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Gasification of Algae for the Production of CNG

Item Type text; Electronic Thesis

Authors Seamans, Kimberly Anne

Publisher The University of Arizona.

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Link to Item http://hdl.handle.net/10150/146671

http://hdl.handle.net/10150/146671

Abstract:

This project involved designing a process that converts algae into compressed natural gas

(CNG). The algae enter the process in a water slurry at 2.5% algae by mass. The algae are

then converted into methane and other gases at supercritical conditions in a reactor. The

gas mixture is then purified by removing the carbon dioxide and hydrogen gases. The

sellable products this process produces are CNG and hydrogen. In addition, the carbon

dioxide and salts are recycled back to the algae farms. This makes the process

environmentally responsible as the carbon dioxide created in the process is not released

to the atmosphere. Before the process is built, it is recommended that the reactor and salt

separator technologies are tested on a pilot scale to ensure their viability on an industrial

scale. A process hazard analysis was completed as a part of this project in order to

eliminate any safety issues with the process. Despite the process and equipment

optimizations that were performed, the economic analysis suggested that the process is

not economically viable. However, it may become viable in the future through the

development of the catalyst and equipment technologies.

Roles and Responsibilities of Group Members

My Honors Thesis is a part of a project completed with my senior design team.

My team members are Zachary Ronan, Kyle Kryger, and Amanda Rubio. Each team

member contributed to various sections and aspects of the project.

As the Project Manager, I was responsible for compiling the report and being the

main point of contact with our professor and our mentor. I also maintained continuous

communication between all members of the group and consistently issued group updates

with the most updated project information to ensure that everyone was aware of the

progress made, impeding internal and external deadlines, and individual assignments.

In the beginning of the semester, I took the initiative and did the majority of the

initial research for our project, finding the critical journal article on which our process

was primary based. I also created the original stream tables and equipment tables that

Kyle was later responsible for managing.

For the design aspect of the project, I assisted Zac Ronan by calculating the duties

of the heat exchangers, volume of the reactor, and mass of the catalyst. This involved

using fluid data tables to determine the heat capacities and researching the process and

catalyst. I also assisted Zac by typing up his calculations, which represents over fifty

pages of calculations given in Appendix A.

Another large part of my involvement in the project was writing several sections

of the report. I was responsible for writing the executive summary, introduction,

conclusions and recommendations, and the Process Hazard Analysis sheets (PHAs).

Writing the introduction section of the report required background research of the process

and of the related topic of the biodiesel industry. Writing the PHAs involved

understanding the safety hazards of the equipment and the contents of the equipment in

our process. This involved evaluating various failures or upsets, both direct and indirect,

that could affect our system, assess the consequences, and determine the corrective

actions to be taken to prevent safety hazards from occurring.

Writing the executive summary and conclusions and recommendations required that I

have knowledge of every aspect of the report and understanding of the decisions made

and analysis of the safety, environmental, and economic aspects of the project. This was

more easily accomplished because I had been the project manager and had been editing

and compiling each section of the report as the project developed.

Aside from writing my own sections of the report, I also edited all other sections

of the reports. This involved checking for spelling and grammatical errors, as well as

checking for consistency with numbers presented in the report and consistency amongst

various sections.

After all sections of the report were written, I was responsible for compiling the

report and formatting it to the specifications required. While doing this, I also wrote all of

the references by putting them in MLA format. This also involved going through the

entire paper and ensuring consistent numbering of the references to ensure that after

compiling the report all references were properly referred to and listed in the number in

which they appeared in the report. I also ensured correct numbering and formatting of all

of the figures and graphs.

Zac Ronan was mainly responsible for the design of the process. He completed the

majority of the equipment design and calculations. While designing the equipment, Zac

also determined what type of equipment to use and the cost of the equipment using cost

calculations and thus was responsible for determining the majority of the information on

the equipment tables. He was also responsible for determining the catalyst used and

conditions of the reactor in the process, which involved extensive research of the process

and evaluating various tradeoffs. Zac also wrote the equipment description, rationale, and

optimization section of the report, as well as the nomenclature section of the report. He

also was responsible for checking the overall mass and energy balances of the process

and compiled this into Appendix C. Zac also wrote Appendix E, which included the web

printouts he found for equipment that could be purchased from industrial suppliers. Zac

also helped edit sections of the report.

Kyle Kryger was responsible for maintaining the BFD and PFDs for the process. He

also managed the stream tables, equipment tables, and raw materials and utility tables. He

also designed the pieces of equipment that Zac did not design. Kyle was also responsible

for writing the process description and economics sections of the report. He also did all of

the calculations that required ChemCAD and compiled his work in Appendix B. He also

wrote Appendix B and helped in editing some parts of the report.

Amanda Rubio was responsible for the safety and environmental sections of the

report. She was also the team secretary and took notes at all of the team meetings, the

important of which she wrote and included in Appendix F of the report.

Gasification of Algae for the Production of CNG

Submitted by: CHEE 443 Team Kazaam

April 30, 2010

_______________________ _______________________

Kyle Kryger Zac Ronan

_______________________ _______________________

Amanda Rubio Kim Seamans

ii

Executive Summary This project examined the feasibility of compressed natural gas (CNG) production from

algae. Detailed process calculations, cost analysis, and optimization of the process and

equipment design were conducted to achieve the most economically viable solution.

The process is designed with 3000 kg/day of algae entering the process in a water slurry.

After being sent through a series of heaters, the algae slurry is sent through a salt

separator where the salts are removed from the stream into a brine solution and further

processed before being sent to the algae farm at a rate of 3080 kg/day. The algae slurry is

then sent to a reactor where it is gasified into methane and other gases. The methane

stream is sent through a CO2 scrubber where CO2 is removed and further processed

before being sent to the algae farm at a rate of 2929 kg/day. The methane stream is then

further compressed into CNG and a hydrogen membrane is used to remove the excess

hydrogen. The hydrogen is further processed and sold as a by-product at 4.63 kg/day for

$2.47 per kg. The CNG produced is comprised of 90.2 mol % methane, 3.6 mol %

ethane, 1.3 mol % propane, and 4.5 mol % of various inert gases.

It should be noted that some assumptions were made while designing this process. Most

importantly, there were several assumptions made concerning the reactor R-101. It was

assumed that the gases behaved ideally and that the density of the stream was the same as

the density of water because the stream was 97.8% water by mass. The required mass

ratio of catalyst and the residence time for R-101 were based limited data. Variation of

these numbers may cause an error in the resulting volumetric flow rate and therefore the

flow rate through the process may need to be altered in order to obtain the 63 minute

iii

residence time required for the reactor. Because this information is based off of one

design experiment, further testing should be done on a pilot scale to verify the values for

the design of the reactor.

It was determined that by selling the CNG and hydrogen from the facility, a total yearly

production of $72,600 could be achieved. However, the NPV of the process was

determined to be ($37,500,000), making the process very unprofitable. In order for the

process to become economically viable, it was determined that the price of natural gas

would have to be $1704.4 per 1000 scf, an increase over today’s prices by a factor of 156,

assuming that the ratio of the prices of natural gas and hydrogen remains constant and

that no operating costs change. Costs can also be mitigated by finding a cheaper catalyst.

The current catalyst used accounts for $4,500,000 a year, which represents over half of

the annual operating costs. Research should also be done to minimize the wall thickness

of the R-101 and S101 vessels thereby reducing equipment costs.

If the catalyst, R-101, and S-101 were all reduced in cost to $100,000 each, the NPV

would increase to ($7,400,000), an increase of $30.1MM. This would significantly

improve likelihood of this project becoming economically viable at a future time.

However, due to the current market of CNG and the high costs of production associated

with producing CNG from algae, it is the recommendation of this report that the proposed

design not be used at this time. However, if in the future CNG gains higher market

demands and the cost of algae and the catalyst decreases, this project can be reconsidered

as the process is both straightforward in design and has a minimal impact on the

environment.

Table of Contents

Executive Summary……………………………………………………………………….ii

Introduction………………………………………………………………………………..2

Overall goal

Current Market Information

Project Premises and Assumptions

Overall Process Description, Rationale and Optimization…………...…………………...6

Equipment Description, Rationale and Optimization………….………………………...19

Quantitative BFD………….………………………..............................................34

PFD………….………………………...................................................................35

Stream Table………….……………………….....................................................39

Major Equipment List………….………………………………………………...45

Raw Material and Utility Requirements………….………………………...........50

Safety/Environmental Factors……………….………………………..……………….....51

Safety Statement

Environmental Impact Statement

PHA documents

Economic Analysis……………………………………………………………………....71

Conclusion and Recommendations………..……………………………………………..78

Nomenclature.…….……...………………………………………………………………80

References………………………………………………………………………………..82

Appendices……………………………………………………………………..…….......87

Appendix A...………………………………………………………………………...87

Final Calculations

Appendix B…...…………………………………………………………………….136

Spread Sheets with Explanations

Appendix C…...…………………………………………………………………….141

Overall Mass and Energy Balances

Appendix D………...……………………………………………………………….143

ChemCad outputs

Appendix E…………………………………………………………………………151

WEB printouts

Appendix F………………………………………………………………………….154

Meeting Logs

2

Introduction As global warming awareness increases and the depletion of petroleum fuels rises, more

and more research is being devoted to alternative fuels such as compressed natural gas

(CNG) from algae. Biofuel from algae is a very sustainable alternative as it does not

interfere with food production like some other biofuel alternatives. Moreover, CNG is a

non-corrosive and non-toxic energy source that has significantly lower air emissions

compared to other fossil fuels and is piped directly to the fueling station, reducing the

cost of transportation (1). From Table 1, it is evident that algae have the greatest potential

in supplementing conventional fuels. Not only do algae have a high photosynthetic

activity, but they only require about two grams of CO2 per gram of biomass generated.

This results in approximately a 78% reduction in carbon dioxide emissions and a 50%

reduction in carbon monoxide emissions (2).

Table 1: Gallons of oil produced per year by various oil sources

Source Gallons of oil per acre per year

Algae 5000-20,000

Coconut 287

Jatropha 207

Canola 127

Peanut 113

Sunflower 102

Safflower 83

Soybeans 48

Hemp 39

Corn 18

Overall Goal

This project examined the feasibility of CNG production from algae through detailed

process calculations, cost analysis, and optimization. The production rate of CNG was

determined to be 364 kg/day, which was based on the assumption of an incoming algae

stream of 3000 kg/day. The composition of the final product is comprised of 90.2 mol %

3

methane, with the remaining gases being ethane, propane, and inert gases. The

specifications for the CNG composition were based on industry specifications which

require that the stream must contain at least 88 mol % of methane (3).

In order to create the most cost-effective scenario, it was determined that the outlet

streams containing the brine and carbon dioxide should be recycled and sold to the algae

farm. This would utilize the flue gas emitted from the process and reduce carbon dioxide

emissions into the atmosphere. There are 3075.6 kg/day of brine solution produced. The

carbon dioxide is retrieved from both the carbon dioxide scrubber and flue gas scrubber

at a total of 2929.10 kg/day. The hydrogen is removed from the process stream using a

hydrogen membrane at 4.63 kg/day can also be sold at $2.47 per kg (4).

Current Market Information

Currently the production of algae as a source for biofuels is still in the pre-commercial

stages of development. Some testing has been done on a lab-scale, but in order for algae

to be produced economically, some financial and technical barriers need to be conquered.

CNG is currently being sold at $2.40 per gasoline gallon equivalent (gge) (5). At this

competitive price, the proposed process needs to compete against CNG acquired through

more cost-effective methods than through the gasification of algae. CNG can be produced

from fossilized natural gas or biogas (6).

Due to the current market of CNG and the high costs of production associated with

producing CNG from algae, it is the recommendation of this report that the proposed

design not be used at this time. However, if in the future CNG gains higher market

4

demands and the cost of algae decreases, this project can be reconsidered as the process is

both straightforward in design and has a minimal impact on the environment.

Project Premises and Assumptions

The original premise of the project was to design a process that produced fuel from algae

with a by-product. CNG from algae was chosen because of the novelty of the process.

Other products, such as biodiesel, already have significant research and studies as to the

economic feasibility of production whereas CNG from algae is in its early stages of

research and development. The decision to use brine and carbon dioxide as the by-

products was made in order to reduce emissions and increase profitability with a minimal

increase in equipment cost.

In the design phase of the project, some assumptions were made in order to define

unknown parameters or factors in the process. An algae source of 3000 kg/day was

chosen in order to produce CNG on a significant scale. In addition, the algae were

assumed to be available for purchase as a slurry with 2.5% algae by mass per a

conversation with Dr. Wayne Seames (see Appendix F). Furthermore, the required mass

ratio of catalyst and the residence time for R-101 was based on only one literature value

(7), which may be slightly different for the designed process. Assumptions were also

made regarding some of the heat exchangers and pumps that may affect the accuracy of

the calculations conducted in Appendix A. Other assumptions include the viability of

processes such as the gravitational salt separator, hydrogen membrane, and availability of

the catalyst, which will all be discussed further in the following sections.

In order to determine if the process was economically viable, equipment choice, cost

analysis, and feedstock and utility requirements needed to be determined. Detailed

5

process calculations were completed for the process equipment in order to determine

equipment sizes and specifications. The final calculations can be found in Appendix A.

Calculations that were done using spreadsheets can be found in Appendix B. Mass and

energy balances were performed throughout the process to evaluate the feedstock

requirements and flow rates. Detailed calculations of the mass and energy balances can

be found in Appendix C. Because of the limited availability of data on the designed

process, CHEMCAD was used for some of the calculations for the pumps, valves, and

heat exchanger duties to check against hand calculations to ensure accuracy. The outputs

from CHEMCAD can be found in Appendix D. Any WEB printouts, such as price

quotes, can be found in Appendix E. Meeting logs were kept in order to track progress

and decisions made throughout the project. A copy of the meeting logs can be found in

Appendix F.

6

Overall Process Description, Rationale, and Optimization The big picture of the process is displayed in the block flow diagram (BFD) in Figure 2;

this diagram represents the general layout of the process as well as showing where mass

enters and leaves the process. Summing the input streams, a total of 156,000 kg/hr enter

the process. Summing the output streams, a total of 156,000 kg/hr leave the process. A

more thorough description of the overall mass balance is provided in Appendix C.

A detailed schematic of the process is presented in the process flow diagrams (PFDs) in

Figures 3-6 below. These diagrams depict the process streams, their temperature and

pressure and their relation to pieces of equipment. The compositions of the streams are

further detailed in the stream tables (see Tables 4-9).

Methane Production Line

The algae source of the process in stream 1 enters the plant at 25 °C and a pressure of

101.3 kPa. The algae slurry then enters P-101 A/B where it exits in stream 2 at a

temperature of 35 °C and a pressure of 31000 kPa as seen in Appendix D-2. This stream

is then fed to E-101 where it absorbs heat from the effluent leaving R-101. The algae

slurry emerges from E-101 in stream 3. At this point in the process, the slurry is

considered a compressible liquid; it has not yet reached its critical temperature of 374 °C

(8). Stream 3 then enters F-101 where it absorbs heat from natural gas fired heater. The

algae slurry exits the furnace in stream 4 at a temperature of 400 °C and a pressure of

30990 kPa. At this point, the fluid is supercritical as it has passed the critical point of

water, 374 °C and 22060 kPa (8).

7

In stream 4, two changes occur; firstly, the dissolved nutrient salts required for algae

production precipitate out. Secondly, the suspended algae begins to degrade into a syngas

precursor, ammonium salts and sulfide salts through mechanisms that are not presently

fully understood (7). Ammonium and sulfide salts exist as solid precipitate suspended in

the water flow; a total of 13.8 hg/hr of salts from algae are formed as shown in Appendix

A-7. Stream 4 then enters S-101.

Salts are removed from the vessel through openings in the bottom, where they are pushed

out by a small flow of supercritical water in stream 15. This stream is a two phase flow,

consisting of both solid and liquid components. The desalted algae slurry exits S-102 in

stream 5 where it is fed to R-101.

The algae slurry has a residence time of 63 minutes in R-101; here it interacts with a

Ru/C catalyst in an 8:1 catalyst:algae mass ratio. The reaction converts the incoming

algae material to a mixture of gasses as described in the stream tables below (see Tables

4-9). No water is consumed in this reaction. This reaction is slightly exothermic,

nominally causing the temperature of the reactor effluent to rise to 401 °C as shown in

Appendix A-6. However, this small increase is assumed to be lost to the surroundings of

the reactor, causing the reactor effluent in stream 6 to have a temperature of 400 °C.

Stream 6 is then fed into the tube side of E-101, where it discharges heat to the incoming

algae slurry. This heat loss drops the temperature of the reactor effluent’s temperature to

117 °C in stream 7. This stream is then fed through the letdown valve V-101; the effluent

adiabatically drops to a pressure of 300 kPa and rises to a temperature of 122 °C in

stream 8 as demonstrated in Appendix A-8.

8

The temperature of the effluent is further dropped in E-102. Next, stream 8 is fed into the

exchanger where it discharges heat to cooling water. The cooled effluent exits the

exchanger as a two phase vapor/liquid flow in stream 9. This stream is then fed to the

knockout drum D-201; in this drum, the vapor and the liquid phases are separated. The

majority of the liquid water exits the bottom of the drum in stream 30. This stream is

released from the process without any further treatment as this water meets EPA

specifications for a petroleum refinery (9). The vapor reactor effluent containing the

remaining vapor exits D-201 in stream 10.

Stream 10 is then fed into S-201, a CO2 scrubber. Also fed into this vessel is process

water. This water enters the process in stream 26. This stream is then passed through P-

201 A/B where it emerges in stream 27. This stream is then fed into S-201. In this vessel,

the two streams intimately mix in a packed bed, coming to thermal equilibrium at 25 °C.

The vapor exits the top of the scrubber in stream 11; 96.7% of the CO2 is removed from

the vapor stream. Justification for this removal rate is shown in Appendix A-14. Stream

28 exits the bottom of the scrubber containing water and dissolved carbon dioxide. This

water is then directed back to the algae farm where it is utilized as a carbon source.

Stream 11 is split into two streams, stream 12 and stream 15. Stream 12 has a mass flow

rate of 18 kg/hr and is sent to F-201 where it is used as an energy source. Stream 15 is

sent for further processing to the 400 unit. Here, stream 15 enters P-401 A/B where it is

compressed to 22000 kPa and 500 °C in stream 45 as shown in Appendix D-2. This

stream is then passed to S-401 where the hydrogen is completely separated from the

methane mixture. Stream 46 exits the membrane free from hydrogen and is then fed to E-

401 where it discharges heat to cooling water; the methane mix exits the exchanger in

9

stream 47 and is cooled further in E-402, where it again discharges heat to cooling water.

The resulting stream exits the exchanger and exits the process in stream 48 at pressure of

21980 kPa. A total of 14.3 kg/hr of CNG are produced by this process.

A secondary product is also produced. Hydrogen gas diffuses across the membrane in S-

401; this mass exits the membrane chamber in stream 49 at a pressure of 50 kPa. The

pressure is maintained at this low level by the vacuum pump P-401 A/B; stream 49 then

passes through the pump. The hydrogen exits the pump in stream 50 at a pressure of 115

kPa. This stream is then fed to E-403 where it discharges heat to a cooling water stream.

The hydrogen then exits the exchanger in stream 51; this stream is then fed to the final

hydrogen cooler E-404 where it again discharges heat to a cooling water stream. The

hydrogen exits the exchanger and the process in steam 52.

Cooling water used to cool the reactor effluent enters the process in stream 23. This water

is fed to P-102 A/B where it is pushed in stream 24 to E-102. Here it absorbs heat before

being discharged from the process in stream 25. Cooling water enters the 400 unit in

stream 44, coming off of the discharge of P-301 A/B. This stream is split into four

streams which are fed to all 400 unit exchangers. Here they absorb heat from product

streams before combining to form stream 61. Stream 61 is exits the process and is

discharged.

Air Inlet/Furnace/Superheater Loop

As described in Appendix A-10, ambient air enters the process in stream 17. This stream

is then drawn into the blower P-202 A/B where it is pressured up to 140 kPa in stream 18.

This stream is then fed to E-201 where the air absorbs heat from flue gas as; the heated

air exits E-201 in stream 19. This stream is then fed to F-101.

10

The energy source for F-101 is derived from the methane stream exiting S-201. Methane

at is taken from S-201 effluent in stream 12. In order to facilitate proper mixing in F-101,

this stream is fed to the letdown valve V-201; stream 13 exits the valve at a temperature

of 24 °C and a pressure of 140 kPa as shown in Appendix D-3. This stream is then fed to

E-202 where it absorbs heat from flue gas. The resulting stream 14 exits E-202 and is

then fed to F-101.

Streams 14 and 19 mix in F-101 where a combustion reaction occurs. The flue gas

transfers heat to the algae slurry before exiting the furnace in stream 20. This stream

passes through E-201 where it discharges heat to the incoming ambient air. This in cooler

flue gas emerges in stream 21, which is then fed to E-202. In this exchanger the flue gas

discharges heat to the methane mixture entering the furnace. The flue gas exits this

exchanger in stream 22. This stream is then sent to the 300 unit for treatment.

Stream 22 is fed to the carbon dioxide scrubber S-301. Also fed into this vessel is process

water. This water enters the process in stream 26. This stream is then passed through P-

201 A/B where it emerges in stream 29. This stream is then fed into S-301. In this vessel,

the two streams intimately mix in a packed bed, coming to thermal equilibrium at 25 °C.

The vapor exits the top of the scrubber in stream 32; this stream has had 90% of its CO2

removed. Justification for this removal rate is shown in Appendix A-19. Stream 31 exits

the bottom of the scrubber and is then directed back to the algae farm where it is utilized

as a carbon source.

11

Brine Treatment

As described above, salts are removed from S-101 through openings in the bottom, where

they are pushed out by a small flow of supercritical water in stream 15. This stream water

and salts in a two phase flow. This stream is then passed through E-301 where it

discharges heat to a cooling water stream. The brine stream 33 leaving the exchanger has

a pressure of 30975 kPa. This stream is then fed through a letdown valve V-301; the

exiting stream 34 has a temperature of 86 °C and a pressure of 110 kPa as shown in

Appendix A-20. This stream is then charged to E-302 where it discharges heat to a stream

of cooling water; stream 35 exits the exchanger and is then routed back to the algae farm

to be used as a nutrient source.

Cooling water enters the process in stream 36 and is fed into P-301 A/B. Streams 38 and

41 branch off from the pump discharge and pass through E-301 and E-302, respectively,

where they absorb heat from the brine stream. The effluent from E-301 and E-302

combine to produce stream 43 which exits the process. The remainder of the P-301 A/B

discharge is routed to the exchangers in the 400 unit in stream 44, as described above.

