final report outline team 2_updated version

Date: May 11, 2016

To: Dr. Jamie Gomez

From: Team 2 with members as follows:

Madelaine S. Chavez, Adrian Ledesma-Mendoza, Thao Pham, Brandi Saavedra, Stephen S. Ulibarri

_______________ _______________ _______________ _______________ _______________

Subject: Sour Water Stripping_________________________________________________________________________________

Letter of Transmittal

Sour water is the wastewater that is normally produced in many refining processes. This

water typically contains hydrogen sulfide (H2S) and ammonia (NH3) along with slight remnants of

hydrocarbons. These compounds must be removed from the water in order to be reused in the

refinery or before being sent to a wastewater system.

For this project, sour water will be obtained from the Deer Park refinery which is owned by

Shell Oil Company and is located in Deer Park, TX. Our operation will be set up within the plant,

allowing us to treat the sour water directly from the Deer Park Refinery’s wastewater stream. This

stream will contain 300-3000 parts-per-million (ppm) NH3, 5 ppm H2S and trace amounts of propane.

The stream will range from 20-50 gallons per minute (GPM). We will remove the hydrocarbons,

ammonia and hydrogen sulfide and return the treated water to the plant for steam generation.

Attached is the full report in the following order:

Executive Summary

Decision Matrix and Innovation Map

Introduction/Background

Process/Project Description

Economic Analysis

Safety

Environmental and Community Awareness

Conclusions

Recommendations

University of New MexicoDepartment of Chemical and Biological Engineering

Sour Water Stripping

Team No. 2

Stephen Ulibarri

Brandi Saavedra

Thao Pham

Adrian Ledesma-Mendoza

Madelaine Chavez

Direct Stripping Services Inc.Report Element Specific Section # in the Report

Project DescriptionEconomics

Safety

Environmental

Capstone Report Type: Final WrittenDate: May 11, 2016

CBE 494LInstructor: Jamie Gomez, Ph.D

2


Table of Contents1. Executive Summary.................................................................................................................................5

2. Decision Matrix and Innovation Map....................................................................................................10

2.1 Innovation Map...............................................................................................................................10

2.2 Decision Matrix...............................................................................................................................10

3. Introduction/Background.......................................................................................................................13

3.1 Introduction.....................................................................................................................................13

3.2 Background.....................................................................................................................................14

4. Process/Project Description...................................................................................................................20

4.1 PFD.................................................................................................................................................20

4.2 List of Equipment............................................................................................................................24

4.3 Process Description.........................................................................................................................25

4.4 Energy Analysis...............................................................................................................................31

4.5 Vessel Sizing...................................................................................................................................34

5. Economic Analysis................................................................................................................................43

5.1 Capital Investment...........................................................................................................................44

5.2 Direct Costs.....................................................................................................................................44

5.3 Indirect Costs...................................................................................................................................46

5.4 Profitability......................................................................................................................................49

5.5 Risk Analysis...................................................................................................................................52

6. Safety.....................................................................................................................................................53

6.1 Plan Layout......................................................................................................................................53

6.2 Control system implementation.......................................................................................................54

6.3 HAZOP analysis..............................................................................................................................57

6.4 Plant Emergency Procedures...........................................................................................................61

7. Environmental & Community Awareness.............................................................................................63

7.1 Regulations......................................................................................................................................63

7.2 Community Awareness....................................................................................................................63

8. Conclusions...........................................................................................................................................64

8.1 Project Goals Satisfaction................................................................................................................64

8.2 Profitability Achievement................................................................................................................65

9. Recommendations.................................................................................................................................65

9.1 Bench-model Experiment................................................................................................................65

3


9.2 Service Expansion...........................................................................................................................66

References.................................................................................................................................................67

Appendix A-1: Vessel Sizing Calculations................................................................................................69

Appendix A-2: Net Enthalpy of Reaction by Vessel.................................................................................69

Appendix B-1: HENSAD Summary table.................................................................................................70

Appendix B-2: HENSAD DT Sensitivity Plot Results..............................................................................71

Appendix B-3: HENSAD Temperature Interval Diagram.........................................................................71

4


1. Executive SummaryProject Scope:

The goal of the project is to remove the contaminants from sour water at Deer Park Refinery in Deer

Park, Texas. The water will come from a tail gas quenching system on a sulfur recovery unit at a rate

of 20-50 gal/min. The sour water will contain 300-3000 ppm of ammonia, 5 ppm of hydrogen

sulfide, and trace amounts of propane (3 to 5 ppm) and will be cleaned to at most 20 ppm of

ammonia, 3 ppm of hydrogen sulfide, and an undetectable amount of propane. These contaminant

concentration requirements will allow for the water to be reused for steam generation at the refinery.

This will allow for the prevention of wastewater dumping into the environment, transportation costs

and the cost of clean water for steam generation. An additional objective is to eliminate the refinery’s

annual disposal costs which will result in an annual savings of $1.4 million for the refinery. An

additional source of profit for the wastewater treatment plant is the production of calcium sulfate

from a side reaction which results in an annual profit of $13 million at a selling price of $7/kg.

Project Description:

The first step in the process is to remove virtually all of the propane. This will be done by using a

series of activated carbon filters. These vessels are made from carbon steel and the bed volume

(vessel size) was determined based on the inlet flowrate of the sour water stream. In order to strip out

the ammonia, calcium hydroxide will be used to raise the pH so that ammonium ions will exist as

dissolved gaseous ammonia. The stripping agent for the ammonia will be air in counter-current flow

with the water. Gaseous air and ammonia will exit the top while the bottoms product will contain 20

ppm of ammonia or less. The size of this vessel was designed using an overall mole balance as well

as a heuristic approach: The Flooding Correlation. This takes into account the packing height for the

ring type and size. For this vessel, Raschig rings were used. After a lime recovery process that

regenerates calcium hydroxide for recycle, a reaction in the process renders the water in the system

acidic, converting bisulfide ions to dissolved hydrogen sulfide gas. The vessels up to this point are

also made of carbon steel and the vessel size was designed with a heuristic method using the

detention time for calcium carbonate. This stream is ultimately sent to a flash evaporator where

hydrogen sulfide volatilizes from solution to an air stream. The bottoms product from the flash is sent

back to the plant for steam generation. The size of this vessel was designed using a heuristic

approach by Phillip C. Wankat, which is also made of carbon steel. The gas from the flash evaporator

is cleaned in a packed tower scrubber with sodium hypochlorite as the stripping agent in a counter-

5


current flow with the gas. The overhead is pure air and the bottoms product is aqueous sulfuric acid

which is then neutralized with calcium oxide forming a side product, calcium sulfate, as well as salt

water. This vessel’s size was designed by using the same heuristic method as the air stripper: The

Flooding Correlation. The vessels up to this point are also made of carbon steel. This material was

chosen because at no part in the process is the sour water corrosive enough to damage the material.

The water recycled for steam generation is purified down to levels below those determined by the

International Electric Committee, which determines contaminant levels for steam generation.

Several alternatives were considered for the design of this project. Propane can be removed via many

methods. One alternative, the gas flotation unit, involves forcing hydrocarbon gas to dissolve in

solution at high pressures and introducing it to the treated water at atmospheric pressure, allowing for

hydrocarbon bubbles to remove propane droplets from the solution. The process requires high

residence times and recirculation, creating a bottleneck in the entire procedure. Skim tanks were

another alternative. The water is held in a tank where the propane is skimmed off of the surface of

the water. This method, like the flotation unit, requires high settling times for the coalescence of

propane droplets and was deemed inappropriate for the project.

Two alternatives to the selected ammonia removal method were breakdown chlorination and ion-

exchange. Breakdown chlorination requires the use of chlorine gas to produce intermediates that

ultimately convert the ammonia to nitrogen gas. If the pH is too high, sodium hydroxide is used to

neutralize the solution. Sodium hydroxide is corrosive and its use is preferably kept to a minimum.

The required chlorine to ammonia weight ratio is 7.6:1, making the bleach input requirement too

large. The ion-exchange method requires the use of highly porous media such as an ion-exchange

resin or zeolite. The ammonia adsorbs into the media and is removed from the system. Periodic

replacement of the media or an ion desorption process is a requirement. The additional waste

requirements and the time requirement for the removal or treatment of the media would create a

bottleneck in the system.

Hydrogen sulfide can be removed by using a mist scrubber or a packed tower scrubber using either

sodium hydroxide or sodium hypochlorite as the stripping agent. The mist scrubber disperses the

stripping agent at high velocities in which the hydrogen sulfide is carried to the bottom while clean

air exits as the overhead. The packed tower scrubber utilizes a packing material in which the surface

area contact between the gas and liquid is maximized. The mist scrubber often has higher energy

requirements because it is also designed to remove particulate solids. A reasonable tower size can be

6


achieved with a lower energy requirement and material cost by using a packed tower. Sodium

hypochlorite was chosen as the stripping agent because it is less corrosive and expensive than

calcium hydroxide.

Waste considerations were made by referring to the Texas Administrative Codes (TAC). These codes

are listed in Section 7.1-Regualtions. OSHA regulations for standard industries were also used for the

safety aspect for the project.

Economic Results:

The refinery pays approximately $2.5 million annually to dispose of its wastewater and obtain clean

water for steam generation. The cost of disposal for wastewater includes transportation to the

treatment plant and treatment fees. The economic evaluation involves two main categories: the

capital investment and the operating costs. The capital investment is the total amount of money

required to construct the sour water treatment plant in the existing refinery facility, which was

generated by CAPCOST at $15.6 million. This amount will be financed 100% by the refinery with an

interest rate of 10%. The operating costs which represent ongoing expenses for day-to-day operations

amount to $1.1 million a year.

A large portion of capital investment comes from equipment costs. The material chosen for all pieces

of equipment was carbon steel to eliminate undesirable effects on the project design such as

corrosion and reaction with reactants or products at the normal operating conditions (pressure <10

bar, temperature between 40-250oC). The vessels have a bare module cost of $11.6 million. The cost

of land is not considered as the sour water treatment plant is built on unoccupied land owned by the

refinery.

The operating costs include the cost of raw materials, utilities and labor costs. Raw materials, which

include calcium hydroxide, sodium hypochlorite and calcium oxide, cost $500 thousand annually.

Utilities, along with the disposal cost of the brine, amount to $3 thousand a year. The number of

operators is based on the staffing requirement for a small sour water treatment plant that runs below

50,000 gallons per day. The plant will be running 16 hours a day, 7 days a week with a stream factor

of 0.67. This was considered appropriate because the sulfur recovery unit at the Deer Park Refinery

doesn’t operate 24 hours a day. There are 5 operators per shift as well as 2 technicians assigned to the

first shift when maintenance and system updates are needed. Total operating labor costs amount to

$500 thousand per year.

7


The economic analysis was run at a 10% annual interest rate and a 42% taxation rate. The total

operating costs the plant proposes will save the refinery $1.4 million a year compared to the sour

water disposal and clean water cost that the refinery is currently paying. The sour water treatment

plant proposes to eliminate transportation fees and disposal costs of the sour water, as well as the

need for wastewater dumping permits and the concern of where the sour water should properly be

disposed. According to the cash flow diagram generated in CAPCOST, the plant has a payback

period of 7.6 years after startup with DCFROR of 11.9%. The net present value within 10 years of

the DSS service is $1.3 million. A Monte-Carlo stimulation carried out the risk analysis and showed

that there is a 60% chance the sour water treatment plant can achieve the NPV of $1.3 million.

Status of Project:

The system has been designed to reduce the content of the above stated contaminants down to their

target concentrations. Once the design was completed, a successful Aspen simulation was developed

to model the system. From these results, flow rates were obtained allowing for energy balances and

vessel sizing. These calculations provided the necessary information to calculate unit prices along

with the utilities necessary to run the operation.

Chemical prices have been obtained from reliable sources and used to estimate the cost of materials

for the economic analysis, such as the side product, calcium sulfate. The cost of electricity in Deer

Park, TX has been researched and the value was used to estimate the cost of operation of the plant

based on the energy flow requirements of the process.

The yearly operating costs for Deer Park have been calculated in order to estimate the cost that the

refinery would have to pay without the DSS stripping process. After the economic analysis, it has

been verified that the DSS stripping process costs less than the net cost for the refinery to dispose of

the sour water and to buy clean water. This takes into account additional costs to transport the clean

and sour water to and from the refinery, respectively.

The safety analysis of the project is complete. A finalized plot plan has been developed where the

location of each vessel is indicated as well as access roads. Basic controllers, such as pressure and

temperature controllers, have also been implemented in the process. A HAZOP analysis has been

performed on two nodes of interest while analyzing the variations in flow rate.

Conclusions:

8


The design of the sour water treatment facility has been completed to successfully realize the project

goals while maintaining profitability. The DSS company will see a net present value of $1.3 million

over the course of the project with an DCFROR of 11.9%. Based on the models employed, the

concentrations of the contaminants in the sour water have been reduced to meet the desired

concentrations in the project objectives. It is a recommended alternative to Deer Park’s current means

of disposing of their sour water and purchasing clean water for steam generation. The Deer Park

refinery would benefit by a net savings of $1.4 million annually, as well as reducing their current

production time. The project also results in the protection of the environment from the dumping of

harmful wastewater streams.

9


2. Decision Matrix and Innovation Map 2.1 Innovation Map

The innovation map was used to help determine the direction of the project. It begins at the

bottom, titled Process Manufacturing Technology. This is broken down into three sections based on

the technology used to clean the water. The row above this is the technical differentiation that makes

this project different. Research had to be conducted in order to determine what options would make

this process different and efficient. The row above this shows the products that are achievable in this

process, which are water and calcium sulfate. The water is reused by the refinery for steam

generation. It is assumed that the refinery has a demineralizer for water they use for steam

generation. Calcium sulfate is a side product generated in the process. The top row is the Consumer

Value Proposition. These are the goals of the project showing the benefits for the investing

consumers. One of the main benefits of this project is the regeneration of calcium oxide. This is

innovative and efficient because it greatly reduces the raw material requirement.

