group g - final report

[ T Y P E T H E C O M P A N Y A D D R E S S ]

2008

Removal of CO2 from a

Hydrogen Plant

Department of Chemical Engineering

Robert St. Pierre, Phung-Minh Dai, Mark Dalton

Removal of CO2 from a Hydrogen Plant

By: Robert St. Pierre, Phung-Minh Dai, and Mark Dalton

Department of Chemical Engineering University of Saskatchewan

2007-2008

i

Abstract

Husky Energy’s Lloydminster Upgrader wanted to determine if removing carbon

dioxide from their Hydrogen plant would be economical using current technology. A

computer model of the plant was developed using the HYSYS process simulator for this

purpose. The model was used to predict the effects of removing the CO2 on the existing

plant. From the effects seen in the model the savings related to lowering the heating costs

could be determined.

To remove the CO2 from the process stream an absorption system using

Monoethanolamine (MEA) was designed. The removal system consists of a 15 tray

absorber, a 17 stage regenerator, and a 12 stage MEA guard absorber along with several

pumps and heat exchangers. The system is able to remove 35 tonnes per hour of CO2 at

94 percent purity, with the remainder being water. The capital cost of this project is

approximately $9.82 million.

The economics of the project were not found to be favourable. The total savings

from removing the CO2 from the gas stream are approximately $859,000 per year. The

cost to operate the amine system is around $16.9 million per year. If a value of $50.43

per tonne CO2 is applied then the net cash flow is zero. If a value of $67.33 per tonne

CO2 is applied then the project will break even after 25 years. It is RPM’s

recommendation that Husky conduct an investigation to determine the value they can

place on the CO2 product before moving forward with this project.

ii

Acknowledgements

RPM is pleased to acknowledge the following people for their contributions and

guidance:

• Tristan Koroscil, Senior Unit Contact Engineer Husky Upgrader

• Bob Brierly, Senior Staff Process Engineer Husky Upgrader

• Les Alberts, Gas Treating Specialist, Dow Chemical Canada

• Dr. Hui Wang, U of S Chemical Engineering Assistant Professor

• Dr. Richard Evitts, U of S Chemical Engineering Associate Professor

• Dr. Ding-Yu Peng, U of S Chemical Engineering Professor

• Dr. Gordon Hill, U of S Chemical Engineer Department Head

• Trey Brown, Vice President of Process Engineering, New Point Gas Services

iii

Table of Contents Abstract ................................................................................................................................ i

Acknowledgements............................................................................................................. ii

List of Tables ..................................................................................................................... vi

List of Figures ................................................................................................................... vii

Nomenclature ................................................................................................................... viii

1. Introduction ................................................................................................................. 1

2. Literature Survey: Alternative Processes .................................................................... 2

2.1 Membrane Separation ......................................................................................... 2

2.2 Hot Potassium Carbonate.................................................................................... 3

2.3 Amine Separation................................................................................................ 4

3. Detailed Qualitative Process Description .................................................................... 7

3.1 HYSYS Simulation............................................................................................. 7

3.1.1 Detailed Model Specifications ........................................................................ 7

3.1.2 Tail Gas Flow Rate Assumption ................................................................... 22

3.2 AMSIM Simulation .......................................................................................... 25

4. Equipment Specification and Design ........................................................................ 29

5. Plant Safety Analysis................................................................................................. 31

6. Economic Analysis .................................................................................................... 33

7. Conclusions and Recommendations .......................................................................... 36

8. References ................................................................................................................. 37

Appendix A: Sample Calculations................................................................................... 39

iv

A.1 Membrane Size ..................................................................................................... 40

A.2 Absorber Size and Cost......................................................................................... 41

A.3 Regenerator Size and Cost .................................................................................... 42

A.4 Condenser Heat Exchanger Size ........................................................................... 44

A.5 Centrifugal Pump Size and Cost ........................................................................... 45

A.6 Amine Holding Tank Size..................................................................................... 47

A.7 Depreciation .......................................................................................................... 48

A.8 Combo Gas Savings .............................................................................................. 48

A.9 Combustion Air...................................................................................................... 48

A.10 Steam Losses....................................................................................................... 49

Appendix B: Safety Document ........................................................................................ 50

B.1 Amine Plant Design Criteria ................................................................................. 51

B.2 HAZOP/Safety Considerations ............................................................................. 52

B.3 Plant Safety ........................................................................................................... 59

B.4 Process Safety Management System..................................................................... 62

B.5 Chemical Hazard Information ............................................................................... 63

B.6 MSDS .................................................................................................................... 65

B.6.1 Hydrogen, H2 ................................................................................................. 65

B.6.2 Monoethanolamine, MEA ............................................................................. 67

B.6.3 Methane, CH4 ................................................................................................ 75

B.6.4 Carbon Dioxide, CO2 .................................................................................... 79

B.6.5 Carbon Monoxide, CO.................................................................................... 85

B.6.6 Nitrogen, N2 ................................................................................................... 91

v

B.7 Dow Fire and Explosion Index.............................................................................. 93

Appendix C: EconExpert Equipment Costing Results .................................................... 95

C.1 Towers ................................................................................................................... 96

C.2 Heat Exchangers.................................................................................................... 98

C.3 Pumps .................................................................................................................. 100

C.4 Storage Vessel ..................................................................................................... 101

C.5 Process Vessels ................................................................................................... 101

Appendix D: Cash Flow Analysis................................................................................... 104

Appendix E: HYSYS Reports......................................................................................... 107

Appendix E1.1: PSA Tail Gas With CO2 ................................................................ 108

Appendix E1.2: Reformer Furnace With CO2......................................................... 112

Appendix E2.1: PSA Tail Gas Without CO2 ........................................................... 124

Appendix E2.2: Reformer without CO2 .................................................................. 127

Appendix F: AMSIM Reports ....................................................................................... 139

vi

List of Tables Table 1: Summary of Process Changes ........................................................................... 22

Table 2: Summary of Equipment Sizing and Specifications ............................................ 29

Table B. 1: Chemical Hazard Information Summary ...................................................... 64

Table C. 1: Ulrich Equipment Costing .......................................................................... 103

Table D. 1: Cash Flow Analysis .................................................................................... 105

Table D. 2: Cash flows at Different Carbon Tax Rates ................................................. 106

Table F. 1: Composition Profile of CO2 in Absorber A................................................. 140

Table F. 2: Vapour Phase Properties in Absorber A...................................................... 140

Table F. 3: Liquid Phase Properties in Absorber A ....................................................... 141

Table F. 4: Composition Profile of CO2 in Regenerator................................................ 141

Table F. 6: Liquid Phase Properties in the Regenerator ................................................ 142

vii

List of Figures Figure 1: CO2 Removal Processes Comparison................................................................. 4

Figure 2: HYSYS Simulation Section 1 ............................................................................. 8

Figure 3: Specified Inlet Gas Compositions ...................................................................... 9

Figure 4: HYSYS Simulation Section 2 ........................................................................... 11



Figure 7: Specified Combo Gas Composition ................................................................. 18

Figure 8: Plant 30 Block Diagram ................................................................................... 24

Figure 9: Amine System Flow Diagram ........................................................................... 27

Figure 10: Cumulative Discounted Cash Flow at Different Carbon Tax Rates in $/tonne

CO2 Emitted .................................................................................................... 35

viii

Nomenclature Symbol Description Units A heat exchanger surface area m2 cP purchased equipment cost $ CBM Bare Module Cost $ CP heat capacity kJ/kgoC D diameter m ed,em electric motor efficiency kJ/s ed,gt gas turbine efficiency kJ/s

iε pump efficiency fq quantity factor

2COF molar flow of carbon dioxide mol/s Fa

BM bare module factor, FM material factor FP pressure factor Hcol height of column m

2COj molar flux of carbon dioxide mol/m2s nact number of stages P pressure barg

emP electric motor power kJ/s

gtP gas turbine power kJ/s Q rate of heat transfer J/s •

q volumetric flowrate m3/s s tray spacing m Tci temperature of the cold stream at inlet oC Tco temperature of the cold stream at outlet oC Thi temperature of the hot stream at inlet oC Tho temperature of the hot stream at outlet oC U overall heat transfer coefficient W/m2oC •

sW shaft power kW

1

1. Introduction

The purpose of this project was to remove CO2 from a methane reformer gas

stream located at Husky Energy’s Lloydminster Upgrader. An economic opportunity was

seen in removing this CO2 for two main reasons. First, the CO2 produced in the

reforming and shift reactions is currently part of the fuel gas for the reformer furnace and

by removing the CO2, and thereby reducing the heating requirements, there are potential

savings in the reformer furnace. The second reason is that the CO2 produced by the

removal process could then be used by Husky for their enhanced oil recovery (EOR)

projects.

The project itself has three major objectives which needed to be achieved in order

for the project to be deemed a success. First a working model of the current Hydrogen

plant using simulation software had to be completed. The CO2 removal system then had

to be designed, the equipment sized, and the costs estimated. Finally, the effects of the

CO2 removal on the Hydrogen plant, including savings due to lower combo gas flows and

changes to any additional flows or conditions, had to be determined. From the changes

observed in the system economics could be performed to determine the project’s

feasibility.

2

2. Literature Survey: Alternative Processes

2.1 Membrane Separation One of the technologies examined for the CO2 removal project was membrane

technology. Membrane separation works by gas molecules permeating through a thin

membrane due to a pressure gradient across the membrane. Different species permeate

the membrane at different speeds due to a large number of factors including the size of

the molecule, the speed of sorption onto the membrane, the ability of the molecule to

dissolve in the membrane, the rigidity of the membrane lattice, and many other factors.

These factors can be summarized by the permeability of the species which is obviously

different for each component through each membrane. Although membranes can

produce the high purities that would be desirable the main design challenge is due to the

large membrane area caused by the high flow rate of the gas stream.

In order to find out if a membrane solution would be feasible a preliminary

calculation was performed. This calculation assumed that only CO2 would permeate the

membrane, an assumption that would be unacceptable for further design. The calculation

also assumed that the membrane had the highest permeability that could be found for CO2

based on Scott’s “Industrial Membrane Separation Technology,” which was a

permeability of 2700 Barrer. (Scott) The next assumption made was that the permeate

gas stream was a perfect vacuum, thereby creating the greatest flux. The final

3

assumption was that the membrane has a thickness of 1mm. For a 600 kgmol/hr

separation of CO2 the minimum surface area, based on the calculated flux of 5.87x10-4

mol/m2s, was found to be 283,736m2 which is far beyond the range of feasibility. Due to

the extremely large ideal area needed the membrane system was not pursued further. A

summary of the calculation is shown in Appendix A.

2.2 Hot Potassium Carbonate

Another alternative technology that was considered was using a physical solvent,

namely hot potassium carbonate. As displayed in Figure 1 the hot potassium carbonate

process could be used based on the composition of our process gas stream as seen in

Appendix E1.1. However, the hot potassium carbonate process requires a high contactor

temperature which was not ideal for our process system. The contactor temperature

required for the process is 110oC and the process gas stream is cooled from 170oC to

69oC by raw gas air coolers. Therefore in order to utilize the hot potassium carbonate

process the gas stream would need to be reheated to 110oC. This additional energy

required is undesired as the energy costs for the process were recognized as a potential

downfall of the project. Therefore alternative solvents were considered, such as primary

and secondary amines.

Other physical solvents are used in industry other than potassium carbonate such

as Dow Chemical’s patented Selexol. Selexol is a popular physical solvent in industry

that could possibly be utilized for this system. However because Selexol’s composition

is proprietary to Dow, simulation and comparison with this solvent was not possible. If

4

this project is to be explored further it is RPM’s recommendation that Selexol or another

physical solvent be considered for the process.

Figure 1: CO2 Removal Processes Comparison (Maddox)

2.3 Amine Separation

The third technology explored was using an amine solvent such as

Monoethanolamine or Diethanolomine to absorb the CO2 from the gas stream. As can be

seen in the Figure 1 the primary and secondary amines are mainly utilized for smaller

acid gas concentrations but can be used in our process as the volume of CO2 is within the

threshold. Amine solvents are common in industry to sweeten acid gas streams by

removing H2S and/or CO2.

5

The amine process comprises of a contacting tower operating at a high pressure

and low temperature. These conditions are ideal for our process as the hydrogen gas

stream operates at 24oC and 2200 kPag at our tie in point. Following absorption in the

contactor the CO2 can be captured by desorbing the CO2 from the amine solution in a

regeneration tower at a low pressure and high temperature. The amine solution can then

be recycled back into the contactor tower to absorb more CO2.

When amine solvents are used to absorb CO2 the process occurs through reactive

absorption. When the gas and absorbent are contacted a reversible chemical reaction

occurs, unlike the more common physical absorption like when water is used as a solvent.

When designing an absorption column which undergoes a reversible chemical reaction

new complexities are added to the task. Upon further research it was discovered that

simulating reversible chemical reactions, “are best handled by computer-aided

calculations.” (J. D. Seader)

There are several amine solutions that can be used to absorb CO2, each with different

strengths and drawbacks. The following is a list of potential amine solutions that could

be utilized:

• Monoethanolamine (MEA)

• Diethanolamine (DEA)

• Methyldiethanolamine (MDEA)

• Triethanolamine (TEA)

• Diglycolamine (DGA)

• Diisopropanolamine (DIPA)

6

MEA was chosen in the design because it is the strongest base of the listed amines

and therefore reacts most rapidly with the acid gases. The rapid reaction was deemed

beneficial to limit the size of our absorber tower. Also it was discovered that, “MEA has

the lowest molecular weight of the common amines (and) it theoretically has the largest

carrying capacity for acid gases on a unit weight or volume basis.” (Maddox) The large

carrying capacity was an important factor in choosing MEA because it would enable the

system to use less solvent. A smaller amount of solvent will result in a lower energy

requirement to regenerate the amine.

7

3. Detailed Qualitative Process Description

3.1 HYSYS Simulation

A major part of the project was determining the effects of the CO2 removal on the

existing system. In order to accurately examine how the rest of Plant 30 would react to

the changes in the system, a HYSYS simulation for the current plant was constructed.

The model deals with the Hydrogen plant from the feedstock until directly after the

pressure swing adsorption (PSA) unit. The following section will go through, in detail,

how the HYSYS simulation works, what assumptions have been made, and how the

model was used to examine how removing the CO2 from after the reformer will affect the

rest of the system. The names of streams and unit operations within the HYSYS

simulation will be referred to in parentheses the first time they are mentioned.

3.1.1 Detailed Model Specifications

The simulation uses the Peng-Robinson equation of state as the fluid package for

the simulation. It is used by HYSYS to calculate the thermodynamic properties

associated with the streams.

8

Figure 2: HYSYS Simulation Section 1

The model begins with the inlet natural gas (Nat gas in) which is at 43oC and

3500 kPa pressure. The gas flows at 23.5 tonnes per hour. The composition of the

stream was provided by Husky and can be seen below in Figure 3.

9

Figure 3: Specified Inlet Gas Compositions

The inlet gas is mixed with recycle Hydrogen at 500oC and 4581 kPa in a mixer

(MIX-101). The recycle H2 is flowing at a rate of 290 std m3 per hour. The mixture is

then heated by convection section gas in a heat exchanger (Hydrocarbon Preheat Coil) to

350oC. The heat exchanger is modelled with no pressure drop, no heat losses, and the Ft

correction factor is not calculated. The size of the heat exchanger is modelled as 60 m2

which is not representative of the actual exchanger, which is much larger. The area,

however, does not matter because the discrepancy between the actual area and the

specified area will be taken into account by the UA term. The UA term was found to be

80900 kJ/oC-h, which is based on the outlet temperature specification for the inlet stream

and duty estimates for the heat exchanger from the P&IDs provided by Husky of

7.13MW.

10

In the plant the heated inlet gas stream (Past Sulphur Guard) will move through a

sulphur guard unit to eliminate any H2S, however in the HYSYS model the system is

ignored. The sulphur guard system is ignored because it does will not have an effect on

the system when it is changed. It also does not appreciably change operating conditions

of the inlet gas stream, which does not contain H2S, and so it has not been included in the

model.

After the sulphur guard the gas is mixed with steam in a static mixer (30-MX-

002) in a molar ratio of 3.348 mol steam to 1 mol of inlet gas. This is accomplished by

the set function (SET-3). The actual molar ratio used is 3.6 mol of steam per mol of

methane. Methane makes up 93 percent of the inlet gas stream and so 93 percent of 3.6,

or 3.348, is used as the set point value. The steam is at a temperature of 330oC and is

produced by flue gas in the Steam Superheat Coil. An explanation of the Steam

Superheat Coil can be found later in this section.

The mixed gas stream from 30-MX-002 (To Convection Section) then moves to a

recycle function (RCY-2). The recycle function allows for the HYSYS model to iterate

until interdependent parts of the model converge. In this case the interdependent parts of

the model are the inlet gas stream and three of the four convection section heat

exchangers. After the iteration is complete the stream (To Convection Section-2) enters

the third preheat heat exchanger (Mixed Feed Preheat Coil).

11


The next heat exchanger heats the mixed feed stream from 350oC to 450oC before

entering the reformer. The exchanger, again, does not model Ft correction factor, heat

loss, or pressure drop. The exchanger heat duty was estimated at 14.32 MW. From the

duty spec and outlet temperature a UA value of 151,700 kJ/oC-h was calculated for the

exchanger. After the feed has been heated the gas enters the reformer furnace on the tube

side.

