LitLion Chemical - Team 15

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LitLion Chemical - Team 15 The Chemical Engineering Department, The Pennsylvania State University,

University Park, PA, 16802

Presented by: Sarah Ramzy, Daniel Cordova, Brian Klapat, Alexander Hatza

Date: April 21st, 2017

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Executive Summary

A pipeline quality natural gas feed is provided to the Philadelphia, Pennsylvania plant at a

temperature of 80 oF and a pressure of 400 psia. The feed comes in at 49269 pounds per hour, and

is converted to 79106 pounds per hour of 99.89% pure methanol. The team investigated many

design options to produce the highest purity methanol product while accounting for financial

constraints. The final process consists of many subunits which play a crucial role in every step.

The Mercury Removal Unit (MRU) is used to adsorb mercury, a corrosive component, from the

feed. Next, the Amine Treatment Tower enables the removal of hydrogen sulfide, a chemical that

is corrosive and can potentially poison catalysts downstream, via the use of MDEA. The pre-

reformer breaks down larger hydrocarbon chains and forms methane which works to maximize the

conversion of hydrocarbons to synthesis gas in the steam reformer. There are two separator units,

a recycle stream, and a few purging points that work together to remove any unnecessary

byproducts and reuse any valuable assets. A fired heater heats the process stream before entry into

the primary Steam Reformer where the hydrocarbons are broken down to synthesis gas. Following

the steam reformer, the process stream enters the Packed Bed Reactor (PBR) where it undergoes

catalyst-facilitated Fischer-Tropsch synthesis to create the methanol product. Finally, the methanol

product is purified by two identical distillation towers in parallel and pumped to product storage.

The kinetic data used for the reactions was found in provided literature materials.

According to the capital estimate for the project, the plan would require an initial capital

investment of $149 million dollars. This figure includes the equipment cost, site development,

contractor costs, catalyst costs, raw material costs and contingency costs for unforeseen

circumstances. The largest equipment cost is the compressors and expanders ($19.2 million) while

the largest OBL cost is raw material purchasing and product storage ($22.7 million). Operating

costs are also very high with significant utility costs coming from a refrigeration unit and large

amounts of costly waste water disposal.

The economic evaluation of the plant does not currently reflect an economically successful

process. The final After Tax Rate of Return is 7.26%, falling short of our target ATROR of 12%.

The net present value of the plant is, accordingly, -$13.4 million. The analysis used a product price

of 12.5 c/lb product and raw material costs of 2.73 c/lb product. These figures were calculated

based on provided values from NitLion Chemicals. In order for the plant to reach the target

ATROR and a positive net present value, the product price would need to increase to 13.92 c/lb

product (an 11.26% increase).

Following a series of sensitivity analyses and an alternative process study, the team would not

recommend that the plant be built as currently designed. Neither the alternative process study nor

the sensitivity analyses revealed any attainable avenue toward economic success. The team would

recommend that Research and Development at NitLion Chemicals work to find a way to increase

reaction conversion for methanol synthesis, avoid the use of costly refrigeration units, and work

to decrease the amount of water required to achieve high conversion rates in the process before

moving forward with the Philadelphia methanol plant.

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Table of Contents

Background Information ................................................................................................................. 5

Project Background ..................................................................................................................... 5

Technical Information ................................................................................................................. 5

Reactions ..................................................................................................................................... 5

Technological Process Challenges .............................................................................................. 6

Economic Process Challenges ..................................................................................................... 7

Base Case Block Flow Diagram ..................................................................................................... 8

Process Overview and Key Design Variables ................................................................................ 9

Reactors ..................................................................................................................................... 10

Separation Units ........................................................................................................................ 12

Process Variable Controls ......................................................................................................... 15

Safety and Environmental ............................................................................................................. 17

Process Flow Diagram (PFD) ....................................................................................................... 21

Part 1 – MRU, ATU, and Pre-reformer ..................................................................................... 21

Part 2 – Steam Reformer and Separator 1 ................................................................................. 22

Part 3 - Methanol Synthesis Reactor and Separator .................................................................. 23

Part 4 – Distillation Columns .................................................................................................... 24

Process Controls............................................................................................................................ 25

Mass Balance ................................................................................................................................ 26

Energy Balance ............................................................................................................................. 30

Equipment Sizing .......................................................................................................................... 33

Reactors ..................................................................................................................................... 33

Separators .................................................................................................................................. 35

Heat Exchangers ........................................................................................................................ 38

Compressors, Expanders, and Pumps ....................................................................................... 42

Distillation Towers .................................................................................................................... 45

Utility Balance .............................................................................................................................. 49

Outside Battery Limits .................................................................................................................. 51

Economic Analysis ....................................................................................................................... 52

Alternate Process Studies .............................................................................................................. 65

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HYSYS Model Verification .......................................................................................................... 69

Project Definition Rating Index (PDRI) ....................................................................................... 70

PDRI 1 ....................................................................................................................................... 70

PDRI 2 ....................................................................................................................................... 71

PDRI 2i ...................................................................................................................................... 72

Assumptions .................................................................................................................................. 73

Outstanding Issues ........................................................................................................................ 74

Conclusions and Recommendations ............................................................................................. 75

Acknowledgements ....................................................................................................................... 76

Research/References ..................................................................................................................... 77

References: ................................................................................................................................ 78

Appendix A ................................................................................................................................... 80

Equilibrium Data ....................................................................................................................... 80

HYSYS Model .......................................................................................................................... 81

HYSYS Model Continued ..................................................................................................... 82

HYSYS Model Verification ...................................................................................................... 83

Chemical Properties Table ........................................................................................................ 85

Component Data ........................................................................................................................ 87

Initial BFD Created – Gate 1..................................................................................................... 92

Code of Ethics ........................................................................................................................... 94

Appendix B ................................................................................................................................... 97

PDRI – Score Chart Summary .................................................................................................. 97

Appendix C- Supporting Files .................................................................................................... 101

Cash Flow Models ................................................................................................................... 101

Capital Equipment List and Sizes ........................................................................................... 107

Detailed Energy Balance ......................................................................................................... 109

Sizing and Specification Sheets .............................................................................................. 121

HYSYS Model Printout .......................................................................................................... 239

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Background Information

Project Background

To design a chemical plant that achieves a 99.85% Methanol product purity from a feed stream of

natural gas that has already been processed to pipeline quality given a feed at 80 oF and

400 psia. Our team was assigned a raw natural gas feed of 25 MMSCFD. This equals about

approximately 40000 lbm/hr of methane. The plant must achieve 90 mol% conversion of this feed.

This is approximately 71500 lbm/hr of methanol product. The methanol product must be delivered

in a liquid phase to be shipped. Natural gas undergoes a process that involves mercury and sulfur

removal, steam reforming, and methanol synthesis.

Technical Information

Due to the feed containing mercury and hydrogen sulfide inherently, pretreatment must be

performed. The mercury removal unit is modeled as an adsorber, while the hydrogen sulfide

removal unit uses the amine MDEA (Methyldiethanolamine). Steam reforming is the process of

converting methane to CO and H2, or syngas, through a catalytic driven reaction. This is known as

steam reforming because steam must be fed with the natural gas. This process is very well known

and is practiced regularly in refineries. This process will be utilizing the Katalco 57-4Q catalyst.

The methanol is synthesized in this process by use of the Fischer-Tropsch process. Here, methanol

is formed from syngas. This is also a very well-known process that is practiced every day. This

process will be utilizing the Katalco 51-7 catalyst. To achieve the purity of the product, distillation

columns will be used. The following reactions were considered in this process.

Reactions

Steam Reforming:

CH4 + H2O ↔ CO + 3H2, ΔHrxn = 206 kJ/mol

C2H6 + 2 H2O ↔ 2CO + 5H2, ΔHrxn = 347 kJ/mol

C3H8 + 3H2O ↔ 3CO + 7H2, ΔHrxn = 497 kJ/mol

CnHm + nH2O ↔ nCO + (m-1) H2

Secondary Steam Reforming:

2H2 + O2 → 2H2O, ΔHrxn = -482 kJ/mol

CH4 + H2O ↔ CO + 3H2, ΔHrxn = 206 kJ/mol

Water Gas Shift (WGS):

CO + H2O ↔ CO2 + H2, ΔHrxn = -41 kJ/mol

Methanol Reaction:

CO + 2 H2 ↔ CH3OH, ΔHrxn = -91 kJ/mol

CO2 + 3 H2 ↔ CH3OH + H2O

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Methyl Formate Formation

2CO + 2H2 ↔ CH3COOH

Ethanol Formation

2CO + 4H2 ↔ CH5OH + H2O

Other information

Sulfur

Sulfur specification is < 1ppm to avoid poisoning of catalyst. If used assume H2S absorption is

irreversible. Beds are not regenerated.

Mercury

Mercury is a catalyst poison. If used, assume Mercury absorption is irreversible. Beds are not

regenerated.

Methanol Synthesis Reaction

The following table shows the selectivity for the MeOH reaction.

Table 1: Selectivity for Methanol Synthesis Reaction

Chemical Name Formula Wt. %

Methanol CH3OH 99.60

Ethanol C2H5OH 0.25

Methyl formate HCOOCH3 0.15

Technological Process Challenges

A majority of the technological process challenges project lie within the reactors. The steam

reforming reactor requires a very high temperature and large amount of steam to get the necessary

conversion in this reactor. Steam reforming is an endothermic reaction. Therefore, not only does

the feed need to be heated to a high temperature, this temperature needs to be maintained

throughout the reactor. Also, due to this high temperature, fired heaters must be used to reach the

process conditions. Fired heaters are both expensive to purchase and to operate. The methanol

synthesis reactor also has its drawbacks. This reactor requires a very high pressure to achieve a

better conversion. To achieve this pressure, the use of compressors is needed. Compressors are

very expensive and add a large cost to the over IBL capital costs. A higher pressure also affects

the downstream equipment, due to needing to purchase equipment that is rated for these operating

conditions. The catalyst used for this system also does not yield a very strong conversion of syngas.

Therefore, other measures must be taken to reach the hard specifications. These operating

conditions are defined, and we cannot deviate from them.

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Economic Process Challenges

The main challenge in this project is the lower product sales price of methanol. There is a large

amount of equipment needed for this process. This leads to a large capital costs. A large upfront

cost is hard to come back from when considering the annual operating costs and the depreciation

of the dollar. If the company cannot make a large enough profit each year, they will end up not

having a positive net present worth at the end of the project. The use of refrigeration adds a huge

annual utility cost, which will be a hard cost to overcome with the current product price.

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Base Case Block Flow Diagram

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Process Overview and Key Design Variables

Hard Specifications

The hard specifications are that the product must be 99.85 wt% pure Methanol at a production rate

of at least 71500 lbm/hour.

The methanol production process has been modeled in ASPEN HYSYS. Images of this simulation

are shown in the Appendix. HYSYS does not show how the actual plant would be modeled, it

shows a simplification of the processes. This allows for the mathematical calculations to be

determined by the software.

Mercury Removal Unit (MRU)

This unit has been implemented because Mercury, even at low levels, has been known to damage

aluminum heat exchangers to the point of catastrophic failure. Not only can it damage some heat

exchangers, but mercury can hinder other separations down the line. Therefore, the MRU was

placed first to try to limit corrosion and damage throughout the process. Mercury is adsorbed to a

substrate which will be replaced over time as the substrate saturates. In HYSYS, the MRU is

modeled as a component splitter. This allows for the components to be separated to a set value.

This value has been determined through literature. The key design variables of this unit include

the run time, number of units, length and diameter. All of this are important to ensure that the

mercury is properly removed from the natural gas feed.

• Based on 7-day operation

• Number of units: 3

• Length (L): 10.42 ft

• Diameter (D): 3.47 ft

• L/D = 3.00

• Volume: 90.65 ft3

Amine Treatment Unit (ATU)

This unit has been implemented to remove the Sulfur from the system. This utilizes amines to

“sweeten” sour gas, creating sweet gas. Sweet gas is necessary to have, because sour gas is

corrosive and sulfur in the stream would poison the reforming catalysts later in the process. This

unit is a single pass process. In HYSYS, the ATU is modeled as a component splitter. This allows

for the components to be separated to a set value. This value has been determined through

literature.

A key design variable is the MDEA solution flow into the ATU.

• Residence time: 15 min

• MDEA flow rate: 93.5 gpm

• Pressure: 385 psig

• MDEA/Water solution Flow Rate: 240.8 gpm

• MDEA Inlet Temperature: 40 °F

• MDEA Outlet Temperature: 52.43 °F

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Reactors

Pre-Reformer

This is used to convert the heavy hydrocarbons contained within the sweet gas into syngas and

methane to increase the overall conversion in the steam reformation reactor. The conversion for

these heavy hydrocarbons is 100%. In HYSYS, this is modeled as a conversion reactor. This allows

for the conversion of the reactions to be set manually. This is actually a packed bed reactor utilizing

a Nickel-based catalyst.

These temperatures have been chosen because this reflects where the data for the pre-reformer is

accurate. Also, this utilizes the pressure of the feed, thus, we do not need to install expensive

equipment to change this condition this unit is cooled by cooling water, and to ensure it remains

isothermal.

• Operating Temperature: 500 °F

• Operating pressure: 385 psig

Steam Reformer

This unit utilizes the sweet gas and a steam feed to create Synthesis gas (Syngas), CO and H2. In

HYSYS, this is modeled as packed bed reactor. This unit is operated at 1100 oF and 15.3 psig. The

temperature was determined through simulation data from HYSYS, finding this temperature to

yield high conversion. The pressure was determined from a literature review performed on the

process. These process conditions were chosen because they are within the range for the kinetic

data, which has been obtained from literature. The steam reforming reaction is endothermic;

therefore, this unit must be heated to maintain the operating temperature. The catalyst used is

Nickel-based. The overall conversion of the methane in this reactor is approximately 89%.

The steam reformer in HYSYS is modeled as a packed bed reactor. These temperatures have been

chosen because this reflects where the data for the reformer is accurate. Increasing the temperature

at the outlet would increase the overall conversion, but the kinetic data is not supported at a

temperature higher than 1139 F. This operating temperature was chosen based on figure 3, showing

the conversion of methane as function of temperature. If the operating temperature could be

increased further, the conversion would increase as well. The pressure chosen reflects the most

applicable value in reference to the literature. This unit requires a very large amount of water to

obtain the necessary conversion of methane. The H2O/CH4 ratio chosen is out of the normal range,

but it is necessary to convert enough methane with the temperature restrictions.

One drawback of operating at these high temperatures is the need for such a large amount of heat

to be added to maintain the isothermal conditions. The fuel for this system is expensive, and adds

a large expense to yearly utility costs. A second drawback is the need to add a refractory inner

coating to the reactor. This is done to withstand the high temperatures within the reactor. This adds

an additional amount to the overall capital cost for this piece of equipment. The final drawback is

the large amount of water required to operate this unit. At these conditions, 150,000 lb/hr of water

needs to be pumped into the system, and the overall recycle within the system is approximately

400,000 lb/hr. This enables for the conversion to be relatively high. However, due to this enormous

amount of water, the size of the reactor must increase. This not only affects the size of the reactor,

but it affects the amount of catalyst needed and other equipment downstream.

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• Inlet Temperature: 1100 °F

• Outlet Temperature: 1100 °F

• Operating Pressure: 15 psig

• H2O/CH4 ratio: 12.35

• Heat Required: 280 MMBTU/hr

Figure 1: SMR Reactor data at varying temperatures.

Methanol Synthesis

This unit is a catalyst packed bed reactor to create Methanol from Syngas. This reactor has a

selectivity of 99.6%. The other 0.4% forms methyl formate and ethanol. This unit is operated at

500 F and 1000 psig. These process conditions were chosen because they reflect the greatest

conversion of CO to methanol, while being within the specifications of the catalyst’s activity. In

HYSYS, this is modeled as a conversion reactor. The conversion of the Methanol synthesis

reaction is currently set at 18%. The CO2 in the stream, which was formed from the Steam

Reformer, also creates methanol. The conversion of this is very high, because this is a very

favorable reaction. In conjunction, the two reactions create approximately 79100 lb/hr of methanol.

The temperatures are key design variable because they are applicable to the model given. This

reactor is modeled as packed bed reactor. This reactor requires a cooling jacket to maintain the

outlet temperature. The kinetic data for the catalyst does not support a value above 520 °F. The

inlet temperature was set to 500 °F to allow for the exothermic reaction to take place, and limit the

cooling needed for the process stream. This system is operated a very high pressure. This allows

for the greatest conversion of syngas to methanol.

The drawback of operating at this pressure is the need to compress the feed stream and expand the

product for further processing. Compressors have the largest capital cost in the design. However,

as pressure decreases, so does the conversion of CO. Thus, this is a necessary action.

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40

50

60

70

80

90

100

800 850 900 950 1000 1050 1100 1150

Met

han

e C

onver

sio

n (

%)

Inlet/Outlet Temperature (F)

SMR Reactor Temperature vs Conversion %

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• Outlet Temperature: 520 °F

• Operating Pressure: 1000 psig

Separation Units

Separator 1 (V-101)

This unit is a two-phase Vapor-Liquid separator drum located after the Steam Reformer. The feed

into this tower has been cooled down to separate out a large portion of the water at this stage, so it

can be recycled. The lighter gases and the remaining water is separated into the vapor phase, where

it is used further in the process. The following are the key design variables for the vapor-liquid

separator. In order to achieve the separation enough residence time should be allowed to have an

effective separation.

• Settling Velocity: 8.14 ft/s

• Residence time: 25 min

• Vessel diameter (d): 10.6 ft


• L/d: 3

• Demister height: 0.5 ft

Separator 2 (V-100)

This unit is a two-phase vapor-liquid separator tower located after the Methanol Synthesis Reactor.

This is used to separate the lighter gases from the heavier liquids. Almost all of the gases in the

system are removed with this separator. This allows for the equipment down the line to be smaller,

due to less mass flow going into it. Also, by removing a majority of the light gases from the system,

the distillation columns have an easier separation, yielding a purer product. The same key design

variables apply to this separator:

• Settling Velocity: 9.71 ft/s

• Residence time (liquid hold-up): 10 min

• Vessel diameter (d): 7.1 ft


• L/d: 3

• Demister height: 0.5 ft

Distillation Columns 1 and 2

These distillation towers are used to separate out the methanol product to a 99.85 wt% purity. Two

distillation towers are required to be operated in parallel for this separation to occur. The feed

stream is split into two identical streams, halving the flowrate. Each column requires 40 trays, and

have reflux ratios of 4.5. These feed stream is fed into the middle of the column at 206.7 F and 15

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psig. This temperature is chosen because it best reflects the temperature of the feed tray. This

prevents excessive flooding in the column. This pressure is chosen to prevent backpressure in the

column during startup. The trays are chosen to be sieve trays, with a tray spacing of 15 inches.

These columns are able to achieve all hard specifications needed.

T-100:

• Theoretical Stages: 40


• Inlet Pressure: 30 psia

• Bottoms Temperature: 228°F

• Bottoms Pressure: 29 psia

• Distillate Temperature: 175 °F

• Distillate Pressure: 24.5 psia

• Reflux Ratio: 4.1

T-103:

• Theoretical Stages: 40


• Inlet Pressure: 30 psia

• Bottoms Temperature: 228 °F

• Bottoms Pressure: 29 psia

• Distillate Temperature: 175 °F

• Distillate Pressure: 24.5 psia

• Reflux Ratio: 4.1

In order to obtain an initial guess for the theoretical number of stages we used the McCabe Thiele

diagram. Based on this diagram, the theoretical number of stages is 11. However, this module is

under the assumption of infinite reflux which is not the case. As a result of this assumption, the

model underestimates the number of theoretical stages. The theoretical number of trays was

determined to be 40 plates using HYSYS.

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In order to obtain the best estimate for the optimal number of trays, we plotted the number of trays

versus the reflux ratio. Based on the tangent on the plot, the optimal number of theoretical stages

is 42. The design we selected has 40 trays, a reboiler, and a condenser bringing the total number

of theoretical stages to 42. The corresponding reflux ratio is 4.1.

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35

40

45

50

55

3 . 4 3 . 6 3 . 8 4 4 . 2 4 . 4 4 . 6 4 . 8

NU

MB

ER O

F TR

AYS

REFLUX RATIO

NUMBER OF TRAYS VERSUS REFLUX RATIO

Figure 3: A plot showing the number of trays versus the reflux ratio.

Figure 2: McCabe Thiele diagram used to estimate an initial guess for the theoretical number of stages.

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In order to determine the optimal feed tray to the distillation column we plotted the reflux ratio

versus the feed tray position. Based on the plot, the optimal feed tray for the distillation column is

at tray 28 (where the minimum in the plot occurs). Tray 28 was thus selected as the feed to the

distillation column.

Figure 4: A plot showing the relationship between the reflux ratio and the feed tray.

We determined the appropriate pressure for the column based on the feed pressure. We also

accounted for a 0.15 psia pressure drop per tray in the column which enabled us to set the pressure

distillate and bottoms. The HYSYS simulation determined the temperature of the distillate and

bottoms based on the pressure input we set.

Process Variable Controls

Heat Exchangers

Heat exchangers are a pivotal tool in this process. The reactor feed into the Steam Reformer must

be heated to 1100 ⁰F. This requires a fired heater to be used. Because the reactor outlet is

maintained at this temperature, the process stream can be used multiple times through counter-

current shell and tube heat exchangers to heat streams. The utilization of process to process heat

transfer is useful because the fired heater requires a large amount of energy and utilities to operate.

When process to process heat transfer does not reach the proper temperature needed for a stream,

cooling water is used. Cooling water is utilized at a temperature range of 90 - 110 oF and 56.3 psig.

3.9

4

4.1

4.2

4.3

4.4

4.5

4.6

4.7

4.8

1 0 1 5 2 0 2 5 3 0 3 5 4 0 4 5

REF

LUX

RA

TIO

FEED TRAY

REFLUX RATIO VERSUS FEED TRAY

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These utility streams are used in multiple counter-current shell and tube heat exchangers to produce

the temperatures desired.

Compressors, Expanders, and Pumps

The pressure of the streams can be controlled through the use of compressors, expanders, and

valves. The natural gas feed comes into the system at 385.4 psig. This must be lowered to enter

the prereformer, due to where the kinetic values have been defined at. An expander is used to

release the pressure of the stream to 15.3 psig, capturing the energy. This energy can be reused

later to power a process. The stream needs to be repressurized, but this time to 1000 psig, to enter

the methanol synthesis reactor. This is done by a two stage compressing and cooling. A pressure

ratio of 6 is maintained through these compressors, to achieve the desired pressure. Another

expander is used to depressurize the stream to 15.6 psig, so it can be separated more efficiently.

Two pumps are required to pump the product and waste water to their respective storage/treatment

facilities. The first pump must pressurize the product streams to 85 psig so it can travel the 1-mile

distance to the storage tank farm, and the second pump must pressurize all waste water to 35 psig

so it can be taken to the bio-treatment area.

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Safety and Environmental

General Safety

A safe working environment is one of the most important factors when designing a chemical plant.

Although the process will be mostly automated, any people on the site will wear proper personal

protective equipment (PPE). Due to the flammability of many chemicals that are present in the

plant, such as methane, all equipment must be properly grounded or bonded. To further reduce the

chance of fire or explosion, smoking will not be permitted on company grounds, and only

authorized personnel will be allowed. Regular preventive maintenance on the pipes and equipment

in the system is necessary to ensure there are no leaks. Personnel will receive the proper training

on how to respond to emergency situations in the plant. Standard operating procedures (SOP)

should be carefully followed whenever the plans needs to shut down or start up.

Major Chemicals

There are many chemicals used in this process, some of these chemicals could be harmful if not

handled properly. These chemicals are described below. All of the chemical data is tabulated in

Appendix A.

Methane:

Methane is a clear, flammable gas. Methane has an LFL of 5%, and an autoignition temperature

of 1076 °F. There is a large amount of methane in the system; thus, it is important to monitor this

chemical. The NIOSH REL-TWA is 800 ppm.

Carbon Monoxide:

Carbon monoxide is a clear, flammable gas. Carbon monoxide has an LFL of 12.5%, and an

autoignition temperature of 1128 °F. This chemical has a chance for asphyxiation. The NIOSH

REL-TWA is 35 ppm.

Ethane:

Ethane is a clear, flammable gas. Ethane has an LFL of 3%, and an autoignition temperature of

959 °F. Although this gas is a risk, there is a very small amount in the system.

Propane:

Propane is a clear, flammable gas. Propane has an LFL of 1.8%, and an autoignition temperature

of 859 °F. Although this gas is a risk, there is a very small amount in the system. . The OSHA PEL

is 1000 ppm.

N-Butane:

N-Butane is a clear, flammable gas. N-Butane has an LFL of 1.6%, and an autoignition temperature

of 761 °F. Although this gas is a risk, there is a very small amount in the system.

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Iso-Butane:

Iso-butane is a clear, flammable gas. Iso-butane has an LFL of 1.6%, and an autoignition

temperature of 890 °F. Although this gas is a risk, there is a very small amount in the system.

N-Pentane:

N-Pentane is a clear, flammable liquid at room temperature. N-Pentane has an LFL of 1.5%, and

an autoignition temperature of 500 °F. Although this is a risk, there is a very small amount in the

system. The OSHA PEL is 1000 ppm.

Iso-Pentane:

Iso-Pentane is a clear, flammable liquid at room temperature. Iso-Pentane has an LFL of 1.4%,

and an autoignition temperature of 788 °F. Although this is a risk, there is a very small amount in

the system. The TLV-TWA is 1000 ppm.

Cyclohexane:

Cyclohexane is a clear, flammable liquid at room temperature. Cyclohexane has an LFL of 1.3%,

and an autoignition temperature of 473 °F. Although this is a risk, there is a very small amount in

the system. The OSHA PEL is 300 ppm.

Nitrogen:

Nitrogen is a clear gas. This chemical has a chance for asphyxiation; thus, it is important to ensure

there are no leaks in the system.

Hydrogen Sulfide:

Hydrogen sulfide is a clear, flammable gas. Hydrogen sulfide has an LFL of 4%, and an

autoignition temperature of 450 °F. The OSHA PEL is 20-50 ppm.

Mercury:

Mercury is a heavy, silver-white, odorless liquid. The boiling point of mercury is 674 °F. Elemental

mercury is toxic to the central and peripheral nervous systems. The inhalation of mercury vapor

can produce harmful effects on the nervous, digestive and immune systems, lungs and kidneys,

and may be fatal. The NIOSH REL TWA is 0.1 mg/m3 and 0.05 mg/m3 for liquid mercury and

mercury vapor, respectively.

Methanol:

Methanol is a clear, flammable liquid. Methanol has an LFL of 6%, and an autoignition

temperature of 878 °F. However, during our process, the methanol will not reach the autoignition

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temperature. The methanol will his is our main product, and it is important to ensure this is safely

transported throughout the plant. The OSHA PEL is 200 ppm.

Methyl Formate:

Methyl formate is a clear, flammable liquid. Methyl formate has an LFL of 4.5%, and an

autoignition temperature of 869 °F. However, during our process, the methyl formate will not

reach the autoignition temperature. Although this is a risk, there is only a small amount of this in

our system. The NIOSH REL-TWA is 1000 ppm.

Ethanol:

Ethanol is a clear, flammable liquid. Ethanol has an LFL of 3.3%, and an autoignition temperature

of 685 °F. However, during our process, the ethanol will not reach the autoignition temperature.

Although this is a risk, there is only a small amount of this in our system. The NIOSH REL-TWA

is 1000 ppm.

Environmental Hazards

Mercury:

Mercury is a highly toxic compound which must be handled with care. The presence of mercury

in a process can lead to catalyst poisoning thus decreasing the efficiency of the plant. This deemed

the presence of a Mercury Removal Unit a necessity. Mercury can be leaked to the environment

through improper disposal techniques or faulty unit design. Such leaks can affect the quality of the

air or surrounding bodies of water. As a result, not only are the lives of humans (workers or

citizens) endangered, but also the aquatics. Solid mercurial waste emitted into the environment

commonly originates from process units dedicated to mercury removal. It is critical that mercury

waste be properly labeled and disposed of according to federal laws and regulations.

Hydrogen Sulfide:

Hydrogen Sulfide is a highly toxic and corrosive chemical. The Presence of hydrogen sulfide can

damage the piping and cause a decrease in plant efficiency as well as leaks to the environment.

Hydrogen sulfide can react with any present moisture to produce sulfuric acid which is highly

corrosive. The Sulfuric acid can lead to detrimental effects on the process and affect the health of

employees and well-being of the environment. As a result, proper precautions must be taken with

regards to designing the Amine Treatment Tower. Since sulfuric acid is highly corrosive to carbon

steel, the tower will be designed using a stainless steel lining instead. By using two materials of

construction, the risk of corrosion is reduced while minimizing the cost for this piece of equipment.

Additionally, temperature and pressure sensors will be installed prior to and immediately after the

ATU to regulate these variables. The addition of valves will ensure adjustment of pressure in case

of deviations.

Methane:

Methane is a highly flammable compound which must be monitored properly. Methane is also a

greenhouse gas with a global warming potential approximately 20 times that of carbon dioxide.

This makes methane leaks a severe threat to the environment. As result, the methane has to be

20

discarded properly. We have elected to flare the excess methane thus converting it to carbon

dioxide and hydrogen gas which has a much less damaging effect on the environment.

Methanol:

Methanol is a highly flammable. It is also considered to be a Volatile Organic Compound (VOC).

VOCs have been linked to many cases of irritations, cancer, increased emergency room visits, and

in some severe cases, premature death. Exposure can occur through inhalation or digestion.

Methanol released into the environment can also affect the wildlife including animals, fish, and

even plants. It can affect the growth of these living things and can even lead to their death.

Methanol can be transported through air or water. This makes it necessary to carry appropriate

measures to ensure proper transportation of methanol product from the plant to the storage

containers and safe delivery to the consumers’ destination. It is also crucial that the storage

containers be constructed from materials that can resist the flammability of methanol.

21

Process Flow Diagram (PFD)

Part 1 – MRU, ATU, and Pre-reformer

22

Part 2 – Steam Reformer and Separator 1

23

Part 3 - Methanol Synthesis Reactor and Separator

24

Part 4 – Distillation Columns

25

Process Controls

Process controls were added in order to control the process safely and profitably. The controllers

added to the process regulate key variables such as flow, temperature, levels, and pressure. The

Process Flow Diagram (PFD) shows key control requirements on major equipment such as heat

exchangers, separator vessels, distillation towers, and reactors.

In the process the MRU and AMU contain a pressure control system that is designed to regulate

vapor pressure. The vapor/liquid separators contain two control systems, level and pressure

control. The level control is essential in order to prevent liquid overflow in the vessel. A pressure

control system was used in order to control overpressure in the separator vessels. All of the heat

exchangers in the process are controlled by temperature controllers to ensure that the process

streams are not overheated to a point that is not safe for the equipment design specifications. Flow

controllers are used to regulate the recycle streams in the process. Distillations columns are

controlled by placing pressure controller on the condenser, temperature control on the reboiler,

and flow control on the reflux. The reactors can be regulated using temperature or pressure control

systems. Reactors also contain relief valves or rupture disks; however, these safety elements are

not shown in the PFD.

26

Mass Balance

The sheet shown in the images below is a full-process mass balance of each of the components

involved in the process. The numbers at the top of many of the columns correspond to stream

numbers from the BFD. All numbered columns are shown in lbm/hr units, but some of the

intermediary columns that involved reaction calculations are shown in lbmol/hr. All assumptions

regarding efficiency, selectivity, and conversion are shown below the rows of the mass balance in

demarcated boxes. The cells of the spreadsheet are defined using references to ensure for ease of

adjustment when changing parameters later on in the project. Each unit of the process is shown

spanning over the columns that are involved in that unit for ease of organization. The mass balance

values have been updated with the conversions, separation efficiencies, and other information from

the HYSYS simulation. The final product of this mass balance is currently a 99.89% pure stream

of MeOH flowing at 79100 lbm/hr. This final product meets both hard specifications.

27

1 2 3 4 5 6 7 8 9 10 11 12 13 14 16 17 18 19 20 21 22 23 24

NG Hg Amine MDEA Tower To aminePre

Steam into WORK CO created H2 created Remaining Water WGSWork Col CO

H2 Water CO2 createdReactor Output Water R-2 Work Col CO to MeOH WGS CO2 to MeOH NH3 R-2 Gas Purge Feed to

Feed Removed Tower feed treatment reformer Pre COL in reformer in reformer HC Used created created Consumed from Recycle Feed and side rxns Output Splitter

Component BP (F) # of C # of H MW Mol % Xi lbm/hr 1.00 Feed feed Reformer lbmol lbmol lbmol lbmol lbmol lbmol lbmol lbmol lbmol lbmol lbmol lbmol R-1 lbmol lbmol lbmol lbmol lbmol lb/hr

H2 -423 2.01588 0.00 0.00 8265.48 8.12 64.58 130.19 130.19 7657.15 15566.09 0.01 0.00 0.00 15566.08 7721.73 6857.35 8728.64 2577.87 2575.29 5191.48 5191.46 0.02 0.01 0.01 0.01 0.00 0.01 0.00 0.00

N2 -320 28.0134 1.51 0.0151 1160.72 1160.72 1160.72 1160.72 0.00 1160.72 1160.72 1160.72 0.00 0.00 0.00 1160.72 41.43 41.43 41.43 41.43 40.58 1136.65 1136.43 0.21 0.11 0.11 0.11 0.00 0.11 0.00 0.00

CO -313 28.0101 0.00 0.00 2823.03 70.58 14.12 395.37 395.40 2388.06 67285.25 0.07 0.04 0.04 67285.17 2402.18 1969.78 98.49 98.49 98.49 2758.69 2758.52 0.18 0.09 0.09 0.09 0.00 0.09 0.00 0.00

CH4 -258.7 1 4 16.0425 90.11 0.9011 39666.98 39666.98 39666.98 39666.98 0.00 2472.62 2472.62 7417.86 0.00 2472.62 2752.45 44156.20 44156.20 2752.45 2449.68 7657.15 2634.54 3868.64 0.00 0.00 0.00 3868.64 3868.64 3868.09 0.55 0.27 0.27 0.27 0.00 0.27 0.00 0.00

C2H6 -128.2 2 6 30.069 4 0.04 3300.37 3300.37 3300.37 3300.37 0.00 109.76 219.52 548.80 0.00 219.52 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

CO2 -110.3 44.0095 1.5 0.015 1811.43 1811.43 1811.43 1811.43 0.00 41.16 97.62 4296.23 4298.68 123.24 9722.44 4.89 2.44 2.44 9717.55 220.81 220.81 2092.10 41.84 41.84 1841.45 1807.42 34.03 17.01 17.01 17.01 0.00 17.01 0.00 0.00

H2S -77 34.0809 1.5 0.015 1402.77 1402.77 1402.77 1402.77 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

C3H8 -43.6 3 8 44.0956 0.9 0.009 1088.98 1088.98 1088.98 1088.98 0.00 24.70 74.09 172.87 0.00 74.09 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

NH3 -28 17.03052 0.00 0.00 0.00 0.00 0.00 0.00 1.72 29.27 13.80 15.47 7.74 7.74 7.74 0.00 7.74 0.00 0.00

iC4H10 11 4 10 58.124 0.2 0.002 318.98 318.98 318.98 318.98 0.00 5.49 21.95 49.39 0.00 21.95 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

nC4H10 31 4 10 58.124 0.1 0.001 159.49 159.49 159.49 159.49 0.00 2.74 10.98 24.70 0.00 10.98 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

iC5H12 82 5 12 72.1488 0.1 0.001 197.98 197.98 197.98 197.98 0.00 2.74 13.72 30.18 0.00 13.72 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

HCOOCH3 89 2 4 60.052 0.00 0.00 0.00 0.00 0.00 0.00 0.20 0.20 0.20 0.20 12.15 4.90 7.24 3.62 3.62 3.62 0.00 3.62 0.00 0.00

nC5H12 97 5 12 72.1488 0.05 0.0005 98.99 98.99 98.99 98.99 0.00 1.37 6.86 15.09 0.00 6.86 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

CH3OH 148.5 1 4 32.0419 0.00 0.00 0.00 0.00 0.00 0.00 430.84 2050.26 2481.10 79499.07 151.49 79347.58 39673.79 39673.79 39511.18 162.61 39511.18 162.61 0.00

C2H5OH 173 2 6 46.06844 0.00 0.00 0.00 0.00 0.00 0.00 0.57 0.57 0.57 0.57 26.45 0.26 26.19 13.09 13.09 13.09 0.00 13.09 0.00 0.00

C6H12 177 6 12 84.16595 0.02 0.0002 46.19 46.19 46.19 46.19 0.00 0.55 3.29 6.59 0.00 3.29 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

H2O 212 18.01528 0.01 0.0001 4.94 4.94 4.94 72062.10 72062.10 4.94 150000.00 8326.54 5503.51 8255.96 8199.50 147716.35 555054.14 30810.19 28175.64 507592.11 471299.06 63961.27 407337.79 36293.05 2014.57 2015.14 143.85 2194.11 2194.11 39527.48 73.82 39453.66 19726.83 19726.83 0.07 19726.76 0.07 19726.76 0.00

MDEA 476.6 119.163 52182.90 52182.90 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

Hg 674 200.592 0.002 0.00002 11.01 11.01 11.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

SUM Total 49269 11 49258 124245 125648 47855 150000 197855 605195 33563 4838 605195 471304 63964 407340 133891 133891 15006 118885 59443 59443 39553 19889 39553 19889 0.0

197855 0 MeOH purity: 0.9989 MeOH purity: 0.9989

New Data for MeOH Reactor mol%

Assume 96% conversion for CO2 to MeOH CO and CO2 to MeOH 0.99445 0.996409231 At 350 F Total Product: 79106 lb/h

89% 0.89 Conversion 93 (Combined weight conversion) 100% CO2

stream 1 NG feed 49269 stream 2 Hg removed 11.01 REFORM CnHm+nH20 --> nCO + (n+1/m)H2 0.2% 0.002 Ethanol 0.00381 0.002655183 Distillate:

4 MDEA/Water 124245 5 S removed 125647.77 CO to Me CO+3 H2 --> CH4 + H2O 0.4% 4.00E-03 Assume 95% conversion for Water Gas Shift Reaction M-Formate 0.00175 0.000935586 MeOH removal (95%) 0.9999 103742

7 Water 150000.000 12 Water Purge 63963.75 WGS CO+H2O --> CO2+H2 Conversion 0.95 Overall Conv Basis 100% of light impurities come out in Distillate 0.1

25 MMSCFD 80 F 17 Gas Purge 15006.18 WGS CO+H2O --> CO2+H2 95.00% CO Bottoms:

2744 lbmol/hr 400 psia 21 MeOH out 1 39553.19 mol% Conversion (%) Assume 90% conversion for N2 to ammonia reaction CO MeOH CO+2 H2 --> CH3OH 18.0% CO Assume 99.99% C6H12 removal 0.95

22 EtOH + water 1 19889.37 Methanol CH3OH 0.996 0.99745 0.9475775 Conversion 0.05 CO2 MeOH CO2+3 H2 --> CH3OH + H2O 98.00% H2 Assume 99.99% H2O removal 0.9999

23 MeOH out 2 39553.19 Ethanol C2H5OH 0.0025 0.00174 0.001653 EtOH 2CO+4 H2 --> C2H5OH + H2O 0.20% CO Assume 99.99% Hg removal 0.9999

24 EtOH + water 2 19889.37 Methyl Formate HCOOCH3 0.0015 0.000801 0.00076095 0.95 Recycle Ratio (from HYSYS) M Form 2 CO+ 2H2 --> HCOOCH3 0.09% CO Assume 99.99% C2H5OH 0.9

323514 323514 NH3 N2+3H2 --> 2 NH3 0.10% H2

H2O % 58%

MDEA % 42% 0.864287302

Fed 56.25 m^3/h

MDEA 52182.9 lbm/h CnHm to CO and H2 1

H2O 72062.1 lbm/h 97.9% methanation CO basis 0.975 Cmpd Separation Fraction Cmpd Sep Frac Cmpd Sep Frac

5% water/gas CO basis 0.02 Methane 3.15E-07 Methane 0.999858662 Methane 0.999990423

Reforming Rxns ΔH 5% unreacted CO, H2 CO basis H2O 9.28E-01 H2O 0.001867591 H2O 3.37504E-06

CH4 206 kJ/mol Ethane 1.46E-08 Ethane 0.999867412 Ethane 0.999996758

C2H6 347 kJ/mol CO2 5.03E-04 CO2 0.981522045 CO2 0.999997535

C3H8 497 kJ/mol Nitrogen 2.72E-06 Nitrogen 0.999812273 Nitrogen 0.999979051

WGS -41 kJ/mol 0.9 Propane 2.21E-10 Propane 0.999959365 Propane 0.999998188

methanating -206 kJ/mol at 206 i-Butane 1.87E-12 i-Butane 0.999994304 i-Butane 0.999998421

n-Butane 2.79E-12 n-Butane 0.999988783 n-Butane 0.99999895

i-Pentane 1.86E-14 i-Pentane 0.999998316 i-Pentane 0.999999266

Rate Conv n-Pentane 2.04E-14 n-Pentane 0.999997249 n-Pentane 0.999999451

CO + 2 H2 80% 0.8 0.95 Cyclohexane 2.50E-13 Cyclohexane 0.999836038 Cyclohexane 0.999999973

CO2 + 3H2 10% 0.1 0.95 CO 1.11E-06 CO 0.999936067 CO 0.999942284

WGS 10% 0.1 0.95 Hydrogen 4.79E-07 Hydrogen 0.999995524 Hydrogen 0.999803972

MDEAmine 0.00E+00 MDEAmine 0 MDEAmine 0

Methanol 0.00E+00 Methanol 0.001905574 Methanol 0.995901422

Ethanol 0.00E+00 Ethanol 0.009750125 Ethanol 0.99999352

M-Formate 0.00E+00 M-Formate 0.40363721 M-Formate 0.999999934

Ammonia 0.00E+00 Ammonia 0.471333638 Ammonia 0.999999601

Air 0.00E+00 Air 0 Air 0

Oxygen 0.00E+00 Oxygen 0 Oxygen 0

10% Efficiency on Recycle

Reactor 2 Reaction Rates

MDEA fed

Pre-reformer Conversion and Efficiency

Separator 1 Separator 2 Distillation Column

Recycle Efficiency for S-2

WGS

CO2 formation

NG FEED specs

Reactor 2 Conversion and Selectivity

Selectivity wt%

95%

Conversion

Overall

Balance (This

should be 0

for all)

NG Feed Composition

Assumptions on MeOH Synthesis Reactor

Overall Mass Balance Reactor 1 Conversion and Selectivity

Inlet streams (lb/hr) Outlet streams (lb/hr) Methane Conversion: Distillation Column 1+2 Efficiency

Feed to D-1 Feed to D-2Methanol

Product

EtOH and

Waste Water

Methanol

Product

EtOH and

Waste Water

Primary RXNs

methanati

ngLeaving Pre-

reformerFeed to Reactor

Waste Water

PurgeRecycle to Feed

Amine Treatment Tower (ATT) Steam Reformation Reactor (R-1) MeOH Synthesis Reactor (R-2) Distillation Column (D-1)

Mercury Removal Unit (MRU) Pre-Reformer Separator-1 (S-1) Separator-2 (S-2) Distillation Column (D-2)

28

1 2 3 4 5 6 7 8 9 10 11 12 13 14 16 17 18 19 20 21 22 23 24






H2 -423 2.01588 0.00 0.00 8265.48 8.12 64.58 130.19 130.19 7657.15 15566.09 0.01 0.00 0.00 15566.08 7721.73 6857.35 8728.64 2577.87 2575.29 5191.48 5191.46 0.02 0.01 0.01 0.01 0.00 0.01 0.00 0.00

N2 -320 28.0134 1.51 0.0151 1160.72 1160.72 1160.72 1160.72 0.00 1160.72 1160.72 1160.72 0.00 0.00 0.00 1160.72 41.43 41.43 41.43 41.43 40.58 1136.65 1136.43 0.21 0.11 0.11 0.11 0.00 0.11 0.00 0.00

CO -313 28.0101 0.00 0.00 2823.03 70.58 14.12 395.37 395.40 2388.06 67285.25 0.07 0.04 0.04 67285.17 2402.18 1969.78 98.49 98.49 98.49 2758.69 2758.52 0.18 0.09 0.09 0.09 0.00 0.09 0.00 0.00

