separating methane and carbon dioxide

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Separating methane and carbon dioxide

from enhanced oil recovery

An Economic Evaluation of Four Common Separation Methods

Authors: Chandler Stiles Water Absorption

James Gamble Jr. MDEA Absorption

Quoc Tran Membrane Separation

Logan Renninger Cryogenic Distillation

Date: April 25, 2016


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1 TABLE OF CONTENTS

2 Executive Summary .....................................................................................................................3

3 Introduction ...............................................................................................................................4

3.1 Natural Gas ................................................................................................................................... 4

3.2 Carbon Dioxide .............................................................................................................................. 4

3.3 Enhanced Oil Recovery ................................................................................................................. 4

4 Water Absorption .......................................................................................................................6

4.1 Cost Considerations ...................................................................................................................... 6

4.1.1 Column Cost .............................................................................................................................. 6

4.1.2 Water Cost ................................................................................................................................ 6

4.1.3 Pressurization Cost .................................................................................................................... 6

4.2 Design Analysis .............................................................................................................................. 6

4.2.1 Choosing L ................................................................................................................................. 7

4.2.2 Choosing Temperature ............................................................................................................. 8

4.2.3 Choosing Pressure ..................................................................................................................... 8

4.2.4 Column Height .......................................................................................................................... 9

4.3 Summary ....................................................................................................................................... 9

5 MDEA Absoprtion ..................................................................................................................... 10

5.1 Initial Design and Assumptions ................................................................................................... 11

5.2 Cost Guidelines ........................................................................................................................... 13

5.3 Relevant Variable Analysis .......................................................................................................... 14

5.3.1 Optimization 1 ......................................................................................................................... 14

5.3.2 Optimization 2 ......................................................................................................................... 14

5.4 Conclusion ................................................................................................................................... 15

6 Membrane Separation .............................................................................................................. 16

6.1 Cost Considerations .................................................................................................................... 16

6.1.1 Capital Costs ............................................................................................................................ 16

6.1.2 Operating Costs ....................................................................................................................... 16

6.2 Design Analysis ............................................................................................................................ 17

6.2.1 Single Stage Membrane System .............................................................................................. 17

6.2.2 Single Stage Membrane System with Compressor .................................................................. 20

6.2.3 Two Stage Counter-Current Membrane System with Compressors ........................................ 22


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2 EXECUTIVE SUMMARY

This report conducts a technical and economic analysis to determine the most profitable course

of action for a natural gas stream from an enhanced oil recovery process. Composing of 18%

methane and 82% percent carbon dioxide, the raw feed can be sold for $ 8.3 million per year.

By separating the stream into a methane-rich stream and a carbon dioxide-rich stream, higher

profits may be achieved. This report evaluates four common gas separation methods: water

absorption, MDEA absorption, membrane separation, and cryogenic distillation. Using

membrane separation to produce high energy natural gas was found to be the most profitable

option generating a net $ 44.5 million per year. Before initiating plans for construction, we

recommend further research to ensure the accuracy of our evaluation and the sustainability of

such a venture.


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3 I

NTRODUCTION

3.1

NATURAL GAS

Natural gas serves as a source for safe, clean, and efficient energy apart from traditional fossil

fuels and coal. This alternative clean-burning fuel can take different qualities of energy. Its

versatility heats homes, fuels industries, and serves as raw material for products while emitting

a low amount greenhouse gases.1 In the United States, the total natural gas consumption was

estimated to be 27,473,081 MMcf in 2015.2 Energy consumption and natural gas demand is

expected to rapidly grow in the future. In order to meet growing demands of energy

worldwide, exploration and extraction of this abundant resource will help supply countries with

natural gas.

3.2 CARBON DIOXIDE

When emitted to the atmosphere from human activities, carbon dioxide is considered a

greenhouse gas. Sources of carbon dioxide emissions include electricity, transportation, and

industry.3 To reduce the emissions of carbon dioxide, capturing and purifying this gas may be

beneficial for the food and beverage industry. Additionally, plants such as algae may use

carbon dioxide to produce bio-fuel.4

3.3

ENHANCED OIL RECOVERY

Oil recovery can occur up to three phases: primary, secondary, and tertiary. The primary phaseof oil recovery occurs when the natural pressure of the reservoir produces oil. Only 10% of the

original amount of oil in a reservoir is recovered of the primary phase. When the reservoir is no

longer pressurized to produce on its own, secondary recovery is utilized to repressurize the

reservoir through the use of water flooding techniques or gas injection. These methods can

recover 20% to 40% of the original reservoir. Efforts to extend the effective life of a reservoir

after secondary recovery is known as tertiary or enhanced oil recovery (EOR). EOR can recover

30% to 60% of the oil in the reservoir through various techniques includes practices such as

thermal recovery, gas injection, and chemical injection.5

One type of EOR includes the gas injection of carbon dioxide (CO2) into the reservoir from an

injection well. The CO2 expands throughout the reservoir and encounters pushes trapped oil

1 “Natural Gas.” Chevron Corporation. May 20152 “Natural Gas.” U.S. Energy Information Administration. April 21, 2016. 3 “Carbon Dioxide Emissions.” U.S. Environmental Protection Agency. April 15, 2016.4 “Carbon Dioxide.” The Linde Group. 20165 “Enhanced Oil Recovery.” U.S. Department of Energy . April 21, 2016.


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toward the producing well. Not only does this injection push crude oil, but it can also drive

residual natural gas up for production as well. Consequently, the natural gas from this method

of EOR can contain varying concentrations of CO2 and CH4 over time.

Depending on the composition, separation of the production containing both CO2 and CH4 can

yield various products. Natural gas can be sold as low energy gas (10-40% CH4), medium energygas (40-70% CH4), high energy gas (70-95% CH4), pipeline quality (95% CH4), and liquefied

natural gas (99.995 %). Figure 1.1 shows the price decks are used to sell different qualities of

natural gas product.

Product Grade Purity Price

Waste Gas < 10% -

Low Energy Gas 10 - 40 % $ 2.50

Medium Energy Gas 40 - 70 % $ 3.00

High Energy Gas 70 - 95 % $ 3.50

Pipline Quality Gas 95% $ 4.00Liquefied Natural Gas (LNG) < 50 ppm CO2 $ 5.20

Figure 1.1. Price estimates by methane concentration in natural gas

Despite a limited, market, CO2 product (>99.5% CO2) can be sold and used in the beverage

industry to make soft drinks. Additionally, pure CO2 can be used in chemical processes such as

EOR as described above.

From the production of EOR, a feed of 60 MSCFD or with a composition of 82% CO2 and 18%

CH4 will be analyzed with different unit operations to determine economic degrees of

separation. These unit operations include methyl diethanolamine (MDEA) absorption, waterabsorption, cryogenic distillation, and gas membrane purification. For the separation to be

economically feasible, the overall value of the products must be greater than the value of the

feed and cost to separate with the given unit operations. The following equation gives the

value of the feed sold as low energy gas at a price deck of $2.5/MMBTU6:

Feed Value

= molhr 0.18 mol CH 890.7

mol 0.947817 BTU

1 kJ 1 MMBTU10 BTU

$2.5MMBTU

6900 hryear

= $8,284,318.39 per yearThis equation uses the heat of combustion, CH4 composition, and operating time to calculatethe value of the feed stream as low energy gas. Since the feed is below 82% CO2, it cannot be

sold as CO2 product. Moreover, if the profit from separation is greater than the feed’s original

value of $8,284,318.39 per year, then the unit operation will be determined as economical.

6 Pittam, D. A.; Pilcjer, G. J. Chem. Soc. Faraday Trans. 1972


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4 WATER ABSORPTION

Absorption is a process by which a substance in a gas state is transferred into another

substance in a liquid state. Water can be used in absorption of carbon dioxide from a gas

mixture containing carbon dioxide and methane. The reason for this is because the carbon

dioxide is water-soluble while the methane is water-insoluble. The amount of carbon dioxide

absorbed is related to the amount of contact between the gas and liquid phases. For this

reason, many absorption columns contain packing material to increase the contact area. In this

analysis, 1” Rashig rings will be used as packing material in the column.