2.2 Decision MatrixBecause the sour water treatment facility involves three major sections (propane removal,

ammonia removal, and hydrogen sulfide removal), three decision matrices were constructed when

10

Stripping out hydrogen sulfide

in scrubber

Air stripping of ammonia

Adsorption of hydrocarbons

Calcium SulfateClean Water

Regeneration and recycling Calcium

Hydroxide

High product yieldUtilizing pH instead of catalyst to remove

contaminants

Lower costMeets EPA regulations

RecyclingEnvironmentally responsible

Process Manufacturing

Technology

Technical Differentiation

Product

Consumer Value Proposition


considering the alternatives for each section. An in-depth discussion on the reasoning behind each

choice is visited in more detail in section 3.2 – Background.

Table 1 – Decision Matrix for Propane Removal

Flotation Tank Skim Tank Activated Carbon FilterCriteria Weight Rating Score Rating Score Rating Score

Cost 0.25 1 0.25 1 0.25 2 0.5Time

Requirement 0.75 1 0.75 1 0.75 2 1.5

Total 2 1.0 2 1.0 4 2.0Notes Best Option

The time requirement for the propane removal system is the most important criterion because

the high throughput of sour water should not be interrupted. The lower the time requirement (or

additional time required for the method to work), the higher the score of the method. Thus the time

requirement was given a weighting of 75%. The flotation and skim tank methods require the

stoppage of flow to allow for propane coalescence, which imposes a severe time penalty for the

process. For this reason, these methods were given a score of 1, while the activated carbon filter

system allows the water flow to continue uninterrupted, so it was given a rating of 2. The cost is the

second most important criterion and was given a weighting of 25%. Flotation and skim tanks are

large, requiring more material. They also contain moving parts which require energy input and hence

have higher operating costs. Thus these methods were each given ratings of 1 for the cost criterion.

The activated carbon filter system only requires periodic the replacement of the carbon filters, so it

was given a cost rating of 2. With a total score of 2.0, the activated carbon filter system is the

preferred method for the removal of propane.

Table 2 – Decision Matrix for Ammonia Removal

Breakdown Chlorination Ion Exchange Air Stripping

Criteria Weight Rating Score Rating Score Rating ScoreSafe 0.25 1 0.25 2 0.5 3 0.75Cost 0.6 1 0.6 1 0.6 3 1.8Time

Requirement 0.15 2 0.3 1 0.15 3 0.45

Total 4 1.15 4 1.25 9 1.38Notes Best Option

The removal of ammonia involves three criteria: safety, cost and time requirement. The

magnitude of the cost differences between the three alternatives is larger for this section than for the

propane removal section. For this reason, cost is given a weighting of 60%. Because the removal of

11


propane can be done via chemical means, input chemical requirements can affect the safety of the

selected method. Safety was given a weighting of 25%. Finally, interruptions to the flow of sour

water can greatly affect production time, the time requirement was given a weighting of 15%.

Breakdown chlorination requires a large amount of sodium hypochlorite, so it was given a safety

rating of 1. The ion exchange method does not require hazardous chemicals, but the need to

frequently replace the zeolite or ion-exchange resin introduces additional hazards and potential

accidents so it was given a rating of 2. Air stripping does not require chemical input or component

replacement and can be operated continuously, so it was given a safety rating of 3. The chemical

requirement of the breakdown chlorination method results in a high operating cost, so its cost rating

is 1. The ion-exchange method requires frequent replacement or maintenance (depending on whether

a disposable resin or a reusable zeolite is used, as explained in section 3.2 – Background), so its cost

rating is also 1. Air stripping merely requires an air input and thus its cost rating is 3. Finally,

breakdown chlorination requires a long enough residence time for the numerous chemical reactions

to occur in which ammonia is ultimately converted to molecular nitrogen and thus its time

requirement rating is 2. Ion exchange requires longer and more frequent interruptions to the flow of

the operation and thus its time requirement rating is 1. Air stripping requires no interruptions and

continuous flow is possible, so its time requirement rating is 3. The total score for the air stripping

method is the highest, at 1.38, so this method was chosen for the project.

Table 3 – Decision Matrix for Hydrogen Sulfide

Mist Scrubber Packed Tower ScrubberCriteria Weight Rating Score Rating ScoreEnergy 0.7 1 0.7 2 1.4

Efficiency 0.3 1 0.3 2 0.6Total 1.0 2.0Notes Best Option

Two methods for the removal of hydrogen sulfide were considered. The use of a mist

scrubber was compared to the use of a packed tower scrubber. The energy requirement was

considered the most important criterion and was given a weighting of 70%. The mist scrubber is

designed to remove particulates as well as gaseous components and thus has a higher energy

requirement than the packed tower scrubber. For this reason, the mist scrubber was given a rating of

1 and the packed tower scrubber was given a rating of 2. The efficiency was the second criterion

chosen, with a weighting of 30%. Because of the wasted energy input into the mist scrubber due to

its ability to remove particulates, it was given a rating of 1. The packed tower scrubber does not

waste energy in this manner. In addition, the use of packing rings to maximize the surface area

12


contact between the gas and liquid phases increases the efficiency of the mass transfer in the packed

tower scrubber, so it was given a rating of 2. With the highest total score of 2.0, the packed tower

scrubber was the chosen method for the removal of hydrogen sulfide.

3. Introduction/Background 3.1 Introduction

The largest contributor to waste streams produced in oil refining and gas processing is

water[1]. Depending on the environmental regulations and the contaminants in the water and their

concentrations, this water may be dealt with in a number of ways. 95% of wastewater from the

aforementioned processes is either pumped back into the production reservoir to enhance oil

recovery, or into depleted or underground reservoirs. Other forms of disposal involve ocean and

stream dumping. Texaco has dumped 18 billion gallons of wastewater into surface streams [2]. In

most cases, wastewater disposed of in this manner is only minimally treated, and thus damaging to

the environment and harmful to aquatic life. Recycling and reusing wastewater in refining and gas

processing is one alternative to the above environmentally harmful methods of disposal.

In gas processing, feed streams can contain as much as 12,000 ppm H2S and 8,000 ppm

NH3[1]. Much of this sulfur can be recovered via the Claus process, the industry standard in sulfur

recovery [3]. The Claus process involves the recovery of elemental sulfur from gaseous hydrogen

sulfide in the following overall reaction:

2 H2 S+O2 →2 S+H2O (1)

Tail gas from the Claus process still contains sulfur contaminants and is hydrogenated to

improve the efficiency of sulfur removal units. Quench systems for the tail gas create waste streams

containing H2S, NH3, and potentially trace amounts of other components. Of wastewater produced in

upstream applications in gas processing plants, this is the primary source [1].

13


The Deer Park refinery in Deer Park, Texas, owned by Shell, produces and ships sulfur as

one of its primary products [4]. The refinery sits adjacent to the Houston Ship Canal, which flows into

Trinity Bay and ultimately the Gulf of Mexico. Dumping of wastewater into the canal has a negative

impact on the environment. Sending the wastewater to treatment plants involves transportation costs,

treatment fees and labor costs. In addition, to replace the discarded water, fresh water must be

purchased. An economic and environmentally safe method to deal with this wastewater would be to

treat it on site and reuse it. While sour water cannot be cleaned to the potable level, it can be

sufficiently treated to steam generation quality. Of the above solutions, the latter is clearly the most

desirable. This report is a detailed design of such a solution, including a complete description of the

process, an economic analysis, a safety analysis, and an environmental and community awareness

analysis.

3.2 BackgroundThe sour water produced by a tail gas quenching system at the Deer Park Refinery will be

subject of this analysis. The water will be assumed to leave the sulfur recovery unit at a rate of 20-50

gal/min. As sour water is characterized as either phenolic (containing phenols) or non-phenolic, the

treatment process must be designed accordingly. The sour water in this analysis is non-phenolic, as

the water from sulfur recovery unit tail gas quenching systems typically contains H2S, NH3, and trace

14

Figure 1- Deer Park Refinery


amounts of other molecules such as light hydrocarbons [1]. For the purpose of this design project, such

hydrocarbons will consist solely of propane in trace amounts.

While many methods of hydrocarbon removal from water exist, the method for this project

must be appropriately selected for the situation. Namely, the stream must flow continuously during

operation hours, therefore processes such as skim tanks and gas flotation devices which require high

residence times and little or no flow are not feasible. These types of devices are heavily dependent on

coalescence, or the phenomenon of smaller oil droplets colliding and forming larger droplets. The

time required to allow a small droplet to grow into a large droplet is:

t=( dd )4

2 f v K s

(2) [5]

where

dd = droplet diameterfv= volume fraction of the dispersed phaseKs= empirical settling constant

In other words, the residence time is proportional to the fourth power of the droplet diameter.

In addition, the residence time must be even greater for more dilute solutions. A skim tank would

therefore be large and exhibit poor flow behavior considering the high throughput requirement for

this project.

One example of a gas flotation unit is the dissolved gas unit [5]. At high pressures (20 to 40

psig), gas (such as natural gas) is forced to dissolve in aqueous solution. This highly pressurized

solution is introduced to the treated water at atmospheric pressure. The dissolved gas molecules then

bubble out of solution and rise to the top of the vessel, coalescing with the dissolved contaminants on

their way. The treated water is typically recirculated back to the vessel for further treatment. Due to

the high pressures required for this method which would result in a high energy input, along with the

potential for bottlenecking of the entire project due to the long residence times required, this

alternative was rejected.

Adsorptive activated carbon filters remove small hydrocarbons in high capacity [6]. They

involve no mechanically moving parts and require no energy input. Because they don’t require low

flow rates to be effective, they can handle the high throughput of sour water for the purposes of

removing trace amounts of propane. In addition, any other small organic molecules that may be

15


present will also adsorb to the activated carbon component [7]. For these reasons, this method best

suits the project and will be utilized.

Hydrogen sulfide and ammonia ionize to some extent in aqueous solution, behaving as a

weak acid and a weak base, respectively [8]. Ammonia tends to deprotonate hydrogen sulfide,

considerably lowering its Henry’s law constant [1]. This has the effect of “tying up” the molecules, or

making them more difficult to remove [1,9]. Since the contaminants can only be removed in their

“parent” (un-ionized) form, the pH must be altered to accomplish the treatment. Specifically, a basic

environment is necessary to ensure that all ammonium molecules are deprotonated and removable in

the form of ammonia. This has the effect of tying up the sulfide molecules. Conversely, an acidic

environment is necessary to ensure that bisulfide ions are protonated and removable in the form of

hydrogen sulfide. For these reasons, the plant must be separated into two major sections: a basic

section for the removal of ammonia and an acidic section for the removal of hydrogen sulfide.

One common method to remove ammonia from water is breakdown or breakpoint

chlorination. The term "breakpoint" is used to describe conditions in which ammonia and chlorine

are both at a minimum. Chlorine oxidizes ammonia in a series of reactions that ultimately result in

the conversion of ammonia to molecular nitrogen which is removed in its gaseous form. To reach the

breakpoint, a typical weight ratio requirement of Cl to NH3 is 7.6:1, corresponding to a molar ratio of

1.5:1 [10]. The molar flow rate of ammonium in the sour water analyzed in this report is 0.945 kmol/hr,

thus the amount of sodium hypochlorite required to achieve the breakpoint, according to the

following calculation, is:

0.945 kmolhr

× 1.5mol Cl

mol NH 4+¿× 1 mol NaOCl

mol Cl× 74.44 kg

kmol=105.53 kg NaOCl /hr ¿

In addition to sodium hypochlorite, pH levels must be regulated for the breakpoint to be

achieved. The pH is generally between 7.0 and 8.0 [10] and to keep it this low, sodium hydroxide must

be added. The amount of chemical reagent necessary for the breakpoint to be achieved is far greater

than that required to remove ammonia via other means, as will be shown later.

Another method of ammonia removal is ion-exchange. The process consists of two stages. In

the first stage, wastewater passes through a highly porous media such as a zeolite or an ion-exchange

resin. For ammonia removal, a resin is selected that has an affinity for NH 4+ ions. Wastewater is

pumped through the system until the purity requirements are fulfilled and the porous matrix is

16


saturated with ammonium. The wastewater flow must then be interrupted for the second stage, where

either the resin is disposed of or the NH4+ ions are removed from zeolite via bioregeneration. In

bioregeneration, the ammonium is oxidized by nitrifying bacteria. Because the bacteria can only

oxidize ammonium in solution, cations are pumped into the zeolite matrix to desorb ammonium back

into solution [10]. Because the second stage of the ion-exchange method involves the cease and

buildup of the wastewater flow, additional waste removal, potential addition of cationic solutions and

nitrifying bacteria if the zeolite recovery method is employed, as well as additional workers to

perform the frequent removal of materials and operation of equipment, this method is not appropriate

for the treatment of the sour water in this project.

Air stripping of ammonia is a third alternative. When the pH of the wastewater is raised to

approximately 12, virtually all of the ammonia exists in the form of dissolved gas [11]. A pH of

approximately 11 is sufficient for the removal of gaseous ammonia at low temperatures. The

wastewater is brought into contact with air to allow for the ammonia to transfer into the vapor phase.

This is achieved by dispersing the water over internal packing media which greatly increases surface

area contact. Air containing the stripped ammonia is released into the atmosphere. The process

requires no heating and continuous, uninterrupted flow at a high flowrate is possible. If calcium

hydroxide is the base used to raise the pH of the solution, the potential exists for the regeneration of

calcium hydroxide for recycle, improving the efficiency of the ammonia removal process. For these

reasons, air stripping was the method chosen for this project.

