an economic cost evaluation of landfill gas[1]

8/3/2019 An Economic Cost Evaluation of Landfill Gas[1]

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An Economic Cost Evaluation of Landfill Gas

Kyle ZezenskiAbsorption with MDEAJeffrey ZugatesAbsorption with Water

Sagarika BadyalCryogenic Distillation

Anuj UpadhyayGas Separation Membranes

ChE 410Mass Transfer Operations

Dr. Michael Janik

December 09, 2011


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Table of Contents

Introduction...3

Executive Summary..4

Separation Processes

A. Absorption with MDEA5

B. Absorption with Water..10

C. Cryogenic Distillation....16

D. Gas Separation Membranes21

Overall Recommendation...24

Appendix/References..25


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Introduction

In todays world, people are progressively looking for alternative forms of energy, and one

interesting source is the use of landfill gas. This rise in the use of landfill gas can be attributed toa variety of factors:

High energy prices make landfill gas economical compared to other sources of renewableenergy

There is a demand from consumers for a greener form of energy.

Landfill gas is generated when microorganisms in the ground break down organic material in thelandfill, and this gas contains approximately fifty to sixty percent methane and forty to fiftypercent carbon dioxide. This gas is eventually released into the air. At certain times, it couldenter buildings, houses, even those that are not directly over buried waste. In rare cases, highlevels of landfill gas may cause a fire if a spark were present. The gas from the decomposition ofwaste in the landfill gas can be used as a source of fuel and can eventually be converted toelectricity.

For our project, we were given the following parametersLandfill gas with a carbon dioxide mole percent of 62, a feed flow rate of 10 MMSCFD, a feedpressure of 25 psig and an operating time of 2000 hr/yr.


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Executive Summary

Adsorption of CO2 from a CH4/CO2 mixture into an MDEA solution was found to be the

most profitable process for producing liquefied natural gas (LNG). This process resulted in a

methane revenue of 421,936,104/yr and a CO2 revenue of $46,315/yr. The operating costsrequired for this process include the cost for a resupply of MDEA each year which was

determined to be $2,085,000/yr and the compression costs to bring the LNG up to pressurized

standards (600psi) which was determined to be $865,000/yr. The capital investment required for

this process includes only the cost of the 1 Rashig rings to fill the adsorption column which was

calculated to be $733,000. Therefore, the total revenue for this process is $19,032,153/yr and

will pay for its initial investment in about 27 days of time.

Absorption using water to produce a medium energy gas yielded the most profitable

separation process with respect to investment costs. The revenue associated with this product is

$1.1 million per year. The absorption tower requires $221,648 each year for operation and $8.3million in packing material. An income of $893,305 could be acquired each year. The

separation process requires a tower height of approximately 15 m and a pure water flow rate of

64 gallons/second. Pressurizing the gas feed to 10 atm provided the best separation. Increasing

the feed pressure beyond this value would not be feasible for this process. Absorption towers

were designed for high energy, pipeline quality, and liquefied natural gas products. These

separation units were not possible due to the high water flow rates and excessive tower heights

needed. The costs associated with operating these towers were much greater than the costs

required to produce a medium energy gas.

The distillation process used 2 trays to produce a fractional concentration of 0.98 ofmethane. An XY equilibrium plot of Methane depicts this is a one stage easy separation process.

This pipeline quality gas can be produced using a two stage process and generating a profit of

$964,129.5 while taking operating costs into considerations. Although, selling the initial

unprocessed low energy gas would be more profitable since it does not take the operating costs

into the consideration while generating high revenue. The distillate product flow is

427.3lbmol/hr and methane fraction of 0.98 is produced.

The membrane separation for the landfill gas stream ran on an annual loss of $204,363.00

per year. This loss is due to the fact that the membrane area was 234,683.08 ft2 which greatly

increased the cost. The total cost of production was the cost of the membrane and was $938,732.34. The permeate stream did not have a concentration of methane that was high enough

to be sold commercially so the revenue from that stream was $0.The retentate stream as sold as a

Medium Energy gas at $3.3/MMBTU and the revenue was $734,369.34.


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Adsorption with MDEA Solutions

Overview

The use of methyldiethanolamine (MDEA) solutions in the process industries for

absorbing gasses such as carbon dioxide from bulk gas streams is a very popular alternative toother processes such as absoption with water, distillation, and gas separation membranes for

many reasons. MDEA solutions has a very high solubility for carbon dioxide, are less prone to

evaporative losses, and are resistant to thermal and chemical degradation.

These traits make MDEA an ideal choice when choosing an adsorbent for carbon dioxide

and allowing the process to bring the product gas to be brought to very high concentrations.

Bringing the product gas to a high purity is fundamental in creating the most valuable product

from your process and therefore producing a high profit.