4.1

C

OST

C

ONSIDERATIONS

When it comes to the cost and efficiency of an absorption column, the success of absorption

itself is the greatest factor. Things that affect the success of absorption are the size of the

column, life span of the packing material, amount of water used, pressure, and temperature. All

of which contribute their own cost to the process.

4.1.1 Column Cost

The column cost with packing material can be estimated to be $1800/ft3. The packing material

needs to be replaced every 30 years which gives a yearly cost of $60/ft3-yr.

4.1.2 Water Cost

The cost of water is one of the dominating costs in this process. Since water is not the best at

absorbing carbon dioxide like MDEA is, it requires a much larger volumetric flow rate. The

water cost is roughly $0.15/1000 gal.

4.1.3

Pressurization Cost

Depending on the pressure that is optimal to run the feed at, it may be beneficial to change the

pressure of the feed stream. To calculate how much power this requires the equation is

Power=VV(R*T*ln(P2/P1) for the gas stream. Also the pressure of the water stream needs to be

accounted for and adjusted to the same pressure of the gas feed to make sure both streams

flow through the column properly. The equation for this is Power=VL(P2-P1). The price for the

electricity to power the pumps is around $0.08/kW-hr.

4.2 DESIGN ANALYSIS

The design of the absorption column was based off of a few assumptions:

1. The water feed into the column contains no carbon dioxide.

2. The absorption column runs at room temperature.

3. The Henry’s law constant is valid because carbon dioxide would be dilute in water

and could be considered constant throughout the column for carbon dioxide at 25oC

which is .8317.


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4.

Ideal gas law was used whenever needed.

5.

Pressure drop in column is negligible.

With these assumptions the absorption column shown in figure 4.1 can help to solve the

equations 4.1 to 4.5.

Figure 4.1. Absorption Column Design

4.2.1 Choosing L

Figure 4.2 shows that the optimal flow rate for the water stream happens between a multiple

of 1 and 1.2 of Lmin. This figure also shows that a product feed containing 99.995% methane will

maximize profit. When obtaining a product stream of this concentration you want to use

L=1.3*Lmin for max profit.7

7 Appendix : Table 9.4.1

Vin=60

MMSCFD

YV,CO2=.82

Vout

YV,CH4

L XL,CO2=0

L XL,CO2

Equations

4.1 CA= ∗ ∗

4.2 Vout=∗,,

,,

4.3 Lmin =

∗,, ( ∗ ,, ,,) 4.4 CL,out,CO2 = ( ∗ ,, ∗,,) 4.5 A=

∗∗∗−∗

∗ ,,−,,−∗∗,,+∗∗,,−,,


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4.2.2 Choosing Temperature

Changing the temperature requires energy which comes at a cost. To increase the profit the

temperature would need to decrease but to decrease the temperature the gas needs to be

cooled. To cool the gas stream it would require more energy and more water which would

increase the cost way more than it helps the profit. So for this absorption process the

temperature is left at room temperature since it creates the smallest cost.

4.2.3 Choosing Pressure

In the absorption process increasing the pressure increases the amount of carbon dioxide that

can dissolve into the water. Since a product feed of 99.995% methane is the most profitable it is

worth looking at pressure change to this process to find the most profitable pressure to operate

at. Figure 4.3 graphs the pressure versus profit to find the maximum profit. This shows that the

most profitable pressure is at 1500 psi.

$0.00

$2,000,000.00

$4,000,000.00

$6,000,000.00

$8,000,000.00

$10,000,000.00

$12,000,000.00

$14,000,000.00

$16,000,000.00

1.0 1.2 1.4 1.6 1.8 2.0 2.2

P r o f i t

Multiple of Lmin

CH4=.4

CH4=.5

CH4=.7

CH4=.95

CH4=.99995

Figure 4.2 Optimization of Liquid Flow Rate


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Figure 4.3 Pressure that Maximizes Profit

4.2.4 Column Height

The optimal column height is when the height/diameter is equal to 6. For this process the

optimal conditions give a column height of 18m and a diameter of 3m. This is a little tall but at

the industrial level this column height would be fine.

4.3

S

UMMARY

All of the revenue in this absorption process is from the sale of methane and no profit comes

from carbon dioxide because it is not pure enough.

Optimal Conditions:

Pressure-1500 psi

Water Flow Rate-1122 m3/hr

Temperatue-298 K

Height-18 M

With the feed we were given and the optimal conditions listed above, we get a profit of $13.8

million per year which is not bad for absorption but other processes would be able to generate

more profit and are discussed later in this report.

$13,300,000.00

$13,400,000.00

$13,500,000.00

$13,600,000.00

$13,700,000.00

$13,800,000.00

$13,900,000.00

$14,000,000.00

$14,100,000.00

$14,200,000.00

0 500 1000 1500 2000

P r o f i t

Pressure (psi)

CH4=.99995


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5 MDEA ABSOPRTION

Separation of CO2 from Natural Gas is facilitated by using an aqueous solution of MDEA as the

solvent due to its reactivity with the acid gasses. When considering enhanced oil recovery the

acid gas is CO2 and the term acid gas is used because it is attracted to the basicity of the MDEA.

The basic nature of MDEA is as a result of its amine functional group. This acid-base interaction

between MDEA and CO2 allows for a better separation than in a neutral solvent like water. The

hydroxyl groups on the MDEA allow it to be soluble in water forming an aqueous solution.8 This

aqueous solution formation is pertinent to the separation because the MDEA can’t react

directly with CO2, but they can react with each other in-situ.1 Other attributes that make CO2 a

great solvent are that it has a very slow degradation rate and a rather high loading capacity.9

The direct effect on the profit of the slow degradation be seen in the cost analysis portion of

this operation.

The specific MDEA used was 2M and 23.3 wt% MDEA in solution at 313K. This type of MDEA

was employed due to its high loading capacity over a range of partial pressures of CO2. Figure 1

below shows the loading capacity ranges of MDEA at different temperatures.

Figure 5.1. Plot of Partial Pressure CO2 vs Loading1

This plot exhibits the large capacity of MDEA at 313k hence why the separation proceeded at313K. The principal disadvantage of using MDEA is the large cost associated, but detrimental

financial impact is thwarted due to the long life of MDEA.

8 Vrachnos A., Kontogeorgis G., Voutsas E, Ind. Eng. Chem. Res. 2006, 45, 5148-51549 STEPHEN T. DONNELLY Bryan Research & Engineering, Inc., Bryan, Texas, Propak Systems, Inc.,

Lakewood, Colorado

Loading (mol CO2/mol MDEA)

P a r t i a l p r e s s

u r e C O 2 k P a


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5.1

I

NITIAL

D

ESIGN AND

A

SSUMPTIONS

When designing the MDEA separation principal objective was to yield a product significantly

more pure than our feed stream. The MDEA separation design was governed by a few major

simplifying assumptions. The first being that the methane and MDEA are completely immiscible

at the pressures and flow rates investigated which means that all of the CH4 in the feed is also

in the purified natural gas product. The application of this assumption allows for the easy

computation of the outgoing vapor molar flow rates. The calculation for a 40% mole percent

methane product can be seen below.

(,) = (,) 877.87 . 18 = . 40 → = 395.04

The second assumption is that the liquid flow rate is constant throughout the column and thatthe MDEA feed is supplied at 313k. The last two design restrictions are the temperature and

MDEA used. As has been previously stated the 23.3 wt% MDEA has the best loading at 313 so

the design parameters will abide by these conditions. The counter-flow absorption column can

be seen in figure 5.2 below.

Figure 5.2. Design of counter-flow absorbance column

As can be seen in the diagram above the vapor inlet and liquid outlet are at the same end of the

absorption column which allows the developing of an equilibrium relationship through the

Henry’s Law constant. The same happens at the top of the absorption column at the vapor

Vapor Out

Vapor InLiquid Out

Liquid In


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outlet/liquid inlet. The Henry’s law constant provides an avenue for relating the concentrations

in the vapor and liquid phases respectfully. The Henry’s Law constants can be seen in equations

2 and 3 below.