CH4 -258.7 1 4 16.0425 90.11 0.9011 39666.98 39666.98 39666.98 39666.98 0.00 2472.62 2472.62 7417.86 0.00 2472.62 2752.45 44156.20 44156.20 2752.45 2449.68 7657.15 2634.54 3868.64 0.00 0.00 0.00 3868.64 3868.64 3868.09 0.55 0.27 0.27 0.27 0.00 0.27 0.00 0.00

C2H6 -128.2 2 6 30.069 4 0.04 3300.37 3300.37 3300.37 3300.37 0.00 109.76 219.52 548.80 0.00 219.52 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

CO2 -110.3 44.0095 1.5 0.015 1811.43 1811.43 1811.43 1811.43 0.00 41.16 97.62 4296.23 4298.68 123.24 9722.44 4.89 2.44 2.44 9717.55 220.81 220.81 2092.10 41.84 41.84 1841.45 1807.42 34.03 17.01 17.01 17.01 0.00 17.01 0.00 0.00

H2S -77 34.0809 1.5 0.015 1402.77 1402.77 1402.77 1402.77 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

C3H8 -43.6 3 8 44.0956 0.9 0.009 1088.98 1088.98 1088.98 1088.98 0.00 24.70 74.09 172.87 0.00 74.09 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

NH3 -28 17.03052 0.00 0.00 0.00 0.00 0.00 0.00 1.72 29.27 13.80 15.47 7.74 7.74 7.74 0.00 7.74 0.00 0.00

iC4H10 11 4 10 58.124 0.2 0.002 318.98 318.98 318.98 318.98 0.00 5.49 21.95 49.39 0.00 21.95 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

nC4H10 31 4 10 58.124 0.1 0.001 159.49 159.49 159.49 159.49 0.00 2.74 10.98 24.70 0.00 10.98 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

iC5H12 82 5 12 72.1488 0.1 0.001 197.98 197.98 197.98 197.98 0.00 2.74 13.72 30.18 0.00 13.72 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

HCOOCH3 89 2 4 60.052 0.00 0.00 0.00 0.00 0.00 0.00 0.20 0.20 0.20 0.20 12.15 4.90 7.24 3.62 3.62 3.62 0.00 3.62 0.00 0.00

nC5H12 97 5 12 72.1488 0.05 0.0005 98.99 98.99 98.99 98.99 0.00 1.37 6.86 15.09 0.00 6.86 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

CH3OH 148.5 1 4 32.0419 0.00 0.00 0.00 0.00 0.00 0.00 430.84 2050.26 2481.10 79499.07 151.49 79347.58 39673.79 39673.79 39511.18 162.61 39511.18 162.61 0.00

C2H5OH 173 2 6 46.06844 0.00 0.00 0.00 0.00 0.00 0.00 0.57 0.57 0.57 0.57 26.45 0.26 26.19 13.09 13.09 13.09 0.00 13.09 0.00 0.00

C6H12 177 6 12 84.16595 0.02 0.0002 46.19 46.19 46.19 46.19 0.00 0.55 3.29 6.59 0.00 3.29 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

H2O 212 18.01528 0.01 0.0001 4.94 4.94 4.94 72062.10 72062.10 4.94 150000.00 8326.54 5503.51 8255.96 8199.50 147716.35 555054.14 30810.19 28175.64 507592.11 471299.06 63961.27 407337.79 36293.05 2014.57 2015.14 143.85 2194.11 2194.11 39527.48 73.82 39453.66 19726.83 19726.83 0.07 19726.76 0.07 19726.76 0.00

MDEA 476.6 119.163 52182.90 52182.90 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

Hg 674 200.592 0.002 0.00002 11.01 11.01 11.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

SUM Total 49269 11 49258 124245 125648 47855 150000 197855 605195 33563 4838 605195 471304 63964 407340 133891 133891 15006 118885 59443 59443 39553 19889 39553 19889 0.0













323514 323514 NH3 N2+3H2 --> 2 NH3 0.10% H2

H2O % 58%

MDEA % 42% 0.864287302

Fed 56.25 m^3/h






C2H6 347 kJ/mol CO2 5.03E-04 CO2 0.981522045 CO2 0.999997535








CO2 + 3H2 10% 0.1 0.95 CO 1.11E-06 CO 0.999936067 CO 0.999942284











MDEA fed




WGS

CO2 formation

NG FEED specs


Selectivity wt%

95%

Conversion

Overall

Balance (This

should be 0

for all)

NG Feed Composition





Product

EtOH and

Waste Water

Methanol

Product

EtOH and

Waste Water

Primary RXNs

methanati

ngLeaving Pre-


Waste Water




1 2 3 4 5 6 7 8 9 10 11 12 13 14 16 17 18 19 20 21 22 23 24






H2 -423 2.01588 0.00 0.00 8265.48 8.12 64.58 130.19 130.19 7657.15 15566.09 0.01 0.00 0.00 15566.08 7721.73 6857.35 8728.64 2577.87 2575.29 5191.48 5191.46 0.02 0.01 0.01 0.01 0.00 0.01 0.00 0.00

N2 -320 28.0134 1.51 0.0151 1160.72 1160.72 1160.72 1160.72 0.00 1160.72 1160.72 1160.72 0.00 0.00 0.00 1160.72 41.43 41.43 41.43 41.43 40.58 1136.65 1136.43 0.21 0.11 0.11 0.11 0.00 0.11 0.00 0.00

CO -313 28.0101 0.00 0.00 2823.03 70.58 14.12 395.37 395.40 2388.06 67285.25 0.07 0.04 0.04 67285.17 2402.18 1969.78 98.49 98.49 98.49 2758.69 2758.52 0.18 0.09 0.09 0.09 0.00 0.09 0.00 0.00

CH4 -258.7 1 4 16.0425 90.11 0.9011 39666.98 39666.98 39666.98 39666.98 0.00 2472.62 2472.62 7417.86 0.00 2472.62 2752.45 44156.20 44156.20 2752.45 2449.68 7657.15 2634.54 3868.64 0.00 0.00 0.00 3868.64 3868.64 3868.09 0.55 0.27 0.27 0.27 0.00 0.27 0.00 0.00

C2H6 -128.2 2 6 30.069 4 0.04 3300.37 3300.37 3300.37 3300.37 0.00 109.76 219.52 548.80 0.00 219.52 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

CO2 -110.3 44.0095 1.5 0.015 1811.43 1811.43 1811.43 1811.43 0.00 41.16 97.62 4296.23 4298.68 123.24 9722.44 4.89 2.44 2.44 9717.55 220.81 220.81 2092.10 41.84 41.84 1841.45 1807.42 34.03 17.01 17.01 17.01 0.00 17.01 0.00 0.00

H2S -77 34.0809 1.5 0.015 1402.77 1402.77 1402.77 1402.77 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

C3H8 -43.6 3 8 44.0956 0.9 0.009 1088.98 1088.98 1088.98 1088.98 0.00 24.70 74.09 172.87 0.00 74.09 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

NH3 -28 17.03052 0.00 0.00 0.00 0.00 0.00 0.00 1.72 29.27 13.80 15.47 7.74 7.74 7.74 0.00 7.74 0.00 0.00

iC4H10 11 4 10 58.124 0.2 0.002 318.98 318.98 318.98 318.98 0.00 5.49 21.95 49.39 0.00 21.95 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

nC4H10 31 4 10 58.124 0.1 0.001 159.49 159.49 159.49 159.49 0.00 2.74 10.98 24.70 0.00 10.98 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

iC5H12 82 5 12 72.1488 0.1 0.001 197.98 197.98 197.98 197.98 0.00 2.74 13.72 30.18 0.00 13.72 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

HCOOCH3 89 2 4 60.052 0.00 0.00 0.00 0.00 0.00 0.00 0.20 0.20 0.20 0.20 12.15 4.90 7.24 3.62 3.62 3.62 0.00 3.62 0.00 0.00

nC5H12 97 5 12 72.1488 0.05 0.0005 98.99 98.99 98.99 98.99 0.00 1.37 6.86 15.09 0.00 6.86 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

CH3OH 148.5 1 4 32.0419 0.00 0.00 0.00 0.00 0.00 0.00 430.84 2050.26 2481.10 79499.07 151.49 79347.58 39673.79 39673.79 39511.18 162.61 39511.18 162.61 0.00

C2H5OH 173 2 6 46.06844 0.00 0.00 0.00 0.00 0.00 0.00 0.57 0.57 0.57 0.57 26.45 0.26 26.19 13.09 13.09 13.09 0.00 13.09 0.00 0.00

C6H12 177 6 12 84.16595 0.02 0.0002 46.19 46.19 46.19 46.19 0.00 0.55 3.29 6.59 0.00 3.29 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

H2O 212 18.01528 0.01 0.0001 4.94 4.94 4.94 72062.10 72062.10 4.94 150000.00 8326.54 5503.51 8255.96 8199.50 147716.35 555054.14 30810.19 28175.64 507592.11 471299.06 63961.27 407337.79 36293.05 2014.57 2015.14 143.85 2194.11 2194.11 39527.48 73.82 39453.66 19726.83 19726.83 0.07 19726.76 0.07 19726.76 0.00

MDEA 476.6 119.163 52182.90 52182.90 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

Hg 674 200.592 0.002 0.00002 11.01 11.01 11.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

SUM Total 49269 11 49258 124245 125648 47855 150000 197855 605195 33563 4838 605195 471304 63964 407340 133891 133891 15006 118885 59443 59443 39553 19889 39553 19889 0.0













323514 323514 NH3 N2+3H2 --> 2 NH3 0.10% H2

H2O % 58%

MDEA % 42% 0.864287302

Fed 56.25 m^3/h






C2H6 347 kJ/mol CO2 5.03E-04 CO2 0.981522045 CO2 0.999997535








CO2 + 3H2 10% 0.1 0.95 CO 1.11E-06 CO 0.999936067 CO 0.999942284











MDEA fed




WGS

CO2 formation

NG FEED specs


Selectivity wt%

95%

Conversion

Overall

Balance (This

should be 0

for all)

NG Feed Composition





Product

EtOH and

Waste Water

Methanol

Product

EtOH and

Waste Water

Primary RXNs

methanati

ngLeaving Pre-


Waste Water




29

1 2 3 4 5 6 7 8 9 10 11 12 13 14 16 17 18 19 20 21 22 23 24






H2 -423 2.01588 0.00 0.00 8265.48 8.12 64.58 130.19 130.19 7657.15 15566.09 0.01 0.00 0.00 15566.08 7721.73 6857.35 8728.64 2577.87 2575.29 5191.48 5191.46 0.02 0.01 0.01 0.01 0.00 0.01 0.00 0.00

N2 -320 28.0134 1.51 0.0151 1160.72 1160.72 1160.72 1160.72 0.00 1160.72 1160.72 1160.72 0.00 0.00 0.00 1160.72 41.43 41.43 41.43 41.43 40.58 1136.65 1136.43 0.21 0.11 0.11 0.11 0.00 0.11 0.00 0.00

CO -313 28.0101 0.00 0.00 2823.03 70.58 14.12 395.37 395.40 2388.06 67285.25 0.07 0.04 0.04 67285.17 2402.18 1969.78 98.49 98.49 98.49 2758.69 2758.52 0.18 0.09 0.09 0.09 0.00 0.09 0.00 0.00

CH4 -258.7 1 4 16.0425 90.11 0.9011 39666.98 39666.98 39666.98 39666.98 0.00 2472.62 2472.62 7417.86 0.00 2472.62 2752.45 44156.20 44156.20 2752.45 2449.68 7657.15 2634.54 3868.64 0.00 0.00 0.00 3868.64 3868.64 3868.09 0.55 0.27 0.27 0.27 0.00 0.27 0.00 0.00

C2H6 -128.2 2 6 30.069 4 0.04 3300.37 3300.37 3300.37 3300.37 0.00 109.76 219.52 548.80 0.00 219.52 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

CO2 -110.3 44.0095 1.5 0.015 1811.43 1811.43 1811.43 1811.43 0.00 41.16 97.62 4296.23 4298.68 123.24 9722.44 4.89 2.44 2.44 9717.55 220.81 220.81 2092.10 41.84 41.84 1841.45 1807.42 34.03 17.01 17.01 17.01 0.00 17.01 0.00 0.00

H2S -77 34.0809 1.5 0.015 1402.77 1402.77 1402.77 1402.77 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

C3H8 -43.6 3 8 44.0956 0.9 0.009 1088.98 1088.98 1088.98 1088.98 0.00 24.70 74.09 172.87 0.00 74.09 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

NH3 -28 17.03052 0.00 0.00 0.00 0.00 0.00 0.00 1.72 29.27 13.80 15.47 7.74 7.74 7.74 0.00 7.74 0.00 0.00

iC4H10 11 4 10 58.124 0.2 0.002 318.98 318.98 318.98 318.98 0.00 5.49 21.95 49.39 0.00 21.95 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

nC4H10 31 4 10 58.124 0.1 0.001 159.49 159.49 159.49 159.49 0.00 2.74 10.98 24.70 0.00 10.98 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

iC5H12 82 5 12 72.1488 0.1 0.001 197.98 197.98 197.98 197.98 0.00 2.74 13.72 30.18 0.00 13.72 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

HCOOCH3 89 2 4 60.052 0.00 0.00 0.00 0.00 0.00 0.00 0.20 0.20 0.20 0.20 12.15 4.90 7.24 3.62 3.62 3.62 0.00 3.62 0.00 0.00

nC5H12 97 5 12 72.1488 0.05 0.0005 98.99 98.99 98.99 98.99 0.00 1.37 6.86 15.09 0.00 6.86 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

CH3OH 148.5 1 4 32.0419 0.00 0.00 0.00 0.00 0.00 0.00 430.84 2050.26 2481.10 79499.07 151.49 79347.58 39673.79 39673.79 39511.18 162.61 39511.18 162.61 0.00

C2H5OH 173 2 6 46.06844 0.00 0.00 0.00 0.00 0.00 0.00 0.57 0.57 0.57 0.57 26.45 0.26 26.19 13.09 13.09 13.09 0.00 13.09 0.00 0.00

C6H12 177 6 12 84.16595 0.02 0.0002 46.19 46.19 46.19 46.19 0.00 0.55 3.29 6.59 0.00 3.29 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

H2O 212 18.01528 0.01 0.0001 4.94 4.94 4.94 72062.10 72062.10 4.94 150000.00 8326.54 5503.51 8255.96 8199.50 147716.35 555054.14 30810.19 28175.64 507592.11 471299.06 63961.27 407337.79 36293.05 2014.57 2015.14 143.85 2194.11 2194.11 39527.48 73.82 39453.66 19726.83 19726.83 0.07 19726.76 0.07 19726.76 0.00

MDEA 476.6 119.163 52182.90 52182.90 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

Hg 674 200.592 0.002 0.00002 11.01 11.01 11.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

SUM Total 49269 11 49258 124245 125648 47855 150000 197855 605195 33563 4838 605195 471304 63964 407340 133891 133891 15006 118885 59443 59443 39553 19889 39553 19889 0.0













323514 323514 NH3 N2+3H2 --> 2 NH3 0.10% H2

H2O % 58%

MDEA % 42% 0.864287302

Fed 56.25 m^3/h






C2H6 347 kJ/mol CO2 5.03E-04 CO2 0.981522045 CO2 0.999997535








CO2 + 3H2 10% 0.1 0.95 CO 1.11E-06 CO 0.999936067 CO 0.999942284











MDEA fed




WGS

CO2 formation

NG FEED specs


Selectivity wt%

95%

Conversion

Overall

Balance (This

should be 0

for all)

NG Feed Composition





Product

EtOH and

Waste Water

Methanol

Product

EtOH and

Waste Water

Primary RXNs

methanati

ngLeaving Pre-


Waste Water




1 2 3 4 5 6 7 8 9 10 11 12 13 14 16 17 18 19 20 21 22 23 24






H2 -423 2.01588 0.00 0.00 8265.48 8.12 64.58 130.19 130.19 7657.15 15566.09 0.01 0.00 0.00 15566.08 7721.73 6857.35 8728.64 2577.87 2575.29 5191.48 5191.46 0.02 0.01 0.01 0.01 0.00 0.01 0.00 0.00

N2 -320 28.0134 1.51 0.0151 1160.72 1160.72 1160.72 1160.72 0.00 1160.72 1160.72 1160.72 0.00 0.00 0.00 1160.72 41.43 41.43 41.43 41.43 40.58 1136.65 1136.43 0.21 0.11 0.11 0.11 0.00 0.11 0.00 0.00

CO -313 28.0101 0.00 0.00 2823.03 70.58 14.12 395.37 395.40 2388.06 67285.25 0.07 0.04 0.04 67285.17 2402.18 1969.78 98.49 98.49 98.49 2758.69 2758.52 0.18 0.09 0.09 0.09 0.00 0.09 0.00 0.00

CH4 -258.7 1 4 16.0425 90.11 0.9011 39666.98 39666.98 39666.98 39666.98 0.00 2472.62 2472.62 7417.86 0.00 2472.62 2752.45 44156.20 44156.20 2752.45 2449.68 7657.15 2634.54 3868.64 0.00 0.00 0.00 3868.64 3868.64 3868.09 0.55 0.27 0.27 0.27 0.00 0.27 0.00 0.00

C2H6 -128.2 2 6 30.069 4 0.04 3300.37 3300.37 3300.37 3300.37 0.00 109.76 219.52 548.80 0.00 219.52 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

CO2 -110.3 44.0095 1.5 0.015 1811.43 1811.43 1811.43 1811.43 0.00 41.16 97.62 4296.23 4298.68 123.24 9722.44 4.89 2.44 2.44 9717.55 220.81 220.81 2092.10 41.84 41.84 1841.45 1807.42 34.03 17.01 17.01 17.01 0.00 17.01 0.00 0.00

H2S -77 34.0809 1.5 0.015 1402.77 1402.77 1402.77 1402.77 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

C3H8 -43.6 3 8 44.0956 0.9 0.009 1088.98 1088.98 1088.98 1088.98 0.00 24.70 74.09 172.87 0.00 74.09 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

NH3 -28 17.03052 0.00 0.00 0.00 0.00 0.00 0.00 1.72 29.27 13.80 15.47 7.74 7.74 7.74 0.00 7.74 0.00 0.00

iC4H10 11 4 10 58.124 0.2 0.002 318.98 318.98 318.98 318.98 0.00 5.49 21.95 49.39 0.00 21.95 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

nC4H10 31 4 10 58.124 0.1 0.001 159.49 159.49 159.49 159.49 0.00 2.74 10.98 24.70 0.00 10.98 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

iC5H12 82 5 12 72.1488 0.1 0.001 197.98 197.98 197.98 197.98 0.00 2.74 13.72 30.18 0.00 13.72 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

HCOOCH3 89 2 4 60.052 0.00 0.00 0.00 0.00 0.00 0.00 0.20 0.20 0.20 0.20 12.15 4.90 7.24 3.62 3.62 3.62 0.00 3.62 0.00 0.00

nC5H12 97 5 12 72.1488 0.05 0.0005 98.99 98.99 98.99 98.99 0.00 1.37 6.86 15.09 0.00 6.86 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

CH3OH 148.5 1 4 32.0419 0.00 0.00 0.00 0.00 0.00 0.00 430.84 2050.26 2481.10 79499.07 151.49 79347.58 39673.79 39673.79 39511.18 162.61 39511.18 162.61 0.00

C2H5OH 173 2 6 46.06844 0.00 0.00 0.00 0.00 0.00 0.00 0.57 0.57 0.57 0.57 26.45 0.26 26.19 13.09 13.09 13.09 0.00 13.09 0.00 0.00

C6H12 177 6 12 84.16595 0.02 0.0002 46.19 46.19 46.19 46.19 0.00 0.55 3.29 6.59 0.00 3.29 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

H2O 212 18.01528 0.01 0.0001 4.94 4.94 4.94 72062.10 72062.10 4.94 150000.00 8326.54 5503.51 8255.96 8199.50 147716.35 555054.14 30810.19 28175.64 507592.11 471299.06 63961.27 407337.79 36293.05 2014.57 2015.14 143.85 2194.11 2194.11 39527.48 73.82 39453.66 19726.83 19726.83 0.07 19726.76 0.07 19726.76 0.00

MDEA 476.6 119.163 52182.90 52182.90 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

Hg 674 200.592 0.002 0.00002 11.01 11.01 11.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

SUM Total 49269 11 49258 124245 125648 47855 150000 197855 605195 33563 4838 605195 471304 63964 407340 133891 133891 15006 118885 59443 59443 39553 19889 39553 19889 0.0













323514 323514 NH3 N2+3H2 --> 2 NH3 0.10% H2

H2O % 58%

MDEA % 42% 0.864287302

Fed 56.25 m^3/h






C2H6 347 kJ/mol CO2 5.03E-04 CO2 0.981522045 CO2 0.999997535








CO2 + 3H2 10% 0.1 0.95 CO 1.11E-06 CO 0.999936067 CO 0.999942284











MDEA fed




WGS

CO2 formation

NG FEED specs


Selectivity wt%

95%

Conversion

Overall

Balance (This

should be 0

for all)

NG Feed Composition





Product

EtOH and

Waste Water

Methanol

Product

EtOH and

Waste Water

Primary RXNs

methanati

ngLeaving Pre-


Waste Water




30

Energy Balance

The sheets shown in the images below show an energy balance around each unit operations in the

plant. This shows the necessary heat flow needed, in BTU/hr, for each major and minor unit. The

values in the energy balance were calculated in most instances using the heat capacity, temperature,

and mass flowrate. The heat capacities and heat of vaporizations of the streams were found using

HYSYS. The mass flowrates for these streams are calculated from the mass balance. The duty of

each heat exchanger, the heats needed and generated by each reaction, and the duties of each

condenser and reboiler are shown in the sizing sheets. Where available, heat of reaction was taken

from given data, however for the reformation of longer chain hydrocarbons, the heat of reaction

was calculated manually using enthalpy of formation values from the NIST Chemistry WebBook.

The following equations were used to complete an energy balance on the process units.

Equation 1: 𝑄 (𝐵𝑇𝑈

ℎ𝑟) = 𝑚 ∗ 𝐶𝑝,𝑎𝑣𝑔 ∗ ∆𝑇

• 𝑚 𝑖𝑠 𝑚𝑎𝑠𝑠 𝑓𝑙𝑜𝑤 𝑟𝑎𝑡𝑒 (𝑙𝑏

ℎ𝑟)

• 𝐶𝑝 𝑖𝑠 𝑡ℎ𝑒 𝑎𝑣𝑒𝑟𝑎𝑔𝑒 𝑠𝑝𝑒𝑐𝑖𝑓𝑖𝑐 ℎ𝑒𝑎𝑡 𝑜𝑓 𝑡ℎ𝑒 𝑠𝑡𝑟𝑒𝑎𝑚 (𝐵𝑇𝑈

𝑙𝑏−℉)

• ∆𝑇 𝑖𝑠 𝑡ℎ𝑒 𝑡𝑒𝑚𝑝𝑒𝑟𝑎𝑡𝑢𝑟𝑒 𝑐ℎ𝑎𝑛𝑔𝑒


ℎ𝑟) = 𝑚 ∗ ∆𝐻𝑣𝑎𝑝


ℎ𝑟)

• ∆𝐻𝑣𝑎𝑝 𝑖𝑠 𝑡ℎ𝑒 ℎ𝑒𝑎𝑡 𝑜𝑓 𝑣𝑎𝑝𝑜𝑟𝑖𝑧𝑎𝑡𝑖𝑜𝑛


ℎ𝑟) = 𝑚 ∗ ∆𝐻𝑟𝑥𝑛


ℎ𝑟)

• ∆𝐻𝑟𝑥𝑛 𝑖𝑠 𝑡ℎ𝑒 ℎ𝑒𝑎𝑡 𝑜𝑓 𝑟𝑒𝑎𝑐𝑡𝑖𝑜𝑛

Below is a sample calculation for how the necessary heat flow was calculated for a stream that had

vaporized, utilizing equations 1 and 2.

Unit: E-102

𝑄 = 𝑚 ∗ 𝐶𝑝,𝑎𝑣𝑔 ∗ ∆𝑇 + 𝑚∆𝐻𝑣𝑎𝑝

Here, there are two components vaporizing, H2O and Methanol. These values are taken from the

energy balance. The mass fraction of vaporizing component was obtained from HYSYS.

𝑄 = 𝑚 ∗ 𝐶𝑝,𝑎𝑣𝑔 ∗ ∆𝑇 + 𝑚𝑣𝑎𝑝,𝐶𝐻3𝑂𝐻∆𝐻𝑣𝑎𝑝 + 𝑚𝑣𝑎𝑝,𝐻2𝑂∆𝐻𝑣𝑎𝑝

𝑄 = 1.19𝐸5 ∗1.053 + 0.9811

2∗ (206 − 30) + 3922 ∗ 40.66 + 11674 ∗ 504.5

𝑄 = 27.3 𝑀𝑀𝐵𝑇𝑈/ℎ𝑟

31

STREAM: 1 2 3 4 5 6 10 6_1 10_1 6_1 7 Reactor Feed 8 8_1 13_1 9 10_1 9_1 10_2 9_1 9_2 10 10_2 CW1 CW2 10_3 CW3 CW4 10_4 11 14 11 13 12 14 14_1 CW5 CW6 14_2 14_3 18 18_1 14_4 CW7 CW8 14_5 16 16_1 R1 R2 16_2 17 18 18_1 19 20 17 17_1 17_2 20 23 24 19 21 22 12 24 25 21 23 26 25 25_1 26 26_1 CW9 CW10 26_2

REPEAT REPEAT REPEAT REPEAT REPEAT REPEAT REPEAT REPEAT REPEAT REPEAT REPEAT REPEAT REPEAT

Flow 4.93E+04 1.10E+01 4.93E+04 1.24E+05 1.26E+05 4.79E+04 6.05E+05 4.79E+04 6.05E+05 4.79E+04 1.50E+05 1.98E+05 1.98E+05 1.98E+05 4.07E+05 6.05E+05 6.05E+05 6.05E+05 6.05E+05 6.05E+05 6.05E+05 6.05E+05 6.05E+05 5.93E+06 5.93E+06 6.05E+05 1.67E+06 1.67E+06 6.05E+05 4.71E+05 1.34E+05 4.71E+05 4.07E+05 6.40E+04 1.34E+05 1.34E+05 1.94E+06 1.94E+06 1.34E+05 1.34E+05 1.19E+05 1.19E+05 1.34E+05 5.19E+05 5.19E+05 1.34E+05 1.34E+05 1.34E+05 N/A N/A 1.34E+05 1.50E+04 1.19E+05 1.19E+05 5.94E+04 5.94E+04 1.50E+04 7.40E+03 7.60E+03 5.94E+04 3.96E+04 1.99E+04 5.94E+04 3.96E+04 1.99E+04 6.40E+04 1.99E+04 1.04E+05 3.96E+04 3.96E+04 7.91E+04 1.04E+05 1.04E+05 7.91E+04 7.91E+04 2.02E+05 2.02E+05 7.91E+04

T (F) 80.00 80.00 80.00 40.00 52.43 72.00 1100.00 500.00 1063.56 500.00 500.00 474.70 480.40 218.90 160.00 216.80 1063.00 520.00 637.19 520.00 1100.00 1100.00 637.19 90.00 110.00 280.00 90.00 110.00 160.00 160.00 160.00 160.00 160.00 160.00 160.00 695.41 90.00 110.00 280.00 891.86 30.00 206.00 609.27 90.00 110.00 500.00 520.00 185.00 -44.00 -44.00 30.00 30.00 30.00 206.00 206.30 206.30 30.00 30.00 30.00 206.30 174.60 238.50 206.30 174.60 238.50 160.00 238.50 192.80 174.60 174.60 174.60 192.80 195.80 174.60 175.00 90.00 110.00 120.00

P (psia) 400.00 400.00 400.00 400.00 400.00 400.00 30.00 400.00 30.00 400.00 400.00 400.00 400.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 70.00 70.00 30.00 70.00 70.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 180.00 70.00 70.00 180.00 1015.00 30.00 30.00 1015.00 70.00 70.00 1015.00 1015.00 30.00 63.00 63.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 24.50 29.00 30.00 24.50 29.00 30.00 29.00 29.00 24.50 24.50 24.50 29.00 50.00 24.50 100.00 70.00 70.00 100.00

Cp (BTU/lbm-F) 0.5409 0.3149 0.5410 0.8371 0.8380 0.5477 0.5762 0.6899 0.5731 0.6899 0.6805 0.5715 0.5793 0.5080 1.0010 0.7587 0.5731 0.4966 0.5835 0.4966 0.6116 0.6409 0.5835 0.9980 0.9982 0.5109 0.9980 0.9982 0.9287 1.0010 0.6724 1.0010 1.0010 1.0010 0.6724 0.7075 0.9980 0.9982 0.6829 0.7350 1.0530 0.9811 0.7096 0.9980 0.9982 0.7069 0.6401 1.0930 0.5590 0.5590 0.9733 1.3510 1.0530 0.9811 1.0580 1.0580 1.3510 1.3510 1.3510 1.0580 0.9565 1.0110 1.0580 0.9565 1.0110 1.0010 1.0110 1.0040 0.9565 0.9565 0.9565 1.0040 1.0040 0.9565 0.9558 0.9980 0.9982 0.9555

dh vap (BTU/lb) 446.0 162.8 226.5 587.0 596.3 207.6 1502.0 207.6 1502.0 207.6 780.6 1025.0 1455.0 1293.0 951.9 1455.0 1502.0 -786.0 1502.0 -786.0 1455.0 -1003.0 1502.0 908.1 908.1 -1014.0 908.1 908.1 1502.0 951.9 948.9 951.9 951.9 951.9 948.9 1850.0 908.1 908.1 977.0 1255.0 638.7 638.7 1425.0 908.1 908.1 1425.0 -800.9 -800.9 174.0 174.0 -543.9 1283.0 638.7 638.7 664.9 664.9 1283.0 1283.0 1283.0 664.9 482.1 954.5 664.9 482.1 954.5 951.9 954.5 954.1 482.1 482.1 482.0 954.1 928.8 482.0 423.6 908.1 908.1 423.6

Heat Flow (BTU/hr) 2.13E+06 2.77E+02 2.13E+06 4.16E+06 5.52E+06 1.89E+06 3.84E+08 1.65E+07 3.69E+08 1.65E+07 5.10E+07 5.37E+07 5.51E+07 2.20E+07 6.52E+07 9.95E+07 3.69E+08 1.56E+08 2.25E+08 1.56E+08 4.07E+08 4.27E+08 2.25E+08 5.32E+08 6.51E+08 8.66E+07 1.50E+08 1.83E+08 8.99E+07 7.55E+07 1.44E+07 7.55E+07 6.52E+07 1.02E+07 1.44E+07 6.59E+07 1.75E+08 2.14E+08 2.56E+07 8.78E+07 3.76E+06 2.40E+07 5.79E+07 4.66E+07 5.70E+07 4.73E+07 4.46E+07 2.71E+07 7.00E+07 7.00E+07 3.91E+06 6.08E+05 3.76E+06 2.40E+07 1.30E+07 1.30E+07 6.08E+05 3.00E+05 3.08E+05 1.30E+07 6.61E+06 4.80E+06 1.30E+07 6.61E+06 4.80E+06 1.02E+07 4.80E+06 2.01E+07 6.61E+06 6.61E+06 1.32E+07 2.01E+07 2.04E+07 1.32E+07 1.32E+07 1.81E+07 2.21E+07 9.07E+06

P-101

E-109TEE-101

TEE-102

T-100

T-103

MIX-103

MIX-104

P-100

K-104

E-103

K-105

E-102

E-105

MeOH Synthesis

K-102

E-104

V-100FH-100

Reformer

E-101

E-108

V-101

TEE-100

MRU

ATU

E-100

MIX-100

Prereformer

K-100

MIX-101

E-107




T (F) 80.00 80.00 80.00 40.00 52.43 72.00 1100.00 500.00 1063.56 500.00 500.00 474.70 480.40 218.90 160.00 216.80 1063.00 520.00 637.19 520.00 1100.00 1100.00 637.19 90.00 110.00 280.00 90.00 110.00 160.00 160.00 160.00 160.00 160.00 160.00 160.00 695.41 90.00 110.00 280.00 891.86 30.00 206.00 609.27 90.00 110.00 500.00 520.00 185.00 -44.00 -44.00 30.00 30.00 30.00 206.00 206.30 206.30 30.00 30.00 30.00 206.30 174.60 238.50 206.30 174.60 238.50 160.00 238.50 192.80 174.60 174.60 174.60 192.80 195.80 174.60 175.00 90.00 110.00 120.00

P (psia) 400.00 400.00 400.00 400.00 400.00 400.00 30.00 400.00 30.00 400.00 400.00 400.00 400.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 70.00 70.00 30.00 70.00 70.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 180.00 70.00 70.00 180.00 1015.00 30.00 30.00 1015.00 70.00 70.00 1015.00 1015.00 30.00 63.00 63.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 24.50 29.00 30.00 24.50 29.00 30.00 29.00 29.00 24.50 24.50 24.50 29.00 50.00 24.50 100.00 70.00 70.00 100.00

Cp (BTU/lbm-F) 0.5409 0.3149 0.5410 0.8371 0.8380 0.5477 0.5762 0.6899 0.5731 0.6899 0.6805 0.5715 0.5793 0.5080 1.0010 0.7587 0.5731 0.4966 0.5835 0.4966 0.6116 0.6409 0.5835 0.9980 0.9982 0.5109 0.9980 0.9982 0.9287 1.0010 0.6724 1.0010 1.0010 1.0010 0.6724 0.7075 0.9980 0.9982 0.6829 0.7350 1.0530 0.9811 0.7096 0.9980 0.9982 0.7069 0.6401 1.0930 0.5590 0.5590 0.9733 1.3510 1.0530 0.9811 1.0580 1.0580 1.3510 1.3510 1.3510 1.0580 0.9565 1.0110 1.0580 0.9565 1.0110 1.0010 1.0110 1.0040 0.9565 0.9565 0.9565 1.0040 1.0040 0.9565 0.9558 0.9980 0.9982 0.9555

dh vap (BTU/lb) 446.0 162.8 226.5 587.0 596.3 207.6 1502.0 207.6 1502.0 207.6 780.6 1025.0 1455.0 1293.0 951.9 1455.0 1502.0 -786.0 1502.0 -786.0 1455.0 -1003.0 1502.0 908.1 908.1 -1014.0 908.1 908.1 1502.0 951.9 948.9 951.9 951.9 951.9 948.9 1850.0 908.1 908.1 977.0 1255.0 638.7 638.7 1425.0 908.1 908.1 1425.0 -800.9 -800.9 174.0 174.0 -543.9 1283.0 638.7 638.7 664.9 664.9 1283.0 1283.0 1283.0 664.9 482.1 954.5 664.9 482.1 954.5 951.9 954.5 954.1 482.1 482.1 482.0 954.1 928.8 482.0 423.6 908.1 908.1 423.6


P-101

E-109TEE-101

TEE-102

T-100

T-103

MIX-103

MIX-104

P-100

K-104

E-103

K-105

E-102

E-105

MeOH Synthesis

K-102

E-104

V-100FH-100

Reformer

E-101

E-108

V-101

TEE-100

MRU

ATU

E-100

MIX-100

Prereformer

K-100

MIX-101

E-107




T (F) 80.00 80.00 80.00 40.00 52.43 72.00 1100.00 500.00 1063.56 500.00 500.00 474.70 480.40 218.90 160.00 216.80 1063.00 520.00 637.19 520.00 1100.00 1100.00 637.19 90.00 110.00 280.00 90.00 110.00 160.00 160.00 160.00 160.00 160.00 160.00 160.00 695.41 90.00 110.00 280.00 891.86 30.00 206.00 609.27 90.00 110.00 500.00 520.00 185.00 -44.00 -44.00 30.00 30.00 30.00 206.00 206.30 206.30 30.00 30.00 30.00 206.30 174.60 238.50 206.30 174.60 238.50 160.00 238.50 192.80 174.60 174.60 174.60 192.80 195.80 174.60 175.00 90.00 110.00 120.00

P (psia) 400.00 400.00 400.00 400.00 400.00 400.00 30.00 400.00 30.00 400.00 400.00 400.00 400.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 70.00 70.00 30.00 70.00 70.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 180.00 70.00 70.00 180.00 1015.00 30.00 30.00 1015.00 70.00 70.00 1015.00 1015.00 30.00 63.00 63.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 24.50 29.00 30.00 24.50 29.00 30.00 29.00 29.00 24.50 24.50 24.50 29.00 50.00 24.50 100.00 70.00 70.00 100.00

Cp (BTU/lbm-F) 0.5409 0.3149 0.5410 0.8371 0.8380 0.5477 0.5762 0.6899 0.5731 0.6899 0.6805 0.5715 0.5793 0.5080 1.0010 0.7587 0.5731 0.4966 0.5835 0.4966 0.6116 0.6409 0.5835 0.9980 0.9982 0.5109 0.9980 0.9982 0.9287 1.0010 0.6724 1.0010 1.0010 1.0010 0.6724 0.7075 0.9980 0.9982 0.6829 0.7350 1.0530 0.9811 0.7096 0.9980 0.9982 0.7069 0.6401 1.0930 0.5590 0.5590 0.9733 1.3510 1.0530 0.9811 1.0580 1.0580 1.3510 1.3510 1.3510 1.0580 0.9565 1.0110 1.0580 0.9565 1.0110 1.0010 1.0110 1.0040 0.9565 0.9565 0.9565 1.0040 1.0040 0.9565 0.9558 0.9980 0.9982 0.9555

dh vap (BTU/lb) 446.0 162.8 226.5 587.0 596.3 207.6 1502.0 207.6 1502.0 207.6 780.6 1025.0 1455.0 1293.0 951.9 1455.0 1502.0 -786.0 1502.0 -786.0 1455.0 -1003.0 1502.0 908.1 908.1 -1014.0 908.1 908.1 1502.0 951.9 948.9 951.9 951.9 951.9 948.9 1850.0 908.1 908.1 977.0 1255.0 638.7 638.7 1425.0 908.1 908.1 1425.0 -800.9 -800.9 174.0 174.0 -543.9 1283.0 638.7 638.7 664.9 664.9 1283.0 1283.0 1283.0 664.9 482.1 954.5 664.9 482.1 954.5 951.9 954.5 954.1 482.1 482.1 482.0 954.1 928.8 482.0 423.6 908.1 908.1 423.6


P-101

E-109TEE-101

TEE-102

T-100

T-103

MIX-103

MIX-104

P-100

K-104

E-103

K-105

E-102

E-105

MeOH Synthesis

K-102

E-104

V-100FH-100

Reformer

E-101

E-108

V-101

TEE-100

MRU

ATU

E-100

MIX-100

Prereformer

K-100

MIX-101

E-107




T (F) 80.00 80.00 80.00 40.00 52.43 72.00 1100.00 500.00 1063.56 500.00 500.00 474.70 480.40 218.90 160.00 216.80 1063.00 520.00 637.19 520.00 1100.00 1100.00 637.19 90.00 110.00 280.00 90.00 110.00 160.00 160.00 160.00 160.00 160.00 160.00 160.00 695.41 90.00 110.00 280.00 891.86 30.00 206.00 609.27 90.00 110.00 500.00 520.00 185.00 -44.00 -44.00 30.00 30.00 30.00 206.00 206.30 206.30 30.00 30.00 30.00 206.30 174.60 238.50 206.30 174.60 238.50 160.00 238.50 192.80 174.60 174.60 174.60 192.80 195.80 174.60 175.00 90.00 110.00 120.00

P (psia) 400.00 400.00 400.00 400.00 400.00 400.00 30.00 400.00 30.00 400.00 400.00 400.00 400.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 70.00 70.00 30.00 70.00 70.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 180.00 70.00 70.00 180.00 1015.00 30.00 30.00 1015.00 70.00 70.00 1015.00 1015.00 30.00 63.00 63.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 24.50 29.00 30.00 24.50 29.00 30.00 29.00 29.00 24.50 24.50 24.50 29.00 50.00 24.50 100.00 70.00 70.00 100.00

Cp (BTU/lbm-F) 0.5409 0.3149 0.5410 0.8371 0.8380 0.5477 0.5762 0.6899 0.5731 0.6899 0.6805 0.5715 0.5793 0.5080 1.0010 0.7587 0.5731 0.4966 0.5835 0.4966 0.6116 0.6409 0.5835 0.9980 0.9982 0.5109 0.9980 0.9982 0.9287 1.0010 0.6724 1.0010 1.0010 1.0010 0.6724 0.7075 0.9980 0.9982 0.6829 0.7350 1.0530 0.9811 0.7096 0.9980 0.9982 0.7069 0.6401 1.0930 0.5590 0.5590 0.9733 1.3510 1.0530 0.9811 1.0580 1.0580 1.3510 1.3510 1.3510 1.0580 0.9565 1.0110 1.0580 0.9565 1.0110 1.0010 1.0110 1.0040 0.9565 0.9565 0.9565 1.0040 1.0040 0.9565 0.9558 0.9980 0.9982 0.9555

dh vap (BTU/lb) 446.0 162.8 226.5 587.0 596.3 207.6 1502.0 207.6 1502.0 207.6 780.6 1025.0 1455.0 1293.0 951.9 1455.0 1502.0 -786.0 1502.0 -786.0 1455.0 -1003.0 1502.0 908.1 908.1 -1014.0 908.1 908.1 1502.0 951.9 948.9 951.9 951.9 951.9 948.9 1850.0 908.1 908.1 977.0 1255.0 638.7 638.7 1425.0 908.1 908.1 1425.0 -800.9 -800.9 174.0 174.0 -543.9 1283.0 638.7 638.7 664.9 664.9 1283.0 1283.0 1283.0 664.9 482.1 954.5 664.9 482.1 954.5 951.9 954.5 954.1 482.1 482.1 482.0 954.1 928.8 482.0 423.6 908.1 908.1 423.6


P-101

E-109TEE-101

TEE-102

T-100

T-103

MIX-103

MIX-104

P-100

K-104

E-103

K-105

E-102

E-105

MeOH Synthesis

K-102

E-104

V-100FH-100

Reformer

E-101

E-108

V-101

TEE-100

MRU

ATU

E-100

MIX-100

Prereformer

K-100

MIX-101

E-107




T (F) 80.00 80.00 80.00 40.00 52.43 72.00 1100.00 500.00 1063.56 500.00 500.00 474.70 480.40 218.90 160.00 216.80 1063.00 520.00 637.19 520.00 1100.00 1100.00 637.19 90.00 110.00 280.00 90.00 110.00 160.00 160.00 160.00 160.00 160.00 160.00 160.00 695.41 90.00 110.00 280.00 891.86 30.00 206.00 609.27 90.00 110.00 500.00 520.00 185.00 -44.00 -44.00 30.00 30.00 30.00 206.00 206.30 206.30 30.00 30.00 30.00 206.30 174.60 238.50 206.30 174.60 238.50 160.00 238.50 192.80 174.60 174.60 174.60 192.80 195.80 174.60 175.00 90.00 110.00 120.00

P (psia) 400.00 400.00 400.00 400.00 400.00 400.00 30.00 400.00 30.00 400.00 400.00 400.00 400.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 70.00 70.00 30.00 70.00 70.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 180.00 70.00 70.00 180.00 1015.00 30.00 30.00 1015.00 70.00 70.00 1015.00 1015.00 30.00 63.00 63.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 24.50 29.00 30.00 24.50 29.00 30.00 29.00 29.00 24.50 24.50 24.50 29.00 50.00 24.50 100.00 70.00 70.00 100.00

Cp (BTU/lbm-F) 0.5409 0.3149 0.5410 0.8371 0.8380 0.5477 0.5762 0.6899 0.5731 0.6899 0.6805 0.5715 0.5793 0.5080 1.0010 0.7587 0.5731 0.4966 0.5835 0.4966 0.6116 0.6409 0.5835 0.9980 0.9982 0.5109 0.9980 0.9982 0.9287 1.0010 0.6724 1.0010 1.0010 1.0010 0.6724 0.7075 0.9980 0.9982 0.6829 0.7350 1.0530 0.9811 0.7096 0.9980 0.9982 0.7069 0.6401 1.0930 0.5590 0.5590 0.9733 1.3510 1.0530 0.9811 1.0580 1.0580 1.3510 1.3510 1.3510 1.0580 0.9565 1.0110 1.0580 0.9565 1.0110 1.0010 1.0110 1.0040 0.9565 0.9565 0.9565 1.0040 1.0040 0.9565 0.9558 0.9980 0.9982 0.9555

dh vap (BTU/lb) 446.0 162.8 226.5 587.0 596.3 207.6 1502.0 207.6 1502.0 207.6 780.6 1025.0 1455.0 1293.0 951.9 1455.0 1502.0 -786.0 1502.0 -786.0 1455.0 -1003.0 1502.0 908.1 908.1 -1014.0 908.1 908.1 1502.0 951.9 948.9 951.9 951.9 951.9 948.9 1850.0 908.1 908.1 977.0 1255.0 638.7 638.7 1425.0 908.1 908.1 1425.0 -800.9 -800.9 174.0 174.0 -543.9 1283.0 638.7 638.7 664.9 664.9 1283.0 1283.0 1283.0 664.9 482.1 954.5 664.9 482.1 954.5 951.9 954.5 954.1 482.1 482.1 482.0 954.1 928.8 482.0 423.6 908.1 908.1 423.6