Mass Transfer Model

In this process the inlet gas feed flow is specified as 10MMSCFD with a carbon dioxide

mole fraction of 0.62 and a methane mole fraction of 0.38. The MDEA liquid feed flow was not

specified, however, four assumptions are made regarding the MDEA entering the column:

The MDEA feed contains no carbon dioxide.

Methane does not readily adsorb into the MDEA solution.

The MDEA solution is 23% wt in water (2M)

Regeneration of the MDEA solution costs no money and recovers 100% of the CO2 from

the leaving MDEA stream.

Equilibrium data was found at a temperature of 313K (1) and a model assuming a concentrated

gas mixture was used.

There were originally two choices as to the adsorption column design. A counter current

design and a concurrent design were both plausible options. However, the counter-current gives

the maximum concentration gradient and in turn provides a higher separation in a shorter column

length. The following equations 1 and 2 provide the equations for the operating line of the

adsorption column

(

) (

) (

) (

) (A-1)L2*x2 + V1*x1 = L1*x1 + V2*y2 (A-2)

Equations 1 and 2 provide a method for calculating the CO2 mole fraction in the exit gas

stream, y2, when the other variables are either known are assumed to be some value. One

assumption that eases with the optimization of this process is the calculation of Lmin. Lmin is the

minimum amount of MDEA liquid flow that will produce a column that has an x1 in equilibrium


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with y1. According to the equilibrium data for MDEA at 313K, x1,eq was equal to 0.4617 when it

was in equilibrium with y1of 0.62. Using this equilibrium data a function of Lmin was created

that was dependent upon y2. This equation is listed as equation 3.

( )

(A-3)

L values could then be estimated using equation 3 and in turn all of the variables listed in

equations 1 and 2 became dependent on only one variable, y2.

Once the gas/liquid feed/exit streams were specified a long with their corresponding mole

fractions of CO2 it was possible to begin the iterative calculation of the adsorption column

volume. Using the equilibrium data and the following equations one can determine the final

volume of the column. First equation 4 is used to determine the slope of the line that will pass

through a given x,y point on the operating line. This allows for the interface compositions, xi

and yi, to be calculated then using equations 5 and 6. Finally the volume of the column is

calculated by integrating equation 7 from the starting composition y1 to the final composition y2.

A more detailed listing of all the variables can be viewed in table M-1.

(A-4)

(A-5)

(A-6)

(A-7)

(A-8)

Table M-1 - Variables

L' (mol/s) - MDEA solution Liquid Flow RateV' (mol/s) 55.60 Methane Gas Flow Rate

V (mol/s) - Total Gas Flow Rate

x - Operating Line Liquid Mole Fraction CO2

x1 - Exit Liquid Mole Fraction CO2

x2 0 Feed Liquid Mole Fraction CO2

xi - Interface Liquid Mole Fraction CO2


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Y - Operating Line Gas Mole Fraction CO2

y1 0.62 Feed Gas Mole Fraction CO2

y2 - Exit Gas Mole Fraction CO2

yi - Interface Gas Mole Fraction CO2k'x (mol/m2

s) 0.86 Liquid Film Mass Transfer Coefficientk'y (mol/m2s) 0.56 Gas Film Mass Transfer CoefficientK'y (mol/m2s) 0.412 Total Film Mass Transfer Coefficient

a (m2/m3) 206.7 Interfacial Area between Phases

S (m2) - Tower Cross Sectional Area

Z (m) - Tower Height

vol (m3) - Tower Volume

Economic Assumptions and Conditions:

Costs include:

o The post-tower pump power for the exiting gas stream. The power calculation

and pricing information can be found in the appendix. The post-tower pump is

only used to pressurize the 98% methane and the 99.9995% methane from the

initial 39.7psia to 600psia.

o Yearly replacement of the MDEA solution at $1.6/lb for a 1day stock that is

continuously used for one year.

o The initial purchase of the Rashig rings at $2,200/ft3 of tower volume.

Profits

o The price of methane is varying according to its purity. The prices can be

reviewed in the appendix.

o The price of CO2 is constant and can be reviewed in the appendix.

Economic Maximizations

The adsorption process economic efficiency was maximized by modifying the desired

product composition and the liquid flow rate in the column while adhering to the previously

stated assumptions.

As the desired product composition was increased, the compression costs of the exit gas

increased, the amount of MDEA required increased, and the column volume increased. These

three factors surely took away from the total profit but the reduction was small in comparison to

the large gain in profit from purifying the methane as much as possible. The increase in profit

being directly proportional to the increase in methane purity can be seen in figure 1. It should be

noted that for figure 1, all of the profits are based on a liquid MDEA flow rate equal to 1.5*Lmin.


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Figure M-1. Profit vs. Methane Purity

As can be seen in figure 1, the maximum profit occurs when the CH4/CO2 stream ispurified to 99.9995% CH4. This is a clear indication that in order to optimize the profit of this

process, the liquid MDEA flow must be altered at the point of 99.9995% CH4 product

composition. According to Transport Processes and Separation Process Principles by

Geankoplis, the most economically efficient adsorption column is produced when the liquid

MDEA flow is 1.2-1.6x greater than the minimum liquid MDEA flow required for x1 to be in

equilibrium with y1. The correlation between the profit and the liquid MDEA flow can be

viewed in figure 2.