= ,

,

= ,, These equations hold true, but when H is not constant H = √ can be used, but forthis design another key assumption was that when calculating volume and monetary values

only the inlet figures were used. This helped simplify the design calculations because varying H

values, flowrates, etc. did not have to be considered. The necessary volume of the column

increases exponentially with the purity of CH4 in the product stream. The height is a function of

area as can be seen in equation 4 below.

ℎ = ∗ − ⁄ ,−, ⁄

− ⁄ ,+ ⁄ ∗,−, ⁄

Since the height is a function of the area, every time the height is varied at the keeping

everything else the same, the area would just compensate the change in height to keep the

volume constant. As a result of the area being a component of the height formula, the volume

was just calculated by multiplying both sides by the area. This produces hA which equals

volume, on the left side of the equation while cancelling out the area term on the right side of

the equation. Since computation of volume is possible without the height or area, the volumes

of the absorption column and its effect on price were investigated instead of the height effect

on profit. A plot of volume vs CH4 product stream concentration at qL = 1.5 qL,min can be seen in

figure 5.4 below.

< 2 >

< 3 >

< 4 >


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Figure 5.3: Plot of Column Volume vs CH4 purity

The large jump at the in volume when concentration approaches 1 is the principle reason why

the LNG product was not economically feasible. The cost of separation due to the volume is an

analyzed in the next section of this report. For sizing purposes, the smallest volume possible

should be used and the ratio of h/d should be as close to 6 as possible assuming a cylinder

shaped absorbance column.

5.2 COST GUIDELINES

Numerous costs had to be taken into consideration when determining the profit that a certain

stream could produce. The first cost considered was the capital cost for the volume of thecolumn due to the Raschig Rings used for separation. The price of Raschig Rings is $1800/ft3 of

column volume. This value changed dramatically over the different desired methane gas

product streams. The cost of Raschig Rings per year is relatively large, but since their lifetime is

approximately thirty years and assuming not much maintenance is needed, they will eventually

pay for themselves. The second and largest cost considered for any profitable product stream is

MDEA. The price of the MDEA is approximately 15749104 x qL, but again, as was previously

stated, the large cost is compensated for by the regeneration and long life of the MDEA. The

two costs involved with this separation are pressurization costs. The MDEA has to be

condensed to the same pressure as feed in order for them to mix. This pressurization cost is a

not a recursive cost and ranges from .2*677008 - .4*667008, where .2 and .4 are molar flow

rates in units of m3/s based on the intended methane product stream.

0

2000

4000

6000

8000

10000

12000

14000

16000

0 0.2 0.4 0.6 0.8 1 1.2

C o l u m V o l u m e

( m 3 )

XCH4


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5.3

R

ELEVANT

V

ARIABLE

A

NALYSIS

With the given conditions of our design the two variables that can be analyzed for their direct

impact on the profit are pressure and flowrate. There is a flowrate at which the profit is

maximized and will be reviewed in the optimization portion of this operation. The profit is

maximized when the costs are minimized. Since there are two variables being analyzed there is

an optimal situation for flowrate at constant pressure and there is an optimal pressure trend

we will investigate at constant flowrate.

5.3.1 Optimization 1

The first variable optimized was the flowrate of MDEA. The pressure of the system was held at

a constant value of 420 psig which is the feed pressure. The flowrate term is involved in the

MDEA cost as well as the Raschig Ring cost for the separation. Since increasing the flowrate

doesn’t vary the MDEA cost in a linear fashion like it does for Raschig Ring cost, but rather in a

polynomial fashion, the profit vs flowrate plot, didn’t have a simple linear trend. The two

streams that created the largest revenue were the lowest purity option of medium energy gas

(40% CH4) and the lowest purity form of medium energy gas (70% CH4). The optimization of

profit plot for 40% and 70% methane product streams can be seen below in figure 5.3 below.

Figure 5.4: Plot of profit vs flowrate for 40% and 70% methane product streams

Plotting profit vs flowrate allowed for the discovery of the optimum flowrates for 40% methane

product stream and 70% product stream which are ql=1.18qL,min and qL=1.34qL,min respectively.

5.3.2 Optimization 2

The second optimization is varying pressure while keeping the flowrate constant. Since there is

no cap on pressure a general trend was developed based on a flowrate of 1.18qL at three

different pressures. 1.8qL,min is being employed because it produces the optimum profit when

selling a 40% CH4 product stream. The plot and table of profit vs pressure are below in table 5.1

and figure 5.5.

2.00E+06

2.50E+06

3.00E+06

3.50E+06

4.00E+06

4.50E+06

5.00E+06

5.50E+06

0.9 1 1.1 1.2 1.3 1.4 1.5 1.6

P r o f i t ( $ )

qL(m3/s)

40% Opt

70% Opt


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Table 5.1: Profit of 40% CH4 product at various pressures

Pressure(psig) qL(m3/S) Profit

150 0.223 3.01E+06

300 0.223 4.69E+06

420 0.223 5.09E+06600 0.223 5.34E+06

750 0.223 5.42E+06

Figure 5.5: Plot of Profit vs Pressure at qL = 1.18 qL,min

This chart and plot demonstrates the general trend that as pressure increases with everything

else held constant the profit increases. The pressure data was calculated up to 750psig for

simplicity sake because at higher pressures other complications occur and there is a need forother safeguarding equipment just to be able to run properly. These complications and extra

equipment obviously incur other cost so pressures above 750 psig were not considered.

5.4 CONCLUSION

The optimization allows for the increase of profit by finding the most efficient flowrates at

specific purities of CH4 product streams. The pressure trend also shows that the profit increases

with pressure. With the optimal flowrate a profit of $5.089 million/yr can be made using MDEA

separation. This figure can be increased with pressure and amount of separations, but only one

separation was considered for simplicity. Due to costs only medium and high energy gas couldbe produced. The recommendation for an engineer with the given conditions using MDEA

separation is to produce and sell 40% CH4 rich gas.

0.00E+00

1.00E+06

2.00E+06

3.00E+06

4.00E+06

5.00E+06

6.00E+06

0 200 400 600 800

P

r o f i t ( $ / y r )

Pressure(psig)


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6 MEMBRANE SEPARATION

Membrane plants have been installed to treat natural gas production through the removal of

carbon dioxide. Historically, membranes that have been used in this process include selective

cellulose acetate membranes with CO2/CH4 selectivities of 12-15 in small scale operations (5-40

MMSCFD). Large membrane systems (>40 MMSFD), however, have also been employed on

offshore platforms and enhanced oil recovery operations. Since the overall composition of

production from carbon dioxide flooding projects may change over time, polyimide and

polyaramide membranes with CO2/CH4 selectivities of 20-25 are being used to replace cellulose

acetate membranes for better carbon dioxide removal. Performance of this particular

separation process is largely dictated by the membrane’s permeability and selectivity of gas

components.10 Since economic advantages associated with gas membrane separation include

operating flexibility and low capital investment, a 60 MMSCFD natural gas feed of 82% CO2 and

18% CH4 will be analyzed with models using polysulfone membranes to determine the optimal

separation design.11

6.1 COST CONSIDERATIONS

Gas separation membrane includes capital investment of membranes and operating costs of

pressurization of the feed/or permeate streams. Operating costs will also include the cost to

heat the feed to carbon dioxide’s critical temperature of 304.14 K.12

6.1.1 Capital Costs

Capital costs with gas separation membrane is the investment of the membrane. The costs are

$4/ft2 with an expected life-time of one year. The permeances of the membranes are given as

follows:

K(CH4) = 1.04 x 10-4 mol/psia ft2 hr

K(CO2) = 0.00778 mol/psia ft2 hr

Since the permeance of carbon dioxide is greater than the permeance of methane, carbon

dioxide will be removed through the polysulfone membrane.