The tube side of the reformer furnace (30-F-001-A) is modelled as a conversion

reactor. The reformation reaction converts hydrocarbons and steam to CO and H2. A

reformer shift reaction also occurs within the reformer which converts CO and steam to

12

CO2 and H2. The components and mole balances for the multiple reforming reactions

and the shift reaction can be seen in Appendix E..2 and E2.2.

The reforming reaction is assumed to occur for all of the higher hydrocarbons

with 100 percent conversion. Methane is converted with 72.8 percent conversion and the

shift reaction is assumed to have 58 percent conversion. These conversions were

determined by trial and error in order to have the proper methane concentration in the

stream out of the reactor (To Waste HEX) on a water free molar basis, as given in the

simplified process flow diagram provided by Husky. The water free composition is

checked by splitting the water from the stream and checking the composition. This is

done by the hypothetical splitter (X-100) and checked with the top stream. (Hypothetical)

The stream is then remixed (MIX-100) before being sent to the waste heat exchanger.

(To Waste HEX2) The liquid stream on the reformer (DNE2) does not exist but is

required by HYSYS to operate.


13

The waste heat exchanger (E-103) cools the gas from 795oC to 365oC by using

steam from the 4500kPa steam drum. The steam flow is set at 5600 kgmol/h or 100,900

kg/h and was specified because it gives a saturated vapour at the exit of the exchanger.

(To SD1) The steam enters as a saturated liquid at 4500 kPa and 258.5oC and exits the

exchanger at 259.9oC as a vapour. The exchanger itself has a tube side pressure drop of

1200 kPa based on the flow diagrams and a shell side assumed pressure drop of zero. No

heat loss modelled and the Ft correction factor is not calculated. Again the default area of

heat transfer is 60 m2 and the UA value is 493,200 kJ/oC-h. The UA value was found by

using the outlet temperature gas stream and specifying a value for the steam flow rate.

From the reformer waste heat exchanger the gas then moves to the shift converter.

The shift converter (30-R-004) is a reactor that uses the shift reaction seen above

to convert CO and steam to Hydrogen and CO2. The reactor is modeled as a conversion

reactor, again because kinetic reaction data was not available. A conversion of 65.5

percent gives the proper composition for the tail gas stream after the PSA unit (PSA Tail

Gas-1) and so it is used. The outlet temperature of the gas stream is 403oC, as calculated

by the heat of reaction. The liquid flow (DNE3) out of the reactor does not exist but must

be included for HYSYS to operate. The shift gas now moves to through a pair of heat

exchangers.

The next step in the process is a series of two heat exchangers which, in order, are

the shift waste heat exchanger (30-E-004) and the boiler feed water pre-heater. (30-E-

005) Both exchangers do not model heat loss and do not have a pressure drop over either

the tube or shell side. The water flow rates are 230,000 kg/h and 125,000 kg/h for the

shift waste heat exchanger and the boiler feed water pre-heater respectively. The water

14

for the shift waste heat exchanger (From P1) enters at a specified 125 oC and exits at

162.1oC. The boiler feed water (From P2) enters at the same specified temperature of

125 oC and exits at a temperature of 174.7 oC. The overall heat transfer coefficients, UA,

are 204,600 kJ/oC-h and 421,700 kJ/oC-h for the shift waste heat exchanger and the boiler

feed water pre-heater respectively. Both the inlet temperatures and the flow rates of both

exchangers are specified in the Husky process flow diagrams. The pressure of both water

streams is 5200kPa which is specified in the process P&IDs.

The gas stream proceeds from the boiler feed water preheater to the first

condensate drum, the hot condensate drum. (30-D-007) The drum is modelled as a

component splitter that removes water from the stream. The splitter removes 2 percent of

the water in the stream and sends it to the dearator which is not shown in the HYSYS

model. The stream exits at 177.1oC and proceeds to the fin fan cooler section.

The raw gas coolers (30-E-007 A/B/C/D/E/F) are a set of large fans that lower the

temperature by moving air over a large number of tubes. The coolers are simply

modelled as a cooler. The cooler lowers the temperature of the gas stream from 177.1oC

to a 50oC, as stated in the Plant 30 documents supplied by Husky. The model predicts a

power of 40.67 MW which is feasible because they are designed for 56.72 MW. The

power was specified by setting the outlet temperature of the stream to 50 oC and copying

the calculated power.

15


After the raw gas coolers the next step is through a second condensate drum. (30-

D-080) The drum, like the other drum, is modelled as a component splitter and removes

50 percent of the water in the stream, which amounts to 24,140 kg/h. The number is just

an estimate since no flow data was supplied and the flow will not have an appreciable

effect on the rest of the system. The next stage in the process is another heat exchanger.

The heat exchanger that follows the condensate drum is the raw gas trim cooler.

(30-E-008) The cooler uses 20oC water at 550kPa to lower the gas temperature from

50oC to 23oC. The water used, according to the model, flows at a rate of 667,700 kg/h

which may not be accurate since the water removed by the condensate drum may not be

accurate. The water on the outlet of the exchanger is specified in the flow diagrams as

16

being 23oC. A pressure drop of 99kPa on the tube side and 210kPa on the shell side are

assumed from the process flow sheets. The exchanger, like all of the others within the

model, uses a default heat transfer area of 60m2 and has a UA value of 778,800 kJ/oC-h.

The next step for the gas is the final cold condensate drum. (30-D-008) The drum

is again modelled as a splitter and removes all of the water left in the stream. The water,

like in the last condensate drum, goes to the dearator which is not modelled at a flow of

241,400 kg/h. The gas then moves to the PSA unit where the gas is purified.

The pressure swing absorption unit (30-PK-004) is made up of 13 process vessels

that purify the Hydrogen. The unit is quite complex and for this reason it was decided

that it should be modelled as a mole and heat balance. The Hydrogen product stream

(Hydrogen Product) is assumed to be 99.5 percent pure, with the remaining .5 percent

being CO. The stream is specified as having a temperature of 25oC and a pressure of

2180kPa. The flow rate of the stream is determined from the inlet gas flow rate using the

set function (SET-2) which makes the molar flow of the outlet hydrogen stream 2.48

times the molar flow of the inlet gas stream; a specification which was provided by

Husky. The PSA tails gas (PSA Tail Gas-1) has a specified pressure of 40kPa and a

calculated temperature of 17.29oC. The flow rate of the gas is approximately 48,400 std

m3/h which does not agree with the value specified by Husky of 35,000 std m3/h, this is

discussed later in Section 3.1.2. The compositions of the stream are determined by the

mole balance and reflect the compositions of the actual gas stream. The removed CO2

(CO2 Out) also attaches to the PSA unit in the model simply because the tie in point for

the CO2 removal system is directly before the PSA unit and the system does not have an

appreciable affect on the stream composition beyond removing CO2.

17

The tail gas from the PSA unit still has a large heating value and so is sent back to

the reformer furnace as a feed for the burners. The tail gas can be sent to vent if it is

deemed necessary and so a tee (TEE-100) has be placed in the model. The tee currently

assumes that no gas is sent to vent, however it is available if more accurate process

information is obtained. The tail gas then moves through a second recycle function.

(RCY-1) The function, like the other recycle, allows HYSYS to iterate. In this case the

recycle function is used to ensure that the temperature around the loop that begins and

ends with the mixed feed preheat coil (Mixed Feed Preheat Coil) is consistent. The tail

gas (PSA Tail Gas-2) is now sent to the reformer burners.

The furnace part of the reformer (30-F-001-B) is modeled, in a similar fashion to

the tube side of the reformer, as a conversion reactor. The combustion section of the

reformer has three main feed streams: the PSA tail gas, a combo gas make-up stream, and

combustion air. The combo gas (Combo Gas) enters at a temperature of 55oC and a

pressure of 146.3kPa. The flow rate is also specified at 3900 std m3/h. The composition

of the combo gas was provided in a similar format to that of the inlet gas stream and can

be seen below in Figure 7.

18

Figure 7: Specified Combo Gas Composition

The combustion air (Combustion Air) is assumed to be 79 mol percent Nitrogen

and 21 mol percent Oxygen. The air is preheated to 300oC in a section prior to the

reformer and is not modeled. The pressure of the stream is 102.8kPa. The gas flow is

regulated by an adjust function (ADJ-1) which changes the flow rate of the combustion

air so that 2 percent oxygen is in the reformer flue gas stream. (Waste Gas) The flow rate

of the air predicted by the model is 194,000 kg/hr which is the same value that Husky

provided for the stream. Once the three feeds are added to the reformer the combustible

products are converted. All of the combustible products are assumed to have 100 percent

conversions in the reactor. The reaction stoichiometry for the reformer furnace can be in

Appendix E1.2 or E2.2 in the HYSYS reports.

19

The energy from the reformer combustion section is used for the energy for the

tube side of the reformer. The energy from the combustion reactions that is not used to

heat the gas stream to the reformer’s 800oC temperature is shown by HYSYS as an

energy stream. (Q-100) There is also a similar stream (Q-102) providing energy to the

tube section of the reformer in 30-F-001-A. At steady state the energy produced by the

combustion reaction should be equivalent to the energy consumed by the tube side of the

reformer plus the heat loss to the atmosphere. This is taken into account by setting the

energy of the tube side stream using a set function (SET-1) with a less than one multiplier

that acts as an efficiency factor. The heat transfer efficiency factor was set by adjusting

the value until an outlet temperature of 795oC was observed from the tube side reformer.

(To Waste HEX) The factor was found to be .8641 which, according to Ulrich, is in an

acceptable region for an industrial furnace (Ulrich, 149). It was assumed that this factor

would remain constant when the CO2 was removed because the reformer should be

running at the same temperature and therefore should have similar heat transfer within the

tubes.

The flue gas from the combustion section of the reformer is at a high temperature

of 800oC and so is used for heating other process streams. The gas has a flow rate of

246,600 kg/h and is at a pressure of 40kPa. The composition of the stream is

approximately 17 percent CO2, 18 percent H2O, 63 percent Nitrogen, and 2 percent

Oxygen. The gas first moves to steam generation coil I (Steam Generation Coil I) to

produce steam for the steam drum 30-D-003. The water enters as a saturated liquid at

5200kPa with a flow rate of 441,700 kg/h as specified on the P&IDs. The exchanger, like

all of the others, does not model heat loss or the Ft correction factor. There is no pressure

20

loss over the tube side which contains the flue gas and a 50kPa drop over the shell side,

again specified in the P&IDs. The UA value is specified at 91,180 kJ/oC-h which is

based off of the heat duty specification of 11.78MW. The outlet gas (WG to MFPC) has

a temperature of 668.4oC. The outlet steam has a vapour fraction of .0512. The gas

stream then moves through the mixed feed preheat coil, which is described above, and is

cooled to 558.3oC. The next step is another steam generation exchanger.

The second steam exchanger (Steam Superheat Coil) takes steam from the steam

drum 30-D-003 and heats it so that it can be used elsewhere and to mix with the inlet feed

gas in 30-MX-002. The exchanger does not have a pressure drop for the shell or tube

side, it does not model heat loss, and the Ft correction factor is not calculated. The UA

value of 80,050kJ/oC-h was found by specifying the flow of steam to elsewhere in the

plant, (To Steam Turbines) the flow rate of steam into the mixer, (Steam) and the

temperature of the steam out of the exchanger at 330oC as specified in the process flow

diagram. Once the UA value was determined an adjust function (ADJ-5) was introduced

to change the inlet flow rate of the stream until the temperature of the outlet steam was

330oC. The flow specification on the other stream was also removed so that changes in

the inlet flow rate would not cause an error. In the exchanger the gas is cooled to

498.5oC. The gas then moves through the first heat exchanger (Hydrocarbon Preheat

Coil) and is cooled to 432.2oC.

The flue gas’s final destination is steam generation coil II (Steam Generation Coil

II) which cools the gas to 332.9oC. The heat exchanger, in the nature of this simulation,

does not show pressure drop over the shell or tube side, does not calculate the Ft

correction factor, and does not model heat loss. The UA value of 273,500kJ/oC-h was

21

determined from the heat transfer specification that 8.26MW of power are transferred to

the steam stream. The vapour fraction of the outlet stream predicted is .0917.

Several additional adjust functions exist within the model, each with its own

purpose when adjusting the model after the CO2 is removed from the system. First is the

adjust function (ADJ-4) for the stream before the sulphur guard. (Past Sulphur Guard)

Although the sulphur guard is not modeled in the simulation the temperature through the

guard must be maintained or the reaction, if one is needed, to remove the sulphur will not

occur. The adjust therefore adjusts the combustion air flow until the proper temperature

of 350oC is obtained. This works against the adjust to maintain the flue gas composition

at 2 percent Oxygen (ADJ-1) and therefore if ADJ-4 is being used ADJ-1 should be

ignored and vice versa. Another adjust function used after the CO2 is removed connects

the combo gas and the outlet flow from the tube section of the reformer. (To Waste HEX)

This adjust function (ADJ-2) adjusts the combo gas flow rate so that the temperature of

the reformer outlet is 795oC, the normal exit gas temperature. Once the CO2 is removed

the combo gas flow rate will have to be changed so that the outlet temperature of the

tubes is maintained. This lowering of combo gas flow rate is the source of revenue for

this project and so the adjust function allows for it to be accurately predicted.

Once the CO2 is removed several process conditions change. These changes

generally are temperatures and flow rates in the initial section of the simulation leading

up to the tube side reformer. The flow rates of the combo gas and combustion air are also

changed to make sure process conditions are maintained as much as possible, specifically

the outlet temperature of the reformer and the temperature of the sulphur guard stream.

22

Detailed analysis of these changes can be seen in the HYSYS reports in Appendix E2. A

summary of the major process changes can be seen below in Table 1.

Table 1: Summary of Process Changes

Initial (kg/hr)

With CO2 Removed (kg/hr)

Savings (M $CAD)

Combo Gas 982.2 831.2 1207784 Combustion Air 194000 232580.217 -306873 Steam Produced 1.5787E+05 1.5771E+05 -7.69E+04

CO2 Emitted 63998 27760 n/a

3.1.2 Tail Gas Flow Rate Assumption

The major assumption made by the model revolves around the flow of the PSA

tail gas stream. The value for the tail gas was provided as being 35,000 std m3/h

however the model predicts a value of 48,400 std m3/h. The issue was discussed with our

industrial contact and he said to assume the losses were due to “leaks in the reformer”

and to send the appropriate amount to vent. It was decided to not follow these directions

for the following reasons:

1. Despite there being many leaks in a reformer, especially because it operates at

below atmospheric pressure and viewing ports are not completely sealed, it does

not seem feasible that over 25 percent of the stream is being sent to vent. A leak

of this size would have been noticed at some point during the operation and it

would be a major safety concern.

2. According to the model the flow rate of 35,000 std m3/h does not produce the

energy needed to heat the stream from 450oC to 795oC, the actual normal

operating temperature of the tube side reformer outlet.

23

3. The combustion air flow rate predicted by the model agrees with the higher tail

gas flow rate and matches both that flow rate and the outlet gas composition of

two percent excess oxygen. The air flow at the measured composition is not

matched at 35,000 std m3/h.

4. The measurement for the tail gas flow occurs on the FD fan in the reformer, but

the P&IDs do not indicate if that flow rate is used to control the operation or if the

outlet gas composition is. It is likely the outlet gas composition that is used since

a feedback loop could be applied. This means that a broken or improperly

calibrated flow meter on the tail gas could go undetected by the operators.

5. The specifications given for the Hydrogen gas stream, inlet gas stream, and

composition of the tail gas stream fully define the flow of the PSA tail gas. As

can be seen below the reformer process up to the PSA unit can be seen as a black

box process. If a balance of Carbon atoms is done over the process it can be seen

that since the inlet flow rate and composition is defined, as is that of the Hydrogen

gas outlet, and the compositions of the tail gas stream the flow rate is also defined.

Also seen below is a rough calculation using the molar flow of methane as being

the only source of carbon with the other hydrocarbons assumed as being

negligible.

24

Figure 8: Plant 30 Block Diagram

hm std

3

333

33321

3

424

43,990

gives Conversion 1860

)066(.)176(.)445(.)005(.3456)93)(.1393(

hkgmol

hkgmol

hkgmol

COCHCOCOCH

F

FFF

yFyFyFyFyF

=

+++=

+++=

The other hydrocarbons in the natural gas feed will increase the flow of the

stream, however it can be seen that the flow is fully defined by the conditions

given.

Natural Gas in 93% CH4

F1 = 1393 kgmol/h

H2O Out

H2O In H2 Out

99.5% H2 .5% CO

F2 = 3456 kgmol/h

PSA Tail Gas 44.5% CO2 17.6% CH4 6.6% CO

F3 = 3456 kgmol/h

25

3.2 AMSIM Simulation

The design of our CO2 extraction process using an amine solution was completed

using a computer simulation as recommended by Seader. It is necessary to use a

computer simulation to accomplish the task of modeling the CO2 extraction system

because the absorption of gas into the solution is a reversible chemical reaction.

AMSIM, AMine treating unit SIMulator, was utilized to simulate and design the

CO2 extraction process. AMSIM is a software package developed by Schlumberger that

simulates the steady state removal of acid gas from process streams using aqueous amine

solutions and physical solvents. In our application no H2S was present in our process gas

stream, therefore only CO2 was absorbed by the process.