Figure M-2. Profit vs. Liquid MDEA Flow Rates

It should be noted that the profit is inversely proportional with the liquid MDEA flow rate. This

is because as you increase the MDEA flow rate it will cost more money to supply the increased

amount of MDEA. However, the liquid MDEA flow rate is also inversely proportional to the

tower volume. As the MDEA flow rate increases, the tower volume decreases. It just so

happens that the tower volume decreases by such a small amount in comparison to the increase

in the MDEA flow rate that it is more economically efficient to use the minimum operating

8000000

10000000

12000000

14000000

16000000

18000000

20000000

38 48 58 68 78 88 98

Profit($/yr

)

Methane Purity (%)

Profit vs. Methane Purity

18200000

18400000

18600000

18800000

19000000

19200000

1.2 1.3 1.4 1.5 1.6

Profit($/yr)

L'=x*L'min

Profit vs Liquid MDEA flow


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liquid MDEA flow rate of 1.2*Lmin. It should be noted that Lmin was determined to be

105.46mol/s when y2 was equal to 0.0005. This corresponds to a liquid MDEA flow rate of

126.55mol/s. You could mathematically go below 1.2*Lmin and keep increasing your profit but

this is not a realistic solution because as you approach Lmin the concentration gradients become

less and less and the tower height becomes larger and larger.

Recommendations

Comparing the initial cost of this process to the profit of the process shows that this is a highly

valuable process. With an initial cost of a little over $700,000, this process could begin making

$19,000,000/yr with an operating time of only 2000 hours/yr. This is a very low risk process

because of the fact that the profit is so much higher than the initial cost.

Table M-2 - 99.9995% purity CH4, 1.2*L'min

V1 (MMSCFD) 10

L' (mol/s) 126.5MDEA ($/yr) 2085278

V' (mol/s) 55.6CH4 ($/yr) 21936104

Volume (ft3) 333.2506Rashig Rings ($) 733151V1 - V2 (mol/s) 90.6

CO2 ($/yr) 46315Feed Pressure (psia) 39.7

Compression Pressure (psia) 600Compression Costs ($/yr) 864988

Profit ($/yr) 19032153Initial Cost ($) 733151

Future Improvements

Due to the assumptions made during the calculations of this process, the profitability

could vary greatly in a declining manner. It is almost certain that some methane will adsorb into

the MDEA solution which will mean a loss in profit. Also, the MDEA regeneration will never

actually be free. Some cost will always be associated with the regeneration of anything and that

will reduce the profitability also. Lastly, the column may not actually be at 313K so there will be

a deviation from the equilibrium data used in the calculations which corresponded to 313K.

Experimentation should be conducted to determine the equilibrium data at a temperature closerto the columns operating temperature. Another improvement would be to increase the columns

operating time/yr. Increasing the amount of hours the column operates will always increase the

profitability of it but may take away from time to perform important maintenance work. In the

end, the adsorption of CO2 from methane into an MDEA solution is a viable separation process

with a high profitability associated with it.


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Absorption using Water as the Absorbent

Overview

The separation of a carbon dioxide/methane stream generated by a landfill was assessedusing a packed absorption tower. The feed stream contained a 62 mole percent carbon dioxidecomposition, a pressure of 25 psig (2.701 atm), and a flow of 10 MMSCFD. The operating timeof the absorption tower was 2000 hr/yr. An analysis using these feed conditions with changes intemperature and pressure was conducted to determine the optimum absorption tower. In additionto altering the feed stream properties, the methane concentration of the gas product was varied todetermine which product concentration yielded the greatest profit. A series of calculations wereevaluated, and the results were translated into revenue, capital, and operating costs.

Engineering Analysis

The equations used to design the packed absorption column were based on certainassumptions. Even though the inlet gas stream was concentrated, it was assumed to be dilute in

order to use Henrys law (Equation B-1).

(B-1)

The Henrys law constants (H) of 1860 and 2330 atm/mol frac were used when evaluating theeffects of raising the feed stream temperature from 30C to 40C. The pressure (P) in Henryslaw was varied from 2.701-10atm to determine the advantages or disadvantages of compressingthe CO2/CH4 inlet stream.

The overall mass transfer coefficient, , was determined from a series of calculations.The Cussler mass transfer correlations and Equation B-2 taken from Geankoplis were used to

obtain an approximate overall mass transfer coefficient of 0.005 mol/m

2

s. It was assumed thisvalue remained constant in all of the calculations. However, in actuality the value will vary atdifferent positions in the column and fluctuations will occur in response to temperature, pressure,and liquid velocity changes.