6.1.2

Operating Costs

The energy costs associated with pressurizing the feed/or permeate streams with a gas

compressor are given as:

Cost to compress/year = ṅRTln PP × $.w × w,, × year (Eq. 6.1) 10 Abedini, R.; Bezhadmogadam, A., S. Petroleum and Coal . 2010 11 Baker, R. Ind. Eng. Chem. Res. 2002 12 Suehiro, N.; J. Chem. Thermodyn., 1996


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where ṅ is molar flow rate in units of mol/hr and 6900 hours per year is the operating time ofthe plant. Operating temperature in units of K and the gas constant R in units of J/mol K were

also used in calculating compression costs.

To prevent the gas mixture from condensing at high operating pressures, heat exchanger costs

will be factored in to heat the appropriate streams above the critical temperature of the carbondioxide to ensure that the carbon dioxide remains in the vapor phase. Heating requirements

are determined by the following equation:

ℎ/ = ∆ × ṅ × × . × × × $. (Eq. 6.2)

where the following assumptions are made:

1. Cp of mixture is approximately 36 J/mol K13 (Heat capacities of methane and carbon dioxide are very similar to one another to so

this approximation is used) 2. Tinitial = 298 K 3. Tfinal = 304.18 K (critical temperature of carbon dioxide)

6.2 DESIGN ANALYSIS

The following cases will be designed and evaluated to determine the most economical gas separation

membrane design: single stage system without compression, single stage system with compression, two

stage countercurrent system with compression, and three stage counter current system with

compression.

Additionally, the following assumptions were made during the overall calculations and analysis of the

four main cases:

1. Feed exhibits ideal gas behavior

2.

Constant Temperature in the overall membrane system unless a heat exchanger is

specifically used to heat streams

3.

No pressure drop occurs between entrance and exit of membrane system

4.

The permeate pressure of the membrane is 15 psia (approximate atmospheric pressure)

6.2.1

Single Stage Membrane System

A preliminary single stage depicted in Figure 6.1 case was analyzed without any compression of

the feed gas to determine the profitable products from the retentate and/or permeate streams.

The only operating cost to heat this feed stream to 304.18 K was calculated to be $13,795.11.Moreover, the capital cost will depend on the size of the membrane. The following parameters

were defined to determine that quality of the retentate and permeate streams that yielded the

most profit:

1. Pr = 434.4 psia (absolute feed pressure)

13 Chase, N.; J. Phys. Chem. Ref. Data, Monograph 9., 1998


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2.

Pp = 15 psia (approximate atmospheric pressure)

3.

yCO2f = 0.82

4.

yCH4f = 0.18

5. Vf = 3,160,379.494 mol/h

6. yCH4r varies from 0.2 to 0.95

Figure 6.1 Base Case-Single Stage Membrane

Furthermore, the parameters were used to solve the following system of six equations to

determine area of the membrane (A), retentate molar flow rate (Vr), permeate molar flow rate(Vp), carbon dioxide mole fraction in retentate (yCO2r), methane mole fraction in permeate

(yCH4p), and carbon dioxide mole fraction in permeate (yCO2p).

= (Eq. 6.3) k ( ) = (Eq. 6.4) k ( ) = (Eq. 6.5)

= (Eq. 6.6) = 1 (Eq. 6.7) = 1 (Eq. 6.8)

By varying the final methane retentate composition, it was found that the mole fraction of carbon

dioxide in the permeate (yCO2p) was greater than 0.995 from yCH4r = 0.2 to 0.26. Between these

mole fraction ranges, the methane retentate can be sold as low energy gas at $2.5/MMBTU and

the permeate can be sold as carbon dioxide product at $0.003/SCF. Figure 6.2 depicts the

maximum profit of approximately $22.6 M/year from the low energy gas and carbon dioxide

products in the range of yCH4r = 0.2 to 0.26.14

14 Appendix: Figure 9.6.1-9.6.2


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Page | 19

Figure 6.2 Base-Case Permeate and Retentate Profit vs ych4r from yCH4r = 0.2 to 0.26

After yCH4r = 0.26, the carbon dioxide mole fraction in the permeate feed drops below 0.995

and can no longer be sold along with methane products. Figure 6.3 illustrates the behavior of

profits after yCH4r = 0.26 to 0.6.

Figure 6.3 Base-Case Retentate Profit vs ych4r from yCH4r = 0.26 to 0.6

$-

$5,000,000.00

$10,000,000.00

$15,000,000.00

$20,000,000.00

$25,000,000.00

0.2 0.22 0.24 0.26 0.28

$ / y e a r

yCH4r

Low energy gas and

Carbon Dioxide

$-

$1,000,000.00

$2,000,000.00

$3,000,000.00

$4,000,000.00

$5,000,000.00

$6,000,000.00

$7,000,000.00

0.2 0.3 0.4 0.5 0.6 0.7

$ / y e a r

yCH4r

Low Energy Gas

Medium Energy Gas


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Page | 20

Profits after yCH4r = 0.6 are negative due to the high capital cost of the membrane and are not

included. It is important to note that only the methane retentate stream can only be sold as

low energy gas and medium energy gas. As a result, the maximum profit is only $5.7 M/year

after yCH4r = 0.26. Consequently, the preliminary design analysis of a single stage membrane

separation indicates that maximum profitability will occur when both low energy gas and

carbon dioxide products can be sold from the retentate and permeate streams respectively.

6.2.2 Single Stage Membrane System with Compressor

A compressor was now utilized to pressurize the feed to drive mass transfer of carbon dioxide

from the retentate side to the permeate side of the membrane as shown in Figure 6.3. To

achieve maximum profit, the parameter of yCO2p was fixed at 0.995 and feed pressure was

increased. The maximum pressure that the membrane can experience before rupturing was

assumed to be 1000 psia in this analysis.15 The pressure of the feed was increased from 434.4

psia to 984.4 psia. This pressure range yielded retentate products with mole fractions of

yCH4r=0.264 to 0.269 which can only be sold as low energy gas. Additionally, the permeate

molar flow rate also increases as compression of the feed occurs.16

Figure 6.4 Single Stage System with Compressor

As pressure was increased, the total capital cost for the membrane decreased while the

operation cost to compress increased. This trend is captured in Figure 6.5 with total costs

ranging from $1.74 M per year to $1.79 M per year.

15 Appendix: Figure 9.6.316 Appendix: Figure 9.6.4-9.6.5

Vf Vr

VpE-1


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Page | 21

Figure 6.5 Single Stage System with Compressor Separation Costs

Although the minimum total cost occurs at 584.4 psia, the maximum profit of $23.8 M occurs at

984.4 psia. The maximum revenue from the carbon dioxide product and low energy gas

outweigh the total separation cost at this pressure. The total revenue, total cost of separation,

and overall profit is shown in Figure 6.6.

Figure 6.6 Single Stage System with Compressor Costs, Revenue, and Profit

In this particular case, the total cost of separation does not increase significantly as pressure

increases whereas there is noticeable increase in revenue from low energy gas and carbon

$-

$200,000.00

$400,000.00

$600,000.00

$800,000.00

$1,000,000.00

$1,200,000.00

$1,400,000.00

$1,600,000.00

$1,800,000.00

$2,000,000.00

300 500 700 900 1100

$ p e r y e a r

Pressure (psia)

Capital Cost

Operation Costs

Total Cost

$-

$5,000,000.00

$10,000,000.00

$15,000,000.00

$20,000,000.00

$25,000,000.00

$30,000,000.00

300 500 700 900 1100

$ p e r y e a r

Pressure (psia)

Total Cost (Operation

and Capital)Revenue from CO2 and

CH4 ProductProfit


23/48

Page | 22

dioxide product. Total revenue from carbon dioxide and low energy gas is substantially greater

than separation costs over points of increased feed pressure. Accordingly, profit is dominated

by the increasing revenue from increasing feed pressure.