P-101

E-109TEE-101

TEE-102

T-100

T-103

MIX-103

MIX-104

P-100

K-104

E-103

K-105

E-102

E-105

MeOH Synthesis

K-102

E-104

V-100FH-100

Reformer

E-101

E-108

V-101

TEE-100

MRU

ATU

E-100

MIX-100

Prereformer

K-100

MIX-101

E-107

32




T (F) 80.00 80.00 80.00 40.00 52.43 72.00 1100.00 500.00 1063.56 500.00 500.00 474.70 480.40 218.90 160.00 216.80 1063.00 520.00 637.19 520.00 1100.00 1100.00 637.19 90.00 110.00 280.00 90.00 110.00 160.00 160.00 160.00 160.00 160.00 160.00 160.00 695.41 90.00 110.00 280.00 891.86 30.00 206.00 609.27 90.00 110.00 500.00 520.00 185.00 -44.00 -44.00 30.00 30.00 30.00 206.00 206.30 206.30 30.00 30.00 30.00 206.30 174.60 238.50 206.30 174.60 238.50 160.00 238.50 192.80 174.60 174.60 174.60 192.80 195.80 174.60 175.00 90.00 110.00 120.00

P (psia) 400.00 400.00 400.00 400.00 400.00 400.00 30.00 400.00 30.00 400.00 400.00 400.00 400.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 70.00 70.00 30.00 70.00 70.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 180.00 70.00 70.00 180.00 1015.00 30.00 30.00 1015.00 70.00 70.00 1015.00 1015.00 30.00 63.00 63.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 24.50 29.00 30.00 24.50 29.00 30.00 29.00 29.00 24.50 24.50 24.50 29.00 50.00 24.50 100.00 70.00 70.00 100.00

Cp (BTU/lbm-F) 0.5409 0.3149 0.5410 0.8371 0.8380 0.5477 0.5762 0.6899 0.5731 0.6899 0.6805 0.5715 0.5793 0.5080 1.0010 0.7587 0.5731 0.4966 0.5835 0.4966 0.6116 0.6409 0.5835 0.9980 0.9982 0.5109 0.9980 0.9982 0.9287 1.0010 0.6724 1.0010 1.0010 1.0010 0.6724 0.7075 0.9980 0.9982 0.6829 0.7350 1.0530 0.9811 0.7096 0.9980 0.9982 0.7069 0.6401 1.0930 0.5590 0.5590 0.9733 1.3510 1.0530 0.9811 1.0580 1.0580 1.3510 1.3510 1.3510 1.0580 0.9565 1.0110 1.0580 0.9565 1.0110 1.0010 1.0110 1.0040 0.9565 0.9565 0.9565 1.0040 1.0040 0.9565 0.9558 0.9980 0.9982 0.9555

dh vap (BTU/lb) 446.0 162.8 226.5 587.0 596.3 207.6 1502.0 207.6 1502.0 207.6 780.6 1025.0 1455.0 1293.0 951.9 1455.0 1502.0 -786.0 1502.0 -786.0 1455.0 -1003.0 1502.0 908.1 908.1 -1014.0 908.1 908.1 1502.0 951.9 948.9 951.9 951.9 951.9 948.9 1850.0 908.1 908.1 977.0 1255.0 638.7 638.7 1425.0 908.1 908.1 1425.0 -800.9 -800.9 174.0 174.0 -543.9 1283.0 638.7 638.7 664.9 664.9 1283.0 1283.0 1283.0 664.9 482.1 954.5 664.9 482.1 954.5 951.9 954.5 954.1 482.1 482.1 482.0 954.1 928.8 482.0 423.6 908.1 908.1 423.6


P-101

E-109TEE-101

TEE-102

T-100

T-103

MIX-103

MIX-104

P-100

K-104

E-103

K-105

E-102

E-105

MeOH Synthesis

K-102

E-104

V-100FH-100

Reformer

E-101

E-108

V-101

TEE-100

MRU

ATU

E-100

MIX-100

Prereformer

K-100

MIX-101

E-107




T (F) 80.00 80.00 80.00 40.00 52.43 72.00 1100.00 500.00 1063.56 500.00 500.00 474.70 480.40 218.90 160.00 216.80 1063.00 520.00 637.19 520.00 1100.00 1100.00 637.19 90.00 110.00 280.00 90.00 110.00 160.00 160.00 160.00 160.00 160.00 160.00 160.00 695.41 90.00 110.00 280.00 891.86 30.00 206.00 609.27 90.00 110.00 500.00 520.00 185.00 -44.00 -44.00 30.00 30.00 30.00 206.00 206.30 206.30 30.00 30.00 30.00 206.30 174.60 238.50 206.30 174.60 238.50 160.00 238.50 192.80 174.60 174.60 174.60 192.80 195.80 174.60 175.00 90.00 110.00 120.00

P (psia) 400.00 400.00 400.00 400.00 400.00 400.00 30.00 400.00 30.00 400.00 400.00 400.00 400.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 70.00 70.00 30.00 70.00 70.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 180.00 70.00 70.00 180.00 1015.00 30.00 30.00 1015.00 70.00 70.00 1015.00 1015.00 30.00 63.00 63.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 24.50 29.00 30.00 24.50 29.00 30.00 29.00 29.00 24.50 24.50 24.50 29.00 50.00 24.50 100.00 70.00 70.00 100.00

Cp (BTU/lbm-F) 0.5409 0.3149 0.5410 0.8371 0.8380 0.5477 0.5762 0.6899 0.5731 0.6899 0.6805 0.5715 0.5793 0.5080 1.0010 0.7587 0.5731 0.4966 0.5835 0.4966 0.6116 0.6409 0.5835 0.9980 0.9982 0.5109 0.9980 0.9982 0.9287 1.0010 0.6724 1.0010 1.0010 1.0010 0.6724 0.7075 0.9980 0.9982 0.6829 0.7350 1.0530 0.9811 0.7096 0.9980 0.9982 0.7069 0.6401 1.0930 0.5590 0.5590 0.9733 1.3510 1.0530 0.9811 1.0580 1.0580 1.3510 1.3510 1.3510 1.0580 0.9565 1.0110 1.0580 0.9565 1.0110 1.0010 1.0110 1.0040 0.9565 0.9565 0.9565 1.0040 1.0040 0.9565 0.9558 0.9980 0.9982 0.9555

dh vap (BTU/lb) 446.0 162.8 226.5 587.0 596.3 207.6 1502.0 207.6 1502.0 207.6 780.6 1025.0 1455.0 1293.0 951.9 1455.0 1502.0 -786.0 1502.0 -786.0 1455.0 -1003.0 1502.0 908.1 908.1 -1014.0 908.1 908.1 1502.0 951.9 948.9 951.9 951.9 951.9 948.9 1850.0 908.1 908.1 977.0 1255.0 638.7 638.7 1425.0 908.1 908.1 1425.0 -800.9 -800.9 174.0 174.0 -543.9 1283.0 638.7 638.7 664.9 664.9 1283.0 1283.0 1283.0 664.9 482.1 954.5 664.9 482.1 954.5 951.9 954.5 954.1 482.1 482.1 482.0 954.1 928.8 482.0 423.6 908.1 908.1 423.6


P-101

E-109TEE-101

TEE-102

T-100

T-103

MIX-103

MIX-104

P-100

K-104

E-103

K-105

E-102

E-105

MeOH Synthesis

K-102

E-104

V-100FH-100

Reformer

E-101

E-108

V-101

TEE-100

MRU

ATU

E-100

MIX-100

Prereformer

K-100

MIX-101

E-107




T (F) 80.00 80.00 80.00 40.00 52.43 72.00 1100.00 500.00 1063.56 500.00 500.00 474.70 480.40 218.90 160.00 216.80 1063.00 520.00 637.19 520.00 1100.00 1100.00 637.19 90.00 110.00 280.00 90.00 110.00 160.00 160.00 160.00 160.00 160.00 160.00 160.00 695.41 90.00 110.00 280.00 891.86 30.00 206.00 609.27 90.00 110.00 500.00 520.00 185.00 -44.00 -44.00 30.00 30.00 30.00 206.00 206.30 206.30 30.00 30.00 30.00 206.30 174.60 238.50 206.30 174.60 238.50 160.00 238.50 192.80 174.60 174.60 174.60 192.80 195.80 174.60 175.00 90.00 110.00 120.00

P (psia) 400.00 400.00 400.00 400.00 400.00 400.00 30.00 400.00 30.00 400.00 400.00 400.00 400.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 70.00 70.00 30.00 70.00 70.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 180.00 70.00 70.00 180.00 1015.00 30.00 30.00 1015.00 70.00 70.00 1015.00 1015.00 30.00 63.00 63.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 24.50 29.00 30.00 24.50 29.00 30.00 29.00 29.00 24.50 24.50 24.50 29.00 50.00 24.50 100.00 70.00 70.00 100.00

Cp (BTU/lbm-F) 0.5409 0.3149 0.5410 0.8371 0.8380 0.5477 0.5762 0.6899 0.5731 0.6899 0.6805 0.5715 0.5793 0.5080 1.0010 0.7587 0.5731 0.4966 0.5835 0.4966 0.6116 0.6409 0.5835 0.9980 0.9982 0.5109 0.9980 0.9982 0.9287 1.0010 0.6724 1.0010 1.0010 1.0010 0.6724 0.7075 0.9980 0.9982 0.6829 0.7350 1.0530 0.9811 0.7096 0.9980 0.9982 0.7069 0.6401 1.0930 0.5590 0.5590 0.9733 1.3510 1.0530 0.9811 1.0580 1.0580 1.3510 1.3510 1.3510 1.0580 0.9565 1.0110 1.0580 0.9565 1.0110 1.0010 1.0110 1.0040 0.9565 0.9565 0.9565 1.0040 1.0040 0.9565 0.9558 0.9980 0.9982 0.9555

dh vap (BTU/lb) 446.0 162.8 226.5 587.0 596.3 207.6 1502.0 207.6 1502.0 207.6 780.6 1025.0 1455.0 1293.0 951.9 1455.0 1502.0 -786.0 1502.0 -786.0 1455.0 -1003.0 1502.0 908.1 908.1 -1014.0 908.1 908.1 1502.0 951.9 948.9 951.9 951.9 951.9 948.9 1850.0 908.1 908.1 977.0 1255.0 638.7 638.7 1425.0 908.1 908.1 1425.0 -800.9 -800.9 174.0 174.0 -543.9 1283.0 638.7 638.7 664.9 664.9 1283.0 1283.0 1283.0 664.9 482.1 954.5 664.9 482.1 954.5 951.9 954.5 954.1 482.1 482.1 482.0 954.1 928.8 482.0 423.6 908.1 908.1 423.6


P-101

E-109TEE-101

TEE-102

T-100

T-103

MIX-103

MIX-104

P-100

K-104

E-103

K-105

E-102

E-105

MeOH Synthesis

K-102

E-104

V-100FH-100

Reformer

E-101

E-108

V-101

TEE-100

MRU

ATU

E-100

MIX-100

Prereformer

K-100

MIX-101

E-107




T (F) 80.00 80.00 80.00 40.00 52.43 72.00 1100.00 500.00 1063.56 500.00 500.00 474.70 480.40 218.90 160.00 216.80 1063.00 520.00 637.19 520.00 1100.00 1100.00 637.19 90.00 110.00 280.00 90.00 110.00 160.00 160.00 160.00 160.00 160.00 160.00 160.00 695.41 90.00 110.00 280.00 891.86 30.00 206.00 609.27 90.00 110.00 500.00 520.00 185.00 -44.00 -44.00 30.00 30.00 30.00 206.00 206.30 206.30 30.00 30.00 30.00 206.30 174.60 238.50 206.30 174.60 238.50 160.00 238.50 192.80 174.60 174.60 174.60 192.80 195.80 174.60 175.00 90.00 110.00 120.00

P (psia) 400.00 400.00 400.00 400.00 400.00 400.00 30.00 400.00 30.00 400.00 400.00 400.00 400.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 70.00 70.00 30.00 70.00 70.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 180.00 70.00 70.00 180.00 1015.00 30.00 30.00 1015.00 70.00 70.00 1015.00 1015.00 30.00 63.00 63.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 30.00 24.50 29.00 30.00 24.50 29.00 30.00 29.00 29.00 24.50 24.50 24.50 29.00 50.00 24.50 100.00 70.00 70.00 100.00

Cp (BTU/lbm-F) 0.5409 0.3149 0.5410 0.8371 0.8380 0.5477 0.5762 0.6899 0.5731 0.6899 0.6805 0.5715 0.5793 0.5080 1.0010 0.7587 0.5731 0.4966 0.5835 0.4966 0.6116 0.6409 0.5835 0.9980 0.9982 0.5109 0.9980 0.9982 0.9287 1.0010 0.6724 1.0010 1.0010 1.0010 0.6724 0.7075 0.9980 0.9982 0.6829 0.7350 1.0530 0.9811 0.7096 0.9980 0.9982 0.7069 0.6401 1.0930 0.5590 0.5590 0.9733 1.3510 1.0530 0.9811 1.0580 1.0580 1.3510 1.3510 1.3510 1.0580 0.9565 1.0110 1.0580 0.9565 1.0110 1.0010 1.0110 1.0040 0.9565 0.9565 0.9565 1.0040 1.0040 0.9565 0.9558 0.9980 0.9982 0.9555

dh vap (BTU/lb) 446.0 162.8 226.5 587.0 596.3 207.6 1502.0 207.6 1502.0 207.6 780.6 1025.0 1455.0 1293.0 951.9 1455.0 1502.0 -786.0 1502.0 -786.0 1455.0 -1003.0 1502.0 908.1 908.1 -1014.0 908.1 908.1 1502.0 951.9 948.9 951.9 951.9 951.9 948.9 1850.0 908.1 908.1 977.0 1255.0 638.7 638.7 1425.0 908.1 908.1 1425.0 -800.9 -800.9 174.0 174.0 -543.9 1283.0 638.7 638.7 664.9 664.9 1283.0 1283.0 1283.0 664.9 482.1 954.5 664.9 482.1 954.5 951.9 954.5 954.1 482.1 482.1 482.0 954.1 928.8 482.0 423.6 908.1 908.1 423.6


P-101

E-109TEE-101

TEE-102

T-100

T-103

MIX-103

MIX-104

P-100

K-104

E-103

K-105

E-102

E-105

MeOH Synthesis

K-102

E-104

V-100FH-100

Reformer

E-101

E-108

V-101

TEE-100

MRU

ATU

E-100

MIX-100

Prereformer

K-100

MIX-101

E-107

33

Equipment Sizing

An example specification sheet for each piece of equipment and excel worksheets used for sizing

are located within Appendix B.

Reactors

The reactors are modeled as packed bed reactors. The working specification sheets are found in

the appendix for this equipment. The following equations illustrate how the steam reformer reactor

was sized. All other reactors followed this model. Catalyst data was given in table 2.

Table 2: Catalyst Data

Reactor Cat Φ

(ft)

β

(ft)

Cost

$/lb

Cat

dens

(lb/ft3)

SV

(ft3/hr

fed)/

(ft3 cat)

SV

(lbmol/hr

fed)/

(ft3cat)

SV

(lbs/hr

fed)/

(lbcat)

Steam

Reformer

57-4

GQ

4.26

E-2

5.58E-2 $6.71 53 13097 14.29 4.58

MeOH 51-7 1.77

E-2

1.71E-2 $16.19 78 10000 717.15 257.4

With the given space velocities (SV), the volume of catalyst needed can be calculated. The space

velocity was chosen in terms of molar flow. This is shown in the Equation below.

𝑉𝑐𝑎𝑡 =𝑀𝑜𝑙𝑎𝑟 𝐹𝑙𝑜𝑤 (

𝑙𝑏𝑚𝑜𝑙ℎ

)

𝑆𝑉 (

𝑙𝑏𝑚𝑜𝑙ℎ − 𝑓𝑒𝑑𝑓𝑡3𝑐𝑎𝑡

)

= 𝑓𝑡3𝑐𝑎𝑡𝑎𝑦𝑙𝑠𝑡

𝑉𝑐𝑎𝑡 =3.33𝐸4

𝑙𝑏𝑚𝑜𝑙ℎ

14.29 (

𝑙𝑏𝑚𝑜𝑙ℎ − 𝑓𝑒𝑑𝑓𝑡3𝑐𝑎𝑡

)

= 2330 𝑓𝑡3𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡

With the volume of catalyst, the length and diameter of the packed bed can be found, by

rearranging the volume of a cylinder equation. The L/D ratio is set to 3, to ensure a proper size

ratio for heating. This is shown in equations 2 and 3. Equation 4 represents how the diameter of

the reactor is found.

𝑉𝑐𝑎𝑡 = (𝜋

4) 𝐷2𝐿

34

𝐿

𝐷= 3 , 𝐿 = 3𝐷

𝐷 = √4𝑉

3𝜋

3

𝐷𝑟𝑒𝑎𝑐𝑡𝑜𝑟 = √4 ∗ 2330 𝑓𝑡3

3 ∗ 𝜋

3

= 9.96 ft

𝐿𝑏𝑒𝑑 = 3 ∗ D = 29.89 𝑓𝑡 With the size of the reactor bed determined, the pressure drop of the bed must be calculated to

ensure it is less than 25 psi. This is done by the Ergun Equation, shown in equation 5.

∆𝑃

𝐿𝑏𝑒𝑑=

4.66 𝜇 𝑉𝑚 (1 − 𝜀)2

𝑑𝑝2 𝜀3(144)

+0.0544 𝜌 𝑉𝑚

2 (1 − 𝜀)

𝑑𝑝 𝜀3(144)

Where,

ΔP = Pressure drop through the packed bed, psi.

ε = bed void fraction (30-40%)

ρ = average density, lb/ft3

μ = average viscosity, lb/ft-sec

Vm = superficial fluid velocity, ft/sec

L = Bed length, ft

dP = spherical equivalent particle diameter, ft.

𝑑𝑝 =3 𝜙 𝛽

2 𝛽+ 𝜙

Where,

φ = catalyst particle diameter, ft

β = catalyst particle length, ft

Δ𝑃 = 29.89 𝑓𝑡 ∗ (4.66 ∗ 1.77𝐸 − 5 ∗ 2.46 ∗ (1 − 0.35)2

(0.0462)2 0.353(144)+

0.0544(3.18𝐸 − 2)(2.46)2 (1 − 0.35)

0.0462 (0.35)3(144) )

Δ𝑃 = 2.43 𝑝𝑠𝑖

Because the pressure drop is less than 25 psi, this design is feasible and the full dimensions of the

reactor can be modeled. The full dimensions include spacing for an inlet distributor for the catalyst,

catalyst supports, and space for catalyst loading and removal. The space for catalyst loading is

defined as 50% of the diameter, this is required at both the top and bottom of the catalyst bed. The

following sizes are shown in table 3 below.

35

Table 3: SMR Full Dimensions

Full dimensions Size (ft)

Catalyst inlet distributor 0.5

Catalyst support 0.5

Catalyst loading space 4.98

Catalyst removal space 4.98

The results of the reactor sizing for all reactors is found in table 4. The overall heating or cooling

is needed to ensure the temperatures at the outlet are maintained. The volume of catalyst inherently

contains the void spacing for the reactor. The length of the reactor is the total size, where the length

of the bed is where the catalyst resides in the unit. The SMR needs to be heated to ensure it stays

isothermal, while the prereformer and the MeOH synthesis reactors need to be cooled to maintain

the temperature conditions chosen.

Table 4: Reactor Sizing Information

Reactor Volume

Cat (ft3)

Diameter

(ft)

Length

of Bed

(ft)

Pressure

drop (psi)

Length

of

Reactor

(ft)

Catalyst

Cost

Heating

Values

(MMBTU/hr)

Prereformer 772 7 21 0.02 29 $ 274,500 -0.303

SMR 2330 10 30 2.45 41 $ 829,000 215

MeOH

Synthesis 6.81 2 4 1.48 7 $ 8,599 -95.8

The price and size of catalyst needed for the Methanol synthesis reactor is far smaller than the

other two reactors in comparison. This is due to the reduced amount of molar flow going through

this reactor, type of catalyst, and space velocity for the catalyst. The separation unit V-101 recycles

a large amount of the water in the system back to the SMR. Thus, the overall flow downstream is

decreased. Also, this catalyst is in a pellet form, in comparison to the cylindrical shape of the

reforming catalyst. Finally, the GHSV of the 51-7 catalyst is 51 times greater than the 54-7 catalyst.

Separators

Vapor/Liquid Separators:

The separators used in the plant were designed as horizontal vessels using the droplet settling

theory. In gravity settling the dispersed drops/bubbles will settle at a velocity that is determined

by equating gravity force on the drop/bubble to the drag coefficient force. The separations in this

case consisted of a vapor/liquid mixture, thus equations for liquid drops in a gas phase were used

36

to size the separators (obtained from document on separator sizing, provided by instructor). For

liquid drops in gas phase, the settling theory yields the equation shown below:

𝑑2 = 5054𝑇𝑍𝑄𝑔

P[(|

ρ𝑔

ρ𝑙

− ρ𝑔

|)𝐶𝐷

𝑑𝑚]

12

Where:

• d = vessel internal diameter, in

• T = operating temperature, °R

• Z = gas compressibility

• Qg = gas flow rate, MMSCFD

• P = operating pressure, psia

• ρg = gas density, lbm/ft3

• ρl = liquid density, lbm/ft3

• dm = bubble or drop diameter, µm (300 µm - assumption)

• CD = drag coefficient (calculated using the equations shown below)

The physical properties needed to solve for the diameter were obtained from the HYSYS model.

The drag coefficient was calculated in order to solve for the vessel internal diameter (d). The

following calculations show how CD was obtained:

The Archimedes (Ar) number was calculated using the equation shown below. The equation was

derived from the balance of drag and buoyancy force:

𝐴𝑟 = 𝑑𝑣

3|ρ𝑐

− ρ𝑑|ρ

𝑐𝑔

µ𝑐

Where:

• dv = dispersed phase drop/bubble size, cm

• ρc = continuous phase density, g/cm3

• ρd = dispersed phase density, g/cm3

• g = gravitational constant, 981 cm/s2

• µc = continuous phase viscocity, g/(cm/sec) = poise

The drag coefficient is a function of the Reynolds number. Thus the Ar number was then used to

calculate the Reynolds number, which was then used to calculate the drag coefficient as follows:

Re = [√19.075 + 2.129√𝐴𝑟 − 4.3675

Re was used to find Cd:

𝐶𝐷 = (0.5423 +4.737

𝑅𝑒12

)2 for Re < 1 and 𝐶𝐷 =24

𝑅𝑒 𝑓𝑜𝑟 𝑅𝑒 > 1

The drag coefficient was used to calculate the inner vessel diameter. The ratio of column length

to diameter range from 3-6. Using this ratio the column length was calculated for both, L/d = 3

and L/d = 6. The length and diameter allowed for volume calculation. The following table shows

37

the results. It was decided to use the design with the smallest diameter to reduce its cost. The

following table summarizes the results for the sizing of the separators.

Table 5: Separator Sizing Summary

Settling

Velocity

(ft/s)

Actual

Diameter

(d) – ft

L (ft)

L/d

Liquid

Hold-Up

Demister

Height

(ft)

Head

Space

(ft)

V-101 8.14 10.6 31.7 3 25 min 0.5 1.0

V-100 9.71 7.1 21.2 3 10 min 0.5 1.0


The Mercury Removal Unit (MRU) was sized based on mercury (Hg) absorption to activated

carbon. The first step in sizing this unit was to calculate the mass of Hg that could be removed

over the period of time that the absorbent operates. The calculation in this case was based on a 7

day run time. It was determined to have tree units, one in operation and two for backup during

maintenance or if the operating unit fails. The Hg capacity in activated carbon is about 0.3,

which was used to calculate the mass of activated carbon. The volume of activated carbon was

then calculated using the mass and its density. It was assumed that the volume of activated

carbon occupied about 50% of the total volume and that the length/diameter ratio was 3.0. Using

this assumptions, L and D were found as it is shown below.

Figure 5: Mercury Removal Unit Sizing Sheet

1 3 2In

Vapor Phase Fraction 1.0 1 0

Temperature (°F) 80 80 80

Pressure (psia) 400 400 400

Mass Flow (lb/hr) 49269 49258 11

Hg mass Absorbed (lb/hr) 11

Mass Flow of Hg x run time of absorbents (lb) 1849.43

Mercury Density (lb/ft^3) 843.60

Mass of Activated Carbon (lb) 6164.75

Activated Carbon Density (lb/ft^3) 125

Hg capacity in activated carbon 0.3

Volume of Activated Carbon (Ft^3) 49.32

L (Ft) 10.42

D (Ft) 3.47

L/D 3.00

Volume of vessel (Ft^3) 98.65


outStreams

Mercury Removal Unit (MRU) - Sizing Calculations

Hg absorption based on 7 days (ONLY BASED ON 1 COLUMN)

38


The amine treatment unit was designed as a one-pass MDEA system where most of the H2S and

some CO2 is removed. In this unit, the MDEA can be regenerated but it was determined that this

could be considered outer battery limit (OBL). In the ATU, regeneration was not taken into

account. The sulfur content in the inlet stream (stream 3) was determined using the gas flow and

% of H2S. The gas velocity can be used to find the cross sectional area of the column and thus

the diameter. The equation used to calculate the gas velocity is shown below:

𝑉𝑔𝑎𝑠 = 0.25 ∗ [(𝜌𝑀𝐷𝐸𝐴 − 𝜌𝑔𝑎𝑠)/𝜌𝑔𝑎𝑠]12

An assumption of length to diameter of 6 was assumed in order to calculate the length and

volume of the column unit. The table below shows a summary of the results:

Figure 6: Amine Treatment Tower Sizing Sheet

Heat Exchangers

The sizing of these components was governed by the desired temperatures of the process stream,

and the utility chosen. The types of utilities chosen were cooling water, other process streams, fuel,

and refrigerant. The most cost efficient method is utilizing the process streams, because this is free

energy that is being used. All of the heat exchangers in the system are either countercurrent shell

and tube, a fired heater, or a refrigeration unit. The first step in sizing a shell and tube heat

exchanger is to determine the process stream temperature desired and the utility you are going to

use. The following equations are the governing equations for sizing the heat exchangers. The duty,

Q, is calculated from the change of temperatures of the process stream.

𝑄 = 𝑚 𝐶𝑝,𝑎𝑣𝑔 ∆𝑇

Streams

3 4 6

Vapor Phase Fraction 1.0 0.0 1.0

Temperature (F) 80 40 72


Mass Flow (lb/hr) 49258 124245 47855

Gas Flow (MMSCFD) 24.95 Sulfur Content 31557.31 lb S/day

Actual Gas Flow 618.5 CFM Circulation rate, MDEA 93.56 gpm

10.31 ft^3/s Gas Velocity 11.50 ft/s

% H2S in feed stream 1.5 L/D = 6 6

% CO2 in feed stream 1.5 Inner Diameter (D) 1.19 ft

Temperature 80 F Length (L) 7.15 ft

Pressure 385.3 psig Volume 7.98 ft^3

K (MMSCFD) 1.25 Retention Time 10 minutes

Gas Density 1.328 lbm/ft3

MDEA Density 62.428 lbm/ft3


Properties Sizing and Calculations

In Out

5

0.0

52

400

125648

39

Where,

m = mass flow rate of utility, lb/hr

Cp, avg = average heat capacity of the inlet and outlet process streams, BTU/lb-F

∆T = change in process stream temperature, F

With the duty needed, the area of the heat exchanger can be determined by rearranging the

following equation.

𝑄 = 𝑈𝐴 ∆𝑇𝐿𝑀

Where,

U = heat transfer coefficient, BTU/F –ft2

∆TLM = the log mean temperature difference

∆𝑇𝐿𝑀 = ∆𝑇𝐴 − ∆𝑇𝐵

𝐿𝑛(∆𝑇𝐴

∆𝑇𝐵)

Where, ∆TA is the largest temperature difference between the two ends of the heat

exchanger.

𝐴 =𝑄

𝑈 ∆𝑇𝐿𝑀∗ 𝜎

where, σ is a multiplicative safety factor, σ = 1.15

The following equations illustrate the sizing for E -108, as an example calculation.

𝑇𝑖𝑛,𝑝𝑟𝑜𝑐𝑒𝑠𝑠 = 280 𝐹, 𝑇𝑜𝑢𝑡,𝑝𝑟𝑜𝑐𝑒𝑠𝑠 = 160 𝐹

𝑇𝑖𝑛,𝑢𝑡𝑖𝑙𝑖𝑡𝑦 = 90 𝐹, 𝑇𝑜𝑢𝑡,𝑢𝑡𝑖𝑙𝑖𝑡𝑦 = 110 𝐹

𝑈 = 50, 𝜎 = 1.15

𝑄 = 597840 (𝑙𝑏

ℎ𝑟) (

0.5109 + 0.92866

2

𝐵𝑇𝑈

𝑙𝑏 − 𝐹) ∗ (280 − 160)𝐹 = 32.684 𝑀𝑀𝐵𝑇𝑈

∆𝑇𝐿𝑀 = (280 − 110) − (160 − 90)

𝐿𝑛 ((280 − 110)

160 − 90 ) = 113 𝐹

𝐴 = (32.684 𝑀𝑀𝐵𝑇𝑈

50 (𝐵𝑇𝑈

𝐹 − 𝑓𝑡2 −) ∗ 113𝐹 ∗ 106) ∗ 1.15 = 6670 𝑓𝑡2

40

Table 5: Summary of Heat Exchanger Sizes

Equipment

name

Process in

(F)

Process out

(F)

Utility Type Utility in

(F)

Utility out

(F)

Area

(ft2)

E -100 72 500 Process 1100 1063 748

E -107 216 520 Process 1063 636 14077

E -101 636 280 CW 90 110 7757

E -108 280 160 CW 90 110 6670

E -103 695 280 CW 90 110 2564

E -102 30 206 Process 892 607 1003

E -105 607 500 CW 90 110 516

E-109 172 120 CW 90 110 1181

The fired heater is fueled by using a process waste stream. This contains methane, which can be

burned to produce enough heat to get the process stream of interest to the desired temperature. The

following equation was used to determine the area needed for the fired heater.

𝑄 = 𝐹𝜎𝛼𝐴𝑐(𝑇𝑡4 − 𝑇𝑔

4)

Where,

F = 0.702, is the exchange factor of the heater, (from Perry’s handbook.)

σ = 1.71 E-9 (BTU/(hr-ft2-R4), a given value.

α = 60%, efficiency of the combustion

Q = Duty, BTU/hr

TT = temperature of the wall of the fired heater, R

Tg = temperature of the gas exhaust, R.

The following calculation represents the sizing for the fired heater.

𝑇𝑡 = 2660 𝑅, 𝑇𝑔 = 1760 𝑅, 𝑄 = 183𝑀𝑀𝐵𝑇𝑈

𝐻𝑅

𝐴𝑐 =183

0.702(0.60) ∗ (1.71𝐸 − 9) ∗ 106 ∗ (26604 − 17604)= 6275 𝑓𝑡2

𝐴 = 1.1 ∗ 6275 𝑓𝑡2 = 6903 𝑓𝑡2

41

Table 6: Summary of Fired Heater Size

Equipment

name

Process in

(F)

Process out

(F)

Utility

Type

Utility in

(F)

Utility out

(F)

Area

(ft2)

FH -100 520 1100 Process 30 1300 6903

The final process temperature control system is the refrigeration unit. The overall refrigeration unit

is an OBL process, but the cost must be approximated and the heat exchange unit must be sized.

The following figure was used to determine the overall cost of the refrigeration. Propane was

chosen as the refrigerant due to the temperature desired. 70 MMBTU is required for this system,

according to the Bluebook. Here, the yellow dot signifies where our refrigeration unit will be

operating.

Table 7: Summary of Refrigeration Unit Cost

Equipment

name

Process

in (F)

Process

out (F)

Utility

Type

Utility

in

(F)

Utility

out (F)

Area

(ft2)

Duty

(MMBTU/

hr)

Cost

($/MMBTU)

E -104 197.2 30 Refrig. -44 -44 4557 70 35

0

10

20

30

40

50

60

70

80

90

100

-150 -100 -50 0 50 100

Cost of

Refr

ig,

$/M

MB

TU

Temperature, F

Refrigeration Cost

Total Cost

Figure 7: Refrigeration Cost

42

Compressors, Expanders, and Pumps

The sizing of the compressors and expanders was governed by the stream flow rates and

required changes in pressure. It is important to size compressors and expanders for the correct

anticipated load as they are an important part of the energy balance within the process.

Compressors are very energy intensive and should be sized properly to avoid wasted input.

Expanders have the potential to capture energy and should be sized properly to ensure capture of

energy for reuse and ability to handle the stream. Both compressors and expanders must be sized

large enough to handle the stream, yet as small as possible so as not to incur extra cost.

The governing equations for sizing these units are as follows:

𝑍𝑎𝑣𝑔 =𝑍1+𝑍2

2 (1)

𝑇2

𝑇1= (

𝑃2

𝑃1)

𝑛−1

𝑛 (2)

𝑛−1

𝑛=

(𝑘−1)

(𝑘∗𝜂𝑝) (3)

𝑘 =𝑐𝑝

𝑐𝑣 (4)

𝐻𝑃 =𝑛

(𝑛−1)∗

𝑍𝑎𝑣𝑔∗𝑅∗𝑇1

𝑀𝑊∗ ((

𝑃2

𝑃1)

(𝑛−1)

𝑛− 1) (5)

𝐵𝐻𝑃 = 𝑤∗𝐻𝑃

33,000∗ 𝜂𝑝 (6)

Values for the variables in these equations were drawn from the HYSYS model.

For compressor K-104, an example calculation can be found below:

𝑍𝑎𝑣𝑔 =1.000+1.002

2= 1.0001 (1)

𝑇2 = 619.67 ∗ (180

30)

.337472

= 1155.16 𝑅 (2)

𝑛−1

𝑛=

(1.368−1)

(1.368∗0.79712)= .3476 (3)

𝑘 =𝑐𝑝

𝑐𝑣 = 1.385 (4)

𝐻𝑃 =1

.3476∗

1.0001∗1545.4∗619.67

11.76∗ ((

180

30)

0.3476

− 1) = 222933.1 (5)

𝑃𝑜𝑤𝑒𝑟 = 2231.52∗222933.1

33,000∗ 0.79972= 18850.52 𝐵𝐻𝑃 (6)

Using these calculations for each compressor and expander followed by unit conversions, a final

power usage in units of kW was determined. The compressor names, pressures, and power usages

can be found in the table below. The expander names, pressures, and power outputs can also be

found in the table below.

43

Table 8: Expander and Compressor Summary:

Unit P1 (psia) P2 (psia) Power (kW)

K-100 (exp) 400 30 -13837.7

K-102 (exp) 1015 30 -13217.8

K-104 30 180 14056.84

K-105 180 1015 16111.87

The sizing of the pumps was governed mainly by the flowrate and density of the fluid traveling

through the pump. The NPSH and efficiency were determined by using the design flow in

combination with the two following charts from the Blue Book and several equations. Pumps are

relatively inexpensive, so backups were easily purchased for each of the pumps in the process. It

is crucial that pumps are sized largely enough to handle the large flowrate of the process.

Figure 8: NPSH determination chart for pump sizing (P-100 shown)

44

Figure 9: Efficiency determination chart for pump sizing (P-100 shown)

The equations for sizing the pumps include the following:

𝐷𝑒𝑠𝑖𝑔𝑛 𝐹𝑙𝑜𝑤 = 𝑃𝑟𝑜𝑐𝑒𝑠𝑠 𝐹𝑙𝑜𝑤 ∗ 1.25 (1)

𝑃𝑠𝑐𝑡𝑛 = 𝑃𝑠𝑜𝑢𝑟𝑐𝑒 +𝑁𝑃𝑆𝐻∗𝜌

144𝑖𝑛2

𝑓𝑡2

− ∆𝑃𝑠𝑐𝑡𝑛 𝑙𝑖𝑛𝑒 − ∆𝑃𝑒𝑛𝑡𝑟𝑎𝑛𝑐𝑒,𝑒𝑥𝑖𝑡 𝑙𝑜𝑠𝑠𝑒𝑠 (2)

𝑃𝑑𝑠𝑐ℎ = 𝑃𝑑𝑒𝑠𝑡𝑖𝑛𝑎𝑡𝑖𝑜𝑛 +𝐻𝑑𝑠𝑐ℎ∗𝜌

144𝑖𝑛2

𝑓𝑡2

+ ∆𝑃𝐻𝑋 + ∆𝑃𝐹𝐸 + ∆𝑃𝐶𝑉 + ∆𝑃𝐿𝑖𝑛𝑒 + ∆𝑃𝑒𝑛𝑡𝑟𝑎𝑛𝑐𝑒,𝑒𝑥𝑖𝑡 𝑙𝑜𝑠𝑠𝑒𝑠 (3)

𝐻𝑒𝑎𝑑 (𝑓𝑡) = (𝑃2 − 𝑃1)𝑝𝑠𝑖𝑎 ∗ 144𝑖𝑛2

𝑓𝑡2 ∗ 𝜌𝑓𝑡3

𝑙𝑏𝑚 (4)

𝐵𝐻𝑃 = 𝑤∗𝐻𝑃

33,000∗ 𝜂𝑝 (5)

45

𝑀𝐻𝑃 =𝐵𝐻𝑃

𝜂𝑝 (6)

An example calculation for pump P-100 can be found below.

272.63 𝐺𝑃𝑀 = 218.10 𝐺𝑃𝑀 ∗ 1.25 (1)

31.5 𝑝𝑠𝑖𝑎 = 29 𝑝𝑠𝑖𝑎 +8.5∗59.7

144𝑖𝑛2

𝑓𝑡2

− 1 𝑝𝑠𝑖𝑎 − 0 𝑝𝑠𝑖𝑎 (2)

79.14 𝑝𝑠𝑖𝑎 = 50 +10∗59.7

144𝑖𝑛2

𝑓𝑡2

+ 0 𝑝𝑠𝑖𝑎 + 3 𝑝𝑠𝑖𝑎 + 10 𝑝𝑠𝑖𝑎 + 12 𝑝𝑠𝑖𝑎 + 0 𝑝𝑠𝑖𝑎 (3)

114.87 𝑓𝑡 = (79.14 − 31.5)𝑝𝑠𝑖𝑎 ∗ 144𝑖𝑛2

𝑓𝑡2∗ 59.7

𝑓𝑡3

𝑙𝑏𝑚 (4)

11.47 = 2175.79∗114.87

33,000∗ 0.66 (5)

12.75 =11.47

0.9 (6)

Using these equations, taking care to balance the pressure on either side of the pump, these

pieces of equipment can be properly sized for the process. The pumps will maintain appropriate

pressure of streams as they exit the distillation towers and move on to product storage or waste

water treatment.

Distillation Towers

In order to carry out the separation required to obtain the maximum methanol product purity, the

design modeled by HYSYS consisted of two identical distillation columns. The feed entering

each distillation column is identical in composition and flow rate. As a result, the product leaving

each distillation column is also identical in composition and flow rate. The distillation columns

separate the methanol product from water. The purity attained is 99.89 % pure methanol with a

flow rate of approximately 37,000 lbm/hr per distillation column. The distillation columns have

40 trays each in addition to a condenser and a reboiler bringing the total number of stages to 42

theoretical stages per distillation column.

The temperature of the feed was chosen to be equal to the temperature of the feed tray in the

distillation column. This was done to limit the flooding of the distillation column. As for the

pressure selection, the pressure of the feed was selected to be 15 psig in order to reduce the

chance of a back pressurization occurring within the column during startup. Initially, the design

only consisted of one distillation column. However, this design failed for several reasons.

Economically, to achieve the necessary separation would require approximately 100 plates. As a

result, the distillation column would not follow the design specifications. Additionally, the vapor

flow rate was too great which did not permit us to use the Glitsch method to estimate the

required diameters for the single-pass, double-pass, and four-pass systems. Another issue that

was encountered with this design was the very high reflux ratio. A high reflux ratio requires

greater energy and size of the overall column; thus, increasing the cost of the column. We elected

to split the feed into two identical distillation columns in order to reduce the flow rate and the

reflux ratio. Based on the design parameters, we were able to achieve a reflux ratio of 4.1. As a

46

result, we were able to minimize the cost and the number of stages required to carry out this

separation. Another important aspect in designing the distillation column is the placement of the

feed relative to the distillation column height. We decided to input the feed streams into the

distillation column towards the center in order to maximize the separation efficiency. Another

significant feature of the distillation column design is the pressure drop per tray which we

designed for the column to have a 0.15 psi pressure drop per plate.

The set of equations below was used to enable utilizing use of the Glitsch method:

Volumetric vapor flow rate: 𝑉𝑙𝑜𝑎𝑑 = ACFS ∗ √ρ𝑣

(ρ𝑙−ρ𝑣) (1)

ACFS= Actual Cubic feet per Second;

Volumetric liquid flow rate: L𝑙𝑜𝑎𝑑 =𝑚𝑙

𝑆𝐺 (2)

Specific gravity: SG =ρ𝑝𝑟𝑜𝑐𝑒𝑠𝑠

ρ𝐻20 (3)

Liquid mass flow rate: 𝐿 = ρ𝑙 ∗ 𝑚𝑙 (4)

Flood Factor: 𝐹𝐹 =80

𝐹𝑙𝑜𝑜𝑑 (5)

Corrected vapor load: 𝑉𝑙𝑜𝑎𝑑,𝑐𝑜𝑟𝑒𝑐𝑡𝑒𝑑 = 𝑉𝑙𝑜𝑎𝑑∗𝐹𝐹

𝑆𝑦𝑠𝑡𝑒𝑚 𝑓𝑎𝑐𝑡𝑜𝑟∗𝑇𝑟𝑎𝑦 𝑠𝑝𝑎𝑐𝑖𝑛𝑔 𝑓𝑎𝑐𝑡𝑜𝑟 (6)

Corrected liquid load: 𝐿𝑙𝑜𝑎𝑑,𝑐𝑜𝑟𝑒𝑐𝑡𝑒𝑑 = 𝐿𝑙𝑜𝑎𝑑 ∗ 𝐹𝐹

𝑆𝑦𝑠𝑡𝑒𝑚 𝑓𝑎𝑐𝑡𝑜𝑟∗𝑇𝑟𝑎𝑦 𝑠𝑝𝑎𝑐𝑖𝑛𝑔 𝑓𝑎𝑐𝑡𝑜𝑟 (7)

𝑉𝑙𝑜𝑎𝑑,𝑐𝑜𝑟𝑒𝑐𝑡𝑒𝑑 =(21.67

𝑓𝑡3

𝑠 ) ∗ (1.14)

(0.75) ∗ (1.00)= 33

𝑓𝑡3

𝑠

𝐿𝑙𝑜𝑎𝑑,𝑐𝑜𝑟𝑒𝑐𝑡𝑒𝑑 =(407.59

𝑓𝑡3

𝑠 ) ∗ (1.14)

(0.75) ∗ (1.00)= 620

𝑓𝑡3

𝑠

By using these equations we were able to calculate the Vload, corrected to be 37.7 ft3/s and the Lload,

corrected to be 620 GPM. Based on these values, we were limited to using a four-pass system. We

obtained a tower diameter of approximately 12 ft as determined by the diagram and equations

below:

𝑉𝑙𝑜𝑎𝑑,4−𝑝𝑎𝑠𝑠 =𝑉𝑙𝑜𝑎𝑑.𝑐𝑜𝑟𝑟𝑒𝑐𝑡𝑒𝑑

2=

37.7

2= 18

𝑓𝑡3

𝑠

𝐺𝑃𝑀4−𝑝𝑎𝑠𝑠 =𝐺𝑃𝑀

2=

664.81

2= 332.4

By getting these values, a line can be drawn across the diagram below. The circle indicates

where the line crosses the two-pass diameter line. The final value of the diameter is found by

multiplying the two-pass diameter by a factor of the square root of 2;

𝐷𝑖𝑎𝑚𝑒𝑡𝑒𝑟 = 8′6" ∗ √2 = 12 𝑓𝑡

A table summarizing the size of the distillation columns can be found below.