(B-2)

The packing material used in the tower was one inch ceramic Raschig rings with a surface area

per unit volume of 190 m2/m3. The pressure drop at flooding () was considered to be2.00 in. of water/ft height of packing since the packing factor was greater than 60. The towercross sectional area was set to 7m2 for each proposed column design. The overall mass balance

around the absorption column was evaluated using the expression in Equation B-3 (taken fromGeankoplis).

( ) ( )

( ) ( ) (B-3)

It was possible to calculate the process water flow rate from Equation B-3 using the minimum liquidflow, , where . The optimum liquid flow rate was taken as 1.2 times the minimumliquid flow. The optimum liquid flow rate of pure water was used in the calculation of process water cost.


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Results

The results of the four processes were analyzed using an Excel spreadsheet. The costcalculations associated with product purity at varying pressure and temperature conditions areprovided for each scenario. The absorption tower that yielded the highest ROI is highlighted inyellow. A summary of the specific design and cost aspects of the best column were gatheredfrom Excel and presented in Tables B-4 and B-5.

Scenario 1

The first scenario considered was designing an absorption tower that could generate aliquefied natural gas product. The target methane concentration in the outlet gas stream was99.995%. The cost data associated with this scenario is presented in Table B-1.

Table B-1: Cost Analysis of Liquefied Natural Gas Product

The column needed to conduct the separation was outside the realm of possibility. The possiblerevenue that could be generated for the outlet stream was $2.2 million per year. This was thehighest revenue value of all the scenarios, but the income and ROI are negative. This operationrequires a large tower height (1321 m) and a realistic target is in the range of 9-21 meters. The

large column height requires an unrealistic amount of packing material which results in anegative ROI. In addition to the high cost of Raschig rings, an electric cost for compressing theproduct to 600 psi adds to the operating costs. This only contributes to a diminishing income. Itwas evident from the design and economic data that this process would not be feasible due to itsimmense size and cost.

Scenario 2

The second scenario was aimed at producing a pipeline quality gas containing a 98%methane concentration. An additional compression cost for the gas product was considered in

the electric cost analysis. A summary of this data is shown in Table B-2.

CH4 Conc (mole fraction) P (atm) T (K) Raschig Rings ($) Process Water ($/yr) Electric ($/yr) Heating ($/yr) Revenue ($/yr) Income ($) ROI (%

0.99995 2.701 303.2 718797625.8 859825.068 83690.67 0 2196121.2 -7.18E+08 -99.

0.99995 5 303.2 718819081 464121.24 114651.7 0 2196121.2 -7.17E+08 -99

0.99995 7.5 303.2 718842463.2 309155.88 135037.1 0 2196121.2 -7.17E+08 -99.

0.99995 10 303.2 718792547.7 231673.2 149500.7 0 2196121.2 -7.17E+08 -99.0.99995 2.701 313.2 718792547.7 1077288.492 86450.92 1530.1047 2196121.2 -7.18E+08 -99.

0.99995 5 313.2 718809662.9 581594.976 118433.1 1530.1047 2196121.2 -7.17E+08 -99.

0.99995 7.5 313.2 718828307 387471.708 139490.8 1530.1047 2196121.2 -7.17E+08 -99.

0.99995 10 313.2 718846985.2 290410.08 154431.5 1530.1047 2196121.2 -7.17E+08 -99.


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Table B-2: Cost Analysis of Pipeline Quality Gas

The best option for this scenario is to compress the gas to 10 atm and enter the feed into thecolumn at ambient temperature. However, the cost of the Raschig rings inflates the capital costand yields a low ROI percentage. A desirable income of $1.4 million is possible for thisseparation, but it is not feasible to acquire due to high investment costs. The tower heightrequired to operate this separation unit is 72 meters which is above the desired interval.

Scenario 3

The third scenario was designed to release a high energy gas product. The desired outletconcentration of methane was considered to be 80%. The cost analysis for this separationprocess is shown in Table B-3.

Table B-3: Cost Analysis of High Energy Gas

The ROI obtained from this process is approximately 6% which is still not an ideal investment.However, decreasing the methane purity in the exiting gas stream lowers the overall operatingand capital costs. The column height of the absorption tower was 36 meters, and the Raschigring cost was around $19 million. The capital cost is still too high relative to the income. It wasobserved from the first three scenarios that this process would need to be constructed to purify a

medium energy product.