Removal of hydrogen sulfide can be done by the use of an H 2S scrubber. Two alternative

designs are mist scrubbers and packed tower scrubbers. Mist scrubbers, also known as wet scrubbers,

are designed to remove gaseous pollutants as well as particulates. Because of their particulate

removal capabilities, these scrubbers require a high energy input, resulting in higher operating costs.

Mist scrubbers have lower overall contaminant removal efficiencies and reuse of spent liquid is a

common necessity. Packed tower scrubbers involve the use of packed beds which increase surface

area contact between the gas and liquid phases. Volatilized gas containing hydrogen sulfide comes

into contact with the stripping agent and reacts to form liquid products while the clean air is removed

from the system. Scrubbing can occur via the use of sodium hydroxide or sodium hypochlorite. The

wastewater in this project does not contain particulates, so a wet scrubber is not appropriate. Packed

tower scrubbing was the method of choice for this project, with sodium hypochlorite for the stripping

agent, as it is less corrosive than sodium hydroxide and the chemistry of the treatment results in the

17


formation of sulfuric acid, which, after the addition of lime, can be converted into sodium sulfate, a

product which can be sold for additional revenue as will be discussed in the process description.

With the above technologies selected for the treatment of the sour water, a list of input chemicals is summarized in Table 4:

Table 4 – Raw Material Requirements for Sour Water Treatment

Raw Material Use Amount Required [kg/hr]Sodium Hydroxide NH4

+ Deprotonation 1.07Calcium Oxide Neutralization of Sulfuric Acid

produced by H2S Scrubber43.2

Sodium Hypochlorite Reacting Agent in H2S Scrubber 228

There are several constraints for this design project. Concentrations of the contaminants

much be sufficiently lowered in order to reuse the water for steam generation. The contaminants and

their respective concentrations in the sour water, along with target concentrations after treatment, are

shown in Table 1. Target concentrations were selected such that contaminant levels remain well

below steam generation limits set by the International Electric Committee [12].

Table 5 – Sour Water Contaminant Levels and Their Target Concentrations

Contaminant Concentration Target Concentration

Hydrogen Sulfide (H2S) 5 ppm < 3 ppm

Ammonia (NH3) 300-3000 ppm < 20 ppm

Propane (C3H8) < 5 ppm Undetectable

Economic feasibility is another constraint of the project. In order for the process to be

affordable and efficient, the design will be implemented as an auxiliary facility within the Deer Park

Refinery. Tail gas quench effluent will enter the system directly, avoiding excessive pipeline

construction or wastewater transportation. The process must be economically feasible. Namely, the

cost of operation of the plant must not exceed the sum of costs currently paid by the Deer Park

Refinery. These costs will be assumed as follows:

1. Sour water effluent from tail gas quenching systems is transported to external wastewater

treatment facilities. Transportation costs and treatment costs apply here.

2. Clean water is purchased by the Deer Park Refinery for steam generation. This involves the

cost of fresh water as well as the cost to transport fresh water to the refinery.

18


3. Untreated wastewater is dumped into the Houston Ship Canal to the extent allowed by

environmental regulations, resulting in the cost of wastewater dumping permits.

Environmental safety imposes additional constraints to the project. Effluent gases from the

NH3 stripper and H2S scrubber must not contain contaminant concentrations in excess of those

enforced by environmental regulations and necessary permits must be purchased in order to release

these gases legally. The stripped ammonia exiting the NH3 scrubber is of primary concern here, as

the air exiting the H2S scrubber will be essentially clean.

19

4. Process/Project Description 4.1 PFD

Figure 2-Complete Process Flow Diagram of the Sour Water Stripper Project. Back up vessels not shown.


Table 6 – Flow Summary Table for the Sour Water Stripping Process Shown in Figure 2

Stream Number 1 2 3 4 5 6 7 8 9 10Temperature [°C] 25 25.00 25.00 25.00 25.00 25.00 25.98 25.00 25.98 25.98Pressure [bar] 1 1.00 1.00 1 1.00 1.00 1.00 1.00 1.00 1.00Vapor Fraction 0 3.07 0.64 0 1.00 0.00 0.003 1.000 1.00 0.003Mole Flow [kmol/hr] 630.30 630.30 0.07 644.06 0.14 644.07 637.65 4442.38 0.95 636.70

Mass Flow [kg/hr] 11369.19 11369.13 1.54 11626.61 0.60 11627.2111511.0

7 11511.07 16.10 11494.98Component flowrate [kmol/hr] Propane 0.00128 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Hydrogen Sulfide 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Bisulfide 0.95 0.95 0.00 0.95 0.00 0.95 0.95 0.00 0.00 0.95 Water 628.40 628.40 0.03 641.67 0.00 641.68 635.26 0.00 0.00 635.26 Ammonia 0.00 0.00 0.00 0.95 0.00 0.95 0.95 0.00 0.95 0.00 Ammonium 0.95 0.95 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Calcium Hydroxide 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Calcium Oxide 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Carbon Dioxide 0.00 0.00 0.00 0.00 0.14 0.008 0.008 0.00 0.00 0.008 Calcium Carbonate 0.00 0.00 0.00 0.00 0.00 0.005 0.00 0.00 0.00 0.00 Sodium hypochlorite 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Sulfuric Acid 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 OH- 0.00 0.00 0.03 0.01 0.00 0.00 0.00 0.00 0.00 0.00 Sodium Chloride 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Calcium Sulfate 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

H30+ 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Air 0.00 0.00 0.00 0.00 0.00 0.00 0.00 4442.38 0.00 0.00 Ca2+ 0.00 0.00 0.01 0.48 0.00 0.48 0.48 0.00 0.00 0.48 H+ 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

21


Table 7 – Flow Summary Table for the Sour Water Stripping Process Shown in Figure 2 (Continued)

Stream Number 11 13 14 15 16 17 18 19 20Temperature [°C] 25.98 25.00 25.00 25.00 25.00 25.00 25.00 25.00 25.00Pressure [bar] 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00Vapor Fraction 0.00 0.001 0.0008 0.00 1.00 0.00 0.00 0.24 0.00Mole Flow [kmol/hr] 6.42 636.19 629.38 6.82 9.91 628.44 3.06 4.00 0.77

Mass Flow [kg/hr] 16.14 11511.3811353.8

0 161.58 292.48 11321.56 228.00 260.24 43.20Component flowrate [kmol/hr] Propane 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Hydrogen Sulfide 0.00 0.94 0.94 0.00 0.94 0.000003 0.00 0.17 0.00 Bisulfide 0.00 0.002 0.002 0.00 0.002 0.000 0.00 0.002 0.00 Water 6.41 634.78 628.44 6.35 0.00 628.44 0.00 0.00 0.00 Ammonia 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Ammonium 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Calcium Hydroxide 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Calcium Oxide 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.77 Carbon Dioxide 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Calcium Carbonate 0.005 0.47 0.00 0.47 0.00 0.00 0.00 0.00 0.00 Sodium Hypochlorite 0.00 0.00 0.00 0.00 0.00 0.00 3.06 0.00 0.00 Sulfuric Acid 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.77 0.00 OH- 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Sodium Chloride 0.00 0.00 0.00 0.00 0.00 0.00 0.00 3.06 0.00 Calcium Sulfate 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

H30+ 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Air 0.00 0.00 0.00 0.00 8.97 0.00 0.00 0.00 0.00 Ca2+ 0.00 0.0008 0.0008 0.00 0.0008 0.00 0.00 0.0008 0.00 H+ 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00

22


Table 8 – Flow Summary Table for the Sour Water Stripping Process Shown in Figure 2 (Continued)

Stream Number 21 22 23 24 25 26 27Temperature [°C] 131.35 25.00 25.00 25.00 38.15 98.11 70.19Pressure [bar] 1.00 1.00 1.00 1.00 1.00 1.00 1.00Vapor Fraction 0.22 0.05 0.00 1.00 1.00 0.00 1.00Mole Flow [kmol/hr] 4.79 4.01 0.77 8.97 0.46 13.22 13.69Mass Flow [kg/hr] 303.44 199.20 104.40 260.24 20.40 255.95 255.95Component flowrate [kmol/hr] Propane 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Hydrogen Sulfide 0.18 0.17 0.00 0.00 0.00 0.00 0.00 Bisulfide 0.002 0.002 0.00 0.00 0.00 0.00 0 Water 0.77 0.77 0.00 0.00 0.00 12.76 12.3 Ammonia 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Ammonium 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Calcium Hydroxide 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Calcium Oxide 0.005 0.005 0.00 0.00 0.00 0.46 0.00 Carbon Dioxide 0.00 0.00 0.00 0.00 0.46 0.00 0.00 Calcium Carbonate 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Sodium Hypochlorite 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Sulfuric Acid 0.00 0.00 0.00 0.00 0.00 0.00 0.00 OH- 0.00 0.00 0.00 0.00 0.00 0.00 0.93 Sodium Chloride 3.06 3.06 0.00 0.00 0.00 0.00 0.00 Calcium Sulfate 0.77 0.00 0.77 0.00 0.00 0.00 0.00 H30+ 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Air 0.00 0.00 0.00 8.97 0.00 0.00 0.00 Ca2+ 0.0008 0.0008 0.00 0.00 0.00 0.00 0.46 H+ 0.00 0.00 0.00 0.00 0.00 0.00 0.00

23

4.2 List of EquipmentTable 9 – Equipment Summary for Stripping Sour Water PFD

Vessel/Towers/Reactors Fr-101 Fr-103 Fr-104 Fr-105 T-101

(A/B)T-103 (A/B)

T-105 (A/B)

Temperature [°C] 25 25 25 25 25.98 25 25Pressure [bar] 1.0 1.0 1.0 1.0 1.0 1.0 0.7

Orientation Vertical Vertical

Vertical Vertical Vertical Vertical Vertica

lMOC CS CS CS CS CS CS CSSizeHeight /length [m] 2.4 3.05 3.05 3.05 3.05 1.5 2.74Diameter [m] 1.9 2.5 3.6 2.7 4 0.3 0.91

Internals packed rings

packed rings 1 tray

Vessel/Towers/Reactors R-101 R-102 R-103 R-104 R-105 R-106

Temperature [°C] 25 98.11 25.98 25 131.35 25Pressure [bar] 1.0 1.0 1.0 1.0 1.0 1.0

Orientation Vertical Vertical

Vertical Vertical Vertical Vertical

MOC CS CS CS CS CS CSSizeHeight /length [m] 3.7 1.6 3.7 3.7 3.7 3.7Diameter [m] 3.6 0.7 0.4 3.0 2.7 1.3Internals

Pumps P-101 P-102 P-103 P-104Heat Exchanger

E-101

Flow [kg/h] 11626 11511 11494 11353 Type Double PipeFluid Density [kg/m3] 1000 1000 1000 1000 Area [m2] 1

Power [kW] 0.373 0.711 0.711 0.711 Duty [MJ/h] 3.6

Type Centrf. Centrf. Centrf. Centrf. ShellMOC CS CS CS CS Temp [°C] 25Temperature [°C] 25 25 25 25 Pres [bar] 1Pressure in [bar] 1 1 1 1 Phase lPressure out [barg] 0.834 0.834 0.834 0.834 MOC CS

TubeTemp [°C] 98Pres [bar] 1Phase lMOC CS

Key:

MOC

Material of construction Centrf. Centrifugal

CS Carbon steel l liquid


4.3 Process Description

Figure 3- The Activated Carbon Filter System and Reactor R-100

Stream 1 exits the tail gas quenching system in the Deer Park Refinery containing ammonium

and bisulfide ions along with trace amounts of propane. This stream enters the activated carbon filter

system, Fr-101 (see Figure 3), consisting of two activated carbon filters in series (not shown). All

propane is adsorbed within the carbon filter system. Stream 2 exits the carbon filter system and enters

R-101. Stream 3, containing calcium hydroxide, enters the reactor along with stream 27, a calcium

hydroxide recycle. In reactor 101 the pH is raised to 11 and the ammonium ions are protonated to

form dissolved ammonia gas in the following reaction:

NH 4

+¿+12Ca (OH )2 → NH3+H 2 O+ 1

2 Ca¿ (3)

25


Figure 4- The Flocculation Tank and Filter System 101

R-101 contains excess calcium hydroxide to ensure that the ammonium deprotonation

reaction proceeds to completion, tying up the bisulfide ions in the process as well as maintaining a

high enough pH to strip the ammonia without any heat input. Many stripping towers in industry raise

the pH to approximately 11 in order to successfully accomplish ammonia stripping at 25 °C [13].

Stream 4 exits R-101 with the ammonia and excess calcium hydroxide and enters R-106 (see Figure

4), the flocculation tank. In this tank, air is passed over the surface of the liquid to introduce CO 2

which causes calcium carbonate to participate in the following reaction:

Ca (OH )2+CO2 →Ca CO3+H 2 O (4)

26


Figure 5-The Ammonia Stripper

The excess calcium hydroxide has now been converted to calcium carbonate which exits via

stream 6 as solid precipitate in solution (lime sludge) and must be filtered out in the Fr-103 filter

system (see Figure 4). This sludge, stream 11, is sent to R-103, the lime recovery system. The

remaining water, stream 7, contains dissolved ammonia gas which is then sent to the NH 3 stripper, T-

101 (see Figure 5). Air enters the NH3 stripper via stream 8. The ammonia gas comes out of the

liquid phase and joins the air on its way up the column, ultimately exiting the stripper via stream 9,

which is vented to the atmosphere. The water exits the bottom of the stripper via stream 10.

