Estimate BETX Absorption in TEG Dehydration Process

BETX stands for benzene, ethylbenzene, toluene, and xylene, a group of compounds that all belong to the broader category of Hazardous Air Pollutants (HAPs). Benzene and ethylbenzene are known carcinogens, and have also been shown to cause blood disorders, impact the central nervous system and the reproductive system.  Additionally, Toluene may affect the reproductive, respiratory and central nervous systems.  Xylene may have respiratory and neurological effects as well [1]. BETX components are present in many natural gas streams and are absorbed by the solvent in glycol dehydration and amine sweetening units.

In gas dehydration service, triethylene glycol (TEG) will absorb limited quantities of BETX from the gas. Based on literature data, predicted absorption levels for BETX components vary from 5-10% for benzene to 20-30% for ethylbenzene and xylene [2]. Absorption is favored at lower temperatures, increasing TEG concentration and circulation rate. The bulk of absorbed BETX is separated from the glycol in the regeneration unit and leaves the system in the regenerator overhead stream.

The emission of BETX components from glycol dehydration units is strictly regulated in most countries.  In the U.S., benzene emissions are limited to 1 ton/year (900 kg/year).  Mitigation of BETX emissions is an important component in the design of a dehydration system. Correctly estimating the quantity of absorbed BETX and understanding the factors that affect absorption levels is critical.

Moshfeghian and Hubbard [3] simulated a contactor column with two theoretical stages and a natural gas feed containing methane through n-decane (C1 through C10) and BETX compounds. The BETX concentrations in the feed gas were 400, 100, 50, and 50 ppmv for benzene, toluene, ethyl benzene, and o-xylene; respectively. The concentration of the lean TEG stream was 99.0 weight % TEG, and it was assumed the lean TEG temperature was the same as the feed gas. The feed gas was saturated with water at feed conditions. For each contactor pressure and temperature, the lean TEG circulation ratio was varied. To perform simulation, they used ProMax [4] with the Soave-Redlich-Kwong [5] equation of state (SRK EOS). Based on their simulation results, they presented two diagrams for quick estimation of BETX absorption in TEG dehydration process. Figures 1 and 2 [3] present the BETX absorption % (on weight basis) as a function of TEG circulation ratio and temperature for two pressures of 300 and 1000 psia (2069 and 6897 kPa).

This tip demonstrates application of these diagrams for quick estimation of BETX absorption in a case study. Specifically, TEG dehydration of a natural gas stream containing BETX compounds was considered and the estimated annual mass rate of BETX compounds absorbed by TEG solution. The graphical and hand calculation results were compared with the ProMax [4] simulation results and good agreement was observed.

Finally, an overview of the most commonly used designs for mitigation of BETX emissions will be provided.

 

.

Figure 1. Approximate BETX Absorption in TEG vs circulation ratio and contactor temperature at 300 psia (2069 kPa) [3]

 

 

Figure 2. Approximate BETX Absorption in TEG vs circulation ratio and contactor temperature at 1000 psia (6896 kPa) [3]

 

 

Example Problem

100 MMscfd (2.832×106 Sm3/day) of natural gas with a composition shown in Table 1 and saturated with water at feed conditions enters a TEG contactor. The inlet temperature is 110°F (43.3°C) and pressure is 1000 psia (6897 kPa). The outlet water dew point spec is 20°F (-6.7°C). The reboiler is operating at 360°F and there is one theoretical stage of random packing in the stripper.

Determine the quantity of BETX in lbm/year (kg/year) which would be vented to the atmosphere if no remedial actions were taken.

Assumptions: Approach temperature = 18°F (10°C) and lean TEG temperature = feed gas temperature

 

 

Solution

Based on the procedure outlined in Chapter 17 of Volume 2 of Gas Conditioning and Processing Book [6], the following parameters were determined:

 

 

 

Using Figure 3 for the lean TEG circulation ratio of 0.17 gpm TEG/MMscfd of gas (1.36 m3/h TEG/106 Sm3/day of gas), BETX absorption weight % for temperatures of 95 and 122°F (35 and 50°C) are presented in Table 2. Note that for 122°F (50°C), the lines were extrapolated by dotted lines for lower flow rates.  By linear interpolation, the estimated values of absorption weight % for temperature of 110°F (43.3°C) are presented in the last column of Table 2.

