater 50 psig

Steam

4.10x104 3.24x108 $ 0.0066 $ 2,140,000

Air Coolers

Air Cooler Utility Type

Power

Requirement (kW)

Yearly Power

Requirement (kW-hr)

Unit Cost ($/kW-hr)

Annualized

Utility Cost ($)

Propane Air

Cooler

Electricity 1.47x102 1.16x106 $ 0.06 $ 70,000

TEG Air Cooler Electricity 9.10x10-1 7.17x103 $ 0.06 $ 430

Pumps

Pump Utility Type

Power Requirement

(kW)

Yearly Power Requirement

(kW-hr)

Unit Cost ($/kW-hr)


($)

TEG Pump Electricity 5.08 4.00x104 $ 0.06 $ 2,000


76

Compressors

Compressor Utility Type

Heat Flow

Rate (MMBTU/hr)

Yearly Heat

Requirement (MMBTU)

Unit Cost ($/MMBTU)

Annualized

Utility Cost ($)


Sales Gas 1.85 1.46x104 $ 4.00 $ 58,000

Propane Compressor 1

Sales Gas 24.0 1.90x105 $ 4.00 $ 760,000

Propane Compressor 2

Sales Gas 32.1 2.53x105 $ 4.00 $ 1,010,000

Sales Compressor

Sales Gas 35.2 2.77x105 $ 4.00 $ 1,110,000


Waste Stream Utility Type

Mass Flow

Rate (lb/hr)

Yearly Mass

Requirement (lb)

Unit Cost

($/lb organic removed)

Annualized

Utility Cost ($)


Wastewater Treatment

0.72 5.71x103 $ 0.15 $ 856

TEG Regeneration

Column

Wastewater

Treatment 35.7 2.82x105 $ 0.15 $ 42,000

Utility Fluid

Utility Fluid Utility Type Volumetric Flow Rate

(bbl/hr)

Required Utility Fluid Volume

(bbl)

Unit Cost ($/bbl)


($)

Propane (Startup)

Refrigeration 4.03x104 4.03x104 $ 50.00 $ 2,010,000

Utility Fluid Utility Type Mass Flow

Rate

(kg/hr)

Required Utility Fluid

Mass (kg)

Unit Cost ($/kg)


($)

TEG (Startup)

Dehydration 6950 6950 $ 0.65 $ 5,000

TEG (Regeneration)

Dehydration 27 2.16x105 $ 0.65 $ 141,000

Total Utilities Cost (First Year Startup) $ 11,800,000

Total Utilities Cost (Second Year Operating and Onward) $ 9,800,000


77

The total utility cost of the natural gasoline re-injection process is cheaper than that of

the expansion plant process, which can be justified by the addition of several extra unit

processes. The total utility cost for the natural gasoline re-injection process is $ 8,850,000. In

comparison, the expansion plant has first-year start up utilities cost of $ 11,800,000.

Subsequent years will be $ 9,800,000, which does not include the previously purchased

propane and TEG. The broader points of economics are discussed in the forthcoming variable

cost section.

Estimation of Capital Investment and Total Product Cost

Rigorous economic analyses provided keen insight on process design and projected

profit margins. All parameters were obtained from the course textbook, Seider’s Product

Process and Design Principles, and from the course notes, compliments of Dr. Sani at 8 AM on

Tuesdays and Thursdays during the fall semester 0f 2009. Analyses were completed in Dr.

Zartman’s Economics 2008 15 yr Oct 08 macro-enabled Excel® spreadsheet, which was

provided for Homework #9 in CHEN 4520 on the CULearn course website. This spreadsheet

will hereafter be referred to as the economics spreadsheet. There were minor disconnects and

semantics between Seider and Zartman’s discussions of profitability; however, both outlines tout

a 50% accuracy range. Rigorous profitability analyses were completed for both processes and

the expansion project is a solid investment in terms of profitability.

Economic Premises

Venture Guidance Appraisal

Site factor for the Yamal Peninsula is 1.65 per Mr. Arendell

Miscellaneous equipment costs are 10% of total engineering equipment/purchased and

delivered

Field maintenance, labor, and insulation are 5%, 10%, and 10%, respectively, of

purchased equipment and delivered cost

Equipment foundations, supports, and platforms are 10% of field maintenance, labor,

and insulation costs

Factored piping, instruments, and electrical were 22%, 9%, and 7%, respectively, of

installed equipment cost


78

The labor material split was 40% and 60%, respectively, of direct installed cost

Freight, quality insurance, and sales taxes were 12% of material costs

Contractor labor distributives were 44% of labor costs

Additional indirect costs were 15% of direct installed cost plus indirect freight, quality

insurance, taxes, and overhead

Buildings and structures were 20% of direct equipment costs

Power, general, and services were 2% of direct equipment costs plus building and

structures costs

Dismantling and rearranging were 2% direct equipment costs plus building and

structures costs

Site development was 5% of direct equipment costs plus building and structures for an

expansion project

Contingency was 35% of the direct permanent investment

Working conditions were 3% of labor costs

Inflation was 2.625% for every year

Start-up spare parts were 10% of total permanent investment

Variable Costs

High-pressure steam (285.3 psig) costs $12.30 per 1,000 kg from interpolated values

Mid-pressure steam (245.3 psig) costs $11.77 per 1,000 kg from interpolated values

Mid-pressure steam (150 psig) costs $10.50 per 1,000 kg

Low-pressure steam (50 psig) costs $6.6 per 1,000 kg

Chilled water costs $1.2 per ton-day

Electricity costs $0.06 per kW-hr


79

Triethylene glycol costs $0.65 per lb (29).

Sales gas costs $4 per MMBTU

Wastewater treatment is $0.15 per lb organic removed

LPG costs $1.31 per US gal

Fixed Costs

Number of operators is 5 for the C5+ process and 15 for the expansion process

Annual wages are $72,800 at $35 per operator-hr

Five shifts per day

Employee benefits are 15% of wages

Operating supervision is 17% of wages at annual wage of $60,000 per operator per shift

Operating supplies are 6% of wages

Maintenance is 3.5% of total permanent investment

Maintenance labor is 25% of total maintenance

Maintenance material is 100% of total maintenance

General overhead is 22.8% of operators’ wages plus maintenance labor plus operator

supervision

Lab and technical support is 6.91% of total permanent investment at $65,000 per

operator per shift

Sales and administration is 2% of total permanent investment

Research and development is 5% of total permanent investment

Insurance and local taxes are 3% of total permanent investment

Cash Flow

C5+ costs $1.90 per US gallon


80

Five years MACR depreciation

Interest of capital is 15%

Salvage percent is zero from the current process that produces 2,500 bpd

Accounts receivable are 30 days

Corporate income tax is 34%

Plant capacity is 50% in the first year, 75% in the second year, and 90% in the following

years

The lifetime of the plant is 15 years

The design period is one year and the construction period is two years

Capital Investment

The purpose of capital investment is to investigate the total capital investment (TCI).

This value factored into the internal rate of return (IRR), net present value (NPV), return of

investment (ROI), payback period (PBP), benefit-cost ratio (BCR), and break-even point (BEP),

which will be discussed in turn, to determine profitability. The net cash flow, the final calculation

of the aforementioned factor, resolved the viability of expanding on a current process to

maintain natural gasoline production and add sales gas and LPG recovery trains.

Cost Indices

Also pertinent to determining the capital investment is the cost index to account for

inflation. These indices are applicable for a month or two and are then obsolete as progress

charges onward. Here, the Chemical Engineering (CE) Plant Cost Index was used with an

overall value of 532.9 from April 2010, which was the most current value (30). The cost index is

used in determining the purchase cost in Equation 30.

Equation 22. Purchase cost adjustment with cost index. I was the current cost index, with a CE value of 532.9

from April 2010

𝐶𝑜𝑠𝑡 = 𝐵𝑎𝑠𝑒 𝐶𝑜𝑠𝑡 (𝐼

𝐼𝑏𝑎𝑠𝑒

)

This index accounts for the entire processing plant, including labor and materials of equipment

fabrication, delivery, and installation.


81

Commodity Chemicals

As aforementioned, the process produced energy products with market-sensitive prices

that were applied in the economic analyses.

Table 42 shows the current value of the natural gasoline product stream.

Table 42. Capacity and current market value of natural gasoline product stream

Unit LNG

Capacity (bpd) 1.21E+04

Capacity (bph) 504

Capacity (USgph) 21189

Sales Price ($/bbl) $ 80.00

Sales ($/h) $ 40,400

Sales ($/yr) $ 318,200,000

Table 43 shows the current value of cumulative product streams based on the flow rates

and market values of sales gas, LPG, and natural gasoline.

Table 43. Current market value of product streams

Unit Sales Gas Unit LPG C5+ Unit

Capacity 8.42E+08 BTU/hr 1662 1.28E+04 bpd

Capacity (units/hr) 842 MMBTU/hr 69 535 bph

Capacity (units/hr) N/A N/A 2908.5 22470 US gal/hr

Capacity (units/yr) 6634386 MMBTU/yr 22930614 177153480 US gal/yr

Sales Price $ 4.00 $/MMBTU $ 1.31 $ 1.90 $/US gal

Sales Price $ 4.00 $/MMBTU $ 55.00 $ 80.00 $/bbl Total

Sales ($/hr) $ 3,400 $/hr $ 4,000 $ 42,800.00 $/hr $ 50,000

Sales ($/yr) $ 26,500,000 $/yr $ 30,000,000 $ 337,400,000 $/yr $ 394,000,000

Total Permanent Investment (TPI)

The total permanent investment (TPI) of the natural gasoline separation process in an

expansion plant reflected a singular expense for the design, construction, and startup. As

published in the course textbook from Busche, the TPI was composed of sixteen separate costs

that covered the aforementioned expenses (31). In brief, any new project, including a grass-

roots plant, has a TPI containing these cumulative costs with the take home message that there

is never a free lunch. The Excel® spreadsheet Example_Economics2008 15 yr Oct 08 provided

values within the Venture Guidance Appraisal (VGA) that were beyond the scope of the course

textbook and were cited as such (32).


82

Both the natural gasoline scale-up and expansion projects were investigated for

profitability. A conservative approach was taken to estimate the bare-module costs of the

equipment; the salvage value of the current process that produces 2,500 bpd was assumed to

be zero. In this way, new equipment was designed for the scaled-up process that produces

10,000 bpd. The salvage value of the equipment used in the natural gasoline process was

assumed to be 100% for overlapping equipment in expanding the plant to produce sales gas

and LPG product streams.

Estimates from the natural gasoline and then the expansion process will be presented

sequentially. The following discussion outlines the methodology behind the VGA.

Bare-Module Cost

Bare-Module Cost (BMC)/Direct Installed Cost (DIC) are reflected in the total bare-

module investment (TBM).The BMC may be primarily divided into process equipment and

fabricated machinery within the VGA sheet within the economic spreadsheet with other related

costs.

Process machinery with standard designs, such as valves and storage tanks, was

chosen from a standard supply list and is shown in Table 44:


83

Table 44. Process machinery with approximate costs

Unit C5+

Approximate Cost (k)

Expansion

Approximate Cost (k)

Valves $ 2 $ 16

Storage Tank $ 435 -

Total Engineered

Equipment/Purchased & Delivered $ 437 $ 16

Associated costs of process equipment are presented in Table 45:

Table 45. Indirect costs associated with purchase and installation

Cost Percentage

(%) Associated Cost Source

C5+ Cost

(k)

Expansion

Cost (k)

Misc Equipment 10

Total Engineered Equipment/Purchased

& Delivered (32) $ 44 $ 2

Subtotal/Purchased Equipment & Delivered $ 481 $ 18

Field Material 5

Subtotal/Purchased Equipment &

Delivered (32) $ 24 $ 1

Labor 10


Delivered (32) $ 48 $ 2

Insulation 10 Subtotal/Purchased

Equipment &

Delivered (32) $ 48 $ 48

Field Erected Equipment

0


Delivered (32) $ 0 $ 0

Equipment Foundations,

Supports,

Platforms

10


Delivered and Field

Mtl/Labor/Insulation

(32) $ 61 $ 61

Installed Equipment $ 672 $ 25


84

Factored Piping 22 Installed Equipment (32) $ 148 $ 5

Factored

Instruments 9 Installed Equipment (32) $ 60 $ 2

Factored Electrical

7 Installed Equipment (32) $ 47 $ 2

No Identified Piping,

Instruments, or

Electrical

0 - (32) - -

Subtotal, Direct Installed Cost $ 927 $ 34

Labor Split 40 Subtotal, Direct

Installed Cost (32) $ 371 $ 14

Material Split 60 Subtotal, Direct

Installed Cost (32) $ 556 $ 20

Freight, Quality Assurance, Sales

Taxes 12 Material (32) $ 67 $ 2

Contractor Labor

Distributives 44 Labor (32) $ 163 $ 6

Subtotal (Direct Installed Cost + Indirect Freight, QA, Taxes, &

Overhead $ 1,157 $ 42

Engg+Home Office (Additional

Indirect) 15

Subtotal (Direct Installed Cost +

Indirect Freight, QA,

Taxes, & Overhead

(32) $ 174 $ 6

Subtotal (DIC Equipment Calculated from Bare Module using PE) $ 1,331 $ 49

Fabricated machinery is specific to the process at hand, such as the distillation columns,

flash drums, air coolers, heat exchangers, and the pump. These unit operation costs contain a

module cost to account for the piece of equipment and the installation, including piping to and

from, concrete foundation, ladders and other supporting structures, instruments, controllers,

lighting, electrical wiring, insulation, and painting. This factor also assumed free on board (f.o.b.)

delivery where the purchase cost did not include the price of delivery to the plant site. This is a

liberal approach where the delivery charge to the Yamal Peninsula may be augmented due to

the remoteness.