In order to calculate the mass transfer process of the amine treating unit, AMSIM

uses a non-equilibrium stage model. The fundamental concept that AMSIM uses is that

the rate of absorption and desorption is considered as a mass-transfer rate process.

(Schlumberger) To simultaneously solve the non-linear stage equations for temperature,

composition and phase rates for each stage of the column AMSIM uses a modified

Newton-Raphson method. The Kent and Eisenberg approach is used as a basis to model

the equilibrium solubility of acid gases in amine solutions. AMSIM also validates the

solubility model with experimental data and proprietary information.

Utilizing AMSIM the following process was designed to remove the CO2 from

the reformer gas stream. The tower contacting the aqueous MEA with the gaseous

reformer stream is illustrated as Absorber A. Absorber A is simulated to be 9.14 m high

and 2.74 m in diameter consisting of 15 bubble cap trays and operates at 2164 kPa and

38oC. Bubble cap trays were chosen to allow a longer interaction time of the amine

26

solution and process gas resulting in a more complete absorption. The amine solution,

which is 30% by weight MEA, circulates at a rate of 407 m3/hr.

Following Absorber A the rich amine is depressurized in a vertical flash drum

sized to be 6.75 m high and 2.25 m in diameter. 31 kmol/hr of gas is flashed off and sent

to the PSA tails gas. The rich amine then flows through a lean/rich amine plate and

frame heat exchanger increasing the temperature from 69.6 oC to 83.5 oC. Once the rich

amine is heated it flows into the regenerator tower. The regenerator tower which is used

to regenerate the rich MEA and capture the absorbed CO2, is 10.36 m high and 3.2 m in

diameter consisting of 17 sieve trays. Sieve trays were chosen as they resulted in the

greatest amount of CO2 captured when simulated with valve or bubble cap and are the

most cost effective. The overhead vapour flow from the regenerator is 1221kmol/hr

which consists of 67% CO2 and 33% H2O on a molar basis. An overhead condenser

condenses the water using a plate and frame heat exchanger with a surface area of 261

m2. Following the condenser 864.9 kmol/hr of acid gas is captured consisting of 94.1%

of CO2 and 5.8% water. Once the water is removed the CO2 can then be used by Husky

for enhanced oil recovery.

Once the amine is regenerated in the tower the lean amine flows through the

lean/rich amine heat exchanger decreasing the temperature from 125 oC to 107 oC. The

lean amine is further cooled by another plate and frame heat exchanger that is sized to be

203 m2. This solvent cooler reduces the temperature of the amine to 37.8 oC before re-

entering the contactor tower, Absorber A.

27

Figure 9: Amine System Flow Diagram

28

Absorber B, the MEA Guard, was added to the simulation in order to ensure no

MEA would be entrained in the process gas stream and contaminate the PSA unit. This

tower uses water as a solvent to absorb MEA from the lean gas stream. The water is

recycled until the MEA concentration increases to 7% by weight at which time it can then

be replaced. The MEA guard is modeled to be 7.31 m high and 1.68 m in diameter using

12 bubble cap trays. The lean gas stream then moves to a water knockout drum which

removes any water in the stream. The lean gas stream then moves to the PSA unit for

Hydrogen purification.

29

4. Equipment Specification and Design

The equipment sizing was done using the text by Ulrich. Sample calculations are

summarized in Appendix A. Table 2 is a summary of the equipment sizing. EconExpert,

a Ulrich based web tool, was used along with the September 2007 Final CEPCI value of

528.2 (Chemical Engineering Journal, Jan 2008) to find the Total Capital Cost for the

new amine system. This cost was determined to be approximately $9.82 million. A

break down of the equipment costs are shown in Table C.1 and a summary of EconExpert

results is shown in Appendix C.

Table 2: Summary of Equipment Sizing and Specifications

Towers Diameter (m) Height (m) # of Trays Material Operating P (barg)

Absorber 2.74 9.14 15 St-St 22 Regenerator 3.20 10.36 17 St-St 2.2 MEA Guard 1.68 7.32 12 St-St 2.1

Heat Exchangers Type Sub Type Surface Area (m2) Material

Regenerator Reboiler Shell & Tube Kettle Reboiler 481 St-St

Regenerator Condenser Plate&Frame Flat Plate 261 St-St Lean/Rich Amine Plate&Frame Flat Plate 1725 St-St Solvent Cooler Plate&Frame Flat Plate 203 St-St

Pumps Pump Shaft Power (kW)

Suction P (barg) Type Material

Pump REGEN to ABS (220 to 2164 kpa) 15.8 2.2 Centrifugal St-St Pump for MEA Guard (2100 to 2150 kpa) 1.06 2.1 Centrifugal St-St

30

Storage Vessels Volume (m3) Material

Amine Holding Tank 33 St-St

Process Vessels Orientation Sub Type Diameter (m)

Height (m) Material

Amine Flash Drum Vertical no packing or trays 2.25 6.75 St-St

Condenser Drum Vertical no packing or trays 1.40 2.50 St-St

31

5. Plant Safety Analysis

A plant safety analysis was conducted for the proposed amine plant. A full report is

included in Appendix B.

The Plant Safety Analysis conducted includes:

• HAZOP of the overall plant design

• Inherently Safe Design

• Process Safety Management System

• Chemical Hazard Information and MSDS

• Dow Fire and Explosion Index

Highlights of the HAZOP recommendations include the selection of stainless steel

material because MEA is highly corrosive. HAZOP analysis can be found in Appendix

B.2. MSDS and Chemical Hazard Information for the materials in the stream are

included in Appendix B.6 and Table B.1. Other recommendations include the addition of

secondary pumps in situations of no-flow. As well, the use of inert gas to purge storage

and process drums to reduce corrosion, amine degradation, and prevent build up of

explosive mixtures. Finally, minimizing the amine flow (407m3/hr) reduces hazardous

material handling onsite.

32

The Process Safety Management System highlights include employee safety training

in the areas of personal protective equipment, steam use, hazardous material handling for

MEA, gas detection, and emergency evacuation.

Documentation is important especially in new employee cases and the inclusion of

operating manuals, equipment specifications, and frequency of maintenance are

important. These are summarized in Appendix B.4 and B.5.

Inherently Safe Design analysis using the concept of Intensification is summarized in

Appendix B.3. The amine system is designed for a minimum amine flow and this

minimizes the amount of amine on site.

The Dow F&EI determined that the amine system would have a degree of hazard of

166 which is a severe risk since it is greater than 159. Since the reformer furnace uses the

same chemicals in similar amounts Husky should already be capable of dealing with the

risk. A summary of the analysis is shown in Appendix B.7.

33

6. Economic Analysis

The economics for this project were assessed based on methods taught in an

engineering economics class. The capital cost for the project was determined in Section

4. The method for determining the operating costs, savings, depreciation, and taxation

can be found in this section.

Operating costs were a major component of the overall cash flow for this project.

The cost of the electrical power needed to run the motors for the pumps was 5.062¢/kW-

h. (Saskpower) The system requirements for power were calculated to be 16.86 kW

resulting in a cost of $7,500 per year. A cost of $323,000 of MEA was calculated based

on 70 m3 of MEA in the process at a cost of MEA of $6.86 per liter. The cost of steam at

the Husky Lloydminster Upgrader is $28.86 per ton. The steam necessary in the reboiler

to regenerate the amine is 64,140 kg/hr resulting in a cost of $16.2 million per year. An

additional operating cost of $253,000 per year resulted from a cost of additional operators

based on the additional equipment.

Depreciation for this project was determined using a sum of the years digits

approach. The depreciation schedule was set so that the plant would be fully depreciated

after 10 years. The values for the depreciation can be seen along with the overall cash

flow in Table D.1.

34

The savings for this system were calculated based on the HYSYS model

explained in Section 3. After the CO2 has been removed several changes are made to

operating conditions at various points in the plant. These changes result in both direct

and indirect costs. The main savings are associated with a reduced need for combo gas

within the process. The main losses are due to lower steam production and increased

need for combustion air. Calculations for these amounts can be seen in Appendix A.8

through A.12. The combo gas savings amount to $1.208 million per year. The expenses

from steam losses and combustion air are $59,000 and $307,000 per year respectively.

Combining these values gives total revenues of $859,000 per year.

The cash flow analysis was based off several different specified conditions. First

it was decided to use a one year build time so the separation system begins operation at

the beginning of year one. The operating costs are higher in year one because additional

amine must be purchased to fill the tower in year one, whereas only a portion of the

amine is replaced in the subsequent years and is estimated as the amount in the amine

holding tank. The tax rate used was 41% which is based on 28% federal tax and 13%

provincial tax. (Canada Revenue Agency) The interest rate used to discount the cash

flow is 8% which is simply based on what was decided to be an acceptable MARR value.

The results of the cash flow analysis can be seen in Table D.1 or Figure 10.

Due to the high operating cost the system shows a negative cash flow. For this

reason an internal rate of return was not calculated. Also, because of the negative cash

flow, it was decided to place a value on the CO2 to determine at what cost the system

would become feasible. Several values were placed on the removed CO2 and the results

can be seen in Figure 10. The important points to note are the point where the project’s

35

revenues match the operating costs at a combined carbon dioxide tax and value of $50.43

per tonne and the point at which the project has a break even point of 25 years, at a value

of $67.33 per tonne.

‐140

‐120

‐100

‐80

‐60

‐40

‐20

0

0 5 10 15 20 25

Time (Years)

Cumulative Discoun

ted Ca

sh Flow ( MM CAD$)

0

10

20

30

40

50.43

67.33

Figure 10: Cumulative Discounted Cash Flow at Different Carbon Tax Rates in $/tonne CO2 Emitted

36

7. Conclusions and Recommendations

A HYSYS model of the existing Hydrogen plant was constructed and used to

predict the effect of CO2 removal. The model predicted savings of $859,000 per year

based on reduced combo gas needed as fuel for the reformer, additional combustion air

needs, and slightly lowered steam production.

A CO2 extraction process utilizing a Monoethanolamine based solvent was

designed and simulated using AMSIM. The system is successful in capturing 35.9

tonnes/hr of CO2 which comprises 91.3% of the CO2 in the original stream. The total

capital costs necessary for this project is estimated at $9.82 Million US. Using Husky’s

current prices this project will reach a break even point after 25 years with a combined

value or tax on CO2 of $67/tonne.

If Husky wants to explore this project further RPM recommends that an internal

study should be conducted to determine the monetary value they can place on the CO2.

If the value of CO2 is economically feasible it is recommended that detailed design be

done for this project.

37

8. References Airgas. 23 February 2008 <http://www.airgas.com/documents/pdf/1013.pdf>.

BOC. 23 February 2008 <http://www1.boc.com/uk/sds>.

Canada Revenue Agency. “Corporation Tax Rates.” Canada Revenue Agency. 2008.

CRA. 02 March 2008 <http://www.cra-

arc.gc.ca/tax/business/topics/corporations/rates-e.html>.

J. D. Seader, Ernest J. Henley. Separation Process Principles. United States of America:

John Wiley & Sons, Inc., 2006.

Maddox, Dr. Robert N. and D. John Morgan. Gas Conditioning and Processing Volume

4: Gas Treating and Sulfur Recovery. Oklahoma: Campbell Petroleum Series,

2006.

Saskatchewan, University of. "CEPCI Value." Chemical Engineering Journal (January

2008).

SaskPower. 02 March 2008

<http://www.saskpower.com/services/busrates/oilfields/doc1.shtml>.

Schlumberger. "AMSIM User's Manual." (2003).

Scott, K. Industrial Membrane Seperation Technology. New York: Chapman and Hall,

1996.

Ulrich, Gael D. Chemical Engineering Process Design and Economics a Practical Guide.

Durham : Process Publishing, 2004.

38

University of Queensland Australia. 23 February 2008

<http://www.cheque.uq.edu.au/ugrad/theses/1998/DaveA/dow.html>.

Valley National Gas Website. 23 Feburary 2008 <http://www.vngas.com>.

Whitaker Oil Company. 23 February 2008

<http://www.whitakeroil.com/MSDS/MEA.pdf>.

39

Appendix A: Sample Calculations

40

The calculations from Ulrich may be slightly different compared with EconExpert

due to the subjective reading of values from the graphs in the textbook.

A.1 Membrane Size

The flux of the CO2 through a membrane was calculated to see if a membrane

system would be feasible. The permeability value is based off of values given in Scott’s

“Industrial Membrane Separation Technology,” the partial pressure defined by the

HYSYS simulation, and an assumed membrane thickness.

smmolj

m

kPakpabarrerbarrerj

LppPj

CO

kPaPasPam

mmol

CO

LiCO

24

16

10874.5

001.

)01000650)(10347.3)(2700(

)(

2

2

2

2

−

⋅−

×=

−⋅×=

−=

A CO2 removal of 600 kgmol/h, about 70% of the CO2 in the stream, was used as

a basis to judge if the system could meet removal specifications.

2

2

24

284000283736

10003600

160010874.5

2

2

mAmA

Asmmol

AF

j

kgmolmol

sh

hkgmol

COCO

=

=

⋅⋅=×

=

−

41

A.2 Absorber Size and Cost

Using Figure 5.44b pg. 387 in Ulrich,

Assuming a vertical vessel with the following data taken from the AMSIM

nact = 15

D = 2.743 m

s = 0.6096 m

Hcol = 9.144 m

From Figure 5.44b gives a purchased equipment cost, cP = $8.5 x104.

Since P < 4 barg use Figure 5.45.

Using stainless steel material due to the corrosive MEA material gives a Material

Factor, FM = 4.0 and a P = 22 barg gives a Pressure Factor, FP = 2.25.

Therefore, the Pressure Factor-Material Factor product, FP x FM = 9.

Using Figure 5.46 with a vertical vessel orientation gives a Bare Module Factor,

FaBM = 18.

Figure 5.44 states that:

CBM = cP x FaBM

CBM = ($8.5 x 104)(18)

C BM = $1.53 x 106

42

Costing the trays using Figure 5.48 gives a Bare Module Cost,

CBM = (cPss x FBM) * nact * fq

CBM = ($4 x 103 ) * (2.2) * (15) * (1.125)

CBM = $1.485 x 105

Total Cost of Absorber and trays = $1.53 x 106 + $1.485 x 105

Total Cost of Absorber and trays ~ $1.6785 x 106

A.3 Regenerator Size and Cost

Using Figure 5.44b pg. 387 in Ulrich,

Assuming a vertical vessel with the following data taken from the AMSIM,

nact = 17

D = 3.2 m

s = 0.6096 m

Hcol = 10.3632 m

From Figure 5.44b gives a purchased equipment cost, cP = $9.2 x104

Since P < 4 barg use Figure 5.45

Using stainless steel material due to the corrosive MEA material gives a Material

Factor, FM = 4.0 and a P = 0.206 barg gives a Pressure Factor, FP = 1.

43

Therefore, the Pressure Factor-Material Factor product, FP x FM = 4.0.

Using Figure 5.46 with a vertical vessel orientation gives a Bare Module Factor,

FaBM = 9.5.

Figure 5.44 states that:

CBM = cP x FaBM

CBM = ($9.2 x 104)(9.5)

C BM = $8.74 x 105

Costing the trays using Figure 5.48 gives a Bare Module Cost,

CBM = cPss * FBM * nact * fq

CBM = ($5.5 x 103 ) * (2.2) * (18) * (1.17)

CBM = $2.5483 x 105

Total Cost of Regenerator and trays = $8.74 x 105 + $2.5483 x 105

Total Cost of Regenerator and trays ~ $1.1288 x 106

44

A.4 Condenser Heat Exchanger Size

Q = UA∆Tlm

Where the following data was taken from AMSIM,

Thi = 89.4oC

Tho = 48.9oC

Tci = 23oC

Tco = 49oC

CP = 4.501 kJ/kgoC

( ) ( )( ) ( )

( ) ( )( ) ( )

CTCCCCT

TTTTTTTT

T

olm

oo

oo

lm

cihocohi

cihocohilm

6.32)239.48/494.89ln(

239.48494.89

)/ln(

=

−−−−−

=

−−−−−

=

The surface area, A:

)6.32)(2200(

W x1018695.89 A

2

3

CCm

W°

°

=

A = 261 m2

45

A.5 Centrifugal Pump Size and Cost (Regenerator to Absorber 220 to 2164 kPa)

Using equation 4.94 pg. 248 in Ulrich,

sm

hrmq

mskgxP

msm

mkgghP

PqWi

s

33

25

23

11306.0407

100461.1

)3632.10)(81.9)(939.1028(

==

=Δ

==Δ

Δ∗=

•

••

ρ

ε

Assume efficiency, iε = 0.75

Therefore, shaft power, •

sW = 15.8 kW.

Using Figure 4.2 pg.121 in Ulrich gives electric motor efficiency, ed,em = 0.93,

and gas turbine efficiency, ed,gt = 0.26.

Power, P, is calculated from equation 4.95 pg. 248 in Ulrich,

kW 64.2)26.0)(75.0(

)100461.1)(11306.0(

kW 18)93.0)(75.0(

)100461.1)(11306.0(

25

3

,

25

3

,

==Δ∗

=

==Δ∗

=

•

•

mskgx

sm

pqP

mskgx

sm

pqP

gtdigt

emdiem

εε

εε

Using Figure 5.49 pg. 390 in Ulrich,

46

Therefore, shaft power, •

sW = 15.8 kW yields a Purchased Cost (electric motor

included), cP = $1.7 x 104 with FM = 1.9 for stainless steel.