As mentioned in the previous paragraph, the lime sludge in stream 11 exiting from filter

system 103 is sent to the lime recovery system, R-103 (see Figure 6). This is a heated vessel which

converts calcium carbonate to calcium oxide and CO2 in the following reaction:

CaCO3 →CaO+CO2 (5)

This reaction is endothermic and requires energy input to occur. R-103 is therefore heated

and sits in a heating jacket. A heat exchanger (not shown) recycles heat from a downstream unit (R-

105) to assist in the heating of the lime recovery system. This will be explained in detail later.

27


Figure 6- The Lime Recovery System and its Relationship to the Slaker, the Recarbonation Tank and the Lime Sludge Inlet Streams

Lime sludge which has been filtered from a downstream recarbonation tank is also sent back

to the lime recovery system as stream 15 (see Figure 6). The calcium oxide produced in the lime

recovery system leaves via stream 26 and enters R-102, the slaker, where it reacts with the water in

the solution to regenerate calcium hydroxide in the following reaction:

H 2O+CaO →Ca (OH )2 (6)

This is the recycled calcium hydroxide, stream 27, mentioned earlier. It returns to R-101 to

deprotonate ammonium ions.

The CO2 produced in the lime recovery system, as shown by reaction equation (5), is sent to

R-104, the recarbonation tank. Stream 10, the bottoms product from the NH3 stripper, also enters the

recarbonation tank, now essentially free of ammonia. In the recarbonation tank, the following

reactions occur:

H 2O+CO2→ H+¿+HCO 3−¿¿ ¿ (8)

Ca2+¿+HC O3−¿→ H

+ ¿+ CaCO3¿¿ ¿ (9)

H+¿+ HS−¿→ H2 S ¿¿ (10)

28


The CO2 from stream 25 reacts with water to form a proton and bicarbonate intermediate.

Calcium ions react with this intermediate to form an additional proton and calcium carbonate, the

source of the lime sludge recycle stream (stream 15) mentioned earlier. This sludge is filtered by Fr-

104 and the remaining water exits via stream 14. The excess protons from the reactions in the

recarbonation tank result in the protonation of bisulfide, forming dissolved H2S gas which is

extremely volatile. The liquid exits filter system 104 as stream 14 and is sent to a flash evaporator, T-

105 (see Figure 7).

Figure 7– The Flash Evaporator

In the flash evaporator, T-105, the H2S gas is removed from the water. The water exits via

stream 17 and is sent back to the Deer Park Refinery for steam generation. The H 2S gas exits the

flash evaporator via stream 16 where it enters the H2S scrubber, T-103 (see Figure 8).

29


Gas from the flash evaporator, containing the gaseous hydrogen sulfide, enters the bottom of

the packed tower scrubber, T-103. Sodium hypochlorite is introduced to the scrubber via stream 18,

where it is sprayed into the packed column as a fine mist. Once it comes into contact with the gas

from the bottom inlet, the gas and the liquid react according to the following reaction:

H 2 S+4 NaClO⇌H 2 SO4+4 NaCl (11)

A sulfuric acid/sodium chloride electrolyte solution trickles down the column and exits via

stream 19. The gas, now virtually free of hydrogen sulfide, exits via stream 24 which is vented to the

atmosphere. Stream 19 enters R-105 where calcium oxide is fed to neutralize the sulfuric acid and

convert it to calcium sulfate according to the following highly exothermic reaction:

CaO+ H 2 SO4 →CaSO4+H 2 O (12)

The calcium sulfate precipitates out of solution and is removed in Fr-105 as stream 23, where

it is packaged and stored for commercial sale. The remaining liquid in filter system 105 is essentially

brine and is dumped into the Houston Ship Canal.

30

Figure 8– The Packed Tower Scrubber, Reactor 202, and Filter System 201


4.4 Energy Analysis

With the exception of the flash evaporation unit (T-105), the lime recovery system (R-103),

the slaker (R-102) and the sulfuric acid neutralization unit (R-105), the plant operates primarily at

25°C and 1 bar. This was verified by calculating the enthalpy of reaction for each chemical reaction

in the plant and determining if there was a significant temperature change in the reactor vessel. The

enthalpy of formation for each of the chemicals participating in any given reaction were retrieved

from literature [8]. The sum of the enthalpies of formation of the products of a given reaction, each

multiplied by its corresponding stoichiometric coefficient, minus the sum of the enthalpies of

formation of the reactants, each multiplied by its corresponding stoichiometric coefficient, gave the

enthalpy of reaction, according to Hess's Law (Equation 13).

∆ H ( reaction )0 ¿ ∆ H f ( products )

0 −∆ H f (reactants )0 (13)[8]

The net enthalpies of reaction for each vessel in the plant can be found in Appendix A-2.

To give an example, the enthalpy of reaction for the neutralization of ammonium ions in

vessel R-101 is –11.63 kJ/mol. The negative sign indicates that the reaction is exothermic, and the

vessel is heated as a result of the ionization. The vessel contains enough calcium hydroxide to

completely neutralize the ammonium entering the vessel, which totals 0.945 kmol/hr ammonium. To

deprotonate this much ammonium, using the enthalpy of reaction stated above, the energy

requirement would be –10,992 kJ/hr (see Appendix A-2). This energy requirement is equal to the

mass flow rate of the liquid in the vessel multiplied by its specific heat capacity (equation 14).

q=m Cp ∆ T (14)[8]

where

q = the heat associated with the temperature change in the vesselm = the mass flow rate of the fluid undergoing the temperature changeC p = the specific heat capacity of the fluid∆ T = change in temperature of the fluid

The specific heat capacity was approximated to be that of water due to the low concentration

of the chemical species in the solution (641.67 kmol/hr of water versus 0.95 kmol/hr of ammonia,

0.48 kmol/hr of calcium ions and 0.01 kmol/hr of hydroxide ions).

Equation (14) requires the temperature of a single inlet stream. Because there are multiple

inlet streams entering R-101 at different temperatures (see Figure 3), the vessel was treated as if

31


streams 2, 3 and 27 combined to a single stream before entering the vessel. The mass flowrate and

temperature of such a stream is calculated with the following equations:

m=m1+m2+…+mn (15)[14]

and

T=(m1 C p1 T 1+ m2 Cp2 T2+…+mnC pnT n) /(m1C p1+ m2 Cp2+…+mnC pn) (16)[14]

where

m = final massm1…n = mass of substancesT = final temperatureT 1…n = temperatures of substances

These formulas were used to give an inlet temperature for equation (14). Using the heat

evolved in the reaction, the outlet temperature was determined to be 25.99 °C. As the vessel and pipe

material is taken to be room temperature (25° C), the temperature gradient between that of the fluid

in the reactor and that of the vessel itself determines that the temperature of the vessel material would

reach no greater than 25.99 °C, according to Fourier’s Law:

qA

=k ∆ TB (17)[16]

where

q = rate of heat flow normal to the surfaceA = surface area∆ T = temperature drop across the layerk = thermal conductivity of the pipe or vessel materialB = thickness of the layer

As the fluid exits the reactor and flows through the pipe, the value of q in equation (17)

diminishes as the temperatures of the fluid and the pipe equilibrate. Because of this, by the time

stream 4 reaches the flocculation tank, it is assumed to have reached room temperature. Thus the

temperature in stream 4 is taken to be 25° C.

A similar analysis was performed on all reaction vessels in the system. For vessels in which

the temperature change was not significant, the outlet streams were taken to be 25° C. For those

vessels in which the temperature change was significant, an energy recovery system was needed. For

example, the lime recovery system, R-103, requires an input of 84,500 kJ/hr and must be run at

32


98.11° C. Conversely, the sulfuric acid neutralization occurring in R-105 is highly exothermic, bringing

the temperature in the outlet stream to 131.35 ° C. Thus a heat exchanger was designed to use the heat

evolved in R-105 to heat the inlet stream to R-103 from 25° C to 98.11° C. For details on the design

of this heat exchanger, refer to the heat exchanger design spreadsheet in the Appendix B on the CD

and see the equipment description section of this report.

The other vessel which undergoes a temperature change significantly greater than 1° C is the

slaker, with an inlet stream of 98.11° C from the lime recovery system and an outlet stream at 70.19°

C. Although this stream (stream 27) is higher than 25° C, its contribution to the temperature change

in R-101, as described above, is less than 1° C.

The pressure for those vessels which do not evolve gas in their reactions is assumed to

remain at 1 bar. CO2 is formed in the lime recovery system via reaction (5), which bubbles to the

surface of the solution and is allowed to flow to the recarbonation tank. The gas evolved by the

reaction is fed in an amount to match the stoichiometric requirement of reaction (8) in the

recarbonation tank, so the flow is driven by the steady production of gas in the lime recovery system

and the consumption of CO2 in the recarbonation tank.

One of the primary ingenuities in the plant is the regeneration of calcium hydroxide. Because

calcium enters the system in the form of calcium hydroxide, it became apparent that solid precipitates

could become a problem in the system, resulting in fouling on the piping system interior. It was

decided that this precipitation could be facilitated intentionally and the precipitates could be filtered

out via settling tanks. One of the reactions occurring in the recarbonation tank is caused by the

introduction of carbon dioxide. The result is the formation of bicarbonate ions, which then react with

the calcium ions to form calcium carbonate. This is similar to when cave water containing dissolved

calcium and bicarbonate ions drips from the ceilings of caves [15]. On its way down, the water loses

CO2, lowering the pH of the water and causing the calcium carbonate to precipitate. In the case of the

recarbonation tank, the bisulfide ions (HS-) are protonated, consuming protons, thus lowering the pH

and causing the calcium carbonate to precipitate. This sludge along with that filtered from the stream

leaving the flocculation tank enters the lime recovery system where heating facilitates the generation

of calcium oxide and carbon dioxide gas. The carbon dioxide is captured and delivered to the

recarbonation tank and the calcium oxide and water is converted to calcium hydroxide in the slaker,

which is then recycled back to the R-101 unit. This design utilizes the products of all reactions

33


involved, eliminating the need for waste disposal. Figure 9 shows the path of calcium through the

system. The bold arrows depict the calcium recycle loop.

Figure 9 – The Calcium Recycle and Ca(OH)2 Regeneration Loop

4.5 Vessel Sizing

Each reaction vessel was sized based on the incoming flowrate to that vessel and the

detention time of the reaction. This was directly related to the maximum incoming flow rate of 50

gpm for the sour water stream from the tail gas quenching system of the sulfur recovery unit in the

Deer Park Refinery.

Activated Carbon Filter System

The activated carbon filter was sized based off of the daily maximum water throughput to the

plant, which was 72,000 gal/day. This was used to determine the bed volume in ft 3 as well as the

mass of carbon needed in lbs. The bed volume and mass of the adsorbed are given by:

Bed Volume ( BV )= QQ b

(18)[17]

M=( BV ) ( ρs ) (19) [17]

where

BV = bed volumeQ = the design liquid flow rateQb = satisfactory liquid flow rate in bed volumes per unit time

34


M = mass of adsorbentρ s = adsorbent bulk density

Qb is generally within the range of 0.2 to 3.0 bed volumes per hour. A value of 1.67 is

appropriate for activated carbon filter systems [17]. The bulk density of activated carbon is 25 lb/ft3 [17].

The activated carbon filter system was chosen to consist of two filters in series. This accomplishes a

double filtration effect as well as provides the opportunity to replace a filter without the need to stop

the operation at any given moment in time. A backup carbon filter is connected but isolated by closed

valves. By diverting the flow from the filter which needs to be replaced (cutting off flow to the filter

by closing valves) to the backup filter, the filter which has been cut off from the system can be

replaced. Thus two activated carbon filters are in service at any given time.

Reaction Vessels

For the vessels that contain calcium carbonate, detention times range from 1-4 hours [18].

Because the system was sized for the largest flowrate, the smallest detention time (0.5 hour) was

assumed to size the vessels. This is because the amount of water per day was amounted to be very

small (72,000 gal/day) in comparison to other water treatment plants that have a much larger intake

(i.e. 4,000,000 gal/day). The primary, secondary, and tertiary precipitations tanks sizing also have a

detention time of 0.5 hour for the same reason. The remaining two reactors, R-101 and R-104, were

also sized based off of a half hour detention time. This detention time is imparted on the vessels

because they are limited by the detention times of the calcium carbonate-containing vessels. The

flocculation tank has a detention time of 0.25 hours, based on heuristics for most flocculation tanks

in water treatment plants. [13]

The volumetric flowrate for a reactor vessel is related to the velocity of the flow and the

cross-sectional area of the vessel according to the following equation:

Q=vA (20)[18]

where

Q = volumetric flowratev = fluid velocityA = cross-sectional area of the reactor

The volumetric flow rate is related to the volume of the reactor and the detention time

according to the following equation:

35


Vv

=τ (21)

where

V = volume of the reactorv = volumetric flow rate of the fluid τ = detention time of the reaction

The volumetric flow rate into each reactor was used to find the reactor volume. For reactors

in which multiple inlet streams entered the reactor, the sum of the volumetric flowrates was used.

The volumetric flowrate of the inlet to the reactor was then multiplied by the corresponding detention

time and the result was the volume of the vessel. Strategic height-to-diameter ratios were chosen to

as to leave enough floor space for maintenance and plant personnel. Vessel ratios are available in

Appendix A.

Ammonia Stripper and Hydrogen Sulfide Scrubber

The Flooding Correlation was used to calculate the vessel diameter for the ammonia stripper

and the hydrogen sulfide scrubber [19]. The Flooding Correlation is an empirical model for the

flooding velocity of a packed tower. Figure 10 shows the correlation.

Figure 10 – The Flooding Correlation, taken from Fundamentals of Momentum, Heat and Mass Transfer

36


Figure 10 shows the relationship between a few key variables in a packed tower. The

maximum allowable pressure drop for the tower can be chosen and its corresponding curve on the

graph relates the x-axis value to a y-axis value. These values will be used in subsequent calculations

to arrive at G', the superficial velocity of the gas, possessing units of mass/area-time. When the mass

flow rate of the entering gas is divided by G', a tower cross-sectional area is obtained.