CH4 Conc (mole fraction) P (atm) T (K) Raschig Rings ($) Process Water ($/yr) Electric ($/yr) Heating ($/yr) Revenue ($/yr) Income ($) ROI (%

0.98 2.701 303.2 39279139.91 849096.204 85394.37 0 1824469.9 889979.3 2.265

0.98 5 303.2 39294749.99 458329.944 115969.1 0 1824469.9 1250171 3.181

0.98 7.5 303.2 39311760.17 305298.252 136100.1 0 1824469.9 1383072 3.518

0.98 10 303.2 39328807.28 228782.4 150383.3 0 1824469.9 1445304 3.674

0.98 2.701 313.2 39275445.05 1063846.116 88210.81 1530.1047 1824469.9 670882.9 1.708

0.98 5 313.2 39287897.94 574337.856 119793.9 1530.1047 1824469.9 1128808 2.87

0.98 7.5 313.2 39301462.05 382636.848 140588.9 1530.1047 1824469.9 1299714 3.307

0.98 10 313.2 39315049.65 286786.344 155343.1 1530.1047 1824469.9 1380810 3.512

CH4 Conc (mole fraction P (atm) T (K) Raschig Rings ($) Process Water ($/yr) Electric ($/yr) Heating ($/yr) Revenue ($/yr) Income ($) ROI (%

0.8 2.701 303.2 19901063.15 728100 0 0 1486605.1 758505.1 3.8

0.8 5 303.2 19908712.59 393017.928 49938.71 0 1486605.1 1043648.5 5.2

0.8 7.5 303.2 19917046.48 261793.248 82819.33 0 1486605.1 1141992.5 5.73

0.8 10 303.2 19925396.74 196180.908 106148.5 0 1486605.1 1184275.7 5.94

0.8 2.701 313.2 19899252.35 912248.052 0 1530.1047 1486605.1 572826.94 2.870.8 5 313.2 19905355.04 492494.712 51585.77 1530.1047 1486605.1 940994.51 4.72

0.8 7.5 313.2 19912001.27 328111.104 85550.84 1530.1047 1486605.1 1071413.1 5.3

0.8 10 313.2 19918657.91 245919.288 109649.4 1530.1047 1486605.1 1129506.3 5.67


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Henrys law constant,H, which will in turn influence the value of the mass transfer coefficient

However, the procedure used to obtain was very detailed and fallible with many iterations.Therefore, the value was kept constant for all of the calculations.

Another aspect of the design that was analyzed involved the tradeoff between a largetower height and a low water flow rate and vice versa. Figure B-1 displays the relationshipbetween tower height and water flow rate for scenarios 2, 3, and 4. The figure clearly shows thewater flow rate increasing with increasing height. It was expected that the water flow rate woulddecrease as height increases and the opposite case would be observed. Upon further analysis, thepressure and temperature values varied with each specified methane outlet concentration.Competing pressure and temperature effects resulted in the correlation shown in Figure B-1.Therefore, the column with the shortest height and lowest water flow rate proved to be the mostcost effective process. The shorter tower requires less packing and less process water to achievethe desired separation. Therefore, the capital and process water operating costs decrease, whichyields the highest ROI.

Figure B-1: Tower Height vs. Water Flow Rate

Conclusion/Recommendation

An absorption tower that can concentrate the 62 mole percent carbon dioxide/methanestream to a medium energy gas product was the ideal design. This conclusion was based on theeconomic evaluation of the data. The focus of the economic analysis was in the operating costssince these costs were the most accurate relative to capital costs. The only specified capital cost

were Raschig rings. A further analysis of all the capital costs would have to be conducted toobtain more accurate ROI values. Such considerations would be concentrated in the constructionaspects. These types of aspects include civil, mechanical, and electrical areas. The siteexcavation, concrete, structural steel, and roofing contribute to civil costs. Auxiliary equipmentsuch as pumps, compressors, filters for the inlet streams prior to entering the column, piping, andvalves are mechanical costs. The electrical costs associated with operating the column are motorcontrols, motor starters, transformers, field instrumentation, overall process control system, andlighting.

0

20

40

60

80

15000 20000 25000 30000 35000

TowerHeight(m)

Water Flow Rate (mol/s)


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Cryogenic Distillation

Overview

Normally, gaseous mixtures are separated by cryogenic distillation where the feed is allowed toenter the column at low temperatures. The distillation of carbon dioxide and methane is a known

and used process due to the large difference in the boiling points of the two components. The aim

of cryogenic distillation is to produce high-purity steams. On a large scale, gaseous separations

are accomplished economically by cryogenic distillation.Cryogenic separation unit are operated

at extremely low temperature and high pressure to separate components according to their

different boiling temperatures. A good economic and less costly scenario would be to have a

high inlet flow rate, high mole percent methane and a feed at low temperature.

Separation of Carbon Dioxide and Methane by Cryogenic Distillation

The HYSYS program was used to create a PFD for the separation process. Separators were used

to separate the two components. Methane has a higher boiling point of 87K and carbon dioxide

has a boiling point of 243K. Therefore, methane is vented off through the overhead of the

separator and is called the distillate. Liquid carbon dioxide is collected at the bottom of the

separator and is called the bottoms product. An XY plot of Methane and Carbon dioxide is

shown below. A McCabe Thiele Analysis shows that this is a one stage easy separation process.

Figure (C-1). Equilibrium plot for methane.


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Figure (C-2). The PFD on HYSYS.

Assumptions

The inlet temperature for the feed was assumed to be 0F.