The calculation is based on a radial/centrifugal pump “because axial flow or

regenerative units are ultimately negligible on economics” (Ulrich 248).

Using Figure 5.50 pg. 391 in Ulrich,

Suction Pressure, Pi = 2.2 barg give a Pressure Factor, Fp = 1.

Using Figure 5.51 pg. 391 in Ulrich gives the Pressure Factor-Material Factor

Product, FP x FM = 1 * 1.9 = 1.9 which give a Bare Module Factor, FaBM = 4.9.

Using Figure 5.49 pg. 390 gives a Bare Module Cost,

CBM = cPss * Fa

BM

CBM = ($1.7 x 104)(4.9)

CBM = $8.82 x 104

47

A.6 Amine Holding Tank Size

Determine volume of Absorber and Regenerator and add 25% for piping and

assume need a tank to hold 20% of the amine volume in the system.

Volume of Absorber:


D = 2.743 m

H = 9.144 m

3

22

04.54

)144.9(2

743.2

mV

mmHrV

ABS

ABS

=

⎟⎠⎞

⎜⎝⎛== ππ

Volume of Regenerator:


D = 3.048 m

H = 10.3632 m

3Re

22

Re

62.75

)3632.10(2

048.3

mV

mmHrV

gen

gen

=

⎟⎠⎞

⎜⎝⎛== ππ

Total Estimated Volume of MEA,

VMEA = 1.25 * ( ABSV + genVRe )

VMEA = 1.25 * ( 33 62.7504.54 mm + )

VMEA = 162 m3

48

If the tank is to hold 20% of the Total Estimated Volume of MEA the volume of

the tank, VT = 33 m3.

A.7 Depreciation

340,786,1$

869,824,9$12345678910

10Years of SumLeft Years

=

⋅+++++++++

=

⋅=

D

D

PD o

A.8 Combo Gas Savings

The combo gas savings are calculated on the basis of reduced energy consumption

per year and rely on the cost of natural gas per unit energy.

YearGJ

ComboGas

mkJ

hm

ComboGas

ComboGasComboGasComboGas

Eyearday

dayh

kJGJE

ductionHFE

188128

)1534)(.365)(24)(1000000

1)(35891)(3900(

Re%

3

3

=Δ

=Δ

⋅⋅=Δ

&

&

&

6

1

10208.1$

42.6$188128

×=

⋅=

⋅Δ=

Savings

SavingsCostNaturalGasESavings

GJyearGJ

ComboGas&

A.9 Combustion Air

The costs of combustion air are based solely on the amount of steam that is

needed to run the ID fans for the air and ignores the cost of heating the air. So for the

20% increase in air flow the associated cost can be seen below.

49

year

tonnehtonne

yearday

dayh

1

1

000,307$LossesAir

)365)(24)(19899)(.86.28)($1.6(LossesAir

Increase%CostSteamIn SteamLossesAir

=

=

⋅⋅=

A.10 Steam Losses

The losses associated with steam are calculated by using the internal steam cost

and the flow rate of steam in the vapour phase for the two steam flows that are affected

by the flue gas.

year

tonnehtonne

yearday

dayh

1

1

200,59$Losses Steam

)365)(24)(86.28($)]1)(503.6()0918)(.52.24[(Losses Steam

CostSteamFraction)Vapour InSteam(Losses Steam

=

⋅⋅=

⋅⋅Σ=

50

Appendix B: Safety Document

51

B.1 Amine Plant Design Criteria

• Design for minimal MEA in the outlet stream of the amine system to the PSA

unit. MEA in the PSA unit will degrade the catalyst used.

• Design for specified pressure to the PSA unit. The PSA unit runs at 2200 kPa and

is black boxed because it is beyond the scope of this project and there is not

enough information to determine effects to the PSA unit.

• Maximize removal of CO2 before the PSA unit to maximize savings.

• 30 wt% MEA is corrosive so stainless steel material will be used for the design of

the amine system.

• Design for a no flow situation (pump cavitation or failure) and purchase (2)

secondary pumps for the amine system.

52

B.2 HAZOP/Safety Considerations Identify Hazards by considering the following Process Parameters:

i) Flow:

- design for a no flow situation (pump cavitation/failure) and purchase

secondary pumps for the amine system

ii) Time:

- Consider designing tanks with level transmitters to prevent overflow (next

phase of project-detailed plant design and layout)

- Consider time and volume when fill a tank

iii) Frequency:

- Consider frequency of loading amine to the holding tank (next phase of

project-detailed plant design) how many times per year is necessary

iv) Mixing:

- N/A

v) Pressure:

- Design for specified pressure to the PSA unit . The PSA unit runs at 2200 kPa

and is black boxed (beyond the scope of the project). There is not enough

information to determine effects to the PSA unit.

- Absorber operates at high pressure/low temperature (2164-2200 kPa, 38-70C).

Consider piping that can withhold the high pressure to the flash drum and then

to the Regenerator

vi) Composition:

53

- 30 wt% MEA is corrosive and stainless steel material will be used in the

AMSIM design for the amine system

- High CO2 concentrations will come out of the tops of the regenerator and if

the project is to go ahead a CEI will have to be completed on potential of CO2

cloud forming around the plant

vii) Viscosity:

- liquid is used, no slurries or solids, only potential for freezing if system is

down, consider insulation, or placing knock-out tank indoors

viii) Temperature:

- Consider steam traps and safety training for use of high pressure steam in the

regenerator reboiler, steam will condense and when restarted can cause

condensate induced-water hammer

- Consider condensate induced –water hammer, training in prepping condensate

lines, making sure lines are properly drained of cooled liquid before

proceeding with any opening or prep of lines

- Regenerator operates at low pressure/high temperature (200-220 kPa, 90-

125C). Consider piping that can withhold the high temperature corrosive

amine to the Absorber.

ix) pH:

- 30 wt% MEA is corrosive and stainless steel material will be used in the

AMSIM design for the amine system

- “The amines in water solution are basic”. (Maddox)

x) Separation/absorption:

54

- Occurs inside absorber and regenerator columns. Consider using material that

can withstand high temperature and high pressure with corrosive amine.

Stainless steel material is chosen for vessels and equipment

xi) Level:

- Consider designing tanks with level transmitters to prevent overflow (next

phase of project-detailed plant design and layout)

- Consider time and volume when fill a tank, consider the dielectric constant for

amine (MEA) material

xii) Speed:

- Pumps have moving parts, consider operator maintenance manuals for the

pumps (next phase of project-detailed plant design and layout)

xiii) Information:

- Documentation such as operating manuals, equipment specifications, design

criteria, and MSDS to communicate details of amine system to new users

xiv) Reaction:

Amine+ �Amine + H+

CO2 + H2O �HCO3- + H+

H2O �OH- + H+

HCO3- �CO3= + H+

Absorption reaction is exothermic:

2MEA+CO2 �� MEACOO-+MEAH+

55

xv) Operation:

General Operating Problems :

Failure to Sweeten Gas:

1) “Solution circulation too low

2) Regenerator temporarily overloaded after foaming episode

3) Poor regeneration due to :

i. Tray Damage

ii. Too cold in reboiler

iii. Too low pressure

iv. Insufficient regenerator stripping

4) Foaming

5) Gas flow too high

6) Acid gas content too high

7) Leak in contactor dP cell

8) Amine concentration too low

9) Contactor pressure too low or temperature too high” (Maddox).

Corrosion:

- According to Polderman et al. “Corrosion will be most severe at places

where the highest concentrations of acid gases encounter the highest

temperatures. These points will include the amine-amine heat exchanger,

the stripping column and the reboiler.” (Maddox).

56

- “Stress corrosion is prevalent in amine systems. This generally is

associated with residual stresses which result from localized heating during

vessel construction, such as welds in absorbers, strippers and piping. Stress

relieving all major equipment and piping will help to alleviate stress

corrosion” (Maddox).

Solution Degradation:

- “Amine solutions will slowly oxidize when exposed to air or oxygen. The

products of these oxidation reactions are generally considered to cause

corrosion problems. The oxidation can be minimized by use of an inert gas

blanket on amine storage containers and surge drums” (Maddox).

- Consider purging storage containers and surge drums with inert gas to

reduce oxidation

Foaming:

- Foaming can cause several different problems. “Plant gas through put may

be severely reduced and sweetening efficiency may decrease to the point

that pipeline specifications cannot be met. Also amine losses may be

significantly increased” (Maddox).

- Causes of Foaming problems in amine units:

1) “Suspended solids including iron sulphide

2) Hydrocarbon liquids

3) Condensed hydrocarbons:

i) Dew point shift

ii) Retrograde behavior

57

4) Amine degradation products

5) Almost any foreign material such as corrosion inhibitors, valve grease,

or even impurities in make-up water

6) Heavily gas overloaded tower

7) Methanol buildup in regenerator

8) Coatings on some filter cartridges

9) Amine contamination during shipping

10) Excessive antifoam addition

11) Coatings on metal and plastic tower packing

12) Other and unknown causes” (Maddox).

General Considerations:

Inlet Scrubbing:

- Most troubles in the contactor section are from entrained solids or entrained

hydrocarbons

- Insufficient inlet scrub of the sour gas can cause foaming, corrosion, and

reboiler tube burnout because of excess amounts of foreign material in the

amine solution

- The process already has a sulphur guard in place to scrub the sour gas upon

inlet

Amine Losses:

- “A sweet gas scrubber will help eliminate amine losses from unexpected

foaming or surges.” (Maddox).

-

58

Piping Design:

- High velocity causes erosion of pipes

- “advisable to:

1) Maintain liquid velocity below 0.9 m/s [3 ft/sec] in all piping unless

stainless or other appropriate alloy is used.

2) Avoid the use of screwed fittings whenever practical.

3) Use welded fittings with long radius ells; avoid tees when possible.

4) When making up pipe with valves, instruments, etc., avoid the use of

dissimilar metals to avoid bimetallic corrosion” (Maddox).

59

B.3 Plant Safety

i) Safe Design:

See design criteria and HAZOP above

ii) Pollution Prevention:

- Project is to remove the CO2 and will be used to inject into the ground for the

tertiary method called enhanced oil recovery (EOR)

iii) Lifecycle Analysis of Products:

- N/A. Project is to remove CO2 and use of it is beyond the scope of this

project phase.

iv) Inherently Safe Design:

Goal: to eliminate all hazards in the process using the 10 concepts that follow:

1. Intensification:

Use very small amounts of hazardous material so that if there is a leak the

hazard will be small.

- Design to minimize the amine flow in the system and the amount of amine in

the holding tank to minimize the amount of amine on-site.

2. Substitution:

Replace hazardous materials with less hazardous ones.

- Design to use 30 wt% amine and dilute with water

3. Attenuation:

Use hazardous material under the least hazardous conditions.

- Design to minimize loss of amine to reduce the frequency of off-loading

corrosive amine (MEA) by tank truck

60

4. Limitation:

Limit the effects of failures by equipment design or by change in condition

rather than adding on protective equipment.

- N/A

5. Simplification:

Simpler plants provide fewer opportunities for error and less equipment that

can fail.

- Design based on a typical amine system

6. Knock on Effect:

Design so that a domino effect doesn’t happen.

- N/A. There is no inherent domino effect that could occur in the process.

7. Avoid Incorrect Assembly:

- N/A. There is no assembly required at this stage of the project (beyond scope

of this project phase)

8. Status Clear:

It should be possible to see, at a glance, if valves are open or shut, if levels are

ok, if correctly assembled.

- N/A. There is no process controls in the design at this stage of the project

(beyond scope of this project phase)

9. Control:

Control systems should be in place.

- N/A. There is no process controls in the design at this stage of the project

(beyond scope of this project phase)

61

10. Survival:

If a hazard occurs personnel should be protected.

- Determine a fire, explosion, and emergency evacuation plan

62

B.4 Process Safety Management System • Employee safety training which includes:

Personal Protective Equipment(PPE) for hearing protection, protective

clothing, eye protection, and proper footwear (steel toed CSA approved)

Hazards of Steam

Hazardous materials (eg. MEA) used on-site, how and where to locate the

MSDS sheets on the materials and on-site via National Fire Protection

Agency (NFPA) signs

Gas detector training

• Documentation such as operating manuals, equipment specifications, design

criteria to communicate details of amine system to new users

• Employee training in fire, explosion, and emergency evacuation plans

• Maintenance plans for equipment being installed

63

B.5 Chemical Hazard Information Definitions:

LD-50/LC-50:

Lethal Dose 50.

The dose that kills half (50%) of the animals tested.

LEL/UEL:

Lower/Upper Explosive Limit.

Is the limiting/maximum concentration (in air) that is needed for the gas to ignite

and explode.

TLV:

Threshold Limit Value.

The reasonable level to which a worker can be exposed without adverse health

effects.

IDLH:

Immediately Dangerous to Life or Health.

The exposure to airborne contaminants that is likely to cause death or immediate

or delayed permanent adverse health effects or prevent escape from such an

environment.

64

Table B. 1: Chemical Hazard Information Summary

Component LD-50 or LC-50 (route/species) LEL/UEL TLV/Exposure Limits Reactivity

Hydrogen H2 No known toxicological effects from this product

Flammability Range: 4-75 vol% in air

No known toxicological effects from this product

Can form explosive mixture with air. May react violently with oxidants.

Monoethanolamine MEA

Oral: believed to be >1.00-2.00g/kg(rat) moderately toxic Inhalation: Not determined Dermal: >1.00g/kg (rabbit) slightly toxic

Flammable Limits %: 5/17

6 ppm STEL-ACGIH; 3 ppm TWA-OSHA; 6 ppm STEL-OSHA; 3 ppm TWA-ACGIH

Reacts violently with: Air, water, heat, strong oxidizers, acids,

Methane CH4 No known toxicological effects from this product

Flammability Range: 5-15 vol% in air

No known toxicological effects from this product

Can form explosive mixture with air. May react violently with oxidants.

Carbon Dioxide CO2 IDLH: 40000ppm Non-Flammable

ACGIH TLV. STEL: 54000 mg/m3 15 minute(s). Form: All forms; STEL: 30000 ppm 15 minute(s). Form: All forms; TWA: 9000 mg/m3 8 hour(s). Form: All forms; TWA: 5000 ppm 8 hour(s). Form: All forms

Stable

Carbon Monoxide CO 1807 ppm/4H (rat) 12.5%/74% PEL-OSHA: 50 ppm TWA; TLV-ACGIH: 25 ppm TWA

Stable; Incompatible with oxidizers

Nitrogen N2 No known toxicological effects from this product Non-Flammable No known toxicological

effects from this product Stable under normal conditions

65

B.6 MSDS

B.6.1 Hydrogen, H2 (BOC)

67

B.6.2 Monoethanolamine, MEA (Whitaker Oil Company)

75

B.6.3 Methane, CH4 (BOC)

79

B.6.4 Carbon Dioxide, CO2 (Airgas):

85

B.6.5 Carbon Monoxide, CO (Valley National Gas Website)

91

B.6.6 Nitrogen, N2 (BOC)

93

B.7 Dow Fire and Explosion Index (Univ. of Queensland)

95

Appendix C: EconExpert Equipment Costing Results

96

Equipment Costing Source: Econ Expert: www.ulrichvasudesign.com/econ.html (usask05, design05)

C.1 Towers Absorber

Regenerator

97

MEA Guard

98

C.2 Heat Exchangers Regenerator Reboiler

Regenerator Condenser Plate & Frame HX

99

Lean/Rich Heat Exchanger

Solvent Cooler Plate & Frame HX to Absorber:

100

C.3 Pumps Pump from Regenerator to Absorber (220 to 2164 kpa)

Pump for MEA Guard (2100 to 2150kPa)

101

C.4 Storage Vessel Amine Holding Tank

C.5 Process Vessels Amine Flash Drum

102

Condenser Drum

103

Table C. 1: Ulrich Equipment Costing

Equipment Cost $ US

Towers Absorber 1420847 Regenerator 1115933 MEA Guard 406301

Heat Exchangers Regenerator Reboiler 977485 Regenerator Condenser Plate & Frame HX 80076 Lean/Rich Amine Plate & Frame HX 356583 Solvent Cooler Plate & Frame HX 66607

Pumps Pump REGEN to ABS (220 to 2164 kpa) 133408 Pump for MEA Guard (2100 to 2150 kpa) 48964

Storage Vessels Amine Holding Tank 146487

Process Vessels Amine Flash Drum 406530 Condenser Drum 140121

Total Bare Module Cost $ 5,299,342 Contingency and Fee $ 937,469

Total Module Cost $ 6,236,811 Auxiliary Facilities $ 249,472

Installation Costs $ 3,338,585 Total Capital $ 9,824,869

104

Appendix D: Cash Flow Analysis

105

Table D. 1: Cash Flow Analysis

year Depreciation (M CAD$)

Savings (M CAD$)

CO2 Tax (M $/kg)

Yearly Operating Costs (M CAD$)

Net Cash Flow (M CAD$)

Tax Savings

(M CAD$)

Net Expenses

(M CAD$)

After Tax Cash

Flow (M CAD$)

Discounted After Tax Cash Flow (M CAD$)

Cumulative ATCF (M

CAD$)

Cumulative ATCF (MM

CAD$)