A simplified diagram of a packed tower is shown in Figure 11. In this diagram, L and G

denote liquid and gas flowrates, respectively. Lowercase and uppercase x’s and y’s refer to molar and

mass fractions, respectively. The arrows indicate whether the streams are entering or leaving the

tower. This diagram was used as an aide in the calculations that follow. While in the ammonia

stripper, the solute is ammonia and is transferring from the liquid stream (sour water) to the gas

stream (air), in the scrubber, the solute is hydrogen sulfide and is being transferred from the gas

stream (from the flash evaporator) to the liquid stream. Due to these differences, the calculations

were adjusted accordingly when performing the analysis on the stripper versus the scrubber. The

overall technique, however, was the same for both units, and the calculations for the stripper will be

shown here.

Figure 11 – Streams Entering and Exiting the Packed Tower

The gas and liquid flowrates at the bottom of the tower will be deemed AG 1 and AL1,

respectively. The inlet molar flowrate of the gas stream was obtained from the ideal gas law:

37


AG1=V PRT (22)[19]

where

AG1 = molar flowrate of the gas at the bottom of the towerV = volumetric flowrate of the gas at the bottom of the towerP = pressure of the gas at the bottom of the towerT = temperature of the gas at the bottom of the towerR = the universal gas constant

The average molecular weight of the inlet gas is determined by

M G1= y A 1 M A+(1− y A1)M B (23)

where

M G1 = average molecular weight of the inlet gasy A 1 = mole fraction of ammonia in the inlet gasM A = molecular weight of ammoniaM B = molecular weight of air

Here, the ammonia content in the inlet gas stream is zero. However, when sizing the

scrubber, the inlet gas stream contains hydrogen sulfide gas and its corresponding mole fraction is

not zero.

The mass flowrate of the entering gas is thus given by:

AG1' =AG1 ∙ M G1 (24)

The average molecular weight of the entering liquid stream, as well as its molar flow rate was

calculated using the following equations:

M L 2=x A 2 M A+(1− xA 2) M B (25)

where

M L 2 = average molecular weight of the inlet liquidx A2 = mole fraction of ammonia in the inlet liquidM A = molecular weight of ammoniaM B = molecular weight of water

AL2=AL2

'

M L2 (26)

where

AL2 = mass flow rate of ammonia in the liquid inlet streamAL2 = molar flow rate of ammonia in the liquid inlet streamM L 2 = average molecular weight of the inlet liquid

38


A balance of the solute (ammonia) around the terminal streams gives:

AL1=AL2(1−x A2)

1−x A 1 (27)

AG2=( AL2−AL1 )+AG1 (28)

x A2 AL2= y A 2 AG2+x A 1 AL1 (29)

where

AL1 = molar flow rate of ammonia in the liquid outlet streamAL2 = molar flow rate of ammonia in the liquid inlet streamx A2 = mole fraction of ammonia in the liquid inlet streamx A 1 = mole fraction of ammonia in the liquid outlet streamAG1 = molar flow rate of ammonia in the gas inlet streamAG2 = molar flow rate of ammonia in the gas outlet stream

The average molecular weight of the gas outlet stream is given by:

M G2= y A 2 M A+(1− y A2)M B (30)

where

M G2 = average molecular weight of outlet air streamM A = molecular weight of ammoniaM B = molecular weight of airy A 2 = mole fraction of ammonia in outlet air stream

And the mass flow rate of the outlet gas stream is given by:

AG2' =AG2 ∙ M G2 (31)

where

AG2' = mass flowrate of the gas outlet stream

AG2 = molar flowrate of the gas outlet streamM G2 = average molecular weight of the gas outlet streamy A 2 = mole fraction of ammonia in outlet air stream

Finally, with the mass flowrates known, the x-axis on the Flooding correlation could be

determined. The ratio of the liquid to gas-mass flowrate at the bottom of the tower is

L'

G'=AL'

AG' (32)

39


This ratio was calculated to be 0.089 ~ 0.01 for the ammonia stripper. Thus the x-axis on the flooding

correlation is:

L'

G' ( ρG

ρL−ρG)

1/2

(33)

where ρG and ρL are the gas and liquid densities, respectively. This value was calculated to be 0.003.

Depending on the allowable gas pressure drop desired in the tower, different y-axis values on the

Flooding Correlation chart will be obtained. Thus a sensitivity study was conducted in which

different values for ΔP were chosen and different y-axis values were calculated. The y-axis value is

equivalent to the following expression:

y axis=(G ' )2 c f (μL)

0.1 JρG(ρL−ρG)gc

(34)

where

G' = the superficial velocity of the gasc f = a packing characteristic (dependent on the type of packing rings chosen)μL = the liquid-phase viscosityJ = constant for the flooding correlationgc = the gravitation constantρG = the gas densityρL = the liquid density

Solving the above equation for the superficial velocity of the gas, G':

G'=√ ( y axis)ρG (ρL−ρG) gc

c f (μL)0.1 J

(35)

The cross-sectional area of the tower can then be determined from G' with the following

relationship:

A= AG '

G' (36)

and thus the tower diameter is

D=√ 4 Aπ

(37)

The results of the sensitivity study, choosing various values for the maximum allowable

pressure drop, resulting in different values for the y-axis and consequently G', A, and ultimately D,

are summarized in Table 10.

40


Table 10 – Sensitivity Analysis Results for the Diameter Calculations of the Stripper

ΔP [ Nm2 ] Rings [¿2] y G' A [m2] D [m]

200 65 0.055 0.9157 37.8 6.9

400 65 0.09 1.82 19.03 4.9

800 65 0.15 2.35 14.7 4.32

The smallest diameter is obtained with an allowable pressure drop of 800 N/m2 per meter of

packed depth. This results in the most affordable packed tower of the three options, thus this is the

value chosen.

The same analysis was performed on the hydrogen sulfide scrubber, and the following values

were obtained:

x-axis = 0.01y-axis = 0.055

G'=0.9161 kgm2 s

A = 0.08 m2

D = 0.31 m

Flash Evaporator

To size the flash evaporator, an approach by Phillip C. Wankat was chosen [20]. This approach

involves the use of an empirical model coupled by rule of thumb values to arrive at a vessel size that

agrees with experimentally obtained data. Because the molar and mass flow rates of the liquid and

gaseous hydrogen sulfide are known, sizing can be performed based on these values.

The first step was to calculate the permissible vapor velocity. This is the maximum vapor

velocity in at the maximum cross-sectional area, and is given by

uperm=K drum√ ρL− ρV

ρV

(38)

After experimenting with various pressure values, it was determined that 0.7 atmosphere

(0.7093 bar) within the flash drum gives the desired separation at the temperature and concentration.

Kdrum is an empirical constant, calculated with the following formula:

41


Kdrum=exp [ A+B ln F lv+C ( ln F lv )2+D ( ln F lv )3+E ( ln F lv )4 ] (39)

where

F lv=W L

W V √ ρV

ρL

(40)

W L and W V are mass flow rates of the liquid and vapor streams, respectively. The constants

are:

A=−1.877478097 C=−0.1870744085 E=−0.0010148518

B=−0.8145804597 D=−0.0145228667

K drum typically ranges from 0.1 to 0.35. Choosing a W V of 320.36 kg/hr, a K drum of 0.37 was

attainable at a temperature of 25 °C and the pressure stated above.

Next, the horizontal area of the flash drum was calculated using the following relationship:

V ( lbmoleshr )=

u perm( fts )( 3600 s

hr ) Ac ( ft2 ) ρV ( lbmft3 )

MW vapor ( lbmlbmole )

(41)

By adding 27.88 kg/hr of air via a blower to the flash drum, the desired separation can occur

with a reasonable horizontal area. With the added air stream, the total vapor flow rate becomes

3230.36 kg/hr and the mole fraction of hydrogen sulfide becomes

0.94 kmolhr

× 34.081 kgkmol

320.36 kg/hr=0.1

Thus the mole fraction of air is 0.9. An average molecular weight could be calculated:

MW vapor= y H 2 S MW H 2 S+ yair MW air (42)

where

MW vapor = average molecular weight of vapory H2 S = vapor mole fraction of hydrogen sulfideyair = vapor mole fraction of airMW H 2 S = molecular weight of hydrogen sulfide

42


MW air = molecular weight of air

The horizontal area for the flash drum was calculated to be 5.25 ft 2. From this, the diameter

was calculated with:

D=√ 4 ( Ac )π

(43)

The diameter was calculated to be 2.59 ft. The diameter is typically rounded up to the next

six-inch increment, so this value becomes 3 ft or 0.91 m. According to Wankat, as a rule of thumb

the ratio of the height to the diameter for a flash drum should range between 3.0 and 5.0. To keep

costs at a minimum, this ratio was chosen to be 3.0. Thus the height was calculated to be 9.0 ft or

2.74 m.

Heat Exchanger

The heat exchanger was sized using a program called HENSAD (Heat Exchanger Network-

Synthesis, Analysis, and Design). This program allows the user to construct a network of heat

exchangers based on the minimum temperature approach of the hot stream and cold stream inputs.

This program also allows for the characterization of the heat exchanger. This is done by entering the

flow rate, inlet and outlet temperature, heat capacity, and film coefficient of both the hot and cold

streams. The information from stream 23 was used as the hot stream and stream 11 was used as the

cold stream for the values entered into HENSAD. With this information, several items were

achieved, the first being a summary table resulting in the number of utilities for the network and the

size of the heat exchanger. This summary can be found in Appendix B-1. A sensitivity analysis can

be done with this as well by changing the minimum temperature approach. The minimum approach

temperature was 10°C. The next item is a cascade can be generated shown in Figure 12.

43


Figure 12 – Cascade Diagram for the heat exchanger designed by HENSAD

With this analysis, it was found that one exchanger would be sufficient for this design

process. Only a cold utility is needed for this network of exchangers with three temperature intervals.

The energy available from the hot stream is 2.6 kW. The cold stream has an available energy of -1.4

kW and therefore requires 1.2 kW for this temperature differential. This program also constructs a

temperature interval diagram, which displays the same data as the cascade diagram but in a different

format. At the end of this analysis and summary, the heat exchanger area is estimated to be 0.13

square-meters. The HENSAD displays 0 m2 for the area of the heat exchanger, however HENSAD

does not display the actual size. This can be found in another part of HENSAD: DT sensitivity plot

(Refer to Appendix B-2). Here the area is displayed and this is the value that was used for the size of

the exchanger. There is one other discrepancy: the number of heat exchangers. HENSAD reports that

two heat exchangers are needed for this process, but this is not the case. In the DT sensitivity plot,

the reported value for the hot utility is 0 kW because a hot utility is not needed. This is also depicted

in Figure 12, indicating that this part of the process requires only one heat exchanger.

5. Economic Analysis The economic evaluation for the sour water stripping facility was taken into careful

consideration with two main categories involved, the capital investment and the operating costs. The

capital investment is the total amount required to purchase, build and other expenses of the facility.

The operating costs are ongoing expenses for day-to-day operations. CAPCOST was utilized to

analyze economic feasibility of this project because it is a computer program that analyzes and

estimates the cost of building a complex plant with variety of equipment selection options.

44


CAPCOST offers a variety of equipment specification options in order to customize the cost

prediction. In each equipment input, CAPCOST allows for specifying capacity, operating pressure

and materials of construction which models the real system very adequately.

5.1 Capital Investment

The capital investment has two primary parts, the fixed capital investment and the working

capital. The fixed capital investment is the money required to purchase and install all equipment

required to operate the plant. This value is sub-divided into direct and indirect costs. Direct costs

include equipment and installation costs. Indirect costs include engineering costs, legal fees,

contingency, general expenses and other fees of this nature.

The project has 2 years of construction and a 10-year project life. Using a design estimate

method with CAPCOST software, the estimated fixed capital investment for this sour water plant

was calculated at approximately $15.6 million and the required working capital was estimated at $1.7

million. These values were calculated using an interest rate of 10% along with a taxation rate of 42%.

This taxation rate is including federal and other local taxes. The CEPCI (chemical engineering plant

cost index), which accounted for inflation over the project life is 579.8 (value in 2015) [31]. The

design for the sour water facility was considered preliminary with no cost of land accounted for due

to the land provided by Deer Park Refinery. In order to reach this evaluation, direct and indirect

project expenses and fees were included in the considered cost.

5.2 Direct Costs

Calculations of the direct costs began with a review of the process flow diagram to determine

all of the equipment needed for the process and their respective sizes, operating temperatures and

pressures was completed. The contents of each piece of equipment were carefully reviewed in

conjunction with the operating temperatures and pressures to identify the appropriate material. The

material chosen for all pieces of equipment was carbon steel. This was chosen because corrosion in

the plant is not a concern at the normal operating conditions (pressure <10 bar, temperature between

40-250oC) and this is the most cost effective material for the vessels. As stated earlier in vessel sizing

the detention time was cut down to 0.5 hours in order to make the facility more profitable compared

to 4-hour detention time by running a sensitivity test.

The utilities and waste were then taken into consideration, from looking at the costs of

utilities in Texas, it was investigated that the Texas electric companies opened to customer choice in

2002. Since then the majority of businesses have taken the opportunity to choose which company

45


they want to get their electricity supply from. The company Change Energy has been chosen for this

facility at 6.7 cents per kWh [32]. A small waste stream that was stated earlier as salt water coming

from the mist scrubber needed to be taken into account for the facility. Due to salt water not being a

harmful waste to the environment or the communities surrounding the facility, the disposal cost was

calculated to be around $.02 per gallon [33].