85

The bare-modules used here in Table 46 were obtained from Guthrie (33).

Table 46. Bare-module factors for process to account for design costs above those encountered in

purchasing process machinery

Unit

Bare-Module Factor

(FBM)

C5+ PE

Cost (k)

Expansion

PE Cost (k)

C5+ Bare-Module

Cost (k)

Expansion Bare-Module

Cost (k)

Distillation

Columns 4.16 $ 199 $ 1,132 $ 828 $ 1,701

Compressors 2.15 $ 22,547 - $ 48,476 -

Flash Drums 4.16 $ 121 $ 1,052 $ 503 $ 4,376

Air Coolers 2.17 - $ 518 - $ 1,124

Heat Exchangers

(Double Pipe) 1.8 $ 2 $ 7 $ 4 $ 13

Heat Exchangers

(Shell and Tube) 3.17 $ 69 $ 223 $ 219 $ 707

Pumps 3.3 - 8 - $ 26

Subtotal (DIC from Total Bare Module Cost w/FBM Factors) $ 50,030 $ 10,955

As can be seen in Figure 12, compression is the primary bare module cost for the

natural gasoline process.


86

Figure 12. Bare-module distribution for natural gasoline process

The compressors serve to re-inject the overhead gas. These compressors were designed to the

maximum limit of the provided costing equations. Therefore, compressors with a larger power

capacity may be available and this is a conservative cost estimate. Fewer and larger

compressors would decrease the presented price.

By salvaging the compressors and other process units from the natural gasoline

process, the expansion process has a more evenly distributed bare module cost allocation, as

seen in Figure 13:

Compressors96%

Distillation Columns2%

Flash Drums1%

Total Engineered Equipment / Purchased &

Delivered1% Heat Exchnagers

(Shell and Tube)0%

Heat Exchangers (Double Pipe)

0%

Bare Module Costs for C5+ Process


87

Figure 13. Bare-module distribution for expansion process

The drums used in the propane refrigeration cycle were designed as vessels without internals;

however, lower price options may be available.

Associated costs of fabricated equipment are presented in Table 47:

Table 47. Fabricated equipment-associated miscellaneous cost

Cost Percentage


C5+ Cost

(k)

Expansion

Cost (k)

Miscellaneous

Equipment 10

Subtotal (DIC from Total Bare Module Cost

w/FBM Factors) (32) $ 5,003 $ 1,096

Subtotal (DIC Equipment from Bare Module Costs) or Subtotal (DIC

Equipment Costs) $ 56,363 $ 12,100

Purchase costs were estimated using the Excel® spreadsheet Formated cost eqns Basis

CE500 Oct 2008, which was provided for Homework #9 on the CULearn course website and

matche.com, an online sizing and costing resource (29). All unit operating parameters aligned

Heat Exchangers (Double Pipe)

0%

Total Engineered Equipment /

Purchased

& Delivered0%

Pump0% Heat Exchnagers

(Shell and Tube)7%

Air Coolers10%

Flash Drums40%

Distillation Columns43%

Bare Module Costs for Expansion Process


88

with or were below the prescribed ranges, and the purchase costs were adjusted to the current

CE index.

Site Preparation

Site preparation included land surveys, dewatering and drainage, surface clearing, rock

blasting, excavation, grading, and piling (25). Upon construction, fencing, roads, sidewalks,

railroad sidings, sewer lines, fire protection facilities, and landscaping were also included in this

cost. Table 48 tabulates the site preparation costs for an expansion plant:

Table 48. Site preparation for an expansion plant

Cost Percentage


C5+ Cost

(k)

Expansion

Cost (k)

Buildings,

Structure 20

Subtotal (DIC

Equipment Costs) (32) $ 11,273 $ 2,240

Subtotal $67,636 $ 24,520

Service Facilities

Service facilities included utility lines, control rooms, laboratories for quality control,

maintenance shops, and other buildings (25). For the expansion process at hand, the growth of

administrative offices, medical facilities, cafeterias, garages, and warehouses was also

considered, as shown in Table 49:

Table 49. Service facilities for an expansion plant

Cost Percentage

(%) Associated

Cost Source

C5+ Cost (k)

Expansion Cost (k)

Power, General, &

Services (PG&S) 2 Subtotal (32) $ 1,353 $ 290

Dismantling &

Rearranging (D&R) 2 Subtotal (32) $ 1,353 $ 290

Site Development 5 Subtotal Expansion

(25) $ 3,382 $ 726

Subtotal (DPI) $ 73,723 $ 15,826


89

Contingencies and Contractor’s Fee

Contingencies and contractor’s fee were unanticipated costs incurred during

construction that were augmented from 15% of DPI to 35% for a student team design (25), as

shown in Table 50:

Table 50. Contingencies and contractor’s fee for an expansion plant

Cost Percentage

(%) Associated

Cost Source

C5+ Cost (k)

Expansion Cost (k)

Contingency 35 Subtotal

(DPI) Student design

team (34) $ 25,803 $ 5,539

Subtotal $ 99,526 $ 21,366

Working Conditions 3 Labor

Contractor’s Fees useful

estimate (34) $ 1,194 $ 256

Net Total $ 100,720 $ 21,622

Minor Changes, Field Indirects,

Spares and

Portables

0 Subtotal (32) - -

Direct Total $ 100,720 $ 21,622

Total Equipment, Total (Current USGC) $ 100,720 $ 21,622

Investment Site Factor

The investment site factor, FISF, accounted for the nuances of location, such as

availability of labor, the efficiency of the workforce, local rules and customs, and union status

among other contributing factors (25). The proposed plant operating site is on the Yamal

Peninsula, Russia with a FISF of 1.65, thus augmenting the total permanent investment.

Equation 23 shows how the site affects the total permanent investment.

Equation 23. Corrected total permanent investment to account for building and operating in the Yamal

Peninsula, Russia with a FISF of 1.65

𝐶𝑇𝑃𝐼𝑐𝑜𝑟𝑟𝑒𝑐𝑡𝑒𝑑 = 𝐹𝐼𝑆𝐹𝐶𝑇𝑃𝐼

Table 51 illustrates the contribution of the site factor to total cost:


90

Table 51. Site factor for plant operating in the Yamal Peninsula, Russia (FISF=1.65) (25)

Cost Percentage


C5+ Cost

(k)

Expansion

Cost (k)

Site

Factor 100

Total (Current

USGC)

Siberia, Russia

(1.65) $ 166,188 $ 35,676

Inflation

Inflation is the change in value of currency over time and serves as a predictive measure

for the long-term viability of the process. Seeing as depreciation allowances are not adjusted for

inflation, an inflation analysis is required. Furthermore, revenues and costs increase with

inflation, causing gross earning to increase, yielding a higher income tax. Average inflation rates

for pertinent goods are shown in Table 52:

Table 52. Average inflation rates

Cost Inflation (%)

Raw materials and price of products 2.5

Utilities 2.5

Processing Equipment 2.5

Hourly labor 3.0

Average 2.625

The effect of inflation according to Equation 24 on the process is shown in Table53.

Equation 24.Inflation calculation

𝐹 = 𝑃(1 + 𝑖)𝑛

Where F is future worth, P is corrected CTPI, i is inflation rate, and n is number of years.


91

Table 53. Inflation table with average inflation rates

Cost Percentage


C5+ Cost

(k)

Expansion Cost

(k)

Inflation 2.625 for 1

year

Total (Current

USGC) (25) $ 170,550 $ 36,612

Scope

Growth 0 Inflation (32) - -

Total Project-Level Cost $ 170,550 $ 36,612

GRAND TOTAL (TPI) $ 170,600 $ 36,600

The TPI represents a major component of the appraisal and later serves as a benchmark for

sensitivity analyses.

Working Capital (WC)

Working capital funds covered expenses incurred during the startup period before a

profit was realized, i.e. year four. These expenses included cost of inventory and funds to cover

accounts receivable, and the values are shown in Table 54:

Table 54.Working capital

Cost Quantity Associated

Cost Source

C5+ Cost

(k)

Expansion

Cost (k)

TEG 7000 lb $ 0.65 (29) - $ 5

Refrigeration Propane

40,231 bbl

$ 50 Project

Proposal - $ 2,011

Total - $ 2,015

Start-Up Spare Parts

10 % GRAND TOTAL

(TPI)

Typical Estimate

(34) $ 17,060 $ 3,660

Total Working Capital $ 17,060 $ 5,675

Operating Cost

The total annual cost of manufacture (COM) reflects the sum of (1) direct manufacturing

stocks: utilities; (2) operating overhead: labor-related operations and maintenance; and (3) fixed

costs: property taxes, insurance, and depreciation.


92

Variable Cost

Utilities were assessed on a consumption basis. The itemized list may be referenced in

Table 55:

Table 55. Utility summary in annual costs and normalized costs to natural gasoline production

Utility C5+ Annual

Cost (k) C5+ Cost per Gal

of Product Expansion Annual

Cost (k)

Expansion Cost per Gal of

Product

Sales Gas $5,640,000 $0.03 $2,940,000 $ 0.02

LP Steam

(50 psig) $2,135,000 $0.01 $2,254,000 $ 0.01

HP Steam (150 Psig)

- - $1,830,000 $ 0.01

HP Steam (245.3 Psig)

$1,054,000 $0.01 $1,054,000 $ 0.01

Chilled Water $4,000 $0.00 $1,393,000 $ 0.01

Electricity - - $72,000 $ 0.00

TEG - - $141,000 $ 0.00

HP Steam

(285.3 Psig) - - $46,000 $ 0.00


$1,000 $0.00 $43,000 $ 0.00

Subtotal Utilities $8,834,000 $0.053 $ 9,773,000.00 $ 0.055

A distribution of the utilities within the two processes is illustrated in Figures 14 and 15,

respectively:


93

Figure 14. Utility distribution within the natural gasoline process.

Sales gas required to run the re-injection compressors dominated the cost. On the other

hand, the expansion process manifests a more even distribution of utilities, as seen in Figure

15:

SALES GAS64%

HP STEAM (245.3 psig)

12%

LP STEAM (50 psig)

24%

CHILLED WATER0% WASTE WATER

TREATMENT0%

Utilities Cost of C5+ Process


94

Figure 15. Utility distribution within the expansion process

These utility allocations merit additional consideration for an integrated heat exchange

approach with process streams or an expanded propane cycle.

The sales gas and LPG product streams were accounted for as byproduct credits within

the variable cost of the profitability analysis. Table 56 illustrates the gained revenue in isolating

the sales gas and LPG products in the expansion process:

Table 56. Byproduct credits gained from sales gas and LPG streams in the expansion process

Byproduct Expansion Annual Cost (k) Expansion Cost per Gal of Product

Sales Gas -$ 26,500 -$ 0.15

LPG -$ 30,000 -$ 0.17

Total Variable Costs

-$ 0.264

($46,838)

The additional credits from the byproducts offset the utility costs and are demarcated as

a profit.