47

Figure 10. Glitsch Method diagram

48

Table 9: Summarization of Distillation Tower Sizes

Specification Value

Column Diameter (ft) 12

Total Column Height (ft) 141

Reflux Drum Height (ft) 9

Pressure (Top/Btm), psia 24.5/29

Temperature (Top/Btm), ºF 175/228

Number of

Theoretical/Actual Stages

40/71

Reflux Drum Diameter (ft) 3

Reboiler Height (ft) 15

Reboiler Diameter (ft) 5.5

Theoretical Stages 42

Actual Trays 71

Condenser Duty

(MMBTU/hr) 16.6

Reboiler Duty

(MMBTU/hr) 46.2

Reflux Ratio 4.1

Feed Tray Location 28 (theoretical)

49

Utility Balance

The utility balance was constructed to display the yearly costs of each of the plants utility services.

This table shows the MMBtu/hr usage of utilities like steam and cooling water of various pressures

and sources respectively. For the fired heater, the needed energy is obtained from utilizing a waste

gas stream containing methane and hydrogen. This stream is split and also used to heat the

Reformer. However, more energy is required and the remaining energy needed is from purchased

fuel. The overall cost of the Refrigeration unit was estimated from Figure 1. The total cost to

operate the unit is approximately $32 per MMBTU/hr. The overall yearly cost sums to 28.99

MM$/yr. The main cost of this is from the fuel needed for the fired heater. Below is a summary of

the total costs, and the following page includes a summary of the utilities needed for each unit.

Table 9: Summary of Overall Costs for Utilities

Utility Total Cost (MM$/yr)

Fuel Gas 3.93

MPS 1.77

CW 0.69

Refrigeration 18.82

Power 3.78

Total 28.99

50

Figure 11: Microsoft Excel Spreadsheet for the Overall Utilities needed for the plant per unit.

kw

System Fuel Gas MPS CW Refrig Power3 Comments

Temperature, F 388 90--->110 -441

Pressure, psia 200 70 63

MRU

ATU

E-100 process

E-107 process

E-101 111.0

E-103 39.2

E-108 32.7

E-102 process

E-105 10.0

E-104 70.0 Refrigeration

E-109 4.3

Fire Heater (FH-100) Waste Stream as fuel

MIX-100

MIX-101

MIX-102

TEE-100 process (waste stream)

Expander: K-100 -9686.4 Power Generated

Compressor: K-104 14056.8

Compressor: K-105 16111.9

Expander: K-102 -9252.5 Power Generated

Pump: P-100 8.6

Pump: P-101 17.3

burn CH4

Prereformer 0.3 Exothermic

Steam Reformer 55 Endothermic

MeOH Synth 93.80 Exothermic

T-100 19 19.33

T-103 19 19.33

Total, MMBTU/hr 55 38 330 70 11256

Cost, $/MMBTU 2 8.50 5.50 0.25 32.00 0.04

Total Cost, MM$/yr 3.93 1.77 0.69 18.82 3.78

28.99

MMBTU/hr

Utility Balances

Total Utilities Cost, MM$/yr

51

Outside Battery Limits

Estimates were provided for the cost of site development, fuel utility storage, product storage, and

waste water treatment.

With a pilot plant on site, the site must be upgraded and developed to withhold all of the new

equipment and necessary utilities. Therefore, an OBL cost for site development must be

considered. Due to this pilot plant already being installed on the site, it was assumed that they were

using process water for their processes already. To obtain process water for things such as heat

exchangers, a DI water system would need to be on site to reduce the salt deposit formations when

creating the steam. However, due to the increasing the size of the operation, the size of the system

would need to be increased as well. This is a factor in site development.

An estimate was provided for the product storage. The Storage Tank Farm is located 1-mile away

from the plant. There is a 1.5 psi pressure drop across 100 ft of the line. Therefore, the product

must be pumped to approximately 85 psig to reach the tanks at atmospheric pressure. The product

streams must also be cooled to 120 °F, so the vapor pressure isn’t too high for the storage tank.

The estimation was given at capacity of 70000 lb/hr of methanol, where our capacity is

approximately 79100 lb/hr. Therefore, the estimate needed to be scaled both for our fiscal year and

for the capacity of our product. This value is shown in the following capital cost estimate. The

utilities followed this trend as well. Costs were provided for the OBL estimates, and these values

were sized for our necessary capacity and fiscal year. It is assumed that there is already a bio-

treatment area installed on the site for waste water. However, this value needs to be scaled as well,

to our capacity. The bio-treatment area for waste water requires at least a pressure of 50 psia.

Therefore, that is why a pump is installed after the distillation columns. The above mentioned

values are summarized in the capital cost estimation below.

52

Economic Analysis

Capital Estimate

Utilizing the Aspen Process Economic Analyzer (PEA), the cost of all the equipment were able to

be determined for our base case. The costs of the equipment were found using the sizes calculated

and the known variables for each unit. The cost for each unit was taken to be the “Total installed

and manpower cost” from the PEA software. Once all the estimates were determined, the

equipment capital cost for the plant was calculated to be approximately

34.026 MM$. The most expensive pieces of equipment were the compressors/expanders. These

are necessary in our process to meet the proper operating conditions of our process. An itemized

list of all the equipment is found in the appendix.

Table 10: Cost of Equipment

Unit Cost ($MM)

Reactors/Towers 4.218

Heat Exchangers 9.070

Compressors/Expanders/Pumps 19.200

Separators 1.539

Total 34.026

Additional capital costs would need to be factored into the calculation to determine the total capital

cost of the plant. The additional costs that need to be considered include outside battery limit items,

team costs, preoperational and start-up costs, and initial costs of catalysts. Because this is a new

plant, site development, product/raw material storage, and waste water treatment operations need

to be considered for OBL costs. An OBL bid summary for the necessary items was obtained. These

bids need to be scaled to the required capacity and escalated to 2017 dollar values. This was done

using an annual escalation rate of 3%, and a capital estimate exponent of 0.65. A sample

calculation of this is shown below. The total cost of these OBL estimates was approximately

63,620 M$.

𝑀𝑒𝑂𝐻 𝑃𝑟𝑜𝑑𝑢𝑐𝑡 𝑆𝑡𝑜𝑟𝑎𝑔𝑒 𝑆𝑐𝑎𝑙𝑖𝑛𝑔:

𝑂𝑟𝑖𝑔𝑖𝑛𝑎𝑙 𝑏𝑖𝑑: 𝐶𝑜𝑠𝑡1 = 16000 𝑀$, 𝑦𝑒𝑎𝑟 = 2011, 𝐶𝑎𝑝𝑎𝑐𝑖𝑡𝑦1 = 70,000𝑙𝑏

ℎ𝑟

𝑆𝑐𝑎𝑙𝑒𝑑 𝑏𝑖𝑑: 𝐶𝑜𝑠𝑡2 = 𝐶𝑜𝑠𝑡1 ∗ (𝐶𝑎𝑝𝑎𝑐𝑖𝑡𝑦2

𝐶𝑎𝑝𝑎𝑐𝑖𝑡𝑦1)

0.65

∗ 𝐹𝑡𝑖𝑚𝑒 ∗ 𝐹𝑙𝑜𝑐𝑎𝑡𝑖𝑜𝑛

Where,

𝐹𝑡𝑖𝑚𝑒 = (1 + 3%)∆𝑦𝑒𝑎𝑟𝑠

∆𝑦𝑒𝑎𝑟𝑠 = 2017 − 2011 = 6 𝑦𝑒𝑎𝑟𝑠

𝐹𝑡𝑖𝑚𝑒 = (1 + 3%)6 = 1.19

𝐹𝑙𝑜𝑐𝑎𝑡𝑖𝑜𝑛 = 1.10

𝐶𝑜𝑠𝑡2 = 16000 ∗ (79100

70000)

0.65

∗ 1.19 ∗ 1.10 = 22.753 𝑀$

The following table shows the scaled values for the OBL bid summary.

53

Table 11: Summarization Scaled OBL Costs

OBL Item Bid Year Initial Bid

M$

Initial

Capacity Ftime

Our

Capacity

Scaled

Cost (M$)

Site

Development 2015 12000 - 1.0609 - 14004

NG Fuel

Utilities 2013 19500

3000

lbmol/hr 1.1255

524

lbmol/hr 7766

Product Storage 2011 16000 70000 lb/hr 1.1941 79100 lb/hr 22753

Environmental 2010 8000 45000 lb/hr 1.2299 103500

lb/hr 18598

Due to all the reactors being packed with catalyst, initially charging the reactors contributes to

overall capital costs of the plant. The amount of catalyst needed for each reactor is calculated when

sizing the reactors. The mass of catalyst is found by utilizing the volume of catalyst needed and

the catalyst density. The price of the catalyst per pound was provided in the project. The following

table summarizes the catalyst capital costs.

Table 12: Summarization of the Catalyst Capital Costs.

Catalyst

Cost of

Catalyst

($/lb)

Volume of

catalyst

needed (ft3)

Catalyst

Density

(lb/ft3)

Mass of

Catalyst (lb)

Total Cost

(M$)

54-7 Q $6.71 3102 53 164,406 1,100

51-7 $16.19 78 6.81 509 8.2

The other costs related to the team, other miscellaneous costs, and pre-operation and start-up. Team

costs are approximated to be 8% of the total contractor costs (TCC), miscellaneous costs are

approximated to be 4% of the TCC, and the pre-op and S/U are approximated to be 7% of the

TCC. These values are equivalent to approximately 18060 M$.

The final costs that need to be considered escalation and project contingency. Because this startup

takes longer than 1 year, escalation needs to be taken into account. It will be assumed that there is

a 2% escalation annually for the 3 year startup period. The cost of escalation was found to be

approximately 7,150 M$. The project contingency is estimated to be 20% of the current total plant

cost. This value is equal to 24,800 M$.

The total plant capital cost was found to be 148,800 M$. This information is summarized in the

pie chart and the following Capital Estimation Summary below.

54

Figure 12: Itemized Capital Cost Breakdown for the Total Plant Cost

Table 13: Capital Estimation Summary

Total Plant Cost Breakdown

TCC IBL

TCC OBL

Misc

Escalation

Project

Contingency

Capital Estimate : MeOH Plant

Base Case : 49800 lb/hr NG FeedBasis Rev 0 4/21/17

2017

665 MMlb/yr MeOH Base CaseNortheastern US

TCC, M$

IBL Reactors/Towers 4,218

Heat Exchangers 9,070

Compressors/Pumps 19,200

Separators 1,539

TCC IBL 34,026

OBL Site Development 14,004

Utilities 8,263

Raw Matl/Prod Storage 22,753

Environmental 18,598

TCC OBL 63,617

Total Contractor Cost, M$ 97,644

Team Costs 8% of TCC 7,812

Other Costs 4% of TCC 3,906

PreOp & S/U (of TCC) 7% of TCC 6,347

Reforming Catalyst $6.71 164000 1,100

MeOH Synthesis Catalyst 16.19$ 540 9

Current Point Estimate, M$ 116,817Years Escalation

Escalation 6% 3 2% 7,150

Project Contingency 20% of CPE+Esc 24,793

Total Plant Cost, M$ 148,761

55

Operating Costs

The operating costs for the base case consists of three primary factors – Raw Materials, Fixed

Costs, and Utilities.

Raw Materials

Below is a complete list of all raw materials and catalysts required for this process:

• Natural Gas Feed

• Water

• Amine Tower Feed

• 57-4 Q catalyst

• 51-7 catalyst

A calculation was completed for the needed utilities, but the values were normalized by utilizing

the amount of product produced. Currently, we are simulating a production of 79160 lb of

MeOH/hr. The cost of the raw materials was used in cents/unit. This allows for the total cash costs

to be calculated for the plant. This value can be compared to the price of product in cents/unit, to

estimate the total profit per year. The largest cost per lb of product is the natural gas feed. The total

raw material cash costs annually were estimated to be 2.73 ¢/lb of product. The following table

summarizes the costs of raw materials. The catalyst is only charged every 3 years.

Table 14: Costs of Raw Materials

Raw Material Amount

Needed

(lb/hr)

Unit/lb of

product

¢/unit Total Cost

(¢/lb of product)

Natural Gas 49269 0.6224 4.3 2.68

Water 150000 1.8949 0.0064 0.01

MDEA/Water feed 124500 1.5728 18 0.003

57-4 Q catalyst 110000 (lb) 1E-3 671 0.04

51-7 catalyst 540 (lb) 3E-4 1619 4.4E-4

Utilities

With the necessary needed utilities calculated in the utility balance, the total cost of utilities for

the annual cash costs could be calculated. Just as with the raw materials, these values can be

normalized to the amount of product created. In conjunction with the raw materials, the profit per

year can be estimated by comparing these values to sales price of the product. The cost of

refrigeration is the single greatest utility cost. The utilities are the largest contributor to the annual

cash costs, at 5.14 ¢/lb of product. The following table and graph summarize the cost of utilities.

56

Table 15: Breakdown of Utility Cash Costs.

Utility Amount

Needed

(MMBTU)

Unit/lb of

product

¢/unit Total Cost

(¢/lb of product)

Fuel 55 0.0007 850 0.59

MPS 38 0.0005 550 0.26

Refrigeration 70 0.0009 3200 2.83

Power 11256 (kw) 0.1422 4 0.57

CW 330 0.0042 25 0.10

Waste Water 12446 (gal) 0.16 5 0.78

Figure 13: Annual Cash Costs of Utilities. Refrigeration has the largest impact.

Fixed Costs

The fixed costs included labor, plant overhead, maintenance, property taxes, and insurance. The

labor costs include a yearly salary of $80,000 for 20 plant operators working a total of 4 shifts.

The 20 workers are made up of 1 equipment operator, 1 control room operator, and 2 laborers to

assist with operations per shift. The other 4 workers are made up of maintenance workers and FTE

(full time equivalent) workers. The overhead costs, estimated as 1.5 times the labor cost, the repair

and maintenance was estimated as 1.5% of the TCC, the property tax was estimated as 2.5% of the

Annual Cash Cost of Utilities(¢/lb product)

Fuel

MPS

Refrigeration

Power

CW

Waste Water

57

TCC, and the insurance was estimated as 0.5% of the TCC. The property tax was the greatest

single fixed cost. The total of the fixed costs was estimated to be approximately 1.61 (¢/lb of

product). The following table summarizes these values.

Table 16: Summarization of the Fixed Cash Costs

Fixed Cost ¢/lb of

product

Labor 0.24

Overhead 0.36

Rep. and

Maintenance

0.34

Property Tax 0.56

Insurance 0.11

The following values were summed, to yield a total annual cash cost. This value was equal to 9.47

¢/lb of product, excluding SG&A fees. The greatest single cost in the annual costs were the

utilities. If these could be reduced, the overall wealth of the plant can be increased. The following

figure summarizes the ongoing costs. This value can be used in the economic analysis of the plant,

to determine the overall predict economics of the plant over time.

58

Figure 14: Annual Ongoing Capital Costs (¢/lb product)

Economic Analysis

Basis

The 665 MMlb per year natural gas to methanol plant, which was added to a pilot plant of Nitlion

chemical in Philadelphia, Pennsylvania, costs approximately 149 MM$ over a three-year

construction of the plant. The projected start up is slated for April 2020. The total project lifespan

is 15 years of operation. The annual ongoing cash costs for the plant, including SG&A, are

approximately 9.60 ¢/lb of product. These values were calculated using the costs of utilities and

raw materials at this current time. The price of methanol was defined at 12.5 ¢/lb. Some

assumptions and parameters used in this economic analysis were:

• Yearly Operating Hours = 8400 hours/year

• Project Life = 15 Operating years

• Capital Spending = 15%/35%/50% over the three year building phase.

• Market Build = 25%/70%/100% over the first three operation years.

• Escalation/Inflation = 2% per year.

• Discount Rate = 9%

• Income Tax Rate = 35%

• Working Capital = 10% of revenues.

• SG&A = 1% of sales

Base Case

2.73

4.36

0.78

1.61

Ongoing Costs (¢/lb product)

Raw Materials

Utilities

Waste Water

Fixed Costs

59

The base case economic evaluation for the process of converting a natural gas feed to methanol

was performed using current values of the raw materials and utilities. This was done to give an

accurate estimate based on the startup date of April 2020. The price of methanol was defined at

12.5 ¢/lb. After analysis, it was found that due to depreciation the plant would not get back the

initial capital costs of 149 MM$ over its lifetime.

For the base case, an ATROR, or after tax return on investment, was found to be 7.26 % at the

current product price. This value is approximately half of the target 15%. An analysis of the

cumulative cash flow was performed to see the economics of the plant over time. This is shown in

the figure below.

Figure 15: Cumulative Cash Flow Model for the Life of the Plant

This figure shows that on a yearly basis, the plant will make a profit. However, due to the

depreciation of a dollar over time, the plant does not ever make back the initial capital costs. The

net present worth (NPW) can be defined as how much money would be made for the lifetime of

the plant, scaled to the present value of a dollar. The NPW for the lifetime of the plant was found

to be -13.4 MM$.

Due to the large capital costs and annual operating costs, the plant will not be profitable over its

lifetime. The summary of this economic analysis can be found on the following page.

-200

-150

-100

-50

0

50

100

150

0 5 10 15 20

Cum

mula

tive

Cas

h F

low

(M

M$

)

Years

Cumulative Cash Flow Model

Startup Period

Operating Years

60

Figure 16: Summary of the Economic Analysis

Sensitivity Analyses

An economic evaluation was performed on the plant for operation in Philadelphia, PA from 2017

to 2035. This economic analysis has revealed that the plant will not break even by the end of the

15 operating years. The analysis was performed using cost and inflation data provided by

NitLion Chemicals. To examine the effect of these parameters, several sensitivity analyses were

performed.

MeOH Plant

Base Case Input & Summary Sheet

Capacity, MM lb/yr MeOH 79160 lb/hr 665 Economic Assumptions

Capital, MM$ 149 April 2020 Startup Discount Rate 9%

Income Tax Rate 35%

Working Capital, % of Revenues 10%

Units unit/lb ¢/unit ¢/lb Prod Inflation/Escalation 2%

Raw Materials SG&A, % of Sales 1%

Natural Gas lb 49269 0.6224 4.3 2.68 hours/year 8400

CO2 lb 0 0.0000 1.0 0.00

Water lb 150000 1.8949 0.0064 0.01

Reformer Catalyst lb 110000 0.0001 671.0 0.04

MeOH Catalyst lb 540 0.0000003 1619 0.00

MDEA/Water feed lb 14.8 0.0002 18 0.003

Total Raw Materials 2.73

Utilities

Fuel MMBTU 55 0.0007 850 0.59

Hot Oil MMBTU - 0.0000 - 0.00

HPS MMBTU - 0.0000 - 0.00

MPS MMBTU 38 0.0005 550 0.26

LPS MMBTU - 0.0000 - 0.00

Refrigeration MMBTU 70 0.0009 3200 2.83

Power kw 11256 0.1422 4.00 0.57

CW MMBTU 330 0.0042 25.00 0.10

Waste Water gal 12446 0.16 5 0.78

Total Utilities 5.14

Fixed Costs Workers Shifts

Labor, m/shift 5$80,000/yr/operator 4 80,000.00$ 0.24

Overhead (1.5 x labor) 0.36

Rep & maint. % Capital = 1.5% 0.34

Property tax % Capital = 2.5% 0.56

Insurance % Capital = 0.5% 0.11

Fixed Costs 1.61

TOTAL CASH COSTS 9.47

SG&A 0.13

Return on Capital CFPO, yrs = 7.70 2.90

Required Product Price for 7.3% ATROR 12.50

NPW ($MM) -13.4 ATROR, % 7.26

61

Five sensitivity analyses were performed to investigate the effects of the relevant economic

parameters. Each of the analyses measured the adjustments to the chosen parameter against the

overall ATROR percentage, or discount rate.

The first sensitivity analysis investigates the effects of an increase to initial raw material costs.

The maximum initial cost of raw materials that will yield a non-negative ATROR value is 4.23

c/lb product. The second analysis investigates the effects of changes to initial product price. The

minimum initial product sale price that will yield a non-negative ATROR is 10.99 c/lb product.

The third analysis investigates the effects of changes to the initial utility cost. The maximum

initial utility cost that yields a non-negative ATROR is 6.63 c/lb product. The fourth analysis

investigates the effects of adjustments to the length of the construction schedule of the plant. The

fifth analysis investigates the effect of changing plant capacity. As product capacity was

increased, values for relevant costs were increased with the amount of product. The cost values

were decreased linearly, but increased at a reduced rate. Aside from the parameter(s) adjusted for

each analysis, the other parameters were held constant at the base case values. Plots of these

adjusted parameters versus ATROR percentage can be found below in Figure 15.

Viewing these analyses, it is clear that the plant’s value is closely tied to each of the adjusted

parameters. The economics appear to be able to tolerate roughly the same amount of variation in

terms of increased utility cost, increased raw material cost, and decreased product price before

reaching a negative discount rate. All three of the values affect the economic success of the plant

significantly. The economics could be greatly improved by speeding up construction (without

added cost) or increasing the overall capacity of the plant which will be impossible given the

current natural gas feed. At this point, it would be prudent to research technological advances

that could improves reaction conversions and reduce capital costs in this plant.

62

Figure 17: Sensitivity analyses for varying economic conditions

63

Required Netback

To achieve a target economic performance of 12% After Tax Rate of Return (ATROR) and a

positive net present value (NPV). The price of the product would have to rise 11.26% to 13.92

c/lb product. If it were possible to sell methanol at this price, the plant would be performing very

well given the current specifications, as shown below.

Figure 18: Required Net Back for 12% ATROR

To achieve an economic performance of 0% ATROR or greater, the required netback for the

product is only 10.99 c/lb product. This price of the product would yield a 0% ATROR and a

negative NPV, leaving the process in an unfavorable position, as shown below.

MeOH Plant




Income Tax Rate 35%





CO2 lb 0 0.0000 1.0 0.00

Water lb 150000 1.8949 0.0064 0.01


MeOH Catalyst lb 540 0.0000003 1619 0.00



Utilities

Fuel MMBTU 55 0.0007 850 0.59

Hot Oil MMBTU - 0.0000 - 0.00

HPS MMBTU - 0.0000 - 0.00

MPS MMBTU 38 0.0005 550 0.26

LPS MMBTU - 0.0000 - 0.00


Power kw 11256 0.1422 4.00 0.57

CW MMBTU 330 0.0042 25.00 0.10

Waste Water gal 12446 0.16 5 0.78








Fixed Costs 1.61


SG&A 0.14



NPW ($MM) 26.3 ATROR, % 12.00

64

Figure 19: Required Net Back for 0% ATROR

With the current price of 12.5 c/lb product, the plant maintains a positive ATROR, but a negative

NPV12. The present economic situation would require an increase in product price to achieve

target economic performance and financial viability.

MeOH Plant




Income Tax Rate 35%





CO2 lb 0 0.0000 1.0 0.00

Water lb 150000 1.8949 0.0064 0.01


MeOH Catalyst lb 540 0.0000003 1619 0.00



Utilities

Fuel MMBTU 55 0.0007 850 0.59

Hot Oil MMBTU - 0.0000 - 0.00

HPS MMBTU - 0.0000 - 0.00

MPS MMBTU 38 0.0005 550 0.26

LPS MMBTU - 0.0000 - 0.00


Power kw 11256 0.1422 4.00 0.57

CW MMBTU 330 0.0042 25.00 0.10

Waste Water gal 12446 0.16 5 0.78








Fixed Costs 1.61


SG&A 0.11



NPW ($MM) -55.5 ATROR, % 0.03

65

Alternate Process Studies

Alternate Process Study 1- Changing the Pressure Ratio

This process study entailed changing the pressure ratio of the compressors between the SMR and

MeOH reactors to 2.45. This requires the use of 4 compressors, instead of the 2 previously

determined. One possible benefit of this is decreasing the size and cost of the compressors, due to

decreasing the duty they perform. Also, another possible benefit to this system could be decreasing

the size and amount of heat exchangers needed between compressors. The smaller pressure ratio

causes a smaller temperature increase in the outlet streams of the compressors. It is predicted this

would allow for only 1 heat exchanger to be used during this overall compression. The costs of

this alternate design can be found in needing to purchase and operate 2 additional compressors.

Also, the operating costs of these compressors need to be taken into account. The diagram on the

next page illustrates this process study.

A quantitative analysis was completed on this process study because the largest single unit capital

cost in the system is the compressors. To begin this study, the new capital costs of the compressors

were determined using ASPEN Process Economic Analyzer software. The following table shows

the proposed compressor specifications and their corresponding prices:

Compressor

Name

Inlet T

(°F)

Inlet P

(psia)

Outlet T

(°F)

Outlet P

(psia)

Price

MM$

K-1 160 30 385 73.5 2.39

K-2 385 73.5 693 180 2.33

K-3 280 180 549 441 2.37

K-4 549 441 886 1015 2.35

Total costs for equipment and installation = $ 9.44 MM.

The heat exchanger would be the same size and require the same amount of cooling duty than the

base case exchanger that is part of the compressor system. While the cooling duty is similar, the

capital cost of the equipment has increased significantly from $ 7.71 MM to $ 9.44 MM.

Additionally, the power required to operate the smaller compressors is negligibly less (13900 kW

as compared to 14100 kW) than the base case operating costs. At this rate, it would take 22 years

to recoup the capital cost disparity. However, as these numbers are estimates for power

consumption, capital cost, and power utility cost, a period of roughly 30 years will be assumed.

As the lifetime of our project simulation is fewer than 30 years, this will be an untenable decision.

In combination, these factors lead to an overall increase in cost for this alternative process study.

66

Additionally, the increased number of compressors will serve to increase the chance for equipment

failure. Overall, this is an undesirable alteration to the process, and the team will be maintaining

the original base case design.

Proposed Design:

67

Alternate Process Study 2 – Changing to a Three-Phase Separator

This case study entailed removing separator 1, and replacing separator 2 with a three phase

separator. This allows for a large amount of water to be recycled back to the Steam Reformer.

There are many potential benefits for this. One potential benefit is increasing the overall

conversion in the SMR. The more water fed, the more conversion of methane is obtained. This

leads to an increase of methanol product obtained. Another benefit is reducing the number of heat

exchangers and separators needed. By utilizing this design, 3 less heat exchangers and 1 less

separator is needed. This could increase the size of the equipment, increase the flowrate of utility

streams, and increase the duties needed. The attached BFD shows the overall process for this study.

Benefits

• More water can be recycled, increasing the overall efficiency of the SMR.

• Only one Separator needs to be purchased.

Costs

• With more water flowing through the process than the base case, the size of all units will

increase. This may offset the price of not purchasing a second separator.

• Compressor sizes will increase due to more flow going into them. This will increase the

price drastically. Compressors have the greatest effect on the capital cost.

This process was not chosen to be analyzed quantitatively because compressors had the greatest

effect on the IBL capital cost. Therefore, it was more desirable to research a cost saving measure

with these units.

68

Proposed Design:

69

HYSYS Model Verification

The model was created, such that, the conditions of the system can be changed, and the system

will react to these changes. There are two hard specifications that must be followed. These are that

the product must be 99.85 wt% pure Methanol at a production rate of at least 71500 lbm/hour. This

is achieved when using the base case data. The following tests were performed to determine how

robust the model is.

Table 17: Model Verification Overview

Test Performed Result from HYSYS Action Required

Decrease NG Flow by 10% System Converged, but did

not reach production rate.

No action is possible. With the

reduced flow of natural gas

feed, the overall product flow

is decreased.

Inlet Steam Reformer Temp

<= 50 F

Temperature of inlet and

outlet were set to 950F.

System converged.

Production rate was not met.

The system did not reach the

production rate. This shows

the importance of maintaining

a temperature of 1100 F in this

reactor.

Changing H2O/CH4 ratio in

steam reformer

H2O/CH4 ratio changed from

12.35 to 5.05. System

converged. Less product than

base case.

The system converged, but the

production was far less than

the base case. This shows the

importance of maintaining

such a high ratio.

Change Methanol Reactor

temp by >= 40F.

Reactor inlet and outlet set to

435 F. System converged, but

less product was created. The

Methanol synthesis reaction

from CO decreased.

Although the system did

converge, less product was

created. This is due to the

production of Methanol from

CO being a function of

temperature.

In summary, the model reacted to these changes relatively well. It converged with nearly instantly

when the changes were implemented. Because the heat exchangers were sized out of the simulation

environment, HYSYS does not reflect the changes to these units. However, this model does reflect

the changes of these tests. Therefore, it can be assumed that this model is a robust simulation of

the kinetics of our design. Screenshots of the Microsoft Excel file used for this verification are

found in Appendix A.

70

Project Definition Rating Index (PDRI)

PDRI 1

The Project Definition Rating Index (PDRI) was used to measure the project’s scope definition for

completeness up to this point. The normalized PDRI score obtained was 653, which indicates that

the project still needs more development prior to detailed design and construction. A summary of

the results is presented in Appendix A. It is important to note that due to time constraints this

project will not be fully defined for detailed design. A PDRI consists of 3 sections and the

following paragraphs discuss more in detail the elements that need to be further defined.

Section I – Basis of Project Decision

The definition level for this section is about 49%. To improve this section we can look into items

that could improve the reliability and affordability of the project. The items could include

production cost, environmental considerations, and operating life cycle for the plant. In terms of

the project scope, the team is progressing to achieve the purpose of this plant. The process is

currently simplified to a point that allowed to create a working HYSYS model. Moving forward,

the team will be looking into ways to simplify the process to reduce cost and avoid using

unnecessary equipment. A more in depth cost analysis will be performed by looking at design and

appropriate materials for equipment.

Section II – Basis of Design

This section has a definition level of 27%, which indicates that, at this point, it is poorly defined.

The low definition level might be attributed to the fact that the team will not have time to work on

some of the elements in this section. For the process/mechanical category, for instance, the team

won’t get to work on Piping and Instrumentation Diagrams (P&IDs) nor instrument index list. To

further define this section, the team will do a further analysis on site information, such as site

location and environmental assessment. Additionally, process flow sheets, heat & material

balances will be improved. The instrument and electrical category will also be reviewed and further

defined for the next gate.

Section III – Execution Approach

This section needs a lot more work, since it only has a definition level of 16%. To improve this

section, the team will be working to further define our HYSYS model. We will also be looking

more into project control requirements and risk analysis. The project execution plan category can’t

really be further defined since elements, such as owner approval requirements, construction plan

& approach, and training requirements are out of the scope of this project due to time constraints.

Section DescriptionPDRI 1

Score

Min

Score

Max

ScoreDef

1

(%)

I BASIS OF PROJECT DECISION 269 29 499 49%

II BASIS OF DESIGN 318 36 423 27%

III EXECUTION APPROACH 66 5 78 16%

Total 653 70 1,000 37%

653 PDRI TOTAL MAXIMUM SCORE = 1000

Normalized Score:

71

PDRI 2

A second PDRI was conducted to determine the improvement of the project scope definition. The

following picture shows the results obtained. The PDRI score for this report is lower than the score

obtained in PDRI 1, which indicates that our project scope definition was improved.

For Section I, Basis of Project Decision, the team was able to increase the definition of this

category by about 20%. The team was able to size the major equipment used in the process and

thus simplify our design by looking at different alternatives. The team was also able to look more

into project scope, more characteristics and design criteria was obtained. Moving forward, we

would need to look more into the project and lead scope of work to further define this section.

Additionally, more materials of constructions and alternatives to the progress should be analyzed

to see of the process can be further simplified.

The Basis of Design section is now 37% defined, a 10% increase compared to the previous report.

The site information category was improved as the team looked into utilities, such as river water,

available at the site location. The process flow diagram was improved by incorporating important

process data and equipment numbers. The PFDs will be improved by adding important control

loops where necessary in the process. The major equipment was sized and the utilities needed and

costs were calculated. To improve this section, the team will be looking more into infrastructure,

such as water treatment requirements.

Section III, the execution approach, did not have a significant improvement. The main reason that

the definition of this section is relatively low is because the project execution approach is beyond

the scope of this class. However, one of the things that could be done to further define this section

is to develop a procurement strategy for the project.


Score

Min

Score

Max

ScoreDef

1

(%)




Total 516 70 1,000 52%

PDRI TOTAL MAXIMUM SCORE = 1000

Normalized Score: 516

72

PDRI 2i

A final PDRI analysis was completed in order to determine the scope definition of the project. The

score for this PDRI analysis was 397 which indicates 65% definition overall. The results are

summarized in the table shown below.

An improvement was observed in section I of this PDRI analysis. In this section the affordability

and feasibility was examined by conducting an economic analysis of the plant. This analysis

included capital and operating cost versus profitability. Additionally, the project strategy was

improved by looking at environmental sustainability and safety aspects in the plant. To further

define this section so that it is ready for detailed design and construction, improvements on design

analysis and lead/discipline scope of work should be analyzed.

The Basis of Design (Section II) was also improved. In this section the team further analyzed the

site location and also conducted a high level environmental assessment. Additionally, process flow

sheets were improved to show all of the equipment in the plant including process controls. Water

treatment requirements were accounted for and equipment utilities were analyzed. The heat and

material balances were also fully completed. In order to improve this section permit requirements

should be obtained and Piping and Instrumentation Diagrams (P&IDs) should be developed.

Mechanical equipment list, line list, and tie-in lists should be created for the process. Additionally,

civil/structural requirements should be analyzed along with architectural requirements to build the

plant.

The last section, Execution Approach, is now at 44% scope definition. This is the section that

would require the most work before the plant can be constructed. A project execution plan needs

to be developed. This includes an engineering/construction plan & approach, startup requirements,

and training requirements.

Section DescriptionPDRI 2i

Score

Min

Score

Max

ScoreDef

1

(%)




Total 397 70 1,000 65%



73

Assumptions

Methanol synthesis reaction

• Other side products such as propanol and butanol will be assumed to be negligible.

• Defined selectivity for methanol synthesis.

• High conversion of CO2 to methanol.

Catalysts used

• Catalyst life is 3 years.

Steam Reforming

• Full conversion of ethane and larger molecules.

• Same catalyst is used for water gas shift reaction.

Pre-Reformer

• Pre-reformer has 100% efficiency on decomposing heavy hydrocarbons.

• 95% conversion in the methanation reaction.

• The same catalyst is used in the Steam Reformer.

Mercury Removal Unit

• 99.9995% of Mercury removed.

74

Outstanding Issues

Utilities

• Too much water is being wasted, yields a large cost in our utilities.

• Requires a large amount of water to be used throughout the system, this yields larger

equipment than initially predicted.

• Price of methanol is low in comparison to our capital and annual cash costs. Therefore, the

plant is not profitable.

• Large reactor size also requires a large catalyst cost.

• Deionized water was included in the OBL costs, due to having an existing plant on site.

Reactor

• Maintaining heating and cooling for reactors

• Sizing of the jackets could not be completed. The necessary duties have been

factored into the project, but not the jacketing itself.

• Preforming and steam reforming reactions very endothermic. This would require

some type of outer heat source.

• Shell and tube design packed with catalyst may need to be researched.

• MeOH reactor is very exothermic. This will require a jacketing around the reactor.

75

Conclusions and Recommendations

Based on the results of the economic analysis, we decided that the plant should not be

constructed provided the given constraints. Over the operating years, the company does make a

net profit, but when scaling the value to present worth, the NPW (lifetime) is approximately

-$13 million. We have attributed the poor economic outlook of the plant to several crucial factors

pertaining to the high costs associated with: raw materials, refrigeration, waste water disposal,

and low efficiency of methanol synthesis reactor. The ATROR target is 12 percent; however, the

ATROR value for our design is significantly lower. As a result, the price of the product would

have to be elevated. Based on the economic analysis coupled with the sensitivity analyses and

the alterative processes studies, the team has decided to not build this plant.

Some recommendations for the process are:

1. Further optimize the refrigeration unit or explore replacement options as the refrigeration

unit is the main contributor to utility costs.

2. Minimize the system water flow rate while maintaining the necessary conversion. This

will reduce the waste water and thus reduce the costs associated with it.

3. Investigate different catalysts with a higher conversion and different operating conditions

for the methanol synthesis.

4. Finally, run a statistical analysis to supplement the sensitivity analyses, such as a Monte

Carlo Simulation, to determine if economic conditions will be favorable to our plant in

the future.

76

Acknowledgements

Team 15, Litlion Chemicals, would like to extend our deepest gratitude to our mentors, Mr.

Nicholas Baran and Mr. Ryan Vanston, who were always able to answer our questions and provide

technical information when needed. Their insights and experience were beyond helpful in this

project. We would also like to extend gratitude to Professor Dawn McFadden for all the technical

assistance, particularly with HYSYS. Without their hard work and dedication, this project would

not have been completed successfully.

77

Research/References

To “sweeten” sour gas (remove S impurities), it is possible to treat with amine solution in a tower.

The amine will bond with the Sulphur compounds, drawing the impurities out of the natural gas.

(Monoethanolamine (MEA), Diethanolamine (DEA), Methylenedioxyethylamphetamine

(MDEA) solutions are possible candidates) [1, 2]. Reference [11] contains HYSYS data for the

separation. Sweet gas is roughly 4 ppm H2S. Reference [12] is a report that recommends lower

amine temperatures for optimal H2S removal.

Potential rate of 42% MDEA solution for LitLion amine treatment. This number was scaled

linearly from the model given in reference [12] to achieve ~4 ppm H2S.

Research: Methods previously used

• Synthetic methanol production first began in 1923 at BASF’s

o Zinc-chromium oxide catalyst was utilized

o Operated at 25-35 MPa and 320-450°C

o Suffered from high capital and compression energy costs

• Imperial Chemical Industries (ICI) developed more active copper-zinc-alumina catalyst

o Operated at 5-10 MPa and 210-270°C

o Higher selectivity and stability

o More energy efficient and cost effective

Mercury Removal Units (MRU) [7, 8, 9]

• Can be regenerative or non-regenerative

• Non-regenerative: adsorb Hg to metal surface, frequently copper

• Good to remove Hg as quickly as possible in MRU, Hg is corrosive and will damage

process equipment

• Non-regenerative processes are simple, but have high installation costs

78

• Mercury has caused numerous aluminum exchanger failures. It amalgamates with

aluminum, resulting in a mechanical failure and gas leakage. Since the level of mercury

that can be tolerated is not established, most operators want to remove it “all.” That is,

remove it to a level where it cannot be detected with the available analytical capability.

Currently, this means reducing the mercury to less than 0.01 µg/Nm3, which

corresponds to about 1 ppt by volume.

Pre reforming [16]:

• Steam must also enter the pre-reformer to reform the hydrocarbons that are larger than

methane.

CnHm + nH2O ⟶ nCO + (n+½m) H2 (1)

CO + 3H2 ⇌ CH4 + H2O (2)

CO + H2O ⇌ CO2 + H2 (3)

References:

1. ChE 470 Course Material. Professor Dawn McFadden.

2. "Amine Treating." Newpoint Gas. n.d. Web. 14 Apr. 2017.

3. "Sour Gas Treating." Bechtel. n.d. Web. 14 Apr. 2017.

4. SYNGAS CONVERSION TO METHANOL." Liquid Fuels. National Energy Technology

Laboratory, n.d. Web. 14 Apr. 201

5. Zhang, Chundong, Ki-Won Jun, Geunjae Kwak, Yun-Jo Lee, and Hae-Gu Park. "Efficient

Utilization of Carbon Dioxide in a Gas-to-methanol Process Composed of CO2/steam-

mixed Reforming and Methanol Synthesis." Journal of CO2 Utilization16 (2016): 1-7.

Web.

6. English, A., Brown E&C, J., Rovner, J., Davies, S. and by Staff, U. 2015. Methanol. Kirk-

Othmer Encyclopedia of Chemical Technology. 1–19

7. "Hydrogen Separations in Syngas Processes." Membrane Technology and Research, Inc,

n.d. Web. 14 Apr. 2017

8. Eckersley, Neil, and Satyam Mishra. "Mercury Removal Options in Hydrocarbon

Processing Facilities." Honeywell UOP. N.p., 2012. Web. 14 Apr. 2017.

9. "UOP Mercury Removal for Natural Gas Production." Honeywell. UOP, n.d. Web. 14 Apr.

2017.

10. Markovs, John. "OPTIMIZED MERCURY REMOVAL IN GAS PLANTS." UOP

Honeywell. Honeywell, n.d. Web. 14 Apr. 2017.

11. Phillips, Steven D., and Joan K. Tarud. "Gasoline from Wood via Integrated Gasification,

Synthesis, and Methanol-to Gasoline Technologies." DEP. NREL, n.d. Web. 14 Apr. 2017.

12. Dyment, Jennifer. "Acid Gas Cleaning Using DEPG Physical Solvents: Validation with

Experimental and Plant Data." Aspentech. Aspen Technology, Inc, n.d. Web. 14 Apr. 2017

13. Satyadileep, D. "Improve Amine Unit Efficiency by Optimizing Operating

Conditions." Gas Processing. Gulf Publishing Company, n.d. Web. 14 Apr. 2017.

79

14. "PRE-REFORMING FOR HYDROGEN." Johnson Matthey Process Technologies. JM,

n.d. Web. 14 Apr. 2017

15. "Sour Gas Sweetening." PetroWiki. N.p., n.d. Web. 14 Apr. 2017

16. Cross, J. "An Introduction to Pre-reforming Catalysis." Pre-Reforming in Syngas Plants.

JM, n.d. Web. 14 Apr. 2017.

80

Appendix A

Equilibrium Data

Table 18: Equilibrium Constants for Gas Phase Reactions [1]

Equilibrium Constants (Keq) for Gas Phase Reactions

Temp.

( F ) 446 644 752 923 1139 1508 1652

CH4 Steam

Reform 1.5E-09 3.9E-06 0.0001 7.1E-03 5.5E-01 1.8E+02 1.1E+03

C2H6 Steam

Reform 2.4E-10 1E-4 2.2E-02 2.5E+01 3.3E+04 4.6E+08 1E+10

C3H8 Steam

Reform 4.7E-12 1E-4 1.1E-01 1.1E+03 1.3E+07 3.1E+12 1.7E+14

Water Shift 100 2.4E+01 1.4E+01 6.22 2.84 1 7.1E-01

NH3

Synthesis 2.5E-01 3.6E-02 1.6E-02 5.5E-03 1.8E-03 4.4E-04 2.7E-04

CH3OH

Synthesis CO 9.1E-3 1.8E-4 3.9E-05 5.2E-06 7.7E-07 7.8E-08 3.9E-08

CH3OH

Synthesis

CO2

1.4E+26 8.8E+19 2.6E+17 1.7E+14 1.5E+11 3.2E+7 2.6E+6

81

HYSYS Model

82

HYSYS Model Continued

83

HYSYS Model Verification

Fresh water purity 1.00 wt frac.

lbs/lbmol

Methane 16.043

Water 18.015

Methanol 32.042

Ammonia 17.031

Natural Gas Comments

Target flow 2744.00 mols/hr

NG Flow in 49289.11 lbs/hr

2744.00 mols/hr At Design Rate

Methane

Mole Fraction in NG 0.9011

Methane In 39668.22 lbs/hr

2472.62 mols/hr

Fresh Water

Fresh water purity 1.00 wt frac.

Fresh water in 150000.00 lbs/hr

Recycle to Reformer

Recycle flow 400000.00 lbs/hr

H2O in recycle 400000.00 lbs/hr

CH4 in recycle 0.00 lbs/hr

Water flow to reformer 550000.00 lbs/hr

30530.11 mols/hr

CH4 flow to reformer 39668.22 lbs/hr

2472.62 mols/hr

H2O/CH4 12.35 Outside Normal Range

Other Feeds

MDEA/H2O 124245.00 lbs/hr

lbs/hr

Steam Reformer Reactor

T at inlet (all feeds) 1100.00 F In Range of Data

T at product outlet 1100.00 F In Range of Data

Methanol Reactor

T at inlet (all feeds) 500.00 F In Range of DataT at product outlet 520.00 F In Range of Data

CH3OH feed to MeOH R 0.00 lbs/hr 99.6C2H5OH feed to MeOH R 0.00 lbs/hr 0.25HCOOCH3 feed to MeOH 0.00 lbs/hr 0.15CH3OH out of MEOH R 77237.98 lbs/hr

C2H5OH out of MEOH R 250.00 lbs/hr

HCOOCH3 out of MEOH R50.00 lbs/hr

CH3OH selectivity 0.9961 wt frac Acceptable Selectivity

C2H5OH selectivity 0.0032 wt frac Acceptable Selectivity

HCOOCH3 selectivity 0.0006 wt frac Acceptable Selectivity

Methanol

Target Product purity 0.9985 wt. frac.