6.2.3

Two Stage Counter Current Membrane System with Compressors

Further optimization from the base case in this design analysis considered a two stage counter-current membrane system with two compressors as shown in Figure 6.7. The second permeate

stream was recycled back into the first stage’s retentate side. Recycling the carbon dioxide rich

permeate stream causes a higher concentration gradient for carbon dioxide transport across

the membrane. Two compressors were needed for this operation to occur where one

compressor pressurized the feed pressure from its initial pressure of 434.4 psia up to 984.4

psia. To obtain consistent pressures throughout the system, the second compressor was

employed to pressurize the second permeate stream from its estimated atmospheric pressure

of 15 psia to matching feed pressure for proper recycle.17 Consequently, the second retentate

increases in methane concentration from yCH4r=0.576 to 0.597. The quality of the second

retentate stream would be sold as medium energy gas at a price deck of $3 per MMBTU.18

Figure 6.7 Two Stage Counter-Current System with Compressors

Since there two membranes and two compressors were considered in this design optimization,

both the total capital and operating costs increased significantly as captured in Figure 6.8. In

this two stage separation system, the minimum total separation costs occurs when the feed

and permeate stream 2 are pressurized to 984.4 psia. Total separation costs ranged from $4.66

M per year to $5.43 M per year depending on the pressurization of the permeate and feed

streams.

17 Appendix: Figure 9.6.618 Appendix: Figure 9.6.7- Figure 9.6.8

Vf Vr1

Vp1

Vr2

Vp2

E-3


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Page | 23

Figure 6.8 Two Stage Counter-Current System with Compressors Separation Costs

In this particular case, the maximum profit of $41.3 M per year occurs when the total

separating costs are at a minimum depicted in Figure 6.9.

Figure 6.9 Two Stage Counter-Current System with Compressor Costs, Revenue, and Profit

$1,000,000.00

$1,500,000.00

$2,000,000.00

$2,500,000.00

$3,000,000.00

$3,500,000.00

$4,000,000.00

$4,500,000.00

$5,000,000.00

$5,500,000.00

$6,000,000.00

400 500 600 700 800 900 1000

$ / y e a r

Pressure (psia)

Total Capital Cost

Total Operating Cost

Total Cost

$1,000,000.00

$6,000,000.00

$11,000,000.00

$16,000,000.00

$21,000,000.00

$26,000,000.00

$31,000,000.00

$36,000,000.00

$41,000,000.00

$46,000,000.00

$51,000,000.00

400 500 600 700 800 900 1000

$ / y e a r

Pressure (psia)

Profit

Total Cost

Total Revenue


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Page | 24

In the two stage counter-current design, the total revenue does not increase significantly in

comparison to the observable decrease in total separating costs. As result, profit is dominated

by the decreasing cost to separate with increasing feed pressure.

6.2.4

Three Stage Counter Current Membrane System with Compressors

A three stage counter-current membrane design system with compressors was assessed todetermine any significant increase in profits from previous designs as depicted in Figure 6.10.

The second and third carbon dioxide rich permeate streams were fed back into the retentate

side of the membrane in stages one and two respectively. In addition, there were a total of

three compressors used in this design to pressurize the feed from its initial pressure of 434.4

psia to 984.4 psia.19 The second permeate stream and third permeate streams were

pressurized from atmospheric pressure of 15 psia to matching feed pressure for proper recycle.

This arrangement also drives carbon dioxide mass transport by increasing the concentration

gradient on stages 1 and 2. It was calculated that the methane mole fraction in the third

retentate stream ranged from yCH4r =0.818 to 0.843.20 This stream can then be sold as high

energy gas at a price deck of $4 per MMBTU.

Figure 6.10 Three Stage Counter-Current System with Compressors

Three stages include three membranes and three compressors which result in an increase in

capital and operating costs as reflected in Figure 6.11. Depending on the pressurization of the

feed and permeate streams, total costs ranged from $7.76 M per year to $11.8 M per year.

19 Appendix: Figure 9.6.920 Appendix: Figure 9.6.10- Figure 9.6.11

Vf Vr1

Vp3

Vr2

E-4

Vr3

Vp2

Vp1


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Page | 25

Figure 6.11 Three Stage Counter-Current System with Compressors Separation Costs

As compressors were used to pressurize the streams, the overall capital cost for area decreases

rapidly while the total operating costs slowly increased. In a three stage counter-current

system, the maximum profit $44.5 M per year also occurs when the total separating costs are at

a minimum shown in Figure 6.12.

Figure 6.12 Three Stage Counter-Current with Compressor Costs, Revenue, and Profit

In the three stage counter-current design, the total revenue does not increase significantly in

comparison to the observable decrease in total separating costs. Similar to the two stage

$1,000,000.00

$3,000,000.00

$5,000,000.00

$7,000,000.00

$9,000,000.00

$11,000,000.00

$13,000,000.00

400 600 800 1000

$ / y e a r

Pressure (psia)

Total Capital Cost

Total Operating Cost

Total Cost

$-

$10,000,000.00

$20,000,000.00

$30,000,000.00

$40,000,000.00

$50,000,000.00

$60,000,000.00

400 500 600 700 800 900 1000

$ / y e a r

Pressure (psia)

Profit

Total Cost

Total Revenue


27/48

Page | 26

counter-current system, profit is dominated by the separation costs associated with increasing

feed pressure.

6.3 SUMMARY

Based on the maximum yearly profits of the different designs, it was found that the three stagecounter-current membrane system was the most economical as shown in Figure 6.13.

Figure 6.13 Maximum Profit of Designs compared to Value of Feed

The three stage system was found to have the maximum yearly profit of $44.5 M which is much

higher than the feed’s value of $8.2 per year. This analysis reflects that multiple stage

membrane systems are effective and profitable in separating natural gas streams with high cut

carbon dioxide. Operating flexibility allows for an economical separation as both methane and

carbon dioxide products can be sold. It should be noted that selling carbon dioxide product

with low-high energy gas dominated revenue while pressurizing feeds substantially decreased

capital costs. Moving forward, suggestions to continue optimizing these systems include placing

compressors between stages to reduce membrane area and thus decrease capital cost.

$-

$5

$10

$15

$20

$25

$30

$35

$40

$45

$50

$ M / y e a r


28/48

Page | 27

7 CRYOGENIC DISTILLATION

Conventionally, carbon dioxide is removed from natural gas streams via absorption or

membrane permeation. However, if the feed concentration of carbon dioxide is high enough,

these conventional methods can become less economically viable.21 Alternatively, cryogenic

distillation can be used to separate gases. Distillation is a process that separates liquids based

on their boiling points. Because methane and carbon dioxide are both gases at standard

conditions, they must be brought cryogenic (very low) temperatures. Naturally, operating at

such conditions can be quite energy intensive, and therefore quite expensive.

7.1

C

OST

C

ONSIDERATIONS

The utility of cryogenic distillation faces two major cost constraints: capital costs, which include

the column and operating costs, which are dominated by the cost of refrigeration.

7.1.1 Capital Costs

Capital costs for the distillation column are estimated as $420,000 per distillation plate per

year. These costs include maintenance and depreciation on the column.

7.1.2 Refrigeration Costs

Refrigeration costs dominate the operating cost of the column. Both the heat exchanger used

to cool the feed before it enters the column and the reflux condenser require high amounts of

electricity. The cost of refrigeration depends on the temperature and the mass flowrate of the

fluid being cooled. The costs of refrigeration can be approximated graphically from Figure 7.1.

21 Maqsood, K.; Mullick, A.; Ali, A.; Kargupta, K.; Ganguly, S. Rev Chem Eng. 2014


29/48


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Page | 29

7.2.1 Assumptions of the Model

The model created in Aspen HYSIS comprised of a four primary assumptions.

1. The reboiler requires zero energy . Liquid at the bottom of the column can simply be re-

vaporized with an ambient air vaporizer.

2.

The column exhibits minimal pressure drop. The pressure drop through the column andcondenser was taken as 2.00 psi and the pressure drop through the reboiler was

assumed to be zero.

3.

The feed is “pre-compressed.” For consistency, in cases where the feed was compressed

above the feed original pressure, the cost of compression was calculated in the same

manner as in the separation methods above (rather than modeled with a compressor in

HYSIS).

4. The Lee-Kesler Method accurately describes the state of the system. Generally this is a

valid assumption. The SRK Method was also tested and it yielded similar results. For

more a more extensive analysis, other equations of state should be also be considered.