Process Rational and Optimization

After an analysis of many available process options, the final process presented in this

report represents an optimal blend of practicality and energy efficiency. Some

innovations were made as described below.

Reactor Conditions

Perhaps the biggest innovation incorporated into the process is the implementation of a

large scale reactor to convert algae into CNG. Supercritical conditions are required in the

12

reactor; the high temperatures and pressures initiate the breakdown of algae and provide

conditions that favor the production of methane over other products such as carbon

dioxide. Reactor conditions for the supercritical gasification of S. platensis were defined

by Stucki, et. al (7). In their experiments, they varied catalyst, algae concentration in the

feed, the weight ratio of catalyst to algae, and the residence time in the reactor. They

published the results of their research, which totaled 13 different combinations of the

aforementioned factors (7).

In selecting the reactor conditions most desirable for the process, the first factor

considered was catalyst choice. Two options were presented: Ru/C and Ru/ZrO2. Ru/C is

a commercially available catalyst (10). Ru/ZrO2 is not commercially available; this

catalyst was manufactured by the researchers in the laboratory (11). Because of the

impracticality of manufacturing large quantities of Ru/ZrO2 in house, Ru/C was selected

to be the catalyst used in our process.

All of the experiments using Ru/C were then compared to see which one produced the

most methane. The reactor was designed based on the best results as follows:

Catalyst: Ru/C

Feed concentration: 2.5 mass % algae

8:1 catalyst:algae mass ratio

63 minute residence time

The reactor effluent from this experiment had the following composition

42.7 vol % methane

1.7 vol % ethane

0.6 vol % propane

49 vol % carbon dioxide

0.1 vol % carbon monoxide

5.8 vol % hydrogen

13

Salt Separator

Another major innovation incorporated into the process is the inclusion of a supercritical

salt separator. Salts must be separated from the algae stream or else they would function

as a catalyst poison (7); a supercritical salt separator is the best alternative when

compared to traditional salt separation techniques such as utilizing an RO membrane.

Algae contains both nitrogen and sulfur; these elements form salts in a supercritical fluid

(7). Therefore, after the algae is taken supercritical, certain ammonium and sulfide salts

will form. These salts are insoluble in supercritical water (7); therefore, both the nutrient

salts and the salts formed from algae will exist as a solid in supercritical conditions. This

also implies that salt separation must occur after the solution is taken to a supercritical

state. Otherwise, the catalyst would be poisoned by the ammonium and sulfide salts.

Thus, salt separation should occur once the solution is in a supercritical state.

S-101 is a gravitational salt separator. With a residence time of 1 hour, the fluid velocities

through this vessel are low as seen in Appendix A-7. This allows the solid material to

sink to the bottom of the vessel, effectively separating the salts from the algae slurry

stream. This approach introduces the operational risk that the solid material could

accumulate and block exchanger tubes or process pipes. Furthermore, this approach

carries the risk that precipitate could foul the exchanger tubes, leading to a loss of heat

exchange. From a process safety standpoint, pipes and exchangers should be sized so that

fluid velocities are high enough to minimize the risk of precipitate blockage and scaling.

A gravitational salt separator was chosen as the best of three options, the other two

choices being a reverse flow gravity separator (RFGS) and a supercritical membrane

separator. In a RFGS, the algae solution would enter the vessel as a compressible liquid.

14

The slurry would be heated to a supercritical state inside the vessel, allowing the salts to

precipitate out as described above (12). From a process safety perspective, utilizing a

RFGS would be advantageous because it minimizes the risk or precipitate blocking pipes

or exchanger tubes; this is accomplished by having the salts precipitate and

gravitationally settle out of the fluid in the same vessel. Utilizing a RFGS would also

minimize the risk of precipitate scaling in E-102, thus preventing loss of effective heat

transfer. For reasons described in the Equipment Description section, a traditional heat

exchanger and a gravitational salt separator were selected over a RFGS. However, from a

process safety standpoint, if pipes and exchangers are sized so that fluid velocities are

high enough as described above, the process safety advantage a RFGS has over the

system selected is greatly reduced.

A supercritical membrane separator was not utilized for a number of reasons. First and

foremost, a membrane in supercritical application is typically used to separate solute

dissolved in a solvent in the product purification phase in order to eliminate the

depressurization and cooling steps usually required (13). Seeing as our process already

has the salts that are precipitated, there is no solute needed to be separated. A membrane

in the context of the process presented here would act as a physical barrier, forcing

suspended solids to separate from the fluid. The team believes that because of the low

fluid velocities in S-101 as described above, no membrane is needed in order to assist

with the separation.

However it should be noted that the gravitational salt separator presented here is an

untested piece of equipment. Thus, it is recommended that a pilot scale salt separator be

constructed and tested to ensure reliable equipment operation.

15

Algae Source

When selecting an algae variety to use as a feedstock, the typical biofuel process is

concerned with the lipid content of species. In a biodiesel plant, these lipids are then

extracted and converted into biodiesel (14). The process presented here is unique in that

the lipid content of the algae is irrelevant; because the entire cell is gasified, lipids were

not a factor in deciding the algae species to be used.

As such, the two biggest factors in considering algae source were availability of

information and growth rate of the algae. The algae selected must be able to reproduce

quickly in order to minimize the size of the algae farm associated with the process. After

consideration, the microalgae Spirulina platensis was selected. This algae is thought to

have a molecular formula proportional to C1.0H1.71O0.48N0.19S0.005 (7). The primary reason

that this algae was selected was that it is the only one with sufficient information

available describing the products formed by supercritical gasification and methanation.

(7). S. platensis was found to have a growth rate of 0.04 g/(L day) in open ponds, a

relatively high growth rate (15).

A total of 3 tons/day of algae enter the process, as was determined from a meeting with

Dr. Wayne Seames (see Appendix F). This algae is piped in directly from the algae farm

in the form of a slurry with 2.5 mass % algae (7). This solution also contains 22.8 g/L

nutrient salts; this number assumes that half of the nutrient salts needed to encourage

growth in solution are consumed (16).

16

Product Quality

A total of 14.3 kg/hr of CNG is produced by this process. At 21980 kPa, the product is

well within the acceptable pressure range for CNG of 200-250 bar (17). The product is

also well within composition specifications, as detailed by Table 2. Calculations

justifying these numbers are shown in Appendix A-14.

Table 2: Composition Specifications and Product Specifications of CNG

Compound Specification(3)

Product

Methane 88 mol % minimum 90.2 mol %

Ethane 6 mol % maximum 3.6 mol %

Propane 3 mol % maximum 1.3 mol %

Inert Gasses 4.5 mol % maximum 4.5 mol %

The hydrocarbon content of the stream is defined by the reactor R-101 and is relatively

difficult to change through traditional methods such as gas absorption. The hydrocarbons

are produced in the correct proportions and do not required blending or further reaction

before sale of the product. The inert gasses, however, are initially present in too great of

an amount to meet product quality specifications. These inert gasses include water,

carbon dioxide and carbon monoxide.

Carbon dioxide is the biggest contaminant in the reactor effluent. As such, the vapor

effluent is sent to S-201, a carbon dioxide scrubber. Here, 96.7% of the carbon dioxide is

removed from the effluent stream as shown in Appendix A-14. This brings the final mole

percent of CO2 in the CNG product to 3.1%. Water is introduced to the stream in S-201;

the water comes to equilibrium with the gas. As such, the final water mole percent in the

CNG product is 1.2%. The final CO mole percent in the product is 0.2. Summing these

together, the final mole percent of inert gasses in the product is 4.5%. By running the

product at the maximum allowable specification, money is saved by reducing the amount

17

of water needed in S-201. This minimizes the size of P-201 A/B and S-201, thereby

minimizing their costs as well. In the end, an on-spec product is delivered for minimal

cost.

No H2 gas is allowed to be in the final CNG product; however, 0.19 kg/hr of the gas are

in the stream sent to the 400 unit. This gas must be removed from the final product in

order to make the gas suitable for use in motor vehicles. This is accomplished by passing

the gas past S-401, a membrane engineered to remove hydrogen gas from a process

stream. The amount of hydrogen in the CNG stream leaving S-401 is negligible, thus

establishing a pure, on spec CNG product. The hydrogen exits S-401 in its own stream;

this secondary stream can also be sold as a fuel source.

Heat Exchanger Network

In order to maximize the amount of CNG produced by the process, the amount of energy

wasted must be minimized. This was accomplished through the construction of a

thorough heat exchanger network. The following incoming process streams are the only

places where opportunity exists to recover heat from outgoing process streams:

1. Algae stream from inlet to E-101

2. Air stream from inlet to F-201

3. Methane inlet to F-201

The first of these streams absorbs 2300 kW of heat from the outgoing reactor effluent. It

reaches a temperature of 372 °C; of all the process streams available to further heat this

while maintaining appropriate temperature approaches (18), none have a high enough

temperature. Thus, this stream has absorbed all the heat it can.

18

In order to minimize the amount of methane required for combustion in the furnace, both

the incoming methane and air streams must be preheated. The inlet air stream is heated in

E-201 to a temperature of 465 °C. There are no other streams available with a high

enough temperature to heat this stream any further. The incoming methane stream is

heated in E-202 to a temperature of 131 °C; while the brine stream could be used to

transfer a small amount of additional heat to this stream, this exchanger was not

implemented because the amount of heat recovered would not justify the cost of the

exchanger. Because this brine stream is the only steam available to heat up the methane

going into the furnace, the methane stream is as hot as it can get. Because all streams

cannot accept anymore heat from process effluent, the heat exchanger network is deemed

to have maximum efficiency. All exchanger calculations can be found in Appendix A.

Furnace

The primary heat source for the process is the methane combusted in F-101. In an ideal

world, no methane would be combusted and high pressure steam would be used to

superheat the algae slurry. However, because the algae slurry is required to be at 400°C,

steam cannot be used (18). The only other option for heating the stream is with a fired

furnace (18). Rather than purchase outside fuel sources, methane from the process was

chosen to be burned in the furnace. For a given volume of natural gas, it is cheaper to use

natural gas from the process than natural gas purchased from external sources.

19

Equipment Description, Rationale, and Optimization

E-101 and E-201

E-101 and E-201 are carbon steel floating head countercurrent shell-and-tube heat

exchangers. The duties and areas of the heat exchangers are 2038 kW and 436 ft2 for E-

101 and 80 kW and 240 ft2 for E-201 as shown in Appendix A-1 and A-2, respectively.

The reactor effluent is on the tube-side and the algae slurry is on the shell-side of E-101

in order to minimize heat loss to the surroundings. The rationale for this is that having the

higher temperature stream on the tube-side of the heat exchanger decreases the

temperature gradient between the shell-side of the exchanger and the surroundings,

thereby decreasing the heat lost to the surroundings (19). The flue gas is on the tube-side

and the air is on the shell-side of E-201 for the same reason.

The two streams flow in a countercurrent fashion for both of the heat exchangers in order

to allow for more heat to be transferred between the streams than would be obtained from

concurrent flow (20). For E-101, the countercurrent design allows the algae slurry stream

to exit the heat exchanger 28 oC less than the incoming reactor effluent stream, which

enters at a temperature of 400 oC. Similarly for E-201, the exiting temperature of the air

stream is set to be 28 oC less than the temperature of the entering flue gas, which enters at

a temperature of 493 oC. The exiting temperatures of the shell-side streams are set to be

28 oC less than the temperature of the entering tube-side streams because of the minimum

recommended temperature approach for temperatures above 150 oC (18). Having a

temperature approach less than 28 oC would cause the area of the heat exchanger to

increase dramatically per unit decrease in temperature. This would then cause the price of

the heat exchanger to increase at an exponential rate (18).

20

E-101 and E-201 are a shell-and-tube heat exchanger because of their relatively large

areas required for heat exchange. Areas of 436 ft2 and 240 ft

2 are outside the

recommended size for a double-pipe heat exchanger (18). E-101 and E-201 are floating

head shell-and-tube heat exchangers because of the ability of the floating head design to

handle large temperature differences between the tube and shell-side fluids (21). Floating

head designs are more expensive than other designs, however the 2038 kW and 80 kW of

energy conserved by E-101 and E-201 respectively more than justify their expense.

E-102, E-202, E-301, E-302, E-401, E-402, E-403, and E-404

E-102, E-202, E-301, E-302, E-401, E-402, E-403, and E-404 are all carbon steel double-

pipe heat exchangers. The duties and areas of these exchangers are 518 kW and 156 ft2

for E-102, 1.4 kW and 5.4 ft2 for E-202, 28 kW and 1.8 ft

2 for E-301, 3.8 kW and 1.7 ft

2

for E-302, 4.9 kW and 6.1 ft2 for E-401, 0.14 kW and 2.4 ft

2 for E-402, 0.34 kW and 0.5

ft2 for E-403, and 0.031 kW and 0.3 ft

2 for E-404 as shown in Appendix A.

The reactor effluent flows in the inner tube and the cooling water flows through the outer

tube of E-102. The cooling water flows through the outer tube of the heat exchanger in

order to prevent the safety issues involved with having the hot reactor effluent in the

outer tube. The outer temperature of the heat exchanger can be as hot as the fluid flowing

in the outer tube. Having the outside temperature of E-102 at 122 oC, the temperature of

the reactor effluent, would be hazardous to those working around the exchanger. The

cooling water flows through the outer tubes of E-301, E-302, E-401, E-402, E-403, and

E-404 for the same reason. The colder stream for E-202, the methane gas, flows through

the outer tube of the heat exchanger for the same reason. Also, this configuration for E-

202 reduces energy lost to the surroundings.

21

The two streams flow in a countercurrent fashion in all of the exchangers in order to

allow for more heat to be transferred between the streams than would be obtained from

concurrent flow (20).

For E-102, E-301, E-302, E-401, E-402, E-403, and E-404 the exiting temperature of the

cooling water is set to 49 oC by an environmental limit (18). Setting the exiting

temperature of the water to its maximum allows for the minimum amount of cooling

water to be used.

For E-102, E-302, E-402, and E-404 the exiting temperatures of the inner tube streams

are set at 43 oC, which is 11

oC above the temperature of the incoming cooling water.

This is because of the 11 oC minimum recommended temperature approach for streams

above ambient temperature but below 150 oC (18). For the same reason the exiting

temperature of the methane stream from E-202 is set to be 11 oC less than the temperature

of the entering flue gas. For E-401 and E-403, the exiting temperatures of the gas streams

were set to be 28 oC more than the temperature of the entering cooling water. This is

because of the minimum recommended temperature approach for temperatures above 150

oC (18). Similarly to E-101, the minimum approach temperatures are set in order to

maintain a reasonable heat exchanger area and cost.

The exiting temperature of the brine solution stream in E-301 is set to be 80 oC in order to

have a large approach temperature while keeping the water below its boiling point at a

pressure of 110 kPa. Increasing the temperature approach above the minimum

recommended approach, which is 28 oC for E-301 because temperature of the entering

brine is above 150 oC, decreases the area required for the heat exchanger thereby

22

decreasing its cost (18). At the same time, cooling the brine solution stream down to 80

oC keeps the water in the stream in the liquid phase after passing through V-301, as

shown in Appendix A-20. This prevents the salts from being left behind as the water

cannot change into the vapor phase.

E-102, E-202, E-301, E-302, E-401, E-403, E-403, and E-404 are double pipe heat

exchangers because of their relatively small areas. Double pipe heat exchangers are

recommended for areas up to 200 ft2 and are cheaper than shell-and-tube exchangers (18).

F-101

F-101 is a carbon steel methane fueled furnace. The duty of the furnace is 277 kW as

shown in Appendix A-3. A methane fueled furnace is used in this application because of

the abundant supply of methane that the process produces. A furnace is preferred over

other methods of heating because of the ability to integrate the furnace inputs and outputs

in other areas of the process in order to conserve energy. As shown in the 200 unit PFD,

the methane is sent through the methane preheater E-202 where 1.4 kW of energy are

conserved from the furnace flue gas. In addition, 80.1 kW are conserved from the furnace

flue gas through E-201 where the air is preheated. The combined energy conservation

allows the furnace to heat the algae slurry while using only 18 𝑘𝑔

𝑕𝑟 of methane, as shown in

Appendix A-3. Without this energy conservation, the furnace would not have been as

economical as a tubular heater, which would have required an estimated power source of

350 kW (22).

P-101A/B

P-101A/B are carbon steel reciprocating pumps with a required head of 10600 ft of water

and a required flow rate of 240 𝑔𝑎𝑙

𝑚𝑖𝑛 as shown in Appendix A-4. Reciprocating pumps are

23

used in this application because of the large head required by the pump. Reciprocating

pumps are recommended for applications requiring a head between 3200 and 20000 ft of

water (18). Reciprocating pumps are more expensive than centrifugal pumps, but a

centrifugal pump cannot be used in this application because of the large head required in

the process.

The electric motors that drive the pumps require a break horsepower of 82 hp and a

power consumption of 220 hp, as shown in Appendix A-4. The motors are enclosed in an

open, drip-proof enclosure. Open, drip-proof enclosures are designed to prevent the

entrance of liquid and dirt particles, but not airborne moisture or corrosive fumes, into the

internal working parts of the motor. This type of enclosure works in this particular

application because there are not any corrosive fumes being sent through the pump, as

shown Tables 4-9. Open, drip-proof enclosures are cheaper than other types of

enclosures, such as totally enclosed and explosion-proof enclosures, which protect the

inner working parts of the motor from fumes (18).

P-101A/B consists of two pumps because one of them serves as a backup for the other.

Since it costs $59,000 for one pump and one motor, as shown in Appendix A-4, which is

relatively cheap compared to the approximate $6,000,000 cost of all of the equipment, the

low cost justifies having a backup. The main goal of having a backup is so that the entire

process does not need to shut down in the event of the pump breaking.

P-102A/B, P-201A/B, and P-301A/B

P-102A/B, P-201A/B, and P-301A/B are cast iron centrifugal pumps. Centrifugal pumps

are used in these applications because of the small heads and flow rates required.

Centrifugal pumps are recommended for applications requiring a head less than 3200 ft

24

of water (18). Centrifugal pumps are used rather than reciprocating pumps because of

their lower cost and lower power consumption (18).

P-102 A/B has a required head of 4.0 ft of water and a required flow rate of 110 𝑔𝑎𝑙

𝑚𝑖𝑛 as

shown in Appendix A-5. The exact pumps used can be obtained from an industrial

supplier. The pumps are Dayton 2 hp centrifugal pumps. The pumps meet the head and

flow rate requirements of the process, with a maximum head of 63 ft of water and a

maximum flow rate of 170 𝑔𝑎𝑙

𝑚𝑖𝑛 (9). The exact pumps can be seen in Appendix E.

P-201A/B has a required head of 65 ft of water and a required flow rate of 560 𝑔𝑎𝑙

𝑚𝑖𝑛 as

shown in Appendix A-5. The electric motors that drive the pumps require a break

horsepower of 12 hp and a power consumption of 15 hp, as shown in Appendix A-5. The

motors are enclosed in an open, drip-proof enclosure for the same reasons as P-101A/B.

P-301A/B has a required head of 4.0 ft of water and a required flow rate of 8.6 𝑔𝑎𝑙

𝑚𝑖𝑛 as

shown in Appendix A-18. The exact pumps used in this application can be obtained from

an industrial supplier. The pumps are Dayton 1

3 hp self priming centrifugal pumps. The

pumps meet the head and flow rate requirements of the process, with a maximum head of

41 ft of water and a maximum flow rate of 43 𝑔𝑎𝑙

𝑚𝑖𝑛 (23). The exact pumps can be seen in

Appendix E.

P-102A/B, P-201A/B, and P-301A/B consist of two pumps for the same reason as P-

101A/B.

25

R-101

R-101 is a carbon steel packed bed reactor. Physically speaking, it is a high pressure

vertical vessel that is 3 m in height, has a 2.4 m inside diameter, and a volume of 13.7 m3.

It weighs 500,000 lbs and has a wall thickness of 24 inches. These sizes were chosen in

order to give the process fluid a residence time of 63 min as shown in Appendix A-6. A

residence time of 63 min is required in order to achieve the maximum conversion of the

algae into methane (7).

Assuming the same flow rate, a residence time smaller than 63 min would be a cheaper

option because the reactor would be smaller and therefore cost less. Also, less catalyst

would be required and the catalyst is one of the most expensive parts of the process at

$151 per 25 g of catalyst (24). However, the decreased residence time would cause

decreased methane production and an increased amount of undesirable products, such as

larger alkanes. The increased amount of the larger alkanes could result in the inability to

sell the product of this process, as the methane is required to be at a purity of 88% in

order to sell it (3), and separating the methane from the other alkanes would be difficult,

particularly for the ethane and propane. Assuming the same flow rate, a residence time

larger than 63 min would needlessly increase the weight and therefore the cost of the

reactor. In addition, the reactor would require more catalyst if the reactor size is

increased. Therefore, the volume of the reactor is set to be 13.7 m3 in order to have a

residence time of exactly 63 min so that the optimum combination of cost and product

purity is achieved.

Ru/C catalyst is used in this reactor in order to facilitate the gasification of the algae (7).

Ru/C is used over other ruthenium catalysts because of its availability. Although it is

26

expensive, it can be purchased from chemical suppliers (24). Other ruthenium catalysts,

such as Ru/ZrO2, selectively produce more methane than the Ru/C. In fact, the methane

conversion could be increased by as much as 20% with the use of Ru/ZrO2 (7). However,

Ru/ZrO2 is not available for commercial purchase and therefore would have to be made

on-site, thereby requiring a chemical plant to be added to the site. In addition, the

Ru/ZrO2 catalyst requires a residence time of 360 min, which would require six times as

much catalyst as the Ru/C catalyst (7). Therefore, Ru/C catalyst is used because of its

availability from commercial sources and for the lower volume of catalyst it requires.

S-101

S-101 is a gravitational salt separator. Physically speaking, it is a high pressure vertical

vessel that is 3 m in height, has a 2.4 m inside diameter, and a volume of 13.3 m3. It

weighs 371,000 lbs and has a wall thickness of 19 in. These sizes were chosen to give the

fluid inside the vessel a residence time of 1 hour as shown in Appendix A-7. Thus, the

fluid inside the vessel has a low fluid velocity as described in the process description in

order to facilitate effective phase separation. This size configuration minimizes the

amount of material used in construction, thus minimizing the cost of the unit as well.