 

 

Figure 3. Estimating BETX Absorption (weight %) in TEG for example at psia (6896 kPa) [3]

 

 

 

For lean TEG concentration of 99.5 weight % and circulation rate of 17 gpm (3.861 m3/hr), the BETX absorption weight % from ProMax simulation results are shown in Table 3. As shown in Tables 2 and 3, there is a good agreement between the estimated BETX absorption weight % from Figure 3 and the ProMax simulation results. The ProMax calculated dry gas water dewpoint was also 20°F (-6.7°C) that matches the specified value.

 

 

The annual mass rate of each BETX compound entering the contactor in lbm/yr (kg/yr) can be determined from the following equations.

 

 

 

In the above equations q = 100 MMscfd (2.832×106 Sm3/day), MW (molecular weight) and ppmv (concentration) are given in Table 4. For each BETX compound, the absorbed mass rate of is determined by multiplying mass rate entering the contactor times the corresponding absorption weight % divided by 100.  The calculation results are presented in Table 4. The last two columns of Table 4 present the annual mass rate of each component absorbed in the TEG and subsequently regenerated and vented with the water.

 

 

The bulk of absorbed BETX will be vented with the water vapor at the top of the regenerator. The most common emission mitigation strategies are to [3]:

1. Condense the regenerator overhead vapor in a partial condenser and combust the remaining vapor.  The uncondensed vapors are typically routed to an incinerator or, if a direct-fired reboiler is used, routed to the reboiler fuel gas. The liquid hydrocarbons are collected and disposed of by blending into a crude oil or condensate stream.  The condensed water is typically routed to produced water disposal.

2. Route the regenerator overhead vapors to another process stream in the facility. This is typically a low-pressure stream such as flash vapors from the last stage of a crude or condensate stabilization system.

From an operational point of view, minimizing circulation ratio is the most effective way of decreasing the absorption of BETX components. This also minimizes reboiler duty and the size of the regeneration skid. Lower TEG circulation rates require more theoretical stages in the contactor to meet outlet water content specifications, but the additional cost of a taller contactor is often offset by savings in the regeneration package. Care should be taken that the glycol circulation rate is sufficient to ensure adequate liquid distribution over the packing. Packing vendors can provide minimum circulation guidelines.

 

 

Summary

Figures 1 and 2 present a simple tool for quick estimation of absorption of BETX compounds in TEG dehydration process. For the case study considered in this tip, the estimated absorption weight % for each of BETX compounds matched well with the ProMax simulation results.

To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), G5 (Advanced Applications in Gas Processing),  and PF49 (Troubleshooting Oil & Gas Processing Facilities) courses.

PetroSkills offers consulting expertise on this subject and many others. For more information about these services, visit our website at http://petroskills.com/consulting, or email us at consulting@PetroSkills.com.

One response to “Estimate BETX Absorption in TEG Dehydration Process”

  1. Alfred King says:

    1. What is the effect of high levels of CO2 in the feed gas which bypass an upstream sweeting unit for emergency situations.

    2. What is the effect of high levels of mercury on the BTEX absorpation and in the glycol closed loop system. Supposedly can reach 40 microgram/m3

    3. What is the effect on the TEG system and BTEX absorption due tomsalts in the feed gas carried out from inlet separation that supposedly separate the formation water

    4. How can we deal with BTEX in the vent stripper rather than destructive and Drizo techniques.

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Low Temperature Methane Gas Water Content

In the past tips of the month (October, November, December 2007, January 2011, February, September 2014. And April 2015), we studied in detail the water phase behaviors of sweet and sour natural gases and acid gas systems. We evaluated the accuracy of different methods for estimating the water content of sweet and sour natural gases as well as for acid gas systems. In addition, correlations to estimate vapor water content of lean sweet and sour gases were presented.