Target CH4 Conversion 0.900 mol frac

Target CH3OH rate 71304.87 lbs/hr

CH3OH in product 76767.49 lbs/hr

Total product rate 76846.57 lbs/hr Meets Target

Product Purity 0.9990 Meets Spec

Waste/Other Streams

Mercury Separation 11.01 lbs/hr

H2S Separation 125653.50 lbs/hr

Recycle Purge 16122.07 lbs/hr

Waste Water 104899.20 lbs/hr

lbs/hr

lbs/hr

Overall Mass Balance 0.00% Mass Balanced

Team target

HYSYS Values - should = HYSYS

Stream Labels to change

Base Case

Natural Gas Comments Natural Gas Comments

Target flow 2469.60 mols/hr Target flow 2744.00 mols/hr

NG Flow in 44360.20 lbs/hr NG Flow in 49289.11 lbs/hr

2471.00 mols/hr At Target Rate 2744.00 mols/hr At Design Rate

Methane Methane

Mole Fraction in NG 0.9011 Mole Fraction in NG 0.9011

Methane In 35721.63 lbs/hr Methane In 39668.22 lbs/hr

2226.62 mols/hr 2472.62 mols/hr

Fresh Water Fresh Water

Fresh water purity 1.00 wt frac. Fresh water purity 1.00 wt frac.

Fresh water in 150000.00 lbs/hr Fresh water in 150000.00 lbs/hr

Recycle to Reformer Recycle to Reformer

Recycle flow 350000.00 lbs/hr Recycle flow 400000.00 lbs/hr

Fraction H2O 350000.00 lbs/hr Fraction H2O 400000.00 lbs/hr

Fraction CH4 lbs/hr Fraction CH4 lbs/hr

Water flow to reformer 500000.00 lbs/hr Water flow to reformer 550000.00 lbs/hr

27754.65 mols/hr 30530.11 mols/hr

CH4 flow to reformer 35721.63 lbs/hr CH4 flow to reformer 39668.22 lbs/hr

2226.62 mols/hr 2472.62 mols/hr

H2O/CH4 12.46 Matches Base Case H2O/CH4 12.35 Matches Base Case

Other Feeds Other Feeds

MDEA/H2O 124245.00 lbs/hr MDEA/H2O 124245.00 lbs/hr

lbs/hr lbs/hr

Steam Reformer Reactor Steam Reformer Reactor

T at inlet (all feeds) 1100.00 F Matches Base Case T at inlet (all feeds) 950.00 F In Range of Data

T at product outlet 1100.00 F T at product outlet 950.00 F In Range of Data

OK Change

Methanol Methanol

Target Product purity 0.9985 wt. frac. Target Product purity 0.9985 wt. frac.

Target CH4 Conversion 0.900 mol frac Target CH4 Conversion 0.900 mol frac

Target CH3OH rate 64210.77 lbs/hr Target CH3OH rate 71304.87 lbs/hr

CH3OH in product 69231.23 lbs/hr CH3OH in product 49387.59 lbs/hr

Total product rate 69302.32 lbs/hr Meets Target Total product rate 49433.60 lbs/hr Does not meet the production target

Product Purity 0.9990 Meets Spec Product Purity 0.9991 Meets Spec

Methanol Reactor Methanol Reactor

T at inlet (all feeds) 500.00 F Matches Base Case T at inlet (all feeds) 500.00 F Matches Base Case

T at product outlet 520.00 F T at product outlet 520.00 F

Waste/Other Streams Waste/Other Streams

Mercury Separation 11.01 lbs/hr Mercury Separation 11.01 lbs/hr

H2S Separation 125653.50 lbs/hr H2S Separation 125653.50 lbs/hr

Gas Fuel Recycle 14437.47 lbs/hr Recycle Purge 26950.25 lbs/hr

Waste Water 109341.42 lbs/hr Waste Water 121485.00 lbs/hr

lbs/hr lbs/hr

lbs/hr lbs/hr

Overall Mass Balance 0.04% Mass Balanced Overall Mass Balance 0.00% Mass Balanced

Model Verification

Change Inlet Steam Reformer Temperature >= 50FDecrease NG flow in by 10%

Model Verification

84

Natural Gas Comments Natural Gas Comments

Target flow 2744.00 mols/hr Target flow 2744.00 mols/hr

NG Flow in 49289.11 lbs/hr NG Flow in 49289.00 lbs/hr

2744.00 mols/hr At Design Rate 2744.00 mols/hr At Design Rate

Methane Methane

Mole Fraction in NG 0.9011 Mole Fraction in NG 0.9011

Methane In 39668.22 lbs/hr Methane In 39668.22 lbs/hr

2472.62 mols/hr 2472.62 mols/hr

Fresh Water Fresh Water

Fresh water purity 1.00 wt frac. Fresh water purity 1.00 wt frac.

Fresh water in 75000.00 lbs/hr Fresh water in 150000.00 lbs/hr

Recycle to Reformer Recycle to Reformer

Recycle flow 150000.00 lbs/hr Recycle flow 400000.00 lbs/hr

Fraction H2O 150000.00 lbs/hr Fraction H2O 400000.00 lbs/hr

Fraction CH4 lbs/hr Fraction CH4 lbs/hr

Water flow to reformer 225000.00 lbs/hr Water flow to reformer 550000.00 lbs/hr

12489.59 mols/hr 30530.11 mols/hr

CH4 flow to reformer 39668.22 lbs/hr CH4 flow to reformer 39668.22 lbs/hr

2472.62 mols/hr 2472.62 mols/hr

H2O/CH4 5.05 Normal Range H2O/CH4 12.35 Matches Base Case

OK Change

Other Feeds Other Feeds

MDEA/H2O 124245.00 lbs/hr MDEA/H2O 124245.00 lbs/hr

lbs/hr lbs/hr

Steam Reformer Reactor Steam Reformer Reactor

T at inlet (all feeds) 1100.00 F Matches Base Case T at inlet (all feeds) 1100.00 F Matches Base Case

T at product outlet 1100.00 F T at product outlet 1100.00 F

Methanol Methanol

Target Product purity 0.9985 wt. frac. Target Product purity 0.9985 wt. frac.

Target CH4 Conversion 0.900 mol frac Target CH4 Conversion 0.900 mol frac

Target CH3OH rate 71304.87 lbs/hr Target CH3OH rate 71304.87 lbs/hr

CH3OH in product 59695.83 lbs/hr CH3OH in product 74384.04 lbs/hr

Total product rate 59752.17 lbs/hr Does not meet the production targetTotal product rate 74471.68 lbs/hr Meets Target

Product Purity 0.9991 Meets Spec Product Purity 0.9988 Meets Spec

Methanol Reactor Methanol Reactor

T at inlet (all feeds) 500.00 F Matches Base Case T at inlet (all feeds) 435.00 F In Range of Data

T at product outlet 520.00 F T at product outlet 435.00 F In Range of Data

OK Change

Waste/Other Streams Waste/Other Streams

Mercury Separation 11.01 lbs/hr Mercury Separation 11.01 lbs/hr

H2S Separation 125653.50 lbs/hr H2S Separation 125653.50 lbs/hr

Recycle Purge 22818.73 lbs/hr Recycle Purge 19734.64 lbs/hr

Waste Water 40297.26 lbs/hr Waste Water 103664.06 lbs/hr

lbs/hr lbs/hr

lbs/hr lbs/hr

Overall Mass Balance 0.00% Mass Balanced Overall Mass Balance 0.00% Mass Balanced

Change Methanol Reactor Temperature >= 40FChange H2O/CH4 Ratio by atleast 1

85

Chemical Properties Table

Chemical Property Table Methanol Nitrogen Carbon Dioxide Methane Ethane Propane Hydrogen Iso-Butane n-Butane Iso-Pentane

Formula CH3OH N2 CO2 CH4 C2H6 C3H8 H2 iC4H10 nC4H10 iC5H12

MW 32.05 28 44.01 16.05 30.08 44.11 2.02 58.1 58.1 72.17

NBP, F 148.5 -320.4 -110.3 -258.7 -128.2 -43.6 -423.4 11 31 82

VP, mmHg 126.96 N/A 43683 N/A 28841 6397 N/A 2356 1558 N/A

Freeze, F -144 -345.8 -110.3 -305.7 -305.7 -305.7 -434.5 -255 -217 -255.6

H2O Solubility, % 1000 g/L N/A N/A N/A N/A N/A N/A Slight Slight N/A

Flammability Non-flammable Non-Flammable

UFL, % 44 N/A N/A 15 13 8.4 76 8.4 8.4 8.3

LFL, % 6 N/A N/A 5 3 1.8 4 1.6 1.6 1.4

Flash Point, F 49.5 N/A N/A -306.7 -155.2 -155.2 N/A N/A N/A -59.8

Class Flam Liquid N/A N/A Flam. Gas Flam. Gas Flam Gas Flam. Gas Flam Gas Flam Gas Flam Gas/Liq

Toxicity O2 Asphyxiation

Hazard

O2 Asphyxiation

Hazard

NIOSH REL-TWA

800 ppm

TLV-TWA, ppm 200 N/A 5000 N/A N/A N/A N/A N/A N/A 1000

TLV-C, ppm N/A N/A N/A N/A N/A N/A N/A N/A N/A N/A

OSHA PEL, ppm 200 N/A 5000 N/A N/A 1000 N/A 800 N/A N/A

IDLH, ppm N/A N/A 40000 N/A N/A 2100 N/A N/A N/A N/A

Incompatibles

Oxidizers X X X X X X X

Caustic/Acid X

Metal Chlorides X

Peroxides

Amines

Alcohols

NH3, HCN, H2S

Reactivity

Polymerization

Peroxide

Formation

86

Chemical Property Table (Continued) N-Pentane Cyclohexane Water Hydrogen Sulfide Mercury Methyl Formate Ethanol Carbon Monoxide Ammonia

Formula nC5H12 C6H12 H2O H2S Hg HCOOCH3 CH3CH2OH CO NH3

MW 72.17 84.2 18.02 34.1 200.6 60.1 46.1 28 17

NBP, F 97 177 212 -77 674 89 173 -313 -28

VP, mmHg 420 77 17.25 13376 0.0012 476 44 >26600 6460

Freeze, F -202 44 32 -122 -38 -148 -173 337 -108

H2O Solubility, % 0.04 Insoluble 100% 0.04 Insoluble 30 Miscible 2 34

Flammability

UFL, % 7.8 8 N/A 44 N/A 23 19 74 28

LFL, % 1.5 1.3 N/A 4 N/A 4.5 3.3 12.5 15

Flash Point, F -57 0 N/A N/A N/A -2 55 N/A N/A

Class Flam Gas/Liq Flam Liq N/A Flam Gas N/A Flam Liq Flam Liq Flam Gas Flam. Gas

Toxicity NIOSH REL-

TWA

100 ppm

NIOSH REL-TWA

1000 ppm

NIOSH REL-TWA

35 ppm

NIOSH REL-

TWA

25 ppm

TLV-TWA, ppm N/A N/A N/A N/A 0.00305 N/A N/A N/A N/A

TLV-C, ppm N/A N/A N/A N/A N/A N/A N/A N/A N/A

OSHA PEL, ppm 1000 300 N/A 20-50 (10 min) 0.01219 100 ppm 1000 ppm 50 ppm 50 ppm

IDLH, ppm 1500 1300 N/A 100 N/A 1200 ppm N/A

Incompatibles KOH,BrF4,Pt,Na,CH3COCl,CH3COBr BrF3,ClF3,Li Halogens

Oxidizers X X X X X X X X

Caustic/Acid X X X

Metal Chlorides X X

Peroxides

Amines X X

Alcohols

NH3, HCN, H2S X X

Reactivity

Polymerization

Peroxide Formation

Health Effects

Carcinogen

87

Component Data

TXY and XY diagrams are useful tools when separating components. The following figures are

implemented to determine how well a separation would occur. The Methanol needs to be purified

to reach the standards for sale. Also, the waste streams can be separated and possibly sold for

profit.

Figure 20:

Figure 22: XY Diagram for Methanol and Water. Figure 23: XY Diagram for Methanol and Water

Figure 21: TXY diagram for Methanol and Water. This is useful when determining the number of stages needed in a distillation

column to reach purity standards.

88

Figure 24: TXY plot for Methane and Carbon Dioxide.

Figure 25: XY Plot for Methane and Carbon Dioxide

89

Figure 5: TXY plot for Methane and Nitrogen. Figure 26:TXY plot for Methane and Nitrogen.

Figure 27: XY plot for Methane and Nitrogen.

90

Figure 28: TXY plot for Methanol and Methyl Formate.

Figure 29: XY plot for Methanol and Methyl Formate.

91

Figure 31: XY Diagram for Methanol and Ethanol

Figure 30: TXY diagram for Methanol and Ethanol. This will be useful in determining the number of stages necessary to reach

the methanol purity needed.

92

Initial BFD Created – Gate 1

1. Initial BFD design by Alexander Hatza

2. Initial BFD by Brian Klapat

93

3. Initial BFD design by Daniel Cordova

4. Initial BFD design by Sarah Ramzy

94

Code of Ethics

“LIT-Lion Chemical”

Preamble

On this, the 31st day of January, 2017, “LIT-Lion Chemical” has assembled to construct a code

ethics, detailing expectations and goals for the Ch E 470 project as well as procedures and

schedules for regular meetings. Failure to comply with these standards will result in the strict

proceedings detailed in Section III [Retribution].

Section I. Main Body

All members of “LIT-Lion Chemical” will hold paramount the integrity and dignity of both

the engineering profession and The Pennsylvania State University by:

● Providing honest and truthful work;

● Being attentive, respectful, and helpful to their mentor, Professor McFadden, and

teammates;

● Being willing to provide assistance to teammates to help further the overall progress of

the project;

● Providing the deliverables for projects within the allotted time frame;

● Completing an equal amount of work for the given assignment;

● Seeking help from other teammates when needed, but never asking for someone to

complete the work;

● Treating the project seriously and always working to their fullest potential to ensure the

completion of the project or assignment in a professional and timely manner;

Section II. Regulations

● All team members shall attend the team meetings. The team meetings are scheduled for

Tuesdays 3:00 pm to 6:00 pm. If need be, additional time blocks will be scheduled for

Wednesday evenings and/ or weekends. For these meetings, all team members must agree

95

on the time to be scheduled to ensure availability of all parties.

● If it is unavoidable for a team member to miss a meeting, that meeting will not count

towards absence. These reasons include: emergencies, religious ceremonies, exams,

injury, and illness.

● Missing a meetings for reasons such as: not setting an alarm, missing the bus, forgetting

the meeting time, etc. are not excusable and will have consequences.

● All team members must have their share of the assignment done at least one full day

before the assignment is due. This ensures assignment completion and revision by

groupmates.

● If for some reason a team member will not be able to achieve this, it is his/her

responsibility to update the team in a timely fashion.

● During meetings, all team members should be attentive. Avoid cell phone use, side

conversations, and other distractions such as social media. Team members must be

prepared for the meetings in order to avoid delays in team progress.

● Any questions, comments, or concerns should be brought up during the meeting or in

the GroupMe.

● Ideas and suggestions presented during meetings should be respected and analyzed by all

team members.

● Team members will not falsify any data or materials in pursuit of project excellence.

● All team members will strive to maintain human and environmental health in designing

the project plant.

Section III. Retribution

The following penalties shall apply to any member of “LIT-Lion Chemical” found to be in

violation of the code of ethics by his or her teammates. A two to one majority, among the peers

of the accused, must be found in group proceedings in order to determine culpability for any

individual infraction. Abstentions in voting leading to a tie vote shall acquit the defendant.

● First offense: Apologize (informal).

● Second offense: Apologize (formal, written; GroupMe is acceptable).

96

● Third and further offenses: Do the most annoying part of the next gate section.

● If the non-offending parties can unanimously agree that a particular infraction would not

be redeemed with any of the aforementioned punishments, a replacement course of action

can be chosen unanimously on a case by case basis.

Signature:

Daniel Cordova Bernal, Alexander Hatza, Brian Klapat, Sarah Ramzy

97

Appendix B

PDRI – Score Chart Summary

Note: *

Project: Senior Design Project

Type of Project: Methanol Synthesis Plant

Project Location: Enter on Cover Page

Owner/Client: Enter on Cover Page

Project No.: Gate 3

Other Control No.: Enter on Cover Page

Stage PDRI 1:

Feasibility

PDRI 2:

Concept

PDRI 2i:

Detailed

Scope

PDRI 3:

Detailed

Scope

Normalized

Score653 516 397 NA

Score 653 516 397 -

Max Score 1,000 1,000 1,000 -

Total Max Score 1,000 1,000 1,000 1,000

Typical Min * 550 450 300 150

Typical Max * 800 600 450 250

Typical range of scores are based on experience on PDRI tools since 1996.

-

100

200

300

400

500

600

700

800

900

1,000

No

rmal

ize

d S

core

Front End Planning Stage

PDRI Score Chart Summary

Normalized Score Typical Min * Typical Max *

98

PDRI 1 – Table Summary of Results


Score

Min

Score

Max

ScoreDef

1

(%)




Total 653 70 1,000 37%

Category DescriptionPDRI 1

Score

Min

Score

Max

ScoreDef

1

(%)A MANUFACTURING OBJECTIVES CRITERIA 28 3 45 40%

B BUSINESS OBJECTIVES 97 11 213 57%

C BASIC DATA RESEARCH & DEVELOPMENT 56 4 94 42%

D PROJECT SCOPE 75 11 120 41%

E VALUE ENGINEERING 13 - 27 52%

F SITE INFORMATION 78 8 104 27%

G PROCESS/MECHANICAL 135 15 196 34%

H EQUIPMENT SCOPE 18 3 33 50%

I CIVIL, STRUCTURAL, & ARCHITECTURAL 19 2 19 0%

J INFRASTRUCTURE 22 3 25 14%

K INSTRUMENT & ELECTRICAL 46 5 46 0%

L PROCUREMENT STRATEGY 15 1 16 7%

M DELIVERABLES 6 - 9 33%

N PROJECT CONTROL 10 1 17 44%

P PROJECT EXECUTION PLAN 35 3 36 3%

Total 653 70 1,000 37%

Top 10

ElementsDescription

PDRI 1

Score

Min

Score

Max

ScoreDef

1

(%)B.1 Products 22 1 56 62%

B.5 Capacities 21 2 55 64%

C.1 Technology 39 2 54 29%

C.2 Processes 17 2 40 61%

G.1 Process Flow Sheets 8 2 36 82%

F.1 Site Location 18 2 32 47%

G.3 Piping and Instrumentation Diagrams (P&IDs) 31 2 31 0%

D.3 Site Characteristics Available vs. Required 22 2 29 26%

B.2 Market Strategy 16 2 26 42%

D.1 Project Objectives Statement 14 2 25 48%

Total 208 19 384 48%

Notes:

1 - Definition percentage indicates the level of completeness of specific Section, Category and Element.

653 PDRI TOTAL MAXIMUM SCORE = 1000

Normalized Score:

99

PDRI 2 – Table Summary of Results


Score

Min

Score

Max

ScoreDef

1

(%)




Total 516 70 1,000 52%

Category DescriptionPDRI 2

Score

Min

Score

Max

ScoreDef

1








H EQUIPMENT SCOPE 21.00 3.00 33 40%

I CIVIL, STRUCTURAL, & ARCHITECTURAL 16.00 2.00 19 18%

J INFRASTRUCTURE 19.00 3.00 25 27%

K INSTRUMENT & ELECTRICAL 35.00 5.00 46 27%





Total 516 70 1,000 52%

Top 10

ElementsDescription

PDRI 2

Score

Min

Score

Max

ScoreDef

1

(%)B.1 Products 1 1 56 100%










Total 119 19 384 73%

Notes:




100

PDRI 2i – Table Summary of Results

Section DescriptionPDRI 2i

Score

Min

Score

Max

ScoreDef

1

(%)




Total 397 70 1,000 65%

Category DescriptionPDRI 2i

Score

Min

Score

Max

ScoreDef

1








H EQUIPMENT SCOPE 21 3 33 40%

I CIVIL, STRUCTURAL, & ARCHITECTURAL 13 2 19 35%

J INFRASTRUCTURE 14 3 25 50%

K INSTRUMENT & ELECTRICAL 28 5 46 44%





Total 397 70 1,000 65%

Top 10

ElementsDescription

PDRI 2i

Score

Min

Score

Max

ScoreDef

1

(%)B.1 Products 1 1 56 100%










Total 100 19 384 78%

Notes:




2i Summary Definition

101

Appendix C- Supporting Files

Cash Flow Models

Base Case

MeO

H P

lan

tC

as

h F

low

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de

l

Dis

co

unt R

ate

9%

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me

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x R

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35

%P

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uc

t P

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e, c

/lb

12

.5

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rkin

g C

ap

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l, %

of R

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nue

s1

0%

Ca

sh

Flo

w, M

M$

11

2

Infla

tio

n2

%R

aw

Ma

tl C

os

ts, c

/lb

Pro

du

ct

2.7

PW

15

, M

M$

-13

SG

&A

, %

of S

ale

s1

%B

yP

rod

uc

t C

red

it, c

/lb

Pro

du

ct

0.0

AT

RO

R, %

7.3

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lity

Co

sts

, c

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du

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5.1

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nt

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1.6

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17

20

18

20

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20

20

20

21

20

22

20

23

20

24

20

25

20

26

20

27

20

28

20

29

20

30

20

31

20

32

20

33

20

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35

Op

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g C

ap

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102

Sensitivity Analysis 1: Raw Material Cost = 3.5 c/lb product

MeO

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31

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-22

-52

-74

.47

.82

0.2

19

.31

6.3

16

.51

3.8

11

.01

1.2

11

.41

1.6

11

.91

2.1

12

.41

2.6

12

.92

.0

Cum

ula

tive

Ca

sh F

low

, M

M$

-22

-74

-14

9-1

41

-12

1-1

01

-85

-69

-55

-44

-33

-21

-10

21

42

73

95

25

4

PW

Facto

r (a

t th

e D

iscount R

ate

)0.9

174

0.8

417

0.7

722

0.7

084

0.6

499

0.5

963

0.5

470

0.5

019

0.4

604

0.4

224

0.3

875

0.3

555

0.3

262

0.2

992

0.2

745

0.2

519

0.2

311

0.2

120

0.1

945

Pre

se

nt W

ort

h (

Annua

l)-2

0-4

4-5

76

13

12

98

65

44

44

33

33

0

Cum

ula

tive

PW

, M

M$

-20

-64

-12

2-1

16

-10

3-9

2-8

3-7

4-6

8-6

3-5

9-5

5-5

1-4

8-4

4-4

1-3

8-3

5-3

5

103

Sensitivity Analysis 2: Product Price = 11.5 c/lb product

MeO

H P

lan

tC

as

h F

low

Mo

de

l

Dis

co

unt R

ate

9%

Inco

me

Ta

x R

ate

35

%P

rod

uc

t P

ric

e, c

/lb

11

.5

Wo

rkin

g C

ap

ita

l, %

of R

eve

nue

s1

0%

Ca

sh

Flo

w, M

M$

38

Infla

tio

n2

%R

aw

Ma

tl C

os

ts, c

/lb

Pro

du

ct

2.7

PW

15

, M

M$

-41

SG

&A

, %

of S

ale

s1

%B

yP

rod

uc

t C

red

it, c

/lb

Pro

du

ct

0.0

AT

RO

R, %

2.9

Uti

lity

Co

sts

, c

/lb

Pro

du

ct

5.1

Pla

nt

Ca

pa

cit

y, M

Mlb

/yr

66

5F

ixe

d/ C

os

ts, c

/lb

Pro

du

ct

1.6

Pla

nt

Ca

pit

al, M

M$

14

9S

G&

A C

os

ts, c

/lb

Pro

du

ct

0.1

Ye

ar

20

17

20

18

20

19

20

20

20

21

20

22

20

23

20

24

20

25

20

26

20

27

20

28

20

29

20

30

20

31

20

32

20

33

20

34

20

35

Op

era

tin

g C

ap

ac

ity

, %

0%

0%

0%

25

%7

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

0%

Ca

pit

al S

pe

nd

ing

15

%3

5%

50

%0

%

Ye

ar

12

34

56

78

91

01

11

21

31

41

51

61

71

81

9

Op

era

tin

g Y

ea

r1

23

45

67

89

10

11

12

13

14

15

Cu

mu

lati

ve

In

fla

tio

n/E

sc

ala

tio

n1

.00

01

.02

01

.04

01

.06

11

.08

21

.10

41

.12

61

.14

91

.17

21

.19

51

.21

91

.24

31

.26

81

.29

41

.31

91

.34

61

.37

31

.40

01

.42

8

Ca

pita

l Co

st, M

M$

22

52

74

00

00

00

00

00

00

00

00

De

pre

cia

ble

Ba

sis

, M

M$

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

9

Sa

les R

eve

nue

s, M

M$

/yr

00

02

05

88

48

68

89

09

19

39

59

79

91

01

10

31

05

10

70

Ra

w M

atl

Co

sts

, M

M$

/yr

00

05

14

20

20

21

21

22

22

23

23

23

24

24

25

25

0

ByP

rod

uct C

red

it, M

M$

/yr

00

00

00

00

00

00

00

00

00

0

Utilit

y C

osts

, M

M$

/yr

00

09

26

38

38

39

40

41

42

42

43

44

45

46

47

48

0

Fix

ed

/ C

osts

, M

M$

/yr

00

01

11

21

21

21

21

31

31

31

31

41

41

41

41

51

50

SG

&A

Co

sts

, M

M$

/yr

00

00

11

11

11

11

11

11

11

0

Op

era

ting

Co

sts

, M

M$

/yr

00

02

55

27

07

27

37

57

67

87

98

18

28

48

68

88

90

Befo

re T

ax

Reve

nues fro

m O

pera

tion, M

M$

00

0-5

61

41

41

51

51

51

61

61

61

61

71

71

71

80

De

pre

cia

tio

n R

ate

, %

20

.0%

32

.0%

19

.2%

11

.5%

11

.5%

5.8

%

De

pre

cia

tio

n, M

M$

29

.74

47

.58

28

.55

17

.10

17

.10

8.6

2

Be

fore

Ta

x In

co

me

, M

M$

00

0-3

5-4

1-1

5-3

-26

15

16

16

16

16

17

17

17

18

0

Inco

me

Ta

xes

00

0-1

2-1

4-5

-1-1

25

56

66

66

66

0

After

Tax

Reve

nues fro

m O

pera

tion, M

M$

00

0-2

3-2

7-9

-2-2

41

01

01

01

01

11

11

11

11

20

To

tal W

ork

ing

Ca

pita

l, M

M$

0.0

0.0

0.0

-0.5

10

.62

1.4

01

.43

1.4

61

.49

1.5

21

.55

1.6

1.6

1.6

1.7

1.7

1.7

1.8

0.0

Ne

w W

ork

ing

Ca

pita

l, M

M$

0.0

0.0

0.0

-0.5

11

.13

0.7

90

.03

0.0

30

.03

0.0

30

.03

0.0

0.0

0.0

0.0

0.0

0.0

0.0

Wo

rkin

g C

ap

ita

l Re

co

very

, M

M$

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

1.8

AT

Op

era

ting

Ca

sh, M

M$

-22

-52

-74

.47

.12

0.7

19

.11

5.3

15

.51

2.7

9.9

10

.11

0.3

10

.51

0.7

10

.91

1.1

11

.31

1.6

0.0

AT

Ca

sh F

low

, M

M$

-22

-52

-74

.47

.61

9.5

18

.31

5.3

15

.51

2.7

9.8

10

.01

0.2

10

.51

0.7

10

.91

1.1

11

.31

1.5

1.8

Cum

ula

tive

Ca

sh F

low

, M

M$

-22

-74

-14

9-1

41

-12

2-1

03

-88

-73

-60

-50

-40

-30

-19

-92

13

25

36

38

PW

Facto

r (a

t th

e D

iscount R

ate

)0.9

174

0.8

417

0.7

722

0.7

084

0.6

499

0.5

963

0.5

470

0.5

019

0.4

604

0.4

224

0.3

875

0.3

555

0.3

262

0.2

992

0.2

745

0.2

519

0.2

311

0.2

120

0.1

945

Pre

se

nt W

ort

h (

Annua

l)-2

0-4

4-5

75

13

11

88

64

44

33

33

32

0

Cum

ula

tive

PW

, M

M$

-20

-64

-12

2-1

16

-10

4-9

3-8

4-7

7-7

1-6

7-6

3-5

9-5

6-5

2-4

9-4

7-4

4-4

2-4

1

104

Sensitivity Analysis 3: Utility Cost = 6 c/lb product

MeO

H P

lan

tC

as

h F

low

Mo

de

l

Dis

co

unt R

ate

9%

Inco

me

Ta

x R

ate

35

%P

rod

uc

t P

ric

e, c

/lb

12

.5

Wo

rkin

g C

ap

ita

l, %

of R

eve

nue

s1

0%

Ca

sh

Flo

w, M

M$

47

Infla

tio

n2

%R

aw

Ma

tl C

os

ts, c

/lb

Pro

du

ct

2.7

PW

15

, M

M$

-38

SG

&A

, %

of S

ale

s1

%B

yP

rod

uc

t C

red

it, c

/lb

Pro

du

ct

0.0

AT

RO

R, %

3.5

Uti

lity

Co

sts

, c

/lb

Pro

du

ct

6.0

Pla

nt

Ca

pa

cit

y, M

Mlb

/yr

66

5F

ixe

d/ C

os

ts, c

/lb

Pro

du

ct

1.6

Pla

nt

Ca

pit

al, M

M$

14

9S

G&

A C

os

ts, c

/lb

Pro

du

ct

0.1

Ye

ar

20

17

20

18

20

19

20

20

20

21

20

22

20

23

20

24

20

25

20

26

20

27

20

28

20

29

20

30

20

31

20

32

20

33

20

34

20

35

Op

era

tin

g C

ap

ac

ity

, %

0%

0%

0%

25

%7

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

0%

Ca

pit

al S

pe

nd

ing

15

%3

5%

50

%0

%

Ye

ar

12

34

56

78

91

01

11

21

31

41

51

61

71

81

9

Op

era

tin

g Y

ea

r1

23

45

67

89

10

11

12

13

14

15

Cu

mu

lati

ve

In

fla

tio

n/E

sc

ala

tio

n1

.00

01

.02

01

.04

01

.06

11

.08

21

.10

41

.12

61

.14

91

.17

21

.19

51

.21

91

.24

31

.26

81

.29

41

.31

91

.34

61

.37

31

.40

01

.42

8

Ca

pita

l Co

st, M

M$

22

52

74

00

00

00

00

00

00

00

00

De

pre

cia

ble

Ba

sis

, M

M$

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

9

Sa

les R

eve

nue

s, M

M$

/yr

00

02

26

39

29

49

59

79

91

01

10

31

05

10

81

10

11

21

14

11

60

Ra

w M

atl

Co

sts

, M

M$

/yr

00

05

14

20

20

21

21

22

22

23

23

23

24

24

25

25

0

ByP

rod

uct C

red

it, M

M$

/yr

00

00

00

00

00

00

00

00

00

0

Utilit

y C

osts

, M

M$

/yr

00

01

13

04

44

54

64

74

84

95

05

15

25

35

45

55

60

Fix

ed

/ C

osts

, M

M$

/yr

00

01

11

21

21

21

21

31

31

31

31

41

41

41

41

51

50

SG

&A

Co

sts

, M

M$

/yr

00

00

11

11

11

11

11

11

11

0

Op

era

ting

Co

sts

, M

M$

/yr

00

02

75

67

77

88

08

28

38

58

68

89

09

29

49

69

70

Befo

re T

ax

Reve

nues fro

m O

pera

tion, M

M$

00

0-5

71

51

51

61

61

61

71

71

71

81

81

81

91

90

De

pre

cia

tio

n R

ate

, %

20

.0%

32

.0%

19

.2%

11

.5%

11

.5%

5.8

%

De

pre

cia

tio

n, M

M$

29

.74

47

.58

28

.55

17

.10

17

.10

8.6

2

Be

fore

Ta

x In

co

me

, M

M$

00

0-3

5-4

1-1

4-2

-27

16

17

17

17

18

18

18

19

19

0

Inco

me

Ta

xes

00

0-1

2-1

4-5

-1-1

36

66

66

66

77

0

After

Tax

Reve

nues fro

m O

pera

tion, M

M$

00

0-2

3-2

7-9

-1-1

51

11

11

11

11

11

21

21

21

20

To

tal W

ork

ing

Ca

pita

l, M

M$

0.0

0.0

0.0

-0.4

90

.68

1.5

01

.53

1.5

61

.59

1.6

21

.65

1.7

1.7

1.8

1.8

1.8

1.9

1.9

0.0

Ne

w W

ork

ing

Ca

pita

l, M

M$

0.0

0.0

0.0

-0.4

91

.17

0.8

20

.03

0.0

30

.03

0.0

30

.03

0.0

0.0

0.0

0.0

0.0

0.0

0.0

Wo

rkin

g C

ap

ita

l Re

co

very

, M

M$

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

1.9

AT

Op

era

ting

Ca

sh, M

M$

-22

-52

-74

.47

.22

1.1

19

.71

5.9

16

.11

3.3

10

.51

0.7

11

.01

1.2

11

.41

1.6

11

.91

2.1

12

.30

.0

AT

Ca

sh F

low

, M

M$

-22

-52

-74

.47

.71

9.9

18

.91

5.9

16

.11

3.3

10

.51

0.7

10

.91

1.1

11

.41

1.6

11

.81

2.1

12

.31

.9

Cum

ula

tive

Ca

sh F

low

, M

M$

-22

-74

-14

9-1

41

-12

1-1

02

-86

-70

-57

-46

-36

-25

-14

-29

21

33

45

47

PW

Facto

r (a

t th

e D

iscount R

ate

)0.9

174

0.8

417

0.7

722

0.7

084

0.6

499

0.5

963

0.5

470

0.5

019

0.4

604

0.4

224

0.3

875

0.3

555

0.3

262

0.2

992

0.2

745

0.2

519

0.2

311

0.2

120

0.1

945

Pre

se

nt W

ort

h (

Annua

l)-2

0-4

4-5

75

13

11

98

64

44

43

33

33

0

Cum

ula

tive

PW

, M

M$

-20

-64

-12

2-1

16

-10

3-9

2-8

3-7

5-6

9-6

5-6

1-5

7-5

3-5

0-4

6-4

3-4

1-3

8-3

8

105

Sensitivity Analysis 4: Construction delayed to 4 years

MeO

H P

lan

tC

as

h F

low

Mo

de

l

Dis

co

unt R

ate

9%

Inco

me

Ta

x R

ate

35

%P

rod

uc

t P

ric

e, c

/lb

12

.5

Wo

rkin

g C

ap

ita

l, %

of R

eve

nue

s1

0%

Ca

sh

Flo

w, M

M$

91

Infla

tio

n2

%R

aw

Ma

tl C

os

ts, c

/lb

Pro

du

ct

2.7

PW

15

, M

M$

-21

SG

&A

, %

of S

ale

s1

%B

yP

rod

uc

t C

red

it, c

/lb

Pro

du

ct

0.0

AT

RO

R, %

6.1

Uti

lity

Co

sts

, c

/lb

Pro

du

ct

5.1

Pla

nt

Ca

pa

cit

y, M

Mlb

/yr

66

5F

ixe

d/ C

os

ts, c

/lb

Pro

du

ct

1.6

Pla

nt

Ca

pit

al, M

M$

14

9S

G&

A C

os

ts, c

/lb

Pro

du

ct

0.1

Ye

ar

20

17

20

18

20

19

20

20

20

21

20

22

20

23

20

24

20

25

20

26

20

27

20

28

20

29

20

30

20

31

20

32

20

33

20

34

20

35

Op

era

tin

g C

ap

ac

ity

, %

0%

0%

0%

0%

25

%7

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

0%

Ca

pit

al S

pe

nd

ing

15

%1

5%

25

%4

5%

Ye

ar

12

34

56

78

91

01

11

21

31

41

51

61

71

81

9

Op

era

tin

g Y

ea

r1

23

45

67

89

10

11

12

13

14

15

Cu

mu

lati

ve

In

fla

tio

n/E

sc

ala

tio

n1

.00

01

.02

01

.04

01

.06

11

.08

21

.10

41

.12

61

.14

91

.17

21

.19

51

.21

91

.24

31

.26

81

.29

41

.31

91

.34

61

.37

31

.40

01

.42

8

Ca

pita

l Co

st, M

M$

22

22

37

67

00

00

00

00

00

00

00

0

De

pre

cia

ble

Ba

sis

, M

M$

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

9

Sa

les R

eve

nue

s, M

M$

/yr

00

00

22

64

94

95

97

99

10

11

03

10

51

08

11

01

12

11

41

16

0

Ra

w M

atl

Co

sts

, M

M$

/yr

00

00

51

42

02

12

12

22

22

32

32

32

42

42

52

50

ByP

rod

uct C

red

it, M

M$

/yr

00

00

00

00

00

00

00

00

00

0

Utilit

y C

osts

, M

M$

/yr

00

00

92

63

83

94

04

14

24

24

34

44

54

64

74

80

Fix

ed

/ C

osts

, M

M$

/yr

00

01

11

21

21

21

21

31

31

31

31

41

41

41

41

51

50

SG

&A

Co

sts

, M

M$

/yr

00

00

01

11

11

11

11

11

11

0

Op

era

ting

Co

sts

, M

M$

/yr

00

01

12

65

37

27

37

57

67

87

98

18

38

48

68

88

90

Befo

re T

ax

Reve

nues fro

m O

pera

tion, M

M$

00

0-1

1-3

11

22

22

23

23

24

24

24

25

25

26

26

27

0

De

pre

cia

tio

n R

ate

, %

20

.0%

32

.0%

19

.2%

11

.5%

11

.5%

5.8

%

De

pre

cia

tio

n, M

M$

29

.74

47

.58

28

.55

17

.10

17

.10

8.6

2

Be

fore

Ta

x In

co

me

, M

M$

00

0-4

1-5

1-1

75

51

42

32

42

42

42

52

52

62

62

70

Inco

me

Ta

xes

00

0-1

4-1

8-6

22

58

88

99

99

99

0

After

Tax

Reve

nues fro

m O

pera

tion, M

M$

00

0-2

7-3

3-1

13

39

15

15

16

16

16

17

17

17

18

0

To

tal W

ork

ing

Ca

pita

l, M

M$

0.0

0.0

0.0

-1.1

3-0

.35

1.1

42

.17

2.2

22

.26

2.3

12

.35

2.4

2.4

2.5

2.5

2.6

2.6

2.7

0.0

Ne

w W

ork

ing

Ca

pita

l, M

M$

0.0

0.0

0.0

-1.1

30

.79

1.4

81

.04

0.0

40

.04

0.0

50

.05

0.0

0.0

0.0

0.0

0.1

0.1

0.1

Wo

rkin

g C

ap

ita

l Re

co

very

, M

M$

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

2.7

AT

Op

era

ting

Ca

sh, M

M$

-22

-22

-37

.2-6

3.9

14

.41

7.4

20

.12

0.4

17

.71

5.0

15

.31

5.6

15

.91

6.2

16

.61

6.9

17

.21

7.6

0.0

AT

Ca

sh F

low

, M

M$

-22

-22

-37

.2-6

2.7

13

.61

5.9

19

.12

0.4

17

.71

4.9

15

.21

5.6

15

.91

6.2

16

.51

6.8

17

.21

7.5

2.7

Cum

ula

tive

Ca

sh F

low

, M

M$

-22

-45

-82

-14

5-1

31

-11

5-9

6-7

6-5

8-4

3-2

8-1

24

20

36

53

70

88

91

PW

Facto

r (a

t th

e D

iscount R

ate

)0.9

174

0.8

417

0.7

722

0.7

084

0.6

499

0.5

963

0.5

470

0.5

019

0.4

604

0.4

224

0.3

875

0.3

555

0.3

262

0.2

992

0.2

745

0.2

519

0.2

311

0.2

120

0.1

945

Pre

se

nt W

ort

h (

Annua

l)-2

0-1

9-2

9-4

49

91

01

08

66

65

55

44

41

Cum

ula

tive

PW

, M

M$

-20

-39

-68

-11

2-1

04

-94

-84

-73

-65

-59

-53

-48

-42

-38

-33

-29

-25

-21

-21

106

Sensitivity Analysis 5: Plant capacity at 50,000 lb product/hr (reduced utility, raw material costs)

MeO

H P

lan

tC

as

h F

low

Mo

de

l

Dis

co

unt R

ate

9%

Inco

me

Ta

x R

ate

35

%P

rod

uc

t P

ric

e, c

/lb

12

.5

Wo

rkin

g C

ap

ita

l, %

of R

eve

nue

s1

0%

Ca

sh

Flo

w, M

M$

10

6

Infla

tio

n2

%R

aw

Ma

tl C

os

ts, c

/lb

Pro

du

ct

1.9

PW

15

, M

M$

-16

SG

&A

, %

of S

ale

s1

%B

yP

rod

uc

t C

red

it, c

/lb

Pro

du

ct

0.0

AT

RO

R, %

6.9

Uti

lity

Co

sts

, c

/lb

Pro

du

ct

3.5

Pla

nt

Ca

pa

cit

y, M

Mlb

/yr

42

0F

ixe

d/ C

os

ts, c

/lb

Pro

du

ct

2.5

Pla

nt

Ca

pit

al, M

M$

14

9S

G&

A C

os

ts, c

/lb

Pro

du

ct

0.1

Ye

ar

20

17

20

18

20

19

20

20

20

21

20

22

20

23

20

24

20

25

20

26

20

27

20

28

20

29

20

30

20

31

20

32

20

33

20

34

20

35

Op

era

tin

g C

ap

ac

ity

, %

0%

0%

0%

25

%7

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

10

0%

0%

Ca

pit

al S

pe

nd

ing

15

%3

5%

50

%0

%

Ye

ar

12

34

56

78

91

01

11

21

31

41

51

61

71

81

9

Op

era

tin

g Y

ea

r1

23

45

67

89

10

11

12

13

14

15

Cu

mu

lati

ve

In

fla

tio

n/E

sc

ala

tio

n1

.00

01

.02

01

.04

01

.06

11

.08

21

.10

41

.12

61

.14

91

.17

21

.19

51

.21

91

.24

31

.26

81

.29

41

.31

91

.34

61

.37

31

.40

01

.42

8

Ca

pita

l Co

st, M

M$

22

52

74

00

00

00

00

00

00

00

00

De

pre

cia

ble

Ba

sis

, M

M$

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

91

49

14

9

Sa

les R

eve

nue

s, M

M$

/yr

00

01

44

05

85

96

06

26

36

46

56

76

86

97

17

27

40

Ra

w M

atl

Co

sts

, M

M$

/yr

00

02

69

99

99

10

10

10

10

10

11

11

11

0

ByP

rod

uct C

red

it, M

M$

/yr

00

00

00

00

00

00

00

00

00

0

Utilit

y C

osts

, M

M$

/yr

00

04

11

16

17

17

17

18

18

18

19

19

19

20

20

21

0

Fix

ed

/ C

osts

, M

M$

/yr

00

01

11

21

21

21

21

31

31

31

31

41

41

41

41

51

50

SG

&A

Co

sts

, M

M$

/yr

00

00

01

11

11

11

11

11

11

0

Op

era

ting

Co

sts

, M

M$

/yr

00

01

72

93

73

83

94

04

04

14

24

34

44

54

54

64

70

Befo

re T

ax

Reve

nues fro

m O

pera

tion, M

M$

00

0-4

11

21

21

22

22

22

23

23

24

24

25

25

26

26

0

De

pre

cia

tio

n R

ate

, %

20

.0%

32

.0%

19

.2%

11

.5%

11

.5%

5.8

%

De

pre

cia

tio

n, M

M$

29

.74

47

.58

28

.55

17

.10

17

.10

8.6

2

Be

fore

Ta

x In

co

me

, M

M$

00

0-3

3-3

7-8

44

13

22

23

23

24

24

25

25

26

26

0

Inco

me

Ta

xes

00

0-1

2-1

3-3

12

58

88

88

99

99

0

After

Tax

Reve

nues fro

m O

pera

tion, M

M$

00

0-2

2-2

4-5

33

91

51

51

51

51

61

61

61

71

70

To

tal W

ork

ing

Ca

pita

l, M

M$

0.0

0.0

0.0

-0.3

51

.07

2.0

72

.11

2.1

52

.20

2.2

42

.29

2.3

2.4

2.4

2.5

2.5

2.6

2.6

0.0

Ne

w W

ork

ing

Ca

pita

l, M

M$

0.0

0.0

0.0

-0.3

51

.43

1.0

00

.04

0.0

40

.04

0.0

40

.04

0.0

0.0

0.0

0.0

0.0

0.1

0.1

Wo

rkin

g C

ap

ita

l Re

co

very

, M

M$

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

2.6

AT

Op

era

ting

Ca

sh, M

M$

-22

-52

-74

.48

.12

3.6

23

.51

9.7

20

.01

7.3

14

.61

4.9

15

.21

5.5

15

.81

6.1

16

.41

6.7

17

.10

.0

AT

Ca

sh F

low

, M

M$

-22

-52

-74

.48

.52

2.2

22

.51

9.7

19

.91

7.3

14

.51

4.8

15

.11

5.4

15

.71

6.0

16

.41

6.7

17

.02

.6

Cum

ula

tive

Ca

sh F

low

, M

M$

-22

-74

-14

9-1

40

-11

8-9

6-7

6-5

6-3

9-2

4-9

62

13

75

36

98

61

03

10

6

PW

Facto

r (a

t th

e D

iscount R

ate

)0.9

174

0.8

417

0.7

722

0.7

084

0.6

499

0.5

963

0.5

470

0.5

019

0.4

604

0.4

224

0.3

875

0.3

555

0.3

262

0.2

992

0.2

745

0.2

519

0.2

311

0.2

120

0.1

945

Pre

se

nt W

ort

h (

Annua

l)-2

0-4

4-5

76

14

13

11

10

86

65

55

44

44

1

Cum

ula

tive

PW

, M

M$

-20

-64

-12

2-1

16

-10

1-8

8-7

7-6

7-5

9-5

3-4

7-4

2-3

7-3

2-2

8-2

4-2

0-1

6-1

6

107

Capital Equipment List and Sizes

Detailed Capital Estimate

Design Direct

P, psig T, F MOC Installed Indirects Total

Packing Head Motor Comments s/t s/t s/t Cost Eng/Pro Rack/Sewers Contractor

No. of Diam Height Tan-Tan Area GPM ft HP Cost

Items ft ft ft ft2 (SCFM) (psi) (BHP) s/t = shell/tube HEx M$ M$ M$ M$

(1) (2) (3)