This gravitational salt separator was chosen over

a reverse flow gravity separator (RFGS) (10). A

cross section of a RFGS can be seen in Figure 1.

If an RFGS were used, the algae slurry would not

be heated in F-101; instead, stream 3 would feed

directly into the interior chamber of the vessel (shown in figure as textured). Flue gas at a

temperature of 1708 °C as shown in Appendix A-3 would be piped into the outer

Interior

Chamber

Exterior

Chamber

Heat Transfer

Wall

Figure 1: Cross section of RFGS

27

chamber. The vessel would function as a large heat exchanger, passing heat from the flue

gas to the algae slurry in the inner chamber through the heat transfer wall. Thus, the algae

slurry would reach a supercritical state once inside the chamber. This would cause the

salts to precipitate out inside the vessel and fall to the bottom. The fluid, however, would

drop suddenly in density and rise up and out of the chamber, effectively separating the

two phases.

A gravitational salt separator and furnace combination was chosen over a RFGS for one

main reason. The flue gas at 1708 °C exceeds the safe operating limit of steel (18). To

operate the vessel with flue gas at that temperature could be dangerous. Lowering the

temperature of the flue gas is not an option; lowering the temperature would invariably

increase the amount of methane required to be combusted in the furnace due to the

increase in air flow required. A gravitational salt separator minimizes the amount of

methane consumed in the furnace; thus, the gravitational salt separator was deemed the

best choice.

V-101

V-101 is a letdown valve that adiabatically decreases the pressure of the reactor effluent

by 30670 kPa. The valve is used in the process in order to decrease the pressure of the

process stream without vaporizing the water. At the conditions of the process stream as it

exits the valve, 300 kPa and 122 oC, the liquid water does not vaporize as shown in Table

4. Vaporizing the water would greatly increase the volumetric flow rate of the process

stream and therefore would require the cooler, E-102, to have a much larger area. This

increased area is associated with an increased cost of the cooler. Therefore, to reduce the

28

cost of E-102 the exiting pressure of the valve is selected to keep the water in the liquid

phase.

D-201

D-201 is a carbon steel gravitational knockout drum. It is a low pressure vessel with a

height of 3 m, an inside diameter of 2.4 m, and a volume of 13.3 m3. It weighs 6,100 lbs

and has a wall thickness of 0.375 inches as shown in Appendix A-9. The volume of the

vessel is selected in order to allow for proper separation between the liquid and vapor

phases of the incoming process stream as shown in Table 4. Gravity pulls the denser

liquid water phase out of the bottom of the vessel while allowing the gases produced in

the reactor to flow freely out of the top.

P-202A/B

P-202A/B are aluminum centrifugal blowers with a required flow rate of 310 𝑓𝑡 3

𝑚𝑖𝑛 of air a

discharge pressure of 140 kPa as shown in Appendix A-13. A blower is used in this

application because of the required pressure increase of 40 kPa. Fans would be cheaper

alternatives, however fans are recommended for pressure increases of 10.6 kPa or less.

Compressors are not used because of their higher cost, and because they are

recommended for pressure increases of more than 206 kPa. A centrifugal blower is

chosen over a straight-lobe blower because of their lower purchase cost and higher

mechanical efficiencies. Centrifugal blowers have mechanical efficiencies of 70-80%, in

contrast to 50-70% efficiencies for straight-lobe blowers (18). Therefore, the centrifugal

blowers have a lower initial cost and a lower power consumption. The blowers require a

break horsepower of 9.7 hp and have a power consumption of 13 hp as shown in

Appendix A-13. P-202A/B consists of two blowers for the same reason as P-101A/B.

29

S-201

S-201 is a carbon dioxide scrubber; its task is to remove CO2 from natural gas. It removes

96.7% of the carbon dioxide from the incoming methane mixture, ensuring that the final

CNG product meets specifications as shown in Appendix A-14. The vessel itself has a

height of 2.5 m, an inside diameter of 1.9 m, and a volume of 7.2 m3. Made of carbon

steel, the vessel has a weight of 4,000 lbs and a wall thickness of 0.375 inches as shown

in Appendix A-14. The inside of the vessel is filled with 7 m3 of 1.5 inch ceramic

Raschig rings. Raschig rings of this size were chosen because of their low initial purchase

cost (18) and because they provide adequate area for the gas and liquid to interact. The

rest of the volume is taken up by distribution and redistribution apparatuses.

This vessel was sized to minimize the volume and amount of packing required while still

meeting the operational specification. See Appendix A-14 for calculations justifying

vessel sizing. In the end, S-201 was optimized by minimizing initial and operating costs

while still keeping the CNG product within specification.

V-201

V-201 is a letdown valve that adiabatically decreases the pressure of the methane by 145

kPa. The valve is used in the process in order to decrease the pressure of the methane so

that it has the same pressure as the air stream when it reaches the furnace F-101 as shown

in Table 5. The methane and air streams are required to be at the same pressure in the

furnace in order to prevent the backflow of one of the streams. If one stream is at a higher

pressure than the other, then the higher pressure stream may flow down the tube of the

lower pressure stream because of the pressure gradient between the streams. The

alternative to decreasing the pressure of the methane is to increase the pressure of the air.

30

However, to increase the pressure of the air to the pre-valve pressure of the methane

would require a compressor rather than the blowers P-202A/B. Compressors are

significantly more expensive than blowers and therefore are not preferable (18).

Therefore, the process lets the pressure of the methane down through the valve in order to

prevent the extra cost of adding a compressor.

S-301

S-301 is a flue gas scrubber; its task is to remove CO2 from flue gas before it is emitted to

the atmosphere. It removes 90% of the carbon dioxide from the incoming flue gas as

shown in Appendix A-19; this is an arbitrary number that lessens the carbon footprint of

the process without being too expensive of a unit. The vessel itself has a height of 2 m, an

ID of 1.5 m, and a volume of 3.5 m3. Made of carbon steel, the vessel has a weight of

2,070 lbs and a wall thickness of 0.3125 inches as shown in Appendix A-19. The inside

of the vessel is filled with 3.3 m3 of 1.5 inch ceramic Raschig rings. Raschig rings of this

size were chosen because of their low initial purchase cost (18) and because they provide

adequate area for the gas and liquid to interact. The rest of the volume is taken up by

distribution and redistribution apparatuses.

This vessel was sized to minimize the volume and amount of packing required while still

removing 90% of the carbon dioxide. See Appendix A-19 for calculations justifying

vessel sizing. In the end, S-301 was optimized by minimizing initial and operating costs

while reducing carbon emissions.

V-301

V-301 is a letdown valve that adiabatically decreases the pressure of the brine solution by

30865 kPa as shown in Table 7. The pressure of the brine solution needs to be let down to

31

near atmospheric pressure so that the brine solution can be safely sent back to the algae

farms. If the brine were sent back at too high of a pressure it could not be added to the

algae farm because of the large pressure gradient that would exist between the brine and

the algae slurry.

P-401A/B

P-401A/B are carbon steel screw compressors with a required flow rate of 4.6 𝑓𝑡 3

𝑚𝑖𝑛 and a

discharge pressure of 500 kPa as shown in Appendix A-25. A screw compressor is used

in this application because of their low purchase cost. Screw compressors are cheaper

than other types of compressors such as centrifugal and reciprocating compressors.

Centrifugal and reciprocating compressors are more flexible in their operation, and

therefore cost more (18). However, no flexibility is required by the compressor in the

process, and therefore screw compressors are chosen for their lower cost.

The motor driving the compressors are electric motors with a break horsepower of 8.4 hp

and a power consumption of 11 hp as shown in Appendix A-25. Electric motors are used

because of their higher efficiency. Electric motors have an efficiency of up to 95% at

1000 hp, in contrast to 65% and 35% for steam turbines and gas turbines respectively

(18).

Typically four compression stages with intercoolers would be used when compressing a

gas to a pressure that is 77 times its initial pressure (18). However, only one compression

stage is used in the process in order to increase the temperature of the gas to 500 oC, as

shown in Appendix D-1. The reason for this is because of the temperature requirements

of the hydrogen membrane S-401. Four compression stages with intercoolers would not

32

increase the temperature and therefore a heater would need to be added to heat the gas

stream. Adding a heater would increase the electricity use of the process as well as the

total cost of the process equipment. To eliminate this needless cost, only one compressor

without a cooler is used. P-401A/B consists of two compressors for the same reason as

P-101A/B.

P-402A/B

P-402A/B are carbon steel single stage rotary vane air-cooled vacuum pumps. The exact

pump used in this application can be obtained from an industrial supplier. The pumps are

TorrVac B series Rotary Vane Vacuum pumps. They are capable of displacing 21 𝑓𝑡 3

𝑚𝑖𝑛 of

gas and have a 1.5 hp motor (25), which meets the required flow rate of 0.43 𝑓𝑡 3

𝑚𝑖𝑛 as

shown in Appendix A-26. The exact pump can be seen in Appendix E. P-402A/B consists

of two pumps for the same reason as P-101A/B.

S-401

S-401 is a palladium hydrogen membrane cast on a ceramic support structure. The

purpose of this membrane is to separate hydrogen from the CNG stream; a dense metallic

membrane was chosen because of its superior properties. Table 3 summarizes

characteristics of the five main types of hydrogen membranes available today as

compiled by S.C.A. Kluiters (26).

Table 3: Characteristics of Various Hydrogen Membranes

Dense

Polymer

Microporous

Ceramic

Dense

Metallic

Porous

Carbon

Dense

Ceramic

Temperature Range (°C) <100 200-600 300-600 500-900 600-900

H2 Selectivity Low 5-139 >1000 4-20 >1000

H2 flux

(10-3

mol/m2s at dP=1 bar)

Low 60-300 60-300 10-200 6-80

33

The membrane must operate at a temperature of 500 °C as seen in Table 8. This clearly

eliminates the dense polymer and dense ceramic membranes from consideration. The

porous carbon membrane is eliminated as well because process upsets could drop the

temperature temporarily below 500, thus ruining the membrane.

Hydrogen selectivity was the next item considered in the choice of a membrane. The

selectivity shown in Table 3 is a unitless number defining the relative ease at which

hydrogen diffuses through a membrane compared to other compounds. (26) Palladium

membranes only let hydrogen diffuse through them; no other compound is physically

capable of moving through. (27) For this reason, a palladium membrane was chosen. By

utilizing a palladium membrane with a substantial pressure drop across the membrane, all

of the hydrogen is removed from the CNG stream as shown in Appendix A-27. As a

result, pure hydrogen and a maximum amount of CNG are produced by utilizing a

palladium membrane cast onto a ceramic support structure.

34

Figure 2. BFD for CNG Production from Algae

35

Figure 3. PFD for CNG Production from Algae

36

Figure 4. PFD for CNG Production from Algae (Continued)

37


38


39

Table 4. Steam Table Summary for CNG Production Process shown in Figures 3-6

Stream Number 1 2 3 4 5 6 7 8 9 10

Temperature (oC) 25 35 372 400 400 400 117 122 43 43 Presure (kPa) 101 31000 30995 30990 30980 30975 30970 300 295 290

Solid Fraction 0.025 0.025 0.025 0.025 0 0 0 0 0 0

Liquid Fraction 0.975 0.975 0.975 0.975 0.978 0.978 0.978 0.978 0.978 0 Vapor Fraction 0 0 0 0 0.022 0.022 0.022 0.022 0.022 1

Mass Flow (kg/hr) 5239 5239 5239 5239 5071 5071.28 5071.28 5071.28 5071.28 107.38

Component Mass Flow (kg/hr)

Algae 125 125 125 125 111 0 0 0 0 0 Air 0 0 0 0 0 0 0 0 0 0

Water 5000 5000 5000 5000 4960 4960 4960 4960 4960 1.40 Salts 114 114 114 114 0 0 0 0 0 0 Carbon Dioxide 0 0 0 0 0 81.7 81.7 81.7 81.7 76.4 Carbon Monoxide 0 0 0 0 0 0.21 0.21 0.21 0.21 0.21 Methane 0 0 0 0 0 26.0 26.0 26.0 26.0 26.0 Ethane 0 0 0 0 0 1.94 1.94 1.94 1.94 1.94 Propane 0 0 0 0 0 1.00 1.00 1.00 1.00 1.00 Hydrogen 0 0 0 0 0 0.43 0.43 0.43 0.43 0.43

40

Table 5. Steam Table Summary for CNG Production Process shown in Figures 3-6 (Continued)

Stream Number 11 12 13 14 15 16 17 18 19 20

Temperature (oC) 25 25 24 131 25 400 16 16 465 493 Presure (kPa) 285 285 140 135 285 30980 101.3 140 135 130

Solid Fraction 0 0 0 0 0 0.755 0 0 0 0

Liquid Fraction 0 0 0 0 0 0.245 0 0 0 0 Vapor Fraction 1 1 1 1 1 0 1 1 1 1

Mass Flow (kg/hr) 32.52 18.03 18.03 18.03 14.4 169.7 629.1 629.1 629.1 647.58


Algae 0 0 0 0 0 0 0 0 0 0 Air 0 0 0 0 0 0 621 621 621 548

Water 0.40 0.22 0.22 0.22 0.18 41.7 8.1 8.1 8.1 45.8 Salts 0 0 0 0 0 128 0 0 0 0 Carbon Dioxide 2.54 1.41 1.41 1.41 1.13 0 0 0 0 53.6 Carbon Monoxide 0.21 0.12 0.12 0.12 0.09 0 0 0 0 0.18 Methane 26.0 14.4 14.4 14.4 11.5 0 0 0 0 0 Ethane 1.94 1.08 1.08 1.08 0.86 0 0 0 0 0 Propane 1.00 0.56 0.56 0.56 0.45 0 0 0 0 0 Hydrogen 0.43 0.24 0.24 0.24 0.19 0 0 0 0 0

41


Stream Number 21 22 23 24 25 26 27 28 29 30

Temperature (oC) 142 136 32 32 49 25 25 25 25 43 Presure (kPa) 125 120 101.3 110 105 101.3 290 285 290 290

Solid Fraction 0 0 0 0 0 0 0 0 0 0

Liquid Fraction 0 0 1 1 1 1 1 1 1 1 Vapor Fraction 1 1 0 0 0 0 0 0 0 0

Mass Flow (kg/hr) 647.52 647.52 23300 23300 23300 125000 75300 75373.8 49200 4965.35


Algae 0 0 0 0 0 0 0 0 0 0 Air 548 548 0 0 0 0 0 0 0 0

Water 45.8 45.8 23300 23300 23300 125000 75300 75300 49200 4960 Salts 0 0 0 0 0 0 0 0 0 0 Carbon Dioxide 53.6 53.6 0 0 0 0 0 73.8 0 5.35 Carbon Monoxide 0.12 0.12 0 0 0 0 0 0 0 0 Methane 0 0 0 0 0 0 0 0 0 0 Ethane 0 0 0 0 0 0 0 0 0 0 Propane 0 0 0 0 0 0 0 0 0 0 Hydrogen 0 0 0 0 0 0 0 0 0 0

42


Stream Number 31 32 33 34 35 36 37 38 39 40

Temperature (oC) 25 25 80 86 43 32 32 32 49 32 Presure (kPa) 115 115 30975 110 105 101.3 110 110 105 110

Solid Fraction 0 0 0.485 0.485 0.485 0 0 0 0 0

Liquid Fraction 1 0 0.515 0.515 0.515 1 1 1 1 1 Vapor Fraction 0 1 0 0 0 0 0 0 0 0

Mass Flow (kg/hr) 49248.2 560.33 169.7 169.7 169.7 1900 1900 1440 1440 463


Algae 0 0 0 0 0 0 0 0 0 0 Air 0 548 0 0 0 0 0 0 0 0

Water 49200 6.85 41.7 41.7 41.7 1900 1900 1440 1440 463 Salts 0 0 128 128 128 0 0 0 0 0 Carbon Dioxide 48.2 5.36 0 0 0 0 0 0 0 0 Carbon Monoxide 0 0.12 0 0 0 0 0 0 0 0 Methane 0 0 0 0 0 0 0 0 0 0 Ethane 0 0 0 0 0 0 0 0 0 0 Propane 0 0 0 0 0 0 0 0 0 0 Hydrogen 0 0 0 0 0 0 0 0 0 0

43


Stream Number 41 42 43 44 45 46 47 48 49 50

Temperature (oC) 32 49 49 32 500 500 60 43 500 500 Presure (kPa) 110 105 105 110 22000 21990 21985 21980 50 115

Solid Fraction 0 0 0 0 0 0 0 0 0 0

Liquid Fraction 1 1 1 1 0 0 0 0 0 0 Vapor Fraction 0 0 0 0 1 1 1 1 1 1

Mass Flow (kg/hr) 191 191 1630 272 14.4 14.21 14.21 14.21 0.19 0.19


Algae 0 0 0 0 0 0 0 0 0 0 Air 0 0 0 0 0 0 0 0 0 0

Water 191 191 1630 272 0.18 0.18 0.18 0.18 0 0 Salts 0 0 0 0 0 0 0 0 0 0 Carbon Dioxide 0 0 0 0 1.13 1.13 1.13 1.13 0 0 Carbon Monoxide 0 0 0 0 0.09 0.09 0.09 0.09 0 0 Methane 0 0 0 0 11.5 11.5 11.5 11.5 0 0 Ethane 0 0 0 0 0.86 0.86 0.86 0.86 0 0 Propane 0 0 0 0 0.45 0.45 0.45 0.45 0 0 Hydrogen 0 0 0 0 0.19 0 0 0 0.19 0.19

44


Stream Number 51 52 53 54 55 56 57 58 59 60 61

Temperature (oC) 60 43 32 49 32 49 32 49 32 49 49 Presure (kPa) 110 105 110 105 110 105 110 105 110 105 105

Solid Fraction 0 0 0 0 0 0 0 0 0 0 0

Liquid Fraction 0 0 1 1 1 1 1 1 1 1 1 Vapor Fraction 1 1 0 0 0 0 0 0 0 0 0

Mass Flow (kg/hr) 0.19 0.19 247.3 247.3 17.3 17.3 7.09 7.09 0.66 0.66 272


Algae 0 0 0 0 0 0 0 0 0 0 0 Air 0 0 0 0 0 0 0 0 0 0 0

Water 0 0 247.3 247.3 17.3 17.3 7.09 7.09 0.66 0.66 272 Salts 0 0 0 0 0 0 0 0 0 0 0 Carbon Dioxide 0 0 0 0 0 0 0 0 0 0 0 Carbon Monoxide 0 0 0 0 0 0 0 0 0 0 0 Methane 0 0 0 0 0 0 0 0 0 0 0 Ethane 0 0 0 0 0 0 0 0 0 0 0 Propane 0 0 0 0 0 0 0 0 0 0 0 Hydrogen 0.19 0.19 0 0 0 0 0 0 0 0 0

45

Major Equipment List

Table 10: Equipment Table - Pump

Pump P-101 A/B P-102 A/B(28)

P-201 A/B

Type Reciprocating Centrifugal Centrifugal

Motor Open drip-proof CSTR

Open-drip-

proof

Component Algae/water Water Water

Flow [kg/hr] 5239.4 23321.9 124470.04

Fluid Density [kg/m3] 977 977 977

Efficiency 0.75 0.75 0.75

Power consumption (hp) 221 2 14.6

Break power (hp) 82.3 --- 11.94

MOC Carbon steel Cast Iron Carbon steel

Vendor --- Grainger ---

Pump P-301 A/B(29)

P-402 A/B(30)

Type Centrifugal Vacuum

Motor CSCR Rotary vane

Component Water Hydrogen

Flow [kg/hr] 1903.16 0.193

Fluid Density [kg/m3] 977 0.0899

Efficiency 0.75 0.75

Power (hp) 0.333 1.5

MOC Cast iron Carbon steel

Vendor Grainger US Vacuum

46

Table 11: Equipment Table - Heat Exchanger

Heat Exchanger E-101 E-102 E-201

Type Shell and tube Double pipe Shell and tube

Direction of Flow Counter-current

Counter-

current

Counter-

current

Duty (kW) 2308.381 518.265 32.601

Area (ft2) 435.71 156.07 233.8

MOC Carbon steel Carbon steel Carbon steel

Tube (Hot)

Component Reactor effluent

Reactor

effluent Flue gas

Inlet Temperature (oC) 400 122 493

Outlet Temperature (oC) 117 43 142

Pressure (kPa) 30975 295 125

Phase Liquid/vapor Liquid/vapor Vapor

Shell (Cold)

Component Algae slurry Cooling water Air



Pressure (kPa) 30995 105 135

Phase Liquid Liquid Vapor


Type Double pipe Double pipe Double pipe


Counter-

current

Counter-

current

Duty (kW) 1.24 28.3 3.8

Area (ft2) 5.4 1.8 1.7


Tube (Hot)

Component Flue gas Brine Brine



Pressure (kPa) 120 30975 105

Phase Vapor Liquid/solid Liquid/Solid

Shell (Cold)

Component Methane mix Cooling water Cooling Water



Pressure (kPa) 135 105 105

Phase Vapor Liquid Vapor

47


Type Double pipe Double pipe Double pipe


Counter-

current

Counter-

current

Duty (kW) 4.88 0.14 0.34

Area (ft2) 6.1 2.4 0.5


Tube (Hot)

Component CNG CNG Hydrogen



Pressure (kPa) 21985 21980 110

Phase Vapor Vapor Vapor

Shell (Cold)

Component Cooling water Cooling water Cooling water



Pressure (kPa) 105 105 105

Phase Liquid Liquid Liquid

Heat Exchanger E-404

Type Double pipe


Duty (kW) 0.013

Area (ft2) 0.3

MOC Carbon steel

Tube (Hot)

Component Hydrogen

Inlet Temperature (oC) 60

Outlet Temperature (oC) 43

Pressure (kPa) 105

Phase Vapor

Shell (Cold)

Component Cooling water

Inlet Temperature (oC) 32

Outlet Temperature (oC) 49

Pressure (kPa) 105

Phase Liquid

48

Table 12: Equipment Table – Heater

Heater F-101

Type Furnace

Component Air/Methane

Temperature (oC) 493

Pressure (kPa) 135

Power (kW) 277.2

Phase Vapor

MOC Carbon Steel

Table 13: Equipment Table - Blower

Blower P-202 A/B

Type Centrifugal

Component Air

Efficiency 0.75

Power consumption (hp) 12.9

Break power (hp) 9.86

MOC Aluminum

Table 14: Equipment Table – Reactor

Reactor R-101 A/B

Type Vertical Vessel

Catalyst Ru/C

Volume (m3) 13.67


Pressure (kPa) 30980

Phase Liquid

MOC Carbon Steel

Table 15: Equipment Table - Salt Separator

Salt Separator S-101

Type Gravitational

Volume (m3) 13.29


Pressure (kPa) 30990

Phase Liquid/Solid

MOC Carbon Steel

49

Table 16: Equipment Table - Scrubber

Scrubber S-201 S-301

Type Gas Absorber Gas Absorber

Volume (m3) 7.2 3.5

Temperature (oC) 25 25

Pressure (kPa) 285 115

Phase

Mixed

Liquid/Vapor

Mixed

Liquid/Vapor

Packing

1.5 in. Raschig

Rings

1.5 in. Raschig

Rings

MOC Carbon Steel Carbon Steel

Table 17: Equipment Table - Valve

Valve V-101 V-201 V-301

Type Letdown Letdown Letdown

Pressure Drop (kPa) 30670 145 30865

Phase Liquid Vapor Liquid/Solid

Table 18: Equipment Table -Drum

Drum D-201

Type

Vertical

Vessel

Volume (m3) 3.29

Temperature (oC) 43

Pressure (kPa) 295

MOC Carbon Steel

Table 19: Membrane

Membrane S-401

Type Dense Metallic

Hydrogen Flux [kg/hr] 0.193


Pressure drop (kPa) 21950

MOC Palladium/Ceramic

Table 20: Equipment Table - Compressor

Compressor P-401 A/B

Type Screw

Component CNG

Efficiency 0.75

Power (hp) 11.2

MOC Carbon Steel

50

Raw Material and Utility Requirements

Table 21 summarizes the annual utility requirements of the process. Calculations showing

yearly consumption are shown in Appendix C.