For normal gas conditioning processes, the water vapor content of natural gases in equilibrium with water is commonly estimated from McKetta and Wehe [1] based charts like Figure 6.1 of Campbell book [2] or Figure 20.4 of Gas Processors and Suppliers Association [3]. This tip presents a chart for estimating methane gas water content at low temperatures, e.g. cryogenic processes.

Hydrate formation is a kinetic (time dependent) process. During this transient “hydrate formation period” the liquid water present is termed “metastable liquid.” Metastable water is liquid water that, at equilibrium, will exist as a hydrate. At temperatures below the hydrate temperature of the gas, the “condensed” phase will be a solid (hydrate). The water content of a gas in equilibrium with a hydrate will be lower than equilibrium with a metastable liquid. The water content of gases in the hydrate region is a strong function of composition. Page 153 of Campbell book [2] provides more detail.

Figures 1A (SI, international system of units) and 1B (FPS field system of units) present a chart for methane gas water content for temperature range of 0 °C to -80 °C (32 °F to -110 °F) and pressures of 0.345 MPa, 3.45 MPa, and 6.9 MPa (50 psia, 500 psia, and 1000 psia). The chart is based on the experimental data reported in GPA-Midstream RR 187 [4] for pressures of 3.45 and 6.9 MPa (500 and 1000 psia) and the results generated using ProMax [5] freeze model (Solid-Vapor-Equilibrium) at low pressure of 0.345 MPa (50 psia).

 

Figure 1A (SI). Low temperature methane gas water content
(Dashed lines are GPA-Midstream RR 187 data [4], solid line predicted using ProMax [5])

 

 

Figure 1B (FPS). Low temperature methane gas water content
(Dashed lines are GPA-Midstream RR 187 data [4], solid line predicted using ProMax [5])

 

 

Based on the methane hydrate formation diagram shown in Figure 2, the above-mentioned low temperatures are generally below the hydrate formation temperature at the corresponding pressures. Therefore, the water contents reported in Figures 1A and 1B are in equilibrium with hydrates.

 

Figure 2. Methane hydrate formation temperature as a function of pressure

 

 

Summary

Figure 1 presents a simple chart to estimate methane gas water content in equilibrium with hydrates. The water content of gases in the hydrate region is a strong function of composition. Where experimental data is unavailable, utilization of an EOS-based correlation that has been tuned to empirical data can provide an estimate of water content in equilibrium with hydrates.

To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), G5 (Practical Computer Simulation Applications in Gas Processing), and PF49 (Troubleshooting Oil & Gas Processing Facilities), courses.

PetroSkills offers consulting expertise on this subject and many others. For more information about these services, visit our website at petroskills.com/consulting, or email us at consulting@PetroSkills.com.

 

By: Dr. Mahmood Moshfeghian

 


 

Reference:

  1. McKetta, J. J., and Wehe, A. H., “Use This Chart for Water Content of Natural Gases,” Petroleum Refiner (Hydrocarbon Processing), Vol. 37, No. 8, p. 153, August 1958.
  2. Campbell, J.M., Gas Conditioning and Processing, Volume 1: The Basic Principles, 9th Edition, 2nd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2014.
  3. GPSA Engineering Data Book, Section 20, Volume 2, 13th Edition, Gas Processors and Suppliers Association, Tulsa, Oklahoma, 2012.
  4. Song, K. Y., Yarrison, M., Kobayashi, R., and W. Chapman, “Low Temperature V-L-E Data for Water, CO2, and Light Hydrocarbon Systems,” Gas Processors Association Research Report RR 187 Tulsa, Oklahoma, 2005.
  5. ProMax 4.0, Bryan Research and Engineering, Inc., Bryan, Texas, 2017.

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Oil Dehydration Design Error Found in Most Designs Worldwide

This tip of the month (TOTM) discusses how to “avoid a costly design error found in most oil dehydration systems worldwide”.

  • Are you having reverse emulsion problems?
  • Are you having problems meeting your discharge oil in water specifications.

 

An area to look for immediately is the Process Flow Diagram (PFD).  Normally systems work as they were designed. In cases where I have visited sites experiencing problems…the problem is the design.