Reactors/Towers Reactors/Towers #

Prereformer 1 7 21 29 425 530 CS Prereformer 1 328 56 69 453

Steam Reformer 1 10 30 41 Refractory Lining 45 1150 CS Steam Reformer1 412 70 86 568

MeOH Synth 1 2 4 7 1102 570 CS MeOH Synth 1 117 20 25 161

T-100 1 11 122 40 trays 18 300 CS T-100 1 1100 187 230 1518

T-103 1 11 122 40 trays 18 300 CS T-103 1 1100 187 230 1518

3057 520 640 4218

Heat Exchangers Heat Exchangers #

E-100 1 748 425/45 550/1150 CS E-100 1 168 29 35 232

E-107 1 14077 425/45 570/1113 CS E-107 1 372 63 78 513

E-101 1 7757 45/70 690/160 CS E-101 1 269 46 56 371

E-108 1 6670 45/70 330/160 CS E-108 1 258 44 54 356

E-103 1 2564 1030/70 745/160 CS E-103 1 183 31 38 252

E-102 1 1003 1030/45 950/256 CS E-102 1 156 27 33 215

E-105 1 516 1030/70 657/160 CS E-105 1 102 17 21 141

E-104 1 4700 45/90 250/10 CS E-104 1 211 36 44 291

E-109 1 1181 45/70 230/160 CS E-109 1 132 22 28 182

Condenser 2 5880 40/70 225/160 CS Condener 2 504 86 106 695

Reboiler 2 5272 45/70 438/275 CS Reboiler 2 437 74 92 603

FH-100 1 6903 45 1400 CS FH-100 1 3783 644 792 5219

6575 1119 1377 9070

Expanders/Compressors

K-100 1 4413 0 530 425 CS Expanders/Compressors #

K-104 1 11481 18851 195 746 CS K-100 1 3150 536 660 4346

K-105 1 2806 254838 1102 941 CS K-104 1 2785 474 583 3842

K-102 1 1185 0 1102 570 CS K-105 1 4916 836 1030 6782

P-100 2 219 115 11 65 245 CS K-102 1 2819 480 591 3889

P-101 2 214 306 23 115 225 CS P-100 2 122 21 26 168

P-101 2 125 21 26 172

Separators 13917 2368 2915 19200

MRU 1 4 11 Mercury 425 200 SS304

ATU 3 2 8 SS lining 425 200 CS/SS304 Separators #

V-101 1 11 35 45 210 CS MRU 1 180 31 38 248

V-100 1 7 30 45 80 CS ATU 3 378 64 79 521

V-101 1 345 59 72 476

V-100 1 212 36 44 293

1115 190 234 1539

Total IBL 24664 4007 4933 34026

Fill in what you know

P-Design = P-Operating + 30 psi or P-Operating * 1.10, whichever is larger

T-Design = T-Operating + 50F (good enough for now…in real life, must be analyzed for different situations) 1 IPE-EQUIP.ICS

2 IPE-Capital Cost Reports-Project Summaries-Account Basis

3 Delta - (CapCost Report DIRECTS - EQUIP.ICS)

4 per RN-20% of Total Installed Cost

Indirects Sewers

Eng/Pro Piperacks

(2) 3960 From Icarus

(3) 236.30

(4) 4933

4196 4933

MeOH Plant (LitLion Chemical)

Equipment List

Allocated Costs

Allocations

108

Detailed Capital Estimate

Design Direct

P, psig T, F MOC Installed Indirects Total

Packing Head Motor Comments s/t s/t s/t Cost Eng/Pro Rack/Sewers Contractor

No. of Diam Height Tan-Tan Area GPM ft HP Cost

Items ft ft ft ft2 (SCFM) (psi) (BHP) s/t = shell/tube HEx M$ M$ M$ M$

(1) (2) (3)

Reactors/Towers Reactors/Towers #

Prereformer 1 7 21 29 425 530 CS Prereformer 1 328 56 69 453

Steam Reformer 1 10 30 41 Refractory Lining 45 1150 CS Steam Reformer1 412 70 86 568

MeOH Synth 1 2 4 7 1102 570 CS MeOH Synth 1 117 20 25 161

T-100 1 11 122 40 trays 18 300 CS T-100 1 1100 187 230 1518

T-103 1 11 122 40 trays 18 300 CS T-103 1 1100 187 230 1518

3057 520 640 4218

Heat Exchangers Heat Exchangers #

E-100 1 748 425/45 550/1150 CS E-100 1 168 29 35 232

E-107 1 14077 425/45 570/1113 CS E-107 1 372 63 78 513

E-101 1 7757 45/70 690/160 CS E-101 1 269 46 56 371

E-108 1 6670 45/70 330/160 CS E-108 1 258 44 54 356

E-103 1 2564 1030/70 745/160 CS E-103 1 183 31 38 252

E-102 1 1003 1030/45 950/256 CS E-102 1 156 27 33 215

E-105 1 516 1030/70 657/160 CS E-105 1 102 17 21 141

E-104 1 4700 45/90 250/10 CS E-104 1 211 36 44 291

E-109 1 1181 45/70 230/160 CS E-109 1 132 22 28 182

Condenser 2 5880 40/70 225/160 CS Condener 2 504 86 106 695

Reboiler 2 5272 45/70 438/275 CS Reboiler 2 437 74 92 603

FH-100 1 6903 45 1400 CS FH-100 1 3783 644 792 5219

6575 1119 1377 9070

Expanders/Compressors

K-100 1 4413 0 530 425 CS Expanders/Compressors #

K-104 1 11481 18851 195 746 CS K-100 1 3150 536 660 4346

K-105 1 2806 254838 1102 941 CS K-104 1 2785 474 583 3842

K-102 1 1185 0 1102 570 CS K-105 1 4916 836 1030 6782

P-100 2 219 115 11 65 245 CS K-102 1 2819 480 591 3889

P-101 2 214 306 23 115 225 CS P-100 2 122 21 26 168

P-101 2 125 21 26 172

Separators 13917 2368 2915 19200

MRU 1 4 11 Mercury 425 200 SS304

ATU 3 2 8 SS lining 425 200 CS/SS304 Separators #

V-101 1 11 35 45 210 CS MRU 1 180 31 38 248

V-100 1 7 30 45 80 CS ATU 3 378 64 79 521

V-101 1 345 59 72 476

V-100 1 212 36 44 293

1115 190 234 1539

Total IBL 24664 4007 4933 34026

Fill in what you know

P-Design = P-Operating + 30 psi or P-Operating * 1.10, whichever is larger

T-Design = T-Operating + 50F (good enough for now…in real life, must be analyzed for different situations) 1 IPE-EQUIP.ICS

2 IPE-Capital Cost Reports-Project Summaries-Account Basis

3 Delta - (CapCost Report DIRECTS - EQUIP.ICS)

4 per RN-20% of Total Installed Cost

Indirects Sewers

Eng/Pro Piperacks

(2) 3960 From Icarus

(3) 236.30

(4) 4933

4196 4933

MeOH Plant (LitLion Chemical)

Equipment List

Allocated Costs

Allocations

109

Detailed Energy Balance

Streams in m (lbm/h) T (F) P (psia) Cp (BTU/lbm-F) dh vap (BTU/lb) Heat Flow (BTU/hr)

1 4.93E+04 80 400 0.5409 446 2.13E+06

Streams out m (lbm/h) T (F) P (psia) Cp (BTU/lbm-F) dh vap (BTU/lb) Heat Flow (BTU/hr)

2 1.10E+01 80 400 0.3149 162.8 2.77E+02

3 4.93E+04 80 400 0.541 226.5 2.13E+06

Overall Flow -195

MRU


3 4.93E+04 80 400 0.541 226.5 2.13E+06

4 1.24E+05 40 400 0.8371 587 4.16E+06


5 1.26E+05 52.43 400 0.838 596.3 5.52E+06

6 4.79E+04 72 400 0.5477 207.6 1.89E+06

Overall Flow -1.12E+06

Requires 1.12E+06 Btu/hr

ATU

Streams in m (lbm/h) T (F) P (psia) Cp (BTU/lbm-F) dh vap (BTU/lb)

6 4.79E+04 72 400 0.5477 207.6 Heat Flow (BTU/hr)

10 6.05E+05 1100 30 0.5762 1502

1.27E+07 6

Streams out m (lbm/h) T (F) P (psia) Cp (BTU/lbm-F) dh vap (BTU/lb) -1.27E+07 10

6_1 4.79E+04 500 400 0.6899 207.6

10_1 6.05E+05 1063.556 30 0.5731 1502

E-100

110

Streams in m (lbm/h)T (F) P (psia) Cp (BTU/lbm-F) dh vap (BTU/lb) Heat Flow (BTU/hr)

7 1.50E+05 500 400 0.6805 780.6 5.10E+07

6_1 4.79E+04 500 400 0.6899 207.6 1.65E+07

Streams out m (lbm/h)T (F) P (psia) Cp (BTU/lbm-F) dh vap (BTU/lb) Heat Flow (BTU/hr)

Reactor Feed 1.98E+05 474.7 400 0.5715 1025 5.37E+07

Overall 1.39E+07

Remove 1.39E+07

to lower temp to 474.7 F

MIX-100


Reactor Feed 1.98E+05 474.7 400 0.5827 1025 5.47E+07

Streams out m (lbm/h) T (F) P (psia) Cp (BTU/lbm-F) dh vap (BTU/lb)

8 1.98E+05 480.4 400 0.5793 1455 5.51E+07

Overall Flow -3.34E+05

RXN dH (kJ/mol) dh(BTU/lbmol) lbmol/hr gmol/hr kJ/hr BTU/hr

WGS -41 N/A 64.58 29294.10499 -1201058 -1138603.27

ref CH4 206 N/A 2472.62 1121554.98 2.31E+08 219026228.9

ref C2H6 347 N/A 109.76 49786.0384 17275755 16377416.05

ref C3H8 497 N/A 24.70 11201.85864 5567324 5277822.909

methanating -206 N/A 2752.45 1248484.485 -2.6E+08 -243814038

iC4H10 N/A 283484.394 5.49 N/A N/A 1555762.354

nC4H10 N/A 279787.0556 2.74 N/A N/A 767735.6806

iC5H12 N/A 348315.0734 2.74 N/A N/A 955776.5613

nC5H12 N/A 345348.6042 1.37 N/A N/A 473818.2849

C6H12 N/A 392251.4915 0.55 N/A N/A 215267.6185

REACTION TOTAL -302812.969

-3.03E+05 Btu/hr

Remove 3.03E+05 to make reactor isothermal

Prereformer


8 1.98E+05 480.4 400 0.5793 1455 Heat Flow (BTU/hr)

-2.81E+07


8_1 1.98E+05 218.9 30 0.508 1293

K-100

111

Streams in m (lbm/h)T (F) P (psia) Cp (BTU/lbm-F) dh vap (BTU/lb) vapor flow Heat Flow (BTU/hr)

13_1 4.07E+05 160 30 1.001 951.9 0 6.52E+07

8_1 1.98E+05 218.9 30 0.508 1293 1.21E+05 2.20E+07


9 6.05E+05 216.8 30 0.7587 1455 1.15E+05 9.95E+07

Overall -4.16E+06

Add -4.16E+06

To heat to 228.7

MIX-101


9 6.05E+05 216.8 30 0.9126 -786 Heat Flow (BTU/hr)

10_1 6.05E+05 1063 30 0.5731 1502

1.49E+08 9

Streams out m (lbm/h) T (F) P (psia) Cp (BTU/lbm-F) dh vap (BTU/lb) -1.49E+08 10

9_1 6.05E+05 520 30 0.4966 -786

10_2 6.05E+05 637.1928 30 0.5835 1502

The component that is condensing in stream 10 is only H2O

There are 485370 lb/hr condensing flow

dHvap for H2O 40.66 BTU/lb 19735144.2 BTU/hr

E-107

112

Streams in m (lbm/h)T (F) P (psia) Cp (BTU/lbm-F) dh vap (BTU/lb)

9_1 6.05E+05 520 30 0.4966 1455 Heat Flow (BTU/hr)

1.86E+08

Streams out m (lbm/h)T (F) P (psia) Cp (BTU/lbm-F) dh vap (BTU/lb)

9_2 6.05E+05 1100 30 0.5606 1455

BURNING METHANE BURNING H2

dHcomb dH comb

CH4 + O2 -> CO2 + H20 -802.34 kJ/gmol LHV 2 H2+O2 -> 2 H2O -244 kJ/gmol LHV

-760.4714918 Btu/gmol -231.267 Btu/gmol

-344942.264 Btu/lbmol -104901 Btu/lbmol

To heat sufficiently, burn:

537.9034353 lbmol/hr CH4 1768.776 lbmol/hr H2

8629.315861 lbm/hr CH4 3565.641 lbm/hr H2

Have 3868.09 lbm/hr available 5191.46 lbm/hr available

241.1151 lbmol/hr available 2575.28 lbmol/hr available

Possible Heat: -3.53E+08 Btu/hr total --> 15006.18 lbm/hr

-1.68E+08 Btu/hr extra --> 7125.681 lbm/hr

7880.5 lbm/hr used in FH

Verifying that stream has dH comb > 8000 Btu/lbm

frac

CH4 -21501.8 Btu/lbm 0.257766427

H2 -52037.1 Btu/lbm 0.3459545

Mix -23544.9 Btu/lbm (HYSYS = -24779.86 Btu/lbm)

FH-100


9_2 6.05E+05 1100 30 0.6116 1455 4.07E+08


10 6.05E+05 1100 30 0.6409 -1003 4.27E+08

RXN dH (kJ/mol) lbmol/hr gmol/hr kJ/hr BTU/hr

ref CH4 206 2449.68 1111151 228897148.6 2.17E+08

ref C2H6 347 0.00 0 0 0

ref C3H8 497 0.00 0 0 0

WGS -41 123.24 55900.77 -2291931.585 -2172751

Total 2.15E+08

Add 2.15E+08 BTU/hr to maintain isothermal in reformer

Reformer

113


10_2 6.05E+05 637.1928 30 0.5835 -1014

CW1 5.93E+06 90 70 0.998 908.1 Heat Flow (BTU/hr)

-1.18E+08 10

Streams out m (lbm/h) T (F) P (psia) Cp (BTU/lbm-F) dh vap (BTU/lb) 1.18E+08 CW

10_3 6.05E+05 280 30 0.51092 -1014

CW2 5.93E+06 110 70 0.9982 908.1

E-101


10_3 6.05E+05 280 30 0.51092 1502 Heat Flow (BTU/hr)

CW3 1.67E+06 90 70 0.998 908.1 -3.33E+07 10

3.33E+07 CW


10_4 6.05E+05 160 30 0.9287 1502

CW4 1.67E+06 110 70 0.9982 908.1

The component that is condensing in stream 10 is only H2O



E-108

Streams inm (lbm/h)T (F) P (psia) Cp (BTU/lbm-F)dh vap (BTU/lb)

10_4 6.05E+05 160 30 0.9287 1502 Heat Flow (BTU/hr)

3.86E+04

Streams outm (lbm/h)T (F) P (psia) Cp (BTU/lbm-F)dh vap (BTU/lb)

11 4.71E+05 160 30 1.001 951.9

14 1.34E+05 160 30 0.6724 948.9

V-101

Streams inm (lbm/h) T (F) P (psia) Cp (BTU/lbm-F) dh vap (BTU/lb) Heat Flow (BTU/hr)

11 4.71E+05 160 30 1.001 951.9 7.55E+07

Streams outm (lbm/h) T (F) P (psia) Cp (BTU/lbm-F) dh vap (BTU/lb) Heat Flow (BTU/hr)

13 4.07E+05 160 30 1.001 951.9 6.52E+07

12 6.40E+04 160 30 1.001 951.9 1.02E+07

Overall 0.00E+00

TEE-100

114


14 1.34E+05 160 30 0.6472 947 4.86E+07


14_1 1.34E+05 695.4077 180 0.7075 1850

Calculate T2

T1 619.67 R T2 1155.077685 R

k 1.385 695.4076848 F

np 0.79981

n/n-1 0.34755547

K-104


14_1 1.34E+05 695 180 0.71424 977 -3.88E+07 14

CW5 1.94E+06 90 70 0.998 908.1 3.88E+07 CW


14_2 1.34E+05 280 180 0.6829 977

CW6 1.94E+06 110 70 0.9982 908.1

E-103


14_2 1.34E+05 280 180 0.6829 1185 5.81E+07


14_3 1.34E+05 891.9 1015 0.735 1255

Calculate T2

T1 739.67 R T2 1351.532281 R

k 1.385 891.8622809 F

np 0.79765

n/n-1 0.3484966

K-105

115


14_3 1.34E+05 891.8623 1015 0.735 1255 Heat Flow (BTU/hr)

18 1.19E+05 30 30 1.053 638.7 -2.73E+07 14

2.73E+07 18


14_4 1.34E+05 609.2688 1015 0.7096 1425

18_1 1.19E+05 206 30 0.9811 638.7

H2O vaporization



CH3OH vaporization


dHvap for CH3OH 504.498 BTU/lb 5889509.652 BTU/hr

E-102


14_4 1.34E+05 609.2688 1015 0.70959 1255 -1.04E+07 14

CW7 5.19E+05 90 70 0.998 908.1 1.04E+07 CW


14_5 1.34E+05 500 1015 0.7069 1425

CW8 5.19E+05 110 70 0.9982 908.1

E-105


14_5 1.35E+05 500 1015 0.7069 1425 4.78E+07


16 1.34E+05 520 1015 0.6401 -800.9 4.46E+07

Overall 3.26E+06

RXN dH (kJ/mol) dH (BTU/lbmol) lbmol/hr gmol/hr kJ/hr BTU/hr

WGS -41 N/A 1871.29 848800.5028 -34800820.6 -3.3E+07

CO to MeOH -91 N/A 430.84 195424.2338 -17783605.3 -1.7E+07

CO2 to MeOH -49.43 N/A 2050.26 929976.662 -45968746.4 -4.4E+07

Ammonia -92 N/A 1.72 779.5305691 -71716.8124 -67987.5

Methyl Form N/A -49802.2885 0.20 91.7475136 N/A -10073.5

Ethanol N/A -109529.781 0.57 260.3785477 N/A -62874.4

TOTAL -9.36E+07

MeOH Synthesis

116


16 1.34E+05 520 1015 0.6401 -800.9 -3.89E+07


16_1 1.34E+05 185 30 1.093 -800.9

K-102


16_1 1.34E+05 185 30 1.093 -8.01E+02 Heat Flow (BTU/hr)

R1 N/A -44 63 0.559 1.74E+02

-2.14E+07 16

Streams out m (lbm/h) T (F) P (psia) Cp (BTU/lbm-F) dh vap (BTU/lb) 70000000 R

16_2 1.34E+05 30 30 0.9733 -5.44E+02

R2 N/A -44 63 0.559 1.74E+02 Suficient refrigeration from unit is

2.14E+07 Btu/hr

E-104


16_2 1.34E+05 30 30 1.087 1853 4.37E+06


17 1.50E+04 30 30 1.351 1283 6.08E+05

18 1.19E+05 30 30 1.053 638.7 3.76E+06

Overall 2.41E+03

V-100


18_1 1.19E+05 206.3 30 1.058 664.9 2.59E+07


19 5.94E+04 206.3 30 1.058 664.9 1.30E+07

20 5.94E+04 206.3 30 1.058 664.9 1.30E+07

Overall 0.00E+00

TEE-101

117


17 1.50E+04 30 30 1.351 1283 6.08E+05


17_1 7.40E+03 30 30 1.351 1283 3.00E+05

17_2 7.60E+03 30 30 1.351 1283 3.08E+05

Overall 0.00E+00

TEE-102


12 6.40E+04 160 30 1.001 951.9 1.02E+07

22 1.99E+04 238.5 29 1.011 954.5 4.80E+06

24 1.99E+04 238.5 29 1.011 954.5 4.80E+06


25 1.04E+05 192.8 29 1.004 954.1 2.01E+07

Overall -2.46E+05

MIX-103


21 3.96E+04 174.6 24.5 0.9565 482.1 6.61E+06

23 3.96E+04 174.6 24.5 0.9565 482.1 6.61E+06


26 7.91E+04 174.6 24.5 0.9565 482 1.32E+07

Overall 0.00E+00

MIX-104


25 1.04E+05 192.8 29 1.004 954.1 2.01E+07


25_1 1.04E+05 195.8 50 1.004 928.8 2.04E+07

Overall -3.12E+05

P-100

118


26 7.91E+04 174.6 24.5 0.9565 482 1.32E+07


26_1 7.91E+04 175 100 0.9558 423.6 1.32E+07

Overall -2.06E+04

P-101


26_1 7.91E+04 172 100 1.0002 423.6 -4.02E+06 14

CW9 2.02E+05 90 70 0.998 908.1 4.02E+06 CW


26_2 7.91E+04 120 100 0.9555 423.6

CW10 2.02E+05 110 70 0.9982 908.1

E-109

119

Stre

ams

inm

(lb

m/h

)T

(F)

P (

psi

a)C

p (

BTU

/lb

m-F

)d

h v

ap (

BTU

/lb

)

205.

94E+

0420

6.3

301.

058

664.

9

Stre

ams

ou

tm

(lb

m/h

)T

(F)

P (

psi

a)C

p (

BTU

/lb

m-F

)d

h v

ap (

BTU

/lb

)

233.

96E+

0417

4.6

24.5

0.95

6548

2.1

241.

99E+

0423

8.5

291.

011

954.

5

Mas

s Fl

ow

(lb

m/h

r)

Tem

p o

f

Tray

(F)

Re

bo

ile

r/

Co

nd

en

ser

Ou

tle

t (F

)

Tem

p

Ch

ange

(F)

Cp

inp

ut

(BTU

/lb

-F)

Cp

ou

tpu

t

(BTU

/lb

-F)

Cp

ave

rage

(BTU

/lb

-F)

Tem

p C

han

ge

Ene

rgy

(BTU

/hr)

dh

vap

(BTU

/lb

)

Ph

ase

Ch

ange

Ene

rgy

(BTU

/hr)

Tota

l En

erg

y

Pe

r U

nit

(BTU

/hr)

con

de

nse

r (2

3)3.

96E+

0417

4.80

1.75

E+02

-2.0

0E-0

11.

06E+

009.

57E-

011.

01E+

00-7

.97E

+03

4.82

E+02

1.91

E+07

1.91

E+07

reb

oil

er

(24)

1.99

E+04

234.

902.

39E+

023.

60E+

001.

06E+

001.

01E+

001.

03E+

007.

41E+

049.

55E+

021.

90E+

071.

91E+

07

TOTA

L3.

81E+

07

T-10

0

120

Stre

ams

inm

(lb

m/h

)T

(F)

P (

psi

a)C

p (

BTU

/lb

m-F

)d

h v

ap (

BTU

/lb

)

195.

94E+

0420

6.3

301.

058

664.

9

Stre

ams

ou

tm

(lb

m/h

)T

(F)

P (

psi

a)C

p (

BTU

/lb

m-F

)d

h v

ap (

BTU

/lb

)

213.

96E+

0417

4.6

24.5

0.95

6548

2.1

221.

99E+

0423

8.5

291.

011

954.

5

Mas

s Fl

ow

(lb

m/h

r)

Tem

p o

f

Tray

(F)

Re

bo

ile

r/

Co

nd

en

ser

Ou

tle

t (F

)

Tem

p

Ch

ange

(F)

Cp

inp

ut

(BTU

/lb

-F)

Cp

ou

tpu

t

(BTU

/lb

-F)

Cp

ave

rage

(BTU

/lb

-F)

Tem

p C

han

ge

Ene

rgy

(BTU

/hr)

dh

vap

(BTU

/lb

)

Ph

ase

Ch

ange

Ene

rgy

(BTU

/hr)

Tota

l En

erg

y

Pe

r U

nit

(BTU

/hr)

con

de

nse

r (1

9)3.

96E+

0417

4.80

1.75

E+02

-2.0

0E-0

11.

06E+

009.

57E-

011.

01E+

00-7

.97E

+03

4.82

E+02

1.91

E+07

1.91

E+07

reb

oil

er

(20)

1.99

E+04

234.

902.

39E+

023.

60E+

001.

06E+

001.

01E+

001.

03E+

007.

41E+

049.

55E+

021.

90E+

071.

91E+

07

TOTA

L3.

81E+

07

T-10

3

121

Sizing and Specification Sheets

Compressors:

K-104

MW 10.69 K [Cp/Cv] 1.385

Z1 1 Efficiency 75 Pt Eff 79.972

Z2 1.002 Zavg 1.001

T1 160 F 619.67 R

P1 30 psia

T2 696 F 1155.67 R

P2 180 psia

T2 Calculated 1155.159 R

Polytropic Head

R 144.565

(n-1)/n 0.347595

Flow 133891.2 lbm/hr 2231.52 lbm/min

HP 222933.1

Power 18850.52 BHP 14056.84 kW

4.86E+07 Btu/hr 1.42E+04 kW DUTY

122

Expander

K-105

MW 10.69 K [Cp/Cv] 1.381

Z1 1.001 Efficiency 75 Pt Eff 79.757

Z2 1.012 Zavg 1.0065

T1 280 F 739.67 R

P1 180 psia

T2 891.9 F 1351.57 R

P2 1015 psia


Polytropic Head

R 144.565

(n-1)/N 0.345909


HP 254837.7

Power 21606.37 BHP 16111.87 kW

5.81E+07 Btu/hr 1.70E+04 kW DUTY

123

K-100

MW 17.72 K [Cp/Cv] 1.336


Z2 0.9902 Zavg 0.9674

T1 480.4 F 940.07 R

P1 400 psia

T2 218.9 F 678.57 R

P2 30 psia


Polytropic Head

R 87.21219

(n-1)/N 0.343135

Flow 197855 lbm/hr 3297.583 lbm/min

HP -136109

NRG OUT Power -18556.7 BHP -13837.7 kW

-2.81E+07 Btu/hr -8.24E+03 kW DUTY

124

Reactor Sizing

K-102

MW 17.6 K [Cp/Cv] 1.383


Z2 0.9898 Zavg 0.9492

T1 520 F 979.67 R

P1 1015 psia

T2 185 F 644.67 R

P2 30 psia


Polytropic Head

R 87.80682

(n-1)/N 0.38348


HP -157747

NRG OUT Power -17725.4 BHP -13217.8 kW

-3.89E+07 Btu/hr -1.14E+04 kW DUTY

GHSV=SV

SV 14.29 GHSV L/D 3

57-4 53 LB/FT^3 V 4.81E+02 ft^3 Eps 0.35

Density 0.757440884 lb/ft^3 D 5.886243 ft dp 0.046280786 ft

mol flow 6867 lbmol/h L 17.65873 ft Vm 8.86E-02 ft/s

Mass flow 1.23E+05 lb/hr mu* 1.53E-02 Cp

ft^3/h 1.62E+05 ft^3/hr Dist 0.5 ft L 17.65873026 ft

Cat Volume 4.81E+02 ft^3 Cat support 0.5 ft mu 1.03E-05 lb/ft-s

Catalyst d 0.04265092 ft top 2.943122 rho 0.757440884 lb/ft^3

L 0.05577428 ft bottom 2.943122

V cat 7.96857E-05 ft^3 dP 0.015400798 psi

Cat mass 2.55E+04 lb cat L' 24.54497 ft

Cat Cost 6.71$ 1/lb cat

V 667.9268 ft^3

COST 170,892.96$ 7.41E+04

A 508.3148

Prereformer Sizing Pressure Drop

125

GHSV=SV

SV 14.29 GHSV L/D 3

57-4 53 LB/FT^3 V 9.17E+02 ft^3 Eps 0.35

Density 2.42E-02 lb/ft^3 D 7.301823 ft dp 0.046280786 ft

mol flow 1.31E+04 lbmol/h L 21.90547 ft Vm 3.44E+00 ft/s




Catalyst d 0.04265092 ft top 3.650911 rho 2.42E-02 lb/ft^3

L 0.05577428 ft bottom 3.650911

V cat 7.96857E-05 ft^3 dP 0.972322875 psi



V 1264.923 ft^3

COST 326,214.61$

A 776.6853

Steam Reformer Pressure DropSizing

GHSV=SV

SV 717.15 GHSV lbmol/hr/ft^3 L/D 3

51-7 78 LB/FT^3 V 1.45E+01 ft^3 Eps 0.35

Density 1.25E+00 lb/ft^3 D 1.833557 ft dp 0.017492276 ft

mol flow 10416 lbmol/h L 5.500671 ft Vm 5.65E-01 ft/s




Catalyst d 0.017716536 ft top 0.916778 rho 1.25E+00 lb/ft^3

L 0.017060368 ft bottom 0.916778

V cat 4.20568E-06 ft^3 dP 0.755380688 psi



V 22.00615 ft^3

COST 18,341.54$

A 53.28847

Methanol Reactor Sizing Pressure Drop

126

Pump Sizing

P-100

→ Manually Input

Flowrate (lb/hr) 1.04E+05 *sum of all 3 H Ex utility flows

Density (lb/ft3) 59.7

Flow (GPM) 218.1035

Design Flow (GPM) 272.6293

NPSH (ft) 8.5 *from fig. on pg. 6-14

Psource (psia) 29

ΔPSctn Line (psia) 1

ΔPEnt/Exit (psia) 0

PSctn (psia) 31.52396

Pdestination (psia) 50

Skirt Height (ft) 10

Feed height (ft) 0 *assume the pump is on the ground, the reflux enters back in at the top of the tower

HDSCH 10 *estimated from summing the skirt height and how high above the tower bottom the feed is

ΔPHEx (psia) 0

ΔPFE (psia) 3

ΔPCV (psia) 10

ΔPDSCH Line (psia) 12


PDSCH (psia) 79.14583

Pressure Head (ft) 114.8668

w (lb/min) 2175.788

Efficiency 0.66 *from fig on pg. 6-18

Motor Efficiency 0.9

BHP (HP) 11.47502

MHP (HP) 12.75002

Psctn Calculation

PDSCH Calculation

Power Calculation

127

P-101

→ Manually Input

Flowrate (lb/hr) 8.02E+04 *sum of all 3 H Ex utility flows

Density (lb/ft3) 45.33

Flow (GPM) 220.4497

Design Flow (GPM) 275.5622

NPSH (ft) 8.5 *from fig. on pg. 6-14

Psource (psia) 24.5

ΔPSctn Line (psia) 1


PSctn (psia) 26.17573

Pdestination (psia) 14.7

Skirt Height (ft) 10

Feed height (ft) 40 *assume the pump is on the ground, the reflux enters back in at the top of the tower

HDSCH 50 *estimated from summing the skirt height and how high above the tower bottom the feed is

ΔPHEx (psia) 0

ΔPFE (psia) 3

ΔPCV (psia) 10

ΔPDSCH Line (psia) 79.2


PDSCH (psia) 122.6396

Pressure Head (ft) 306.4371

w (lb/min) 1669.841

Efficiency 0.67 *from fig on pg. 6-18

Motor Efficiency 0.9

BHP (HP) 23.14343

MHP (HP) 25.71492

Psctn Calculation

PDSCH Calculation

Power Calculation

128

Heat Exchanger Sizing

Utility Dowtherm A Service (5-30) Vapor-vapor

IN Temp (F) 72.0 Mass Flow lbm/hr 232364 47855 U (5-30) 25.000 10 to 40

EXIT Temp (F) 500 Cp (Btu/(lb F)) 0.65426 0.548 Safety Factor 1.15

Utility Temp IN (F) 1100

Duty (MMBTU/hr) 11.22

11217971

Calculations

Utility Temp Out (F) 1026 Q-T Diagram

Area ft^3 676 Q 0 11 Q T

Utility Stream 1026 1100 Process Stream 0 72

11 500

Feed

E-100

0

200

400

600

800

1000

1200

0 2 4 6 8 10 12

T (F

)

Q (MMBTU/hr)

Utility Stream

Process Stream

129

Utility Dowtherm A Service (5-30) Vapor-vapor

IN Temp (F) 216.0 Mass Flow lbm/hr 232364 232364 U (5-30) 25.000




Calculations 58883204


Area ft^3 6233 Q 0 59 Q T


59 550

Feed

E-107

0

200

400

600

800

1000

1200

0 10 20 30 40 50 60 70

T (F

)

Q (MMBTU/hr)

Process Stream

Utility Stream

Utility Dowtherm A Service (5-30) Vapor-Liquid

IN Temp (F) 621.0 Mass Flow lbm/hr 232364 U (5-30) 50.000

EXIT Temp (F) 238 Cp (Btu/(lb F)) 0.64087 Safety Factor 1.15



Calculations


Area ft^3 4485 Q 0 57 Q T


57 238

Feed

E-101

0

100

200

300

400

500

600

700

0 10 20 30 40 50 60

T (F

)

Q (MMBTU/hr)

Utility Stream

Process Stream

130






Calculations


Area ft^3 2025 Q 0 30 Q T


30 280

Feed

E-103

0

100

200

300

400

500

600

700

0 5 10 15 20 25 30 35

T (F

)

Q (MMBTU/hr)

Utility Stream

Process Stream





Duty (MMBTU/hr) 21.736 The component that is vaporizing in stream 10 is only H2O

There are 95090 lb/hr vaporizing flow

Calculations dHvap for H2O 41 BTU/lb 3866359 BTU/hr


Area ft^3 4436 Q 0 22 Q T


22 160

Feed

E-108

0

50

100

150

200

250

300

0 5 10 15 20 25

T (F

)

Q (MMBTU/hr)

Utility Stream

Process Stream

131


IN Temp (F) 30.0 Mass Flow lbm/hr 1.16E+05 112317 U (5-30) 50.000



Duty (MMBTU/hr) 22

Calculations 21534005


Area ft^3 858 Q 0 22 Q T


22 206

Feed

E-102

0

100

200

300

400

500

600

700

800

900

0 5 10 15 20 25

T (F

)

Q (MMBTU/hr)

Utility Stream

Process Stream


IN Temp (F) 565.8 Mass Flow lbm/hr 115814 U (5-30) 50.000 dh Vapor Flow

EXIT Temp (F) 500 Cp (Btu/(lb F)) 1.05466 Safety Factor 1.15 908 6343



Calculations


Area ft^3 733 Q 0 14 Q T


14 500

Feed

E-105

0

100

200

300

400

500

600

0 2 4 6 8 10 12 14 16

T (F

)

Q (MMBTU/hr)

Utility Stream

Process Stream

132

Molar Flowrate 13108.02 lbmole/hr

Cp 13.48 BTU/lbmole-F

ΔT 680 F

Q 120130316.26 BTU/hr 120 MMBTU/hr

σ 1.71E-09 Btu/(hr-ft2-R

4)

α 0.54

F 0.7

Tt 2660 R

Tg 1760 R

Ac 4560 ft2

Area needed 5016 ft3

Fired Heater (FH-100)

𝑄 = 𝐹𝜎𝛼𝐴𝑐(𝑇𝑡4 − 𝑇𝑔

4)

Utility Dowtherm A Service (5-30)Refrigerant

IN Temp (F) 197.2 Utility Temp IN (F) -44 U (5-30) 125.000

EXIT Temp (F) 30 Duty (MMBTU/hr) 70.000 Safety Factor 1.15

Calculations

Utility Temp Out (F) -44 Q-T Diagram

Area ft^3 4557 Q 0 70 Q T

Utility Stream -44 -44 Process Stream 0 197

70 30

E-104 Refrigeration

Feed

0

10

20

30

40

50

60

70

80

90

100

-150 -100 -50 0 50 100

Cost

of

Refrig

, $/M

MB

TU

Temperature, F

Cost of Refrigeration

Effect of Temperature Level

Total Cost

Power Cost

-100

-50

0

50

100

150

200

250

0 10 20 30 40 50 60 70 80

T (F

)

Q (MMBTU/hr)

133

Separator Sizing

MRU

Streams 1 3 2

Vapor Phase Fraction 1.0 1 0

Temperature (°F) 80 80 80


Mass Flow (lb/hr) 49269 49258 11

Hg mass Absorbed (lb/hr) 11

Mass Flow of Hg x run time of absorbents (lb) 1849.43

Mercury Density (lb/ft^3) 843.60

Mass of Activated Carbon (lb) 6164.75

Activated Carbon Density (lb/ft^3) 125

Hg capacity in activated carbon 0.3

Volume of Activated Carbon (Ft^3) 49.32

L (Ft) 10.42

D (Ft) 3.47

L/D 3.00

Volume of vessel (Ft^3) 98.65

Hg absorption based on 7 days (ONLY BASED ON 1 COLUMN)

Mercury Removal Unit (MRU) - Sizing

ATT Streams

3 4 6 5

Vapor Phase Fraction 1.0 0.0 1.0 0.0

Temperature (F) 80 40 72 52

Pressure (psia) 400 400 400 400

Mass Flow (lb/hr) 49258 124245 47855 125648

Gas Flow (MMSCFD) 24.95 Sulfur Content 31557.31 lb S/day

Actual Gas Flow 618.5 CFM Circulation rate, MDEA 93.56 gpm

10.31 ft^3/s Gas Velocity 11.50 ft/s

% H2S in feed stream 1.5 L/D = 6 6

% CO2 in feed stream 1.5 Inner Diameter (D) 1.19 ft

Temperature 80 F Length (L) 7.15 ft

Pressure 385.3 psig Volume 7.98 ft^3

K (MMSCFD) 1.25 Retention Time 10 minutes

Gas Density 1.328 lbm/ft3

MDEA Density 62.428 lbm/ft3

Amine Flow Rate

1.04 g/cm^3

52190 lbm/hr

8.679221 lbm/gal

6013.212

100.2202 gpm

Properties Sizing and Calculations

In Out

134

V-100

16-2 18 17 Settling Velocity (u t ) 2.96E+00 m/s

Vapor Fraction 0.3941 0.0 1.0 Minimum vessel diameter (Dv)

Temperature [F] 30.0 30.0 30.0 1.314 m

Pressure [psia] 30.0 30.0 30.0 4.310 ft

Molar Flow [lbmole/hr] 6318 4857 1461 Archimedes Number (Ar) 176.86

Mass Flow [lb/hr] 133891 118885 15006 Reynolds Number (Re) 6.332

Drag Coefficient (Cd)

for Re > 1 3.790

Needed to size separator (Properties from hysys) for Re <1 5.879

Gas rate (Qg) 26.39 MMSCFD Actual vessel Diameter (d)

Actual Gas Flow (ACFM) 8488 CFM d2 7223

4.006 m^3/s d 7.082 ft

Gas z-factor (Z) 1.0

Gas density (ρg) 3.06E-02 lbm/ft^3

0.491 kg/m^3 L (for L/d = 3) 21.247 ft

Liquid rate (Std. Cond) 55.31 m^3/h L (for L/d = 5) 42.494 ft

Actual Liquid Flow Rate 0.017 m^3/s Volume 1 (from HYSYS) 836.900 ft3

Liquid Density 54.65 lbm/ft^3 Volume 2 (from HYSYS) 1674.000 ft3

875.41 kg/m^3

Operating pressure (P) 30 psia

Operating temperature (T) 30 F

490 R

Drop Diameter (dm) 300 um

Liquid Viscosity 0.01071 P

0.001071 kg/m-s

Gravity Constant (g) 9.81 m/s^2

Nozzle Parameters

16-2 18 17

Diameter [ft] 1.06235 1.06235 1.06235

Elevation (Base) [ft] 10.6235 21.247 0

Elevation (Ground) [ft] 10.6235 21.247 0

Ratio of Length/Diameter ~ 3-5 range

Assume Length/Diameter:

Streams Sizing and Calculations

135

V-101

Streams In

10-4 11 14 Settling Velocity (u t ) 2.48E+00 m/s

Vapor Phase Fraction 0.3252 0 1.0 Minimum vessel diameter (Dv):

Temperature [F] 160 160 160 3.335517 m

Pressure [psia] 30 30 30 10.943297 ft

Mass Flow [lb/hr] 605195 471304 133891 Archimedes Number (Ar) 1606.43

Reynolds Number (Re) 34.227392

Drag Coefficient (Cd):

for Re > 1 0.7011928

for Re <1 1.8278654

Gas rate (Qg) 113.1 MMSCFD Actual vessel Diameter (d):

Actual Gas Flow (ACFM) 4.60E+04 CFM d216108.735

21.685985 m^3/s d 10.576683 ft

Gas z-factor (Z) 1.00

Gas density (ρg) 4.82E-02 lbm/ft^3

7.72E-01 kg/m^3 L Length (L/d = 3) 31.730 ft 3

Liquid rate (Std. Cond) 42.51 m^3/h L Length (L/d = 5) 52.883 ft

Actual Liquid Flow Rate 6.03E-02 m^3/s Volume 1 (from HYSYS) 2788.000 ft3

Liquid Density 60.66 lbm/ft^3 Volume 2 (from HYSYS) 4646.000 ft3

971.68079 kg/m^3

Operating pressure (P) 30 psia

Operating temperature (T) 160 F

620 R

Drop Diameter (dm) 300 um

Liquid Viscosity 0.003944 P

0.0003944 kg/m-s

g 9.81 m/s^2

10-4 11 14

Diameter [ft] 1.5865 1.5865 1.5865

Elevation (Base) [ft] 15.865 31.73 0

Elevation (Ground) [ft] 15.865 31.73 0

Sizing and Calculations

Ratio of Height/Diameter ~ 3-5 range

Assume Length/Diameter:

Nozzle Parameters

Needed to size separator (properties from HYSYS)

Out

136

Distillation Sizing

1) Copy from Volume flow of Tower

Key Light Key Heavy

Net L

(ft^3/hr)

Net V

(ft^3/hr)

Liq Density

(lb/ft^3)

Vap Density

(lb/ft^3)

Net L

(ft^3/s)

Net V

(ft^3/s)

Total

(ft^3/s)

Liq Mass Flow

(lb/hr)

Vap Mass

Flow

(lb/hr)

Liquid flow

(m^3/hr)

Vapor flow

(m^3/hr)

Methyanol

K Value

Water

K Value

Light Viscosity

(cP)