Table 21: Raw Material and Utility Requirements

Component Yearly Consumption Unit Cost Total Yearly Cost

Algae 1,095 tons $1.00/ton $1,095

Electricity 1,722,177.6 kWh $0.06/kWh(18)

$103,330.66

Cooling Water 220,971.53 m3 $0.02/m

3(18) $4419.43

Process Water 1,090,357.55 m3

$0.20/m3(18)

$218,071.51

Total Yearly Utility Cost $325,821.60

Total Yearly Raw

Material Cost

$1,095

51

Safety/Environmental Factors

Safety Statement

The safety of the plant is a key component to consider when constructing a prosperous

plant design, and safety requirements were reviewed to ensure there were no violations.

An examination of the chemicals, using several sources such as Material Safety Data

Sheets (MSDS) and the Environmental Protection Agency, was completed to evaluate the

potential health effects, PPE requirements, first aid measures, and handling and storage.

Accessible eye washes and body emergency baths will be located in each area of the

plant to ensure immediate access. When using the eye wash bath, it is important that the

individual lifts the upper and lower eyelids and removes contacts to improve rinsing. In

the event of a chemical spill on the body, an individual should remove any exposed

clothing to minimize the level of contamination and thoroughly rinse in the body bath.

The entire plant will be ventilated but a system will be integrated that will monitor the air

quality for contaminant levels. Safety meetings will be held to address any concerns that

may arise and to keep all personnel in compliance with all safety requirements. They

should only eat in designated areas after hands have been meticulously washed.

Table 22 lists all the chemicals that are used in the process and the possible symptoms of

exposure and the corrective actions. Hydrogen, salt, algae, ethane and methane do not

pose a threat to the safety of the personnel and the listed symptoms occur in rare cases. If

there is any exposure to these chemicals, including ingestion, the corrective actions

should be taken and medical attention is required, even if symptoms desist. Although

there are no fatal side effects to chronic exposure to these chemicals it is best to minimize

52

any exposure. Extra precautions should be taken when handling these chemicals and

employees should inform others when they are going to handle the chemicals in case of

an accident and to prevent leaks. It was determined that carbon dioxide carbon monoxide,

and propane were the most dangerous chemicals in the plant process. Gaseous methane

and ethane are simple asphyxiates but carbon dioxide, carbon monoxide, and propane can

cause suffocation at high exposure levels. Carbon monoxide has Permissible Exposure

Limit (PEL) set by the Occupational Safety and Health Administration (OSHA) of 50

ppm for an eight hour period (31). Carbon dioxide and propane has a PEL level of 5000

ppm and 1000ppm, respectively. These levels are not present in our process therefore

there is not a high concern (32, 33).

53

Table 22: Exposure Hazards and Appropriate Action

Inhalation Eyes Skin Ingestion

Compound Symptoms Action Symptoms Action Symptoms Action Symptoms Action

CO (31) Nausea, Vomiting,

Headache

Fresh Air,

Medical Attention Irritation

Flush with

water (15 min.) Irritation

Flush with cool


Medical

Attention

CO2 (32) Nausea, Vomiting,

Headache

Fresh Air,


Flush with


Flush with cool


Do Not

Induce

Vomiting

H2 (34) Nausea, Vomiting,

Headache

Fresh Air,

Medical Attention

Irritation,

Redness

Water, Medical

Attention Irritation

Flush with


Medical

Attention

C2H6 (33) Shortness of Breath Fresh Air,


Flush with


Flush with


Medical

Attention

CH4 (35) Rapid Breathing Fresh Air,


Flush with


Flush with


Medical

Attention

C3H8 (36) Suffocation Relocate to Fresh

Air

Irritation,

Burning

Flush with


Flush with


Medical

Attention

Salt (37) Irritation Relocate to Fresh

Air Irritation

Flush with


Flush with


Drink

Plenty of

Water

Algae (38) None Relocate to Fresh

Air Irritation

Flush with


Flush with


Drink

Plenty of

Water

H2O (39) None N/A None N/A None N/A None N/A

Air (40) None N/A None N/A None N/A None N/A

54

The most hazardous chemicals were further evaluated for their flammability, reactivity,

incompatibly and any other special hazard and organized in Table 23. The storage of all

the chemicals will be in a cool, dry, well-ventilated location away from its incompatible

compounds. For the chemicals in our process the most incompatible compounds are

oxidizers which could react with the gases to cause a fire. Since there are no oxidizers

present in the plant the relative threat is minimal, but all precautions should remain. If a

gas base fire was to occur, it must be extinguished with foam, dry chemical or carbon

dioxide, but not water.

Table 23: General Hazards for Process Chemicals

Compound Flammability Reactivity Incompatibilities Special Hazards

CO (31) High Medium Oxidizers, Barium

Peroxide

Use dry chemical, foam, or

carbon dioxide to quench fire

CO2 (32) None Low Reactive Metals,

Hydrides N/A

H2 (34) Extremely Medium Oxidizers,

Chlorine, Lithium



C2H6 (33) Extremely Low Oxidizers Use dry chemical, foam, or


CH4 (35) Extremely Low Heat, Sparks,

Flames, Oxidizers



C3H8 (36) Extremely Low Heat, Sparks,

Flames



Another safety consideration is the hazard that the equipment presents. A Process

Hazards Analysis (PHA) was conducted for most pieces of equipment and can be

reviewed below. The biggest hazard identified are upsets such as unit overpressures or

ruptures that release natural gas or hydrogen to the atmosphere, potentially causing a fire

or explosion. To minimize the number of ignition sources available, all electronics in

hazardous locations will have Specific Class I, Division 1 Groups B and D electrical

classification. Another risk is vessel failure when the vessel contents are supercritical; the

equipment is functioning at very high pressures and should be well maintained to prevent

55

any malfunction, rupture and resulting steam explosion in the equipment. Pressure

sensors will be installed on all necessary equipment and they will be hardwired to

appropriate alarms and automated. Many of the pieces of the equipment have a spare that

can be automatically utilized during maintenance or malfunction. Pump, compressor or

blower malfunctions could cause an upset down the line as in many of the cases for the

equipment, leading to fire or explosion. Safeguards against this possibility include

incorporating minimum flow lows and low flow alarms wired to the control house.

Steps were also taken to reduce temperature exposure safety concerns with individual

pieces of equipment. All of the shell-and-tube heat exchangers were designed to have the

higher temperature fluids on the tube-side. This design allows the outside of the heat

exchanger to be cooler than if the higher temperature streams were on the shell-side. All

of the double-pipe heat exchangers were designed to have the higher temperature stream

in the inner tube for the same reason. However, the outside temperatures of several pieces

of equipment are still at high temperatures, and therefore precautions need to be made in

order to prevent accidental exposure to the hot surfaces. Barriers will be set in place to

ensure these surfaces cannot be accidentally touched and heat-resistant clothing, such as

Nomex, will be required when working around these hot pieces of equipment.

Environmental Impact Statement

The effect that a chemical may have on the environment is an important aspect to

research when implementing a new process design. Throughout the process design, the

chemicals that were proposed were reconsidered if there was a negative impact on the

environment, such as a high toxicity level.

56

For all chemicals when an accidental spill or leak occurs it needs to be attended to

immediately to avoid further contamination, for the safety of the personnel and

environment. All piping will have a non-corrosive coating, nickel chrome alloy, targeted

to protect against gas corrosion (41). Chemical waste is either taken to a Resource

Conservation and Recovery Act (RCRA) approved disposal facility or incinerator. When

cleaning a chemical spill, all personnel will be required to wear the personal protective

equipment (PPE) appropriate to the situation and the supply source of the chemical

should be immediately shut off. If large amounts of any chemical are released into the

environment the Environmental Protection Agency should be informed. Providing a fire

was to start, the proper equipment needed to extinguish it would be available to minimize

byproduct discharge. Table 24 lists all the chemicals with their appropriate measures in

an accidental release, appropriate disposal method, and harmful degradation by-products.

Water and Air were not included in the table because they do not have any environmental

hazards.

Table 24: Appropriate Chemical Measures

Compound Accidental Spill Measures Disposal Method Degradation

Byproducts

Carbon

Dioxide (32)

Contained ventilated area,

collect in appropriate container

Observe State and

Local Regulations

Carbonic acid,

Carbon monoxide

Carbon

Monoxide

(31)

Contained ventilated area,


Observe State and

Local Regulations Carbon Dioxide

Methane (35) Contained ventilated area,


Send to a RCRA

approved incinerator

or facility

None

Ethane (36) Contained ventilated area,


Observe State and

Local Regulations None

Propane (33) Contained ventilated area,


Sealed Container to

BOC Gas Location None

Hydrogen (34) Contained ventilated area,


Sealed Container to

BOC Gas Location None

57

Further research was conducted to evaluate the environmental effect were there to be a

release of the chemical from the plant. Carbon dioxide is a gas that is abundantly found in

the atmosphere, occurring naturally and man produced. Although the majority of plant

and animal life rely on carbon dioxide, the amount that is released into the air has exceed

the amount can be naturally sequestered (42). With excess carbon dioxide in the air more

heat is captured and remains in the atmosphere known as the greenhouse effect. The

estimated amount of the mass of the greenhouse gas that can contribute to capturing heat

is known as Global Warming Potential (GWP), and carbon dioxide has a GWP of 1.

Another chemical gas that is dealt with in the process design in methane and it has a

GWP of 21(43). This is more elevated potential than carbon dioxide and all safety

measures to prevent a leak should be taken. Carbon dioxide also is released from the

plant due to the use of electricity for several pieces of our unit operation. The estimated

amount of carbon dioxide emissions from the use of electricity for one year was

determined to be 1,720,000 kg (44).

The discharge of a flammable chemical or material into the environment is a hazard that

should be avoided. Ethane, propane and the catalyst, ruthenium on carbon, are all highly

flammable materials that need to have a properly functioning shut off valve in case of

fire. When any of these chemicals burn it produces carbon dioxide and the possible

incomplete combustion can produce carbon monoxide. If a fire is to occur, remove the

source of the ignition and contain the ventilated gases.

58

Company:

Kazaam, LLC.

Plant: University of Arizona- Spring 2010 Chemical Engineering Department

Site: TBD Unit:100 System:

Method: What-if Type: Continuous Reactor Design Intent: To remove various salt components

Number:

Team Members: Kyle Kryger, Kimberly Seamans, Zachary Ronan, Amanda Rubio

No.: 1 Description: S-101 Salt Separator

Item What if...? Root Causes/Related Questions Consequences Safeguards Action Items

1.1 What if the vessel over-pressurizes?

Valve failure P-101 A/B disturbance Unchecked salt accumulation in process pipes What if the vessel has excessive salt accumulation?

Vessel rupture Steam explosion Loss of effective salt separation R-101 catalyst poisoning and product loss

Pressure relief valves Pressure sensors hardwired to quench system Emergency shut-off system

1.2 What if the vessel overheats?

F-101 upset Unintended runaway chemical reaction Fire nearby caused by failure of other pieces of equipment

Increase in temperature in downstream units, leading to product quality loss or fire Over-pressuring of vessel Overheating of brine sent to algae farm, killing process feedstock

Temperature sensor hardwired to quench system Emergency shut-off system

1.3 What if there is excessive salt accumulation in the vessel?

Valve blockage E-301 upset Corrosion What if the vessel over-pressurizes?

Loss of effective salt separation R-101 catalyst poisoning and product loss

Emergency shut-off/Isolation Low-flow alarms wired to control house Regular maintenance

1.4 What if the temperature drops too low?

E-101 tube leak F-101 upset Vessel insulation failure Accidental quench system triggering

Loss of effective salt separation Loss of conversion of algae to syngas Product loss Poisoning of R-101 catalyst

Temperature sensor hardwired to emergency high pressure steam heater Use of slop tank to pass cold slurry to in order to spare R-101 Vessel isolation system

59

Company

Kazaam, LLC.


Site: TBD Unit:100 System:

Method: What-if Type: Continuous Reactor Design Intent: To gasify algae biomass & create methane

Number:


No.: 2 Description: R-101 Gasification & Methanation Reactor


2.1 What if the reactor ruptures? V-101 failure causes overpressure Upstream upset Runaway chemical reaction causes overpressure External impact Corrosion by supercritical conditions What if the reactor leaks?

Loss of containment Steam explosion Methane fire/explosion

Pressure relief valve Pressure sensors hardwired to quench system Regular inspections of reactor to look for corrosion or leaks Corrosion allowance Stress relief

2.2 What if the reactor overheats? Runaway chemical reaction S-101 overheating What if the reactor ruptures?

Destruction of carbon catalyst support, leading to catalyst loss Overheating of downstream units, leading to off-spec CNG

Temperature sensor hardwired to quench system

2.3 What if the reactor leaks? Corrosion by supercritical conditions Weakened or melted by high temperatures External Impact Metal fatigue

Loss of containment Fire/explosion caused by methane leak

Regular inspections to look for potential trouble spots Periodic nondestructive testing Deluge system Protective barriers

Specific Class I, Division 1 Group B and D electrical classification for Hazardous Locations

2.4 What if there is catalyst poisoning/loss?

S-101 upset Overheating of reactor Corrosion Retaining mechanism failure

Product Loss Unintended chemical reactions Loss of feed to F-101 Loss of effective salt separation in S-101

Proper catalyst containment Temperature sensor hardwired to quench system Corrosion allowance Knockout drum to absorb S-101 upsets

60

Company:

Kazaam, LLC.


Site: TBD Unit: 100 System:

Method: What-if Type: Heat Exchanger Design Intent: To increase the temperature of the algae slurry to be able to reach supercritical values before entering the salt separator

Number:


No.: 3 Description: E-101 Preheater & E-202 Superheater


3.1 What if there is a tube leak or rupture?

Corrosion Stress cracking Weakened/melted by exposure to high temperature Tube cracks because of differential thermal expansion

Loss of containment Methane fire/explosion Corrosion downstream S-101 has influx of new unwanted compounds, potentially causing upsets

Corrosion allowance Periodic nondestructive testing Deluge system


3.2 What if the tubes are subjected to excessive pressure?

R-101 upset Tubes blocked during startup Heat applied with tubes blocked in What if the shell is over-pressurized?

Tube leak or rupture Loss of containment Methane fire/explosion

Pressure relief valves on each pass Low flow alarms wired into control room Deluge system Exchanger bypass system


3.3 What if the exchanger shell is over-pressurized?

P-101 A/B deadheaded Blocked exit valve F-101 upset induces high pressure in outlet line

Loss of containment Methane fire/explosion Tube leak or rupture

Pressure relief valves Pressure alarms wired to control room P-101 A/B over-speed trip


3.4 What if the exchanger loses cooling stream feed?

P-101 A/B trip Algae farm loses feed Valve closure What if there is a tube leak or rupture?

High temperature induced failures, fires or explosions in downstream units Loss of reactor feed Loss of product Supercritical unit depressurization

A/B redundant pumps Temperature alarms wired to control room Low flow alarms wired to control room Emergency cooling water source available


61

Company:

Kazaam, LLC.



Method: What-if Type: Heater Design Intent: To heat air entering the E-102 heat exchanger

Number:


No.: 4 Description: F-101 Furnace


4.1 What if firebox leaks or ruptures?

Corrosion by flue gasses Localized damage by flame impingement External impact Stress corrosion cracking

Release of hot flue gases Loss of heat transfer to algae Poisoning of catalyst in R-101 by loss of salt removal in S-101 Product loss

Automatic burner shutdown Vacuum relief system Double block valves, caps, or plugs in all sample and drain connections

4.2 What if the firebox over-pressurizes?

Closed stack damper on forced draft furnace Liquid carryover in fuel gas What if the firebox leaks or ruptures?

Loss of containment Fire/explosion Tube failure

Pressure relief valve Furnace trip signals Firebox air intake separated from likely sources of flammable gas leaks

4.3 What if the air stream cannot reach the furnace?

P-202 A/B failure Valve closure Air inlet blockage (debris, etc.) What if the air supply starts up again and there is accumulation of uncombusted fuel in the firebox?

Loss of heat transfer to slurry Poisoning of catalyst in R-101 by loss of salt removal in S-101 Explosion

Automated shutoff of fuel gas source Spare blower Low flow alarms wired to control house Furnace isolation

4.4 What if the flame is extinguished?

Low fuel pressure What if there are still explosive vapors inside the furnace? What if the air stream cannot reach the furnace?

Explosion caused by sudden reignition in firebox Release of natural gas into atmosphere Fire downstream of S-301

Temperature sensors hardwired to automatic valve shutoff system Deluge system Automatic reignition system


62

Company:

Kazaam, LLC.



Method: What-if Type: Condenser Design Intent: To remove water from the process stream

Number:


No.: 5 Description: E-401, E-402, E-403, E-404 Coolers



Corrosion Stress cracking Weakened/melted by exposure to high temperature Tube cracks because of differential thermal expansion

Loss of containment Product leak to environment Methane or hydrogen fire/explosion

Corrosion allowance Periodic nondestructive testing Deluge system


P-401 A/B upset Tubes blocked during startup Heat applied with tubes blocked in What if there is a tube leak or rupture?

Tube leak or rupture Loss of containment Methane or hydrogen fire/explosion

Pressure relief valves Low flow alarms integrated into control room Deluge system Exchanger bypass system Incorporation of semi-automated back-flushing system

5.3 What if the exchanger shell is over-pressured?

P-301 A/B deadheaded Blocked exit valve What if the tubes are subjected to excessive pressure?

Loss of containment Methane or hydrogen fire/explosion Tube leak or rupture

Pressure relief valves Pressure alarms wired to control room P-301 A/B over-speed trip

5.4 What if the exchanger lost cooling stream feed?

P-301 A/B trip Cooling water source disrupted Valve closure

Product produced at too high a temperature Loss of product Fire/explosion

Spare pump Temperature alarms wired to control room Low flow alarms wired to control room Emergency cooling water source available

63

Company:

Kazaam, LLC.



Method: What-if Type: Scrubber Design Intent: To remove CO2 and excess water from the process stream

Number:


No.: 6 Description: S-201 Scrubber


6.1 What if the scrubber leaks or ruptures?

External impact V-101 failure Blockage of tower packing Stress fractures causes by temperature swings

Release of methane to atmosphere Fire/Explosion CO2 release to atmosphere CO2 feed to algae farm partially disrupted

Pressure relief valve Shielding from external impacts Annual tower scans to detect packing abnormalities Inspection of scrubber shell to find potential leak points/corrosion


6.2 What if the scrubber reaches too high a temperature?

Upstream heat exchanger failure Furnace upset Fouling in upstream exchangers Upstream exchangers bypassed What if the scrubber loses water feed?

Possible damage to tower packing Inadvertent reactions between vapor and water Loss of product quality

Temperature sensors hardwired to quench system Heat resistant packing Regular cleaning of exchangers

6.3 What if the scrubber loses water feed?

P-201 A/B failure Process water intake obstruction Inadvertent valve closure Leak in P-201 A/B discharge line

Loss of CO2 feed to algae farm Loss of product quality

Spare pump Regular line inspections to look for leaks/corrosion Low flow alarms wired to control house

6.4 What if the scrubber loses packing material?

Improper securing of packing Mechanical impact Packing degradation by high temperatures What if the scrubber loses water feed?

Loss of CO2 feed to algae farm Loss of product quality

Shielding from external impacts Heat resistant packing Maintenance of packing retention systems as needed

64

Company:

Kazaam, LLC.



Method: What-if Type: Scrubber Design Intent: Removes CO2 from flue gas

Number: S-201


No.: 7 Description: S-301 Flue Gas Scrubber


7.1 What if the scrubber leaks or ruptured?

External impact Blockage of tower packing Stress fractures causes by temperature swings Corrosion

CO2 release to atmosphere CO2 feed to algae farm disrupted

Pressure relief valve Shielding from external impacts Annual tower scans to detect packing abnormalities Inspection of scrubber shell to find potential leak points/corrosion

7.2 What if the scrubber reaches too high a temperature?

Upstream heat exchanger failure Furnace upset Fouling in upstream exchangers Upstream exchangers bypassed What if the scrubber loses water feed?

Possible damage to tower packing Inadvertent reactions between vapor and water Loss of product quality

Temperature sensors hardwired to quench system Heat resistant packing Regular cleaning of exchangers

7.3 What if the scrubber loses water feed?

P-201 A/B failure Process water intake obstruction Inadvertent valve closure

Loss of CO2 feed to algae farm CO2 release to atmosphere

Spare pump Regular line inspections to look for leaks/corrosion Low flow alarms wired to control house

7.4 What if the scrubber loses packing material?

Improper securing of packing Mechanical impact Packing degradation by high temperatures What if the scrubber loses water feed?

Loss of CO2 feed to algae farm CO2 release to atmosphere

Shielding from external impacts Heat resistant packing Maintenance of packing retention systems as needed

65

Company:

Kazaam, LLC.