Why does this occur? The world we deal with is more complex than process simulation models. They are simply thermodynamic models to predict the phase separation of multicomponent systems and perform heat and energy balances. They are useful and very well proven tools. Most designs however; do not consider the various chemicals, presented in in Table 1, that are added to a system on a routine basis.

 

Table 1. Example of various problems and their typical treatment

 

Most of these chemicals are added to a system to improve a designs performance in oil water and water oil separation. Most of these chemicals are used in parts per million (ppm) fashion using proprietary formulations from Chemical Companies. Most assets will spend $5-$10 Million dollars per year on the use of these formulations. The ability to add the chemicals is included in most designs, but their effects are not normally included in the designs and it’s left to Field Operations and the Chemical Companies to come up with a solution.

So, What’s the design error? It’s mixing fluids that have come into contact with both oil treating and water treating chemicals and then recycling them back to the front of the process.

 

 

What are these streams?

  • Oil skims from WEMCOs (induced gas floatation manufacture by Western Equipment Manufacturing Company) or Flotation Cells are routinely re-routed back into the Oil Dehydration Train.
  • Hydrocyclone Oil Skims
  • Closed Drain Oil Skims

 

I would encourage you to check this out at your asset using the real fluids. I perform this demonstration during the Petroskills PF-3 Course. I take a 50% water cut sample of a 45 API Gulf of Mexico Crude Oil and produced water. I shake it up and the emulsion resolves in less than 30 seconds back into oil and water. I then add 5 -10 ml of Demulsifier and Water Clarifier being used on the Platform. Shake this up and voila… instant chocolate milk emulsion. This emulsion will not resolve in days. How is the emulsion resolved? We send it to a heater treater. In class I use a coffee cup of hot water. Let it heat soak for a couple of minutes and the sample resolves into three phases as shown in Figure 1: Oil / a scrambled eggs rag layer (complex emulsion- reverse and normal) / Water. The rag layer is normally about 25% of the total sample. This rag layer is very stable, and I have had samples stay stable for over one year.

 

Figure 1. Oil water complex emulsion

 

Never send these skims back into your main process. Consider sending them directly to the sales oil LACT. Most contracts allow you 3-5% BS&W. Use the contract to your advantage. You will not be able to break these emulsions with your existing process and they will continue to build. Some operators send them to the beach as hazardous waste for further processing. Other operators have a separate slop oil system for treating with higher temperatures, and then send streams to the sales LACT (Lease Automatic Custody Transfer), but never back into the process.

Every process needs a trash can for off spec product. Offshore it’s the closed drain tank / open drain tank. Never recycle this material back into the main oil dehydration system. Why? Let’s look at what fluids enter the closed and open drain tanks?

  • Engine Lube Oils – these contain polymers that will generate emulsions
  • Oxygenated water from rain. Oxygen will corrode your main production system.
  • Oily Deck Drains from skid pans

 

 

How to become an Expert?

Use sources such as SPE’s One Petro. Chances are that sometime, somewhere, someone has faced the same issue you are trying to solve.  In your company is it easy to publish a paper or are there controls and bureaucracy? In most instances, it’s not easy. What does it say to you about the person who has taken the time and written a paper. They want to communicate!  What do you think their reaction will be if you contact them about their paper to get further information? They usually will be very happy to discuss the paper and give you even more insights / tips / advice / recommendations for your situation. It’s a way of leveraging your time, and getting to solutions quickly.

I once made a trip to an FPSO Offshore Brazil that was having water treating problems. They made the same design mistakes discussed in this TOTM. I searched One Petro and found another FPSO Offshore Brazil that faced similar issues and how they were solved [1]. The paper is listed in the references of this TOTM. Here is the recommendation from SPE 90409 paper:

Oil Reject Skid. To provide a positive path for removing oily solids and chemically stabilized emulsions from the process, an oil reject skid was installed on 135 D to receive hydrocyclone reject liquids and skimming from the surge drum, as well as skimming from the sparger IGF vessels. In the oil reject vessel (Fig. 9), larger solids are removed from the cone bottom, while oily solids and chemically stabilized emulsion separates into an oil-continuous phase that is pumped off the platform without recycling back into the process. The volume of this emulsion is low, a few barrels per day, so the impact on the BS&W of crude leaving the platform is low, increasing it by approximately 0.05 %. The clarified water is returned to one of the degasser vessels, from which it reenters the water treatment process. The oil reject skid installation has been one of the major contributors to the ability of 135 D to reliably reduce the TOG of the produced water to < 20 mg/L.”