1 3840 1784529 45.2 0.12 1.07 495.70 496.8 1.74E+05 2.11E+05 109 50532 1 1.00 0.49 0.27

2 3843 1783492 46.1 0.12 1.07 495.41 496.5 1.74E+05 2.12E+05 109 50503 2 1.00 0.47 0.27

feed 3 3845 1776351 46.1 0.12 1.07 493.43 494.5 1.74E+05 2.12E+05 109 50301 3 1.00 0.46 0.27

4 3847 1768931 46.1 0.12 1.07 491.37 492.4 1.74E+05 2.12E+05 109 50091 4 1.00 0.46 0.27

5 3849 1761558 46.1 0.12 1.07 489.32 490.4 1.74E+05 2.12E+05 109 49882 5 1.00 0.46 0.27

6 3851 1754249 46 0.12 1.07 487.29 488.4 1.74E+05 2.12E+05 109 49675 6 1.00 0.46 0.27

7 3853 1747005 46 0.12 1.07 485.28 486.3 1.74E+05 2.12E+05 109 49470 7 1.00 0.46 0.27

8 3855 1739824 46 0.12 1.07 483.28 484.4 1.74E+05 2.12E+05 109 49266 8 1.00 0.45 0.27

9 3856 1732708 46 0.12 1.07 481.31 482.4 1.74E+05 2.13E+05 109 49065 9 1.00 0.46 0.27

10 3857 1725658 46 0.12 1.07 479.35 480.4 1.74E+05 2.13E+05 109 48865 10 1.00 0.46 0.27

11 3858 1718676 46 0.12 1.07 477.41 478.5 1.74E+05 2.13E+05 109 48667 11 1.00 0.46 0.27

12 3856 1711767 45.9 0.12 1.07 475.49 476.6 1.74E+05 2.13E+05 109 48472 12 1.00 0.46 0.27

13 3852 1704941 45.9 0.12 1.07 473.59 474.7 1.74E+05 2.13E+05 109 48279 13 1.00 0.46 0.27

14 3844 1698214 45.9 0.13 1.07 471.73 472.8 1.74E+05 2.12E+05 109 48088 14 1.01 0.46 0.27

15 3827 1691619 45.9 0.13 1.06 469.89 471.0 1.73E+05 2.12E+05 108 47901 15 1.01 0.47 0.27

16 3798 1685207 45.9 0.13 1.05 468.11 469.2 1.72E+05 2.11E+05 108 47720 16 1.02 0.47 0.27

17 3749 1679062 45.9 0.13 1.04 466.41 467.4 1.70E+05 2.10E+05 106 47546 17 1.03 0.49 0.27

18 3674 1673276 45.8 0.12 1.02 464.80 465.8 1.67E+05 2.08E+05 104 47382 18 1.04 0.51 0.27

19 3572 1667863 45.8 0.12 0.99 463.30 464.3 1.63E+05 2.06E+05 101 47229 19 1.06 0.54 0.27

20 3450 1662568 45.8 0.12 0.96 461.82 462.8 1.59E+05 2.02E+05 98 47079 20 1.08 0.58 0.27

21 3321 1656752 45.8 0.12 0.92 460.21 461.1 1.54E+05 1.97E+05 94 46914 21 1.11 0.62 0.27

22 3196 1649666 45.8 0.12 0.89 458.24 459.1 1.50E+05 1.93E+05 91 46713 22 1.12 0.67 0.27

23 3085 1641037 45.8 0.11 0.86 455.84 456.7 1.46E+05 1.88E+05 87 46469 23 1.13 0.70 0.27

24 2989 1631282 45.9 0.11 0.83 453.13 454.0 1.42E+05 1.84E+05 85 46193 24 1.14 0.73 0.27

25 2909 1621102 45.9 0.11 0.81 450.31 451.1 1.39E+05 1.81E+05 82 45904 25 1.15 0.75 0.27

26 2843 1611051 46 0.11 0.79 447.51 448.3 1.37E+05 1.78E+05 80 45620 26 1.16 0.76 0.27

27 2787 1601406 46.2 0.11 0.77 444.83 445.6 1.35E+05 1.75E+05 79 45347 27 1.17 0.77 0.27

28 2740 1592231 46.4 0.11 0.76 442.29 443.0 1.33E+05 1.73E+05 78 45087 28 1.18 0.78 0.27

29 2696 1583539 46.8 0.11 0.75 439.87 440.6 1.31E+05 1.71E+05 76 44841 29 1.18 0.78 0.27

30 3710 1577473 47.2 0.11 1.03 438.19 439.2 1.81E+05 1.70E+05 105 44669 30 1.20 0.78 0.26

31 3672 1496690 47.8 0.11 1.02 415.75 416.8 1.80E+05 1.61E+05 104 42382 31 1.20 0.79 0.27

32 3642 1485375 48.4 0.11 1.01 412.60 413.6 1.79E+05 1.60E+05 103 42061 32 1.20 0.79 0.27

33 3608 1477976 49 0.11 1.00 410.55 411.6 1.78E+05 1.59E+05 102 41852 33 1.21 0.80 0.27

34 3568 1470550 49.8 0.11 0.99 408.49 409.5 1.76E+05 1.58E+05 101 41641 34 1.21 0.80 0.27

35 3520 1462753 50.8 0.11 0.98 406.32 407.3 1.74E+05 1.56E+05 100 41421 35 1.22 0.80 0.27

36 3460 1454420 52.8 0.11 0.96 404.01 405.0 1.72E+05 1.54E+05 98 41185 36 1.23 0.81 0.27

37 3381 1445315 59.2 0.11 0.94 401.48 402.4 1.69E+05 1.52E+05 96 40927 37 1.25 0.81 0.27

38 3273 1435008 59.2 0.10 0.91 398.61 399.5 1.65E+05 1.49E+05 93 40635 38 1.28 0.81 0.27

39 3108 1422443 59.2 0.10 0.86 395.12 396.0 1.59E+05 1.45E+05 88 40279 39 1.35 0.81 0.27

40 2281 1399364 59.2 0.10 0.63 388.71 389.3 1.28E+05 1.39E+05 65 39626 40 3.30 0.66 0.29

Distillation Column

K Values

137

T-100

2) Vapor and Liqud Flowrates

Max Stage 1

Vapor ACFS 495.7 Actual ft 3̂/s Largest from Total Column above

Liquid Flow 0.97 ft 3̂/s

Vapor Density 0.12 lb/ft 3̂

Liquid Density 47 lb/ft 3̂

Liquid Mass Flow 164133 lb/hr

Specific Gravity 0.75

V_Load 24.74 Actual ft 3̂/s

GPM 436.28 GPM

TS, inches TS

12 0.65

3) Correction Factors 15 0.75

18 0.84

Tray Spacing 18 inches 21 0.92

Tray Spacing Factor 0.75 24 1.00

System Non-foaming 27 1.06

System Factor 1.00 30 1.12

Flood 70 70 for new tower 36 1.15

Flood Factor 1.14 FF = 80/Flood System SF

Non Foaming 1.00

Absorbers 0.85

37.70 Actual ft^3/s Vacuum Towers 0.85

664.81 GPM Amine/Glycol/CO2/H2S 0.80

Amine/Glycol/CO2/H2S 0.60

4) Tower Diameter obtained using with Glitsch Method

Four Pass:

Vapor Load 18.85 Actual ft 3̂/s

Liquid Load 332 GPM

Diameter from 2 pass N/A feet

Diameter for 4 pass 12.3 feet

*round up to nearest half foot

N/A

N/A

12

5) Actual Stages and Length

Tray Type 4 pass

Tower Diameter 12.3 feet (chosen from above)

alpha 2 (avg K Light)/(avg K Heavy)

avg µ 0.27

Efficiency 56.12%


Actual Trays 71

Number of Feeds 1

Liquid Level Time 15 min (feed to reboiler pump)

Bot Product 20997 lb/hr http://www.cbu.edu/~rprice/lectures/distill7.html#eff

Bot Product Density 58 lb/ft 3̂

Calc Liquid Level 0.8 feet (3 feet min)

Actual Liquid Level 3 feet

Tower Sizing:

Reflux 3.0 feet

Trays and Feed 108.4 feet

Reboiler 6.0 feet

Liquid Level 3.0 feet

Tower Bottom 0.5 feet

Manways 20 feet

Total Height 141 feet

Diameter 12 feet Max Diameter is 16 feet

Height 141 feet Max Height is 200 ft

L/D 12 Avoid L/D >30

Tower Diameter (4 pass)

Corrected Vapor Load

Corrected Liquid Load




Max Stage 1








GPM 436.28 GPM

TS, inches TS

12 0.65


18 0.84







Non Foaming 1.00

Absorbers 0.85





Four Pass:


Liquid Load 332 GPM




N/A

N/A

12


Tray Type 4 pass



avg µ 0.27

Efficiency 56.12%


Actual Trays 71

Number of Feeds 1






Tower Sizing:

Reflux 3.0 feet


Reboiler 6.0 feet



Manways 20 feet










138

1) Copy from Volume flow of Tower

Key Light Key Heavy

Net L

(ft^3/hr)

Net V

(ft^3/hr)

Liq Density

(lb/ft^3)

Vap Density

(lb/ft^3)

Net L

(ft^3/s)

Net V

(ft^3/s)

Total

(ft^3/s)

Liq Mass Flow

(lb/hr)

Vap Mass

Flow

(lb/hr)

Liquid flow

(m^3/hr)

Vapor flow

(m^3/hr)

Methyanol

K Value

Water

K Value

Light Viscosity

(cP)

1 3840 1784529 45.2 0.12 1.07 495.70 496.8 1.74E+05 2.11E+05 109 50532 1 1.00 0.49 0.27

2 3843 1783492 46.1 0.12 1.07 495.41 496.5 1.74E+05 2.12E+05 109 50503 2 1.00 0.47 0.27

feed 3 3845 1776351 46.1 0.12 1.07 493.43 494.5 1.74E+05 2.12E+05 109 50301 3 1.00 0.46 0.27

4 3847 1768931 46.1 0.12 1.07 491.37 492.4 1.74E+05 2.12E+05 109 50091 4 1.00 0.46 0.27

5 3849 1761558 46.1 0.12 1.07 489.32 490.4 1.74E+05 2.12E+05 109 49882 5 1.00 0.46 0.27

6 3851 1754249 46 0.12 1.07 487.29 488.4 1.74E+05 2.12E+05 109 49675 6 1.00 0.46 0.27

7 3853 1747005 46 0.12 1.07 485.28 486.3 1.74E+05 2.12E+05 109 49470 7 1.00 0.46 0.27

8 3855 1739824 46 0.12 1.07 483.28 484.4 1.74E+05 2.12E+05 109 49266 8 1.00 0.45 0.27

9 3856 1732708 46 0.12 1.07 481.31 482.4 1.74E+05 2.13E+05 109 49065 9 1.00 0.46 0.27

10 3857 1725658 46 0.12 1.07 479.35 480.4 1.74E+05 2.13E+05 109 48865 10 1.00 0.46 0.27

11 3858 1718676 46 0.12 1.07 477.41 478.5 1.74E+05 2.13E+05 109 48667 11 1.00 0.46 0.27

12 3856 1711767 45.9 0.12 1.07 475.49 476.6 1.74E+05 2.13E+05 109 48472 12 1.00 0.46 0.27

13 3852 1704941 45.9 0.12 1.07 473.59 474.7 1.74E+05 2.13E+05 109 48279 13 1.00 0.46 0.27

14 3844 1698214 45.9 0.13 1.07 471.73 472.8 1.74E+05 2.12E+05 109 48088 14 1.01 0.46 0.27

15 3827 1691619 45.9 0.13 1.06 469.89 471.0 1.73E+05 2.12E+05 108 47901 15 1.01 0.47 0.27

16 3798 1685207 45.9 0.13 1.05 468.11 469.2 1.72E+05 2.11E+05 108 47720 16 1.02 0.47 0.27

17 3749 1679062 45.9 0.13 1.04 466.41 467.4 1.70E+05 2.10E+05 106 47546 17 1.03 0.49 0.27

18 3674 1673276 45.8 0.12 1.02 464.80 465.8 1.67E+05 2.08E+05 104 47382 18 1.04 0.51 0.27

19 3572 1667863 45.8 0.12 0.99 463.30 464.3 1.63E+05 2.06E+05 101 47229 19 1.06 0.54 0.27

20 3450 1662568 45.8 0.12 0.96 461.82 462.8 1.59E+05 2.02E+05 98 47079 20 1.08 0.58 0.27

21 3321 1656752 45.8 0.12 0.92 460.21 461.1 1.54E+05 1.97E+05 94 46914 21 1.11 0.62 0.27

22 3196 1649666 45.8 0.12 0.89 458.24 459.1 1.50E+05 1.93E+05 91 46713 22 1.12 0.67 0.27

23 3085 1641037 45.8 0.11 0.86 455.84 456.7 1.46E+05 1.88E+05 87 46469 23 1.13 0.70 0.27

24 2989 1631282 45.9 0.11 0.83 453.13 454.0 1.42E+05 1.84E+05 85 46193 24 1.14 0.73 0.27

25 2909 1621102 45.9 0.11 0.81 450.31 451.1 1.39E+05 1.81E+05 82 45904 25 1.15 0.75 0.27

26 2843 1611051 46 0.11 0.79 447.51 448.3 1.37E+05 1.78E+05 80 45620 26 1.16 0.76 0.27

27 2787 1601406 46.2 0.11 0.77 444.83 445.6 1.35E+05 1.75E+05 79 45347 27 1.17 0.77 0.27

28 2740 1592231 46.4 0.11 0.76 442.29 443.0 1.33E+05 1.73E+05 78 45087 28 1.18 0.78 0.27

29 2696 1583539 46.8 0.11 0.75 439.87 440.6 1.31E+05 1.71E+05 76 44841 29 1.18 0.78 0.27

30 3710 1577473 47.2 0.11 1.03 438.19 439.2 1.81E+05 1.70E+05 105 44669 30 1.20 0.78 0.26

31 3672 1496690 47.8 0.11 1.02 415.75 416.8 1.80E+05 1.61E+05 104 42382 31 1.20 0.79 0.27

32 3642 1485375 48.4 0.11 1.01 412.60 413.6 1.79E+05 1.60E+05 103 42061 32 1.20 0.79 0.27

33 3608 1477976 49 0.11 1.00 410.55 411.6 1.78E+05 1.59E+05 102 41852 33 1.21 0.80 0.27

34 3568 1470550 49.8 0.11 0.99 408.49 409.5 1.76E+05 1.58E+05 101 41641 34 1.21 0.80 0.27

35 3520 1462753 50.8 0.11 0.98 406.32 407.3 1.74E+05 1.56E+05 100 41421 35 1.22 0.80 0.27

36 3460 1454420 52.8 0.11 0.96 404.01 405.0 1.72E+05 1.54E+05 98 41185 36 1.23 0.81 0.27

37 3381 1445315 59.2 0.11 0.94 401.48 402.4 1.69E+05 1.52E+05 96 40927 37 1.25 0.81 0.27

38 3273 1435008 59.2 0.10 0.91 398.61 399.5 1.65E+05 1.49E+05 93 40635 38 1.28 0.81 0.27

39 3108 1422443 59.2 0.10 0.86 395.12 396.0 1.59E+05 1.45E+05 88 40279 39 1.35 0.81 0.27

40 2281 1399364 59.2 0.10 0.63 388.71 389.3 1.28E+05 1.39E+05 65 39626 40 3.30 0.66 0.29

Distillation Column

K Values

139

T-103


Max Stage 1








GPM 436.28 GPM

TS, inches TS

12 0.65


18 0.84







Non Foaming 1.00

Absorbers 0.85





Four Pass:


Liquid Load 332 GPM




N/A

N/A

12


Tray Type 4 pass



avg µ 0.27

Efficiency 56.12%


Actual Trays 71

Number of Feeds 1






Tower Sizing:

Reflux 3.0 feet


Reboiler 6.0 feet



Manways 20 feet











Max Stage 1








GPM 436.28 GPM

TS, inches TS

12 0.65


18 0.84







Non Foaming 1.00

Absorbers 0.85





Four Pass:


Liquid Load 332 GPM




N/A

N/A

12


Tray Type 4 pass



avg µ 0.27

Efficiency 56.12%


Actual Trays 71

Number of Feeds 1






Tower Sizing:

Reflux 3.0 feet


Reboiler 6.0 feet



Manways 20 feet










140

Specification sheets

Identification: Date 4/15/2017

Item No.

No. required 2 by: AKH

Function:

Operation:

Purchased Cost M$ Equipment Cost Only

Installed Cost 7701 M$ Installation, not including InDirects or Escalation/Contingency

Type:

Stage 1 2 3 4

Material of construction CS CS Component lb/hr w t frac

Compression ratio 6 5.639 Methane 5084.46 0.038

Volumetric flowrate, ACFM 45969 9151 H2O 36099.27 0.272

Pressure, psia Ethane 0.033 0.000

Inlet 30 180 CO2 11089.648 0.083

Outlet 180 1015 Nitrogen 1160.159 0.009

Design Propane 0.001 0.000

Temperature, F i-Butane 0.003 0.000

Inlet 160 280 n-Butane 0.002 0.000

Outlet 696 891.9 i-Pentane 0.002 0.000

Design n-Pentane 0.001 0.000

Cyclohexane 0.000 0.000

Power: CO 64326.30 0.484

Driver (Turbine or Motor) Hydrogen 15176.29 0.114

Brake horsepower 18850 21606

Total Power consumption kW 132936.2 1.00

Intercooling/Aftercooling:

Outlet temperature, F 280

Outlet pressure, psia 180

Utility Flowrate, lb/hr 2E+06

Duty, MMBTU/hr 4E+07

Area (ft 3̂) 2025

Mean temp. difference

Auxilaries:

Safety factor

Corrosion allowance

Utilities:

Controls:

Insulation:

Tolerances:

Comments:

COMPRESSORS

K 104, K 105

Compress streams before the reformer

Feed Composition

Turbine

30169

141


Item No. T-100/T-103

No. required 2 by: Sarah Ramzy

Function:

Operation:

Purchased Cost M$ Equipment Cost Only (from IPE)


Design: Column

Material of construction A515 Height t-t 163 ft

Pressure Inner diameter 12 ft

Max Operating 15.3 psig Yield stress

Design 30 psig Joint efficiency

Temperature Corrosion allowance 1/8 inches

Max Operating 240 F Shell thickness 0.4673 inches

Design 400 F Shell weight

Head Information

Head type Joint efficiency

Material of construction Corrosion allowance

Inner diameter Head thickness

Thickness Head weight

Yield stress

Trays

Tray Type Sieve (valve, sieve, bubble)

No. of trays 40

Material of construction A285C

Tray spacing 18 inches

Diameter 9.2 ft

Tray thickness

Weir height 2 inches

Flooding factor 70 %

No. of passes 4

Downcomer width

Pressure drop/ tray 0.150 psi

Tray Design

Sieve Valve Bubble cap

Hole area Number Number

Hole diam. Size Size

Utilities: Reboiler MPS identify utilities used (e.g. LPS, MPS, CW, -40Fref, etc)

Condenser CW

Controls:

Insulation:

Tolerances:

Comments:

DISTILLATION COLUMNS

Distillation Tower

separate PG and EG into two prodcuct streams

Notes:-Not drawn to scale-40 trays

142


Item No.


Function:

Operation:

Purchased Cost M$ Equipment Cost Estimate

Installed Cost 3150 M$ Installation estimate, not including InDirects or Escalation/Contingency

Type:

Stage 1 2 3 4

Material of construction CS Component lb/hr w t frac

Compression ratio 0.075 Methane 44219.44 0.223

Volumetric flowrate, ACFM 4406 H2O 147554.1 0.746

Pressure, psia Ethane 0.033 2E-07

Inlet 400 CO2 4799.000 0.024

Outlet 30 Nitrogen 1160.160 0.006

Design Propane 0.001 6E-09

Temperature, F i-Butane 0.003 2E-08

Inlet 480.4 n-Butane 0.002 8E-09

Outlet 218.9 i-Pentane 0.002 1E-08

Design n-Pentane 0.001 5E-09

Cyclohexane4.62E-05 2E-10

Power: Hydrogen 136.853 7E-04

Driver (Turbine or Motor)

Brake horsepower -18557

Total Power consumption 197869.6 1.000

Intercooling/Aftercooling:

Outlet temperature, F N/A

Outlet pressure, psia

Utility Flowrate, lb/hr

Duty, MMBTU/hr

Area

Mean temp. difference

Auxilaries:

Safety factor

Corrosion allowance

Utilities:

Controls:

Insulation:

Tolerances:

Comments:

Expander

Expander K 100

Expands stream after prereformer

Feed Composition

Turbine

-13837.7

143


Item No.

No. required 1 by: BK

Function:

Operation:

Purchased Cost M$ Equipment Cost Only (from IPE)


Type:

Heat exchanger type: ST BEM, etc 0.000 Duty 32.684 MMBTU/hr

Number of Shells 1 Outside area 6670 f t2

Heat Transfer Coef (U) 50 BTU/hr-ft2-F

Delta T-LM 113 F

Tube side: IN OUT Tubes:

Fluid handled Stream 10-4 Stream 10-5 Material of construction A 214

Flow rate, lb/hr Outer diameter

Liquid 0 131678

Vapor 597840 466162

Pressure, psig Thickness

Operating 15.3 15.3 Length, ft 20 from IPE

Design 15.3 No. of tubes 136 from IPE

Temperature, F Tube pattern Triangular from IPE

Operating 237.7 170 Pitch, in 1.25

Design 170 Pressure drop, psi 10

No. of Passes Fouling factor

Inlet nozzle ID Corrosion allowance

Outlet nozzle ID

Shell side: IN OUT Shell:

Fluid handled CW3 CW4 Material of construction A285C

Flow rate, lb/hr Length

Liquid N/A N/A

Vapor 0 0

Pressure, psig Inner diameter 18 from IPE

Operating 55 55 Thickness

Design 85 Total Weight

Temperature, F Shell

Operating 90 110 Full of water

Design 160 Bundle

No. of Passes Pressure drop 10

Inlet nozzle ID Fouling

Outlet nozzle ID Corrosion allowance 1/8 inches

Baffles:

Type Spacing

Material of construction Inner

Center

Outer

Utilities: Process

Controls:

Insulation:

Tolerances:

Comments: Provide Q-T Diagram

HEAT EXCHANGERS

E-108

Cooling process stream before compression

0

50

100

150

200

250

300

-35 -30 -25 -20 -15 -10 -5 0

T (F

)

Q (MMBTU/hr)

Utility Stream

Process Stream

144


Item No.


Function:

Operation:


Installed Cost 62.3 M$ Installation, not including InDirects or Escalation/Contingency

Type: Centrif Centrif or Recip

Material of construction

Volumetric flowrate, GPM 275.56

Required Head, ft 306.44

NPSH, ft 8.5

Pressure, psig

Operating 100

Design 115.3

Temperature, F

Operating

Design

Efficiency, % 75

Motor:

Driver (Turbine or Motor) Motor

Motor efficiency 90%

RPM 1800

Pump Horsepower 23.14 hP

Motor Horsepower 25.71 hP

Impeller diameter

Auxilaries:

Safety factor

Utilities:

Controls:

Insulation:

Tolerances:

Comments:

PUMPS

P-101

Pumps liquid from distillate to product stream

Carbon Steel

145


Item No.

No. required 1 by: BK

Function:

Operation:



Design Data:

Vessel

Material of construction Height t-t 41 ft

Shape Inner diameter 30 ft

Configuration H H or V Thickness 0.25 in (from IPE)

Pressure, psig Volume 3185 ft3

Operating 15 % full N/A

Design 45

Temperature, F

Operating 1100

Design 1150

Heads

Material of construction Joint efficiency

Head type Corrosion allowance

Inner diameter Weight

Thickness

Stress values

Auxilaries

Platforms

Ladders

Inlet location

Outlet location

Mist separator

Utilities:

Controls:

Insulation:

Tolerances:

Comments: Provide attached sketch.

PRESSURE VESSELS

Steam Reformer

Reform remaining hydrocarbons

CS W/ refractory lining

146

ASPEN PEA Sizing Information

ITEM REPORT

Processing Date : Mon Apr 17 01:16:09 PM 2017

Version : Aspen Process Economic Analyzer 34.1.0(Build 3457)

List of Items :

Project: base case

Scenario: Base Case

Pre-reformer

Project : BASECASE

Scenario : BASECASE

Pre-reformer

Item Code: DTW PACKED

Internal Name : DTW PACKED #1#PFR

Sizing Data

Design Data

Summary Costs

Sizing Data

Description Value Units

Design Data

Parameter Value Units

Item type PACKED

Number of identical items 1

EQUIPMENT DESIGN DATA

Application ABSORB

ASME design basis D1NF

Liquid volume 8300.000 GALLONS

147

Design gauge pressure 425.000 PSIG

Design temperature 530.000 DEG F

Operating temperature 480.000 DEG F

GENERAL DESIGN DATA

Number of platforms 2

COLUMN DATA

Shell material CS

Diameter option ID

Vessel diameter 7.0000 FEET

Vessel tangent to tangent height 29.000 FEET

Head type ELLIP

MECHANICAL DESIGN DATA

Wind or seismic design W+S

Fluid volume 20.000 PERCENT

Weld efficiency 100.000 PERCENT

Base material thickness 1.3750 INCHES

Corrosion allowance 0.125000 INCHES

THICKNESSES REQUIRED

Circumferential stress thickness 1.1735 INCHES

Long tensile stress thickness 0.482159 INCHES

Long compressive stress thick 0.129581 INCHES

PACKING DATA

Number of distributor plates 4

Number of packed sections 2

Section height 10.500 FEET

Cross sectional area 38.485 SF

SECTION 1

Total packing height 21.000 FEET

Packing volume 808.175 CF

148

Packing volume per unit height 38.485 CF/FT

VESSEL SKIRT DATA

Skirt material CS

Skirt height 11.000 FEET

Skirt thickness 0.500000 INCHES

NOZZLE AND MANHOLE DATA

Nozzle ASA rating 300 CLASS

Nozzle material A285C

Nozzle A Quantity 1

Nozzle A Diameter 8.0000 IN DIAM

Nozzle A Location S

Nozzle B Quantity 1

Nozzle B Diameter 6.0000 IN DIAM

Nozzle B Location S

Nozzle C Quantity 1

Nozzle C Diameter 10.000 IN DIAM

Nozzle C Location S

Nozzle D Quantity 1

Nozzle D Diameter 12.000 IN DIAM

Nozzle D Location S

Nozzle E Quantity 5

Nozzle E Diameter 2.0000 IN DIAM

Nozzle E Location S

Number of manholes 3

Manhole diameter 30.000 INCHES

PROCESS DESIGN DATA

Molecular weight Overhead prod 30.000

WEIGHT DATA

Shell 36400 LBS

149

Trays and supports 2600 LBS

Heads 7400 LBS

Nozzles 590 LBS

Manholes and Large nozzles 5100 LBS

Skirt 5000 LBS

Base ring and lugs 770 LBS

Ladder clips 110 LBS

Platform clips 250 LBS

Fittings and miscellaneous 160 LBS

Total weight less packing 58400 LBS

VENDOR COST DATA

Material cost 65375 DOLLARS

Shop labor cost 23405 DOLLARS

Shop overhead cost 24106 DOLLARS

Office overhead cost 19191 DOLLARS

Profit 19823 DOLLARS

Total cost 151900 DOLLARS

Cost per unit weight 2.6010 USD/LBS

Cost per unit height or length 7233.333 USD/FT

Cost per unit volume 187.954 USD/CF

Cost per unit area 3947.043 USD/SF

Summary Costs

Item Material(USD) Manpower(USD) Manhours

Equipment&Setting 151900. 4355. 138

Piping 42440. 20064. 654

Civil 3741. 4125. 165

Structural Steel 14479. 2482. 86

150

Instrumentation 28440. 11870. 377

Electrical 1887. 1102. 37

Insulation 18818. 15190. 652

Paint 1185. 2707. 118

Subtotal 262890 61895 2227

Total material and manpower cost=USD 324800.

ITEM REPORT



List of Items :

Project : basecase

Scenario : Base Case

SMR

Project : BASECASE

Scenario : BASECASE

SMR



Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type PACKED

151



Application ABSORB






GENERAL DESIGN DATA


COLUMN DATA

Shell material CS

Diameter option ID



Head type HEMI











PACKING DATA


Number of packed sections 2

152

Section height 15.000 FEET


SECTION 1




VESSEL SKIRT DATA

Skirt material CS






Nozzle A Quantity 1


Nozzle A Location S

Nozzle B Quantity 1


Nozzle B Location S

Nozzle C Quantity 1


Nozzle C Location S

Nozzle D Quantity 1


Nozzle D Location S

Nozzle E Quantity 6


Nozzle E Location S


153


PROCESS DESIGN DATA


WEIGHT DATA

Shell 19800 LBS


Heads 4000 LBS

Nozzles 770 LBS


Skirt 7300 LBS






VENDOR COST DATA











Summary Costs

154



Piping 67272. 25598. 836

Civil 6391. 6035. 242




Insulation 35083. 26642. 1144

Paint 1628. 3607. 157

Subtotal 317338 83362 3074


ITEM REPORT



List of Items :

Project : basecase


MeOH Synth

Project : BASECASE

Scenario : BASECASE

MeOH Synth



Sizing Data

Design Data

Summary Costs

Sizing Data


155

Design Data


Item type PACKED



Application ABSORB






GENERAL DESIGN DATA


COLUMN DATA

Shell material CS

Diameter option ID



Head type ELLIP








PACKING DATA

156



SECTION 1




VESSEL SKIRT DATA

Skirt material CS






Nozzle A Quantity 5


Nozzle A Location S

Nozzle B Quantity 1


Nozzle B Location S

Nozzle C Quantity 1


Nozzle C Location S

Nozzle D Quantity 1


Nozzle D Location S



PROCESS DESIGN DATA


157

WEIGHT DATA

Shell 2100 LBS


Heads 480 LBS

Nozzles 160 LBS


Skirt 490 LBS


Ladder clips 40 LBS



VENDOR COST DATA





Profit 4177 DOLLARS






Summary Costs



Piping 12799. 14254. 463

Civil 1012. 1319. 54

158




Insulation 5701. 4712. 202

Paint 581. 1460. 64

Subtotal 81160 35911 1236


ITEM REPORT



List of Items :

Project : basecase


T-100

Project : BASECASE

Scenario : BASECASE

T-100

Item Code: DTW TRAYED

Internal Name : DTW TRAYED dist

Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


159

Item type TRAYED



Tray type SIEVE

Application DISTIL





GENERAL DESIGN DATA


COLUMN DATA

Shell material A285C

Diameter option ID



Head type ELLIP





Thickness Average 0.395833 INCHES






TRAY STACK DATA

Tray material A285C

160

Number of trays 40

Tray spacing 15.000 INCHES

Tray thickness 0.187500 INCHES


VESSEL SKIRT DATA

Skirt material CS






Nozzle A Quantity 3


Nozzle A Location S

Nozzle B Quantity 1


Nozzle B Location S

Nozzle C Quantity 1


Nozzle C Location S

Nozzle D Quantity 1


Nozzle D Location S

Nozzle E Quantity 10


Nozzle E Location S

Nozzle F Quantity 1

Nozzle F Diameter 8.0000 IN DIAM

Nozzle F Location S

161



PROCESS DESIGN DATA


WEIGHT DATA

Shell 86200 LBS


Heads 5700 LBS

Nozzles 1100 LBS


Skirt 21100 LBS





Total weight 206300 LBS

VENDOR COST DATA

Cost per unit Trays and supports 3388 DOLLARS

Labor per unit Trays and support 39 MH/TRAY








Cost per tray 15165 DOLLARS

Cost per unit height or length 4302 USD/FT

162

Summary Costs



Piping 106409. 40506. 1328

Civil 15600. 13239. 525



Electrical 6122. 3174. 106

Insulation 50340. 45775. 1967

Paint 2801. 5980. 260

Subtotal 924058 171779 6212


ITEM REPORT



List of Items :

Project : basecase


T-100

Project : BASECASE

Scenario : BASECASE

T-103

Item Code: DTW TRAYED

Internal Name : DTW TRAYED dist

Sizing Data

Design Data

Summary Costs

Sizing Data

163


Design Data


Item type TRAYED



Tray type SIEVE

Application DISTIL





GENERAL DESIGN DATA


COLUMN DATA

Shell material A285C

Diameter option ID



Head type ELLIP





Thickness Average 0.395833 INCHES



164




TRAY STACK DATA

Tray material A285C

Number of trays 40

Tray spacing 15.000 INCHES

Tray thickness 0.187500 INCHES


VESSEL SKIRT DATA

Skirt material CS






Nozzle A Quantity 3


Nozzle A Location S

Nozzle B Quantity 1


Nozzle B Location S

Nozzle C Quantity 1


Nozzle C Location S

Nozzle D Quantity 1


Nozzle D Location S

Nozzle E Quantity 10

165


Nozzle E Location S

Nozzle F Quantity 1

Nozzle F Diameter 8.0000 IN DIAM

Nozzle F Location S



PROCESS DESIGN DATA


WEIGHT DATA

Shell 86200 LBS


Heads 5700 LBS

Nozzles 1100 LBS


Skirt 21100 LBS






VENDOR COST DATA

Cost per unit Trays and supports 3388 DOLLARS

Labor per unit Trays and support 39 MH/TRAY






166



Cost per tray 15165 DOLLARS

Cost per unit height or length 4302 USD/FT

Summary Costs



Piping 106409. 40506. 1328

Civil 15600. 13239. 525



Electrical 6122. 3174. 106

Insulation 50340. 45775. 1967

Paint 2801. 5980. 260

Subtotal 924058 171779 6212


ITEM REPORT



List of Items :

Project : basecase


E-100

Project : BASECASE

Scenario : BASECASE

167

E-100

Item Code: DHE U TUBE

Internal Name : DHE U TUBE he

Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type U TUBE


GENERAL DESIGN DATA

TEMA type BEU

Heat exchanger design option STAND

Heat exchanger design+cost tool ECON

Heat transfer area 748.000 SF

Number of shells 1

Number of tube passes 2

Number of shell passes 1

Vendor grade HIGH

SHELL DATA

Shell material SS304

Shell diameter 18.000 INCHES

Shell length 20.000 FEET

Shell design gauge pressure 150.000 PSIG

Shell design temperature 650.000 DEG F

Shell operating temperature 650.000 DEG F

168

Shell corrosion allowance 0.0 INCHES

Shell wall thickness 0.188000 INCHES

ASA rating Shell side 300 CLASS

Number of baffles 16

Shell fabrication type PIPE

Expansion joint NO

TUBE DATA

Tube material 304LW

Number of tubes per shell 72

Tube outside diameter 1.0000 INCHES

Tube length extended 40.000 FEET

Tube design gauge pressure 150.000 PSIG

Tube design temperature 650.000 DEG F

Tube operating temperature 650.000 DEG F

Tube corrosion allowance 0.0 INCHES

Tube wall thickness 0.049000 INCHES

Tube gauge 18 BWG

Tube pitch symbol TRIANGULAR

Tube pitch 1.2500 INCHES

Tube seal type SEALW

TUBE SHEET DATA

Tube sheet material 304L

Tube sheet thickness 1.5000 INCHES

Tube sheet corrosion allowance 0.0 INCHES

Channel material 304L

TUBE SIDE HEAD DATA

Head material Tube side 304L

ASA rating Tube side 300 CLASS

Head thickness Tube side 0.250000 INCHES

169

SHELL SIDE HEAD DATA

Head material Shell side SS304


Head thickness Shell side 0.250000 INCHES

WEIGHT DATA

Shell 730 LBS

Tubes 1500 LBS

Heads 70 LBS

Internals and baffles 220 LBS

Nozzles 180 LBS

Flanges 340 LBS


Tube sheet 100 LBS

Saddles 80 LBS



VENDOR COST DATA





Profit 4038 DOLLARS




Summary Costs


170


Piping 75514. 19461. 634

Civil 1157. 1645. 67



Electrical 0. 0. 0

Insulation 11880. 7622. 327

Paint 0. 0. 0

Subtotal 134354 33873 1194


ITEM REPORT



List of Items :

Project : basecase


E-107

Project : BASECASE

Scenario : BASECASE

E-107


Internal Name : DHE U TUBE #11#he

Sizing Data

Design Data

Summary Costs

Sizing Data


171

Design Data


Item type U TUBE


GENERAL DESIGN DATA

TEMA type BEU




Number of shells 1



Vendor grade HIGH

SHELL DATA

Shell material CS










Shell fabrication type PLATE

Expansion joint NO

TUBE DATA

Tube material A 214

172









Tube gauge 16 BWG




TUBE SHEET DATA

Tube sheet material A 515



Channel material A 515

TUBE SIDE HEAD DATA

Head material Tube side A 515




Head material Shell side CS



WEIGHT DATA

Shell 9600 LBS

Tubes 35000 LBS

Heads 1100 LBS

173


Nozzles 1400 LBS

Flanges 3000 LBS


Tube sheet 1600 LBS

Saddles 560 LBS



VENDOR COST DATA









Summary Costs



Piping 110623. 26432. 863

Civil 2495. 2709. 109



Electrical 0. 0. 0

Insulation 26907. 16337. 701

Paint 1164. 2934. 127

174

Subtotal 317953 54280 1991


ITEM REPORT



List of Items :

Project : basecase


E-101

Project : BASECASE

Scenario : BASECASE

E-101



Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type U TUBE


GENERAL DESIGN DATA

TEMA type BEU


175



Number of shells 1



Vendor grade HIGH

SHELL DATA

Shell material CS











Expansion joint NO

TUBE DATA

Tube material A 214









176

Tube gauge 16 BWG




TUBE SHEET DATA





TUBE SIDE HEAD DATA








WEIGHT DATA

Shell 6200 LBS

Tubes 19300 LBS

Heads 550 LBS


Nozzles 790 LBS

Flanges 1700 LBS


Tube sheet 720 LBS

Saddles 340 LBS



177

VENDOR COST DATA





Profit 8511 DOLLARS




Summary Costs



Piping 83602. 23015. 751

Civil 1895. 2246. 91



Electrical 0. 0. 0

Insulation 21883. 13716. 588

Paint 1047. 2643. 114

Subtotal 221870 47043 1720


ITEM REPORT



List of Items :

Project : basecase


E-108

178

Project : BASECASE

Scenario : BASECASE

E-108



Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type U TUBE


GENERAL DESIGN DATA

TEMA type BEU




Number of shells 1



Vendor grade HIGH

SHELL DATA

Shell material CS


179










Expansion joint NO

TUBE DATA

Tube material A 214









Tube gauge 16 BWG




TUBE SHEET DATA





180

TUBE SIDE HEAD DATA








WEIGHT DATA

Shell 5800 LBS

Tubes 16600 LBS

Heads 400 LBS


Nozzles 790 LBS

Flanges 1500 LBS


Tube sheet 580 LBS

Saddles 300 LBS



VENDOR COST DATA





Profit 7874 DOLLARS




181

Summary Costs



Piping 83602. 23015. 751

Civil 1784. 2156. 87



Electrical 0. 0. 0

Insulation 21592. 13461. 578

Paint 1047. 2643. 114

Subtotal 211168 46698 1706


ITEM REPORT



List of Items :

Project : basecase


E-103

Project : BASECASE

Scenario : BASECASE

E-103



Sizing Data

Design Data

Summary Costs

182

Sizing Data


Design Data


Item type U TUBE


GENERAL DESIGN DATA

TEMA type BEU




Number of shells 1



Vendor grade HIGH

SHELL DATA

Shell material CS











183

Expansion joint NO

TUBE DATA

Tube material A 214









Tube gauge 16 BWG




TUBE SHEET DATA





TUBE SIDE HEAD DATA








WEIGHT DATA

184

Shell 2900 LBS

Tubes 6400 LBS

Heads 240 LBS


Nozzles 390 LBS

Flanges 920 LBS


Tube sheet 290 LBS

Saddles 190 LBS



VENDOR COST DATA





Profit 5227 DOLLARS




Summary Costs



Piping 61697. 19159. 624

Civil 1432. 1862. 75



185

Electrical 0. 0. 0

Insulation 17319. 10428. 447

Paint 905. 2304. 100

Subtotal 143686 38898 1412


ITEM REPORT



List of Items :

Project : basecase


E-102

Project : BASECASE

Scenario : BASECASE

E-102



Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type U TUBE


186

GENERAL DESIGN DATA

TEMA type BEU




Number of shells 1



Vendor grade HIGH

SHELL DATA












Expansion joint NO

TUBE DATA

Tube material A 214






187




Tube gauge 16 BWG




TUBE SHEET DATA





TUBE SIDE HEAD DATA





Head material Shell side SS304



WEIGHT DATA

Shell 940 LBS

Tubes 2500 LBS

Heads 80 LBS


Nozzles 170 LBS

Flanges 410 LBS


Tube sheet 100 LBS

188

Saddles 90 LBS



VENDOR COST DATA





Profit 3495 DOLLARS




Summary Costs



Piping 65279. 18964. 618

Civil 1225. 1708. 69



Electrical 0. 0. 0

Insulation 14562. 8526. 366

Paint 438. 1106. 48

Subtotal 120610 35449 1267


ITEM REPORT



189

List of Items :

Project : basecase


E-105

Project : BASECASE

Scenario : BASECASE

E-105



Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type U TUBE


GENERAL DESIGN DATA

TEMA type BEU




Number of shells 1



190

Vendor grade HIGH

SHELL DATA

Shell material CS











Expansion joint NO

TUBE DATA

Tube material A 214









Tube gauge 16 BWG




TUBE SHEET DATA

191





TUBE SIDE HEAD DATA








WEIGHT DATA

Shell 1000 LBS

Tubes 1300 LBS

Heads 70 LBS


Nozzles 180 LBS

Flanges 280 LBS


Tube sheet 60 LBS

Saddles 70 LBS



VENDOR COST DATA





192

Profit 2356 DOLLARS




Summary Costs



Piping 30881. 14093. 458

Civil 1131. 1620. 66



Electrical 0. 0. 0

Insulation 11434. 7213. 309

Paint 622. 1603. 69

Subtotal 72134 29543 1064


ITEM REPORT



List of Items :

Project : basecase


E-104

Project : BASECASE

Scenario : BASECASE

193

E-104



Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type U TUBE


GENERAL DESIGN DATA

TEMA type BEU




Number of shells 1



Vendor grade HIGH

SHELL DATA

Shell material CS






194






Expansion joint NO

TUBE DATA

Tube material A 214









Tube gauge 16 BWG




TUBE SHEET DATA





TUBE SIDE HEAD DATA




195





WEIGHT DATA

Shell 4000 LBS

Tubes 11700 LBS

Heads 420 LBS


Nozzles 790 LBS

Flanges 1600 LBS


Tube sheet 610 LBS

Saddles 310 LBS



VENDOR COST DATA





Profit 7079 DOLLARS




Summary Costs


196


Piping 61697. 19159. 624

Civil 1821. 2186. 88



Electrical 0. 0. 0

Insulation 18087. 11210. 481

Paint 905. 2304. 100

Subtotal 171143 40004 1459


ITEM REPORT



List of Items :

Project : basecase


E-109

Project : BASECASE

Scenario : BASECASE

E-109



Sizing Data

Design Data

Summary Costs

Sizing Data


197

Design Data


Item type U TUBE


GENERAL DESIGN DATA

TEMA type BEU




Number of shells 1



Vendor grade HIGH

SHELL DATA

Shell material CS











Expansion joint NO

TUBE DATA

Tube material A 214

198









Tube gauge 16 BWG




TUBE SHEET DATA





TUBE SIDE HEAD DATA








WEIGHT DATA

Shell 1700 LBS

Tubes 2900 LBS

Heads 110 LBS

199


Nozzles 250 LBS

Flanges 450 LBS


Tube sheet 100 LBS

Saddles 100 LBS



VENDOR COST DATA





Profit 3064 DOLLARS




Summary Costs



Piping 45190. 16488. 537

Civil 1272. 1751. 71



Electrical 0. 0. 0

Insulation 14639. 8619. 370

Paint 769. 1965. 85

200

Subtotal 97757 33968 1229


ITEM REPORT



List of Items :