The feed pressure of 39.70 psia was increased by 200 psi for first compressor.

The first cooling heat exchanger is used to cool the gas mixture to -120F.

The second cooling heat exchanger is used to cool the gas mixture to -175F, since at this

temperature the maximum methane is collected in the vapor form.

The heat exchanger works at 100% efficiency.

Figure (C-2). First and second cooling heat exchangers

First heat exchanger

Second heat

exchanger

First Compressor Second Com ressorFirst Se arator Second Separator


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The landfill feed( CO2 and Methane) containing 62 mole percent carbon dioxide and an inlet

molar flow rate of 10 MMCSFD was allowed to enter a compressor at temperature of 0F in

order to compress the gas to a pressure 239.7 psia. A low temperature was assumed so as to have

favorable economic condition. The compressed gas was then cooled in a cooling heat exchanger

to a temperature of -120F. The exiting gas was then sent to the first separator where the mixture

was separated and the fractional concentration of methane and carbon dioxide produced was

found to be 0.88 and 0.11 respectively.

In order to get a more purified concentration of methane, the methane product from the first

separator was sent into another separator passing through a cooling heat exchanger again. The

purpose of using a 2 stage process is to increase purity while simultaneously keeping the

operating costs at minimum. This exiting gas from the first reflux drum was further sent to a

cooling heat exchanger and cooled to a temperature of -175F. The mixture was then sent to a

second separator and methane was compressed to the given specification of 600 psi to produce

methane gas at a fractional concentration of 0.98 and carbon dioxide was produced at a fractional

concentration of 0.0186.

The two separators can be assumed to be two trays where the gas mixture is being separated. The

use of two separators instead of one was to get a higher purity concentration for methane.

Cost/Data Analysis

In order to calculate the operating cost for the entire system, the electric costs are taken into

consideration since the system is dominated by the cost of refrigeration. The electric costs if the

two heat exchangers and the two compressors to get the operating cost. The two tables below

show the cooling duty and the compressor duty electric cost.

HeatExchangerDuty (Btu/hr)

Heat ExchangerDuty (kwh/hr)

ElectricCost ($/yr)

OperatingTime (hr/yr)

Cooling DutyElectric cost($/yr)

HeatExchanger 1

8781000 2572.833 0.11 $2,000 $566,023.3

HeatExchanger 2

618400 181.1912 0.11 $2,000 $39,862.1

Table (C-1). Heat Exchangers Cost Analysis

CompressorDuty (Btu/hr)

CompressorDuty (kwh/hr)

ElectricCost ($/yr)

OperatingTime (hr/yr)

CompressorElectric cost($/yr)

Compressor1

2900000 849.7 0.11 $2,000 $186,934

Compressor2

618400 49.1654 0.11 $2,000 $10,816.388

Table (C-2). Compressors Cost Analysis


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The total operating cost is the sum of the cooling duty electric cost of the two heat exchangers

and the compressor electric cost of the two compressors.

Total OperatingCost

$803,635.7

Cooling DutyElectric cost($/yr)$566,023.3

Cooling DutyElectric cost($/yr)$39,862.1

CompressorElectric cost($/yr)$186,934

CompressorElectric cost($/yr)$10,816.388

The end methane product is of 98% purity, which is a pipeline quality gas and can be sold at a

price of $5.4/MMBtu. From the HYSYS simulation the flow rate of methane gas from the

second separator was found to be found to be 427.7 lbmol/hr. Using the following data, the total

cost of selling the methane per year was calculated using the heat of combustion of methane tobe 0.383061MMBtu/lbmol.

(Equation C-1)

Therefore, at a total operating cost of $803,635.7, the total cost of selling a feed of 98% methane

would cost$964,129.5 (Total cost at 98% purity Total operating cost). The total cost of sellingmethane at the feed concentration was calculated to be $2,271,865.84 without any operating

costs. The bar chart below shows the cost analysis of selling the methane product at 98% purity

and 38% purity.


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Figure (C-4). Cost Analysis Bar Graph.

Therefore, even though a profit is generated using cryogenic distillation, there is net loss of

$1,307,736.3 is generated when compared to selling the landfill feed methane. This proves that

Distillation would not be an ideal process.

Recommendations

With the feed conditions given, using cryogenic distillation is a profitable process for separating

methane and carbon dioxide but selling methane from the landfill feed was calculated to be more

profitable. Therefore, a net loss of $1,307,736.3 was generated if distillation was used as a

separation process. Therefore, Distillation may not be an ideal choice of process and the landfill

gas (feed) can be sold a higher price without purifying it.

$0.0

$500,000.0

$1,000,000.0

$1,500,000.0

$2,000,000.0

$2,500,000.0

Cost Analysis

$5.4/MMBtu $2.7/MMBtu


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Gas Separation

Overview

Membrane filtration is a very important process for separating a gaseous mixture into twocomponents which are purer streams. Modern examples of membrane separation in industry

other than methane and carbon dioxide include removal of H2S from natural gas, recovery of H2

in oil refinery processes, as well oxygen enrichment for metallurgical processes.