0 -9824.87 -9824.9 -9824.9 -9825 -9.825 1 -1786.34 -17123 -18909.3 7752.8 -11156.5 -12942.9 -11984 -21809 -21.809 2 -1607.71 859.34 9523.35 -16868 -8093.02 3318.1 -4774.9 -6382.6 -5472 -27281 -27.281 3 -1429.07 859.34 9523.35 -16868 -7914.39 3244.9 -4669.5 -6098.6 -4841 -32122 -32.122 4 -1250.44 859.34 9523.35 -16868 -7735.76 3171.7 -4564.1 -5814.5 -4274 -36396 -36.396 5 -1071.80 859.34 9523.35 -16868 -7557.12 3098.4 -4458.7 -5530.5 -3764 -40160 -40.160 6 -893.17 859.34 9523.35 -16868 -7378.49 3025.2 -4353.3 -5246.5 -3306 -43466 -43.466 7 -714.54 859.34 9523.35 -16868 -7199.85 2951.9 -4247.9 -4962.4 -2896 -46362 -46.362 8 -535.90 859.34 9523.35 -16868 -7021.22 2878.7 -4142.5 -4678.4 -2528 -48889 -48.889 9 -357.27 859.34 9523.35 -16868 -6842.59 2805.5 -4037.1 -4394.4 -2198 -51088 -51.088

10 -178.63 859.34 9523.35 -16868 -6663.95 2732.2 -3931.7 -4110.4 -1904 -52992 -52.992 11 859.34 9523.35 -16868 -6485.32 2659.0 -3826.3 -3826.3 -1641 -54633 -54.633 12 859.34 9523.35 -16868 -6485.32 2659.0 -3826.3 -3826.3 -1519 -56152 -56.152 13 859.34 9523.35 -16868 -6485.32 2659.0 -3826.3 -3826.3 -1407 -57559 -57.559 14 859.34 9523.35 -16868 -6485.32 2659.0 -3826.3 -3826.3 -1303 -58862 -58.862 15 859.34 9523.35 -16868 -6485.32 2659.0 -3826.3 -3826.3 -1206 -60068 -60.068 16 859.34 9523.35 -16868 -6485.32 2659.0 -3826.3 -3826.3 -1117 -61185 -61.185 17 859.34 9523.35 -16868 -6485.32 2659.0 -3826.3 -3826.3 -1034 -62219 -62.219 18 859.34 9523.35 -16868 -6485.32 2659.0 -3826.3 -3826.3 -958 -63177 -63.177 19 859.34 9523.35 -16868 -6485.32 2659.0 -3826.3 -3826.3 -887 -64063 -64.063 20 859.34 9523.35 -16868 -6485.32 2659.0 -3826.3 -3826.3 -821 -64884 -64.884 21 859.34 9523.35 -16868 -6485.32 2659.0 -3826.3 -3826.3 -760 -65644 -65.644 22 859.34 9523.35 -16868 -6485.32 2659.0 -3826.3 -3826.3 -704 -66348 -66.348 23 859.34 9523.35 -16868 -6485.32 2659.0 -3826.3 -3826.3 -652 -67000 -67.000 24 859.34 9523.35 -16868 -6485.32 2659.0 -3826.3 -3826.3 -603 -67603 -67.603 25 859.34 9523.35 -16868 -6485.32 2659.0 -3826.3 -3826.3 -559 -68162 -68.162

106

Table D. 2: Cash flows at Different Carbon Tax Rates

Carbon Tax

($/tonne) 0 10 20 30 40 50.43 67.33

Year Cumulative ATCF (MM CAD$)

0 -9.825 -9.825 -9.825 -9.825 -9.825 -9.825 -9.825 1 -21.81 -21.81 -21.81 -21.81 -21.81 -21.81 -21.81 2 -32.10 -30.49 -28.89 -27.28 -25.68 -24.00 -21.29 3 -41.40 -38.31 -35.21 -32.12 -29.03 -25.80 -20.58 4 -49.80 -45.33 -40.87 -36.40 -31.93 -27.27 -19.71 5 -57.39 -51.65 -45.90 -40.16 -34.42 -28.43 -18.72 6 -64.24 -57.31 -50.39 -43.47 -36.54 -29.32 -17.62 7 -70.41 -62.40 -54.38 -46.36 -38.34 -29.98 -16.43 8 -75.98 -66.95 -57.92 -48.89 -39.86 -30.44 -15.18 9 -80.98 -71.02 -61.05 -51.09 -41.12 -30.73 -13.88

10 -85.49 -74.66 -63.82 -52.99 -42.16 -30.86 -12.55 11 -89.54 -77.91 -66.27 -54.63 -43.00 -30.86 -11.19 12 -93.29 -80.91 -68.53 -56.15 -43.77 -30.86 -9.935 13 -96.77 -83.70 -70.63 -57.56 -44.49 -30.86 -8.771 14 -99.98 -86.28 -72.57 -58.86 -45.16 -30.86 -7.694 15 -103.0 -88.66 -74.37 -60.07 -45.77 -30.86 -6.696 16 -105.7 -90.87 -76.03 -61.18 -46.34 -30.86 -5.772 17 -108.3 -92.92 -77.57 -62.22 -46.87 -30.86 -4.916 18 -110.6 -94.81 -79.00 -63.18 -47.36 -30.86 -4.124 19 -112.8 -96.57 -80.32 -64.06 -47.81 -30.86 -3.391 20 -114.8 -98.19 -81.54 -64.88 -48.23 -30.86 -2.712 21 -116.7 -99.70 -82.67 -65.64 -48.62 -30.86 -2.083 22 -118.5 -101.1 -83.72 -66.35 -48.98 -30.86 -1.501 23 -120.1 -102.4 -84.69 -67.00 -49.31 -30.86 -0.961 24 -121.6 -103.6 -85.59 -67.60 -49.62 -30.86 -0.462 25 -122.9 -104.7 -86.42 -68.16 -49.90 -30.86 0.000

107

Appendix E: HYSYS Reports

108

Appendix E1.1: PSA Tail Gas With CO2

------------------------------------------------------------------------------- PSA Tail Gas-1 (Material Stream): Conditions, Properties, Composition, Attachments ------------------------------------------------------------------------------- Material Stream: PSA Tail Gas-1 Fluid Package: Basis-1 Property Package: Peng-Robinson CONDITIONS Overall Vapour Phase Vapour / Phase Fraction 1.0000 1.0000 Temperature: (C) 17.29 17.29 Pressure: (kPa) 40.00* 40.00 Molar Flow (kgmole/h) 2047 2047 Mass Flow (kg/h) 5.163e+004 5.163e+004 Std Ideal Liq Vol Flow (m3/h) 91.28 91.28 Molar Enthalpy (kJ/kgmole) -1.962e+005 -1.962e+005 Molar Entropy (kJ/kgmole-C) 175.7 175.7 Heat Flow (kJ/h) -4.016e+008 -4.016e+008 Liq Vol Flow @Std Cond (m3/h) --- --- PROPERTIES Overall Vapour Phase Molecular Weight 25.22 25.22 Molar Density (kgmole/m3) 1.658e-002 1.658e-002 Mass Density (kg/m3) 0.4181 0.4181 Act. Volume Flow (m3/h) 1.235e+005 1.235e+005 Mass Enthalpy (kJ/kg) -7778 -7778 Mass Entropy (kJ/kg-C) 6.967 6.967 Heat Capacity (kJ/kgmole-C) 34.06 34.06 Mass Heat Capacity (kJ/kg-C) 1.350 1.350 Lower Heating Value (kJ/kgmole) 2.328e+005 2.328e+005 Mass Lower Heating Value (kJ/kg) 9232 9232 Phase Fraction [Vol. Basis] --- 1.000 Phase Fraction [Mass Basis] 4.941e-324 1.000 Partial Pressure of CO2 (kPa) 17.81 --- Cost Based on Flow (Cost/s) 0.0000 0.0000

109

Act. Gas Flow (ACT_m3/h) 1.235e+005 1.235e+005 Avg. Liq. Density (kgmole/m3) 22.43 22.43 Specific Heat (kJ/kgmole-C) 34.06 34.06 Std. Gas Flow (STD_m3/h) 4.840e+004 4.840e+004 Std. Ideal Liq. Mass Density (kg/m3) 565.6 565.6 Act. Liq. Flow (m3/s) --- --- Z Factor 0.9992 0.9992 Watson K 10.54 10.54 User Property --- --- Partial Pressure of H2S (kPa) 0.0000 --- Cp/(Cp - R) 1.323 1.323 Cp/Cv 1.325 1.325 Heat of Vap. (kJ/kgmole) --- --- Kinematic Viscosity (cSt) 34.21 34.21 Liq. Mass Density (Std. Cond) (kg/m3) --- --- Liq. Vol. Flow (Std. Cond) (m3/h) --- --- Liquid Fraction 0.0000 0.0000 Molar Volume (m3/kgmole) 60.32 60.32 Mass Heat of Vap. (kJ/kg) --- --- Phase Fraction [Molar Basis] 1.0000 1.0000 Surface Tension (dyne/cm) --- --- Thermal Conductivity (W/m-K) 4.211e-002 4.211e-002 Viscosity (cP) 1.430e-002 1.430e-002 Cv (Semi-Ideal) (kJ/kgmole-C) 25.74 25.74 Mass Cv (Semi-Ideal) (kJ/kg-C) 1.021 1.021 Cv (kJ/kgmole-C) 25.71 25.71 Mass Cv (kJ/kg-C) 1.019 1.019 Cv (Ent. Method) (kJ/kgmole-C) --- --- Mass Cv (Ent. Method) (kJ/kg-C) --- --- Cp/Cv (Ent. Method) --- --- Reid VP at 37.8 C (kPa) --- --- True VP at 37.8 C (kPa) --- --- Liq. Vol. Flow - Sum(Std. Cond) (m3/h) 0.0000 0.0000 COMPOSITION Overall Phase Vapour Fraction 1.0000 COMPONENTS MOLAR FLOW MOLE FRACTION MASS FLOW MASS FRACTION LIQUID VOLUME LIQUID VOLUME (kgmole/h) (kg/h) FLOW (m3/h) FRACTION Methane 360.5000 0.1761 5783.4649 0.1120 19.3172 0.2116

110

Ethane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Propane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 CO 136.7177 0.0668 3829.5870 0.0742 4.7906 0.0525 CO2 911.3458 0.4452 40108.0559 0.7769 48.5961 0.5324 Hydrogen 613.9684 0.2999 1237.7603 0.0240 17.7180 0.1941 H2O 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Nitrogen 23.7409 0.0116 665.0549 0.0129 0.8247 0.0090 Oxygen 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 i-Butane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 n-Butane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 i-Pentane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 n-Pentane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 n-Hexane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Ethylene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 tr2-Butene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 1-Butene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Helium 0.8802 0.0004 3.5233 0.0001 0.0284 0.0003 Propene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Total 2047.1530 1.0000 51627.4464 1.0000 91.2751 1.0000 Vapour Phase Phase Fraction 1.000 COMPONENTS MOLAR FLOW MOLE FRACTION MASS FLOW MASS FRACTION LIQUID VOLUME LIQUID VOLUME (kgmole/h) (kg/h) FLOW (m3/h) FRACTION Methane 360.5000 0.1761 5783.4649 0.1120 19.3172 0.2116 Ethane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Propane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 CO 136.7177 0.0668 3829.5870 0.0742 4.7906 0.0525 CO2 911.3458 0.4452 40108.0559 0.7769 48.5961 0.5324 Hydrogen 613.9684 0.2999 1237.7603 0.0240 17.7180 0.1941 H2O 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Nitrogen 23.7409 0.0116 665.0549 0.0129 0.8247 0.0090 Oxygen 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 i-Butane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 n-Butane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 i-Pentane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 n-Pentane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 n-Hexane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Ethylene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 tr2-Butene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 1-Butene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000

111

Helium 0.8802 0.0004 3.5233 0.0001 0.0284 0.0003 Propene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Total 2047.1530 1.0000 51627.4464 1.0000 91.2751 1.0000 UNIT OPERATIONS FEED TO PRODUCT FROM LOGICAL CONNECTION Tee: TEE-100 Balance: 30-PK-004 UTILITIES ( No utilities reference this stream ) ------------------------------------------------------------------------------- Hyprotech Ltd. Aspen HYSYS Version 2006 (20.0.0.6728)

112

Appendix E1.2: Reformer Furnace With CO2

------------------------------------------------------------------------------- 30-F-001-A (Conversion Reactor): Design, Reactions, Worksheet ------------------------------------------------------------------------------- Conversion Reactor: 30-F-001-A CONNECTIONS Inlet Stream Connections Stream Name From Unit Operation To Reformer Mixed Feed Preheat Coil Heat Exchanger Outlet Stream Connections Stream Name To Unit Operation To Waste HEX Component Splitter: X-100 DNE2 Energy Stream Connections Stream Name From Unit Operation Q-102 PARAMETERS Physical Parameters Optional Heat Transfer: Heating Delta P Vessel Volume Duty Energy Stream 0.0000 kPa --- 3.160e+008 kJ/h Q-102 User Variables REACTION DETAILS Reaction: Meth Reform Component Mole Weight Stoichiometric Coeff. Methane 16.04 -1.000 H2O 18.02 -1.000 CO 28.01 1.000 Hydrogen 2.016 3.000

113

Reaction: Eth Reform Component Mole Weight Stoichiometric Coeff. Ethane 30.07 -1.000 H2O 18.02 -2.000 CO 28.01 2.000 Hydrogen 2.016 5.000 Reaction: Prop Reform Component Mole Weight Stoichiometric Coeff. Propane 44.10 -1.000 H2O 18.02 -3.000 CO 28.01 3.000 Hydrogen 2.016 7.000 Reaction: Reformer Shift Component Mole Weight Stoichiometric Coeff. CO 28.01 -1.000 H2O 18.02 -1.000 CO2 44.01 1.000 Hydrogen 2.016 1.000 Reaction: 1-butene ref Component Mole Weight Stoichiometric Coeff. 1-Butene 56.11 -1.000 H2O 18.02 -4.000 CO 28.01 4.000 Hydrogen 2.016 8.000 Reaction: ethyl ref Component Mole Weight Stoichiometric Coeff. Ethylene 28.05 -1.000 H2O 18.02 -2.000 CO 28.01 2.000 Hydrogen 2.016 4.000 Reaction: i-but ref Component Mole Weight Stoichiometric Coeff. i-Butane 58.12 -1.000 H2O 18.02 -4.000

114

CO 28.01 4.000 Hydrogen 2.016 9.000 Reaction: i-pent ref Component Mole Weight Stoichiometric Coeff. i-Pentane 72.15 -1.000 H2O 18.02 -5.000 CO 28.01 5.000 Hydrogen 2.016 11.000 Reaction: n-but ref Component Mole Weight Stoichiometric Coeff. n-Butane 58.12 -1.000 H2O 18.02 -4.000 CO 28.01 4.000 Hydrogen 2.016 9.000 Reaction: n-hex ref Component Mole Weight Stoichiometric Coeff. n-Hexane 86.18 -1.000 H2O 18.02 -6.000 CO 28.01 6.000 Hydrogen 2.016 13.000 Reaction: n-pent ref Component Mole Weight Stoichiometric Coeff. n-Pentane 72.15 -1.000 H2O 18.02 -5.000 CO 28.01 5.000 Hydrogen 2.016 11.000 Reaction: propene ref Component Mole Weight Stoichiometric Coeff. Propene 42.08 -1.000 H2O 18.02 -3.000 CO 28.01 3.000 Hydrogen 2.016 6.000 Reaction: tr-but ref Component Mole Weight Stoichiometric Coeff.