Project : basecase


Reboiler

Project : BASECASE

Scenario : BASECASE

Reboiler


Internal Name : DHE U TUBE #1#E-109

Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type U TUBE


GENERAL DESIGN DATA

TEMA type BEU


201



Number of shells 1



Vendor grade HIGH

SHELL DATA

Shell material CS











Expansion joint NO

TUBE DATA

Tube material A 214









202

Tube gauge 16 BWG




TUBE SHEET DATA





TUBE SIDE HEAD DATA








WEIGHT DATA

Shell 4200 LBS

Tubes 13100 LBS

Heads 550 LBS


Nozzles 790 LBS

Flanges 1700 LBS


Tube sheet 710 LBS

Saddles 340 LBS



203

VENDOR COST DATA





Profit 7522 DOLLARS




Summary Costs



Piping 61697. 19159. 624

Civil 1895. 2246. 91



Electrical 0. 0. 0

Insulation 18233. 11350. 487

Paint 905. 2304. 100

Subtotal 178263 40204 1468


ITEM REPORT



List of Items :

Project : basecase


Condenser

204

Project : BASECASE

Scenario : BASECASE

Condenser


Internal Name : DHE U TUBE #2#E-109

Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type U TUBE


GENERAL DESIGN DATA

TEMA type BEU




Number of shells 1



Vendor grade HIGH

SHELL DATA

Shell material CS


205










Expansion joint NO

TUBE DATA

Tube material A 214









Tube gauge 16 BWG




TUBE SHEET DATA





206

TUBE SIDE HEAD DATA








WEIGHT DATA

Shell 4400 LBS

Tubes 14600 LBS

Heads 630 LBS


Nozzles 880 LBS

Flanges 2000 LBS


Tube sheet 890 LBS

Saddles 380 LBS



VENDOR COST DATA





Profit 8056 DOLLARS




207

Summary Costs



Piping 81462. 21917. 714

Civil 2009. 2336. 94



Electrical 0. 0. 0

Insulation 20200. 12517. 537

Paint 1009. 2561. 111

Subtotal 208013 44476 1622


ITEM REPORT



List of Items :

Project : basecase


FH-100

Project : BASECASE

Scenario : BASECASE

FH-100

Item Code: EFU HEATER

Internal Name : EFU HEATER fe

Sizing Data

Design Data

Summary Costs

208

Sizing Data


Design Data


Item type HEATER

Material 304LW

Duty 183.000 MMBTU/H



Standard gas flow rate 26213.00 CFM

Source of quote SG

Process type GAS


Summary Costs



Piping 187676. 45683. 1490

Civil 40994. 24658. 996



Electrical 30606. 9105. 301

Insulation 34151. 46738. 2009

Paint 1130. 2348. 103

Subtotal 3552337 230272 8395

209


ITEM REPORT



List of Items :

Project : basecase


MRU

Project : BASECASE

Scenario : BASECASE

MRU

Item Code: DVT CYLINDER

Internal Name : DVT CYLINDER MRU

Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type CYLINDER






210



Fluid specific gravity 1.0000

SHELL DATA


Diameter option ID



Head type ELLIP






Head thickness Top 0.625000 INCHES

Head thickness Bottom 0.625000 INCHES



VESSEL LEG DATA

Leg material CS

Vessel leg height 4.0000 FEET



Nozzle material SS304

Nozzle A Quantity 1


Nozzle A Location S

Nozzle B Quantity 1


211

Nozzle B Location S

Nozzle C Quantity 1


Nozzle C Location S

Nozzle D Quantity 5


Nozzle D Location S


Manhole diameter 18 INCHES

WEIGHT DATA

Shell 4400 LBS

Heads 1200 LBS

Nozzles 160 LBS



Ladder clips 40 LBS


Structural steel 160 LBS



VENDOR COST DATA





Profit 8149 DOLLARS



Cost per unit liquid volume 51.455 USD/GALL

212

Summary Costs



Piping 41596. 17000. 552

Civil 1990. 2104. 86




Insulation 4796. 4791. 206

Paint 146. 221. 10

Subtotal 147935 32150 1115


ITEM REPORT



List of Items :

Project : basecase


ATU

Project : BASECASE

Scenario : BASECASE

ATU


Internal Name : DVT CYLINDER #3#MRU

Sizing Data

Design Data

Summary Costs

213

Sizing Data


Design Data


Item type CYLINDER









SHELL DATA

Shell material CS

Diameter option ID



Head type ELLIP









214


CLADDING DATA

Cladding material SS304

Cladding thickness 0.500000 INCHES

VESSEL LEG DATA

Leg material CS





Nozzle A Quantity 5


Nozzle A Location S

Nozzle B Quantity 1


Nozzle B Location S

Nozzle C Quantity 1


Nozzle C Location S



WEIGHT DATA

Shell 2000 LBS

Heads 430 LBS

Nozzles 280 LBS



Ladder clips 40 LBS


215



VENDOR COST DATA





Profit 4537 DOLLARS




Summary Costs



Piping 21596. 16574. 537

Civil 958. 1271. 52




Insulation 6131. 4620. 198

Paint 108. 167. 7

Subtotal 95787 30263 1041


ITEM REPORT



List of Items :

216

Project : basecase


V-101

Project : BASECASE

Scenario : BASECASE

V-101



Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type CYLINDER









SHELL DATA

Shell material CS

217

Diameter option ID



Head type HEMI













VESSEL SKIRT DATA

Skirt material CS






Nozzle A Quantity 1


Nozzle A Location S

Nozzle B Quantity 1


Nozzle B Location S

218

Nozzle C Quantity 1


Nozzle C Location S

Nozzle D Quantity 1


Nozzle D Location S

Nozzle E Quantity 6


Nozzle E Location S



WEIGHT DATA

Shell 18600 LBS

Heads 4900 LBS

Nozzles 580 LBS


Skirt 7500 LBS






VENDOR COST DATA







219



Summary Costs



Piping 63770. 22903. 747

Civil 7494. 6779. 272




Insulation 17402. 15550. 668

Paint 1532. 3431. 149

Subtotal 282713 62734 2297


ITEM REPORT



List of Items :

Project : basecase


V-100

Project : BASECASE

Scenario : BASECASE

V-100



220

Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type CYLINDER









SHELL DATA

Shell material CS

Diameter option ID



Head type ELLIP






221








VESSEL LEG DATA

Leg material CS





Nozzle A Quantity 2


Nozzle A Location S

Nozzle B Quantity 1


Nozzle B Location S

Nozzle C Quantity 1


Nozzle C Location S

Nozzle D Quantity 6


Nozzle D Location S



WEIGHT DATA

Shell 8500 LBS

222

Heads 1500 LBS

Nozzles 300 LBS








VENDOR COST DATA





Profit 6341 DOLLARS




Summary Costs



Piping 41647. 18323. 597

Civil 3599. 4015. 160




Insulation 0. 0. 0

223

Paint 2806. 5284. 232

Subtotal 174768 37405 1311


ITEM REPORT



List of Items :

Project : basecase


K 104 IG

Project : BASECASE

Scenario : BASECASE

K 104 IG

Item Code: DGC CENTRIF IG

Internal Name : DGC CENTRIF IG#1#K 104 IG

Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type CENTRIF IG


PROCESS DESIGN DATA

Gas type option AIR

224

Molecular weight 10.690

Specific heat ratio 1.4000

Compressibility factor Inlet 1.0000

Compressibility factor Outlet 1.0000

Flow rate 133920.00 LB/H

Maximum interstage temperature 270.000 DEG F

Intercooler outlet temperature 90.000 DEG F

OPERATING CONDITIONS

SUCTION

Design gauge pressure Inlet 45.300 PSIG

Design temperature Inlet 68.000 DEG F

Actual gas flow rate Inlet 19695.70 CFM

DISCHARGE

Design gauge pressure Outlet 195.300 PSIG

Design temperature Outlet 250.227 DEG F

Actual gas flow rate Outlet 7781.000 CFM


Casing material CS

Polytropic head 124717.59 FEET

Number of impellers 2

Shop assembly option FULL

Intercooler required YES

Aftercooler required NO

Frame size MSG9

Driver power 18850.00 HP

Driver speed 1200.000 RPM

Motor type ODP

Driver type MOTOR

Gear reducer type YES

225

Lube oil system YES

GENERAL DESIGN DATA

Power per gas flow rate 0.957062 HP/CFM

Power per stage 9425.000 HP/STAGE

Power per liquid flow rate 7.1045 LB/H/HP

Differential pressure 150.000 PSIG

Pressure ratio Output to Input 3.5000 P2/P1

WEIGHT DATA

Motor 113100 LBS

Compressor 37900 LBS


VENDOR COST DATA

Motor cost 708869 DOLLARS

Lube and seal oil cost 605590 DOLLARS



Cost per unit gas flow rate 112.705 USD/CFM

Cost per unit power 117.761 USD/HP

Cost per unit flow rate 59672.04 USD/LB/H

Cost per stage 1109900. DOLLARS

Summary Costs



Piping 120669. 33542. 1089

Civil 39807. 22602. 901



226

Electrical 143936. 64836. 2152

Insulation 8494. 6292. 270

Paint 2111. 4530. 197

Subtotal 2596260 188970 6431


ITEM REPORT



List of Items :

Project : basecase


K 105 H

Project : BASECASE

Scenario : BASECASE

K 105 H

Item Code: DGC CENTRIF

Internal Name : DGC CENTRIF #1#K 105 H

Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type CENTRIF


227

PROCESS DESIGN DATA




Compressibility factor Outlet 1.0000

Flow rate 133920.00 LB/H

Maximum interstage temperature 350.000 DEG F

Intercooler outlet temperature 90.000 DEG F

Interstage pressure drop 5.0000 PSIG

OPERATING CONDITIONS

SUCTION




DISCHARGE


Design temperature Outlet 215.117 DEG F

Actual gas flow rate Outlet 1354.000 CFM


Casing material CS

Polytropic head 155830.00 FEET

Number of impellers 12

Compressor speed 17291.00 RPM

Frame size 15MB

Number of stages 2

Number of nozzles 4

ASA rating Inlet 150 CLASS

Inlet diameter 14.000 IN DIAM

ASA rating Outlet 600 CLASS

228

Outlet diameter 8.0000 INCHES


Driver speed 1800.000 RPM

Motor type ODP

Driver type MOTOR

Gear reducer type YES

Lube oil system YES

INTERCOOLER(S), AFTERCOOLER

Intercooler required NO

Aftercooler required NO

GENERAL DESIGN DATA

Power per gas flow rate 3.8400 HP/CFM

Power per stage 10803.18 HP/STAGE

Power per liquid flow rate 6.1982 LB/H/HP

Differential pressure 906.500 PSIG

Pressure ratio Output to Input 5.3167 P2/P1

WEIGHT DATA

Flanges 230 LBS

Gear reducer 46400 LBS

Motor 42200 LBS

Compressor 20500 LBS



VENDOR COST DATA

MATERIAL COST INCLUDES DRIVER


Lube and seal oil cost 801864 DOLLARS

Testing cost 221499 DOLLARS

Fabrication labor 15518 HOURS

229








Cost per unit gas flow rate 795.913 USD/CFM


Cost per unit flow rate 120384.41 USD/LB/H

Cost per stage 373191.66 DOLLARS

Summary Costs



Piping 58150. 18171. 591

Civil 17575. 10074. 402



Electrical 144437. 65135. 2162

Insulation 3938. 3435. 147

Paint 1202. 2667. 116

Subtotal 4751634 164498 5490


ITEM REPORT



List of Items :

230

Project : basecase


P 100

Project : BASECASE

Scenario : BASECASE

P 100

Item Code: DCP API 610

Internal Name : DCP API 610 #1#P 100

Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type API 610



Casing material CS



Fluid head 225.000 FEET

ASA rating 150 CLASS


Speed 3600.000 RPM

Driver type MOTOR

231

Motor type TEFC

Pump efficiency 75.000 PERCENT

SEAL DATA

Seal type SNGL

Primary seal pipe plan 11

Pipe plan pipe type WELD

Pipe plan material type A 106

PROCESS DESIGN DATA

Liquid flow rate 272.630 GPM


Fluid viscosity 1.0000 CPOISE

Power per liquid flow rate 0.042072 HP/GPM

Liquid flow rate times head 61341 GPM -FT

WEIGHT DATA

Pump 320 LBS

Motor 230 LBS

Base plate 70 LBS



VENDOR COST DATA






Profit 3283 DOLLARS



Cost per unit liquid flow rate 86.197 USD/GPM

232


Summary Costs



Piping 12722. 6268. 204

Civil 305. 751. 31




Insulation 0. 0. 0

Paint 629. 1311. 58

Subtotal 47927 13102 448


ITEM REPORT



List of Items :

Project : basecase


P 101

Project : BASECASE

Scenario : BASECASE

P 101

Item Code: DCP API 610

Internal Name : DCP API 610 #1#P 101

Sizing Data

233

Design Data

Summary Costs

Sizing Data


Design Data


Item type API 610



Casing material CS



Fluid head 225.000 FEET

ASA rating 150 CLASS


Speed 3600.000 RPM

Driver type MOTOR

Motor type TEFC

Pump efficiency 75.000 PERCENT

SEAL DATA

Seal type SNGL

Primary seal pipe plan 11

Pipe plan pipe type WELD

Pipe plan material type A 106

PROCESS DESIGN DATA

Liquid flow rate 275.560 GPM


234

Fluid viscosity 1.0000 CPOISE

Power per liquid flow rate 0.083974 HP/GPM

Liquid flow rate times head 62001 GPM -FT

WEIGHT DATA

Pump 320 LBS

Motor 320 LBS

Base plate 70 LBS



VENDOR COST DATA






Profit 3341 DOLLARS



Cost per unit liquid flow rate 88.184 USD/GPM


Summary Costs



Piping 12722. 6268. 204

Civil 442. 945. 39



235


Insulation 0. 0. 0

Paint 629. 1311. 58

Subtotal 48872 13413 459


ITEM REPORT



List of Items :

Project : basecase


K 100 TE

Project : BASECASE

Scenario : BASECASE

K 100 TE

Item Code: DTURTURBOEXP

Internal Name : DTURTURBOEXP #1#K 100 TE

Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type TURBOEXP


236


Material A3003





Power output 9654.221 HP




Isentropic efficiency 75.000 PERCENT

Number of spare cartridges 0

WEIGHT DATA


VENDOR COST DATA



Summary Costs



Piping 198883. 41577. 1342

Civil 20386. 11554. 460



Electrical 0. 0. 0

Insulation 14881. 6988. 306

Paint 40. 103. 5

237

Subtotal 3055230 101209 3421


ITEM REPORT



List of Items :

Project : basecase


K 102 TE

Project : BASECASE

Scenario : BASECASE

K 102 TE

Item Code: DTURTURBOEXP

Internal Name : DTURTURBOEXP #1#K 102 TE

Sizing Data

Design Data

Summary Costs

Sizing Data


Design Data


Item type TURBOEXP



Material A3003


238




Power output 8683.142 HP




Isentropic efficiency 75.000 PERCENT

Number of spare cartridges 0

WEIGHT DATA


VENDOR COST DATA



Summary Costs



Piping 152270. 34869. 1123

Civil 18559. 10526. 419



Electrical 0. 0. 0

Insulation 15344. 7481. 328

Paint 40. 103. 5

Subtotal 2726653 92122 3125


239

HYSYS Model Printout

PENNSYLVANIA STATE UNIVEBedford, MAUSA

Case Name: NEWEST_HYSYS.hsc

Unit Set: Field

Date/Time: Mon Apr 17 17:18:58 2017

Conversion Reactor: Prereformer

CONNECTIONS

Inlet Stream Connections

Stream Name From Unit OperationReactor Feed MIX-100Mixer

Outlet Stream Connections

Stream Name To Unit Operation8 Expander: K-100

x

Energy Stream Connections

Stream Name From Unit Operation

PARAMETERS

Physical Parameters Optional Heat Transfer

Delta P Vessel Volume Duty Energy Stream0.0000 psi 10.00 ft3 * 0.0000 Btu/hr

User Variables

Conversion Reactor: MeOH Synth

CONNECTIONS

Inlet Stream Connections

Stream Name From Unit Operation14-5 E-105Heat Exchanger

Outlet Stream Connections

Stream Name To Unit Operation16 Expander: K-102

xxx

Energy Stream Connections

Stream Name From Unit OperationQ-103

PARAMETERS

Physical Parameters Optional Heat Transfer

Delta P Vessel Volume Duty Energy Stream0.0000 psi 20.00 ft3 * -1.002e+008 Btu/hr Q-103

User Variables

Component Splitter: ATU

CONNECTIONS

Inlet Stream

Aspen Technology Inc. Aspen HYSYS Version 8.8 (34.0.1.8909) Page 1 of 22

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* Specified by user.Licensed to: PENNSYLVANIA STATE UNIVE



Unit Set: Field


Component Splitter: ATU (continued)

CONNECTIONS

STREAM NAME FROM UNIT OPERATION

3 MRUComponent Splitter

4

Outlet Stream

STREAM NAME TO UNIT OPERATION

6 Heat Exchanger: E-100

5

Energy Stream


Q-100

PARAMETERS

Stream Specifications

Overhead Pressure: 400.0 psia Overhead Vapour Fraction: 1.0000

Bottoms Pressure: 400.0 psia Bottoms Vapour Fraction: 0.0000

SPLITS

Component Fraction To Overhead

Component

Methane

H2O

Ethane

CO2

Nitrogen

Propane

i-Butane

n-Butane

i-Pentane

n-Pentane

Cyclohexane

CO

Hydrogen

Mercury

H2S

MDEAmine

Methanol

Ethanol

M-Formate

Ammonia

Air

Oxygen

Split Basis

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Split Type

FeedFrac. to Products






















6

1.0000 *

0.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

0.0000 *

0.0000 *

0.0000 *

0.0000 *

0.0000 *

0.0000 *

0.0000 *

0.0000 *

0.0000 *

User Variables


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2

3

4

5

6

7

8

9

10

11

12

13

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15

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63




Unit Set: Field


Component Splitter: MRU

CONNECTIONS

Inlet Stream


1

Outlet Stream


3 Component Splitter: ATU

2

Energy Stream


Q-101

PARAMETERS

Stream Specifications

Overhead Pressure: 400.0 psia Overhead Vapour Fraction: 1.0000

Bottoms Pressure: 400.0 psia Bottoms Vapour Fraction: 0.0000 *

SPLITS

Component Fraction To Overhead

Component

Methane

H2O

Ethane

CO2

Nitrogen

Propane

i-Butane

n-Butane

i-Pentane

n-Pentane

Cyclohexane

CO

Hydrogen

Mercury

H2S

MDEAmine

Methanol

Ethanol

M-Formate

Ammonia

Air

Oxygen

Split Basis

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Molar

Split Type























3

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

1.0000 *

0.0000 *

1.0000 *

1.0000 *

0.0000 *

0.0000 *

0.0000 *

0.0000 *

0.0000 *

0.0000 *

User Variables


1

2

3

4

5

6

7

8

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11

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63




Unit Set: Field


Expander: K-100

CONNECTIONS

Inlet Stream


8 PrereformerConversion Reactor

Outlet Stream


8-1 MIX-101Mixer

Energy Stream


Q-105

PARAMETERS

Duty: 1.0359e+04 hp Speed: ---

Adiabatic Eff.: 75.00 PolyTropic Eff.: 73.29

Adiabatic Head: 1.382e+005 ft Polytropic Head: 1.414e+005 ft

Adiabatic Fluid Head: 1.382e+005 lbf-ft/lbm Polytropic Fluid Head: 1.414e+005 lbf-ft/lbm

Polytropic Exp. 1.123 Isentropic Exp. 1.147 Poly Head Factor 0.8136

User Variables

Expander: K-102

CONNECTIONS

Inlet Stream


16 MeOH SynthConversion Reactor

Outlet Stream


16-1 E-104Heat Exchanger

Energy Stream


Q-111

PARAMETERS

Duty: 1.0944e+04 hp Speed: ---





User Variables

Compressor: K-104

DESIGN


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Unit Set: Field


Compressor: K-104 (continued)

Connections

Inlet Stream


14 V-101Separator

Outlet Stream



Energy Stream


Q-121

Parameters

Speed: --- Duty: 1.8589e+04 hp





User Variables

Compressor: K-105

DESIGN

Connections

Inlet Stream



Outlet Stream



Energy Stream


Q-123

Parameters

Speed: --- Duty: 2.1408e+04 hp





User Variables


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Unit Set: Field


Distillation: T-103

CONNECTIONS

Inlet Stream

STREAM NAME Stage FROM UNIT OPERATION

Q-11921

Reboiler30__Main Tower TEE-101Tee

Outlet Stream

STREAM NAME Stage TO UNIT OPERATION

Q-1182827

CondenserReboilerCondenser

MIX-103

MIX-104

Mixer

Mixer

MONITOR

Specifications Summary

Specified Value Current Value Wt. Error

Reflux Ratio 4.500 * 4.500 1.092e-008

Reflux Rate 4576 lbmole/hr * 5395 lbmole/hr 0.1790

Comp Fraction 1.400e-003 * 1.022e-003 -9.343e-002

Comp Fraction - 2 3.475e-025 * 2.822e-003 2.822e-003

28 Rate 474.5 lbmole/hr * 1107 lbmole/hr 1.333

27 Rate 1220 lbmole/hr * 1199 lbmole/hr -1.748e-002

Comp Fraction - 3 9.538e-002 * --- ---

Comp Flow 100.0 lb/hr * 100.1 lb/hr 6.630e-004

Wt. Tol. Abs. Tol. Active Estimate Used

Reflux Ratio 1.000e-002 1.000e-002 On On On

Reflux Rate 1.000e-002 2.205 lbmole/hr Off On Off

Comp Fraction 1.000e-002 1.000e-003 Off On Off

Comp Fraction - 2 1.000e-002 1.000e-003 Off On Off

28 Rate 1.000e-002 2.205 lbmole/hr Off On Off

27 Rate 1.000e-002 2.205 lbmole/hr Off On Off

Comp Fraction - 3 1.000e-002 1.000e-003 Off On Off

Comp Flow 1.000e-002 2.205 lb/hr On On On

SPECS

Column Specification Parameters

Reflux Ratio

Fix/Rang: Ranged Prim/Alter: Primary Lower Bnd: 5.000 * Upper Bnd: 5.000e-002 *

Stage: Condenser Flow Basis: Molar Liquid Spec: ---

Reflux Rate

Fix/Rang: Fixed Prim/Alter: Primary Lower Bnd: --- Upper Bnd: ---


Comp Fraction

Fix/Rang: Ranged Prim/Alter: Primary Lower Bnd: --- Upper Bnd: ---

Stage: Flow Basis: Mass Fraction Phase: Liquid

Components:H2O Ethanol Ammonia

CO2 CO M-Formate

Comp Fraction - 2

Fix/Rang: Ranged Prim/Alter: Primary Lower Bnd: --- Upper Bnd: ---

Stage: Flow Basis: Mole Fraction Phase: Liquid

Components: Methanol


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Unit Set: Field


Distillation: T-103 (continued)Column Specification Parameters

28 Rate


Stream: 28 @COL2 Flow Basis: Molar

27 Rate



Comp Fraction - 3


Stage: Flow Basis: Mass Fraction Phase: Liquid

Components: CO2

Comp Flow


Draw: 28 @COL2 Flow Basis: Mass Phase: Liquid


SUBCOOLING

Degrees of Subcooling

Subcool to

Condenser

---

---

User Variables

Distillation: T-100

CONNECTIONS

Inlet Stream

STREAM NAME Stage FROM UNIT OPERATION

Q-10620

Reboiler30__Main Tower TEE-101Tee

Outlet Stream

STREAM NAME Stage TO UNIT OPERATION

Q-1042423

CondenserReboilerCondenser

MIX-103

MIX-104

Mixer

Mixer

MONITOR

Specifications Summary

Specified Value Current Value Wt. Error

Reflux Ratio 4.500 * 4.500 -1.139e-008

Reflux Rate --- 5395 lbmole/hr ---

Btms Prod Rate --- 1107 lbmole/hr ---

Distillate Rate --- 1199 lbmole/hr ---

Comp Flow 100.0 lb/hr * 100.1 lb/hr 6.760e-004

Wt. Tol. Abs. Tol. Active Estimate Used

Reflux Ratio 1.000e-002 1.000e-002 On On On

Reflux Rate 1.000e-002 2.205 lbmole/hr Off On Off

Btms Prod Rate 1.000e-002 2.205 lbmole/hr Off On Off

Distillate Rate 1.000e-002 2.205 lbmole/hr Off On Off


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Unit Set: Field


Distillation: T-100 (continued)Specifications Summary

Comp Flow 1.000e-002 2.205 lb/hr On On On

SPECS

Column Specification Parameters

Reflux Ratio



Reflux Rate



Btms Prod Rate



Distillate Rate



Comp Flow


Draw: 24 @COL1 Flow Basis: Mass Phase: Liquid


SUBCOOLING

Degrees of Subcooling

Subcool to

Condenser

---

---

User Variables

Mixer: MIX-100

CONNECTIONS

Inlet Stream



Outlet Stream


Reactor Feed PrereformerConversion Reactor

PARAMETERS

User Variables


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Unit Set: Field


Mixer: MIX-101

CONNECTIONS

Inlet Stream


8-113-1

K-100

RCY-2

Expander

Recycle

Outlet Stream


9 E-107Heat Exchanger

PARAMETERS

User Variables

Mixer: MIX-103

CONNECTIONS

Inlet Stream


222412

T-103

T-100

TEE-100

Distillation

Distillation

Tee

Outlet Stream


27 P-100Pump

PARAMETERS

User Variables

Mixer: MIX-104

CONNECTIONS

Inlet Stream


2123

T-103

T-100

Distillation

Distillation

Outlet Stream


29 P-101Pump

PARAMETERS

User Variables


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Unit Set: Field


Recycle: RCY-2

CONNECTIONS

Inlet Stream


13 TEE-100Tee

Outlet Stream

Stream Name To Unit Operation

13-1 MIX-101Mixer

TOLERANCE

Vapour Fraction: 10.00 * Temperature: 10.00 * Pressure: 10.00 *

Flow: 0.1000 * Enthalpy: 10.00 * Composition: 0.1000 *

NUMERICAL

Acceleration Type: Wegstein Iteration Type: Nested

Maximum Iterations: 10 * Iteration Count: 0 *

Wegstein Count: 3 * Q Minimum: -20.00 * Q Maximum: 0.0000 *

Iteration History

Iteration

0 *

Variable

Converged

Outlet Value

---

Inlet Value

---

User Variables

Recycle: RCY-1

CONNECTIONS

Inlet Stream


17 V-100Separator

Outlet Stream


17-1

TOLERANCE

Vapour Fraction: 10.00 * Temperature: 10.00 * Pressure: 10.00 *

Flow: 10.00 * Enthalpy: 10.00 * Composition: 10.00 *

NUMERICAL

Acceleration Type: Wegstein Iteration Type: Nested

Maximum Iterations: 10 * Iteration Count: 0 *

Wegstein Count: 3 * Q Minimum: -20.00 * Q Maximum: 0.0000 *

Iteration History

Iteration

1 *

Variable

Converged

Outlet Value

---

Inlet Value

---

User Variables


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Unit Set: Field


Plug Flow Reactor: Steam Reformer

CONNECTIONS

Inlet Stream


9-2 FH-100Fired Heater

Outlet Stream



Energy Stream


Q-1000

PARAMETERS

Physical Parameters

Type : User Specified Pressure Drop: 0.0000 psi *

Heat Transfer

Type : Direct Q Value Energy Stream : Q-1000 Duty : -2.323e+008 Btu/hr

Dimensions

Total Volume: 1924 ft3 Length: 50.00 ft * Diameter: 7.000 ft * Number of Tubes: 1 *

Wall Thickness: 1.640e-002 ft * Void Fraction: 1.0000 * Void Volume: 1924 ft3

Reaction Info

Reaction Set: Set-1 Initialize From: Current

Integration Information

Number of Segments: 20 * Minimum Step Fraction: 1.0e-06 * Minimum Step Length: 5.0e-05 ft

User Variables

Tee: TEE-100

CONNECTIONS

Inlet Stream


11 V-101Separator

Outlet Stream


1213

MIX-103

RCY-2

Mixer

Recycle

PARAMETERS

12

13

Flow Ratios

0.1397

0.8603

Dynamic Valve Openings

13.97

86.03


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Unit Set: Field


Tee: TEE-100 (continued)

PARAMETERS

Valve Control: Multiple Stream

User Variables

Tee: TEE-101

CONNECTIONS

Inlet Stream



Outlet Stream


1920

T-103

T-100

Distillation

Distillation

PARAMETERS

19

20

Flow Ratios

0.5000 *

0.5000

Dynamic Valve Openings

50.00

50.00

Valve Control: Multiple Stream

User Variables

Separator: V-100

CONNECTIONS

Inlet Stream


16-2 Heat Exchanger: E-104

Outlet Stream


17 Recycle: RCY-1

18 Heat Exchanger: E-102

Energy Stream


PARAMETERS

Vessel Volume: --- Level SP: 50.00 % Liquid Volume: ---

Vessel Pressure: 30.00 psia Pressure Drop: 0.0000 psi Duty: 0.0000 Btu/hr Heat Transfer Mode: Heating

User Variables


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2

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Unit Set: Field


Separator: V-101

CONNECTIONS

Inlet Stream


10-4 Heat Exchanger: E-108

Outlet Stream


14 Compressor: K-104

11 Tee: TEE-100

Energy Stream


PARAMETERS

Vessel Volume: --- Level SP: 50.00 % Liquid Volume: ---

Vessel Pressure: 30.00 psia Pressure Drop: 0.0000 psi Duty: 0.0000 Btu/hr Heat Transfer Mode: Heating

User Variables

Heat Exchanger: E-100

CONNECTIONS

Tube Side

Inlet

Name

From Op.

Op. Type

Temp

10

Steam Reformer

Plug Flow Reactor

1100.00 F *

Outlet

Name

To Op.

Op. Type

Temp

10-1

E-107

Heat Exchanger

1063.88 F

Shell Side

Inlet

Name

From Op.

Op. Type

Temp

6

ATU

Component Splitter

72.00 F *

Outlet

Name

To Op.

Op. Type

Temp

6-1

MIX-100

Mixer

500.00 F *

PARAMETERS

Heat Exchanger Model: Simple End Point

Tube Side DeltaP: 0.0000 psi * Shell Side DeltaP: 0.0000 psi * Passes: ---

UA: 1.599e+004 Btu/F-hr Tolerance: 1.0000e-04

Tube Side Data Shell Side Data

Heat Transfer Coeff

Tube Pressure Drop

Fouling

Tube Length

Tube O.D.

Tube Thickness

Tube Pitch

Orientation

Passes Per Shell

Tubes Per Shell

Layout Angle

TEMA Type

---

0.00 psi *

0.00000 F-hr-ft2/Btu

19.69 ft

0.79 in

0.0787 in

1.9685 in

Horizontal

2

160 *

Triangular (30 degrees)

A E L

Heat Transfer Coeff

Shell Pressure Drop

Fouling

Shell Passes

Shell Series

Shell Parallel

Baffle Type

Baffle Cut(%Area)

Baffle Orientation

Spacing

Diameter

Area

---

0.00 psi *


1

1

1

Single

20.00

Horizontal

31.4961 in

29.0964 in

649.26 ft2

SPECS


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Unit Set: Field


Heat Exchanger: E-100 (continued)Spec Value Curr Value Rel Error Active Estimate

E-100 Heat Balance 0.0000 Btu/hr 9.310e-008 Btu/hr 7.500e-015 On Off

E-100 UA --- 1.599e+004 Btu/F-hr --- On Off

Detailed Specifications

E-100 Heat BalanceType: Duty Pass: Error Spec Value: 0.0000 Btu/hr

E-100 UAType: UA Pass: Overall Spec Value: ---

User Variables


CONNECTIONS

Tube Side

Inlet

Name

From Op.

Op. Type

Temp

9

MIX-101

Mixer

228.41 F

Outlet

Name

To Op.

Op. Type

Temp

9-1

FH-100

Fired Heater

230.00 F *

Shell Side

Inlet

Name

From Op.

Op. Type

Temp

10-1

E-100

Heat Exchanger

1063.88 F

Outlet

Name

To Op.

Op. Type

Temp

10-2

E-101

Heat Exchanger

1029.54 F

PARAMETERS





Heat Transfer Coeff

Tube Pressure Drop

Fouling

Tube Length

Tube O.D.

Tube Thickness

Tube Pitch

Orientation

Passes Per Shell

Tubes Per Shell

Layout Angle

TEMA Type

---

0.00 psi *


19.69 ft

0.79 in

0.0787 in

1.9685 in

Horizontal

20 *

140 *


A E L

Heat Transfer Coeff

Shell Pressure Drop

Fouling

Shell Passes

Shell Series

Shell Parallel

Baffle Type

Baffle Cut(%Area)

Baffle Orientation

Spacing

Diameter

Area

---

0.00 psi *


1

10 *

1

Single

20.00

Horizontal

31.4961 in

27.2818 in

568.11 ft2

SPECS

Spec Value Curr Value Rel Error Active Estimate

E-107 Heat Balance 0.0000 Btu/hr -1.552e-008 Btu/hr -1.322e-015 On Off






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Unit Set: Field


Heat Exchanger: E-107 (continued)

User Variables


CONNECTIONS

Tube Side

Inlet

Name

From Op.

Op. Type

Temp

14-3

K-105

Compressor

859.07 F

Outlet

Name

To Op.

Op. Type

Temp

14-4

E-105

Heat Exchanger

531.87 F

Shell Side

Inlet

Name

From Op.

Op. Type

Temp

18

V-100

Separator

30.00 F

Outlet

Name

To Op.

Op. Type

Temp

18-1

TEE-101

Tee

206.44 F

PARAMETERS





Heat Transfer Coeff

Tube Pressure Drop

Fouling

Tube Length

Tube O.D.

Tube Thickness

Tube Pitch

Orientation

Passes Per Shell

Tubes Per Shell

Layout Angle

TEMA Type

---

0.00 psi *


19.69 ft

0.79 in

0.0787 in

1.9685 in

Horizontal

4 *

160 *


A E L

Heat Transfer Coeff

Shell Pressure Drop

Fouling

Shell Passes

Shell Series

Shell Parallel

Baffle Type

Baffle Cut(%Area)

Baffle Orientation

Spacing

Diameter

Area

---

0.00 psi *


1

1

1

Single

20.00

Horizontal

31.4961 in

29.0964 in

649.26 ft2

SPECS







User Variables


CONNECTIONS

Tube Side

Inlet Outlet

Shell Side

Inlet Outlet


1

2

3

4

5

6

7

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10

11

12

13

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Unit Set: Field


Heat Exchanger: E-103 (continued)Name

From Op.

Op. Type

Temp

CW5

90.00 F *

Name

To Op.

Op. Type

Temp

CW6

110.00 F *

Name

From Op.

Op. Type

Temp

14-1

K-104

Compressor

675.75 F

Name

To Op.

Op. Type

Temp

14-2

K-105

Compressor

280.00 F *

PARAMETERS





Heat Transfer Coeff

Tube Pressure Drop

Fouling

Tube Length

Tube O.D.

Tube Thickness

Tube Pitch

Orientation

Passes Per Shell

Tubes Per Shell

Layout Angle

TEMA Type

---

0.00 psi *


19.69 ft

0.79 in

0.0787 in

1.9685 in

Horizontal

2

160 *


A E L

Heat Transfer Coeff

Shell Pressure Drop

Fouling

Shell Passes

Shell Series

Shell Parallel

Baffle Type

Baffle Cut(%Area)

Baffle Orientation

Spacing

Diameter

Area

---

0.00 psi *


1

1

1

Single

20.00

Horizontal

31.4961 in

29.0964 in

649.26 ft2

SPECS







User Variables


CONNECTIONS

Tube Side

Inlet

Name

From Op.

Op. Type

Temp

14-4

E-102

Heat Exchanger

531.87 F

Outlet

Name

To Op.

Op. Type

Temp

14-5

MeOH Synth

Conversion Reactor

500.00 F *

Shell Side

Inlet

Name

From Op.

Op. Type

Temp

CW7

90.00 F *

Outlet

Name

To Op.

Op. Type

Temp

CW8

110.00 F *

PARAMETERS


Tube Side DeltaP: 0.0000 psi Shell Side DeltaP: 0.0000 psi * Passes: ---

UA: 7204 Btu/F-hr Tolerance: 1.0000e-04


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5

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13

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17

18

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63




Unit Set: Field





Heat Transfer Coeff

Tube Pressure Drop

Fouling

Tube Length

Tube O.D.

Tube Thickness

Tube Pitch

Orientation

Passes Per Shell

Tubes Per Shell

Layout Angle

TEMA Type

---

0.00 psi


19.69 ft

0.79 in

0.0787 in

1.9685 in

Horizontal

2

160 *


A E L

Heat Transfer Coeff

Shell Pressure Drop

Fouling

Shell Passes

Shell Series

Shell Parallel

Baffle Type

Baffle Cut(%Area)

Baffle Orientation

Spacing

Diameter

Area

---

0.00 psi *


1

1

1

Single

20.00

Horizontal

31.4961 in

29.0964 in

649.26 ft2

SPECS



E-105 UA --- 7204 Btu/F-hr --- On Off




User Variables


CONNECTIONS

Tube Side

Inlet

Name

From Op.

Op. Type

Temp

CW1

90.00 F *

Outlet

Name

To Op.

Op. Type

Temp

CW2

110.00 F *

Shell Side

Inlet

Name

From Op.

Op. Type

Temp

10-2

E-107

Heat Exchanger

1029.54 F

Outlet

Name

To Op.

Op. Type

Temp

10-3

E-108

Heat Exchanger

190.00 F *

PARAMETERS





Heat Transfer Coeff

Tube Pressure Drop

Fouling

Tube Length

Tube O.D.

Tube Thickness

Tube Pitch

---

0.00 psi *


19.69 ft

0.79 in

0.0787 in

1.9685 in

Heat Transfer Coeff

Shell Pressure Drop

Fouling

Shell Passes

Shell Series

Shell Parallel

Baffle Type

---

0.00 psi *


1

1

1

Single


1

2

3

4

5

6

7

8

9

10

11

12

13

14

15

16

17

18

19

20

21

22

23

24

25

26

27

28

29

30

31

32

33

34

35

36

37

38

39

40

41

42

43

44

45

46

47

48

49

50

51

52

53

54

55

56

57

58

59

60

61

62

63




Unit Set: Field




Orientation

Passes Per Shell

Tubes Per Shell

Layout Angle

TEMA Type

Horizontal

6 *

156 *


A E L

Baffle Cut(%Area)

Baffle Orientation

Spacing

Diameter

Area

20.00

Horizontal

31.4961 in

28.7430 in

633.03 ft2

SPECS







User Variables


CONNECTIONS

Tube Side

Inlet

Name

From Op.

Op. Type

Temp

CW3

90.00 F *

Outlet

Name

To Op.

Op. Type

Temp

CW4

110.00 F *

Shell Side

Inlet

Name

From Op.

Op. Type

Temp

10-3

E-101

Heat Exchanger

190.00 F *

Outlet

Name

To Op.

Op. Type

Temp

10-4

V-101

Separator

160.00 F *

PARAMETERS





Heat Transfer Coeff

Tube Pressure Drop

Fouling

Tube Length

Tube O.D.

Tube Thickness

Tube Pitch

Orientation

Passes Per Shell

Tubes Per Shell

Layout Angle

TEMA Type

---

0.00 psi *


19.69 ft

0.79 in

0.0787 in

1.9685 in

Horizontal

2

160 *


A E L

Heat Transfer Coeff

Shell Pressure Drop

Fouling

Shell Passes

Shell Series

Shell Parallel

Baffle Type

Baffle Cut(%Area)

Baffle Orientation

Spacing

Diameter

Area

---

0.00 psi *


1

1

1

Single

20.00

Horizontal

31.4961 in

29.0964 in

649.26 ft2

SPECS



1

2

3

4

5

6

7

8

9

10

11

12

13

14

15

16

17

18

19

20

21

22

23

24

25

26

27

28

29

30

31

32

33

34

35

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39

40

41

42

43

44

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Unit Set: Field


Heat Exchanger: E-108 (continued)E-108 Heat Balance 0.0000 Btu/hr -1.002e-003 Btu/hr -1.498e-011 On Off





User Variables


CONNECTIONS

Tube Side

Inlet

Name

From Op.

Op. Type

Temp

16-1

K-102

Expander

181.77 F

Outlet

Name

To Op.

Op. Type

Temp

16-2

V-100

Separator

30.00 F *

Shell Side

Inlet

Name

From Op.

Op. Type

Temp

25

-43.74 F *

Outlet

Name

To Op.

Op. Type

Temp

26

-43.70 F *

PARAMETERS





Heat Transfer Coeff

Tube Pressure Drop

Fouling

Tube Length

Tube O.D.

Tube Thickness

Tube Pitch

Orientation

Passes Per Shell

Tubes Per Shell

Layout Angle

TEMA Type

---

0.00 psi *


19.69 ft

0.79 in

0.0787 in

1.9685 in

Horizontal

2

160 *


A E L

Heat Transfer Coeff

Shell Pressure Drop

Fouling

Shell Passes

Shell Series

Shell Parallel

Baffle Type

Baffle Cut(%Area)

Baffle Orientation

Spacing

Diameter

Area

---

0.00 psi *


1

1

1

Single

20.00

Horizontal

31.4961 in

29.0964 in

649.26 ft2

SPECS







User Variables


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2

3

4

5

6

7

8

9

10

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12

13

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63




Unit Set: Field



CONNECTIONS

Tube Side

Inlet

Name

From Op.

Op. Type

Temp

34

90.00 F *

Outlet

Name

To Op.

Op. Type

Temp

35

110.00 F *

Shell Side

Inlet

Name

From Op.

Op. Type

Temp

30

P-101

Pump

174.80 F

Outlet

Name

To Op.

Op. Type

Temp

33

120.00 F *

PARAMETERS





Heat Transfer Coeff

Tube Pressure Drop

Fouling

Tube Length

Tube O.D.

Tube Thickness

Tube Pitch

Orientation

Passes Per Shell

Tubes Per Shell

Layout Angle

TEMA Type

---

0.00 psi *


19.69 ft

0.79 in

0.0787 in

1.9685 in

Horizontal

2

160 *


A E L

Heat Transfer Coeff

Shell Pressure Drop

Fouling

Shell Passes

Shell Series

Shell Parallel

Baffle Type

Baffle Cut(%Area)

Baffle Orientation

Spacing

Diameter

Area

---

0.00 psi *


1

1

1

Single

20.00

Horizontal

31.4961 in

29.0964 in

649.26 ft2

SPECS


E-106 Heat Balance 0.0000 Btu/hr -9.107e-006 Btu/hr -2.264e-012 On Off





User Variables

Fired Heater: FH-100

CONNECTIONS

Combustion Product Exhaust Fuel Gas Feed Air Feed

Economizer Zone Connections

Econ Zone Inlet Econ Zone Outlet # of Tubes 0 *

Convective Zone

Conv Zone Inlet Conv Zone Outlet # of Tubes 0 *

Radiative Zone

Radiant Zone Inlet

9-1

Radiant Zone Oulet

9-2# of Tubes 1 *


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Unit Set: Field


Fired Heater: FH-100 (continued)

PARAMETERS

Min Air Fuel Ratio ---

Air Fuel Ratio ---

Max Air Fuel Ratio ---

Flame Temp ---

Lean Factor 1.000 *

Rich Factor 40.00 *

Heater SS Efficiency 60.00 *

Heater SS Excess Air Percent 40.00 *

User Variables

Pump: P-100

CONNECTIONS

Inlet Stream


27 MIX-103Mixer

Outlet Stream


28

Energy Stream


Q-102

PARAMETERS

Adiabatic Efficiency (%): 75.00 Delta P: 21.00 psi Duty: 3.573 hp

CURVES

Delta P: 21.00 psi Duty: 3.573 hp

Coefficient A: 0.0000 * Coefficient B: 0.0000 * Coefficient C: 0.0000 *

Parameter Preferences Units for Delta P: ft Flow Basis ActVolFlow Units for Flow: barrel/day

User Variables

Pump: P-101

CONNECTIONS

Inlet Stream


29 MIX-104Mixer

Outlet Stream



Energy Stream


Q-107

PARAMETERS

Adiabatic Efficiency (%): 75.00 Delta P: 75.50 psi Duty: 12.41 hp


1

2

3

4

5

6

7

8

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Unit Set: Field


Pump: P-101 (continued)

CURVES

Delta P: 75.50 psi Duty: 12.41 hp

Coefficient A: 40.00 * Coefficient B: 0.0000 * Coefficient C: 0.0000 *

Parameter Preferences Units for Delta P: ft Flow Basis ActVolFlow Units for Flow: barrel/day

User Variables


1

2

3

4

5

6

7

8

9

10

11

12

13

14

15

16

17

18

19

20

21

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Documents

LitLion Chemical - Team 15