To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), G5 (Practical Computer Simulation Applications in Gas Processing), PF3 (Concept Selection and Specification of Production Facilities in Field Development Projects), and PF49(Troubleshooting Oil & Gas Processing Facilities), courses.

PetroSkills offers consulting expertise on this subject and many others. For more information about these services, visit our website at petroskills.com/consulting, or email us at consulting@PetroSkills.com.

 


Reference:

  1. Chifon Yand, Michel Gabrun, and Ted Frankiewicz, “Identification and Resolution of Water Treatment Performance Issues on the 135 D Platform,” SPE 90409, SPE Annual Technical Conference and Exhibition, Houston, Texas, 26-29 September, 2004.

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Get the most from your asset with this checklist

Finding Opportunities Checklist

This Tip of the Month (TOTM) discusses places to look for opportunities to increase production quickly. I’ll discuss the concepts, techniques, and later where to look for and find opportunities….to be able to answer the Key Question: “Are we getting the most from our Asset?”

  • Do you know how your asset makes money?  Is it simply oil price, and gas price?
  • Do the prices of oil and gas change daily?
  • Do you adjust your operating conditions based on the prices of oil and gas?
  • Does fuel gas have any value?

 

 

BENCHMARKING AGAINST PHYSICS™

What are your process Key Performance Indicators?

  • Energy Consumed per boe (barrel of oil equivalent) produced?
  • Theoretical Energy Required per boe produced?
  • Design Energy required per boe produced?

 

 

WHERE ARE THE GAPS?

How do you evaluate the performance of your:

  • Pumps
  • Compressors
  • Processes
    • Oil / Water Dehydration
    • Gas Dehydration
    • Gas Treating

 

 

Figure 1: Technical Limit Diagram or Choke Model

 

 

Looking at the Technical Limit Diagram (Choke Model) shown in Figure 1, where could there be opportunities?

  • How can we use the current equipment to increase production?
  • What’s the cost of one day of downtime?
  • What’s the monetary value of the annual flare volume?
  • How can we increase inflow from the wells to increase production?
    • Can we reduce the inlet pressures to the facilities to increase reservoir inflow?
    • Are we holding backpressure on the facilities, or are we floating the sales gas pipeline?
    • Are there any excessive pressure drops in the flowlines/equipment?
    • Scale
    • Remove choke internals

 

 

COMPRESSORS

  • Energy Efficiency:
  • What’s the position of the recycle valve? Are you wasting energy?
  • Reciprocating: Adjust Pockets / re-cylinder
  • Centrifugal: Consider re-wheeling
  • Can you reduce the suction temperature of the compressors to increase available HP
  • Infinite Recycling of LPG
    • Is this occurring in your compression system?
  • What’s causing downtime of your machines?
  • Could your upstream separators be carrying over liquids?

 

 

HEAT EXCHANGERS

  • Have the overall heat transfer coefficients become lower?
  • Have the pressure drops increased?

 

 

PUMPS

  •  What’s the position of your centrifugal pump discharge control valve?
    • Do you have too much HP, consider re-wheeling
  • Are you shearing and creating a lot of emulsion requiring higher chemical costs?

 

 

SPECIFICATIONS

  • Challenge sales specifications.
    • Are we over treating?
    • How’s our crude vapor pressure (RVP)?
  • Are we over stabilizing?
    • How’s our crude BS&W?
  • Are we over treating?