Gas separation using membranes utilize three streams; the inlet feed, and the outlet

retentate and permeate. The driving force through the membrane is a pressure difference which

provides the separation.

Figure D-1

The amount of separation depends on the selectivity of the membrane and the

Permeability of each component to the membrane. The membranes can be selective based on

molecule size, polarity, and volatility. The permeablities for each component through the

membrane are given as P(CH4) and P(CO2). Throughout this process, it was assumed that the

contents of every stream was completely mixed, which made the calculations of exit values

feasible.

Design

In the situation presented, a 62% by mole mixture of carbon dioxide was fed into the

separator at a flow rate of 10 MMSCFD. The rest of the mixture consisted of methane and the

pressure of the feed was 25 psig. The goal of the process was to separate the feed stream into two

streams that were high in methane content. Based on the concentration methane in the exit

streams, a certain price could be demanded for the cost of the natural gas.

Low Pressure

High Pressure


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Calculations

With the specified conditions given in the scenario, the amount of methane in each exit

stream needed to be calculated in order to do cost analysis. As an intermediate calculation, the

Yp is determined. This is found using intermediate values of a, b, and c which are in turn

calculated from r = p(l)/p(h), Xo, and * = P(CH4)/P(CO2). Ypis found using the quadratic

formula with a, b, and c. From this theta can be found and then the area of the membrane needed

is calculated. The area of the membrane is found using the following equation:

(D-1)

Process Analysis

When pressure on the permeate side is lower, the driving force is greater which means the

separation is greater. Calculations were performed with varying pressures on the permeate side

but as expected, the smallest membrane area was achieved when the pressure was lowest.

Physically, the lowest pressure possible for the permeate side is 1/10 of the high pressure

because anything more would cause too much strain on the membrane and blow it out. Using the

minimum pressure, the following values were determined:

Process Analysis

Xo Yp Theta Am Lo Vp CH4 in Lo CH4 in Vp

CH4 ft^2 mol/s mol/s mol/s mol/s

0.1 0.002262 -2.8648 -8224264 373.3718 -276.763 37.33718 -0.62599

0.4 0.014247 0.051847 234683.1 91.59959 5.008816 36.63984 0.071358

0.7 0.057793 0.498282 5253116 48.47019 48.13822 33.92914 2.782061

0.9 0.32602 0.905956 43121722 9.085475 87.52294 8.176927 28.53427

Table D-1

Table D-1 shows that the renentate methane concentration (Xo) cannot be less than feed

methane concentration (Xf) because that would then mean the membrane area would be

negative. Also, if Xo was less than Xf it would not make sense to use a membrane process and

just selling the straight feed stream would make more money. Other than the exit concentration

Am = ___LfYp______

Pat (PhXoPlyp)


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being 10% methane, every other concentration is possible, but in order to determine which is

best, an economic analysis of the costs and revenues must be done.

Another method that could be utilized would be a multiple-stage membrane process. In

this scenario, the retentate stream is then put into another shell where another membrane is

present to purify the stream even more. The retentate stream from the second shell will be morepurified than the feed; this could be done as many times as needed to get a desired methane

concentration. However this is an interesting idea, the feed flow rate already started very low and

each subsequent stream would only be lower. When a necessary concentration is achieved, the

flow rate would be so low that barely any money could be made from it. Also, the costs of each

shell is not known so that would need to be accounted for before proceeding with this idea.

Economic Analysis

In any process there are certain fixed and variable costs. With this separation, the fixed

cost is only the area of the membrane, which is $4/ft2. Each membrane needs to be replaced

annually, so that cost will be incurred every year this process is conducted. Additionally, there

are various variable costs that may need to be attributed to the overall cost of the process. For

this situation, those variable costs are heating, cooling, electric, and steam. These costs need to

be evaluated if the feed or exit streams need to be pressurized to meet certain conditions.

Pressurizing the inlet feed would increase the driving force and separation but that would

increase the costs and ultimately diminish the profits, so that was not done in this process.

Revenue is made by selling the exit streams which are at certain concentrations. The

above table shows those concentrations with Xo and Yp. For given conditions, the permeate

stream does not have a high enough methane concentration to sell for any value.

Cost Analysis

Xo Membrane cost Retentate Revenue Profit Return on Investment

.4 $938,732.34 $734,369.34 ($204,363.00) (21.77)

.7 $21,012,462.16 $906,718.60 ($20,105,743.56) (95.68)

.9 $172,486,888.73 $218,519.33 ($172,268,369.40) (99.87)

Table D-2

From Table D-2, the exit stream cannot achieve a profit at any of these exit

concentrations. This is due to the fact that there was not enough of a pressure gradient to drive

the feed through the membrane. This caused the area of the membrane to be very large, making it

more expensive than the exit streams could be sold for. When an exit stream of 70% is demanded


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with all other conditions constant, the area of the membrane must increase substantially. With

this increase, there was a slight increase in the price it could be sold for but it was not enough to

offset the costs. With then the increase to 90% methane, there is yet another increase in the area

of membrane needed but there is no increase in selling price so the profit decreases dramatically.