115

tr2-Butene 56.11 -1.000 H2O 18.02 -4.000 CO 28.01 4.000 Hydrogen 2.016 8.000 REACTION RESULTS FOR : Reformer Extents Name Rank Specified Use Default Actual Base Reaction Extent % Conversion % Conversion Component (kgmole/h) Meth Reform 0 72.80 No 72.80 Methane 964.9 Eth Reform 0 100.00 Yes 100.0 Ethane 29.16 Prop Reform 0 100.00 Yes 100.0 Propane 8.390 Reformer Shift 1 58.00 No --- CO 616.4 1-butene ref 0 100.00 Yes --- 1-Butene 0.0000 ethyl ref 0 100.00 Yes --- Ethylene 0.0000 i-but ref 0 100.00 Yes 100.0 i-Butane 0.1789 i-pent ref 0 100.00 Yes 100.0 i-Pentane 0.4969 n-but ref 0 100.00 Yes 100.0 n-Butane 2.016 n-hex ref 0 100.00 Yes 100.0 n-Hexane 0.2413 n-pent ref 0 100.00 Yes 100.0 n-Pentane 0.3407 propene ref 0 100.00 Yes --- Propene 0.0000 tr-but ref 0 100.00 Yes --- tr2-Butene 0.0000 Balance Components Total Inflow Total Reaction Total Outflow (kgmole/h) (kgmole/h) (kgmole/h) Methane 1325 -964.9 360.5 Ethane 29.16 -29.16 0.0000 Propane 8.390 -8.390 0.0000 CO 0.0000 446.4 446.4 CO2 2.569 616.4 619.0 Hydrogen 12.27 3748 3760 H2O 4706 -1679 3027 Nitrogen 23.74 0.0000 23.74 Oxygen 0.0000 0.0000 0.0000 i-Butane 0.1789 -0.1789 0.0000 n-Butane 2.016 -2.016 0.0000 i-Pentane 0.4969 -0.4969 0.0000

116

n-Pentane 0.3407 -0.3407 0.0000 n-Hexane 0.2413 -0.2413 0.0000 Ethylene 0.0000 0.0000 0.0000 tr2-Butene 0.0000 0.0000 0.0000 1-Butene 0.0000 0.0000 0.0000 Helium 0.8802 0.0000 0.8802 Propene 0.0000 0.0000 0.0000 CONDITIONS Name To Reformer DNE2 To Waste HEX Q-102 Vapour 1.0000 0.0000 1.0000 --- Temperature (C) 449.9883 794.9993 794.9993 --- Pressure (kPa) 3501.0000 3501.0000 3501.0000 --- Molar Flow (kgmole/h) 6111.7508 0.0000 8237.2949 --- Mass Flow (kg/h) 108305.6690 0.0000 108306.4336 --- Std Ideal Liq Vol Flow (m3/h) 160.8581 0.0000 231.9613 --- Molar Enthalpy (kJ/kgmole) -1.875e+005 -1.007e+005 -1.007e+005 --- Molar Entropy (kJ/kgmole-C) 184.1 174.5 174.5 --- Heat Flow (kJ/h) -1.1458e+09 0.0000e-01 -8.2978e+08 3.1598e+08 PROPERTIES Name To Reformer DNE2 To Waste HEX Molecular Weight 17.72 13.15 13.15 Molar Density (kgmole/m3) 0.5960 0.3926 0.3926 Mass Density (kg/m3) 10.56 5.162 5.162 Act. Volume Flow (m3/h) 1.025e+004 0.0000 2.098e+004 Mass Enthalpy (kJ/kg) -1.058e+004 -7661 -7661 Mass Entropy (kJ/kg-C) 10.39 13.27 13.27 Heat Capacity (kJ/kgmole-C) 44.71 38.97 38.97 Mass Heat Capacity (kJ/kg-C) 2.523 2.964 2.964 Lower Heating Value (kJ/kgmole) 1.857e+005 1.609e+005 1.609e+005 Mass Lower Heating Value (kJ/kg) 1.048e+004 1.224e+004 1.224e+004 Phase Fraction [Vol. Basis] --- --- --- Phase Fraction [Mass Basis] 4.941e-324 4.941e-324 4.941e-324 Partial Pressure of CO2 (kPa) 1.472 0.0000 263.1 Cost Based on Flow (Cost/s) 0.0000 0.0000 0.0000 Act. Gas Flow (ACT_m3/h) 1.025e+004 --- 2.098e+004 Avg. Liq. Density (kgmole/m3) 37.99 --- 35.51 Specific Heat (kJ/kgmole-C) 44.71 38.97 38.97 Std. Gas Flow (STD_m3/h) 1.445e+005 0.0000 1.948e+005

117

Std. Ideal Liq. Mass Density (kg/m3) 673.3 466.9 466.9 Act. Liq. Flow (m3/s) --- 0.0000 0.0000 Z Factor 0.9770 --- --- Watson K 18.97 14.89 14.89 User Property --- --- --- Partial Pressure of H2S (kPa) 0.0000 0.0000 0.0000 Cp/(Cp - R) 1.228 1.271 1.271 Cp/Cv 1.272 1.277 1.277 Heat of Vap. (kJ/kgmole) 5.007e+004 3.920e+004 3.920e+004 Kinematic Viscosity (cSt) 2.059 0.7359 5.299 Liq. Mass Density (Std. Cond) (kg/m3) 741.0 --- --- Liq. Vol. Flow (Std. Cond) (m3/h) 146.2 0.0000 --- Liquid Fraction 0.0000 1.000 0.0000 Molar Volume (m3/kgmole) 1.678 2.547 2.547 Mass Heat of Vap. (kJ/kg) 2826 2982 2982 Phase Fraction [Molar Basis] 1.0000 0.0000 1.0000 Surface Tension (dyne/cm) --- 0.0000 --- Thermal Conductivity (W/m-K) 7.302e-002 0.1432 0.1885 Viscosity (cP) 2.175e-002 3.799e-003 2.736e-002 Cv (Semi-Ideal) (kJ/kgmole-C) 36.39 30.65 30.65 Mass Cv (Semi-Ideal) (kJ/kg-C) 2.054 2.331 2.331 Cv (kJ/kgmole-C) 35.15 30.51 30.51 Mass Cv (kJ/kg-C) 1.983 2.320 2.320 Cv (Ent. Method) (kJ/kgmole-C) 36.24 --- 30.49 Mass Cv (Ent. Method) (kJ/kg-C) 2.045 --- 2.319 Cp/Cv (Ent. Method) 1.234 --- 1.278 Reid VP at 37.8 C (kPa) --- --- --- True VP at 37.8 C (kPa) --- --- --- Liq. Vol. Flow - Sum(Std. Cond) (m3/h) 146.2 0.0000 0.0000 ------------------------------------------------------------------------------- 30-F-001-B (Conversion Reactor): Design, Reactions, Worksheet ------------------------------------------------------------------------------- Conversion Reactor: 30-F-001-B CONNECTIONS Inlet Stream Connections

118

Stream Name From Unit Operation Combo Gas PSA Tail Gas-2 RCY-1 Recycle Combustion Air Outlet Stream Connections Stream Name To Unit Operation waste gas Heat Exchanger: Steam Generation Coil I DNE Energy Stream Connections Stream Name From Unit Operation Q-100 PARAMETERS Physical Parameters Optional Heat Transfer: Cooling Delta P Vessel Volume Duty Energy Stream 0.0000 kPa --- 3.657e+008 kJ/h Q-100 User Variables REACTION DETAILS Reaction: Meth Combust Component Mole Weight Stoichiometric Coeff. Methane 16.04 -1.000 Oxygen 32.00 -2.000 CO2 44.01 1.000 H2O 18.02 2.000 Reaction: Eth combust Component Mole Weight Stoichiometric Coeff. Ethane 30.07 -1.000 Oxygen 32.00 -3.500 CO2 44.01 2.000 H2O 18.02 3.000 Reaction: Prop combust Component Mole Weight Stoichiometric Coeff. Propane 44.10 -1.000

119

Oxygen 32.00 -5.000 CO2 44.01 3.000 H2O 18.02 4.000 Reaction: H2 combust Component Mole Weight Stoichiometric Coeff. Hydrogen 2.016 -2.000 Oxygen 32.00 -1.000 H2O 18.02 2.000 Reaction: 1-but combust Component Mole Weight Stoichiometric Coeff. Oxygen 32.00 -6.000 CO2 44.01 4.000 H2O 18.02 4.000 1-Butene 56.11 -1.000 Reaction: ethylene comust Component Mole Weight Stoichiometric Coeff. Oxygen 32.00 -3.000 CO2 44.01 2.000 H2O 18.02 2.000 Ethylene 28.05 -1.000 Reaction: i-but combust Component Mole Weight Stoichiometric Coeff. i-Butane 58.12 -1.000 Oxygen 32.00 -6.500 CO2 44.01 4.000 H2O 18.02 5.000 Reaction: i-pent combust Component Mole Weight Stoichiometric Coeff. Oxygen 32.00 -8.000 CO2 44.01 5.000 H2O 18.02 6.000 i-Pentane 72.15 -1.000 Reaction: n-but combust Component Mole Weight Stoichiometric Coeff.

120

Oxygen 32.00 -6.500 CO2 44.01 4.000 H2O 18.02 5.000 n-Butane 58.12 -1.000 Reaction: n-hex combust Component Mole Weight Stoichiometric Coeff. Oxygen 32.00 -9.500 CO2 44.01 6.000 H2O 18.02 7.000 n-Hexane 86.18 -1.000 Reaction: n-pent comust Component Mole Weight Stoichiometric Coeff. Oxygen 32.00 -8.000 CO2 44.01 5.000 H2O 18.02 6.000 n-Pentane 72.15 -1.000 Reaction: propene combust Component Mole Weight Stoichiometric Coeff. Propene 42.08 -1.000 Oxygen 32.00 -4.500 CO2 44.01 3.000 H2O 18.02 3.000 Reaction: tr2-but combust Component Mole Weight Stoichiometric Coeff. Oxygen 32.00 -6.000 CO2 44.01 4.000 H2O 18.02 4.000 tr2-Butene 56.11 -1.000 Reaction: Component Mole Weight Stoichiometric Coeff.

121

REACTION RESULTS FOR : Combust Extents Name Rank Specified Use Default Actual Base Reaction Extent % Conversion % Conversion Component (kgmole/h) Meth Combust 0 100.00 Yes 100.0 Methane 391.6 Eth combust 0 100.00 Yes 100.0 Ethane 3.619 Prop combust 0 100.00 Yes 100.0 Propane 1.653 H2 combust 0 100.00 Yes 100.0 Hydrogen 370.6 1-but combust 0 100.00 Yes 100.0 1-Butene 1.400e-002 ethylene comust 0 100.00 Yes 100.0 Ethylene 0.2444 i-but combust 0 100.00 Yes 100.0 i-Butane 5.982e-002 i-pent combust 0 100.00 Yes 100.0 i-Pentane 4.084e-003 n-but combust 0 100.00 Yes 100.0 n-Butane 8.196e-002 n-hex combust 0 100.00 Yes 100.0 n-Hexane 3.305e-003 n-pent comust 0 100.00 Yes 100.0 n-Pentane 9.938e-003 propene combust 0 100.00 Yes 100.0 Propene 0.3504 tr2-but combust 0 100.00 Yes 100.0 tr2-Butene 2.801e-003 CO Combust 0 100.00 Yes 100.0 CO 136.8 Balance Components Total Inflow Total Reaction Total Outflow (kgmole/h) (kgmole/h) (kgmole/h) Methane 391.6 -391.6 0.0000 Ethane 3.619 -3.619 0.0000 Propane 1.653 -1.653 0.0000 CO 136.8 -136.8 0.0000 CO2 911.4 542.8 1454 Hydrogen 741.2 -741.2 0.0000 H2O 0.0000 1544 1544 Nitrogen 5336 0.0000 5336 Oxygen 1412 -1247 165.5 i-Butane 5.982e-002 -5.982e-002 0.0000 n-Butane 8.196e-002 -8.196e-002 0.0000 i-Pentane 4.084e-003 -4.084e-003 0.0000 n-Pentane 9.938e-003 -9.938e-003 0.0000 n-Hexane 3.305e-003 -3.305e-003 0.0000 Ethylene 0.2444 -0.2444 0.0000 tr2-Butene 2.801e-003 -2.801e-003 0.0000

122

1-Butene 1.400e-002 -1.400e-002 0.0000 Helium 1.020 0.0000 1.020 Propene 0.3504 -0.3504 0.0000 CONDITIONS Name Combo Gas PSA Tail Gas-2 Combustion Air DNE Vapour 1.0000 1.0000 1.0000 0.0000 Temperature (C) 55.0000 17.2199 300.0000 800.0000 Pressure (kPa) 146.3250 40.0000 102.8000 40.0000 Molar Flow (kgmole/h) 164.9437 2047.1530 6724.0000 0.0000 Mass Flow (kg/h) 982.2490 51627.4464 193989.2181 0.0000 Std Ideal Liq Vol Flow (m3/h) 5.8716 91.2751 224.2517 0.0000 Molar Enthalpy (kJ/kgmole) -1.617e+004 -1.962e+005 8248 -8.404e+004 Molar Entropy (kJ/kgmole-C) 142.2 175.7 171.1 216.8 Heat Flow (kJ/h) -2.6663e+06 -4.0156e+08 5.5463e+07 0.0000e-01 Name waste gas Vapour 1.0000 Temperature (C) 800.0000 Pressure (kPa) 40.0000 Molar Flow (kgmole/h) 8501.0097 Mass Flow (kg/h) 246596.8774 Std Ideal Liq Vol Flow (m3/h) 295.4778 Molar Enthalpy (kJ/kgmole) -8.404e+004 Molar Entropy (kJ/kgmole-C) 216.8 Heat Flow (kJ/h) -7.1444e+08 PROPERTIES Name Combo Gas PSA Tail Gas-2 Combustion Air DNE waste gas Molecular Weight 5.955 25.22 28.85 29.01 29.01 Molar Density (kgmole/m3) 5.362e-002 1.658e-002 2.157e-002 4.483e-003 4.483e-003 Mass Density (kg/m3) 0.3193 0.4182 0.6222 0.1300 0.1300 Act. Volume Flow (m3/h) 3076 1.235e+005 3.118e+005 0.0000 1.896e+006 Mass Enthalpy (kJ/kg) -2715 -7778 285.9 -2897 -2897 Mass Entropy (kJ/kg-C) 23.88 6.966 5.932 7.475 7.475 Heat Capacity (kJ/kgmole-C) 31.48 34.05 30.84 38.36 38.36 Mass Heat Capacity (kJ/kg-C) 5.286 1.350 1.069 1.322 1.322 Lower Heating Value (kJ/kgmole) 3.988e+005 2.328e+005 0.0000 0.0000 0.0000 Mass Lower Heating Value (kJ/kg) 6.697e+004 9232 --- --- --- Phase Fraction [Vol. Basis] --- --- --- --- --- Phase Fraction [Mass Basis] 4.941e-324 4.941e-324 4.941e-324 2.122e-314 2.122e-314

123

Partial Pressure of CO2 (kPa) 1.267e-002 17.81 0.0000 0.0000 6.842 Cost Based on Flow (Cost/s) 0.0000 0.0000 0.0000 0.0000 0.0000 Act. Gas Flow (ACT_m3/h) 3076 1.235e+005 3.118e+005 --- 1.896e+006 Avg. Liq. Density (kgmole/m3) 28.09 22.43 29.98 --- 28.77 Specific Heat (kJ/kgmole-C) 31.48 34.05 30.84 38.36 38.36 Std. Gas Flow (STD_m3/h) 3900 4.840e+004 1.590e+005 0.0000 2.010e+005 Std. Ideal Liq. Mass Density (kg/m3) 167.3 565.6 865.1 834.6 834.6 Act. Liq. Flow (m3/s) --- --- --- 0.0000 --- Z Factor 1.000 0.9992 1.000 --- --- Watson K 25.94 10.54 6.042 6.967 6.967 User Property --- --- --- --- --- Partial Pressure of H2S (kPa) 0.0000 0.0000 0.0000 0.0000 0.0000 Cp/(Cp - R) 1.359 1.323 1.369 1.277 1.277 Cp/Cv 1.360 1.325 1.370 1.277 1.277 Heat of Vap. (kJ/kgmole) 7195 --- 5891 --- --- Kinematic Viscosity (cSt) 29.15 31.30 48.58 2.347 328.8 Liq. Mass Density (Std. Cond) (kg/m3) --- --- --- --- --- Liq. Vol. Flow (Std. Cond) (m3/h) --- --- --- 0.0000 --- Liquid Fraction 0.0000 0.0000 0.0000 1.000 0.0000 Molar Volume (m3/kgmole) 18.65 60.31 46.37 223.1 223.1 Mass Heat of Vap. (kJ/kg) 1208 --- 204.2 --- --- Phase Fraction [Molar Basis] 1.0000 1.0000 1.0000 0.0000 1.0000 Surface Tension (dyne/cm) --- --- --- 0.0000 --- Thermal Conductivity (W/m-K) 0.1298 4.210e-002 4.360e-002 7.160e-002 7.326e-002 Viscosity (cP) 9.309e-003 1.309e-002 3.022e-002 3.052e-004 4.276e-002 Cv (Semi-Ideal) (kJ/kgmole-C) 23.16 25.74 22.53 30.05 30.05 Mass Cv (Semi-Ideal) (kJ/kg-C) 3.890 1.021 0.7808 1.036 1.036 Cv (kJ/kgmole-C) 23.14 25.71 22.52 30.04 30.04 Mass Cv (kJ/kg-C) 3.886 1.019 0.7805 1.036 1.036 Cv (Ent. Method) (kJ/kgmole-C) 23.12 --- 22.48 --- --- Mass Cv (Ent. Method) (kJ/kg-C) 3.882 --- 0.7791 --- --- Cp/Cv (Ent. Method) 1.362 --- 1.372 --- --- Reid VP at 37.8 C (kPa) --- --- --- --- --- True VP at 37.8 C (kPa) --- --- --- 1.630e+005 1.630e+005 Liq. Vol. Flow - Sum(Std. Cond) (m3/h) 0.0000 0.0000 0.0000 0.0000 0.0000 ------------------------------------------------------------------------------- Hyprotech Ltd. Aspen HYSYS Version 2006 (20.0.0.6728)