 

 

REDUCE OPEX

  • Chemical Costs: Demulsifier
    • Are we over dosing?
    • Are we adjusting the chemical injection rate base on changes in the inlet gross rate?
    • Have we done a “shootout” with other vendors, and allowed treating to failure point?
    • Are there recycle streams we can eliminate from the process and treat separately?
    • Do we have an OPEX model?
  • Have we identified our fixed and variable costs?
  • What’s the cost of fuel?  Is fuel free?
  • What’s the cost of flaring?
  • Are there operating cost differences between “A crew” and “B crew”?
    • Set up a competition between “A crew” and “B crew”.  Bbls produced, $ energy used per bbl, $ gas flared/bbl

 

 

DOWNTIME

  • How many times have we repaired this pump this year? Or over its life to date?
  • What’s the cost of downtime in $/hr?  How many $ in revenue has we lost this quarter?
  • How long does it take to repair?  Do we need a new seal design?  More spare parts?

 

 

CAPITAL PROJECTS

What could we add to increase production?

  • Is inlet compression profitable? (Lease / Purchase)

 

 

FLARE

Is the flare smokey?

  • This means its rich gas containing propane / butanes worth more than methane.  It’s a place to look for process changes to prevent the loss of these higher valued components.

 

 

WELLS

  • What’s the effect of increasing the tubing size?
  • What’s the effect of optimizing the gas lift?
  • Do all injection wells need the same maximum pressure for gas lift or water injection? Can we segregate the system?
  • Do the wells need stimulation?
  • Do the wells need water shutoff?
  • Manifolds – Are there leaking check valves allowing backflow from an HP well into an LP well?

 


 

Case 1: Different Molecules have different values in the Gas and Oil.  Are you maximizing the value of your products?

Offshore in the Gulf of Mexico, some contracts allow higher vapor pressure in the sales crude oil during the winter months.

 

KEY QUESTION: How can you take advantage of this opportunity?

Generally, selling molecules as oil has been more valuable than selling the molecules as gas.  So how can you send more of smaller LPG molecules into the sales crude oil to increase the number of crude oil barrels sold?

As you learned about vapor pressure and phase diagrams in G4 (Gas Conditioning and Processing course) you simply have to increase the pressure of the final separation, and decrease its temperature.  In winter months by lowering the oil dehydration train temperature, and increasing its pressure you increase the vapor pressure in the sales oil tank, and increase the number of barrels sold.  You can generate millions of dollars per year, by simply adjusting a temperature control setpoint and a pressure control setpoint. For Zero CAPEX.

 


 

Case 2: High Gas Prices

For a short period in 2000 a unique opportunity presented itself.  Gas Prices Spiked!

 

KEY QUESTION: How do we take advantage of this opportunity?

Simply send your molecules into the gas sales by dropping the oil dehydration treater pressure and raising its temperature.  Many Operators in the GOM took advantage of this opportunity. Some just sat and watched!

 


 

To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), G5 (Advanced Applications in Gas Processing), PF3 (Concept Selection and Specification of Production Facilities in Field Development Projects), and PF49 (Troubleshooting Oil & Gas Processing Facilities), courses.

PetroSkills offers consulting expertise on this subject and many others. For more information about these services, visit our website at petroskills.com/consulting, or email us at consulting@PetroSkills.com.

By: James F Langer, P.E.

One response to “Get the most from your asset with this checklist”

  1. James, thanks for the article.

    I support the message you got.

    For the wells part, don’t forget to open/remove well chokes as they are the biggest performance gap of all so far in the industry. The reservoir recovery will also go up, no worries.

    The “can of beans” study clearly explains this: https://www.youtube.com/watch?v=RPrBpKXpcZk&t=8s

    Thanks,
    Mikhail

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Benchmarking Against Physics – Increasing Production Today, Not Tomorrow

This Tip of the Month (TOTM) discusses how to increase production today, quickly, and for zero CAPEX. I’ll discuss the concepts, techniques, and where to look for and find opportunities, so we can answer the key question: “Are we getting the most from our asset?”

 

Many companies use benchmarking data shared anonymously with other participants in the survey. The problem is they will not have exactly the same assets. What’s being recommended here is that you benchmark you asset against physics and not others.