SALES GAS30%

LP STEAM (50 psig)

23%

HP STEAM (150 psig)

19%


11%

CHILLED WATER14%

ELECTRICITY1%

TEG1%


1% WASTE WATER TREATMENT

0%

Utilities Costs for Expansion Process


95

Fixed Cost

Labor-related operations account for the lab hours required to produce the annual

capacity. Annual wages were assessed at the hourly scale for plant operators and at the annual

salary scale for technical assistance and control laboratory operators.

Operating Labor and Benefits

The natural gasoline process was operated as a single operation. On the other hand, the

expansion process was divided into four fluids processing sections, each with a single operator

requirement that varied based on unit multiplicity (25).

Table 57 outlines the operating and pay parameters for a 24-hour, seven-days-a-week

schedule. The required number of weekly shifts is 4.2; however, rounding this figure up to 5

shifts per week was required due to illness, vacations, holidays, training, special assignments,

and overtime during startups (25).

Table 57. Labor-related operations parameters used to calculate fixed costs

The required number of operators was determined by counting the engineered, process

support, and key process pieces of equipment within each separation process. Firstly, it was

assumed that ¼ of an operator was required for each engineered and each process support

piece of equipment. Secondly, it was assumed that each key process piece of equipment, such

as columns, required a whole operator.

Labor-Related Operations Parameters

Plant operation 7 dy/wk

Plant operation 24 hr/dy

Plant operation 168 hr/wk

Full-time workforce 40 hr/(wk-operator)

Required shifts 4.2 shifts/wk

Rounded up required shifts 5 shifts/wk

Hourly wage for operators $ 35.00 /operator-hr

Full-time workforce 40 hr/wk

Full-time pay 52 wk/yr

Full-time pay 2080 hr/yr


96

Operator requirements for the natural gasoline and expansion processes are shown in

Tables 58 and 59, respectively.

Table 58. Operator requirement for the natural gasoline process

Number of Units

Equipment Inlet Separation

Engineered 3

Purchase Support 7

Key 2

Operators 4.5

Table 59. Operator requirement for the expansion process

Number of Units

Equipment Inlet

Separation TEG

Dehydration Propane

Refrigeration Sales/LPG Recovery

Total

Engineered 3 2 2 2 9

Purchase Support 4 5 6 4 19

Key 2 3 1 2 8

Operators 3.75 4.75 3 3.5 15

Table 60 illustrates total labor-related annual cost. The expenses herein include

technical assistance to manufacturing with an annual salary of $60,000 and the control

laboratory with an annual salary of $65,000, each at one operator per shift for five per week,

yielding ten in total (25).


97

Table 60. Operating labor and benefits

Cost Percentage

(%)

Associated

Cost Source

C5+ Cost

(k)

Expansion Cost

(k)

Annual Wages N/A N/A N/A $ 364 $ 1,092

Employee

Benefits 15 Wages (25) $ 55 $ 164

Operating

Supervision

$300𝑘

$1,820𝑘= 16.5%

Wages (25) $ 62 $ 186

Subtotal operating labor $ 480 $1,441

Operating

Supplies 6 Wages (25) $ 22 $ 66

Maintenance

Maintenance is required to keep all processing equipment in acceptable working order

according to a preventative maintenance schedule. This requires spares and parts represented

by material and labor. Table 61 outlines the expenses related to maintenance. Maintenance

labor is best utilized during the down time, here 5% of the year. The purpose of this labor is to

clean heat exchangers to curtail fouling, and the lubrication and replacement of mechanical

seals in pumps, and compressors (25).

Table 61. Maintenance on complete distillation process

Cost Percentage

(%) Associated

Cost Source

C5+ Cost (k)

Expansion Cost (k)

Total maintenance

3.5 Investment (25) $ 5,971 $ 1,281

Maintenance

labor 25

Total

maintenance (25) $ 1,493 $ 320

Maintenance

material 100

Total

maintenance (25) $ 5,971 $ 1,281

Overhead

Overhead costs are non-plant operational expenses. Instead, these costs account for

cafeteria; employment and personnel; fire protection, inspection, and safety; first aid and


98

medical; industrial relations; janitorial; purchasing, receiving, and warehousing; automotive and

other transportation; and recreation. These costs are categorized into four sections that sum to

the general overhead cost as shown in Table 62. Employee appreciation mantivities are

included in the recreation cost within the business services expense (35).

Table 62. Overhead on complete distillation process

Cost Percentage (%) Associated

Cost Source

C5+ Cost (k)

Expansion Cost (k)

General plant overhead 7.1

Operator Wages +

Maintenance Labor + Operator

Supervision

(25) - -

Mechanical department 2.4 (25) - -

Employee relations

department 5.9 (25) - -

Business services 7.4 (25) - -

General Overhead

(sum of above costs) 22.8 (25) $ 437 $ 364

Lab and technical support $325𝑘

$4,700𝑘= 6.9% Investment (25) $ 11,771 $ 2,525

Corporate Overhead

Corporate overhead costs cover sales and administration expenses to ensure that the

products earn a fair market price. Additionally, investments are made in research and

development to maintain a competitive edge and improve efficiencies. The salary of the

proposed CEO, Mr. Wolff MS, is included in this category. These expenses are outlined in Table

63:

Table 63. Corporate overhead on complete distillation process

Cost Percentage

(%)

Associated

Cost Source

C5+ Cost

(k)

Expansion

Cost (k)

Sales and

administration 2 Investment (25)

$ 3,412 $ 732

Research and development

4.8 Investment (25) $ 8,530

$ 1,830

Subtotal corporate overhead $ 11,492 $ 2,562


99

Insurance and Local Taxes

Annual property taxes are levied by the Yamal Peninsula municipality separate from

those from the Russian equivalent to the Internal Revenue Service (IRS) (25). A local tax of 2%

on the investment was assumed (25). In reality, Gazprom’s relations with the Kremlin may

confound this assumption. Insurance is assessed based on pressure and temperature levels of

plant operations. The use of hazardous materials may also augment the insurance cost. In this

way, the process operated in a well-controlled manner for an insurance rate of 1% on the

investment. Table 64 defines these costs:

Table 64.Insurance and local taxes assessed annually

Cost Percentage

(%)

Associated

Cost Source

C5+ Cost

(k)

Expansion

Cost (k)

Insurance and

local taxes 3 Investment (25) $ 5,118 $ 1,098

Royalties 0 Per kg annual

capacity - -

Depreciation 0 Investment (32) - -

Total fixed cost (for cash flow calculations) ($0.21 per Gal)

Total fixed cost (for ROI calculations) ($0.21 per Gal) $ 35,742 $ 9,338

Royalties are null for this process seeing as there were no known intellectual property

infringements. Likewise, the total fixed cost was used in the cash flow calculation; therefore,

depreciation was not assessed. Nevertheless, typical values are 8% of total depreciable capital

for a 12-year plant life.

Profitability Analysis

Suffice it to say that profitability was the crux determination of feasibility. The primary

concern was to determine whether the investment to expand the current operation to yield sales

gas and LPG product streams was favorable over 15-year plant operation expectancy. These

factors will be discussed in turn.


100

Profitability

Profitability was assessed for current market selling prices of natural gasoline, sales gas,

and LPG. Finance terms are defined herein to characterize the arguments.

Cost of Capital

Fittingly, the cost of capital is an annual discount rate that reflects the cost of borrowing.

Therefore, the cost of capital is equal to the lender’s required return on investment.

Net Present Value

Net present value (NPV) gives the value of an investment by using a discount rate and a

series of future payments and income. Equation 25 gives the definition of NPV:

Equation 25. Net present value using cost of capital rate.

𝑁𝑃𝑉 = ∑𝑣𝑎𝑙𝑢𝑒𝑠𝑖

(1 + 𝑟𝑎𝑡𝑒) 𝑖

𝑛

𝑖=1

Here, the rate was the cost of capital, values were net cash flows, and inflation, i, was 2.625%,

as previously described. Each component of the summation represents one discounted cash

flow. The NPV is sensitive to changing interest rates due to the exponential denominator.

Internal Rate of Return

The internal rate of return is the interest rate that yields a net present value of zero

based on payments and income that occur at regular periods, i.e. yearly. This is an iterative

process to yield the value, 𝑁𝑃𝑉{𝑟} = 0. The Excel® function requires a guess that is close to the

expected IRR. The macro breaks down if the guess is too far astray.

Return on Investment

The return on investment (ROI) is the annual interest rate made by the profits of the

original investment. Equation 26 gives the broad definition of ROI (25).

Equation 26. ROI definition

𝑅𝑂𝐼 =𝑛𝑒𝑡 𝑒𝑎𝑟𝑛𝑖𝑛𝑔𝑠

𝑡𝑜𝑡𝑎𝑙 𝑐𝑎𝑝𝑖𝑡𝑎𝑙 𝑖𝑛𝑣𝑒𝑠𝑡𝑚𝑒𝑛𝑡


101

This is a profitability measurement that does not account for the size of the venture. In other

words, it may behoove a large company with hefty capital to invest in a separate solvent waste

recovery process. Whereas the large company possesses the capital to minimize loans and

earn a greater ROI, the smaller venture must borrow and realize a smaller ROI to account for

the loan payments. Here, the third year, the first year of full operating capacity serves as the

benchmark for the ROI.

Break-Even Point

The break-even point (BEP) is the time required for the cumulative annual earnings to

equal the original investment, as defined by the total depreciable capital divided by the cash

flow. BEP is widely used to compare alternatives but not to make final decisions due to the

inability to account for plant operation after the BEP.

Benefit-Cost Ratio

A benefit-cost ratio (BCR) seeks to evaluate the service life of the project by dividing

positive by negative, non-discounted cash flows. A company may rank potential projects by

BCC and reject any project with a value of less than unity.

Depreciation

Depreciation is a tax shield in that a company may treat depreciation as a cost of

production. This cost is the decline of the book value of each piece of capital equipment with

time, thereby reducing income tax liability, although there is no representative cash outflow from

the company. In this way, the age-old mantra that “a dollar today is more valuable than a dollar

tomorrow” supports the notion to wisely invest in a process today with the intent to reap the

profits tomorrow (36). The analysis used Modified Accelerated Cost Recovery System (MACRS)

on a five-year schedule, as shown in Table 65:


102

Table 65. MACRS Tax-Basis Depreciation

Percent of total depreciable capital (CTDC)

Year 5 Year

1 20.00

2 32.00

3 19.20

4 11.52

5 11.52

6 5.76

Total 100.00

MACRS is an initially accelerated depreciation model to allow companies to recoup a

greater percentage of capital investment, compared to straight-line depreciation which is

calculated by dividing the fixed costs by the number of years of operation (25).

Salvage Percent

Salvage percent is the value of the capital equipment at the end of the plant lifetime as a

percentage of the initial investment. A value of zero percent means that there is no worth to the

equipment upon plant retirement. The profitability analyses assumed that there was no salvage

value to the current process that produces 2,500 bpd.

Accounts Receivable

Accounts receivable are cash reserves to cover operating costs while the plant waits for

customers to fulfill obligation for product sales. A 30-day accounts receivable accounts for

8.22% of the annual sales of all products.

Corporate Income Tax

Current corporate income tax rate for companies making over $18,333,333 is 35% (37).

Cash Flow Analyses

The lifetime of the plant for natural gasoline, sales gas, and LPG was 15 years with one

initial design and two construction years. The profitability parameters are juxtaposed in Table

66. The cash flows of both processes were generated using parameters in Table 66.


103

Table 66. Profitability comparison of the different processes

Input C5+ Expansion

Cost of Capital (%) (25) 15 15

Inflation of all Costs (%) (25) 2.625 2.625

Inflation to Selling Price of Product (%) (25) 2.5 2.5

Accounts Receivable (dys) (25) 30 30

Income Tax (%) (37) 35 35

Land (38) - -

Total Capital Cost (k) $ 170,600 $ 36,600

Salvage Percent (%) (32) - -

Selling Price of Natural Gasoline $ 1.90 $ 1.90

Output C5+ Expansion

NPV (Cash Flows at End of Each Period) (k) $ 474,000 $ 823,000

NPV (Cash Flows at Beginning of Each Period) (k) $ 546,000 $ 947,000

IRR (%) 51.7 164.1

ROI (%) 69.5 314.5

Payback Period (yrs) 1.4 0.3

BEP (yrs) 3-4 3-4

BCR 14 87

The five pertinent benchmarks were drawn from the IRR, ROI, payback period, BEP,

BCR, and net profit. Whereas the ROI, payback period, and BEP are initial estimators for

profitability, the IRR, and BCR shows the profitability over the course of the plant lifetime. For

these reasons, it is highly recommended to pursue the expansion project.