7.2.2 Choosing the Condenser

The separation was first simulated to achieve the minimum purity for low energy gas (40%

methane) in the distillate (stream exiting the top of the column) while also obtaining sellable

CO2 (99.5%) in the bottoms (stream exiting the bottom of the column). The simulation was

evaluated using both a total condenser and a full reflux condenser. In all cases, using a full

reflux condenser was more profitable. While a total condenser produces a liquid distillate

stream, a full reflux condenser produces a vapor distillate. Therefore, not having to liquefy the

distillate products, the full reflux condenser enjoys the advantage of considerably lower

refrigeration costs.

7.2.3 Condenser Inlet Phase

The feed was isobarically cooled at its original pressure (420 psig) to saturated vapor (7.372 °F)

and saturated liquid (-96.50 °F).22 In both cases the feed was distilled to the minimum purity of

each gas grade (40%, 70%, 95%, 99.995%). Because cooling the feed to saturated vapor

required significantly less refrigeration, it proved to be more profitable.

7.3 OPTIMIZATION

Profit is optimized when the margin between revenue and costs is at a maximum. The revenue

and costs for this model depend on the following parameters:

Inlet temperature

Inlet pressure

Reflux Rate

Purity

22 Appendix : Figure 9.7.1


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Page | 30

Number of stages

By manipulating these variables to understand their trends, we can arrive at optimum

operating conditions.

7.3.1 Purity Determination

The distillation column was evaluated by taking saturated vapor to the minimum purity for each

gas grade respectively. Evaluating the profitability of compositions between the purity grade

minimums is futile because as the distillate purity increases, the distillate flowrate decreases.

Because the value of the gas remains the same until it reaches the next purity threshold, the

revenue simply decreases. In other words, maximum revenue for each gas grade is achieved at

the minimum purity.

This process was repeated while increasing the number of distillation trays (N) so that a local

maximum could be observed for each gas grade. Through this process, high energy grade

methane (70%) proved to be the most profitable option for cryogenic distillation.23

7.3.2 Effect of Number of Stages

Increasing the number of distillation trays N in the distillation column decreases the load on the

condenser thereby reducing refrigeration costs. However, increasing N also increases capital

costs ($420,000/tray/year). We can therefore assume that there is some value Nopt that

optimizes these two constraints. Figure 7.3 shows the effect of the number of trays on profit. In

this case, Nopt corresponds to four trays.24

Figure 7.3. Profit in million dollars per year verses N for high energy gas at 434.7 psia

23 Appendix. Table 9.7.124 Appendix. Table 9.7.2

$26.6

$35.4

$39.3 $39.4 $39.0 $38.8

0

5

10

15

20

25

30

35

40

45

1 2 3 4 5 6

P r o f i t ( M M $ / y r )

N


32/48

Page | 31

7.3.3 Effect of Entrance Pressure

Another important parameter to be optimized is the pressure of the stream entering the

distillation column. At higher pressures, the lower temperatures for distillation are no longer

required. Consequently, as the entrance pressure is increased, the cost of refrigeration

decreases. However, the increasing pressure requires work from a compressor and therefore

higher pressures mean higher costs. The effect of entrance pressure P is illustrated in Figure 7.4

below.

(NOTE: Because in our case, the stream is entering at saturated vapor, the temperature (Tsat)

determined by the pressure and therefore not an independent parameter.)

Figure 7.4. Refrigeration and compression costs versus entrance pressure

In this case, with all other parameters constant, revenue is independent of P. Consequently,

profit is maximized where total cost is minimized. Here the total cost is only a function of

refrigeration and compression costs as capital costs are constant at constant N. Figure 7.5

illustrates the effect of P on profits. In this case, costs are minimized by operating at 550 psia.

0.0

1.0

2.0

3.0

4.0

5.0

6.0

7.0

8.0

9.0

10.0

400 450 500 550 600 650 700

C o s t s ( M M $ / y r )

P (psia)

Compression

Refrigeration


33/48

Page | 32

Figure 7.5. Profit in million dollars per year versus column entrance pressure

7.3.4 Revenue Analysis

It is important to note that the majority of the revenue generated from this separation is not

from methane, but in fact from carbon dioxide as seen in Figure 7.6. Considering the fact that

the feed stream contained 82% carbon dioxide, these results are less surprising.

Figure 7.6. Revenue composition for optimized product

39.0

39.2

39.4

39.6

39.8

40.0

40.2

40.4

40.6

40.8

41.0

400 450 500 550 600 650 700

P r o f i t ( M M $ / y r )

P (psia)

Methane

23%

Carbon

dioxide

77%


34/48

Page | 33

7.4

S

UMMARY

We have shown that cryogenic distillation holds promise as a profitable means of separation.

High energy methane proved to be the most profitable grade of natural gas to produce with

this method.

The optimal operating conditions were determined to be:

P = 550 psia

T = 96.50 °F

N = 4 trays

Reflux Ratio = 3.518

At these conditions, it is possible to generate an annual profit of $40.1 million from our feed

stream using cryogenic distillation.


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Page | 34

8 RECOMMENDATIONS

Of the four separation methods evaluated in this report, membrane separation was proven tobe the most profitable venture as shown in Table 8.1.

Table 8.1 Summary comparison of separation methods

Separation

Method

Optimal

Grade

High Purity

CO2 Annual Profit

Absorption

(Water)

High

EnergyNo $ 13.8 million

Absorption

(MDEA)

High

EnergyNo $ 5.1 million

Membrane

Separation

High

EnergyYes $ 44.5 million

Cryogenic

Distillation

High

EnergyYes $ 40.1 million

Figure 8.1 depicts the comparison between the maximum profits of each unit operation with

the original feed value.

Figure 8.1 Comparison between Feed Value and Maximum Profits of Unit Operations

$8.3 $8.3 $8.3 $8.3

$5.1

$13.8

$40.1

$44.5

$-

$5.0

$10.0

$15.0

$20.0

$25.0

$30.0

$35.0

$40.0

$45.0

$50.0

MDEA Absorption Water Absoprtion Cryogenic

Distillation

Gas Membrane

Separation

$ M / y e a r

Feed Value Max Profit of Unit Operation


36/48


37/48

Page | 36

9 APPENDIX

Table 9.4.1. Initial economic evaluation of absorption


38/48

Page | 37

Table 9.6.1 Single Stage System Calculated Values

ych4r yco2p

VR

mol/hr

VP

mol/hr

Are a

ft^2

0.2 0.99653 2,838,763.14 321,616.35 123,784.00 0.21 0.99631 2700822.171 459,557.32 179,174.18

0.22 0.99608 2575350.633 585,028.86 231,097.78

0.23 0.99585 2460719.327 699,660.17 280,068.06

0.24 0.99561 2355570.44 804,809.05 326,516.01

0.25 0.99537 2,258,763.26 901,616.23 370,807.00

0.26 0.99511 2169332.428 991,047.07 413,253.47

0.3 0.99403 1,870,564.94 1,289,814.55 569,502.83

0.35 0.99248 1,591,648.33 1,568,731.17 747,676.05

0.4 0.99067 1,380,655.11 1,779,724.39 921,416.81

0.4 0.99067 1,380,655.11 1,779,724.39 921,416.81

0.45 0.98851 1,214,478.83 1,945,900.66 1,102,547.31

0.5 0.98591 1,079,076.78 2,081,302.71 1,302,140.55

0.55 0.98270 965,253.15 2,195,126.34 1,533,020.13

0.6 0.97865 866,478.65 2,293,900.84 1,812,615.26

0.65 0.97337 777,577.68 2,382,801.82 2,167,594.22

0.7 0.96623 693,683.74 2,466,695.76 2,643,086.37

0.7 0.96623 693,683.74 2,466,695.76 2,643,086.37

0.75 0.95603 608,907.46 2,551,472.03 3323978.992

0.8 0.94029 513,519.88 2,646,859.61 4392644.66

0.85 0.91297 385,114.23 2,775,265.26 6323661.632

0.9 0.85518 147,244.43 3,013,135.06 10,811,966.53

0.95 0.67947 (705,584.33) 3,865,963.82 29,267,942.06


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Page | 38

Table 9.6.2 Single Stage System Revenue, Costs, and Profits

V a l u e o f

R e t e n t a t e

C H 4

p e r

y e a r

V a l u e o f P e r m e a t e

C O 2 p e r y e a r

T o t a l R e v e n u e

C o s t o f M e m b

r a n e

p e r y e a r

C o s t o f H e a t

E x c h a n g e r p e r

y e a r

P r o f i t

8 ,

0 7 7 ,

4 0 5 .