Raw materials were chosen from extensive research on less corrosive materials. From the

table below it demonstrates all materials that are needed for the plant to run annually. Calcium sulfate

is an additional source of production that has an annual profit from the table at $4 million a year. The

price of each material was chosen from a manufacturer located in Chinas mainland called the

Shandong Dianmei International Trade Co., in order to get an adequate abundance of the material

needed [34].

46


Table 11 – Raw material pricing

Material Price ($/kg) Annual CostCalcium Oxide $0.15 $38,915Calcium Sulfate ($7) ($4,285,450)

Sodium Hypochlorite $0.33 $438,997Calcium Hydroxide $0.22 $954

Reviewing the process flow diagram and control of the facility, the amount of operators and

hours of run time for the plant were taken into consideration. The run time of the plant was decided

based on the parent plant, Deer Park Refinery. The sulfur recovery unit of the parent plant does not

operate all day, to match the run time, the plant will run for 16 hours a day. This will need two 8 hour

shifts of operators. Two sources were used to decide on the amount of operators per shift. The

equation below shows the first source:

NOL=(6.29+31.7 p2+0.23 N np )0.5 (44) [21]

where

p = the number of processing steps involving the handling of particulate solids

Nnp = the number of non-particulate processing steps and includes compression, heating and cooling,

mixing, and reaction.

For the uses for the facility it has been decided that Nnp, would be calculated by the amount of vessels

the facility has. This equation calculated to have an average of 17 operators per shift, this amount of

operators didn’t seem reasonable. Researching this equation, this required the plant to run 24 hours a

day and 7 days a week, the facility only runs 16 hours a day so this equation could not be used. With

further research it was chosen that only 5 operators were needed per shift along with two technicians

assigned to the first shift when maintenance and system updates are needed [35]. The average salary

for an operator at this facility is $24 per hour along with a technician at $16 per hour [36}. A single

operator/technician will work on average 49 weeks a year with 3 weeks off for vacation and sick

leave, five 8 hour shifts a week. Table 11 summarizes all the costs stated earlier for the annual

operating costs for the facility. These costs are only for the direct costs and no other costs that were

taken into account.

Table 11-Total Direct Costs for the Facility

47


Cost of Operating Labor (COL): $ 533,120.00Cost of Utilities (CUT): $ 1,270.00Cost of Waste Treatment (CWT): $ 1,490.91Cost of Raw Materials (CRM): $ 478,867.00Cost of Equipment: $ 11,575,140.00Total: $ 12,589,887.91

48


5.3 Indirect Costs

The indirect costs include two types of costs, fixed manufacturing costs and the general

expenses. The fixed manufacturing costs are independent of changes in production rate. They include

property taxes, insurance, and depreciation, which are charged at constant rates even when the plant

is not in operation. The general expenses have cost that represents an overhead burden that is

necessary to carry out business functions. They include management, sales, financing, and research

functions. General expenses seldom vary with production level. However, items such as research and

development and selling costs may decrease if extended periods of low production levels occur.

Fixed manufacturing costs were first calculated looking into the depreciation. Since the value

or worth of his facility decreases with time, this value has to be taken into account. Some of the

equipment also wears out and has to be replaced during the life of the facility. Although there are

different types of depreciation, two types were looked at for the sour water facility. The first

depreciation as stated earlier is the fixed capital investment at $15.6 million, taken from CAPCOST.

This number represents the fixed capital investment to build the plant minus the cost of land and

represents the depreciable capital investment. Another type of depreciation taken from CAPCOST is

the salvage value at $1.6 million. This represents the fixed capital investment of the plant, minus the

value of the land, evaluated at the end of the plant life, for this facility plant life was assumed to be

around 10 years. Usually, the equipment salvage value represents a small fraction of the initial fixed

capital investment. For the sour water facility, it is assumed that the salvage value of the equipment is

zero. The total capital depreciation is calculated from the given equation:

D=FC I L−S (45) [21]

where

FCIL = fixed capital investment

S = salvage value

The total capital depreciation from equation 45 is $14 million. From there local taxes and insurance

were able to be taken into account. These costs are the costs that are associated with property taxes

and liability insurance. This costs is based on the plant location and severity of the process, using the

equation below the local taxes and insurance is annually $500 thousand.

(0.014−.05 ) FC I L (46) [21]

49


Plant overhead costs also were taken into account at $940 thousand annually. This cost is

considered the “catch-all costs” that are associated with factory expenses. These costs involve payroll

and accounting services, cafeteria and any recreation facilities, payroll overhead and employee

benefits, general engineering costs and things of this nature. This cost was calculated from using the

equation 47.

0.708 COL+0.036 FC I L (47) [21]

where

COL = cost of operating labor

Table 12 below summarizes the costs stated earlier. Each of these costs represents the fixed

manufacturing costs associated with the facility.

Table 12- Total Fixed Costs

General

expenses were calculated with costs associated with management-level and administrative activities

that is not directly related to the manufacturing process. Table 13 below illustrates the different costs

associated with general expenses. Administration costs are costs that include salaries, other

administration, buildings and other related activities. Equation 48 was used to calculate

administration costs.

50

Fixed Capital Investment (FCI): $ 15,600,000.00Salvage: $ 1,560,000.00Total Capital Depreciation: $ 14,040,000.00Local Taxes & Insurance: $ 499,200.00Plant Overhead Costs: $ 939,048.00Total: $ 15,478,248.00


0.177 COL+0.009 FC I L (48) [21]

Distribution and selling costs are costs of sales and marketing required to sell chemical products.

This cost also includes salaries and other miscellaneous costs. Two equations were used in order to

calculate this expense. The first equation, equation 49 was used in order to calculate total

manufacturing costs (COM). Equation 50 was then used to get the distribution and selling cost.

COM=0.18 FC I L+2.76 COL+1.23 (CUT+CWT+CRM ) (49) [21]

(0.02−0.2 )COM (50) [21]

where

CUT = cost of utilities

Cwt = cost of waste treatment

CRM = cost of raw materials

From equation 49 the cost total manufacturing costs is annually $5 million. This costs was then

plugged into equation 50.

Total distribution and selling costs calculated to be $500 thousand annually. Research and

development are costs of research activities related to the process and product. This includes salaries

and funds for research-related equipment and supplies.

0.05 COM (51) [21]

Equation 51 was then used and gave an annual expense of $200 thousand.

Table 13 - General Expenses

Administration Costs: $ 234,762.00 Distribution and Selling: $ 535,899.00 Research and Development: $ 243,590.00 Total: $ 1,014,252.00

51


5.4 Profitability

The Deer Park Refinery pays approximately $2.5 million annually to dispose its wastewater

and obtain clean water for steam generation. The cost of disposal for the wastewater includes

transportation to the treatment plant and treatment fees. After the sour water treatment plant is

implemented as an auxiliary plant of this refinery, it proposes to save the plant approximately $1.4

million annually. Table 14 verifies that this savings is possible. As demonstrated in the table, before

the sour water facility, the refinery pays most of its yearly operational costs for transportation and

disposal of sour water. The transportation costs include the cost of extra operators. These operators

would be in charge of having to load the water trucks with the sour water, dispose of the sour water

to a water treatment facility and also haul clean water back to Deer Park Refinery. Not only are there

extra operators this transportation cost also includes 4 water trucks that would be needed to dispose

of the sour water and bring back the clean water to the facility. Having to have the water trucks to

use, gas, insurance and maintenance fees would be needed [37]. The disposal cost is also higher than

using the sour water facility at $.04 per gallon [33]. Not only would the facility have to pay for the

disposal of water, but Texas has extra fines just to haul disposal water at $.07 per gallon [33]. This

price is due to permits and problems that Texas has due to the problems with illegal water dumping.

With the sour water facility being an auxiliary plant, the disposal and transportation costs will be

eliminated. The cost of clean water is at 0.04 cents per gallon using the sour water facility [38].

Table 14 - Comparing Operating Costs before the Sour Water Facility and After

Costs Before Us Costs With UsTransportation $1,339,272.00Clean Water Cost $72,576.00 $72,332.00Disposal Water Cost $691,200.00Operational Cost $376,320.00 $533,120.00Utilities $1,270.00Raw Materials $478,867.00Total $2,479,368.00 $1,085,589.00

52


To determine if the sour water facility was profitable, the CAPCOST software was used to input

equipment, all raw materials and direct costs as stated earlier to generate a cash flow diagram and

table illustrated below. Table 15 demonstrates the process of cash flow calculations. The numbers

present value in million dollars. According to the table, at year 12 of the project, the plant will make

a net profit of $1.33 million. The profitability criteria were discounted to the present values (at time 0

of the project) to account for the time value of money and allow for a thorough profitability analysis.

Table 15- Cash Flow Table

Year Investment dk

FCIL-Sdk R COMd

(R-COMd-

dk)*(1-t)+dk

Cash Flow (Non-

discounted)Cash Flow

(discounted)

Cumulative Cash Flow

(discounted)

Cumulative Cash Flow

(Non-discounted)

0 0.00 15.60 0.00 0.00 0.00 0.00 1 9.36 15.60 (9.36) (8.51) (8.51) (9.36)2 7.90 15.60 (7.90) (6.53) (15.04) (17.26)3 3.12 12.48 8.73 4.87 3.55 3.55 2.67 (12.37) (13.71)4 4.99 7.49 8.73 4.87 4.33 4.33 2.96 (9.41) (9.38)5 3.00 4.49 8.73 4.87 3.50 3.50 2.17 (7.24) (5.88)6 1.79 2.70 8.73 4.87 2.99 2.99 1.69 (5.55) (2.89)7 1.79 0.90 8.73 4.87 2.99 2.99 1.54 (4.02) 0.10 8 0.90 - 8.73 4.87 2.62 2.62 1.22 (2.80) 2.72 9 - 8.73 4.87 2.24 2.24 0.95 (1.85) 4.96

10 - 8.73 4.87 2.24 2.24 0.86 (0.98) 7.20 11 - 8.73 4.87 2.24 2.24 0.78 (0.20) 9.43 12 - 8.73 4.87 3.14 4.80 1.53 1.33 14.24

53


Based on the table above, a cash flow diagram was generated to reflect the cumulative discounted cash flow diagram.

Figure 13 - Cash Flow Diagram

54

-1 0 1 2 3 4 5 6 7 8 9 10 11 12 13-16.0-14.0-12.0-10.0

-8.0-6.0-4.0-2.00.02.04.0

Project Life (Years)

Proj

ect V

alue

(mill

ions

of d

olla

rs)


From the figure you can see the first two years is a negative trend in the flow diagram, this is

due to the first two years being the facilities construction period. This period includes the fixed

capital expenditures for purchasing and installing the equipment required to run the plant. At the end

of the second year, construction is finished and the sour water facility is started. At this point, the

additional expenditure for working capital required to float the first few months of operations is

shown. The working capital from CAPCOST is $1,660,000, this is a one-time expense at the startup

of the plant and will be recovered at the end of the project which is a 10-year project life. After

startup, the process begins to generate finished products for sale, in this case for the sour water

facility there is calcium sulfate as a side product and the clean water that is sent back to Deer Park

Refinery for further steam generation. From these products the trend of the cash flow diagram starts

to become positive. The revenue for the first year after startup is less than in subsequent years this is

due to the effect of depreciation. The depreciation that was chosen was MACRS (modified

accelerated cost recovery system) for 5 years. This class life was chosen because the equipment life

in the facility is assumed to have a class life of about 10 years with no salvage value. Therefore, the

capital investment is depreciated over a shorter period of time (5 years). MACRS depreciation was

also chosen because it is better to depreciate an investment as soon as possible. This is because the

more the depreciation is in a given year, the less taxes paid. The discounted payback period is 7.6

years after the 2-year start up shown in figure 13. The payback period means that after 9.6 years

including the construction period, the project will break even and the net present value is 0. After 9.6

years the sour water facility will start making profit. The discounted cash flow rate of return

(DCFROR) was also reported in CAPCOST at 11.9%. DCFROR is the interest rate at which the

project breaks even at NPV = 0, with all cash flow discounted to the present value. Since this rate is

greater than the interest rate proposed in the project, this is another evidence to prove that the sour

water facility will be profitable.

5.5 Risk Analysis

55


In order to examine the economic potential of this project, a Monte-Carlo stimulation was

generated in CAPCOST. This program allowed the risks associated with economic parameters to be

evaluated. The Monte-Carlo analysis method is used to value and analyze the various sources of

uncertainty affecting the facilities value, and determining the facilities average value over the range

of resultant outcomes. These variables include FCIL, working capital, interest rate and other forms of

this nature. The variance accounts for uncertainty in the expected values of these parameters. Among

the uncertainty of the parameters, the largest uncertainty range came from working capital and

salvage value as their lower limits, -50% and -80%, respectively. Using these estimations, the Monte-

Carlo analysis presented the distribution of profitability factors such as NPV and DCFROR.

-8 -6 -4 -2 0 2 4 6 80

250

500

750

1000

Net Present Value (millions of dollars)

Cum

ulat

ive

Num

ber o

f Dat

a P

oint

s

Figure 14 - Monte-Carlo analysis demonstrates the probability distribution of NPV

Figure 14 shows the probability distribution of NPV with a low NPV of -$5.9 million and a

high NPV of $6.9 million. There is 60% chance it can obtain the expected NPV of $1.33 million. At

NPV = 0, the Monte-Carlo analysis predicted a 30% probability of the plant going into deficit and a

40% chance that the project would make more than the expected NPV. Since the chance of the plant

being non-profitable is low, Deer Park Refinery should proceed with the project and expect profit the

in future.

56


0.00 0.05 0.10 0.15 0.20 0.250

250

500

750

1000

DCFROR

Cum

ulat

ive

Num

ber o

f Dat

a P

oint

s

Figure 15 - Monte-Carlo analysis demonstrates the probability distribution of DCFROR

Figure 15 illustrates another profitability criterion that the Monte-Carlo predicted, the DCFROR with

low DCFROR at 2% and high DCFROR at 20%. This figure also demonstrates that there is a 55%

chance that the facility can obtain the expected DCFROR at 11.9% and a 45% chance the project will

get a higher DCFROR which would make the sour water facility even more profitable.