Natural Gasoline Production

The capital investment allocation chart for the natural gasoline process in Figure 16

illustrates the contributions of operating parameters:


104

Figure 16. Capital investment allocation for the natural gasoline process.

The TPI represents the greatest share of the capital investment. Therefore, any savings

in this parameter would enhance profitability.

Non-discounted cash flows of the natural gasoline process are presented in Figure 17:

TPI79%

TOTAL FIXED

COST (for ROI

calculations):17%

TOTAL VARIABLE COST

4%

Capital Investment Allocation for Natural Gasoline

105

Figure 17. Non-discounted cash flows for the natural gasoline and re-injection process with a selling price of $1.90 for natural gasoline

The negative cash flows incurred during the first three years represent the cost of design, construction, and working capital. The cash

flows become positive in 2013 upon production at 50% capacity and then increase as capacity jumps to 90%. This increase in

-$150,000

-$100,000

-$50,000

$0

$50,000

$100,000

$150,000

$200,000

$250,000

2010 2011 2012 2013 2014 2015 2016 2017 2018 2019 2020 2021 2022 2023 2024 2025 2026 2027

Non-Discounted Cash Flow for Natural Gasoline Process

106

capacity reflects the progress of the manufacturing team as unforeseen nuances are resolved.

The overall upward trend upon startup suggests that production ought to continue, so long as

demand exists for a profitable selling price.

Figure 18 illustrates the BEP where the curve crosses the abscissa. Cash flows are

given at the end of the year with no clear delineation of more regular cash flows, i.e. monthly.

107

Figure 18. Non-discounted net values plotted against the years of operation to demonstrate the BEP occurring in the third year

Accordingly, the BEP is somewhere in the third year, a highly favorable estimate for investment.

-150000

-100000

-50000

0

50000

100000

150000

200000

250000

2010 2011 2012 2013 2014 2015 2016 2017 2018 2019 2020 2021 2022 2023 2024 2025 2026 2027

No

n-D

isco

un

ted

Net

Val

ue

($k)

Non-Discounted Net Value vs. Year of Operation for Natural Gasoline and Re-Injection Process

108

Expansion Process

The capital investment allocation chart for the expansion process in Figure 19 illustrates

the contributions of operating parameters:

Figure 19. Capital investment allocation for the expansion process

Again, the TPI represents the primary contributor to capital investment, albeit to a lesser extent

than observed in the natural gasoline process.

Non-discounted cash flows of the natural gasoline process are presented in Figure 20:

TPI66%

TOTAL VARIABLE COST17%

TOTAL FIXED COST (for ROI calculations):

17%

Capital Investment Allocation for Expansion Proces

109

Figure 20. Cash flow for the expansion process with a selling price of $1.90/gal for natural gasoline

The negative cash flows incurred during the first three years represent the cost of design, construction, and working capital. The cash

flows become positive in 2013 upon production at 50% capacity and then increase as capacity jumps to 90%. This increase in

capacity reflects the progress of the manufacturing team as unforeseen nuances are resolved. Again, the overall upward trend upon

startup suggests that production ought to continue, so long as demand exists for a profitable selling price.

-$50,000

$0

$50,000

$100,000

$150,000

$200,000

$250,000

$300,000

$350,000

2010 2011 2012 2013 2014 2015 2016 2017 2018 2019 2020 2021 2022 2023 2024 2025 2026 2027

Ca

sh F

low

($

k)

Non-Discounted Cash Flow for Expansion Process


110

Discounted cash flows represent the current investment required to meet the IRR of the process. These flows are shown in

Figure 21:

Figure 21. Discounted cash flows for the expansion process at an IRR of 164.1%

At an IRR of 164.1%, there is a clear effect of the time value of money such that a modest investment with long-range

foresight continually compounds to match the non-discounted cash flows of the process. These discounted cash flows are used to

calculate the IRR; therefore, the limit at the plant lifetime is zero.

Figure 22 illustrates the BEP where the curve crosses the abscissa. Cash flows are given at the end of the year with no clear

delineation of more regular cash flows, i.e. monthly.

-6000

-4000

-2000

0

2000

4000

6000

8000

2010 2011 2012 2013 2014 2015 2016 2017 2018 2019 2020 2021 2022 2023 2024 2025 2026 2027

Dis

cou

nte

d C

ash

Flo

w (

$k)

Discounted Cash Flows vs. Year of Operation for Expansion Process


111

Figure 22. Non-discounted net values plotted against the years of operation to demonstrate the BEP occurring in the third year for the expansion

process

Accordingly, the BEP is somewhere in the third year, a highly favorable predictor for investment.

-500000

0

500000

1000000

1500000

2000000

2500000

3000000

3500000

4000000

2010 2011 2012 2013 2014 2015 2016 2017 2018 2019 2020 2021 2022 2023 2024 2025 2026 2027

No

n-D

isco

un

ted

Ne

t V

alu

e (

$k)

Non-Discounted Net Value vs. Year of Operation


112

Sensitivity Analysis

Sensitivity analyses demonstrate the strength of the process in the face of changing

variables. Due to the volatility of the Russian political landscape, the sensitivity analyses were

ranged for 100% variability. All analyses were performed with the base parameter of $1.90 per

gallon of natural gasoline for a 164.1% IRR.

Present ROI and IRR for a +/- 100% Variation in TPI

The sensitivity analysis of a ±100% variation in TPI on ROI and IRR is shown in Figure

23:

Figure 23. Variation of ROI and IRR with respect to a 100% variation in TPI

Figure 23 demonstrates that both the ROI and the IRR non-linearly decrease with an

increasing TPI. Nevertheless, 100% variability in TPI, rooted in equipment purchase costs, still

yields favorable IRR and ROI rates.

0.00%

100.00%

200.00%

300.00%

400.00%

500.00%

600.00%

700.00%

800.00%

0 10000 20000 30000 40000 50000 60000 70000 80000

IRR

& R

OI

TPI ($k)

IRR & ROI vs. 100% Variability in TPI Sensitivity Analysis

IRR

ROI


113

Present ROI and IRR for a +/- 100% Variation in Fixed Operating Cost

The sensitivity analysis of a ±100% variation in TPI on ROI and IRR is shown in Figure

24:

Figure 24. Variation of ROI and IRR with respect to a 100% variation in Fixed Costs

The interpretation of these analyses follows the same reasoning as for variations in TPI.

However, here, both relationships are linear. As seen in Equation 26, fixed costs comprise the

denominator in the ROI calculation, thus increasing fixed costs, i.e. hiring more operators,

decreases ROI; there is no variable for worker efficiency and output. The IRR also decreases

with increased fixed costs such as hiring more operators, as this would result in greater costs as

a result of the increase in the number of salaries to be paid out with the same products being

produced, and thus a decrease in the internal rate of return to the company. Again, favorable

rates are observed for 100% variability in fixed costs.

Conclusion

The purpose of this endeavor to investigate the feasibility of expanding a natural

gasoline production facility to produce sales gas and LPG was achieved. The perpetual demand

0%

50%

100%

150%

200%

250%

300%

350%

400%

0 5000 10000 15000 20000

IRR

& R

OI

Fixed Costs ($k)

IRR & ROI vs. 100% Variability in Fixed Costs Sensitivity Analysis

IRR

ROI


114

for these products suggests that the investment is highly favorable. Safety and environmental

concerns are reasonable and practicable through adherence to local and federal requirements.

Two complete designs were generated and assessed for profitability. Both designs converged

with mostly closed material and energy balances. The outlier here was the energy balance on

the expansion process. Dissected balances suggest that recycle parameters and loose column

convergence parameters propagated to yield the energy imbalance. Both designs are capable

of producing product streams within specification. All units comprising the processes were

designed and specified via cited physical properties and assumptions. Neither of the processes

presents glaring manufacturing difficulties. Utility streams remain to be optimized by integration

with process streams for heat exchange demands.

The economic indicators for both processes were highly favorable. The variable costs

comprise a reasonable fraction of operating costs, as discerned by comparison with functioning

processes. It was determined that 15 operators working over five shifts with appropriate

supervision and control laboratory assistance are capable of generating a profit with highly

competitive profitability markers. A rigorous profitability analysis of both designs determined that

re-injection costs would readily be offset by selling byproduct credits on sales gas and LPG

product streams. In fact, these credits offset variable costs completely. The analysis used cited

parameters for each variable; nevertheless, these approaches are accurate to just 50%. In this

way, the ultimate decision selection relies on the nuances of investment. Companies rank

capital investment projects in order of any number of metrics. The presented profitability

discussion reveals that the expansion process yields favorable IRR, ROI, BCR, PBP, and BEP.

Sensitivity analyses with 100% changeability of the expansion process align with theory and

demonstrate the robustness of the process to market variability in TPI and fixed costs. Yotta

Designs strongly recommends the investment in the Yamal Peninsula expansion project.


115

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28. Brownell, Lloyd E and Young, Edwin H. Process Equipment Design. Hoboken : John Wiley & Sons,

Inc., 1959.

29. Matches. Matches' Process Equipment and Cost Estimates. Equip Costs. [Online] Matches, October

15, 2003. http://www.matche.com/EquipCost/.

30. ICIS. INDICATIVE CHEMICAL PRICES A-Z. CHEMICAL PRICES. [Online] ICIS, 2010. [Cited: May 2, 2010.]

http://www.icis.com/staticpages/a-e.htm.

31. Chemical Engineering. Economic Indicators - Chemical Engineering Plan Cost Index (CEPCI). Chemical

Engineering News. April, November 2010, p. 64.

32. Busche, R. M. Venture Analysis: A Framework for Venture Planning - Course Notes. Wilmington,

Delaware : Bio-en-gene-er Associates, 1995.

33. Sani, Robert. Example_Economics2008_15_yr_Oct_08. Boulder, Colorado, U.S.A. : University of

Colorado at Boulder, October 2008.


117

34. Bauman, H. C. Process Plant Estimating, Evaluation, and Control. Solano Beach, California :

Craftsman, 1974.

35. Sani, Robert. CHEN 4520 Chemical Process Synthesis. Class 17 Estimating Total Capital Investment

(C_TCI). Boulder, Colorado, U.S.A. : University of Colorado at Boulder, 2009.

36. Bastar IV, Richard G. Inspiration. University of Colorado at Boulder, Boulder, Colorado : 2009.

37. Emmerling, Joey. Financier. Boulder, Colorado, Decmeber 7, 2009.

38. Dunn, Dave. 2009 Corporate Tax Rates. Samarak. [Online] November 14, 2008. [Cited: December 5,

2009.] http://www.samarak.com/2009-corporate-tax-rates/.

39. Sani, Robert. Minor Design Project. CHEN 4520 Chemical Process Synthesis. Boulder, Colorado,

U.S.A. : University of Colorado at Boulder, Fall 2009.

40. Mogck, Drew. Mini Design. Boulder, Colorado, December 3, 2009.


118

Appendix A: Acronyms

Acronym Stands For

BCR Benefit-Cost Ratio

BEP Break-Even Period

BMC Bare-Module Cost

BTEX Paraffinic and Aromatic Hydrocarbons

BTU British Thermal Unit

CE Chemical Engineering Plant Cost Index

COM Cost of Manufacture

DIC Direct Installed Cost

f.o.b. Free on Board

HC Hydrocarbon

IRR Investor’s Rate of Return

IRS Internal Revenue Service

MACRS Modified Accelerated Cost Recovery

System

MMBTU Million British Thermal Units

MMSCF Million Standard Cubic Feet

MSDS Material Safety Data Sheets

NFPA National Fire Protection Association

NGL Natural Gas Liquids

NPV Net Present Value

PBP Payback Period

PFD Process Flow Diagram

ROI Return on Investment

RVP Reid Vapor Pressure

SCF Standard Cubic Feet

TBM Total Bare-Module Investment

TCI Total Capital Investment

tcm Trillion Cubic Meters

TPI Total Permanent Investment

TVP True Vapor Pressure


119

VGA Venture Guidance Appraisal

WC Working Capital


120

Appendix B: Chemical Information

LPG MSDS (18)

PEG, INC. DBA PACIFIC ENERGY GROUP A DIVISION OF -- LIQUIFIED PETROLEUM GAS, LP-GAS, LPG, PROPANE -- - ===================== Product Identification =====================

Product ID:LIQUIFIED PETROLEUM GAS, LP-GAS, LPG, PROPANE

MSDS Date:08/01/1996

FSC:NIIN:Submitter:N EN

Status Code:A

MSDS Number: CKTXF

=== Responsible Party ===

Company Name:PEG, INC. DBA PACIFIC ENERGY GROUP A DIVISION OF

Address:UNITED LIQUID GAS CO.