0 6

$

5 ,

2 6 6 ,

3 4 4 .

2 9

$

1 3 ,

3 4 3 ,

7 4 9 .

3 5

$

4 9 5 ,

1 3 6 .

0 0

$

1 3 ,

7 9 5 .

1 1

$

1 2

, 8 3 4 ,

8 1 8 .

2 4

$

8 ,

0 6 9 ,

1 5 3 .

8 1

$

7 ,

5 2 5 ,

0 7 4 .

6 8

$

1 5 ,

5 9 4 ,

2 2 8 .

4 8

$

7 1 6 ,

6 9 6 .

7 2

$

1 3 ,

7 9 5 .

1 1

$

1 4

, 8 6 3 ,

7 3 6 .

6 5

$

8 ,

0 6 0 ,

6 8 1 .

4 0

$

9 ,

5 7 9 ,

6 2 2 .

9 2

$

1 7 ,

6 4 0 ,

3 0 4 .

3 2

$

9 2 4 ,

3 9 1 .

1 0

$

1 3 ,

7 9 5 .

1 1

$

1 6

, 7 0 2 ,

1 1 8 .

1 2

$

8 ,

0 5 1 ,

9 7 8 .

8 4

$

1 1 ,

4 5 6 ,

6 6 6 .

5 4

$

1 9 ,

5 0 8 ,

6 4 5 .

3 8

$

1 ,

1 2 0 ,

2 7 2 .

2 4

$

1 3 ,

7 9 5 .

1 1

$

1 8

, 3 7 4 ,

5 7 8 .

0 3

$

8 ,

0 4 3 ,

0 3 6 .

6 3

$

1 3 ,

1 7 8 ,

4 3 9 .

1 7

$

2 1 ,

2 2 1 ,

4 7 5 .

7 9

$

1 ,

3 0 6 ,

0 6 4 .

0 6

$

1 3 ,

7 9 5 .

1 1

$

1 9

, 9 0 1 ,

6 1 6 .

6 3

$

8 ,

0 3 3 ,

8 4 4 .

7 2

$

1 4 ,

7 6 3 ,

6 1 9 .

5 0

$

2 2 ,

7 9 7 ,

4 6 4 .

2 2

$

1 ,

4 8 3 ,

2 2 8 .

0 0

$

1 3 ,

7 9 5 .

1 1

$

2 1

, 3 0 0 ,

4 4 1 .

1 2

$

8 ,

0 2 4 ,

3 9 2 .

5 3

$

1 6 ,

2 2 8 ,

0 1 5 .

1 0

$

2 4 ,

2 5 2 ,

4 0 7 .

6 3

$

1 ,

6 5 3 ,

0 1 3 .

8 7

$

1 3 ,

7 9 5 .

1 1

$

2 2

, 5 8 5 ,

5 9 8 .

6 6

$

7 ,

9 8 3 ,

7 4 6 .

7 9

$

-

$

7 ,

9 8 3 ,

7 4 6 .

7 9

$

2 ,

2 7 8 ,

0 1 1 .

3 2

$

1 3 ,

7 9 5 .

1 1

$

5

, 6 9 1 ,

9 4 0 .

3 6

$

7 ,

9 2 5 ,

5 2 1 .

9 0

$

-

$

7 ,

9 2 5 ,

5 2 1 .

9 0

$

2 ,

9 9 0 ,

7 0 4 .

1 9

$

1 3 ,

7 9 5 .

1 1

$

4

, 9 2 1 ,

0 2 2 .

6 0

$

7 ,

8 5 7 ,

0 2 0 .

8 1

$

-

$

7 ,

8 5 7 ,

0 2 0 .

8 1

$

3 ,

6 8 5 ,

6 6 7 .

2 4

$

1 3 ,

7 9 5 .

1 1

$

4

, 1 5 7 ,

5 5 8 .

4 7

$

9 ,

4 2 8 ,

4 2 4 .

9 8

$

-

$

9 ,

4 2 8 ,

4 2 4 .

9 8

$

3 ,

6 8 5 ,

6 6 7 .

2 4

$

1 3 ,

7 9 5 .

1 1

$

5

, 7 2 8 ,

9 6 2 .

6 4

$

9 ,

3 3 0 ,

3 1 7 .

4 0

$

-

$

9 ,

3 3 0 ,

3 1 7 .

4 0

$

4 ,

4 1 0 ,

1 8 9 .

2 6

$

1 3 ,

7 9 5 .

1 1

$

4

, 9 0 6 ,

3 3 3 .

0 4

$

9 ,

2 1 1 ,

2 0 2 .

0 3

$

-

$

9 ,

2 1 1 ,

2 0 2 .

0 3

$

5 ,

2 0 8 ,

5 6 2 .

2 0

$

1 3 ,

7 9 5 .

1 1

$

3

, 9 8 8 ,

8 4 4 .

7 2

$

9 ,

0 6 3 ,

5 4 0 .

3 6

$

-

$

9 ,

0 6 3 ,

5 4 0 .

3 6

$

6 ,

1 3 2 ,

0 8 0 .

5 2

$

1 3 ,

7 9 5 .

1 1

$

2

, 9 1 7 ,

6 6 4 .

7 4

$

8 ,

8 7 5 ,

7 0 9 .

3 3

$

-

$

8 ,

8 7 5 ,

7 0 9 .

3 3

$

7 ,

2 5 0 ,

4 6 1 .

0 5

$

1 3 ,

7 9 5 .

1 1

$

1

, 6 1 1 ,

4 5 3 .

1 8

$

8 ,

6 2 8 ,

8 1 3 .

7 8

$

-

$

8 ,

6 2 8 ,

8 1 3 .

7 8

$

8 ,

6 7 0 ,

3 7 6 .

8 7

$

1 3 ,

7 9 5 .

1 1

$

( 5 5 ,

3 5 8 .

2 0 )

$

8 ,

2 8 9 ,

9 8 0 .

4 9

$

-

$

8 ,

2 8 9 ,

9 8 0 .

4 9

$

1 0 ,

5 7 2 ,

3 4 5 .

5 0

$

1 3 ,

7 9 5 .

1 1

$

( 2

, 2 9 6 ,

1 6 0 .

1 2 )

$

9 ,

6 7 1 ,

6 4 3 .

9 0

$

-

$

9 ,

6 7 1 ,

6 4 3 .

9 0

$

1 0 ,

5 7 2 ,

3 4 5 .

5 0

$

1 3 ,

7 9 5 .

1 1

$

( 9 1 4 ,

4 9 6 .

7 0 )

$

9 ,

0 9 6 ,

0 5 9 .

7 0

$

-

$

9 ,

0 9 6 ,

0 5 9 .

7 0

$

1 3 ,

2 9 5 ,

9 1 5 .

9 7

$

1 3 ,

7 9 5 .

1 1

$

( 4

, 2 1 3 ,

6 5 1 .

3 7 )

$

8 ,

1 8 2 ,

5 3 7 .

3 1

$

-

$

8 ,

1 8 2 ,

5 3 7 .

3 1

$

1 7 ,

5 7 0 ,

5 7 8 .

6 4

$

1 3 ,

7 9 5 .

1 1

$

( 9

, 4 0 1 ,

8 3 6 .

4 4 )

$

6 ,

5 2 0 ,

0 2 4 .

6 2

$

-

$

6 ,

5 2 0 ,

0 2 4 .

6 2

$

2 5 ,

2 9 4 ,

6 4 6 .

5 3

$

1 3 ,

7 9 5 .

1 1

$

( 1 8

, 7 8 8 ,

4 1 7 .