6. Safety 6.1 Plan Layout

The DSS water treatment plant is based in Deer Park, TX and will function as an auxiliary

plant to the Deer Park Refinery. The water treatment plant will be located in the northeast corner of

the refinery on the south end of the Houston South Canal. Figure 13 shows a plot plan of the DSS

water treatment facility.

57


Figure 16– DSS water treatment facility layout

The water treatment plant was designed in order to comply with safety and environmental

regulations. All of the vessels in our plant, are either reactors, towers or filters which require a

minimum distance to be considered safe. Since inherited safety is the top priority of the modern

chemical process design, more space was allocated between vessels in order to allow machinery for

maintenance, operators and technicians as well as precautionary space in case of an incident. A pipe

rack was added to the design in order to provide structural support to pipes, power lines and

instrument cables. The plant is conveniently located on the intersection of two roads for ease of

maintenance and access.

6.2 Control system implementation

Controller system have become an important aspect in chemical process designing since it

offers a more stable and consistent process, as well as minimization of hazards and potential

incidents. For this design, various controllers were implement in order to avoid any deviations in the

process as well as increasing the protection of the workers on site. Such controllers include flow

indicator, pressure controllers and pH indicators. Figure 17 illustrates a process flow diagrams with

all the implemented controllers.

Various indicators were placed in the system in order to effectively monitor variances of some

parameters. These indicators are equipped with high and low alarms, which will alert the operators.

pH indicators (PhI) were added strategically where pH changes are observed in the system,

58


specifically in R-100 where the system is introduced to calcium hydroxide, in the recarbonation tank

(R-104) where hydrogen sulfide is created and, on stream 21 before the brine is separated from the

calcium sulfate. Flow indicators (FI) were placed in piping where the flow rate is critical for the

operation of the facility. Pressure indicators (PI) were placed in areas were gases are present in the

system, making their surveillance easier to monitor. Level indicators (LI) were placed in areas where

liquids are present and could potentially accumulate if there is an incident in the process.

Inducer controllers were added on some vessels in order to relief pressure or flow. These are

continuous controllers which control a valve to open or close as needed to maintain a constant

condition. For example, the lime recovery system (R-103), which has a heating element associated

with it, has a temperature induced controller (TIC) to control the flow in stream 26 to maintain a

constant temperature. Another example is the recarbonation tank (R-104) which has carbon dioxide

as a feed and produces hydrogen sulfide, both which are gases, has a pressure induced controller

(PIC) to maintain a constant pressure in the vessel by controlling a valve on stream 13.

59

Figure 17 – Process flow diagram of the DSS water treatment facility with implemented controllers. Legend: FI - Flow indicator, LIC - Level induced controller, LI - Level indicator, FIC - Flow induced controller, PhI - pH indicator, TIC - Temperature induced controller, PIC

– Pressure induced controller.

6.3 HAZOP analysis

In order to design a safer chemical engineering process, the Process Safety Management of

Highly Hazardous Chemicals regulation (29 CFR 1910.119) requires employers to perform various

analysis to protect not such its employees, but the community surrounding any plant or facility [21].

The most widely used process hazards analysis technique in the chemical process industries

is HAZOP. This is a modified brainstorming technique for identifying and resolving process hazards

by considering unusual or extreme occurrences at a given point in any process. HAZOP is a very

rigorous technique and may require the action of many people to ensure a process is safe.

The first step when performing a HAZOP analysis is to identify all the nodes, or points of

interest, in a given process and determine what are the normal operating conditions or purpose of this

node. This is called the intention. Next, the relevant conditions of the node are explored by using a

guide word (flow/temperature, etc.) to identify any deviations.

For this design project, the team performed a HAZOP analysis on two nodes of interest by

investigating one specific parameter, which in both cases, was flow rate. The deviations studied were

no flow through the node, less flow through the node and more flow through the node. Figure 18

illustrates a PFD with all the nodes in the process.

Figure 18 – DSS Sour water treatment facility process flow diagram with the identification of nodes for the HAZOP analysis

For the HAZOP analysis on this design, the nodes chosen were node 18 located on stream 16

which is the overhead product of the flash evaporator (T-105 A/B) and node 27 located on stream 25

which is one of the exiting streams from the lime recovery system (R-103). Tables 16 and 17

summarize the results of the HAZOP analysis on nodes 18 and 27, respectively.

Table 16 – HAZOP analysis results for node 18


Table 17 – HAZOP analysis results for node 27

As seen in Table 16, the main hazards related to node 18 are pressure buildup or overfill in T-

105 A/B due to blockages in the flow or complete closure of stream 16. While there are existing

provisions to alert operators of these scenarios, such as pressure and flow indicators, it was suggested

that pressure relief valves be placed on T-105 A/B. Since the indicators are equipped with ow and

high deviation alarms, the activation of the pressure relief valve would serve as a second warning for

operators and would require compliance to established emergency procedures.

64


Similarly, in Table 17, the main hazards associated with node 27 are overfill and pressure

buildup in R-103 due to various reasons such as issues in the flow of stream 25, a malfunction in Fr-

103 or deviations in the operating temperature due to a malfunction in the heat exchanger being

employed in R-103. Some of the existing precautions include temperature and flow indicators on

stream 25 and stream 11 with implementation of level indicators on Fr-103 for suggested action.

By performing a HAZOP analysis periodically throughout the life of the project, safer,

consistent and reliable sour water treatment can be designed which creates a safer environment for

the operators as well as increases the economic profitability of the plant.

6.4 Plant Emergency Procedures

By developing an efficient plot plan, which allows the entry of operators and/or machinery in

order to inspect the units and by performing preventive safety actions such as the usage of automated

controllers and routinely HAZOP analysis on the process, DSS Inc. can establish a safer environment

for the operation in the sour water treatment facility.

Nevertheless, unexpected deviations and incidents in the plant can occur at any time and the

plant personnel must be educated on what to do in an emergency. As previously mentioned, the plant

is equipped with various controllers and sensors to alert the operators from any deviations of the

standard operating conditions. These indicators are equipped with a high and low alarm which will

alert the control tower of any major deviations to the system. This alert system acts as a primary

warning sign and should promptly encourage the operators for investigations on the causes of these

deviations. As summarized in the previous section, one of the first suggested steps to improve the

safety of the facility is to install pressure relief valves on some vessels. These will act as secondary

alarm system for the operators and should mean that the process is near a dangerous operating

condition and if necessary a complete evacuation.

It is worth mentioning that our plant deals with chemicals with mild health concerns, with

sulfuric acid and calcium oxide and sodium hydroxide being the chemicals with most health

concerns. The complete list of chemicals present at any point in the process, along with their

associated properties, are summarized in Table 18.

65


Table 18 – Properties of the various chemicals present in the DSS sour water treatment plantChemical Properties and Hazards

Chemical Name FormulaPhysical

StateS, L, G

NFPA Ratings* Hazardous Precaution

Toxicology*

H F S Sp. TWA PELCalcium hydroxide Ca(OH)2 S 2 0 0 Corrosive, irritant 5 mg/m3 NA

Calcium carbonate CaCO3 S 0 0 0 NA 5-15 mg/m3 NA

Calcium oxide CaO S 3 0 0 Corrosive, irritant 2 mg/m3 NA

Calcium sulfate CaSO4 S 0 0 0 NA 5-15 mg/m3 NA

Carbon dioxide CO2 G 0 0 0 Gases under pressure 9000 mg/m3

5000 ppm

Hydrogen sulfide H2S G 4 4 0Flammable, gases

under pressure, toxic, irritant

15 mg/m3 1-50 ppm

Sodium sulfate Na2SO4 S 0 0 0 NA NA NASodium chloride NaCl S 1 0 0 NA NA NASodium hypochlorite NaOCl L 3 0 0 Corrosive, toxic to

environment NA NA

Sodium hydroxide NaOH S 3 0 0 Corrosive 2 mg/m3 NA

Sulfuric acid H2SO4 L 3 0 0 Corrosive 0.2-1 mg/m3 NA

Ammonia NH3 G 3 0 0Flammable, gases

under pressure, toxic, irritant

18-35 mg/m3

25-50 ppm

Propane C3H8 S 0 4 0 Flammable, irritant 1800 mg/m3

1000 ppm

Water H2O L 0 0 0 NA NA NA*NFPA Ratings: H – Health, F – Flammability, S – Stability, Sp. – Special*Toxicology: TWA – Time Weighted Average, PEL – Permissible Exposure Limit

In case of an emergency, a valve before stream 1 should be closed to restrict the feed of sour

water into the process and any heating elements, mainly the one present for R-103, should be turned

off to avoid any explosion hazards. The only combustible chemicals in the system that have any

concerns are the hydrogen sulfide and the propane. Propane is remove almost immediately by the

activated carbon filters and render this non-flammable. The hydrogen sulfide is obtained in the

recarbonation tank and is eventually turned into calcium sulfide. The pipes where this material is

present will be equipped with pressure and flow indicators which signal the presence of leaks.

In case of an incident, such as a fire or leak, and if not able to be contained immediately, the

heated elements and feed to the system should be turned off and the area should be evacuated

immediately. Once the incident is contained, the vessels should be inspected to ensure there is no

damage to the unit to ensure a successful start-up. It is worth mentioning that the vessels will be

constructed in order to withstand a moderate earthquake (Scale 5.0-5.9) which historically has never

occurred in Texas [22].

66


7. Environmental & Community Awareness7.1 Regulations

During the past decade, the emphasis on designing an environmentally friendly chemical

plant has become a priority. Some of the methods used in current chemical plants to reduce the

damage to the environment is minimizing the generated waste by recycling the excess, selling by

products to create revenue or transportation for waste treatment at another facility.

In order to regulate and monitor the environmental effect of chemical facilities, the U.S.

Environmental Protection Agency (EPA) has designed various regulations that any process must

follow. The EPA also establishes a basis for state specific regulations. For this design project, the

Texas Administrative Code (TAC) was carefully inspected to ensure the facility follows the

regulations established by the EPA. The complete list of regulations specific to this design project are

as follows:

1 Tex. Admin. Code § 15 16 Tex. Admin. Code § 70 28 Tex. Admin. Code § 28.2.42 28 Tex. Admin. Code § 28.2.43 28 Tex. Admin. Code § 28.2.45 28 Tex. Admin. Code § 28.2.47 30 Tex. Admin. Code § 30.1 31 Tex. Admin. Code § 31.9.286 37 Tex. Admin. Code § 13

The OSHA regulations for standard industries (29 CFR §1910) has also been inspected to ensure

following of regulations.

7.2 Community Awareness

This design project has also been designed in order to ensure the safety of the surrounding

community by accomplishing various tasks. The first task was implemented when designing the plot

plan. The facility was strategically placed where residential area is 1.2 miles away from the plant and

was also placed next to the Houston Ship Canal for access in case of a major incident. The plant was

also designed taking into account the wind direction, which for Deer Park, TX is predominately

towards the NW direction. By using this fact, we can ensure that any gas being to the atmosphere,

specifically the ammonia in the air, will be carried away from the residential areas or neighboring

facilities.

67


Another task to minimize the impact to the community is reducing the waste emitted by the

plant. The final products produced by the plant are the water, being sent back to the plant, brine

which is used as a heat source and then disposed to a nearby canal, calcium sulfide, which is sold,

and air with small quantities of ammonia. As mentioned in the previous section, there are various

regulations which limit the amount of waste created. This is one way the waste is controlled. The

other way is by selling the products where revenue can be made. The brine, which is eventually

disposed to a canal has no negative effects to the environment or the surrounding communities.

8. Conclusions

8.1 Project Goals Satisfaction

The design of sour water treatment at Deer Park Refinery was successfully assessed and proven to be

feasible. The project was designed based on the following technical goals:

Reduce contaminant concentrations to safe levels: <20ppm NH3, <3 ppm H2S and traced

amount of propane

Design an effective scrubber and stripper to remove H2S and NH3, respectively, at minimum

costs

Size all vessels, pumps and blowers and position them in an efficient configuration without

jeopardizing the safety of plant personnel, adjacent plant operations and the environment.

Design and implement a complete control system to monitor temperature, pressure and

concentration for the wastewater treatment operation

As results, these goals were achieved and justified in the project based on the models employed.

After an extensive research on waste water treatment, the team came up with a working PFD which

was assessed in Aspen Stimulation Program. The materials used in the plant were chosen to be most

efficient and safe for the process. HAZOP analysis was also evaluated to ensure the safety of the

overall project. By implementing the sour water treatment plant design by DDS team, Deer Park will

be able to treat the sour water coming from its sulfur recovery unit and obtain an efficient amount of

clean for steam generation in a faster production time. The project also results in the protection of the

environment from the dumping of harmful wastewater streams. A few highlights in the design of

sour water treatment should be revisited. Ca(OH)2 was efficiently regenerated from recarbonation

tank to feed the lime recovery system, eventually got recycled to deprotonate NH4+ ions. Heat

68


generated in brine from the exothermic reaction, where sulfuric acid was neutralized, transferred to

the endothermic reaction in the lime recovery system as source of energy.