Box:398

City:PASO ROBLES

State:CA

ZIP:93446

Country:US

Info Phone Num:(800) 726-5747; (805) 239-2182

Emergency Phone Num:(800) 633-8253 (PERS, INC)

CAGE:TO802

=== Contractor Identification ===

Company Name:PEG INC. DBA PACIFIC ENERGY GROUP (DIV OF UNITED LIQUID

GAS)

Box:398

City:PASO ROBLES

State:CA

ZIP:93446

Country:US

Phone:800-726-5747;805-239-2182

CAGE:TO802

============= Composition/Information on Ingredients =============

Ingred Name:PROPANE

CAS:74-98-6

RTECS #:TX2275000

= Wt:92.

OSHA PEL:1000 PPM

ACGIH TLV:SIMPLE ASPHYXIANT

Ingred Name:PROPYLENE

CAS:115-07-1

RTECS #:UC6740000

= Wt:5.

Ingred Name:BUTANE

CAS:106-97-8

RTECS #:EJ4200000

= Wt:3.

===================== Hazards Identification =====================


121

Routes of Entry: Inhalation:YES Skin:YES Ingestion:YES

Reports of Carcinogenicity:NTP:NO IARC:NO OSHA:NO

Health Hazards Acute and Chronic:PRODUCT IS A SIMPLE ASPHYXIANT IN HIGH

CONCENTRATIONS. INHALATION: EXPOSURE MAY PRODUCE RAPID BREATHING,

HEADACHE, DIZZINESS, DISTURBANCES, MUSCULAR WEAKNESS, TREMORS,

NARCOSIS, UNCONCIOUSNESS AND DEA TH, DEPENDING ON DURATION AND

CONCENTRATION OF EXPOSURE. EYE CONTACT: THIS GAS IS NON-IRRITATING,

BUT DIRECT CONTACT WITH LIQUIFIED, PRESSURIZED GAS OR FROST

PARTICLES MAY PRODUCE SEVERE AND POSSIBLE PERMANENT EYE DAMAGE

FROM FREEZE BURNS. SKIN CONTACT: THIS MATERIAL IS NOT EXPECTED TO

BE ABSORBED THROUGH THE SKIN. NON-IRRITATING; BUT SOLID AND LIQUID

FORMS OF THIS MATERIAL AND PRESSURIZED GAS CAN CAUSE FREEZE

BURNS.(EFFECTS OF OVEREXP)

Explanation of Carcinogenicity:PRODUCT IS NOT LISTED AS A CARCINOGEN OR

POTENTIAL CARCINOGEN BY NTP, IARC OR OSHA.

Effects of Overexposure:HEALTH HAZARDS ACUTE AND CHRONIC (CONT):

INGESTION: SOLID AND LIQUID FORMS OF THIS MATERIAL AND THE

PRESSURIZED GAS CAN CAUSE FREEZE BURNS.

Medical Cond Aggravated by Exposure:SPECIAL HEALTH EFFECTS: PERSONNEL

WITH PRE-EXISTING CHRONIC RESPIRATORY DISEASES SHOULD AVOID

EXPOSURE TO THIS MATERIAL.

======================= First Aid Measures =======================

First Aid:INHALATION: IMMEDIATELY MOVE PERSONNEL TO FRESH AIR. FOR

RESPIRATORY DISTRESS, GIVE AIR, OXYGEN OR ADMINISTER CPR

(CARDIOPULMONARY RESUSCITATION) IF NECESSARY. OBTAIN MEDICAL

ATTENTION IF BREATHING DI FFICULTIES CONTINUE. EYE CONTACT:

VAPORSARE NOT EXPECTED TO PRESENT AN EYE IRRITATION HAZARD. IF

CONTACTED BY LIQUID/SOLID, IMMEDIATELY FLUSH EYE(S) GENTLY WITH

WARM WATER FOR AT LEAST 15 MINUTES. SEE K MEDICAL ATTENTION IF PAIN

OR REDNESS PERSISTS. INGESTION: INDUCE VOMITING WITH WARM WATER

(QT.) ONLY IF PATIENT CONSCIOUS. IMMEDIATELY OBTAIN MEDICAL

ATTENTION. SKIN CONTACT: FLUSH WITH COPIOUS AMOU NTS OF WATER.

CONTACT A PHYSICIAN

===================== Fire Fighting Measures =====================

Flash Point:=-104.4C, -156.F

ESTIMATED

Autoignition Temp:=450.C, 842.F

Lower Limits:2.1%

Upper Limits:9.5%

Extinguishing Media:DRY CHEMICAL, WATER SPRAY, FOAM, CO2.

Fire Fighting Procedures:USE NIOSH-APPROVED SCBA AND FULL PROTECTIVE

EQUIPMENT . EVACUATE AREA. SHUT OFF SOURCE OF GAS, IF POSSIBLE.

NOTIFY FIRE DEPTARTMENT. REMAIN UP-WIND OF VAPORS. ALLOW ONLY

PROPERLY PROTECTED PERSO NNEL IN AREA. READILY FORMS EXPLOSIVE WITH

AIR AND OXIDIZERS. ALLOW FIRE TO BURN UNTIL GAS FLOW IS SHUT OFF.

(EXPLO HAZ)

Unusual Fire/Explosion Hazard:FIRE FIGHT PROC (CONT): USE WATER SPRAY

TO COOL EXPOSED EQUIPMENT AND VAPOR SPACE OF CONTAINERS, CONTAINERS

MAY RUPTURE IF EXPOSED TO HEAT OR FLAME. APPROACH A

FLAME-ENVELOPED CONTAINER FROM SIDE NEV ER FROM THE HEAD ENDS. FOR

MASSIVE, UNCONTROLLABLE FIRES AND WHEN FLAME IS IMPINGING ON VAPOR

SPACE (OTHER INFORMATION)


122

================== Accidental Release Measures ==================

Spill Release Procedures:PRODUCT IS EXTREMELY FLAMMABLE. VAPOR IS

HEAVIER THAN AIR AND MAY COLLECT AT LOWER LEVELS. IF THERE IS A

LEAK BUT NO FIRE, DO NOT IGNITE THE ESCAPED GAS. ELIMINATE ALL

IGNITION SOURCES. WATER SPRAY CA N BE USED TO HELP DILUTE VAPOR

CONCENTRATION IN AIR. IF POSSIBLE, REMOVE LEAKING CONTAINER TO SAFE

AREA.

====================== Handling and Storage ======================

Handling and Storage Precautions:STORE AND USE CYLINDERS AND TANKS IN A

WELL VENTILATED AREA, AWAY FROM HEAT AND SOURCES OF IGNITION. NO

SMOKING NEAR STORAGE OR USE. FOLLOW STANDARD PROCEDURES FOR

HANDLING CYLINDERS, TANKS, LOADING/U NLOADING. FIXED STORAGE

CONTAINERS MUST BE GROUNDED AND BONDED DURING TRANSFER OF PRODUCT.

============= Exposure Controls/Personal Protection =============

Respiratory Protection:FOR EXCESSIVE GAS CONCENTRATIONS, USE ONLY NIOSH

(APPROVED - ) SELF-CONTAINED BREATHING APPARATUS.

Protective Gloves:INSULATED, IMPERVIOUS PLASTIC OR NEOPRENE COATED

CANVAS GLOVES.

Eye Protection:USE (ANSI APPROVED - ) CHEMICAL-TYPE GOGGLES AND FACE

SHIELD (SUPP SAFETY)

Other Protective Equipment:EYEWASH AND DELUGE SHOWER MEETING ANSI

DESIGN CRITERIA . USE INSULATED, IMPERVIOUS PLASTIC OR NEOPRENE

COATED CANVAS PROTECTIVE GEAR (APRON, FACE SHIELD, ETC.) TO PROTECT

HANDS & OTHER SKIN AREAS.

Work Hygienic Practices:PREVENT POTENTIAL SKIN CONTACT WITH COLD

LIQUID/SOLID/VAPORS.

Supplemental Safety and Health

EYE PROTECT (CONT): WHEN HANDLING LIQUEFIED GASES. (ANSI APPROVED - )

SAFETY GLASSES AND/OR FACE SHIELD ARE RECOMMENDED WHEN HANDLING

HIGH PRESSURE CYLINDERS AND PIPING SYSTEM AND WHENEVER VAPORS ARE

DISCHARGED.

================== Physical/Chemical Properties ==================

Boiling Pt:=-42.2C, -44.F

Vapor Pres:208 PSIG @ 100 F.

Vapor Density:1.55 AIR=1

Spec Gravity:.508 (H2O=1)

Evaporation Rate & Reference:NA GAS @STD TEMP & PRESS.

Solubility in Water:SLIGHT.

Appearance and Odor:COLORLESS LIQUEFIED PETROLEUM GAS. ODOR: (OTHER

INFO)

Percent Volatiles by Volume:100%

================= Stability and Reactivity Data =================

Stability Indicator/Materials to Avoid:YES

STRONG OXIDIZING AGENTS.

Hazardous Decomposition Products:COMBUSTION MAY PRODUCE CARBON MONOXIDE

AND OTHER HARMFUL SUBSTANCES.


123

=================== Toxicological Information ===================

Toxicological Information:OCCUPATIONAL EXPOSURE LIMITS: ACGIH LISTS AS

A SIMPLE ASPHYIXIANT. OSHA TWA 1000 PPM. STATE REGS (CONT):

REPRODUCTIVE HARM." THE BENZENE AND TOULENE ARE DESTROYED WHEN

PROPANE IS BURNED. RADON DOES NOT BURN BUT IS RELEASED WITH THE

COMBUSTION BY-PRODUCTS. RADON AND PROPANE COMBUSTION BY-PRODUCTS

CAN GENERALLY BE REMOVED THROUGH APPLIANCE VENTS AND OTHER EXHAUST

SYSTEMS. WHEN PROPANE IS PROCESS ED IN SOME DEHYDRATORS, BENZENE

AND TOULENE CAN BE RELEASED TO THE ENVIRONMENT. A WARNING ODORANT

IS ADDED TO PROPANE SO LEAKS OF UNBURNED GAS CAN BE QUICKLY

DETECTED. IF GAS ODOR IS DETECTED, YOUR SU PPLIER SHOULD BE

CONTACTED PROMPTLY.

===================== Ecological Information =====================

Ecological:ENVIRONMENTAL EFFECTS: AVOID UNCONTROLLED RELEASES OF THIS

MATERIAL. LIQUID RELEASE WILL HAVE POSSIBLE EFFECT ON PLANT AND

ANIMAL LIFE. LARGE LIQUID RELEASE WILL QUICKLY VAPORIZE TO PRODUCE

A LARGE V APOR CLOUD. VAPOR CLOUD IS BOTH A FIRE AND ASPHYXIATION

HAZARD.

==================== Disposal Considerations ====================

Waste Disposal Methods:DISPOSAL OF GAS IN ACCORDANCE WITH APPLICABLE

LAWS AND REGULATIONS. VENT VAPOR IN SAFE LOCATION AND INSURE THAT

GAS DISSIPATES BELOW THE LOWER FLAMMABLE LIMIT. CONTROLLED BURNING

IS PREFERRED.

=================== MSDS Transport Information ===================

Transport Information:D.O.T. HAZARD CLASS: FLAMMABLE GAS. D.O.T. ID NO

(UN/NA): UN 1075 LIQUEFIED PETROLEUM GAS (LPG). D.O.T. SHIPPING

NAME: PROPANE OR LIQUEFIED PETROLEUM GAS. IMO SHIPPING NAME:

PROPANE BUTANE. IMO HA ZARD CLASS: 2.1. IMO LABEL: FLAMMABLE GAS.