Method: What-if Type: Compressor Design Intent: To compress methane to CNG

Number: P-202


No.: 8 Description: P-402 A/B Methane Compressors


8.1 What if liquid enters the compressor?

Entrainment from S-201 Condensation of water vapor in suction lines What if the compressor trips offline?

Compressor trip Loss of discharge temperature leads to hydrogen in CNG product and poisoning of S-401 Loss of containment Fire/explosion

Knockout drum on suction line Heat tracing on suction line High level alarm on knockout drum hardwired to trip


8.2 What if compressor discharge pressure increases?

Blocked discharge valve Compressor over-speed Ignition in compressor caused by air leak into compressor

Downstream overpressure leads to S-401 failure and product quality loss Line rupture/leakage Loss of containment Fire/explosion

Suction to discharge relief valve Minimum flow loop Compressor casing design pressure exceeds the maximum suction pressure plus the compressor shutoff pressure Compressor design such that air cannot enter product stream


8.3 What if compressor overheats?

Upstream upset causing compressor to run dry Excessive upstream temperature Coolers are bypassed or lose flow What if compressor discharge pressure increases?

Mechanical failure of compressor High discharge temperature yields off-sped product

Temperature sensor wired to high pressure trip Synthetic nonflammable lubricants Low flow alarms

8.4 What if compressor trips offline?

Electrical loss Mechanical failure of some part of the compressor

Reverse flow and rotation Backup power source Utilization of spare compressor Check valve in compressor discharge

66

Company:

Kazaam, LLC.



Method: What-if Type: Membrane Design Intent: Removal of H2 from final product

Number: S-202


No.: 9 Description: S-401 Hydrogen Membrane


9.1 What if the membrane ruptures?

P-401 A/B overpressure Corrosion Fatigue caused by temperature swings P-402 A/B over-speed induced greater vacuum What if the membrane chamber leaks or ruptures?

Loss of hydrogen separation Product quality loss Methane/hydrogen release through hydrogen product line Fire/Explosion

Pressure sensors hardwired to isolation system Pressure relief valves Bypass system to equalize pressure


9.2 What if the membrane is exposed to too low a temperature?

P-401 A/B failure

Hydrogen poisoning of membrane Loss of hydrogen separation Loss of product quality

Low temperature alarms wired to control room Emergency high pressure steam heating system

9.3 What if the membrane chamber leaks or ruptures?

P-402 A/B failure P-401 A/B overpressure Exit valve blocked What if the membrane ruptures?

Loss of containment Fire/Explosion Loss of ability to remove hydrogen Product loss

Nondestructive testing Deluge system Pressure relief valve High pressure alarm wired to control house Isolation system


67

Company:

Kazaam, LLC.



Method: What-if Type: Heat Exchanger Design Intent: Remove heat from brine solution

Number: E-201


No.: 10 Description: E-301, E-302 Brine Coolers



Corrosion by exposure to brine Stress cracking Weakened/melted by exposure to high temperature Tube cracks because of differential thermal expansion What if the tubes are subjected to excessive pressure?

Loss of containment Release of hot, high pressure product to environment Further corrosion downstream

Corrosion allowance Periodic nondestructive testing


S-101 upset Tubes blocked during startup Heat applied with tubes blocked in V-301 blockage

Tube leak or rupture Loss of containment Release of hot, high pressure product to environment

Pressure relief valves Low flow alarms integrated into control room Exchanger bypass system Incorporation of semi-automated back-flushing system

10.3 What if the exchanger shell were over-pressured?

P-301 A/B deadheaded Blocked exit valve What if the tubes are subjected to excessive pressure?

Loss of containment Release of hot, high pressure product to environment

Pressure relief valves High pressure alarms wired to control room P-301 A/B over-speed trip

10.4 What if the exchanger loses cooling stream feed?

P-301 A/B trip Cooling water source disrupted Valve closure

Brine produced at too high a temperature, killing algae feedstock in the algae farm Overpressure of E-302 tubes

A/B redundant pumps Temperature alarms wired to control room Low flow alarms wired to control room Emergency cooling water source

68

Company:

Kazaam, LLC.



Method: What-if Type: Pump Design Intent: Provide water for heat and CO2 removal

Number: P-201, P-203


No.: 11 Description: P-201 A/B Water Pumps


11.1 What if the pump trips offline?

Loss of electricity Mechanical failure

Loss of process water flow to scrubbers Overheating of CNG product Loss of CO2 scrubbing capacity, inducing product quality loss Release of hot flue gas to atmosphere Packing damage in S-301

Backup power source Spare pump Low flow alarm Pump run light

11.2 What if the pump is subjected to high temperature?

Lube oil coolers are bypassed or lose flow Pump run dry What if the pump overpressures?

High temperature in discharge stream Loss of CO2 scrubbing Product quality loss

Low flow alarms wired to control house Temperature sensors hardwired to lube oil coolers Spare pump

11.3 What if the pump overpressures?

Pump over-speed Leakage through the check valve of the parallel standby pump Blocked discharge valve

Over-pressurization of CO2 scrubbers Loss of containment in S-201 Methane discharge to atmosphere Fire/explosion

Suction to discharge relief valve Minimum flow loop


11.4 What if the pump leaks or ruptures?

Corrosion caused by contaminants Stress corrosion cracking Gasket leak Drain valve open What if the pump trips offline?

Water release to environment Loss of feed to scrubbers Loss of CO2 scrubbing capacity, inducing product quality loss Release of hot flue gas to atmosphere Packing damage in S-301

Corrosion allowance Periodic non-destructive testing Stress relief Provision for flushing out all pumps during startup and shutdown

69

Company:

Kazaam, LLC.


Site: TBD Unit: System:

Method: What-if Type: Utilities and plant services Design Intent: Provide electric power, control system, plant air/steam/refrigeration. vacuum, fuel oil, natural gas, HVAC service, fire and cooling water, etc. to the facility

Number:


No.: 12 Description: Utilities and plant services

Item What if...? Root Causes/Related Questions

Consequences Safeguards Action Items

12.1 What if electric power were lost momentarily or longer?

What if high pressure steam were lost? Severed Cable Lightning strike Offsite utility power loss Overload Transformer fire Turbogenerator trip What if cooling water was lost?

Process stream backflow Loss of nighttime lighting Loss of control system Potential furnace upset

Alternate power source Breakers and protective logic Emergency shutdown procedures Backup power generators Status panel in the motor control center

12.2 What if the control system (DCS, PLC, etc.) were lost?

What if electric power were lost momentarily or longer?

Loss of controls inside the control room Process must by operated directly

All DCS-operated valves stay in their last valid position to keep the unit stable Backup control modules Backup power supply

12.3 What if cooling water was lost? Cooling water pump shuts off Fouling in the cooling water system Header rupture Low level in the cooling tower reservoir What if electric power were lost momentarily or longer?

Potential compressor damage Leaks on pumps with water cooled seals Overheating of process stream

Pressure indication and low pressure alarm Flow indication and low flow alarm Temperature indication and high temperature alarm Parallel backup pumps

12.4 What if fire suppression (water, carbon dioxide, Halon, etc.) were lost?

What if high pressure steam were lost? Pump tripping off Debris plugging intake

Loss of firefighting capability Possible explosion due to unresolved fire

Intake screens Parallel backup pumps Pressure indication and low pressure alarm

70

No.: 12 Description: Utilities and plant services

Item What if...? Root Causes/Related Questions

Consequences Safeguards Action Items

Fouling in the water system Header rupture Low level in the reservoir

12.5 What if there was inadequate drainage? Improper grade/slope Inadequate drain pipe diameter Sand/gravel/shell accumulation Sludge accumulation Collapsed or plugged piping

Drainage backup Flooding Slip hazards

Periodic cleaning

12.6 What if the flare or thermal oxidizer were lost?

Burner control system malfunction High level in knockout pot High wind Loss of assist gas What if fuel gas were lost?

Potential release of unburned methane and carbon dioxide

Independent pilot system Infrared heat sensors Natural gas backup to purge gas Video monitoring of flare Opacity monitoring of stack exhaust

12.7 What if nighttime lighting were lost? What if electric power were lost momentarily or longer? Lightning strike Offsite utility power loss Overload

Potential personnel injuries (falls, etc.) Inability to see equipment clearly

Battery power for selected lights Emergency lighting circuit Flashlights

12.8 Plant people and contractors are not adequately trained?

Inadequate training Lack of organized records maintenance Updates to computer software for process monitoring

Potential personnel injuries Potential equipment failures

Initial safety training with required refresher courses as needed Required work permits

71

Economic Analysis

Statewide Energy Markets

In addition to completing a full technical review, a thorough economic analysis of the

project was performed. To fully understand the economic impact of this process, the

Arizona natural gas market must first be analyzed. In the year 2009, a total of 32.4 billion

scf of natural gas were purchased by commercial consumers (45). As described in

Appendix A-14, 5,865,000 scf per year of CNG will be produced by this process,

representing 0.018% of the Arizona commercial market. This plant will have a very small

impact in this market; therefore, this plant is not expected to have any measurable impact

on the market price of natural gas.

Hydrogen is also produced by this process. The hydrogen market in the US is estimated

to be $1,600,000,000 (46) every year. As shown below, the 1690.68 gge of hydrogen

produced annually by this process has a value of $8500. This amount of hydrogen

amounts to 0.0005% of the American hydrogen market. This plant will have a very small

impact in this market; therefore, this plant is not expected to have any measurable impact

on the market price of hydrogen.

Sales Revenue

After defining the energy markets in Arizona, the next logical step is to find the annual

revenue this process can generate selling fuel in this market. Yearly production, market

prices and anticipated revenue are all summarized in Table 25.

72

Table 25: Yearly Production and Sales Revenue

Product Yearly Production Market Price Sales Revenue

Natural gas 5,865,000 scf $10.93 / 1000 scf(47)

$64,100

Hydrogen 1700 gge $5 / gge(48)

$8,500

Total: $72,600

Total Capital Investment and Operating Costs

After defining projected sales revenue, defining the costs involved in production of the

product was the next step. The total installed costs for all major pieces of equipment were

found and are summarized in Table 26. Calculations justifying all equipment costs are

found throughout Appendix A and in Appendix B-1. All dollar amounts are in 2010

dollars.

Table 26: Total installation Costs for Major Equipment

Name Cost Name Cost Name Cost

D-101 171900 E-402 4100

P-301

A/B 900

E-101 352000 E-403 2100

P-401

A/B 87000

E-102 5400 E-404 2000

P-402

A/B 5500.00

E-201 34700 F-101 101900 R-101 2802200

E-202 3100 P-101 A/B 410000 S-101 2236000

E-301 7600 P-102 A/B 1800 S-201 134400

E-302 2600 P-201 A/B 9400 S-301 94000

E-401 4800 P-202 A/B 14800 S-401 52400

Summing these numbers up, the cost of the equipment is $6,540,000. The other major

item to be purchased upfront is the initial catalyst charge. This has a price of $4,500,000.

As demonstrated in Appendix B-2, these purchase costs were used to find the total capital

investment required to build the plant, CTCI. The CTCI was found to be $18,200,000. As

shown in Appendix B-3, the value for CTCI was used to approximate the annual cost of

manufacturing (COM). The annual COM was found to be $7,230,000.

73

These numbers were used to find the net present value (NPV) of the plant. Table 27

summarizes the factors involved in computing NPV as well as the cumulative NPV. To

arrive at these numbers, a one year construction time was assumed. Depreciation was

modeled assuming no salvage effort will be made; the depreciation was modeled

according to the MACRS model on a seven year basis (18). The depreciation scheme

here is not identical to the depreciation scheme used in calculating COM, as justified in

Appendix B-4. Income taxes were not taken out of the net earnings because no profit is

turned. As explained in Appendix F, an interest rate of 30% was used to calculate the

discounted cash flow because the process is untested and thus inherently risky. Detailed

calculations for all values can be found in Appendix B-4.

74

Table 27: Net Present Value of the Plant

Year Investment Ctdc Investment Cwc D COM-D S Net Earnings Discounted Cash Flow Cash Flow (PV) Cum PV

2010 (14,458,476.54) (602,150.47) (15,060,627.01) (15,060,627.01) (15,060,627.01)

2011 2,066,116.30 6,794,942.86 72,560.67 (8,788,498.49) (6,722,382.19) (5,171,063.23) (20,231,690.24)

2012 3,540,880.90 6,794,942.86 72,560.67 (10,263,263.10) (6,722,382.19) (3,977,740.94) (24,209,431.18)

2013 2,528,787.55 6,794,942.86 72,560.67 (9,251,169.74) (6,722,382.19) (3,059,800.73) (27,269,231.91)

2014 1,805,863.72 6,794,942.86 72,560.67 (8,528,245.91) (6,722,382.19) (2,353,692.87) (29,622,924.77)

2015 1,291,141.95 6,794,942.86 72,560.67 (8,013,524.15) (6,722,382.19) (1,810,532.97) (31,433,457.75)

2016 1,289,696.11 6,794,942.86 72,560.67 (8,012,078.30) (6,722,382.19) (1,392,717.67) (32,826,175.42)

2017 1,291,141.95 6,794,942.86 72,560.67 (8,013,524.15) (6,722,382.19) (1,071,321.29) (33,897,496.71)

2018 644,848.05 6,794,942.86 72,560.67 (7,367,230.25) (6,722,382.19) (824,093.30) (34,721,590.00)

2019 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (633,917.92) (35,355,507.92)

2020 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (487,629.17) (35,843,137.09)

2021 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (375,099.36) (36,218,236.46)

2022 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (288,537.97) (36,506,774.43)

2023 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (221,952.28) (36,728,726.71)

2024 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (170,732.53) (36,899,459.24)

2025 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (131,332.71) (37,030,791.95)

2026 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (101,025.16) (37,131,817.11)

2027 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (77,711.66) (37,209,528.78)

2028 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (59,778.20) (37,269,306.98)

2029 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (45,983.23) (37,315,290.22)

2030 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (35,371.72) (37,350,661.93)

2031 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (27,209.01) (37,377,870.95)

2032 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (20,930.01) (37,398,800.96)

2033 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (16,100.01) (37,414,900.97)

2034 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (12,384.62) (37,427,285.59)

2035 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (9,526.63) (37,436,812.22)

2036 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (7,328.18) (37,444,140.40)

2037 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (5,637.06) (37,449,777.46)

2038 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (4,336.20) (37,454,113.66)

2039 6,794,942.86 72,560.67 (6,722,382.19) (6,722,382.19) (3,335.54) (37,457,449.20)

2040 602,150.47 6,794,942.86 72,560.67 (6,722,382.19) (6,120,231.72) (2,335.97) (37,459,785.17)

75

As can be seen in Table 27, above, the NPV of the process is ($37,500,000), making this

process very unprofitable.

Sensitivity Analysis

The spreadsheets described in Appendices B-3 and B-4 were used to analyze the impact

of fluctuations in the prices of utilities and feedstocks on the overall economics of the

process. Prices for these items were manually changed in the COM calculations in

Appendix B-3; the change in NPV using the spreadsheet in Appendix B-4 was noted in

order to determine how sensitive the process is to changes in input costs.

The price of the algae feedstock was first increased; for comparison’s sake, the price of

the algae was increased to $20 per ton, an increase by a factor of 20. This resulted in the

NPV decreasing to ($37,600,000), a decrease of 0.2%. Because of the small magnitude of

this change, the process economics are not considered to be sensitive to changes in

feedstock cost.

The price of the utilities (process and cooling water, electricity) was then analyzed. As a

basis of analysis, the price of utilities was increased by a factor of 2 to $600,000 per year.

This resulted in the NPV decreasing to ($38,000,000), a decrease of 1%. Because of the

small magnitude of this change, the process economics are not considered to be sensitive

to changes in utility costs. Analysis of the impacts of catalyst cost changes are discussed

later in this section.

Road to Profitability, Market Changes

In order for a project to be considered feasible, the NPV must be greater than or equal to

zero (18). An analysis was performed to see what the selling price of natural gas would

76

have to be in order to reach a NPV of zero, assuming that ratio of the prices of natural gas

and hydrogen remains constant and that no operating costs change. To reach a NPV of

zero, the price of natural gas would have to be $1700 per 1000 scf, an increase over

today’s prices by a factor of 156, an event that will likely never happen in the foreseeable

future. This factor was calculated using the spreadsheet described in Appendix B-4.

Road to Profitability, Catalyst Changes

Looking at the process, there are a few areas where attention should be focused in order

to reduce the high cost associated with production. The biggest opportunity to reduce cost

comes with the catalyst. Because the carbon support structure of the catalyst degrades

over time in the supercritical environment (7), the catalyst will have to be replaced every

year. This catalyst degradation is unavoidable because the reaction must occur in

supercritical conditions. Replacing the catalyst increases the annual operating cost of the

process by $4,500,000, well over half the annual operating budget. If the catalyst could

be reduced in cost to $100,000 per load, the CTCI would decrease to $11,200,000 and the

NPV would increase to ($15,300,000), as calculated using the spreadsheets shown in

Appendix B-4. The $22.2MM savings in NPV is significant, showing that finding an

inexpensive catalyst solution should be a top priority. This could include finding a new,

cheaper catalyst or finding a way to reduce the cost of the Ru/C catalyst.

There are only two metals that are capable of catalyzing the methanation reaction:

ruthenium and rhodium (49). Considering rhodium has a cost of $80,000 per kg (50), it is

not considered a viable option for a large quantity of catalyst. Therefore, an alternative

ruthenium catalyst must be found. Considering that the Ru/C catalyst included in the

process is 5% ruthenium by mass, the total mass of ruthenium in the reactor is 38 kg.

This has a market value of $250,000 (51). Assuming that this amount of ruthenium could

77

be affixed to a non-degradable support structure with negligible cost, the CTCI would

decrease to $11,400,000 and the NPV would increase to ($15,200,000), as calculated

using the spreadsheets shown in Appendix B-4. The $22.3MM savings in NPV is

significant, showing that finding an inexpensive, non-degradable catalyst solution should

be a top priority. One support structure that shows promise is zirconia; zirconia does not

show signs of degradation in supercritical environments (49). This catalyst would not

have to be purchased every year. Therefore, effort should be directed towards developing

a Ru/zirconia catalyst on a commercial scale with minimal cost.

Road to Profitability, Equipment Changes

Another opportunity to reduce costs comes with R-101 and S-101. These vessels, priced

at $2.8MM and $2.2MM, respectively, are more than four times more expensive than the

next most expensive piece of equipment, P-101 A/B. These vessels must operate at

supercritical conditions because the reaction requires the algae slurry to be supercritical.

If the cost of each of these vessels could be decreased to $100,000, the CTCI would

decrease to $10,800,000 and the NPV would increase to ($29,600,000). The $9MM in

savings shows that special attention should be focused on these vessels to find ways to

reduce their wall thicknesses; reducing the amount of material required to make these

vessels could significantly reduce their costs.

If R-101 and S-101 were reduced in cost to $100,000 each and a Ru/zirconia catalyst

were implemented with only a $250,000 initial purchase cost, the NPV would increase to

($7,300,000), an increase of $30.2 MM. This would significantly improve likelihood of

this project becoming economically viable.

78

Conclusions and Recommendations

The design presented in this report is expected to have minimal impact on the Arizona

energy market. With an annual output of 5,865,000 scf of CNG, this process captures

only 0.018% of the state’s commercial natural gas sales. Seeing as the design has a NPV

of ($37,500,000), it is recommended that this design is shelved until technological

improvements are made and economic conditions improve.

Several innovations were included in the design of the process; while the assumptions

relating to these innovations appear reasonable, there are several opportunities for future

work to test the accuracy of the design. The design centers around R-101, the

supercritical gasification and methanation reactor. Until now this technology has only

been tested on a lab scale. It is recommended that a pilot scale reactor be built in order to

confirm the feasibility of this reactor on the scale required for this process. A related

innovation was included in S-101, the supercritical gravitational salt separator. This

technology has also never been tested on an industrial scale. To ensure functionality, the

design team recommends that a pilot scale salt separator be built and tested as well.

Advancements in steel metallurgy should also be tracked to see if a RFGS becomes a

feasible alternative. A final major technological innovation was included in S-401, the

hydrogen separation membrane. The palladium/ceramic membranes included in the

design have only been tested on a lab scale; it is recommended that this technology be

tested on a pilot scale to ensure functionality in the final process.

Other opportunities exist to bring down the cost of the plant. Accounting for a $4.5MM

yearly expenditure, catalyst costs severely depress the NPV. Work should be directed at

creating a low cost Ru/zirconia catalyst that retains the functionality of ruthenium but

79

does not degrade like a carbon support structure. Pursuing this path has the potential to

increase the NPV by $22.3MM. Further costs could be cut by reducing the significant

costs of R-101 and S-101. Both vessels currently have large wall thicknesses due to the

supercritical state their contents are required to be at; future work should be directed at

reengineering the structure, shape and supports of the vessels to reduce their wall

thicknesses. If the vessels could be engineered such that the cost of each was $100,000, it

would increase the NPV by $9MM. If solutions to the high costs of the catalysts and the

high pressure vessels can be found, the NPV of the process would increase to ($7.3MM),

an increase of $32.2MM.

Economically speaking, the financial relationship this plant would have with the

associated algae farm must be more clearly defined. Currently, nutrient salts and carbon

dioxide are recycled to the algae farm in order to reduce the purchase price of the algae

feedstock. This decision should be reevaluated in later phases of engineering to

investigate whether selling the salts on the open market would be more profitable.

Included in this analysis should be the investigation of a brine purifying unit that could

separate the different salts so that a pure product can be sold.

Due to the current market of CNG and the high costs of production associated with

producing CNG from algae, it is the recommendation of this report that the proposed

design not be used at this time. However, if in the future the described technological

advances can be made in and the market cost of natural gas increases, this project can be

reconsidered as the process is both straightforward in design and has a minimal impact on

the environment.