8.2 Profitability Achievement

After the technical goals were achieved successfully. Economic analysis was studied to ensure the

project makes profit and is worth the investment from Deer Park Refinery. Based on the taxation rate

of 42%, annual interest rate of 10% and 5 years MACRS depreciation, the refinery expects the sour

water treatment plant to make a profit of $1.33 million after 2-year start-up and 10 years of project

life at the DCFROR of 11.9%. Sensitivity study of the economic was evaluated in Monte-Carlo

Simulation to account for any possible variation in key economic factors over the course of this

project. Based on the probability distribution in NPV and DCFROR, it has been proven that the

project most likely to gain the expected profit. Moreover, the Deer Park refinery would benefit by a

net annual savings of $1.4 million in operation to choose the implementation of the sour water

facility over current means of disposing of their sour water and purchasing clean water for steam

generation. Another great source of revenue for the plant is the amount of calcium sulfate

precipitated out of the system. In conclusion, the sour water treatment plant is an absolute alternative

solution to clean sour water at the Deer Park Refinery efficiently, safely and economically.

9. Recommendations

9.1 Bench-model Experiment

Throughout the design of the project, wastewater treatment books were referenced to obtain

heuristics for the detention times for the various chemical reactors.[] These detention times affected

the vessel sizes and consequently the cost of materials for the project. Initially, heuristically-

determined detention times of four hours were used for the reactors, resulting in material costs which

prevented the project from being profitable. It was decided to do a sensitivity analysis to investigate

the effect of vessel-sizing on the economics by varying the detention times. The detention times were

cut in half and the reactors were resized accordingly. The economics was run using these vessel sizes

and the process still wasn’t profitable. The detention times were then subsequently lowered and the

process was repeated until the project became profitable. It was then realized that the detention times

initially used were heuristics for larger wastewater treatment facilities (four million gallons/day) as

mentioned in section 4.5 – Vessel Sizing. Because our facility treats 72,000 gal/day, we made the

assumption that by cutting our detention times down, our vessels would still function properly for our

wastewater throughput.

69


One reason the detention times were determined via a heuristic means rather than by using empirical

kinetics data for the chemical reactions involved was that sufficient empirical data wasn’t available

for all of the reactions in the system. One recommendation is to create bench scale models of the

reactors in the system in order to obtain kinetics parameters for the various chemical reactions in

order to calculate more accurate detention times for the reactions. This would result in vessel sizes

more appropriate for each individual reaction or set of simultaneously occurring reactions. It is likely

that these detention times would be close to those assumed for the vessel sizing for this project.

However, if this study were carried out and the detention times were large enough to render the cost

of materials beyond the realm of a profitable operation, the cost of materials could be a potential

show-stopper and the team would have to search for other ways lower costs.

9.2 Service Expansion

Since the sour water treatment plant has a great potential in cleaning sour water efficiently and

making good profit, it is recommended that the service be expanded to become a hub for surrounding

refineries. Within 20-mile radius around the Deer Park Refinery, there are 17 refineries for which the

DDS company can provide the sour water treatment service. In order to expand the service, vessel

sizing will have to be adjusted according to the throughput flow rate. This is possible because the

designed plot plan shows that there is room for an expansion of operating vessels in order to upscale

the vessels to accommodate for the sour water from other pants within the hub. With this expansion,

the plant will expect to earn more profit.

70


References [1] Addington, Fitz, Lunsford, Lyddon. "Sour Water: Where it comes from and how to handle it". Bryan Research and Engineering, Inc. Dr. Ing. Mariana Siwek, Verfahrenstechnik und Automatisierung GmbH. <http://www.bre.com/portals/0/technicalarticles/sws_paper_finala4.pdf >. Accessed 3/18/2016

[2] Environmental Impacts. Chevron Toxico: The Campaign for Environmental Justice in Ecuador. Web. Accessed 4/7/2016. <http://chevrontoxico.com/about/environmental-impacts> .

[3] Gary, J.H. and Handwerk, G.E. (1984). Petroleum Refining Technology and Economics (2nd ed.). Marcel Dekker, Inc.

[4] Shell Deer Park Fact Sheet. About Shell Deer Park. Shell, United States. Web. Accessed 4/7/2016. <file:///C:/Users/Seth/Downloads/shell-deer-park-fact-sheet-2014.pdf> .

[5] Removing Hydrocarbons From Water. PetroWiki. SPE International. Web. Accessed 4/7/2016

[6] Activated Carbon Adsorption Ratings. Sentry Air Systems, Inc. Web. Accessed 4/7/2016

[7] Mazille, F., Spuhler, D. Adsorption (Activated Carbon). Sustainable Sanitation and Water Management. [PDF]. Retrieved from <http://www.sswm.info/content/adsorption-activated-carbon>.

[8] Tro, N. (2011). Chemistry: A Molecular Approach. Upper Saddle River, New Jersey: Pearson Education, Inc.

[9] Stevens, D., Mosher, A. Fundamentals of Sour Water Stripping [PowerPoint]. Retrieved from http://www.kps-engr.com/pdf/Fundamentals%20of%20Sour%20Water%20Stripping.pdf

[10] Brooks, M. (1999). Breakpoint Chlorination as an Alternate Means of Ammonia-Nitrogen Removal at a Water Reclamation Plant. (Master's thesis). Retrieved from Northern Virginia Center.

[11] Branch Environmental Corp. Ammonia Stripping. [PDF]. Retrieved from <http://www.branchenv.com/> .

[12] Zhou, S., Turnbull, A. Steam Turbine Operating Conditions, Chemistry of Condensates, and Environment Assisted Cracking – A Critical Review. [PDF] Retrieved from <http://publications.npl.co.uk/npl_web/pdf/matc95.pdf>.

[13] Water Treatment Plant Design. 2nd ed. New York: McGraw-Hill, 1990. Print.

[14] Mixing Fluids. The Engineering Toolbox. Web. Accessed 4/12/2016. Retrieved from <http://www.engineeringtoolbox.com/mixing-fluids-temperature-mass-d_1785.html>.

[15] McCabe, W., Smith, J., Harriot, P. (2005). Unit Operations of Chemical Engineering. New York, NY: McGraw-Hill.

[16] Fuller, J. What’s the Difference Between Stalactites and Stalagmites? How Stuff Works. Retrieved from <http://science.howstuffworks.com/environmental/earth/geology/stalactite-stalagmite1.htm>.

[17] Reynolds, Tom D., and Paul A. Richards. Unit Operations and Processes in Environmental Engineering. Boston: PWS Pub., 1996. Print.

[18] Weber, Walter J. Physicochemical Processes for Water Quality Control. New York: Wiley-Interscience, 1972. Print.

71


[19] Welty, Rorrer & Foster. Fundamentals of Momentum, Heat and Mass Transfer. River Street, Hoboken, NJ: Wiley & Sons, Inc.

[20] Wankat, P. Separation Process Engineering. 3rd Ed. Upper Saddle River, NJ: Prentice Hall, 2012. Print.

[21] Turton, R., Bailie, R.C., Whiting, W.B., Shaeiwitz, J.A., Bhattacharyya, D. 2014. Analysis, Synthesis, and Design of Chemical Processes. Prentice Hall. 4th Ed. Ch. 26. p. 901

[22] Earthquaketrack.com. Biggest Earthquakes near Texas, US. Web. 4/28/16

[23] McGraw-Hill Series in Water Resources and Environmental Engineering. New York: S.n., 1973. Print.

[24] Water Treatment Plant Design. 2nd ed. New York: McGraw-Hill, 1990. Print.

[25] Sundstrom, Donald W., and Herbert E. Klei. Wastewater Treatment. Englewood Cliffs, NJ: Prentice-Hall, 1979. Print.

[26] Patterson, James W. Industrial Wastewater Treatment Technology. Boston: Butterworth, 1985. Print.

[27] Galbraith, Kate. “Interactive Graphic: Comparing Cost of Water.” Texas Tribune web. June 8, 2012. Date Accessed March 2, 2016. <http://www.texastribune.org/library/data/cheap-water-in-texas/>

[28] Save on Energy. Date accessed March 2, 2016. <http://www.saveonenergy.com/compareoffers.html?zip=77009>

[29] Clean water sources and special waste in Texas. Date accessed March 4, 2016. http://www.tceq.texas.gov/permitting/waste_permitslmfo_permits/msw_specialwaste.html

[30] Silberberg, M. (2007). Principles of General Chemistry. New York, NY: McGraw-Hill Higher Education.

[31] Economic Indicators. About CEPCI number. Web. Accessed 4/11/2016. H

<http://www.isr.umd.edu/~adomaiti/chbe-446/literature/ChECostIndexJan2015.pdf>

[32] Save on Energy. About energy costs, United States. Web. Accessed 3/2/2016.

<http://www.aveonenergy.com/compareoffers.html?zip=77009>

[33] Texas Commission. About disposal costs, United States. Web. Accessed 3/4/2016. <http://www.tceq.texas.gov/permitting/wate_permits/msw_permits/msw_specialwaste.html>

[34] Albaba. Chemical Production, China. Web. Accessed 4/11/2016. <http://www.alibaba.com/product-detail/factory.html>

[35] Baruth, E. E. "25/Preparing a Staffing Plan." Water Treatment Plant Design. 4th ed. New York: McGraw-Hill, 2005. 25.20-5.21. Print.

[36] Salaries. United States. Web. Accessed 4/11/2016. <http://www.salary.com/chemical-plant-operator-hourly-wages.html> .

[37] Houston Gas Prices. About cost of gas, Houston, United States. Web. Accessed 4/14/2016. <http://houstongasprices.com/chevron_Gas_Stations/Houston_-_Houston_-_central/34023/index.aspx>

72


[38] Interactive Graphic: Comparing Cost of Water. Texas Tribune, United States. Web. Accessed 3/2/2016. <http://www.texaatribune.org/library/data/cheap-water-in-texas/>

73


Appendix A-1: Vessel Sizing CalculationsVessel Volumetric

Flow RateTau Volume Area Height Diameter

[L/hr] [hr] [m3] [m2] [m] [m]Fr 103 15451.8 0.5 7.73 2.53 3.05 1.80Fr-104 30792.6 0.5 15.40 5.05 3.05 2.54Fr-105 17154 0.5 8.58 2.81 3.05 1.89R-103 459 0.5 0.23 0.13 1.83 0.40R-104 50538.6 0.5 25.27 6.91 3.66 2.97R-106 15451.2 0.0625 0.97 0.32 3.05 0.64R-102 748.2 0.5 0.37 0.20 1.83 0.51R-105 21612 0.5 10.81 2.95 3.66 1.94R-101 77217.6 0.5 38.61 10.56 3.66 3.67T-101 - - 115.29 37.80 3.05 6.94T-103 - - 0.12 0.08 1.50 0.32T-105 - - - 0.48 2.37 0.79

Appendix A-2: Net Enthalpy of Reaction by VesselVessel Net ΔH°rxn [kJ/hr] Resulting ΔT (°C)R-101 -10,992 0.76R-102 -29,896 27.92R-103 84,528 73.11R-104 -13,656 0.0003R-105 -207,811 106.35R-106 -58,824 0.001

74


Appendix B-1: HENSAD Summary tableMinimum Temperature Approach = 10°C

Hot Stream Data

Mass Flow Cp Temp In Temp Out Stream Enthalpy Film Heat Transf. Coefkg/s kJ/kg/°C °C °C kW W/m2/°C

.0290 4.184 131.3 110.0 2.590 5100.

Cumulative Hot Stream Energy Available = 2.6 kW

Cold Stream Data

Mass Flow Cp Temp In Temp Out Stream Enthalpy Film Heat Transf. Coefkg/s kJ/kg/°C °C °C kW W/m2/°C

.0045 4.184 25.00 98.11 -1.376 5100.

Cumulative Cold Stream Energy Available = -1.4 kW

Data for Generating Temperature Interval Diagram

Number of Temperature Intervals = 3

Interval Temperature Range Excess Heat Cummulative Q°C °C kW kW

A 131.3 110.0 2.590 2.590B 110.0 108.1 0 2.590C 108.1 35.00 -1.376 1.214

This Problem has no Pinch Point

Cold Utility Requirement = 1 kW

Hot Utility Requirement = kW

Minimum Number of Exchanger Required to Accomplish Minimum Utility LoadsIn Special Circumstances the Minimum Required may be Lower than Indicated Below

Number of Exchangers = 2

Data for Composite Enthalpy - Temperature Diagram

Temperature Hot Stream Enthalpy Temperature Cold Stream Enthalpy°C kW °C kW

35.00 .0000 25.00 1.214108.1 .0000 98.11 2.590

110.0 .0000 100.0 2.590131.3 2.590 121.3 2.590

75


Heat Transfer Area for Process Exchangers in Network = 0 m2

Appendix B-2: HENSAD DT Sensitivity Plot Results MinTemp Exch Hot Util Cold Util EAOC Number of Approach Area Load Load Exchangers °C m2 kJ/s kJ/s (1000)$/y

10.00 .13806 0 1.214 .3164 2.000 12.25 .13814 0 1.214 .3165 2.000 14.50 .13806 0 1.214 .3164 2.000 16.75 .13806 0 1.214 .3164 2.000 19.00 .13806 0 1.214 .3164 2.000 21.25 .13806 0 1.214 .3164 2.000 23.50 .13806 0 1.214 .3164 2.000 25.75 .13806 0 1.214 .3164 2.000 28.00 .13806 0 1.214 .3164 2.000 30.25 .13806 0 1.214 .3164 2.000 32.50 .13806 0 1.214 .3164 2.000 34.75 .12462 2.843 1.242 .3074 2.000 37.00 .13339 .0707 1.284 .3312 2.000 39.25 .13888 .1131 1.327 .3506 2.000 41.50 .14794 .1555 1.369 .3743 2.000 43.75 .15715 .1978 1.411 .3979 2.000 46.00 .16650 .2402 1.454 .4213 2.000 48.25 .17598 .2826 1.496 .4446 2.000 50.50 .18324 .3249 1.538 .4653 2.000 52.75 .19000 .3673 1.581 .4852 2.000 55.00 .19678 .4096 1.623 .5050 2.000

76


Appendix B-3: HENSAD Temperature Interval Diagram

77

final report outline team 2_updated version

Documents