===================== Regulatory Information =====================

State Regulatory Information:PROPOSITION 65 [PUBLIC WARNING] THE SAFE

DRINKING WATER AND TOXIC ENFORCEMENT ACT 1986, COMMONLY REFERRED TO

AS PROPOSITION 65, REQUIRES THEN GOVERNOR TO PUBLISH A LIST OF

CHEMICALS "KNOWN TO THE STAT E TO CAUSE CANCER, BIRTH DEFECTS OR

REPRODUCTIVE HARM." IT ALSO REQUIRES CALIFORNIA BUSINESSES TO WARN

THE PUBLIC QUARTERLY OF POTENTIAL EXPOSURE TO THESE CHEMICALS WHICH

RESULT FROM THEIR OPERATIONS. LIQUIFIED PETROLEUM GAS (PROPANE),

IN ITS ORIGINAL STATE, CONTAINS RADON AND BENZENE, CHEMICALS "KN

OWN TO STATE OF CALIFORNIA TO CAUSE CANCER." ALSO CONTAINS TOULENE,

A CHEMICAL "KNOWN TO STATE OF CALIFORNIA TO CAUSE (TOXICOLOGICAL)

======================= Other Information =======================

Disclaimer (provided with this information by the compiling agencies):

This information is formulated for use by elements of the Department

of Defense. The United States of America in no manner whatsoever,

expressly or implied, warrants this information to be accurate and

disclaims all liability for its use. Any person utilizing this

document should seek competent professional advice to verify and


124

assume responsibility for the suitability of this information to their

particular situation.

Natural Gas MSDS (19) MATERIAL SAFETY DATA SHEET

EQUILON MSDS: 55277E-04 11/16/98

CONDENSATE (NATURAL GAS) - FLAMMABLE

TELEPHONE NUMBER:

24 HOUR EMERGENCY ASSISTANCE GENERAL MSDS ASSISTANCE

EQUIVA SERVICES: 877-276-7283 877-276-7285

CHEMTREC: 800-424-9300

NAME AND ADDRESS

EQUILON ENTERPRISES LLC

PRODUCT STEWARDSHIP

P.O. BOX 674414

HOUSTON, TX 77267-4414

_____________________________________________________________________________

__

SECTION I NAME

_____________________________________________________________________________

__

PRODUCT: CONDENSATE (NATURAL GAS) - FLAMMABLE

CHEM NAME: NATURAL GAS CONDENSATE

CHEM FAMILY: PETROLEUM HYDROCARBON

SHELL CODE: 87879 82966 82977 80491 87594 87596 87597 87598

87599 87605 87606 87755 89433 89726 89870

HEALTH HAZARD: 2 FIRE HAZARD: 3 REACTIVITY: 0

_____________________________________________________________________________

__

SECTION II-A PRODUCT/INGREDIENT

_____________________________________________________________________________

__

NO. COMPOSITION CAS NO. PERCENT

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

P CONDENSATE (NATURAL GAS) - FLAMMABLE

1 CONDENSATE* 64741-47-5 100

A NATURAL GAS 8006-14-2 VARIABLE

B BENZENE 71-43-2 VARIABLE

C N-HEXANE 110-54-3 VARIABLE

THIS IS 1 OF 8 MSDS'S BASED ON FLASHPOINT AND SULFUR CONTENT. LABEL CODE IS

0008695.

*THIS CHEMICAL IS A COMPLEX SUBSTANCE WHICH MAY CONTAIN CONSTITUENTS

IDENTIFIED

AS A, B, C, ABOVE THAT ARE NOT INTENTIONALLY ADDED TO THE PRODUCT.

_____________________________________________________________________________

__

SECTION II-B ACUTE TOXICITY DATA

_____________________________________________________________________________

__

NO. ACUTE ORAL LD50 ACUTE DERMAL LD50 ACUTE INHALATION LC50

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

P NOT AVAILABLE

_____________________________________________________________________________

__

SECTION III HEALTH INFORMATION


125

_____________________________________________________________________________

__

THE HEALTH EFFECTS NOTED BELOW ARE CONSISTENT WITH REQUIREMENTS UNDER THE

OSHA HAZARD COMMUNICATION STANDARD (29 CFR 1910.1200).

EYE CONTACT: BASED ON SIMILAR PRODUCT TESTING PRODUCT IS MINIMALLY

IRRITATING TO THE EYES.

SKIN CONTACT: PROLONGED AND REPEATED LIQUID CONTACT CAN CAUSE DEFATTING AND

DRYING OF THE SKIN RESULTING IN SKIN IRRITATION AND DERMATITIS.

INHALATION: WARNING. NATURAL GAS, AND OTHER HAZARDOUS VAPORS MAY EVOLVE

AND COLLECT IN THE HEADSPACE OF STORAGE TANKS OR OTHER ENCLOSED

VESSELS. NATURAL GAS IS EXTREMELY FLAMMABLE AND A SIMPLE

ASPHYXIANT. INHALATION OF OTHER LIGHT HYDROCARBONS MAY CAUSE

PULMONARY IRRITATION AND RESULT IN CNS DEPRESSION. PROLONGED

AND REPEATED INHALATION OF N-HEXANE MAY PRODUCE PERIPHERAL

NEUROPATHY. PRODUCT MAY BE IRRITATING TO THE NOSE, THROAT AND

RESPIRATORY TRACT. PROLONGED AND REPEATED EXPOSURE TO BENZENE

MAY CAUSE SERIOUS INJURY TO BLOOD FORMING ORGANS AND IS LINKED

TO LATER DEVELOPMENT OF ACUTE MYELOGENOUS LEUKEMIA.

INGESTION: THIS PRODUCT MAY BE HARMFUL OR FATAL IF SWALLOWED. INGESTION

OF PRODUCT MAY RESULT IN VOMITING; ASPIRATION (BREATHING) OF

VOMITUS INTO THE LUNGS MUST BE AVOIDED AS EVEN SMALL QUANTITIES MAY

RESULT IN ASPIRATION PNEUMONITIS.

SIGNS AND SYMPTOMS: IRRITATION AS NOTED ABOVE. EARLY TO MODERATE CNS

(CENTRAL NERVOUS SYSTEM) DEPRESSION MAY BE EVIDENCED BY

GIDDINESS, HEADACHE, DIZZINESS AND NAUSEA; IN EXTREME CASES,

UNCONCIOUSNESS AND DEATH MAY OCCUR. ASPHYXIATION AND

H2S TOXICITY MAY BE NOTED BY A SUDDEN LOSS OF CONSCIOUSNESS;

DEATH MAY QUICKLY FOLLOW. ASPIRATION PNEUMONITIS MAY BE

EVIDENCED BY COUGHING, LABORED BREATHING AND CYANOSIS

(BLUISH SKIN); IN SEVERE CASES DEATH MAY OCCUR.

PERIPHERAL NERVE DAMAGE MAY BE EVIDENCED BY MUSCULAR

WEAKNESS AND LOSS OF SENSATION IN THE ARMS AND LEGS.

DAMAGE TO BLOOD FORMING ORGANS MAY BE EVIDENCED BY EASY

FATIGABILITY AND PALLOR (RBC EFFECT), DECREASED

RESISTANCE TO INFECTION (WBC EFFECT) AND EXCESSIVE BRUISING AND

BLEEDING (PLATELET EFFECT).

AGGRAVATED MEDICAL CONDITIONS:

PREEXISTING EYE, SKIN, AND RESPIRATORY DISORDERS OR PREEXISTING IMPAIRED

BLOOD FORMING FUNCTIONS MAY BE AGGRAVATED BY EXPOSURE TO THIS PRODUCT.

OTHER HEALTH EFFECTS:

BENZENE IS LISTED BY THE NATIONAL TOXICOLOGY PROGRAM, THE INTERNATIONAL

AGENCY FOR RESEARCH ON CANCER, AND OSHA AS A CHEMICAL CAUSALLY ASSOCIATED

WITH CANCER (ACUTE MYELOGENOUS LEUKEMIA) IN HUMANS.

SEE SECTION VI FOR ADDITIONAL HEALTH INFORMATION.

_____________________________________________________________________________

__

SECTION IV OCCUPATIONAL EXPOSURE LIMITS

_____________________________________________________________________________

__

COMP OSHA ACGIH

NO. PEL/TWA PEL/CEILING TLV/TWA TLV/STEL OTHER

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

P* 300 PPM 300 PPM 500 PPM 500 PPM**

B 1 PPM 10 PPM*** 5 PPM**

C 50 PPM 50 PPM

*GASOLINE **OSHA PEL/STEL ***CLASSIFIED BY ACGIH AS A "SUSPECTED HUMAN

CARCINOGEN" (A2)


126

_____________________________________________________________________________

__

SECTION V EMERGENCY AND FIRST AID PROCEDURES

_____________________________________________________________________________

__

EYE CONTACT: FLUSH WITH WATER FOR 15 MINUTES WHILE HOLDING EYELIDS OPEN.

GET MEDICAL ATTENTION.

SKIN CONTACT: FLUSH WITH WATER WHILE REMOVING CONTAMINATED CLOTHING AND

SHOES. FOLLOW BY WASHING WITH SOAP AND WATER. DO NOT REUSE CLOTHING

OR SHOES UNTIL CLEANED. IF IRRITATION PERSISTS, GET MEDICAL

ATTENTION.

INHALATION: REMOVE VICTIM TO FRESH AIR AND PROVIDE OXYGEN IF BREATHING IS

DIFFICULT. GIVE ARTIFICIAL RESPIRATION IF NOT BREATHING. GET

MEDICAL ATTENTION.

INGESTION: DO NOT INDUCE VOMITING. IF VOMITING OCCURS SPONTANEOUSLY KEEP

HEAD BELOW HIPS TO PREVENT ASPIRATION OF LIQUID INTO THE LUNGS.

GET MEDICAL ATTENTION.*

NOTE TO PHYSICIAN: *IF MORE THAN 2.0 ML PER KG HAS BEEN INGESTED AND

VOMITING HAS NOT OCCURRED, EMESIS SHOULD BE INDUCED WITH MEDICAL

SUPERVISION. KEEP VICTIM'S HEAD BELOW HIPS TO PREVENT

ASPIRATION. IF SYMPTOMS SUCH AS LOSS OF GAG REFLEX,

CONVULSIONS OR UNCONSCIOUSNESS OCCUR BEFORE EMESIS,

GASTRIC LAVAGE USING A CUFFED ENDOTRACHEAL TUBE SHOULD BE

CONSIDERED.

_____________________________________________________________________________

__

SECTION VI SUPPLEMENTAL HEALTH INFORMATION

_____________________________________________________________________________

__

WHILE THERE IS NO EVIDENCE THAT EXPOSURE TO INDUSTRIALLY ACCEPTABLE LEVELS OF

HYDROCARBON HAVE PRODUCED CARDIAC EFFECTS IN HUMANS, ANIMAL STUDIES HAVE

SHOWN THAT INHALATION OF HIGH LEVELS OF NATURAL GAS VAPORS PRODUCED CARDIAC

SENSITIZATION. SUCH SENSITIZATION MAY CAUSE FATAL CHANGES IN HEART RHYTHMS.

THIS LATTER EFFECT WAS SHOWN TO BE ENHANCED BY HYPOXIA OR THE INJECTION OF

ADRENALIN-LIKE AGENTS. ANIMAL STUDIES ON BENZENE HAVE DEMONSTRATED

IMMUNOTOXICITY, TESTICULAR EFFECTS AND ALTERATIONS IN REPRODUCTIVE CYCLES,

EVIDENCE OF CHROMOSOMAL DAMAGER OR OTHER CHROMOSOMAL CHANGES, AND

EMBRYO/FETOTOXICITY BUT NOT TERATOGENICITY. STUDIES ON N-HEXANE IN LABORATORY

ANIMALS HAVE SHOWN MILD, TRANSITORY EFFECTS ON THE SPLEEN AND BLOOD (WHITE

BLOOD CELLS), AND EVIDENCE OF LUNG DAMAGE. IN ADDITION, FETOTOXICITY HAS

BEEN DEMONSTRATED AT LEVELS PRODUCING MATERNAL TOXICITY. AT HIGH LEVELS,

INHALATION EXPOSURE HAS RESULTED IN TESTICULAR AND EPIDIDYMAL ATROPHY.

_____________________________________________________________________________

__

SECTION VII PHYSICAL DATA

_____________________________________________________________________________

__

BOILING POINT (DEG F): SPECFIC GRAVITY (H2O = 1): VAPOR PRESSURE (MM HG):

-4 TO 356 APPROX. >0.7 7-14.5 PSI

(REID)

MELTING POINT (DEG F): SOLUBILITY IN WATER: VAPOR DENSITY (AIR =

1):

NOT AVAILABLE NEGLIGIBLE >1

% VOLATILE BY

VOL=

100 (@ 415 DEG.