0 1 )

$

2 ,

6 3 9 ,

5 0 2 .

8 1

$

-

$

2 ,

6 3 9 ,

5 0 2 .

8 1

$

4 3 ,

2 4 7 ,

8 6 6 .

1 1

$

1 3 ,

7 9 5 .

1 1

$

( 4 0

, 6 2 2 ,

1 5 8 .

4 1 )

$

( 1

1 ,

4 4 3 ,

7 0 0 .

6 0 )

$

-

$

( 1 1 ,

4 4 3 ,

7 0 0 .

6 0 )

$

1 1 7 ,

0 7 1 ,

7 6 8 .

2 5

$

1 3 ,

7 9 5 .

1 1

$

( 1 2 8

, 5 2 9 ,

2 6 3 .

9 6 )

$


40/48

Page | 39

Table 9.6.3 Single Stage System with Compressor Calculated Values

Pr

psia

Vr

mol/h

Vp

mol/h

A

ft^ 2 yCH4 yCO2p

434.4 2,131,735.48 1,028,644.01 431,597.43 0.264126061 0.995

484.4 2,123,894.66 1,036,484.84 388,563.54 0.265147047 0.995534.4 2,117,563.43 1,042,816.07 353,303.51 0.265971688 0.995

584.4 2,112,344.11 1,048,035.39 323,892.33 0.266651656 0.995

634.4 2,107,967.43 1,052,412.06 298,990.09 0.267221952 0.995

684.4 2,104,244.57 1,056,134.92 277,636.06 0.267707129 0.995

734.4 2,101,039.21 1,059,340.28 259,123.76 0.26812492 0.995

784.4 2,098,250.46 1,062,129.04 242,922.24 0.268488453 0.995

834.4 2,095,802.04 1,064,577.45 228,624.88 0.268807652 0.995

884.4 2,093,635.25 1,066,744.25 215,915.02 0.269090162 0.995

934.4 2,091,704.13 1,068,675.37 204,542.46 0.269341964 0.995

984.4 2,089,972.22 1,070,407.28 194,306.90 0.269567807 0.995

(Eq. 9.6.1)

(Eq. 9.6.2)

(Eq. 9.6.3)

(Eq. 9.6.4)

(Eq. 9.6.5)

(Eq. 9.6.6)


41/48

Page | 40

Table 9.6.4 Single Stage System with Compressor Revenue, Costs, and Profits

T o t a l C a p t i a l

C o s t

C o s t t o

c o m p r e

s s

p e r y e a

r

C o s t t o h e a t

f e e d p e r y e a r

T o t a l o p e r a t i o n

C o s t p e r y e a r

T o t a l C o s t p e r

y e a r

V a l u e o f C O 2

V a l u e o f C H 4

R e v e n u e

P r o

f i t

1 ,

7 2 6 ,

3 8 9 .

7 4

$

-

$

1 3 ,

7 9 5 .

1 1

$

1 3 ,

7 9 5 .

1 1

$

1 ,

7 4 0 ,

1 8 4 .

8 4

$

1 6 ,

8 4 3 ,

6 5 0 .

6 9

$

8 ,

0 1 0 ,

4 5 6 .

7 0

$

2 4 ,

8 5 4 ,

1 0 7 .

3 9

$

2

3 ,

1 1 3 ,

9 2 2 .

5 5

$

1 ,

5 5 4 ,

2 5 4 .

1 5

$

1 3 3

, 5 1 3 .

3 7

$

1 3 ,

7 9 5 .

1 1

$

1 4 7 ,

3 0 8 .

4 8

$

1 ,

7 0 1 ,

5 6 2 .

6 3

$

1 6 ,

9 7 2 ,

0 4 1 .

1 6

$

8 ,

0 1 1 ,

8 4 3 .

8 2

$

2 4 ,

9 8 3 ,

8 8 4 .

9 9

$

2

3 ,

2 8 2 ,

3 2 2 .

3 6

$

1 ,

4 1 3 ,

2 1 4 .

0 2

$

2 5 3

, 8 9 9 .

5 4

$

1 3 ,

7 9 5 .

1 1

$

2 6 7 ,

6 9 4 .

6 4

$

1 ,

6 8 0 ,

9 0 8 .

6 7

$

1 7 ,

0 7 5 ,

7 1 2 .

6 4

$

8 ,

0 1 2 ,

8 0 4 .

4 9

$

2 5 ,

0 8 8 ,

5 1 7 .

1 3

$

2

3 ,

4 0 7 ,

6 0 8 .

4 6

$

1 ,

2 9 5 ,

5 6 9 .

3 1

$

3 6 3

, 5 1 0 .

3 5

$

1 3 ,

7 9 5 .

1 1

$

3 7 7 ,

3 0 5 .

4 6

$

1 ,

6 7 2 ,

8 7 4 .

7 7

$

1 7 ,

1 6 1 ,

1 7 6 .

9 7

$

8 ,

0 1 3 ,

4 8 9 .

3 1

$

2 5 ,

1 7 4 ,

6 6 6 .

2 8

$

2

3 ,

5 0 1 ,

7 9 1 .

5 1

$

1 ,

1 9 5 ,

9 6 0 .

3 6

$

4 6 4

, 1 1 7 .

2 7

$

1 3 ,

7 9 5 .

1 1

$

4 7 7 ,

9 1 2 .

3 8

$

1 ,

6 7 3 ,

8 7 2 .

7 4

$

1 7 ,

2 3 2 ,

8 4 3 .

3 8

$

8 ,

0 1 3 ,

9 8 8 .

9 1

$

2 5 ,

2 4 6 ,

8 3 2 .

2 9

$

2

3 ,

5 7 2 ,

9 5 9 .

5 5

$

1 ,

1 1 0 ,

5 4 4 .

2 4

$

5 5 7

, 0 8 7 .

9 0

$

1 3 ,

7 9 5 .

1 1

$

5 7 0 ,

8 8 3 .

0 1

$

1 ,

6 8 1 ,

4 2 7 .

2 4

$

1 7 ,

2 9 3 ,

8 0 3 .

7 2

$

8 ,

0 1 4 ,

3 6 0 .

2 7

$

2 5 ,

3 0 8 ,

1 6 3 .

9 9

$

2

3 ,

6 2 6 ,

7 3 6 .

7 4

$

1 ,

0 3 6 ,

4 9 5 .

0 3

$

6 4 3

, 5 0 0 .

1 2

$

1 3 ,

7 9 5 .

1 1

$

6 5 7 ,

2 9 5 .

2 3

$

1 ,

6 9 3 ,

7 9 0 .

2 6

$

1 7 ,

3 4 6 ,

2 9 0 .

2 9

$

8 ,

0 1 4 ,

6 4 0 .

5 0

$

2 5 ,

3 6 0 ,

9 3 0 .

7 9

$

2

3 ,

6 6 7 ,

1 4 0 .

5 4

$

9 7 1 ,

6 8 8 .

9 6

$

7 2 4

, 2 1 8 .

5 8

$

1 3 ,

7 9 5 .

1 1

$

7 3 8 ,

0 1 3 .

6 9

$

1 ,

7 0 9 ,

7 0 2 .

6 5

$

1 7 ,

3 9 1 ,

9 5 5 .

1 0

$

8 ,

0 1 4 ,

8 5 4 .

5 8

$

2 5 ,

4 0 6 ,

8 0 9 .

6 9

$

2

3 ,

6 9 7 ,

1 0 7 .

0 3

$

9 1 4 ,

4 9 9 .

5 0

$

7 9 9

, 9 4 7 .

4 5

$

1 3 ,

7 9 5 .

1 1

$

8 1 3 ,

7 4 2 .

5 6

$

1 ,

7 2 8 ,

2 4 2 .

0 6

$

1 7 ,

4 3 2 ,

0 4 6 .

9 6

$

8 ,

0 1 5 ,

0 1 9 .

7 2

$

2 5 ,

4 4 7 ,

0 6 6 .

6 8

$

2

3 ,

7 1 8 ,

8 2 4 .

6 3

$

8 6 3 ,

separating methane and carbon dioxide

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