The only scenario where this process could be profitable at any concentration methane isif the low pressure was extremely close to a vacuum. Theoretically, that would mean that the

permeate side would have maximum ability to pull the feed stream through the membrane. If this

maximum pulling was available, the membrane would not have to be so big and the fixed costs

would greatly decrease. This vacuum idea is not physically possible due to the restrictions of the

membrane as described above.

Recommendation

With the feed already having a concentration of 38%, it is able to be sold for

$3.3/MMBTU and $602,017.83 could be made without any processes done to the feed. Due to

the tables and data above, using membrane separation to purify the given feed with given

conditions should not be done. By trying to purify the feed to a point where it could be sold at a

higher price is not worth it because of the price of the membrane. This process runs at a loss

even when conditions like low pressure are optimized. Even though no single-stage process

proved to be profitable, a multi-stage process could be investigated as a possible way to turn a

profit. Cost research for shells must be done and a second analysis of costs as determined by

area, shells, and pressurization.

Overall Recommendation

Of all four separation processes analyzed the gas membrane separation was the least

effective in producing a profitable exit stream. The minimum loss of the membrane unit was

$204,366. The cryogenic distillation unit generated less profit than selling the initial

unprocessed gas. The distillate contained 98% methane and the profit obtained from this

separation was $964,000. The profit obtained from the unprocessed gas was $2.3 million.

Therefore, it was unreasonable to design a distillation unit. The absorption processes both

generated a considerable profit. However, the absorption process using MDEA yielded thehighest profit possible at $19 million per year for a liquefied natural gas product. Therefore, the

final consensus of constructing an absorption column using MDEA solutions would be the most

economical option.


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Appendix

1.

Price estimates are in the range of $0.002/SCF (SCF = standard cubic feet).

2. Definitions

Heat Exchanger: A piece of equipment used in many chemical plants that is built for efficientheat transfer from one medium to another (i.e. from liquid to gas or vice versa).Separator: A piece of equipment that makes the necessary separation between a gas/ liquid

mixture giving a one gaseous and one liquid product.

Compressor: A mechanical device that is used to compresses a gas to a higher pressure.

References

Athanassios Vrachnos, Georgios Kontogeorgis, Epaminodas Voutsas, Thermodynamic Modeling

of Acidic Gas Solubility in Aqueous Solutions of MEA, MDEA, and MEA-MDEA Blends, Ind.

Eng. Chem. Res. 2006, 45, 5148-5154.

"Carbon Dioxide Properties."Engineering ToolBox. Web. 06 Dec. 2011.

.

"Cryogenic Distillation."Don't Waste Your Green Beer. Web. 05 Dec. 2011..Rousseau, Ronald W. "Distillation."Handbook of Separation Process Technology. New York: J.Wiley, 1987. 230. Print.

Geankoplis, C. J. Transport Processes and Separation Process Principles: (includes Unit

Operations). Upper Saddle River, NJ: Prentice Hall Professional Technical Reference, 2003.

Print.

Product CH4 concentration Price Specifications

Waste Gas < 10% - -Low Energy Gas 10 - 40% $2.7/MMBtu -

Medium Energy Gas 40 - 70% $3.3/MMBtu -

High Energy Gas 70 - 90% $4.4/MMBtu -

Pipeline Quality Gas 98% $5.4/MMBtu 600 psi

Liquified Natural Gas


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Heat of Combustion of methane: http://www.wolframalpha.com/.

"Landfill Gas."Illinois Department of Public Health Home Page. Web. 06 Dec. 2011..

"Landfill Gas to Energy: A Growing Alternative Energy Resource." TreeHugger. Web. 07 Dec.2011. .

"Methane - Specific Heat."Engineering ToolBox. Web. 6 Dec. 2011.

.

Photo courtesy: "Landfill Gas to Energy Overview."ACUA Home Page. Web. 05 Dec. 2011.

.

Photo courtesy: "Looks Into The Carbon Offset Project At The Tontitown Landfill."

TreeHugger. Web. 05 Dec. 2011. .

"Return on Investment (ROI) Definition."Business & Small Business / News, Advice, Strategy /

Entrepreneur.com. Entrepreneur Media, Inc. Web. 07 Dec. 2011.

.

Solomon, Stephen M. Plural Stage Distillation of a Natural Gas Stream. The Lummus

Company; 1976. Bloomfield, NJ. Pat# - 3983711.

"Water - Density and Specific Weight."Engineering ToolBox. Web. 6 Dec. 2011..

an economic cost evaluation of landfill gas[1]

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