80

Nomenclature

Symbol Description Units

a Material Construction Factor none

A Area length2

b Material Construction Factor none

CAlloc Allocated Costs dollars

CB Base Cost dollars

CDPI Direct Permanent Investment dollars

Cp Specific Heat

CP f.o.b. Purchase Cost dollars

CPL

Cost of Vessel Platforms and

Ladders dollars

CV Cost of Vessel dollars

CTBM Total Bare Module Cost dollars

CTCI Total Capital Investment dollars

CTDC Total Depreciable Capital dollars

Di Inner Diameter of Vessel length

E Fractional Weld Efficiency none

F Molar Flux ondsarea

moles

sec

FL Tube Length Correction Factor none

FM Material Factor none

FP Pressure Factor None

FTBM Total Bare Module Factor none

g Gravitational Constant

h Specific Enthalpy

Standard Enthalpy of Formation H Enthalpy energy

head Head of Water I Plant Cost Index none

k Specific Heat Ratio none

L Height length

m Mass mass

Mass Flow Rate

81

n Moles Mol

Molar Flow Rate

P Pressure

Pc Power Consumption energy (hp)

PB Break Horsepower energy (hp)

Pd Internal Design Gauge Pressure

Po Operating Pressure

Q Energy Flow Rate s Pump Size Factor None

S

Maximum Allowable Vessel

Stress None

SC Flow rate of cooling water Gpm

SP Flow rate of process water Gpm

tp Vessel Wall Thickness Inches

T Temperature oC,

oF, K

Log Mean Temperature Difference oC,

oF, K

Average Temperature Difference oC,

oF, K

U Overall Heat Transfer Coefficient

V Volume length3

Volumetric Flow Rate

Average Fluid Velocity W Weight Lbs

Wp Pump Work Z Height Above Datum Plane Length

Symbol Units

Density

η

Mechanical Efficiency of the

Pump None

ηP Pump Motor Fractional Efficiency None

ηM Pump Fractional Efficiency None

τ Residence Time Time

82

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87

Appendices

Appendix A: Final Calculations

A-1: E-101 Preheater Calculations

The minimum temperature approach was set to be 28 oC, the recommended minimum

approach for temperatures up to 150 oC (18).

CCCTTT ooo

approach 37228400min_24

Note: Assume almost all of the energy transfer is between the water in the streams

(because it is approximately 97.8% H2O by mass) and that the stream is entirely water

Note: Literature values used for the specific enthalpies, h (49). Q is the energy flow rate

and 𝑚 is the mass flow rate of the stream.

hr

Btu

hr

kJQ

kg

kJ

kg

kJ

hr

kgQ

hhmhmQ

32.788291816.8310055

23.1743.17604.5239

)( 14

kg

kJh

hkg

kJ

hr

kg

hr

kg

hhmhmQ

43.511

50.20974.523916.8310055

)(

3

3

32

2 MPa

CT o

975.30

4002

MPa

CT o

0.31

351 MPa

CT o

995.30

372?4 4 1

MPa

CT o

970.30

117?3 3

At P = 30.975 MPa, T3 = 117 oC (52)

88

E-101 Cost Calculations

hrftF

BtuU

F

FF

FF

FFFF

T

T

TTT

TUAQ

o

o

oo

oo

oooo

lm

lm

2

1

2

12

200

46.90

956.242

6.701752ln

)956.242()6.701752(

ln

Note: U was determined based on a conservative estimate for water-water systems (18).

Note: The heat transfer design equation was used to size the heat exchanger (20).

2

2

71.435

46.90200

32.7882918

ft

FhrftF

Btu

hr

Btu

TU

QA

o

o

lm

Note: It was determined a floating head counter current shell and tube heat exchanger

worked best for this application (18).

71.16296$

2)ln(09005.0)ln(8709.0667.11 AA

B eC

Where CB is the base cost of the exchanger and A is the area in ft2.

BLMPP CFFFC

Where FP is the pressure factor, FM is the material factor, FL is the tube length correction

factor.

For carbon steel: 1100

0100

0

AAaF

b

M

Note: A conservative estimate of 8 ft was assumed for the tube length.

25.1LF

2.5

49.4481696.14325.101

696.1431000

1000017.0

100018.09803.0

2

P

P

F

psigpsikPa

psikPaP

PPF

61.105929$ BCMPP CFFFC

Where:

P = shell side pressure in psig

89

A-2: E-102 Cooler Calculations

Note: T1 determined from a conservative estimate for cooling water and T4 determined

from an environmental limit (18).

Component m (kg/hr) n (mol/hr) 2h

(kJ/mol) 3h

(kJ/mol)

)ˆˆ( 32 hhmQ

(kJ/hr)

CO2 81.702 1856.44 26.040 22.873 5879.35

CO 0.211 7.53 15.221 12.904 17.45

CH4 25.966 1618.63 18.26 15.221 4919.02

C2H6 1.936 64.39 25.697 20.925 307.27

C3H8 1.001 22.70 35.793 20.951 336.91

H2 0.434 215.29 10.747 8.448 494.95

H2O 4958.333 309896 9.231 3.249 1853797.87

1865752.82

Note: The energy balance above verifies previous assumption that the majority of the

heat transfer occurs in the water as it accounts for 99.3% of the energy.

The amount of cooling water required was determined through an energy balance where

the enthalpies were determined from literature values (49).

hr

m

kg

m

hr

kgv

hr

kg

g

kg

mol

g

hr

molm

hr

moln

mol

kJn

hr

kJ

hhnQ

OH

OH

OH

OH

OH

33

14

59.230010116.09.23321

9.233211000

1639.1457619

39.1457619

)420.2700.3(82.1865752

)ˆˆ(

2

2

2

2

2

MPa

CT o

11.0

321

MPa

CT o

3.0

1222

MPa

CT o

295.0

433

MPa

CT o

105.0

494 1

2

3

4

90


hrftF

BtuU

F

FF

FF

FFFF

T

T

TTT

TUAQ

o

o

oo

oo

oooo

lm

lm

2

1

2

12

200

7.56

90110

120252ln

)90110()120252(

ln



hr

Btu

J

Btu

kJ

J

hr

kgQ 13.1769853

10486.9100082.1865752

4

2

2

07.156

7.56200

13.1769853

ft

FhrftF

Btu

hr

Btu

TU

QA

o

o

lm

Note: It was determined a double pipe heat exchanger worked best for this application

(18).

94.2846$)ln(16.0146.7 A

B eC


BMPP CFFC


0100

0

AAaF

b

M

For pressures less than 600 psig, FP = 1

94.2846$ BMPP CFFC

91

A-3: F-101 Superheater Calculations

The temperature of the algae slurry exiting the superheater must be 400 °C. Note:

Assume almost all of the energy transfer is between the water in the streams (because it is

approximately 97.8% H2O by mass) and that the stream is entirely water. The amount of

heat Q required to be absorbed by the slurry is given by

TCmQ P.

)3600

1)(372400)(5000)(13.7(

s

hrCC

hr

kg

kgK

kJQ

kWQ 2.277

Where the average heat capacity of water at this pressure and range of temperatures is

7.13 (24)

This heat is provided by the combustion of natural gas in the furnace. The preheated air

and methane mix and come to an equilibrium temperature given by the following energy

balance.

)131)(05.18)(34.2()465)(5.629)(84.1)(013.0()5.629)(465)(01.1( CThr

kg

kgK

kJTC

hr

kg

kgK

kJ

kgair

kgwater

hr

kgTC

kgK

kJ

CT 445

The heat released by the combustion of methane is 55687.5 kJ/kg combusted (53).

Therefore, the total heat released is

HmQ

)3600

1)(5.55687)(05.18(

s

hr

kg

kJ

hr

kgQ

s

kJQ 2.279

The heat capacity of the combusted material is assumed to have the same heat capacity as

air. It has a moisture content of 0.149 kg water/kg air. The temperature of the recently

combusted air is shown to be

)445)(5.647)(84.1)(149.0()5.647)(445)(01.1(375.1005159 CThr

kg

kgK

kJ

kgair

kgwater

hr

kgCT

kgK

kJ

hr

kJ

CT 1702

Heat is then transferred to the algae mix. An approach temperature of 93 °C is used (18).

)4931702)(5.647)(84.1)(149.0()5.647)(4931702)(01.1( CChr

kg

kgK

kJ

kgair

kgwater

hr

kgCC

kgK

kJQ

kWQ 2.277

F-101 Cost Calculations

Note: F-101 was priced as a fired heater (18)

))ln(766.032325.0( Q

B eC

92

Where Q is in units of hr

Btu

BMPP

B

CFFC

C

hr

Btu

s

JkWQ

10.53063$

9466272772002.277

FM = 1 because of carbon steel

FP = 1 because P < 500 psig

CP = $53063.10

93

A-4: P-101 A/B Algae Pump Calculations

2

2VgZ

Phead

Note: Assume the difference between the heights of the suction and discharge

connections is negligible (i.e. 0Z )

Assume: m constant (same into pump as out of pump)

AVm

Therefore: A

mV

Note: Assume A and are the same at the inlet and outlet of the pump. Therefore, V

constant

kg

J

m

kg

PaPaPPhead 1.31626

977

)10013.1()101.3(

3

57

12

Note: was determined by assuming it to be water at 25oC and 101.3kPa (49)

kg

Jkg

J

headWp 1.42168

75.0

1.31626

where:

Wp = pump work

mechanical efficiency of the pump

A conservative estimate of 0.75 was made for (18)

hpWs

J

s

hr

hr

kg

kg

JmWP pB 3.826137161371

3600

14.52391.42168

where: PB = break horsepower

P-101 A/B Cost Calculations

ftftm

ft

s

m

kg

J

g

head32001.1058528.3

8.9

1.31626

2

Where: g is the gravitational constant.

where:

head = total head

V = average fluid velocity

Z = height above datum plane

P = pressure

= density of fluid

A = area of pipe

94

min5000

min85.23

17.264977

min60

14.5293

3

3

galgal

gal

m

m

kg

hr

hr

kg

mV

Where: V is the volumetric flow rate

Note: P-101A/B was determined to be a reciprocating pump due the large head (10585 ft)

it is required to generate (18)

Note: Cost determined from empirical correlations for reciprocating pumps (18)

20.29948$2)ln(06718.0)ln(26986.08103.7(

BB PP

B eC

CP = FMCB

For carbon steel, FM = 1.50

CP = $44922.40

Motor Cost Calculations:

mP

C

g

headV

P

33000

Where

g

head, is in feet.

9.0)ln(00182.0)ln(0319.08.0

32.0)ln(01199.0)ln(24015.0316.0

2

2

BBm

P

PP

VV

Where: ηP is the pump fractional efficiency and ηM is the pump motor fractional

efficiency

40.14328$)40.14238)($0.1(

40.14328$

00.221)9.0)(32.0(33000

32.81.10585min

85.23

432 )ln(0035549.0)ln(02828.0)ln(053255.0)ln(13141.08259.5

BTP

PPPP

B

C

CFC

eC

hpgal

lbft

gal

P

CCCC

Note: The motor was chosen to be a drip-proof enclosure, therefore, FT = 1 (18).

Thus,

The total cost for 1 pump and motor: $59250.80

Where:

PC = motor power consumption

= 8.32 lb/gal

𝑉 is in units of gal/min

95

The total cost for 2 pumps and motors: $118501.60

A-5: P-102A/B Cooling Water Pump Calculations

hpWs

J

s

hr

hr

kg

kg

JmWP

kg

Jkg

J

headW

kg

J

m

kg

PaPaPPH

PB

P

073.09.769.763600

9.233219.11

9.1175.0

9.8

9.8

977

)10013.1()101.1(

3

55

12

P-102A/B Cost Calculations

ftftm

ft

s

m

kg

J

g

head320098.328.3

8.9

9.11

2

min5000

min1.105

17.264977

min60

19.23321

3

3

galgal

gal

m

m

kg

hr

hr

kg

mV

Note: P-102A/B was determined to be a centrifugal pump due its small head (3.98 ft) and

flow rate

min1.105

gal (18)

Note: Cost determined from empirical correlations for a centrifugal pump (18)

5.0

g

headVs

Where: s is the pump sizing factor

5.05.075.20998.3

min1.105 ftgpmft

gals

Note: The size factor for this pump was too small to use empirical relationships to cost

the pump. Therefore, a pump was located that could meet the technical requirements of

the process was found.

96

A 2 hp straight center discharge pump with an optimum flow rate of 101 min

gal was

located with a cost of $900.00, including an assumed shipping and handling rate (23).

A-6: R-101 Gasification and Methanation Reactor Calculations

Note: The volume ratios of the gases produced in the reactor were determined from a

literature source (7).

Note: The gases were assumed to behave ideally

Component Vol % Mol % MW (g/mol) Mass

(100 mol sample) Mass %

CH4 42.7 42.7 16.042 684.993 23.34

C2H6 1.7 1.7 30.068 51.116 1.74

C3H8 0.6 0.6 44.094 26.456 0.90

CO2 49.0 49.0 44.010 2156.490 73.44

CO 0.2 0.2 28.010 5.602 0.19

H2 5.8 5.8 2.016 11.693 0.39

TOTAL: 100 100 2936.350 100

Volume of Reactor:

Note: The required residence time for the reactor was determined to be one hour (7).

Note: The density of the stream was assumed to have the same density as water at the

same conditions (because the stream is 97.8% water by mass) (49).

Q

V QV

min217.001.13

61.389

583.5069 33

3

m

hr

m

m

kghr

kg

mQ

367.13)217.0min)(63( mV

Mass of catalyst:

Note: The ratio of the mass of catalyst in the reactor to the mass of dry matter was known

to be eight from a literature source (7).

OHCatalystDMreactor VVVV2

Where mcat is the mass of the Ru/C catalyst, mDM is the mass of the algae, and mH2O is the

mass of water in the reactor.

97

OH

OH

catalyst

DM

DM

DM

OH

OH

catalyst

catalyst

DM

DMmmmmmm

m

2

2

2

28

67.13 3

Note: totalm is the total mass of DM and H2O in the reactor

Note: It was assumed that nonadecaneDM at 30.85MPa and 400oC (49)

DMcat mm 8

totaltotaltotal

total

DMDM mm

kg

kgm

m

mm 022.0

583.5069

25.111

totalOH mm 978.02

382.518

m

kgDM (52)

3260

m

kgcatalyst (Note: from the paper that Kyle sent)

361.389

2 m

kgOH (Note: at 30.85MPa and 400

oC (18))

totaltotaltotaltotal m

m

kg

m

m

kg

m

m

kg

mm 00323.0

61.389

978.0

260

022.810

82.518

022.067.13

333

3

kgmtotal 2.4232

Therefore,

kgmmm totalDMcat 87.744)022.0(88

98

T Across the Reactor

The temperature change across the reactor was determined with the aid of the

thermodynamic cycle illustrated below. A basis of one mole was used. Only one of the

methane producing reactions, the conversion of formaldehyde to methane and carbon

dioxide, was evaluated (7). The gases were assumed to behave ideally, and therefore

experienced no enthalpy change with a change in pressure.

Note: The enthalpy changes the formaldehyde, methane, and carbon dioxide experienced

during the temperature changes were determined using the specific heats of the

compounds (54).

kJH

dTTmolH

dTCnH

deformaldehy

C

C

deformaldehy

C

C

pdeformaldehydeformaldehy

o

o

o

o

26.16

)10268.41028.34(1

25

400

53

25

400

kJH

dTTTmolH

dTCnH

methane

C

C

methane

C

C

pmethanemethane

o

o

o

o

65.8

)103361.010469.51031.34(5.0

400

25

2853

400

25

CH2O(g) (400 oC, 30.85 MPa) 0.5 CH4(g) (400

oC, 30.85 MPa) 0.5 CO2(g)

0.5 CH4(g) (400 oC, 1 atm) 0.5 CO2(g)

0.5 CH4(g) (25 oC, 1 atm) 0.5CO2(g)

CH2O(g) (400 oC, 1 atm)

CH2O(g) (25 oC, 1 atm)

C(s,graphite) +H2(g)+0.5O2(g) C(s,graphite) +H2(g)+0.5O2(g)

0H

0H

0H 0H

kJH 26.16

kJH 9.115 kJH 44.37 kJH 76.196

kJH 65.8 kJH 15.8

99

kJH

dTTTmolH

dTCnH

dioxidecarbon

C

C

dioxidecarbon

C

C

pdioxidecarbondioxidecarbon

o

o

o

o

15.8

)10887.210233.41011.36(5.0

400

25

2853

400

25

Note: The enthalpy changes the formaldehyde, methane, and carbon dioxide experienced

during the formation of the compounds were determined using the standard enthalpy of

formation of the compounds (49).

mol

kJH

mol

kJmolH

hnH

deformaldehy

deformaldehy

o

fdeformaldehydeformaldehy deformaldehy

9.115

9.1151

mol

kJH

mol

kJmolH

hnH

methane

methane

o

fmethanemethane methane

44.37

88.775.0

mol

kJH

mol

kJmolH

hnH

dioxidecarbon

dioxidecarbon

o

fdioxidecarbondioxidecarbon methane

76.196

52.3935.0

100

Summing all of the enthalpy changes listed above yielded the net enthalpy change in the

reaction of one mole of formaldehyde to form one half of a mole of methane and one half

of a mole of carbon dioxide.

kJH

kJkJkJkJkJkJH

rxn

rxn

76.117

15.865.876.19644.379.11526.16

Therefore, the production one mole of methane would produce twice the enthalpy value

given above, and multiplying by the mass flow of methane provided the total energy flow

rate generated by the reaction.

442 37.3670

52.23276.117

CHkg

kJ

CHmol

kJ

OCHmol

kJH rxn

hr

kJ

kg

kJ

hr

kgQ 5.9531937.367097.25

This change in enthalpy could then be converted into a temperature change in the reactor

by assuming the water in the reactor absorbed all of the energy.

mol

kJh

hr

kghr

kJ

h

m

Qh

water

water

water

water

48.08467.26

33.4958

133115

mol

kJh

mol

kJ

mol

kJh

hhh

final

final

MPaCofinal

water

water

waterwaterwater

375.38

48.0895.37

85.30,400

Using a literature source, it was determined that the water would rise in temperature by

1.2 oC, exiting the reactor at a temperature of 401.2

oC (49). This small change in

101

temperature was assumed to be negligible as energy losses from the reactor to the

surroundings would likely compensate for the increase in energy due to the reaction.

Other methane and non-methane producing reactions occur in the reactor that are

exothermic, however it was assumed that they do not produce enough heat to drastically

influence the temperature of the stream exiting the reactor.

P Across the Reactor

The pressure of the stream would drop due to friction losses as it passed through the

catalyst in the reactor. However the pressure would increase due to the formation of the

gas molecules. Therefore, the pressure change across the reactor was assumed to be

negligible. In addition, the reaction rates are not affected greatly by small changes in

pressure (7).

Reactor Cost Calculations

Note: Reactor height, L, set to be 3m = 118.08 in

inmD

mD

m

hD

hrV

i

i

ii

86.9441.2

)3(2

67.13

2

2

3

2

2

Note: Empirical correlations used to determine the reactor wall thickness, tp (18).

d

idp

PSE

DPt

2.12

E =1 for carbon steel thicker than 1.25 in (18)

S = 15000 at T = 400oC (18)

Pd = 1.1Po (for Po > 1000 psig)

int

P

psigpsikPa

psikPaP

p

d

o

16.24

67.4927

7.4479696.143.101

696.1430980

where:

Di = inner diameter of the vessel (in)

tp = wall thinkness to withstand the internal pressure (in)

Pd = internal design gauge pressure (psig)

S = maximum allowable stress of the shell material at the

design temperature (psi)

E = fractional weld efficiency

where:

Po = operating pressure (psig)

102

Note: Empirical correlations used to determine the weight of the reactor

lbsin

lbinininininW

tDhtDW pipi

497640284.0)16.24))(86.94(8.008.118)(16.2486.94(

)8.0)((

3

Where 3

284.0in

lb for carbon steel (18)

PLvMP CCFC

Note: The reactor was designed as a vertical vessel

28.8399$)()(8.361

00.634124$

70684.07396.0

))ln(02297.0)ln(18255.00132.7( 2

LDC

eC

iPL

WW

v

For carbon steel, therefore FM = 1

CP = $642523.28

Where: Di and L are in feet

103

A-7: S-101 Salt Separator Calculations

Note: the molecular formula of algae is approximated to be C1.0H1.7O0.48N0.19S0.005 (7)

Note: All nitrogen and sulfur are removed from the algae.

Element Molecular Weight

Mass in 1 mole algae

Percent of total mass

C 12 12 49.6

H 1 1.71 7.1

O 16 7.68 31.7

N 14 2.66 10.9

S 32 0.16 0.5

11 % of the 125 kg algae/hr is removed. This corresponds to 13.75 kg salts precipitated

per hour.

hr

m

m

kghr

kg

mQ

QV

Q

V

3

3

29.13

20.394

4.5239

Therefore,

33

29.1329.13)1( mhr

mhrV

Fluid velocity 222 45.4)19.1( mmrA

s

m

m

s

hr

hr

m

A

QU 0008.0

45.4

)3600

1)(29.13(

3

3

S-101 Cost Calculations

Note: Separator height set to 3 m = 118.08 in

inmD

mD

m

LD

LrV

i

i

ii

48.9338.2

)3(2

29.13

2

2

3

2

2

d

idp

PSE

DPt

2.12

104

E =1 for carbon steel thicker than 1.25 in (18)

S = 15000 at T = 400oC (18)

Pd = 1.1Po (for Po > 1000 psig)

04.4928

04.4480696.143.101

696.1430990

d

o

P

psigpsikPa

psikPaP

int p 13.19

lbstDLtDW pipi 370690)8.0)((

Where 3

284.0in

lb for carbon steel

PLvMP CCFC

Note: S-101 was designed as a vertical vessel

63.8312$)()(8.361

504374$

70684.07396.0

))ln(02297.0)ln(18255.00132.7( 2

LDC

eC

iPL

WW

v

For carbon steel FM = 1

CP = $512686.63

105

A-8: V-101 Calculations

Note: Assume adiabatic expansion

Note: Assume that the expanding (cooling) gas does not affect the temperature of the

system much (because it is only about 2% gas by mass)

Note: Assume that the enthalpy of the system comes from the water only

kg

kJ

kg

kJ

hh OfHOiH

18.51318.513

22

Therefore, Tf = 122oC

Ti=117oC

Pi=36.97MPa

Tf = ?

Pf=0.3MPa

Q=0

@ P = 0.3 MPa, T = 122oC (52)

106

A-9: D-201 Knockout Drum Calculations

Note: A volume of 13.19 m3 was assumed (the same volume as S-101) in order to

achieve proper stream separation

D-201 Cost Calculations


inmD

mD

m

LD

LrV

i

i

ii

48.9338.2

)3(2

29.13

2

2

3

2

2

Note: int p 375.0

for low pressure vessels with diameters between 6 and 8 ft (18).

lbstDLtDW pipi 31.6056)8.0)((

Where 3

284.0in

lb for carbon steel

PLvMP CCFC


63.8312$)()(8.361

60.31105$

70684.07396.0

))ln(02297.0)ln(18255.00132.7( 2

LDC

eC

iPL

WW

v


CP = $39418.23

107

A-10: E-201 Air Preheater Calculations

Air quantity, based on 18.05 kg of natural gas combusted in furnace and double the

required oxygen

kgairkmolair

kg

kmolO

kmolair

kmolNG

kmolO

kg

kmolkgNGQ 5.629)

29)(

21.0)(

4)(

16)(05.18(

2

2

Note: Incoming air was assumed to be saturated and have a temperature of 16 °C and a

pressure of 101.3 kPa. The moisture content of air is 0.013 kg water per kg of air (20).