F)


127

EVAPORATION RATE (NORMAL BUTYL ACETATE = 1):NOT AVAILABLE

APPEARANCE AND ODOR:AMBER TO DARK COLORED LIQUID. HYDROCARBON ODOR.

PHYS/CHEM PROPERTIES: SEE ABOVE FOR DETAILS

_____________________________________________________________________________

__

SECTION VIII FIRE AND EXPLOSION HAZARDS

_____________________________________________________________________________

__

FLASH POINT AND METHOD: <100 DEG F (PMCC)

FLAMMABLE LIMITS/PERCENT VOLUME IN AIR: LOWER: N/AV HIGHER: N/AV

EXTINGUISHING MEDIA: USE WATER FOG, FOAM, DRY CHEMICAL OR CO2. DO NOT USE A

DIRECT STREAM OF WATER. PRODUCT WILL FLOAT AND CAN BE REIGNITED ON SURFACE

OF WATER. SPECIAL FIRE FIGHTING PROCEDURES AND PRECAUTIONS: WARNING.

FLAMMABLE. CLEAR FIRE AREA OF UNPROTECTED PERSONNEL. DO NOT ENTER

CONFINED FIRE SPACE WITHOUT FULL BUNKER GEAR (HELMET WITH FACE SHIELD, BUNKER

COATS, GLOVES AND RUBBER BOOTS), INCLUDING A POSITIVE PRESSURE NIOSH APPROVED

SELF-CONTAINED BREATHING APPARATUS. COOL FIRE EXPOSED CONTAINERS WITH WATER.

UNUSUAL FIRE AND EXPLOSION HAZARDS: VAPORS ARE HEAVIER THAN AIR ACCUMULATING

IN LOW AREAS AND TRAVELING ALONG THE GROUND AWAY FROM THE HANDLING SITE. DO

NOT WELD, HEAT OR DRILL ON OR NEAR CONTAINER. HOWEVER, IF EMERGENCY

SITUATIONS REQUIRE DRILLING, ONLY TRAINED EMERGENCY PERSONNEL SHOULD DRILL.

_____________________________________________________________________________

__

SECTION IX REACTIVITY

_____________________________________________________________________________

__

STABLITY: STABLE HAZARDOUS POLYMERIZATION WILL NOT OCCUR

CONDITIONS AND MATERIALS TO AVOID: AVOID HEAT, SPARKS, OPEN FLAMES AND STRONG

OXIDIZING AGENTS. PREVENT VAPOR ACCUMULATION. HAZARDOUS DECOMPOSITION

PRODUCTS: THERMAL DECOMPOSITION PRODUCTS ARE HIGHLY DEPENDENT ON THE

COMBUSTION CONDITIONS. A COMPLEX MIXTURE OF AIRBORNE SOLID, LIQUID,

PARTICULATES AND GASES WILL EVOLVE WHEN THIS MATERIAL UNDERGOES PYROLYSIS OR

COMBUSTION. CARBON MONOXIDE AND OTHER UNIDENTIFIED ORGANIC COMPOUNDS MAY BE

FORMED UPON COMBUSTION.

_____________________________________________________________________________

__

SECTION X EMPLOYEE PROTECTION

_____________________________________________________________________________

__

RESPIRATORY PROTECTION:

AVOID BREATHING VAPOR. IF EXPOSURE MAY OR DOES EXCEED OCCUPATIONAL

EXPOSURE LIMITS (SEC. IV) USE A NIOSH-APPROVED RESPIRATOR TO PREVENT

OVEREXPOSURE. IN ACCORD WITH 29 CFR 1910.134 AND 1910.1028 USE EITHER AN

ATMOSPHERE-SUPPLYING RESPIRATOR OR AN AIR-PURIFYING RESPIRATOR FOR ORGANIC

VAPORS. PROTECTIVE CLOTHING AVOID CONTACT WITH EYES. WEAR CHEMICAL GOGGLES

IF THERE IS LIKELIHOOD OF CONTACT WITH EYES. AVOID CONTACT WITH SKIN AND

CLOTHING. WEAR CHEMICAL-RESISTANT GLOVES AND PROTECTIVE CLOTHING.

ADDITIONAL PROTECTIVE MEASURES: USE EXPLOSION-PROOF VENTILATION AS REQUIRED

TO CONTROL VAPOR CONCENTRATIONS.

_____________________________________________________________________________

__

SECTION XI ENVIRONMENTAL PROTECTION

_____________________________________________________________________________

__

SPILL OR LEAK PROCEDURES:

WARNING. FLAMMABLE. ELIMINATE ALL IGNITION SOURCES. HANDLING EQUIPMENT

MUST BE GROUNDED TO PREVENT SPARKING. *** LARGE SPILLS *** EVACUATE THE


128

HAZARD AREA OF UNPROTECTED PERSONNEL. WEAR APPROPRIATE RESPIRATOR AND

PROTECTIVE CLOTHING. SHUT OFF SOURCE OF LEAK ONLY IF SAFE TO DO SO. DIKE

AND CONTAIN. IF VAPOR CLOUD FORMS, WATER FOG MAY BE USED TO SUPPRESS;

CONTAIN RUN-OFF. REMOVE WITH VACUUM TRUCKS OR PUMP TO STORAGE/SALVAGE

VESSELS. SOAK UP RESIDUE WITH AN ABSORBENT SUCH AS CLAY, SAND OR OTHER

SUITABLE MATERIAL; PLACE IN NON-LEAKING CONTAINERS FOR PROPER DISPOSAL.

FLUSH AREA WITH WATER TO REMOVE TRACE RESIDUE; DISPOSE OF FLUSH SOLUTIONS

AS ABOVE. *** SMALL SPILLS *** TAKE UP WITH AN ABSORBENT MATERIAL AND PLACE

IN NON-LEAKING CONTAINERS; SEAL TIGHTLY FOR PROPER DISPOSAL.

_____________________________________________________________________________

__

SECTION XII SPECIAL PRECAUTIONS

_____________________________________________________________________________

__

KEEP LIQUID AND VAPOR AWAY FROM HEAT, SPARKS AND FLAME. SURFACES THAT ARE

SUFFICIENTLY HOT MAY IGNITE EVEN LIQUID PRODUCT IN THE ABSENCE OF SPARKS OR

FLAME. EXTINGUISH PILOT LIGHTS, CIGARETTES AND TURN OFF OTHER SOURCES OF

IGNITION PRIOR TO USE AND UNTIL ALL VAPORS ARE GONE. VAPORS MAY ACCUMULATE

AND TRAVEL TO IGNITION SOURCES DISTANT FROM THE HANDLING SITE; FLASH-FIRE CAN

RESULT. KEEP CONTAINERS CLOSED WHEN NOT IN USE. USE WITH ADEQUATE

VENTILATION. CONTAINERS, EVEN THOSE THAT HAVE BEEN EMPTIED, CAN CONTAIN

EXPLOSIVE VAPORS. DO NOT CUT, DRILL, GRIND, WELD OR PERFORM SIMILAR

OPERATIONS ON OR NEAR CONTAINERS. STATIC ELECTRICITY MAY ACCUMULATE AND

CREATE A FIRE HAZARD. GROUND FIXED EQUIPMENT. BOND AND GROUND TRANSFER

CONTAINERS AND EQUIPMENT. WASH WITH SOAP AND WATER BEFORE EATING, DRINKING,

SMOKING, APPLYING COSMETICS, OR USING TOILET FACILITIES. LAUNDER

CONTAMINATED CLOTHING BEFORE REUSE.

_____________________________________________________________________________

__

SECTION XIII TRANSPORTATION REQUIREMENTS

_____________________________________________________________________________

__

DEPARTMENT OF TRANSPORTATION CLASSIFICATION:

CLASS 3 (FLAMMABLE LIQUID), PACKING GROUP MUST BE DETERMINED ON A

CASE-BY-CASE BASIS.

DOT PROPER SHIPPING NAME:FLAMMABLE LIQUID, N.O.S. (PETROLEUM CONDENSATE)

OTHER REQUIREMENTS:UN1993, GUIDE 128

_____________________________________________________________________________

__

SECTION XIV OTHER REGULATORY CONTROLS

_____________________________________________________________________________

__

THIS PRODUCT IS LISTED ON THE EPA/TSCA INVENTORY OF CHEMICAL SUBSTANCES.

IN ACCORDANCE WITH SARA TITLE III, SECTION 313, THE ENVIRONMENTAL DATA SHEET

(EDS) SHOULD ALWAYS BE COPIED AND SENT WITH THE MSDS.

_____________________________________________________________________________

__

SECTION XV STATE REGULATORY INFORMATION

_____________________________________________________________________________

__

THE FOLLOWING CHEMICALS ARE SPECIFICALLY LISTED BY INDIVIDUAL STATES; OTHER

PRODUCT SPECIFIC HEALTH AND SAFETY DATA IN OTHER SECTIONS OF THE MSDS MAY

ALSO BE APPLICABLE FOR STATE REQUIREMENTS. FOR DETAILS ON YOUR REGULATORY

REQUIREMENTS YOU SHOULD CONTACT THE APPROPRIATE AGENCY IN YOUR STATE.

STATE LISTED COMPONENT CAS NO PERCENT STATE CODE

_____________________________________________________________________________

__


129

NATURAL GAS 8006-14-2 VARIABLE MA, PA

BENZENE 71-43-2 VARIABLE CA, CT, FL, IL,

LA, MA, ME, MN,

NJ, PA, RI,

CA65C/R

N-HEXANE 110-54-3 VARIABLE CA, CT, FL, IL,

LA, MA, ME, MN,

PA, RI

CA = CALIFORNIA HAZ. SUBST. LIST; CA65C, CA65R, CA65C/R = CALIFORNIA SAFE

DRINKING WATER AND TOXICS ENFORCEMENT ACT OF 1986 OR PROPOSITION 65 LIST; CT

=

CONNECTICUT TOXIC. SUBST. LIST; FL = FLORIDA SUBST. LIST; IL = ILLINOIS TOX.

SUBST. LIST; LA = LOUISIANA HAZ. SUBST. LIST; MA = MASSACHUSETTS SUBST.

LIST; ME = MAINE HAZ. SUBST. LIST; MN = MINNESOTA HAZ. SUBST. LIST; NJ =

NEW JERSEY HAZ. SUBST. LIST; PA = PENNSYLVANIA HAZ. SUBST. LIST; RI = RHODE

ISLAND HAZ. SUBST. LIST.

CALIFORNIA PROPOSITION 65 FOOTNOTE: CA65C = THE CHEMICAL IDENTIFIED WITH THIS

CODE IS KNOWN TO THE STATE OF CALIFORNIA TO CAUSE CANCER. CA65R = THE

CHEMICAL IDENTIFIED WITH THIS CODE IS KNOWN TO THE STATE OF CALIFORNIA TO

CAUSE BIRTH DEFECTS OR OTHER REPRODUCTIVE HARM. CA65C/R = THE CHEMICAL

IDENTIFIED WITH THIS CODE IS KNOWN TO THE STATE OF CALIFORNIA TO CAUSE BOTH

CANCER AND BIRTH DEFECTS OR OTHER REPRODUCTIVE HARM.

_____________________________________________________________________________

__

SECTION XVI SPECIAL NOTES

_____________________________________________________________________________

__

MSDS REVISED IN SECTION XV - STATE REGULATORY INFORMATION.

_____________________________________________________________________________

__

THE INFORMATION CONTAINED IN THIS DATA SHEET IS BASED ON THE DATA AVAILABLE

TO US AT THIS TIME, AND IS BELIEVED TO BE ACCURATE BASED UPON THAT DATA. IT

IS PROVIDED INDEPENDENTLY OF ANY SALE OF THE PRODUCT, FOR PURPOSE OF HAZARD

COMMUNICATION. IT IS NOT INTENDED TO CONSTITUTE PRODUCT PERFORMANCE

INFORMATION, AND NO EXPRESS OR IMPLIED WARRANTY OF ANY KIND IS MADE WITH

RESPECT TO THE PRODUCT, UNDERLYING DATA OR THE INFORMATION CONTAINED

HEREIN. YOU ARE URGED TO OBTAIN DATA SHEETS FOR ALL PRODUCTS YOU BUY,

PROCESS, USE OR DISTRIBUTE, AND ARE ENCOURAGED TO ADVISE THOSE WHO MAY

COME IN CONTACT WITH SUCH PRODUCTS OF THE INFORMATION CONTAINED HEREIN.