124

Appendix E2.1: PSA Tail Gas Without CO2 ------------------------------------------------------------------------------- PSA Tail Gas-1 (Material Stream): Conditions, Properties, Composition, Attachments ------------------------------------------------------------------------------- Material Stream: PSA Tail Gas-1 Fluid Package: Basis-1 Property Package: Peng-Robinson CONDITIONS Overall Vapour Phase Vapour / Phase Fraction 1.0000 1.0000 Temperature: (C) 31.98 31.98 Pressure: (kPa) 40.00* 40.00 Molar Flow (kgmole/h) 1231 1231 Mass Flow (kg/h) 1.570e+004 1.570e+004 Std Ideal Liq Vol Flow (m3/h) 47.74 47.74 Molar Enthalpy (kJ/kgmole) -6.438e+004 -6.438e+004 Molar Entropy (kJ/kgmole-C) 167.7 167.7 Heat Flow (kJ/h) -7.924e+007 -7.924e+007 Liq Vol Flow @Std Cond (m3/h) --- --- PROPERTIES Overall Vapour Phase Molecular Weight 12.75 12.75 Molar Density (kgmole/m3) 1.577e-002 1.577e-002 Mass Density (kg/m3) 0.2011 0.2011 Act. Volume Flow (m3/h) 7.805e+004 7.805e+004 Mass Enthalpy (kJ/kg) -5048 -5048 Mass Entropy (kJ/kg-C) 13.15 13.15 Heat Capacity (kJ/kgmole-C) 31.59 31.59 Mass Heat Capacity (kJ/kg-C) 2.477 2.477 Lower Heating Value (kJ/kgmole) 3.873e+005 3.873e+005 Mass Lower Heating Value (kJ/kg) 3.036e+004 3.036e+004 Phase Fraction [Vol. Basis] --- 1.000 Phase Fraction [Mass Basis] 4.941e-324 1.000 Partial Pressure of CO2 (kPa) 3.086 --- Cost Based on Flow (Cost/s) 0.0000 0.0000 Act. Gas Flow (ACT_m3/h) 7.805e+004 7.805e+004 Avg. Liq. Density (kgmole/m3) 25.78 25.78 Specific Heat (kJ/kgmole-C) 31.59 31.59 Std. Gas Flow (STD_m3/h) 2.910e+004 2.910e+004

125

Std. Ideal Liq. Mass Density (kg/m3) 328.8 328.8 Act. Liq. Flow (m3/s) --- --- Z Factor 0.9999 0.9999 Watson K 15.10 15.10 User Property --- --- Partial Pressure of H2S (kPa) 0.0000 --- Cp/(Cp - R) 1.357 1.357 Cp/Cv 1.358 1.358 Heat of Vap. (kJ/kgmole) --- --- Kinematic Viscosity (cSt) 60.91 60.91 Liq. Mass Density (Std. Cond) (kg/m3) --- --- Liq. Vol. Flow (Std. Cond) (m3/h) --- --- Liquid Fraction 0.0000 0.0000 Molar Volume (m3/kgmole) 63.42 63.42 Mass Heat of Vap. (kJ/kg) --- --- Phase Fraction [Molar Basis] 1.0000 1.0000 Surface Tension (dyne/cm) --- --- Thermal Conductivity (W/m-K) 7.500e-002 7.500e-002 Viscosity (cP) 1.225e-002 1.225e-002 Cv (Semi-Ideal) (kJ/kgmole-C) 23.28 23.28 Mass Cv (Semi-Ideal) (kJ/kg-C) 1.825 1.825 Cv (kJ/kgmole-C) 23.27 23.27 Mass Cv (kJ/kg-C) 1.824 1.824 Cv (Ent. Method) (kJ/kgmole-C) --- --- Mass Cv (Ent. Method) (kJ/kg-C) --- --- Cp/Cv (Ent. Method) --- --- Reid VP at 37.8 C (kPa) --- --- True VP at 37.8 C (kPa) --- --- Liq. Vol. Flow - Sum(Std. Cond) (m3/h) 0.0000 0.0000 COMPOSITION Overall Phase Vapour Fraction 1.0000 COMPONENTS MOLAR FLOW MOLE FRACTION MASS FLOW MASS FRACTION LIQUID VOLUME LIQUID VOLUME (kgmole/h) (kg/h) FLOW (m3/h) FRACTION Methane 360.5000 0.2929 5783.4649 0.3684 19.3172 0.4046 Ethane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Propane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 CO 136.7177 0.1111 3829.5870 0.2440 4.7906 0.1003 CO2 94.9458 0.0771 4178.5362 0.2662 5.0628 0.1060

126

Hydrogen 613.9684 0.4989 1237.7603 0.0788 17.7180 0.3711 H2O 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Nitrogen 23.7409 0.0193 665.0549 0.0424 0.8247 0.0173 Oxygen 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 i-Butane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 n-Butane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 i-Pentane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 n-Pentane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 n-Hexane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Ethylene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 tr2-Butene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 1-Butene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Helium 0.8802 0.0007 3.5233 0.0002 0.0284 0.0006 Propene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Total 1230.7530 1.0000 15697.9267 1.0000 47.7418 1.0000 Vapour Phase Phase Fraction 1.000 COMPONENTS MOLAR FLOW MOLE FRACTION MASS FLOW MASS FRACTION LIQUID VOLUME LIQUID VOLUME (kgmole/h) (kg/h) FLOW (m3/h) FRACTION Methane 360.5000 0.2929 5783.4649 0.3684 19.3172 0.4046 Ethane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Propane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 CO 136.7177 0.1111 3829.5870 0.2440 4.7906 0.1003 CO2 94.9458 0.0771 4178.5362 0.2662 5.0628 0.1060 Hydrogen 613.9684 0.4989 1237.7603 0.0788 17.7180 0.3711 H2O 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Nitrogen 23.7409 0.0193 665.0549 0.0424 0.8247 0.0173 Oxygen 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 i-Butane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 n-Butane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 i-Pentane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 n-Pentane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 n-Hexane 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Ethylene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 tr2-Butene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 1-Butene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Helium 0.8802 0.0007 3.5233 0.0002 0.0284 0.0006 Propene 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Total 1230.7530 1.0000 15697.9267 1.0000 47.7418 1.0000 Hyprotech Ltd. Aspen HYSYS Version 2006 (20.0.0.6728)

127

Appendix E2.2: Reformer without CO2

------------------------------------------------------------------------------- 30-F-001-A (Conversion Reactor): Design, Reactions, Worksheet ------------------------------------------------------------------------------- Conversion Reactor: 30-F-001-A CONNECTIONS Inlet Stream Connections Stream Name From Unit Operation To Reformer Mixed Feed Preheat Coil Heat Exchanger Outlet Stream Connections Stream Name To Unit Operation To Waste HEX Component Splitter: X-100 DNE2 Energy Stream Connections Stream Name From Unit Operation Q-102 PARAMETERS Physical Parameters Optional Heat Transfer: Heating Delta P Vessel Volume Duty Energy Stream 0.0000 kPa --- 3.160e+008 kJ/h Q-102 User Variables REACTION DETAILS Reaction: Meth Reform Component Mole Weight Stoichiometric Coeff. Methane 16.04 -1.000 H2O 18.02 -1.000 CO 28.01 1.000 Hydrogen 2.016 3.000

128

Reaction: Eth Reform Component Mole Weight Stoichiometric Coeff. Ethane 30.07 -1.000 H2O 18.02 -2.000 CO 28.01 2.000 Hydrogen 2.016 5.000 Reaction: Prop Reform Component Mole Weight Stoichiometric Coeff. Propane 44.10 -1.000 H2O 18.02 -3.000 CO 28.01 3.000 Hydrogen 2.016 7.000 Reaction: Reformer Shift Component Mole Weight Stoichiometric Coeff. CO 28.01 -1.000 H2O 18.02 -1.000 CO2 44.01 1.000 Hydrogen 2.016 1.000 Reaction: 1-butene ref Component Mole Weight Stoichiometric Coeff. 1-Butene 56.11 -1.000 H2O 18.02 -4.000 CO 28.01 4.000 Hydrogen 2.016 8.000 Reaction: ethyl ref Component Mole Weight Stoichiometric Coeff. Ethylene 28.05 -1.000 H2O 18.02 -2.000 CO 28.01 2.000 Hydrogen 2.016 4.000 Reaction: i-but ref Component Mole Weight Stoichiometric Coeff. i-Butane 58.12 -1.000 H2O 18.02 -4.000

129

CO 28.01 4.000 Hydrogen 2.016 9.000 Reaction: i-pent ref Component Mole Weight Stoichiometric Coeff. i-Pentane 72.15 -1.000 H2O 18.02 -5.000 CO 28.01 5.000 Hydrogen 2.016 11.000 Reaction: n-but ref Component Mole Weight Stoichiometric Coeff. n-Butane 58.12 -1.000 H2O 18.02 -4.000 CO 28.01 4.000 Hydrogen 2.016 9.000 Reaction: n-hex ref Component Mole Weight Stoichiometric Coeff. n-Hexane 86.18 -1.000 H2O 18.02 -6.000 CO 28.01 6.000 Hydrogen 2.016 13.000 Reaction: n-pent ref Component Mole Weight Stoichiometric Coeff. n-Pentane 72.15 -1.000 H2O 18.02 -5.000 CO 28.01 5.000 Hydrogen 2.016 11.000 Reaction: propene ref Component Mole Weight Stoichiometric Coeff. Propene 42.08 -1.000 H2O 18.02 -3.000 CO 28.01 3.000 Hydrogen 2.016 6.000 Reaction: tr-but ref Component Mole Weight Stoichiometric Coeff.

130

tr2-Butene 56.11 -1.000 H2O 18.02 -4.000 CO 28.01 4.000 Hydrogen 2.016 8.000 REACTION RESULTS FOR : Reformer Extents Name Rank Specified Use Default Actual Base Reaction Extent % Conversion % Conversion Component (kgmole/h) Meth Reform 0 72.80 No 72.80 Methane 964.9 Eth Reform 0 100.00 Yes 100.0 Ethane 29.16 Prop Reform 0 100.00 Yes 100.0 Propane 8.390 Reformer Shift 1 58.00 No --- CO 616.4 1-butene ref 0 100.00 Yes --- 1-Butene 0.0000 ethyl ref 0 100.00 Yes --- Ethylene 0.0000 i-but ref 0 100.00 Yes 100.0 i-Butane 0.1789 i-pent ref 0 100.00 Yes 100.0 i-Pentane 0.4969 n-but ref 0 100.00 Yes 100.0 n-Butane 2.016 n-hex ref 0 100.00 Yes 100.0 n-Hexane 0.2413 n-pent ref 0 100.00 Yes 100.0 n-Pentane 0.3407 propene ref 0 100.00 Yes --- Propene 0.0000 tr-but ref 0 100.00 Yes --- tr2-Butene 0.0000 Balance Components Total Inflow Total Reaction Total Outflow (kgmole/h) (kgmole/h) (kgmole/h) Methane 1325 -964.9 360.5 Ethane 29.16 -29.16 0.0000 Propane 8.390 -8.390 0.0000 CO 0.0000 446.4 446.4 CO2 2.569 616.4 619.0 Hydrogen 12.27 3748 3760 H2O 4706 -1679 3027 Nitrogen 23.74 0.0000 23.74 Oxygen 0.0000 0.0000 0.0000 i-Butane 0.1789 -0.1789 0.0000 n-Butane 2.016 -2.016 0.0000 i-Pentane 0.4969 -0.4969 0.0000

131

n-Pentane 0.3407 -0.3407 0.0000 n-Hexane 0.2413 -0.2413 0.0000 Ethylene 0.0000 0.0000 0.0000 tr2-Butene 0.0000 0.0000 0.0000 1-Butene 0.0000 0.0000 0.0000 Helium 0.8802 0.0000 0.8802 Propene 0.0000 0.0000 0.0000 CONDITIONS Name To Reformer DNE2 To Waste HEX Q-102 Vapour 1.0000 0.0000 1.0000 --- Temperature (C) 449.9139 795.0129 795.0129 --- Pressure (kPa) 3501.0000 3501.0000 3501.0000 --- Molar Flow (kgmole/h) 6111.7508 0.0000 8237.2949 --- Mass Flow (kg/h) 108305.6690 0.0000 108306.4336 --- Std Ideal Liq Vol Flow (m3/h) 160.8581 0.0000 231.9613 --- Molar Enthalpy (kJ/kgmole) -1.875e+005 -1.007e+005 -1.007e+005 --- Molar Entropy (kJ/kgmole-C) 184.1 174.5 174.5 --- Heat Flow (kJ/h) -1.1458e+09 0.0000e-01 -8.2978e+08 3.1600e+08 PROPERTIES Name To Reformer DNE2 To Waste HEX Molecular Weight 17.72 13.15 13.15 Molar Density (kgmole/m3) 0.5961 0.3926 0.3926 Mass Density (kg/m3) 10.56 5.162 5.162 Act. Volume Flow (m3/h) 1.025e+004 0.0000 2.098e+004 Mass Enthalpy (kJ/kg) -1.058e+004 -7661 -7661 Mass Entropy (kJ/kg-C) 10.39 13.27 13.27 Heat Capacity (kJ/kgmole-C) 44.71 38.97 38.97 Mass Heat Capacity (kJ/kg-C) 2.523 2.964 2.964 Lower Heating Value (kJ/kgmole) 1.857e+005 1.609e+005 1.609e+005 Mass Lower Heating Value (kJ/kg) 1.048e+004 1.224e+004 1.224e+004 Phase Fraction [Vol. Basis] --- --- --- Phase Fraction [Mass Basis] 4.941e-324 4.941e-324 4.941e-324 Partial Pressure of CO2 (kPa) 1.472 0.0000 263.1 Cost Based on Flow (Cost/s) 0.0000 0.0000 0.0000 Act. Gas Flow (ACT_m3/h) 1.025e+004 --- 2.098e+004 Avg. Liq. Density (kgmole/m3) 37.99 --- 35.51 Specific Heat (kJ/kgmole-C) 44.71 38.97 38.97 Std. Gas Flow (STD_m3/h) 1.445e+005 0.0000 1.948e+005

132

Std. Ideal Liq. Mass Density (kg/m3) 673.3 466.9 466.9 Act. Liq. Flow (m3/s) --- 0.0000 0.0000 Z Factor 0.9770 --- --- Watson K 18.97 14.89 14.89 User Property --- --- --- Partial Pressure of H2S (kPa) 0.0000 0.0000 0.0000 Cp/(Cp - R) 1.228 1.271 1.271 Cp/Cv 1.272 1.277 1.277 Heat of Vap. (kJ/kgmole) 5.007e+004 3.920e+004 3.920e+004 Kinematic Viscosity (cSt) 2.059 0.7359 5.299 Liq. Mass Density (Std. Cond) (kg/m3) 741.0 --- --- Liq. Vol. Flow (Std. Cond) (m3/h) 146.2 0.0000 --- Liquid Fraction 0.0000 1.000 0.0000 Molar Volume (m3/kgmole) 1.678 2.547 2.547 Mass Heat of Vap. (kJ/kg) 2826 2982 2982 Phase Fraction [Molar Basis] 1.0000 0.0000 1.0000 Surface Tension (dyne/cm) --- 0.0000 --- Thermal Conductivity (W/m-K) 7.301e-002 0.1432 0.1885 Viscosity (cP) 2.175e-002 3.799e-003 2.736e-002 Cv (Semi-Ideal) (kJ/kgmole-C) 36.39 30.65 30.65 Mass Cv (Semi-Ideal) (kJ/kg-C) 2.054 2.331 2.331 Cv (kJ/kgmole-C) 35.15 30.51 30.51 Mass Cv (kJ/kg-C) 1.983 2.320 2.320 Cv (Ent. Method) (kJ/kgmole-C) 36.22 --- 30.49 Mass Cv (Ent. Method) (kJ/kg-C) 2.044 --- 2.319 Cp/Cv (Ent. Method) 1.234 --- 1.278 Reid VP at 37.8 C (kPa) --- --- --- True VP at 37.8 C (kPa) --- --- --- Liq. Vol. Flow - Sum(Std. Cond) (m3/h) 146.2 0.0000 0.0000 ------------------------------------------------------------------------------- 30-F-001-B (Conversion Reactor): Design, Reactions, Worksheet ------------------------------------------------------------------------------- Conversion Reactor: 30-F-001-B CONNECTIONS Inlet Stream Connections

133

Stream Name From Unit Operation Combo Gas PSA Tail Gas-2 RCY-1 Recycle Combustion Air Outlet Stream Connections Stream Name To Unit Operation waste gas Heat Exchanger: Steam Generation Coil I DNE Energy Stream Connections Stream Name From Unit Operation Q-100 PARAMETERS Physical Parameters Optional Heat Transfer: Cooling Delta P Vessel Volume Duty Energy Stream 0.0000 kPa --- 3.657e+008 kJ/h Q-100 User Variables REACTION DETAILS Reaction: Meth Combust Component Mole Weight Stoichiometric Coeff. Methane 16.04 -1.000 Oxygen 32.00 -2.000 CO2 44.01 1.000 H2O 18.02 2.000 Reaction: Eth combust Component Mole Weight Stoichiometric Coeff. Ethane 30.07 -1.000 Oxygen 32.00 -3.500 CO2 44.01 2.000 H2O 18.02 3.000 Reaction: Prop combust Component Mole Weight Stoichiometric Coeff. Propane 44.10 -1.000

134

Oxygen 32.00 -5.000 CO2 44.01 3.000 H2O 18.02 4.000 Reaction: H2 combust Component Mole Weight Stoichiometric Coeff. Hydrogen 2.016 -2.000 Oxygen 32.00 -1.000 H2O 18.02 2.000 Reaction: 1-but combust Component Mole Weight Stoichiometric Coeff. Oxygen 32.00 -6.000 CO2 44.01 4.000 H2O 18.02 4.000 1-Butene 56.11 -1.000 Reaction: ethylene comust Component Mole Weight Stoichiometric Coeff. Oxygen 32.00 -3.000 CO2 44.01 2.000 H2O 18.02 2.000 Ethylene 28.05 -1.000 Reaction: i-but combust Component Mole Weight Stoichiometric Coeff. i-Butane 58.12 -1.000 Oxygen 32.00 -6.500 CO2 44.01 4.000 H2O 18.02 5.000 Reaction: i-pent combust Component Mole Weight Stoichiometric Coeff. Oxygen 32.00 -8.000 CO2 44.01 5.000 H2O 18.02 6.000 i-Pentane 72.15 -1.000 Reaction: n-but combust Component Mole Weight Stoichiometric Coeff.