Note: The heat capacity of water in air is taken to be 1.84 (55)

Note: The approach is set at 28 °C (18)

Air:

)16465)(5.629)(84.1)(013.0()5.629)(16465)(01.1( CChr

kg

kgK

kJ

kgair

kgwater

hr

kgCC

kgK

kJQ

kWQ 1.80

Flue gas:

)493)(55.647)(84.1)(149.0()55.647)(493)(01.1(1.80 TChr

kg

kgK

kJ

kgair

kgwater

hr

kgTC

kgK

kJkWQ

CT 142


hrftF

BtuU

F

FF

FF

FFFF

T

T

TTT

TUAQ

o

o

oo

oo

oooo

lm

lm

2

1

2

12

10

0.117

61288

869919ln

)61288()869919(

ln

Note: U was determined based on a conservative estimate for water-air systems (18).


MPa

CT o

13.0

4931

MPa

CT o

14.0

162

MPa

CT o

135.0

4653

MPa

CT o

125.0

1424 1

2

3

4

108

2

2

8.233

0.11710

273538

ft

FhrftF

Btu

hr

Btu

TU

QA

o

o

lm

Note: It was determined a floating head counter current shell and tube heat exchanger

worked best for this application (18).

90.14702$

2)ln(09005.0)ln(8709.0667.11 AA

B eC


BLMPP CFFFC

Where FP is the pressure factor, FM is the material factor, FL is the tube length correction

factor.


0100

0

AAaF

b

M

Note: A conservative estimate of 8 ft was assumed for the tube length.

25.1LF

FP = 1 for low pressure shell and tube heat exchangers

60.18378$ BCMPP CFFFC

109

A-11: E-202 Methane Preheater Calculations

Note: The approach for the exchanger was set at 11 °C (18)

Note: The heat capacity of natural gas is 2.34 kJ/kgK (56)

Methane:

)24131)(05.18)(34.2( CChr

kg

kgK

kJQ

kWQ 24.1

Flue gas:

)142)(55.647)(84.1)(149.0()142)(55.647)(01.1(24.1 TChr

kg

kgK

kJ

kgair

kgwaterTC

hr

kg

kgK

kJkWQ

CT 136


hrftF

BtuU

F

FF

FF

FFFF

T

T

TTT

TUAQ

o

o

oo

oo

oooo

lm

lm

2

1

2

12

10

7.78

75277

268288ln

)75277()268288(

ln

Note: U was determined based on a conservative estimate for water-air systems (18).


2

2

4.5

7.7810

6.4234

ft

FhrftF

Btu

hr

Btu

TU

QA

o

o

lm

MPa

CT o

125.0

1421

MPa

CT o

14.0

242

MPa

CT o

135.0

1313

MPa

CT o

12.0

1364 1

2

3

4

110


(18).

08.1662$)ln(16.0146.7 A

B eC


BMPP CFFC


0100

0

AAaF

b

M


08.1662$PC

111


hps

J

s

hr

hr

kg

kg

JmWP

kg

Jkg

J

headW

kg

J

m

kg

PaPaPPH

PB

P

94.1176.89033600

04.12447052.257

52.25775.0

14.193

14.193

977

)10013.1()109.2(

3

55

12


ftftm

ft

s

m

kg

J

g

head320064.6428.3

8.9

14.193

2

min5000

min9.560

17.264977

min60

104.124470

3

3

galgal

gal

m

m

kg

hr

hr

kg

mV

Note: P-102A/B was determined to be a centrifugal pump due its small head (64.64 ft)

and flow rate

min9.560

gal (18)


5.0

g

headVs


5.05.058.450964.64

min9.560 ftgpmft

gals

BMTP

ss

B

CFFC

eC

93.4133$2)ln(0519.0)ln(6019.07171.9

For FT, it was determined a 1 stage, 3600 shaft rpm, horizontal split case (HSC), 50-900

gpm, 75 max motor hp best fits the application, leading to FT = 1. For FM cast iron was

assumed.

112

FT = 1

FM = 1

Therefore,

CP = $4133.93

Motor Cost Calculations:

mP

C

g

headV

P

33000

87.0)ln(00182.0)ln(0319.08.0

72.0)ln(01199.0)ln(24015.0316.0

2

2

BBm

P

PP

VV

50.1021$)50.1021)($0.1(

50.1021$

6.14)72.0)(54.0(33000

32.898.3min

9.560

432 )ln(0035549.0)ln(02828.0)ln(053255.0)ln(13141.08259.5

BTP

PPPP

B

C

CFC

eC

hpgal

lbft

gal

P

CCCC

Note: The motor was chosen to be a drip-proof enclosure, therefore, FT = 1 (18).

The total cost for 1 pump and motor: $5155.43

The total cost for 2 pumps and motors: $10310.86

113

A-13: P-202A/B Air Blower Calculations

min4.312

min85.8

101325

)289(314.8min

78.361

min78.361

29min10490

min10490

min49.10

min605.629

33

3

ftm

Pa

KKmol

Pammol

P

RTnV

mol

g

molgn

gkghr

hr

kgm

Note: Assume ideal gas

kPakPakPaP 675.38325.101140

Note: A blower was required to gain the necessary pressure increase (18)

11

00436.01k

k

I

O

B

IB

p

ppV

k

kP

Note: Selected a centrifugal blower because they are more efficient at a lower V

Therefore B = 0.7 (18)

Po = 140 kPa = 20.3 psi

PI = 101.325 kPa = 14.696 psi

4.1

029.21

389.29

Kmol

JKmol

J

C

CK

v

P

Note: Literature values used for CP and vC (49)

BMP

P

B

BC

B

CFC

eC

hphpP

P

hpP

C

73.7426$

9.1275.0

68.9

68.9

))ln(79.08929.6(

Note: Aluminum blades used, therefore FM = 0.6

CP = 4457.24

Thus,

The total cost for 1 blower: $4457.24

where:

K = specific heat ratio

PI = inlet pressure

Po = outlet pressure

V = inlet volumetric flow rate

B = mechanical efficiency

114

The total cost for 2 blowers: $8914.48

A-14: S-201 CO2 Scrubber Calculations

Note: Inert gasses must total no more than 4.5% of the final CNG product.

Note: All gasses except carbon dioxide in methane mix are assumed to have negligible

solubility in water (57)

Note: Due to the fact that the water stream has a mass flow rate three orders of magnitude

higher than the gas stream, the scrubber was assumed to operate at a temperature of 25

°C.

Carbon dioxide must be removed such that the final CNG product is no more than 4.5%

inert gasses. The amount of CO2 there dictates how much of the gas must be removed in

S-201.

Component Stream 10 (kg/hr) Stream 11 (kg/hr)

Methane 25.9 25.9

Ethane 1.9 1.9

Propane 1 1

Water 1.4 0.4

Carbon Monoxide 0.2 0.2

Carbon Dioxide 76.4 X

The value X must be found such that the combination of X and water content is 4.5% of

the molar volume of the gas. Assuming the gasses are ideal,

Component Mass flow

(kg/hr)

Molecular Weight Molar flow

(kmol/hr)

Methane 25.9 16 1.6

Ethane 1.9 30 0.06

Propane 1 44 0.02

Water 0.4 18 0.02

Carbon Monoxide 0.2 28 Negligible

Carbon Dioxide X 44 X/44

Using algebra and solving for X, the mass flow rate of 2.5 kg/hr was found. Thus, 73.8

kg/hr of carbon dioxide must be removed in the scrubber.

Note: Carbon dioxide has a solubility of 1.08 g/kg water (57). Thus the minimum flow

rate was found to be:

hr

kgwater

hr

kg

kg

g

g

kgwaterQ 68435)

8.73)(

1

1000)(

08.1(

This value was increased by 10%, making the final flow rate of water 75278.5 kg/hr y*=87x (58)

115

xa=0; y*a=0

03.0;0004.0

)18

(75278

)44

(8.73* bb y

kg

kmolkg

kg

kmolkg

x

yb-yb*=0.44

ya=0.029

ΔyLM=0.15

Noy= Δy/ΔyLM=0..47/0.15=3.1 transfer units

3.1*0.84 m=2.59 m tower height


Note: Separator height set to 2.5 m = 98.4 in

inmD

mD

m

LD

LrV

i

i

ii

2.7591.1

)5.2(2

2.7

2

2

3

2

2

Note: int p 375.0



Where 3

284.0in

lb for carbon steel

PLvMP CCFC


70.6245$)()(8.361

00.24343$

70684.07396.0

))ln(02297.0)ln(18255.00132.7( 2

LDC

eC

iPL

WW

v


CP = $30588.70

116


Note: The temperature of the stream exiting the valve was calculated using ChemCAD,

as described in Appendix D-3.

Note: The cost of the valve was assumed to be negligible.

117

A-16: E-301 Brine Heat Exchanger Calculations

Note: All nutrient salts were assumed to have heat capacities similar to sodium

bicarbonate.

Note: All salts generated by the decomposition of algae were assumed to have heat

capacities similar to ammonium chloride.

Compone

nt m

(kg/hr

)

CP (kJ/kgK) ΔT Q

(kJ/hr)

Water 41.67 4.184 320 55791.13

NaHCO3 114.4 1.046 (59) 320 38158.08

NH4Cl 13.75 1.8 (60) 320 7930.79

101880

𝑄 = 28.3𝑘𝑊 = 𝑚𝐻2𝑂 𝑕3 − 𝑕2

28.3𝑘𝐽

𝑠= 𝑚𝐻2𝑂 205.24

𝑘𝐽

𝑘𝑔− 134.19

𝑘𝐽

𝑘𝑔

𝑚𝐻2𝑂 = 0.4𝑘𝑔

𝑠= 1440

𝑘𝑔

𝑕𝑟

Note: specific enthalpies obtained from literature (49)




MPa

CT o

98.30

4001

MPa

CT o

11.0

322

MPa

CT o

105.0

493

MPa

CT o

975.30

804 1

2

3

4

118

hrftF

BtuU

F

FF

FF

FFFF

T

T

TTT

TUAQ

o

o

oo

oo

oooo

lm

lm

2

1

2

12

200

7.273

90176

120752ln

)90176()120752(

ln


Note: The heat transfer design equation was used to size the heat exchanger (20

2

2

8.1

7.273200

4.96643

ft

FhrftF

Btu

hr

Btu

TU

QA

o

o

lm


(18).

16.1394$)ln(16.0146.7 A

B eC


BMPP CFFC


0100

0

AAaF

b

M

For pressures greater than 600 psig:

9.2600

0198.0600

1292.08510.0

2

PPFP

06.4043$PC

119

A-17: E-302 Brine Heat Exchanger Calculations

Component m

(kg/hr

)

CP (kJ/kgK) ΔT Q

(kJ/hr)

Water 41.67 4.184 43 7496.93

NaHCO3 114.4 1.046 (59) 43 5145.48

NH4Cl 13.75 1.8 (60) 43 1064.25

13706.66

𝑄 = 3.8𝑘𝑊 = 𝑚𝐻2𝑂 𝑕3 − 𝑕2

3.8𝑘𝐽

𝑠= 𝑚𝐻2𝑂 205.24

𝑘𝐽

𝑘𝑔− 134.19

𝑘𝐽

𝑘𝑔

𝑚𝐻2𝑂 = 0.053𝑘𝑔

𝑠= 190.8

𝑘𝑔

𝑕𝑟

Note: specific enthalpies obtained from literature (49)




hrftF

BtuU

F

FF

FF

FFFF

T

T

TTT

TUAQ

o

o

oo

oo

oooo

lm

lm

2

1

2

12

200

8.38

90140

120187ln

)90140()120187(

ln



MPa

CT o

11.0

861

MPa

CT o

11.0

322

MPa

CT o

105.0

493

MPa

CT o

105.0

434 1

2

3

4

120

2

2

7.1

8.38200

8.12976

ft

FhrftF

Btu

hr

Btu

TU

QA

o

o

lm


(18).

47.1381$)ln(16.0146.7 A

B eC


BMPP CFFC


0100

0

AAaF

b

M


47.1381$PC

121


hps

J

s

hr

hr

kg

kg

JmWP

kg

Jkg

J

headW

kg

J

m

kg

PaPaPPH

PB

P

073.029.63600

16.19039.11

9.1175.0

9.8

9.8

977

)10013.1()101.1(

3

55

12


ftftm

ft

s

m

kg

J

g

head320098.328.3

8.9

9.11

2

min5000

min58.8

17.264977

min60

116.1903

3

3

galgal

gal

m

m

kg

hr

hr

kg

mV

Note: P-102A/B was determined to be a centrifugal pump due its small head (3.98 ft) and

flow rate

min1.105

gal (18)


5.0

g

headVs


5.05.075.20998.3

min1.105 ftgpmft

gals

Note: The size factor for this pump was too small to use empirical relationships to cost

the pump. Therefore, a pump was located that could meet the technical requirements of

the process was found.

A 1/3 hp self priming centrifugal pump with an optimum flow rate of 25 min

gal was

located with a cost of $450.00, including an assumed shipping and handling rate (23).

122

A-19: S-301 Flue Gas Scrubber Calculations

Note: 90% of carbon dioxide is to be removed.

Note: All gasses except carbon dioxide in flue gas are assumed to have negligible solubility in

water (57)

Note: Due to the fact that the water stream has a mass flow rate two orders of magnitude higher

than the gas stream, the scrubber was assumed to operate at a temperature of 25 °C

Amount of carbon dioxide to be removed

hr

kg

hr

kg2.48)6.53)(9.0(

Minimum flow rate

hr

kg

L

kg

g

L

kg

g

hr

kg44629)

1)(

08.1)(

1000)(2.48(

hr

kg

hr

kg49231)44629)(1.1( water flow rate

y*=87x (58)

xa=0; y*a=0

03.0;0004.0

)18

(49231

)44

(2.48* bb y

kg

kmolkg

kg

kmolkg

x

yb-yb*=0.0210

ya=0.0573

ΔyLM=0.036

Noy= Δy/ΔyLM=0.05346/0.036=1.5 transfer units

1.5*1.4 m=2.04 m tower height



inmD

mD

mLD

LrV

i

iii

75.5849.1

)2(2

5.32

2

3

2

2

Note: int p 3125.0



Where 3

284.0in

lb for carbon steel

PLvMP CCFC


61.4448$)()(8.361

30.17089$

70684.07396.0

))ln(02297.0)ln(18255.00132.7( 2

LDC

eC

iPL

WW

v


123

CP = $21537.91


Note: An adiabatic expansion was assumed.

Note: The enthalpy of the salts was assumed to be constant over the valve. Therefore, the

enthalpy of the water would also be constant over the valve.

kg

kJ

kg

kJ

HH fi

64.35964.359

ˆˆ

Therefore, Tf = 86oC

Ti=80 oC

Pi=30.975MPa

Tf = ?

Pf=0.11MPa

Q=0

Water @ P = 0.11 MPa, Tf = 86 oC (52)

Water @ P = 30.975 MPa, T = 80oC (52)

124

A-21: E-401 Methane Cooler Calculations

kWs

kJ

hr

Btu

hr

kJ

mol

kJ

hr

molQ

dTTThr

moldTCnQ

T

T

C

C

CHPgas

o

o

88.488.43.166685.1757199.2107.799

103661.010469.51031.3407.7992

1

4

60

500

2853

,

Note: The specific heat of the gas was assumed to be the specific heat of methane, which

is reasonable as the stream is more than 90% methane

Note: The specific heat of the methane was obtained from literature (54)

𝑄 = 4.88𝑘𝑊 = 𝑚𝐻2𝑂 𝑕4 − 𝑕1

17571.5𝑘𝐽

𝑕𝑟= 𝑚𝐻2𝑂 205.24

𝑘𝐽

𝑘𝑔− 134.19

𝑘𝐽

𝑘𝑔

𝑚𝐻2𝑂 = 247.31𝑘𝑔

𝑕𝑟

Note: The specific enthalpies of the water streams came from a literature source (49)


hrftF

BtuU

F

FF

FF

FFFF

T

T

TTT

TUAQ

o

o

oo

oo

oooo

lm

lm

2

1

2

12

10

37.273

90140

120932ln

)90140()120932(

ln

Note: U was determined based on a conservative estimate for water-gas systems (18).


MPa

CT o

11.0

321

MPa

CT o

99.21

5002

MPa

CT o

985.21

603

MPa

CT o

105.0

494 1

2

3

4

125

2

2

1.6

37.27310

3.16668

ft

FhrftF

Btu

hr

Btu

TU

QA

o

o

lm


(18).

81.1694$)ln(16.0146.7 A

B eC


BMPP CFFC


0100

0

AAaF

b

M


5.1600

0198.0600

1292.08510.0

2

PPFP

22.2542$PC

126

A-22: E-402 Methane Cooler Calculations

kWhr

Btu

hr

kJ

mol

kJ

hr

molQ

dTTThr

moldTCnQ

T

T

C

C

CHPgas

o

o

14.054.47741.50363.007.799

103661.010469.51031.3407.7992

1

4

43

60

2853

,

Note: The specific heat of the gas was assumed to be the specific heat of methane, which

is reasonable as the stream is more than 90% methane

Note: The specific heat of the methane was obtained from literature (54)

𝑄 = 0.14𝑘𝑊 = 𝑚𝐻2𝑂 𝑕4 − 𝑕1

503.41𝑘𝐽

𝑕𝑟= 𝑚𝐻2𝑂 205.24

𝑘𝐽

𝑘𝑔− 134.19

𝑘𝐽

𝑘𝑔

𝑚𝐻2𝑂 = 7.09𝑘𝑔

𝑕𝑟



hrftF

BtuU

TUAQ

o

lm

210

Note: The log mean temperature difference was zero for this case, and therefore the

average approach temperature was used



2

2

4.2

2010

54.477

ft

FhrftF

Btu

hr

Btu

TU

QA

o

o

MPa

CT o

11.0

321

MPa

CT o

985.21

602

MPa

CT o

98.21

433

MPa

CT o

105.0

494 1

2

3

4

127


(18).

83.1459$)ln(16.0146.7 A

B eC


BMPP CFFC


0100

0

AAaF

b

M


5.1600

0198.0600

1292.08510.0

2

PPFP

74.2189$PC

128

A-23: E-403 Hydrogen Cooler Calculations

𝑄 = 𝑚𝐻2 𝑕2 − 𝑕3

𝑄 = 0.193𝑘𝑔

𝑕𝑟 10832

𝑘𝐽

𝑘𝑔− 4434.3

𝑘𝐽

𝑘𝑔

𝑄 = 0.34𝑘𝑊

Note: The specific enthalpies of the hydrogen streams came from a literature source (49)

𝑄 = 0.34𝑘𝑊 = 𝑚𝐻2𝑂 𝑕4 − 𝑕1

503.41𝑘𝐽

𝑕𝑟= 𝑚𝐻2𝑂 205.24

𝑘𝐽

𝑘𝑔− 134.19

𝑘𝐽

𝑘𝑔

𝑚𝐻2𝑂 = 17.3𝑘𝑔

𝑕𝑟



hrftF

BtuU

F

FF

FF

FFFF

T

T

TTT

TUAQ

o

o

oo

oo

oooo

lm

lm

2

1

2

12

10

37.273

90140

120932ln

)90140()120932(

ln



2

2

5.0

37.27310

22.1165

ft

FhrftF

Btu

hr

Btu

TU

QA

o

o

lm


(18).

MPa

CT o

11.0

321

MPa

CT o

115.0

5002

MPa

CT o

11.0

603

MPa

CT o

105.0

494 1

2

3

4

129

80.1135$)ln(16.0146.7 A

B eC


BMPP CFFC


0100

0

AAaF

b

M

For pressures less than than 600 psig, FP = 1

80.1135$PC

130

A-24: E-404 Hydrogen Cooler Calculations

𝑄 = 𝑚𝐻2 𝑕2 − 𝑕3

𝑄 = 0.193𝑘𝑔

𝑕𝑟 4434.3

𝑘𝐽

𝑘𝑔− 4189.8

𝑘𝐽

𝑘𝑔

𝑄 = 0.013𝑘𝑊

Note: The specific enthalpies of the hydrogen streams came from a literature source (49)

𝑄 = 47.2𝑘𝐽

𝑕𝑟= 𝑚𝐻2𝑂 𝑕4 − 𝑕1

47.2𝑘𝐽

𝑕𝑟= 𝑚𝐻2𝑂 205.24

𝑘𝐽

𝑘𝑔− 134.19

𝑘𝐽

𝑘𝑔

𝑚𝐻2𝑂 = 0.66𝑘𝑔

𝑕𝑟



hrftF

BtuU

TUAQ

o

lm

210

Note: The log mean temperature difference was zero for this case, and therefore the

average approach temperature was used



2

2

3.0

2010

77.44

ft

FhrftF

Btu

hr

Btu

TU

QA

o

o


(18).

MPa

CT o

11.0

321

MPa

CT o

115.0

602

MPa

CT o

11.0

433

MPa

CT o

105.0

494 1

2

3

4

131

67.1046$)ln(16.0146.7 A

B eC


BMPP CFFC


0100

0

AAaF

b

M

For pressures less than than 600 psig, FP = 1

67.1046$PC

132

A-25: P-401A/B Methane Compressor Calculations

Note: Empirical correlations were used to size and cost the compressors (Seider, 569)

Compound

hr

kgm

mol

gMW

hr

moln

yi (mol

fraction)

H2O 0.179 18 9.94 0.011

CO2 1.130 44 25.68 0.029

CO 0.094 28 3.36 0.004

CH4 11.541 16 721.31 0.806

C2H6 0.860 30 28.67 0.032

C3H8 0.445 44 10.11 0.011

H2 0.193 2 96.5 0.107

895.57 mol/hr =

14.93 mol/min

1.0

min58.4

min13.0

285000

)298(314.8min

93.1433

3

ftm

Pa

KKmol

Pammol

P

RTnV

Note: Assume ideal gas

31.1

440.27

974.35

Kmol

JKmol

J

C

Ck

v

P

Note: CP and Cv determined from a literature source (49)

11

00436.01k

k

I

O

B

IB

p

ppV

k

kP

133

hphpP

P

hpP

psikPaP

psikPaP

BC

B

o

I

B

2.1175.0

36.8

36.8

84.319022000

34.41285

75.0

Note: Screw compressors were chosen because of the low power consumption (18)

BMDP

P

B

CFFC

eC C

60.19336$))ln(79.08929.6(

For an electric motor, FD = 1

FM = 1 for carbon steel

CP = $19336.60

Thus,

The total cost for 1 compressor: $19336.60

The total cost for 2 compressors: $38673.20

134

A-26: P-402A/B Vacuum Pump Calculations

Note: No empirical relationships for determining the size and purchase cost of a vacuum

pump could be found. Therefore, a vacuum pump was located that met the necessary

technical requirements of the process.