TO DETERMINE THE APPLICABILITY OR EFFECT OF ANY LAW OR REGULATION WITH

RESPECT TO THE PRODUCT, YOU SHOULD CONSULT WITH YOUR LEGAL ADVISOR OR THE

APPROPRIATE GOVERNMENT AGENCY. WE WILL NOT PROVIDE ADVICE ON SUCH

MATTERS, OR BE RESPONSIBLE FOR ANY INJURY FROM THE USE OF THE PRODUCT

DESCRIBED HEREIN. THE UNDERLYING DATA, AND THE INFORMATION PROVIDED HEREIN

AS A RESULT OF THAT DATA, IS THE PROPERTY OF EQUIVA SERVICES, LLC AND IS NOT

TO BE THE SUBJECT OF SALE OR EXCHANGE WITHOUT THE EXPRESS WRITTEN CONSENT OF

EQUIVA SERVICES, LLC.

_____________________________________________________________________________

__

ENVIRONMENTAL DATA SHEET

EQUILON EDS: 55277E

CONDENSATE (NATURAL GAS) - FLAMMABLE

TELEPHONE NUMBER:

24 HOUR EMERGENCY ASSISTANCE GENERAL MSDS ASSISTANCE

EQUIVA SERVICES: 877-276-7283 877-276-7285

CHEMTREC: 800-424-9300


130

NAME AND ADDRESS

EQUILON ENTERPRISES

PRODUCT STEWARDSHIP

P.O. BOX 674414

HOUSTON, TX 77267-4414

PRODUCT CODE: 89870

Natural Gasoline MSDS

The hazards of natural gasoline are similar to those of LPG.

Propane MSDS

The hazards of propane are similar to those of natural gas.

TEG MSDS (20)

MSDS Number: T5382 * * * * * Effective Date: 11/09/06 * * * * * Supercedes: 02/12/04

TRIETHYLENE GLYCOL

1. Product Identification

Synonyms: Ethanol, 2,2'-[1,2-ethanediylbis(oxy)]bis-; triglycol; ethylene glycol dihydroxy-diethyl ether

CAS No.: 112-27-6

Molecular Weight: 150.20

Chemical Formula: C6H14O4

Product Codes:

J.T. Baker: W660

Mallinckrodt: 2735

2. Composition/Information on Ingredients

Ingredient CAS No Percent

Hazardous


131

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

------

Triethylene Glycol 112-27-6 90 - 100%

Yes

3. Hazards Identification

Emergency Overview

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

WARNING! CAUSES EYE IRRITATION. MAY CAUSE SKIN IRRITATION.

SAF-T-DATA(tm) Ratings (Provided here for your convenience)

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

Health Rating: 0 - None

Flammability Rating: 1 - Slight

Reactivity Rating: 0 - None

Contact Rating: 2 - Moderate

Lab Protective Equip: GOGGLES; LAB COAT; PROPER GLOVES

Storage Color Code: Green (General Storage)

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

Potential Health Effects

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

Inhalation:

No adverse health effects expected from inhalation.

Ingestion:

No adverse effects expected.

Skin Contact:

Prolonged exposure may cause skin irritation.

Eye Contact:

Splashing in eye causes irritation with transitory disturbances of corneal epithelium. However, these

effects diminish and no permanent injury is expected. Vapors are non-irritating.

Chronic Exposure:

Possible skin irritation.

Aggravation of Pre-existing Conditions:

No information found.

4. First Aid Measures


132

Inhalation:

Remove to fresh air. Not expected to require first aid measures.

Ingestion:

If large amounts were swallowed, give water to drink and get medical advice.

Skin Contact:

In case of contact, immediately flush skin with plenty of water for at least 15 minutes. Remove

contaminated clothing and shoes. Wash clothing before reuse. Call a physician if irritation develops..

Eye Contact:

If splash occurs, immediately flush eyes with plenty of water for at least 15 minutes, lifting upper and

lower eyelids occasionally. Call a physician.

5. Fire Fighting Measures

Fire:

Flash point: 177C (351F) CC

Autoignition temperature: 371C (700F)

Flammable limits in air % by volume:

lel: 0.9; uel: 9.2

Slight fire hazard when exposed to heat or flame.

Explosion:

Above the flash point, explosive vapor-air mixtures may be formed.

Fire Extinguishing Media:

Water spray, dry chemical, alcohol foam, or carbon dioxide. Water or foam may cause frothing.

Special Information:

In the event of a fire, wear full protective clothing and NIOSH-approved self-contained breathing

apparatus with full facepiece operated in the pressure demand or other positive pressure mode.

6. Accidental Release Measures

Ventilate area of leak or spill. Wear appropriate personal protective equipment as specified in Section 8.

Isolate hazard area. Keep unnecessary and unprotected personnel from entering. Contain and recover

liquid when possible. Collect liquid in an appropriate container or absorb with an inert material (e. g.,

vermiculite, dry sand, earth), and place in a chemical waste container. Do not use combustible materials,

such as saw dust. Do not flush to sewer!

7. Handling and Storage


133

Keep in a tightly closed container, stored in a cool, dry, ventilated area. Protect against physical damage.

Isolate from heat, ignition sources and oxidizing agents. Protect from freezing. Containers of this

material may be hazardous when empty since they retain product residues (vapors, liquid); observe all

warnings and precautions listed for the product.

8. Exposure Controls/Personal Protection

Airborne Exposure Limits:

None established.

Ventilation System:

Not expected to require any special ventilation.

Personal Respirators (NIOSH Approved):

Not expected to require personal respirator usage.

Skin Protection:

Wear protective gloves and clean body-covering clothing.

Eye Protection:

Use chemical safety goggles. Maintain eye wash fountain and quick-drench facilities in work area.

9. Physical and Chemical Properties

Appearance:

Clear, colorless liquid.

Odor:

Odorless.

Solubility:

Miscible in water.

Specific Gravity:

1.1274 @ 15C/4C

pH:

No information found.

% Volatiles by volume @ 21C (70F):

100

Boiling Point:

285C (545F)

Melting Point:

-5C (23F)

Vapor Density (Air=1):

5.17

Vapor Pressure (mm Hg):


134

< 0.01 @ 20C (68F)

Evaporation Rate (BuAc=1):

0.01

10. Stability and Reactivity

Stability:

Stable under ordinary conditions of use and storage. Hygroscopic.

Hazardous Decomposition Products:

Carbon dioxide and carbon monoxide may form when heated to decomposition.

Hazardous Polymerization:

Will not occur.

Incompatibilities:

Strong oxidizers.

Conditions to Avoid:

Heat, flames, ignition sources and incompatibles.

11. Toxicological Information

Oral rat LD50: 17 gm/kg; investigated as a reproductive effector.

--------\Cancer Lists\-----------------------------------------------------

-

---NTP Carcinogen---

Ingredient Known Anticipated IARC

Category

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

-

Triethylene Glycol (112-27-6) No No None

12. Ecological Information

Environmental Fate:

When released into the soil, this material is expected to readily biodegrade. When released into the soil,

this material is expected to leach into groundwater. When released into the soil, this material is not

expected to evaporate significantly. When released into water, this material is expected to readily

biodegrade. When released into water, this material is not expected to evaporate significantly. This

material has a log octanol-water partition coefficient of less than 3.0. This material is not expected to

significantly bioaccumulate. When released into the air, this material is expected to be readily degraded

by reaction with photochemically produced hydroxyl radicals. When released into the air, this material is


135

expected to have a half-life of less than 1 day.

Environmental Toxicity:

This material is expected to be slightly toxic to aquatic life. The LC50/96-hour values for fish are between

10 and 100 mg/l.

13. Disposal Considerations

Whatever cannot be saved for recovery or recycling should be managed in an appropriate and approved

waste disposal facility. Processing, use or contamination of this product may change the waste

management options. State and local disposal regulations may differ from federal disposal regulations.

Dispose of container and unused contents in accordance with federal, state and local requirements.

14. Transport Information

Not regulated.

15. Regulatory Information

--------\Chemical Inventory Status - Part 1\-------------------------------

--

Ingredient TSCA EC Japan

Australia

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

-

Triethylene Glycol (112-27-6) Yes Yes Yes Yes

--------\Chemical Inventory Status - Part 2\-------------------------------

--

--Canada--

Ingredient Korea DSL NDSL Phil.

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

Triethylene Glycol (112-27-6) Yes Yes No Yes

--------\Federal, State & International Regulations - Part 1\--------------

--

-SARA 302- ------SARA 313----

--

Ingredient RQ TPQ List Chemical

Catg.

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

--

Triethylene Glycol (112-27-6) No No No No

--------\Federal, State & International Regulations - Part 2\--------------

--


136

-RCRA- -TSCA-

Ingredient CERCLA 261.33 8(d)

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

Triethylene Glycol (112-27-6) No No No

Chemical Weapons Convention: No TSCA 12(b): No CDTA: No

SARA 311/312: Acute: Yes Chronic: No Fire: No Pressure: No

Reactivity: No (Pure / Liquid)

Australian Hazchem Code: None allocated.

Poison Schedule: None allocated.

WHMIS:

This MSDS has been prepared according to the hazard criteria of the Controlled Products Regulations

(CPR) and the MSDS contains all of the information required by the CPR.

16. Other Information

NFPA Ratings: Health: 1 Flammability: 1 Reactivity: 0

Label Hazard Warning:

WARNING! CAUSES EYE IRRITATION. MAY CAUSE SKIN IRRITATION.

Label Precautions:

Avoid contact with eyes, skin and clothing.

Wash thoroughly after handling.

Label First Aid:

In case of contact, immediately flush eyes or skin with plenty of water for at least 15 minutes. Call a

physician.

Product Use:

Laboratory Reagent.

Revision Information:

MSDS Section(s) changed since last revision of document include: 3.

Disclaimer:

*************************************************************************************

***********

Mallinckrodt Baker, Inc. provides the information contained herein in good faith but makes no

representation as to its comprehensiveness or accuracy. This document is intended only as a guide to

the appropriate precautionary handling of the material by a properly trained person using this

product. Individuals receiving the information must exercise their independent judgment in

determining its appropriateness for a particular purpose. MALLINCKRODT BAKER, INC. MAKES NO

REPRESENTATIONS OR WARRANTIES, EITHER EXPRESS OR IMPLIED, INCLUDING WITHOUT LIMITATION

ANY WARRANTIES OF MERCHANTABILITY, FITNESS FOR A PARTICULAR PURPOSE WITH RESPECT TO

THE INFORMATION SET FORTH HEREIN OR THE PRODUCT TO WHICH THE INFORMATION REFERS.


137

ACCORDINGLY, MALLINCKRODT BAKER, INC. WILL NOT BE RESPONSIBLE FOR DAMAGES RESULTING

FROM USE OF OR RELIANCE UPON THIS INFORMATION.

*************************************************************************************

***********

Prepared by: Environmental Health & Safety

Phone Number: (314) 654-1600 (U.S.A.)

Appendix C: Engineering Calculations

Design

The diameters calculated in the column design tables represent examples of how each tray diameter

was calculated. The diameter that was used to cost each column was the largest necessary tray

diameter, and is shown in the following tables.


138


139


140


141


142


143


144


145


146


147


148


149


150


151


152


153


154


155


156


157


158


159


160

Costing

Natural Gasoline Process


161


162


163


164

Natural Gasoline Expansion Plant Process


165


166


167


168


169


170


171


172


173

Appendix D: Computer Process Modeling

Aspen HYSYS

Natural Gasoline Expansion Plant Heat Exchangers


174

Appendix E: Economic Spreadsheets

Total Capital Investment

Total bare-moldule costs CFE

Process machinery CPM

Spares Cspare

Storage and surge tanks Cstorage

Initial catalyst charges Ccatalyst

Computers, software, distributed control systems,

instruments, and alarms Ccomp

Total bare-module investment, TBM CTBM,

CBMC

Cost of site preparation Csite

Cost of service facilities Cserv

Allocated costs for utility plants and related facilities Calloc

Total direct permanent investment, DPI CDPI

Cost of contingencies and contractor’s fee Ccont

Total depreciable capital, TDC CTDC

Cost of land Cland

Cost of royalties Croy al

Cost of plant startup Cstartup

Total permanent investment, TPI CTPI

Working capital, WC CWC

Total capital investment, TC CTCI


175

Natural Gasoline Process


176


177


178


179


180


181

Natural Gasoline Expansion Plant


182


183

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