135

Oxygen 32.00 -6.500 CO2 44.01 4.000 H2O 18.02 5.000 n-Butane 58.12 -1.000 Reaction: n-hex combust Component Mole Weight Stoichiometric Coeff. Oxygen 32.00 -9.500 CO2 44.01 6.000 H2O 18.02 7.000 n-Hexane 86.18 -1.000 Reaction: n-pent comust Component Mole Weight Stoichiometric Coeff. Oxygen 32.00 -8.000 CO2 44.01 5.000 H2O 18.02 6.000 n-Pentane 72.15 -1.000 Reaction: propene combust Component Mole Weight Stoichiometric Coeff. Propene 42.08 -1.000 Oxygen 32.00 -4.500 CO2 44.01 3.000 H2O 18.02 3.000 Reaction: tr2-but combust Component Mole Weight Stoichiometric Coeff. Oxygen 32.00 -6.000 CO2 44.01 4.000 H2O 18.02 4.000 tr2-Butene 56.11 -1.000 Reaction: Component Mole Weight Stoichiometric Coeff. REACTION RESULTS FOR : Combust

136

Extents Name Rank Specified Use Default Actual Base Reaction Extent % Conversion % Conversion Component (kgmole/h) Meth Combust 0 100.00 Yes 100.0 Methane 386.8 Eth combust 0 100.00 Yes 100.0 Ethane 3.059 Prop combust 0 100.00 Yes 100.0 Propane 1.397 H2 combust 0 100.00 Yes 100.0 Hydrogen 360.8 1-but combust 0 100.00 Yes 100.0 1-Butene 1.184e-002 ethylene comust 0 100.00 Yes 100.0 Ethylene 0.2066 i-but combust 0 100.00 Yes 100.0 i-Butane 5.057e-002 i-pent combust 0 100.00 Yes 100.0 i-Pentane 3.452e-003 n-but combust 0 100.00 Yes 100.0 n-Butane 6.928e-002 n-hex combust 0 100.00 Yes 100.0 n-Hexane 2.794e-003 n-pent comust 0 100.00 Yes 100.0 n-Pentane 8.400e-003 propene combust 0 100.00 Yes 100.0 Propene 0.2961 tr2-but combust 0 100.00 Yes 100.0 tr2-Butene 2.368e-003 CO Combust 0 100.00 Yes 100.0 CO 136.8 Balance Components Total Inflow Total Reaction Total Outflow (kgmole/h) (kgmole/h) (kgmole/h) Methane 386.8 -386.8 0.0000 Ethane 3.059 -3.059 0.0000 Propane 1.397 -1.397 0.0000 CO 136.8 -136.8 0.0000 CO2 94.96 535.8 630.7 Hydrogen 721.5 -721.5 0.0000 H2O 0.0000 1512 1512 Nitrogen 6390 7.994e-013 6390 Oxygen 1692 -1223 468.8 i-Butane 5.057e-002 -5.057e-002 0.0000 n-Butane 6.928e-002 -6.928e-002 0.0000 i-Pentane 3.452e-003 -3.452e-003 0.0000 n-Pentane 8.400e-003 -8.400e-003 0.0000 n-Hexane 2.794e-003 -2.794e-003 0.0000 Ethylene 0.2066 -0.2066 0.0000 tr2-Butene 2.368e-003 -2.368e-003 0.0000 1-Butene 1.184e-002 -1.184e-002 0.0000

137

Helium 0.9984 0.0000 0.9984 Propene 0.2961 -0.2961 0.0000 CONDITIONS Name Combo Gas PSA Tail Gas-2 Combustion Air DNE Vapour 1.0000 1.0000 1.0000 0.0000 Temperature (C) 55.0000 31.9294 300.0000 800.0000 Pressure (kPa) 146.3250 40.0000 102.8000 40.0000 Molar Flow (kgmole/h) 139.4231 1230.7530 8057.8742 0.0000 Mass Flow (kg/h) 830.2727 15697.9267 232471.8488 0.0000 Std Ideal Liq Vol Flow (m3/h) 4.9631 47.7418 268.7376 0.0000 Molar Enthalpy (kJ/kgmole) -1.617e+004 -6.439e+004 8248 -4.229e+004 Molar Entropy (kJ/kgmole-C) 142.2 167.7 171.1 211.1 Heat Flow (kJ/h) -2.2538e+06 -7.9244e+07 6.6465e+07 0.0000e-01 Name waste gas Vapour 1.0000 Temperature (C) 800.0000 Pressure (kPa) 40.0000 Molar Flow (kgmole/h) 9002.2071 Mass Flow (kg/h) 248998.0490 Std Ideal Liq Vol Flow (m3/h) 296.1206 Molar Enthalpy (kJ/kgmole) -4.229e+004 Molar Entropy (kJ/kgmole-C) 211.1 Heat Flow (kJ/h) -3.8073e+08 PROPERTIES Name Combo Gas PSA Tail Gas-2 Combustion Air DNE waste gas Molecular Weight 5.955 12.75 28.85 27.66 27.66 Molar Density (kgmole/m3) 5.362e-002 1.577e-002 2.157e-002 4.483e-003 4.483e-003 Mass Density (kg/m3) 0.3193 0.2012 0.6222 0.1240 0.1240 Act. Volume Flow (m3/h) 2600 7.804e+004 3.736e+005 0.0000 2.008e+006 Mass Enthalpy (kJ/kg) -2715 -5048 285.9 -1529 -1529 Mass Entropy (kJ/kg-C) 23.88 13.15 5.932 7.631 7.631 Heat Capacity (kJ/kgmole-C) 31.48 31.59 30.84 36.10 36.10 Mass Heat Capacity (kJ/kg-C) 5.286 2.477 1.069 1.305 1.305 Lower Heating Value (kJ/kgmole) 3.988e+005 3.873e+005 0.0000 0.0000 0.0000 Mass Lower Heating Value (kJ/kg) 6.697e+004 3.036e+004 --- --- --- Phase Fraction [Vol. Basis] --- --- --- --- --- Phase Fraction [Mass Basis] 4.941e-324 4.941e-324 4.941e-324 2.122e-314 2.122e-314 Partial Pressure of CO2 (kPa) 1.267e-002 3.086 0.0000 0.0000 2.803

138

Cost Based on Flow (Cost/s) 0.0000 0.0000 0.0000 0.0000 0.0000 Act. Gas Flow (ACT_m3/h) 2600 7.804e+004 3.736e+005 --- 2.008e+006 Avg. Liq. Density (kgmole/m3) 28.09 25.78 29.98 --- 30.40 Specific Heat (kJ/kgmole-C) 31.48 31.59 30.84 36.10 36.10 Std. Gas Flow (STD_m3/h) 3297 2.910e+004 1.905e+005 0.0000 2.129e+005 Std. Ideal Liq. Mass Density (kg/m3) 167.3 328.8 865.1 840.9 840.9 Act. Liq. Flow (m3/s) --- --- --- 0.0000 --- Z Factor 1.000 0.9999 1.000 --- --- Watson K 25.94 15.10 6.042 6.557 6.557 User Property --- --- --- --- --- Partial Pressure of H2S (kPa) 0.0000 0.0000 0.0000 0.0000 0.0000 Cp/(Cp - R) 1.359 1.357 1.369 1.299 1.299 Cp/Cv 1.360 1.358 1.370 1.299 1.299 Heat of Vap. (kJ/kgmole) 7195 --- 5891 3.433e+004 3.433e+004 Kinematic Viscosity (cSt) 29.15 51.08 48.58 2.408 350.4 Liq. Mass Density (Std. Cond) (kg/m3) --- --- --- --- --- Liq. Vol. Flow (Std. Cond) (m3/h) --- --- --- 0.0000 --- Liquid Fraction 0.0000 0.0000 0.0000 1.000 0.0000 Molar Volume (m3/kgmole) 18.65 63.41 46.37 223.1 223.1 Mass Heat of Vap. (kJ/kg) 1208 --- 204.2 1241 1241 Phase Fraction [Molar Basis] 1.0000 1.0000 1.0000 0.0000 1.0000 Surface Tension (dyne/cm) --- --- --- 0.0000 --- Thermal Conductivity (W/m-K) 0.1298 7.500e-002 4.360e-002 7.018e-002 7.294e-002 Viscosity (cP) 9.309e-003 1.028e-002 3.022e-002 2.985e-004 4.345e-002 Cv (Semi-Ideal) (kJ/kgmole-C) 23.16 23.28 22.53 27.78 27.78 Mass Cv (Semi-Ideal) (kJ/kg-C) 3.890 1.825 0.7808 1.004 1.004 Cv (kJ/kgmole-C) 23.14 23.27 22.52 27.78 27.78 Mass Cv (kJ/kg-C) 3.886 1.824 0.7805 1.004 1.004 Cv (Ent. Method) (kJ/kgmole-C) 23.12 --- 22.48 --- --- Mass Cv (Ent. Method) (kJ/kg-C) 3.882 --- 0.7791 --- --- Cp/Cv (Ent. Method) 1.362 --- 1.372 --- --- Reid VP at 37.8 C (kPa) --- --- --- --- --- True VP at 37.8 C (kPa) --- --- --- 2.042e+005 2.042e+005 Liq. Vol. Flow - Sum(Std. Cond) (m3/h) 0.0000 0.0000 0.0000 0.0000 0.0000 ------------------------------------------------------------------------------- Hyprotech Ltd. Aspen HYSYS Version 2006 (20.0.0.6728)

139

Appendix F: AMSIM Reports

140

Table F. 1: Composition Profile of CO2 in Absorber A

Stage Amine Sol. [Fraction] Vapor [Fraction] 1 0.022656 0.011879 2 0.02514 0.018135 3 0.029085 0.027852 4 0.035361 0.042966 5 0.044617 0.066199 6 0.052562 0.098851 7 0.056291 0.1255 8 0.058416 0.137557 9 0.059852 0.144295

10 0.060906 0.148791 11 0.061711 0.152062 12 0.062343 0.15455 13 0.062871 0.156513 14 0.063431 0.158226 15 0.064456 0.160386

AMSIM 7.2 --- Copyright by Oilphase-DBR, Schlumberger Canada Ltd.

Table F. 2: Vapour Phase Properties in Absorber A

Stage Pressure Temperature Mass Flow

Vol. Flow

Molar Flow

Mol. Weight Density

[ kPa ] [ C ] [ kg/h ] [ m3/h ] [ kmol/h ] [ kg/kmol ] [ kg/m3 ] 1 2163.6 39.7 22174.9 5544.1 4554.758 4.869 3.99972 2 2166.2 42.7 23530.5 5635 4591.706 5.125 4.17575 3 2168.8 47.4 25642 5773.5 4642.399 5.523 4.44134 4 2171.4 54.4 29053.8 5995.3 4725.238 6.149 4.84608 5 2174 63.2 34577 6320.1 4860.23 7.114 5.47097 6 2176.6 69.8 42741.5 6689 5056.142 8.453 6.38986 7 2179.2 72.8 49796.9 6955.6 5222.87 9.534 7.15928 8 2181.8 74.5 53185.8 7086.8 5303.828 10.028 7.50496 9 2184.5 75.7 55153 7163.7 5351.418 10.306 7.69896

10 2187.1 76.5 56496.9 7215.1 5384.172 10.493 7.83042 11 2189.7 77.1 57488.7 7251 5408.421 10.629 7.92837 12 2192.3 77.6 58245.7 7275.6 5426.863 10.733 8.00566 13 2194.9 77.7 58822.1 7287.3 5440.412 10.812 8.07193 14 2197.5 76.7 59210.2 7267.6 5447.056 10.87 8.14714 15 2200.1 70.4 59235.9 7111.8 5434.766 10.899 8.32927


141

Table F. 3: Liquid Phase Properties in Absorber A

Stage Pressure Temperature Mass Flow

Vol. Flow

Molar Flow

Mol. Weight Density

[ kPa ] [ C ] [ kg/h ] [ m3/h ] [ kmol/h ] [ kg/kmol ] [ kg/m3 ] 1 2163.6 39.7 420406.8 409.9 18023.88 23.325 1025.582 2 2166.2 42.7 422518.3 411.3 18074.58 23.376 1027.243 3 2168.8 47.4 425930.1 413.6 18157.42 23.458 1029.801 4 2171.4 54.4 431453.3 417.3 18292.41 23.586 1033.813 5 2174 63.2 439617.8 422.7 18488.32 23.778 1040.047 6 2176.6 69.8 446673.2 427.2 18655.05 23.944 1045.676 7 2179.2 72.8 450062.1 430.9 18736.01 24.021 1044.576 8 2181.8 74.5 452029.3 433.1 18783.6 24.065 1043.682 9 2184.5 75.7 453373.2 434.7 18816.35 24.095 1043.072

10 2187.1 76.5 454365 435.8 18840.6 24.116 1042.622 11 2189.7 77.1 455122 436.7 18859.04 24.133 1042.284 12 2192.3 77.6 455698.4 437.3 18872.59 24.146 1042.045 13 2194.9 77.7 456086.5 437.7 18879.23 24.158 1041.978 14 2197.5 76.7 456112.2 437.5 18866.94 24.175 1042.501 15 2200.1 70.4 455473.2 435.5 18809.79 24.215 1045.839


Table F. 4: Composition Profile of CO2 in Regenerator Stage Amine Sol. [Fraction] Vapor [Fraction]

1 0.000665 0.941076 2 0.059573 0.666759 3 0.055604 0.539271 4 0.050837 0.398938 5 0.046309 0.29988 6 0.04242 0.237417 7 0.039191 0.194647 8 0.036465 0.163104 9 0.034147 0.138419

10 0.032144 0.118575 11 0.030382 0.102133 12 0.028796 0.088182 13 0.027346 0.076009 14 0.025995 0.065202 15 0.02471 0.055401 16 0.023481 0.046308 17 0.02095 0.037835


142

Table F. 5: Vapour Phase Properties in the Regenerator Stage Pressure Temperature Mass Flow Vol. Flow Molar Flow Mol. Weight Density

[ kPa ] [ C ] [ kg/h ] [ m3/h ] [ kmol/h ] [ kg/kmol ] [ kg/m3 ] 1 199.9 48.9 36737.6 11477.1 864.864 42.478 3.20096 2 206.2 92 43279.6 17837.9 1220.853 35.45 2.42627 3 207.1 101 46380.6 21455.7 1440.796 32.191 2.1617 4 208 108.6 51643.5 27281.1 1805.303 28.607 1.89302 5 208.9 113.2 56016.7 32668.1 2148.032 26.078 1.71473 6 209.8 115.9 58180.4 36198.7 2376.248 24.484 1.60726 7 210.7 117.7 59093.7 38464.8 2526.154 23.393 1.53632 8 211.7 119 59557.6 40086.8 2636.715 22.588 1.48572 9 212.6 120 59852.7 41352.5 2725.796 21.958 1.44738

10 213.5 120.8 60060.2 42367.7 2799.826 21.451 1.41761 11 214.4 121.5 60253.3 43230.2 2864.872 21.032 1.39379 12 215.3 122.2 60440.2 43979.1 2923.255 20.676 1.37429 13 216.3 122.7 60619.5 44642.8 2976.664 20.365 1.35789 14 217.2 123.2 60796.7 45242.4 3026.364 20.089 1.34381 15 218.1 123.7 60971.3 45794.2 3073.336 19.839 1.33143 16 219 124.1 61145.9 46313.3 3118.448 19.608 1.32028 17 219.9 125 61490.7 46824.8 3159.503 19.462 1.31321


Table F. 6: Liquid Phase Properties in the Regenerator Stage Pressure Temperature Mass Flow Vol. Flow Molar Flow Mol. Weight Density

[ kPa ] [ C ] [ kg/h ] [ m3/h ] [ kmol/h ] [ kg/kmol ] [ kg/m3 ] 1 199.9 48.9 6545.1 6.6 355.989 18.386 998.6405 2 206.2 92 464022.1 449 19354.28 23.975 1033.371 3 207.1 101 469290.5 456.8 19718.78 23.799 1027.448 4 208 108.6 473650.3 463.4 20061.51 23.61 1022.022 5 208.9 113.2 475807 468.6 20289.73 23.451 1015.443 6 209.8 115.9 476718.3 472.3 20439.63 23.323 1009.263 7 210.7 117.7 477181.5 475.1 20550.19 23.22 1004.37 8 211.7 119 477476.4 477.3 20639.28 23.134 1000.343 9 212.6 120 477683.7 479.1 20713.31 23.062 996.9533

10 213.5 120.8 477876.6 480.7 20778.35 22.999 994.0457 11 214.4 121.5 478063.4 482.2 20836.73 22.943 991.4991 12 215.3 122.2 478242.7 483.5 20890.14 22.893 989.213 13 216.3 122.7 478419.8 484.7 20939.84 22.847 987.1263 14 217.2 123.2 478594.3 485.8 20986.82 22.805 985.1855 15 218.1 123.7 478768.9 486.9 21031.93 22.764 983.3416 16 219 124.1 479113.6 488.1 21072.98 22.736 981.529 17 219.9 125 417622.9 427.8 17913.48 23.313 976.2471


group g - final report

Documents

tonne co2

hour of co2

co2 product

mark dalton removal

dow chemical canada

hydrogen plant

removal system

husky conduct