A TorrVac B Series Rotary Vane Vacuum Pump was located with a maximum flow rate

of 21 min

3ft and a maximum vacuum of 0.5 mm Hg was located (25).

135

A-27: S-401 Hydrogen Membrane Calculations

Note: The hydrogen flux across the membrane is )(

06.02 barsm

mol (26)

Note: The area of the membrane is assumed to be 60cm-1

(26)

The maximum amount of hydrogen that could diffuse through this membrane is given by

FAtdPQ

Where time is in seconds

)22)(1000

1)(

2)(06.0)(3600)(06.0( 2

2bar

g

kg

mole

gm

hr

s

sbarm

molQ

hr

kgQ 57.0

This membrane is able to handle the 0.19 kg/hr of hydrogen produced by the process.


The cost of the membrane is approximated to be $50,000. (61)

136

Appendix B: Spread Sheets with Explanations

Team Kazaam! used spreadsheets exclusively to perform the economic analysis of the

project. The following information demonstrates how the spreadsheets were utilized to

arrive at the final economic data. All calculations were performed on a single

spreadsheet; therefore, cells referenced in figures later in this appendix can be found in

figures towards the beginning of the chapter.

B-1: Calculation of Total Bare-Module Costs

The first step in calculating the costs of the process was to determine the installed cost of

each piece of equipment, CTBM. CTBM was obtained by the following equation:

NFCC TBMPTBM

Figure 7 shows the formulas used to calculate the CTBM of each piece of equipment.

Values for CP were obtained from Appendices A and D. FTBM values were obtained from

Sieder (18). N is the number of units required.

Figure 7: Calculating the CTBM of the equipment

137

The output of this spreadsheet is shown in Figure 8.

Figure 8: Output from calculating the CTBM of the equipment

.

B-2: Calculation of Total Capital Investment

This spreadsheet was then used to calculate the total capital investment required for the

project. Figure 9 shows the formulas used to calculate CTCI; this calculation follows the

model put forth by Sieder (18).

Figure 9: Formulas used to calculate CTCI

Calculated CTBM values were initially given in 2006 dollars. To adjust this to 2010

dollars, the following equation was used.

)(O

OI

ICC

138

C is the cost in 2010 dollars. CO is the cost in 2006 dollars. It has a value of 524.2, the

PCI for December 2009, the most recent data available (62). IO has a value of 500, the

PCI for 2006 (18).

The total CTBM is the sum of the inflation adjusted CTBM, the prices of equipment

purchased at 2010 dollars and the price of the initial catalyst charge. The cost of site

preparation is 5% of the CTBM. The cost of service facilities was assumed to be zero, due

to the fact that this facility will be integrated with a nearby algae farm and coal fired

power plant. The allocated cost associated with process and cooling water facility

preparation was calculated using the following formula 96.068.0 15001000 PCAlloc SSC

Where SC has a value of 110.8 and SP has a value of 546.

The CDPI is given by the sum of the CTBM, cost of site preparation and allocated costs.

The cost of contingency funds and contractor’s fee is approximated by being 18% of the

CDPI. The CDPI and the contingency and contractor’s fee cost are summed together to

produce the CTDC.

The cost of land is approximated as 2% of the CTDC. The cost of royalties is set as zero, as

no patented processes are utilized in the design. The cost of plant startup is approximated

as 20% of CTDC. The sum of these factors together with the CTDC yields the cost of the

total permanent investment.

The working capital is the money required to run the plant for one month. This is 8.33%

of the yearly operating cost which is described in the next section. The sum of the total

permanent investment and the working capital yields the CTCI.

Figure 10 shows the output of the spreadsheet used to find CTCI.

Figure 10: Output of formulas used to calculate CTCI

139

B-3: Calculation of Yearly Operating Costs

This spreadsheet was then used to calculate the yearly operating costs required for the

project. Figure 11 shows the formulas used to calculate this cost; this calculation follows

the model put forth by Sieder (18).

Figure 11: Formulas used to calculate yearly operating costs

The cost of the feedstock was assumed to be $1/ton. The yearly cost for the utilities was

read in from Table 21. The cost of direct wages and benefits (DW&B) is based on one

operator per shift at five shifts at 2080 operator hours/year at $35/hour. The operational

direct salaries and benefits were assumed to be 15% of the DW&B. The operational

supplies were assumed to be 6% of the DW&B.

The cost of maintenance wages and benefits (MW&B) was set at 3.5% of the CTDC. The

salaries and benefits associated with maintenance were set at 25% of MW&B.

Maintenance overhead was set at 5% of MW&B.

The general plant overhead was set at 7.1% of the combined wages and salaries for

operations and maintenance (M&O-SW&B). Mechanical services costs were set at 2.4%

of the M&O-SW&B. The employee relations department amounts to 5.9% of the M&O-

SW&B. Business services account for 7.4% of the M&O-SW&B.

Depreciation was set assuming a linear depreciation of the CTDC with 10% salvage value

at the end of a 30 year plant operation. Due to the small annual sales of the process, no

general expenses were assessed. The total annual production cost is the sum of the above

values.

140

Figure 12 shows the output of the above spreadsheet.

Figure 12: Output of formulas used to calculate yearly operating costs

B-4: Calculation of NPV

This spreadsheet was then used to calculate the net present value of the plant. Figure 13

shows the formulas used to calculate this value; this calculation follows the model put

forth by Sieder (18).

Figure 13: Formulas used to calculate NPV

A one year construction time was assumed. Depreciation values were found according to

the method described in the Economic Analysis section; this depreciation is not linear

and is based on the current U.S. tax code. Linear depreciation is typically accounted for

in the calculation of COM (18). Cost of manufacture minus depreciation is the COM

minus depreciation. Sales is the annual sales revenue. Net earnings is sales minus COM

minus depreciation. Discounted cash flow is the net earnings plus depreciation. Cash flow

is given by the following equation

n

CashFlowDiscountedPV

)30.01(

where n is the number of years beyond 2010. Cumulative PV is given by the previous

year’s cumulative PV plus the PV for that year.

The full table showing the final PV and the output of the full 30 year spreadsheet is

shown in the Economic Analysis section.

141

Appendix C: Overall Mass and Energy Balances

Table 28: Flow Rates of Input Streams

Input Stream Mass Flow Rate 𝑘𝑔

𝑕𝑟

1 5240

17 630

23 23300

26 125000

36 1900

Total 156070

Table 29: Flow Rates of Output Streams

Output Stream Mass Flow Rate 𝑘𝑔

𝑕𝑟

25 23300

28 75400

30 4960

31 49300

32 560

35 170

43 1630

48 14.3

52 0.193

61 272

Total 155606.493

Table 30: Net Flow Rates of Individual Components in the Process

Component Input Flow Rate 𝑘𝑔

𝑕𝑟 Output Flow Rate

𝑘𝑔

𝑕𝑟 Net Flow

𝑘𝑔

𝑕𝑟

Algae 125 0 (125)

Air 621 548 (73.354)

Water 155000 155000 37.470

Salts 114 128 13.750

Carbon Dioxide 0 134 133.883

Carbon Monoxide 0 0.211 0.211

Methane 0 11.5 11.542

Ethane 0 0.860 0.860

Propane 0 0.445 0.445

Hydrogen 0 0.193 0.193

Total 155860 155823.209 0.000

Note: The water mass flow rate is three orders of magnitude larger than any other

component. Therefore the mass flows of the entire process can be approximated through

the mass flows of the water. The mass flows in and out of water are approximately equal

142

with the exception of the water produced in the furnace, which are four orders of

magnitude smaller than the total water flow rate.

Table 31: Overall Energy Balance

Equipment Power Consumption (kW) Annual Power Usage (kWh)

P-101A/B 165 1440000

P-102A/B 1.5 13100

P-201A/B 10.9 95400

P-202A/B 9.6 84300

P-301A/B 0.25 2190

P-401A/B 8.4 73200

P-402A/B 1.1 9810

1718000

Note: The annual power usage of P-101A/B is two orders of magnitude larger than any

other pump. Therefore the annual power usage of the entire process can be approximated

through the annual power usage of P-101A/B.

143

Appendix D: ChemCad Outputs

The following printouts were generated using ChemCAD software. The entire process

was not modeled using ChemCAD; rather, the software was generally used to double

check hand calculations to ensure their accuracy. Certain values were calculated using

ChemCAD. These values are highlighted in yellow and referenced in Process Description

section.

D-1: Pumps

ChemCAD outputs for pumps are shown in Table 32. ChemCAD was used to double

check the calculated break horsepower of each pump, shown in the line labeled

“Calculated power hp.” The hand calculated values are in the line below that labeled

“Hand calculated hp.” The line below that marked “Within 5hp?” indicates whether or

not the ChemCAD value is within 5hp of the hand calculated value. If the values are

within 5hp, then the hand calculation was assumed to be accurate. The other information

is presented to give the reader more information about the pump.

Information is not included for P-301 A/B or P-402 A/B. These pumps are commercially

purchased; thus, the power for these pumps was not calculated.

Table 32: ChemCAD outputs for pumps CHEMCAD 6.1.3

Page 1

Job Name: Date: 04/21/2010 Time: 20:39:00

Pump Summary

Equip. No. P-101 A/B P-102 A/B P-201 A/B P-301

A/B

Name

Output pressure kPa 31000.0000 110.0000 290.0000

110.0000

Efficiency 0.7500 0.7500 0.7500

0.7500

Calculated power hp 79.2513 0.1014 11.7122

0.0083

Hand calculated hp 82.3 0.073 11.94

Within 5hp? Yes No Yes

Calculated Pout kPa 31000.0000 110.0000 290.0000

110.0000

Head ft 10289.6855 2.9260 63.3385

2.9260

Vol. flow rate ft3/hr 182.2432 827.9667 4410.1313

67.6009

Mass flow rate kg/h 5184.3999 23321.9023 124470.0000

1904.1599

Cost estimation flag 1 1 1

1

Pump type 1 0 0

0

Install factor 2.8000 2.8000 2.8000

2.8000

Basic pump cost $ 844 3148 4364

5129

144

Basic motor cost $ 3583 356 649

1954

Total purchase cost $ 4427 3504 5013

7084

Total installed cost 12397 9811 14036

19834

($)

Compressor Summary

Equip. No. P-401 A/B P-402 A/B

Name

Pressure out kPa 22000.0000 115.0000

Type of Compressor 1 1

Efficiency 0.7500 0.7500

Calculated power hp 7.5796 0.2868

Hand calculated hp 11.2

Within 5hp?

Cp/Cv 1.3078 1.3923

Theoretical power hp 5.6847 0.2151

Ideal Cp/Cv 1.3015 1.3925

Calc Pout kPa 22000.0000 115.0000

Install factor 1.3000 1.3000

Basic compressor $ 37868 4973

Basic motor cost $ 464 263

Basic driver cost $ 486 26

Total purchase cost $ 38818 5262

Total installed cost 50463 6841

($)

Cost estimation flag 1 1

Calc. mass flowrate 13 0

(kg/h)

D-2: Pump Outlet Streams

ChemCAD was used to calculate the temperatures of streams 2 and 45. The rest of the

pump outlet streams are included here to demonstrate that the streams do not experience a

significant rise in temperature across the pumps. Stream 50 is shown here to have a

temperature of 769 °C; however, this value is not utilized because P-402 A/B is a

commercial vacuum pump, not a theoretical compressor as shown here. ChemCAD does

not have the appropriate settings to simulate a vacuum pump. The rest of the material is

presented to give the reader a greater understanding of the characteristics of these process

streams. ChemCAD simulation of output streams are shown in Table 33.

The schematic used to calculate both the pumps and the outlet streams are shown in

Figure 14.

Table 33: ChemCAD simulation of pump output streams CHEMCAD 6.1.3

Page 1


Stream No. 2 24 27+29

37

Name

- - Overall - -

145

Molar flow kmol/h 279.1532 1294.5824 6909.2426

105.6986

Mass flow kg/h 5184.4000 23321.9018 124470.0000

1904.1600

Temp C 35.0712 32.0027 25.0604

32.0030

Pres kPa 31000.0000 110.0000 290.0000

110.0000

Vapor mole fraction 0.0000 0.0000 0.0000

0.0000

Enth kJ/h -7.9866E+007 -3.6917E+008 -1.9739E+009 -

3.0142E+007

Tc C 420.1414 374.2000 374.2000

374.2000

Pc kPa 32869.7819 22118.2302 22118.2302

22118.2302

Std. sp gr. wtr = 1 1.008 1.000 1.000

1.000

Std. sp gr. air = 1 0.641 0.622 0.622

0.622

Degree API 8.8663 10.0000 10.0000

10.0000

Average mol wt 18.5719 18.0150 18.0150

18.0150

Actual dens lb/ft3 62.5232 62.0991 62.2215

62.0991

Actual vol ft3/hr 182.8063 827.9673 4410.1987

67.6009

Std liq ft3/hr 181.6187 823.6063 4395.6226

67.2449

Std vap 0 C scfh 220958.2344 1024701.4375 5468876.0000

83663.6563

- - Liquid only - -

Molar flow kmol/h 279.1532 1294.5824 6909.2426

105.6986

Mass flow kg/h 5184.4000 23321.9018 124470.0000

1904.1600

Average mol wt 18.5719 18.0150 18.0150

18.0150

Actual dens lb/ft3 62.5232 62.0991 62.2215

62.0991

Actual vol ft3/hr 182.8063 827.9673 4410.1987

67.6009

Std liq ft3/hr 181.6187 823.6063 4395.6226

67.2449

Std vap 0 C scfh 220958.2344 1024701.4375 5468876.0000

83663.6563

Cp Btu/lbmol-F 18.1855 17.9844 18.0078

17.9844

Z factor 0.3887 0.0010 0.0028

0.0010

Visc cP 0.8024 0.7946 0.9218

0.7946

Th cond Btu/hr-ft-F 0.3520 0.3557 0.3503

0.3557

Surf. tens. dyne/cm 69.8316 70.8904 72.0931

70.8903

CHEMCAD 6.1.3

Page 2

146


Stream No. 45 50

Name

- - Overall - -

Molar flow kmol/h 0.8697 0.0957

Mass flow kg/h 13.4880 0.1930

Temp C 500.4834 768.9864

Pres kPa 22000.0000 115.0000

Vapor mole fraction 1.000 1.000

Enth kJ/h -41519. 2038.5

Tc C -80.2268 -239.8800

Pc kPa 5563.3378 1295.9465

Std. sp gr. wtr = 1 0.300 0.070

Std. sp gr. air = 1 0.535 0.070

Degree API 340.4216 1889.9286

Average mol wt 15.5092 2.0158

Actual dens lb/ft3 3.0354 0.0017

Actual vol ft3/hr 9.7963 254.7803

Std liq ft3/hr 1.5886 0.0974

Std vap 0 C scfh 688.3771 75.7840

- - Vapor only - -


Mass flow kg/h 13.4880 0.1930



Actual vol ft3/hr 9.7963 254.7803

Std liq ft3/hr 1.5886 0.0974

Std vap 0 C scfh 688.3771 75.7840

Cp Btu/lbmol-F 14.5970 7.2554

Z factor 1.0768 1.0002

Visc cP 0.02478 0.02119

Th cond Btu/hr-ft-F 0.0799 0.2575

Figure 14: ChemCAD Pumps

147

D-3: Valve Outlet Streams

. The properties of the streams leaving the three letdown valves of the process were

simulated using ChemCAD. ChemCAD simulations of output streams are shown in Table

34. ChemCAD was used to calculate the temperature of stream 13. Streams 8 and 34 are

shown here as well; the simulated values are within 10% of the hand calculated values,

demonstrating that the hand calculated values are accurate.

The schematic used to calculate the valve outputs is shown in Figure 15.

Table 34: ChemCAD simulations of valve output streams CHEMCAD 6.1.3

Page 1


Stream No. 8 13 34

Name

- - Overall - -

Molar flow kmol/h 279.0179 1.1149 3.8206

Mass flow kg/h 5069.6190 18.0440 169.8160

Temp C 111.1644 23.9800 80.0000

Pres kPa 300.0000 140.0000 110.0000

Vapor mole fraction 0.02686 1.000 0.0000

Enth kJ/h -7.7541E+007 -87903. -1.3658E+006

Tc C 366.0101 -77.8064 984.3835

Pc kPa 21212.5792 5598.5001 63966.1696

Std. sp gr. wtr = 1 0.983 0.313 1.662

Std. sp gr. air = 1 0.627 0.559 1.535

Degree API 12.4558 320.8178 -46.3615

Average mol wt 18.1695 16.1839 44.4476

Actual dens lb/ft3 3.7659 0.0578 101.0561

Actual vol ft3/hr 2967.8679 688.3544 3.7047

Std liq ft3/hr 182.1394 2.0369 3.6083

Std vap 0 C scfh 220851.1875 882.5038 3024.1104

- - Vapor only - -


Mass flow kg/h 177.9338 18.0440



Actual vol ft3/hr 2785.9036 688.3544

Std liq ft3/hr 9.3846 2.0369

Std vap 0 C scfh 5931.3022 882.5038

Cp Btu/lbmol-F 8.9200 8.5563

Z factor 0.9885 0.9976

Visc cP 0.01512 0.01120


- - Liquid only - -


Mass flow kg/h 4891.6841 169.8160



Actual vol ft3/hr 181.9644 3.7047

Std liq ft3/hr 172.7547 3.6083

Std vap 0 C scfh 214919.8594 3024.1104

Cp Btu/lbmol-F 18.2277 25.5928

Z factor 0.0024 0.0220

148

Visc cP 0.2495 1.519


Surf. tens. dyne/cm 56.4453 98.1514

Figure 15: ChemCAD Valves

D-4: Heat Exchanger Duties

The duties of the ten heat exchangers used in the process were simulated using

ChemCAD; the ChemCAD output of these exchangers is found in Table 35; these values

are shown in the line labeled “Calc Ht Duty kJ/s.” Underneath that line are the hand

calculated values in the line labeled “Hand values kJ/s.” All exchangers except for E-301

and E-302 had ChemCAD calculated values within 10% of the hand values; thus, the

hand values were assumed to be accurate. The bigger discrepancy in E-301 and E-302 is

likely caused by different heat capacities of the salts; in these cases, the data used in hand

calculations were assumed to be more accurate.

The ChemCAD schematic used in this simulation is shown in Figure 16.

Table 35: ChemCAD heat exchanger outputs CHEMCAD 6.1.3

Page 1


Heat Exchanger Summary

Equip. No. E-101 E-102 E-201 E-

202

Name

1st Stream dp kPa 5.0000 5.0000 5.0000

5.0000

2nd Stream dp kPa 5.0000 5.0000 5.0000

5.0000

1st Stream T Out C 465.0000

2nd Stream T Out C 372.0000 43.0000

131.0000

Calc Ht Duty kJ/s 2321.4846 556.9088 84.0278

1.2704

149

Hand values kJ/s 2308 518 80

1.24

LMTD (End points) C 39.7675 31.7110 37.7404

43.2850

LMTD Corr Factor 1.0000 1.0000 1.0000

1.0000

1st Stream Pout kPa 30965.0000 105.0000 135.0000

120.0000

2nd Stream Pout kPa 30995.0000 295.0000 125.0000

135.0000

Equip. No. E-301 E-302 E-401 E-

402

Name

1st Stream dp kPa 5.0000 5.0000 5.0000

5.0000

2nd Stream dp kPa 5.0000 5.0000 5.0000

5.0000

2nd Stream T Out C 80.0000 43.0000 60.0000

43.0000

Calc Ht Duty kJ/s 40.6089 4.8854 5.0465

0.1472

Hand values kJ/s 28.3 3.8 4.9

0.14

LMTD (End points) C 150.2101 19.6421 152.0590

10.5497

LMTD Corr Factor 1.0000 1.0000 1.0000

1.0000

1st Stream Pout kPa 105.0000 105.0000 105.0000

105.0000

2nd Stream Pout kPa 30975.0020 105.0000 21985.0020

21980.0020

Equip. No. E-403 E-404

Name

1st Stream dp kPa 5.0000 5.0000

2nd Stream dp kPa 5.0000 5.0000

2nd Stream T Out C 60.0000 43.0000

Calc Ht Duty kJ/s 0.3429 0.0131

Hand values kJ/s 0.34 0.013

LMTD (End points) C 152.1790 10.9607

LMTD Corr Factor 1.0000 1.0000

1st Stream Pout kPa 105.0000 105.0000

2nd Stream Pout kPa 110.0000 105.0000

150

Figure 16: ChemCAD heat exchangers

151

Appendix E: Web Printouts

P-102A/B

152

P-301A/B

153

P-402A/B

154

Appendix F: Meeting Logs

Meeting/ Phone Log 1

Date: February 2, 2010

Members Present (name of senior design group members plus name and title of person

providing information)

Dr. Seames, Professor at University of North Dakota

Members: Kim Seamans, Zac Ronan, Kyle Kryger and Amanda Rubio

Summary of Information, that pertains to the report (costs, flow rates, sizes,

assumptions).

The meeting was very helpful in guiding us with our plant design because Dr. Seames

was well verse in the topic of algae. After we were able to give him some details about

our plant design he was able to provide some assumptions about the algae. It can be

assumed that algae is available as a slurry with 2.5% algae by mass for purchase and that

a reasonable amount of algae entering the process a day is about 3 tons/day.

Meeting/ Phone Log 2

Date: April 22, 2010

Members Present (name of senior design group members plus name and title of person

providing information)

Mike Arnold, Professor at University of Arizona, Team Kazaam

Members: Kim Seamans, Zac Ronan, Kyle Kryger and Amanda Rubio

Summary of Information that pertains to the report (costs, flow rates, sizes, assumptions).

The last meeting with Professor Arnold to review the overall reports and help us find any

discrepancies. There were some minor changes and when he was reviewing the

economic analysis he suggested that we change the interest rate for the discounted cash

flow from 15% to 30% because our product is new and inherently risky.

repository.arizona.edu · abstract: this project involved designing a process that converts algae...

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