 

The good news is that no one can have better performance than the theoretical technical limit. If you determine your performance and compared it with a theoretical model / theoretical maximum and find a gap, then there’s an opportunity to improve.

 

 

 

 

BACKGROUND

You have been asked by the Asset Manager to find opportunities to increase production. In this case the well is flowing at 1200 psig (8.3 MPag) downstream of a wide open choke into a 1 mile (1.6 km) 3inch (76 mm) flowline discharging into a sales line operating at 1000 psig (6.9 MPag), and flowing gas at 10 MMSCFD (0.28×106 std m3/day).

 

 

KEY QUESTION – CASE 1

Is Today a good day? Are we getting the most from this asset?

Start by comparing the following:

  • What is the theoretical maximum given the current conditions – MODEL?
  • What is the current operational performance – ACTUAL PERFORMANCE?

Figure 1 presents a bar chart for BENCHMARKING AGAINST PHYSICS.

 

 

The most common answer is that the well is online. We are selling gas, and the choke is wide open. There is no further investigation or action performed. This is because the gap cannot be seen without performing the engineering.

 

The answer from reading this TOTM should be:

  • We have a process model that indicates that Benchmarking against PhysicsTM the flow line should be able to deliver 13 MMSCD (0.37×106 std m3/day).
  • You visit the field and verify the data. The gap is a 30% increase in production, see Figure 1.
  • What’s causing this? Another block valve in the system? System Scale?
  • How could you increase production quickly? Remove the entire choke body internals? Replace a small ported ball valve with a full ported valve? Recommend cleaning the line? Or simply opening a valve that is partially closed at the sales connection?

 

What’s the incremental value of 3 MMSCFD (0.085×106 std m3/day) with $4/ MMBTU ($3.8/GJ) sales gas per year? $4.35 Million USD/year. Was that worth a field trip? Absolutely!

 

 

Figure 2. Flowline Manifold – Are these valves lined up properly? Could they be leaking?

 

 

In Figure 2, what are the implications of leaking manifold valves? Could HP (high pressure) wells be sending gas backwards into lower pressure wells? Are stronger wells backing out lower pressure wells to sales?  Are a competitor’s higher pressure wells backing out our production?

 

 

KEY QUESTION – CASE 2

Is Anything wrong with our sale gas pipeline?

 

BACKGROUND

The pink points in Figure 3 are the gas sales in MMSCFD on the right axis, and the dark blue points are the pipeline differential pressure. The rate is decreasing over time and the pressure drop is decreasing over time.

 

 

Figure 3. SCADA DATA for pipeline performance gas rate in MMSCFD and differential pressure over time

 

 

 

The pipeline is a 14 inch (356 mm) 60 mile (96 km) from a major offshore asset. If the gas pipeline becomes plugged, then the oil production and gas production needs to be shut down; approximately 100,000 bopd (15,900 m3/day) and 200 MMSCFD (5.65×106 std m3/day) of gas. The operational data are presented in Table 1.

 

Table 1- Operational data performance

 

A UNISIM [1] Model shown in Figure 4 was completed to determine if the current performance matched the actual operational performance.

 

 

Figure 4 – UNISIM PIPELINE PERFORMANCE

 

 

The top pipeline simulation in Figure 4 matched the operating data performance at 200 MMSCFD (5.65×106 std m3/day), however; the pipeline at the lower rate of 169 MMSCFD (4.77×106 std m3/day) did not match the observed performance.  In order to match the performance a choke 4 inches (102 mm) in diameter by 100 ft (31.1 m) was added to the 60 mile (96 km) 14 inch (356 mm) pipeline model.

 

The platform had been experiencing asphaltenes in its separators. They were lucky to have performed the engineering analysis, and have a scheduled shutdown to clean a section of the pipeline at the base of the sales gas riser.  How are you monitoring your critical cash flow pipelines?

 

 

KEY QUESTION – CASE 3

How can we improve the performance of gas pipeline using SCADA data?

 

Here is another quick tip using SCADA Data and Excel.

 

From the GPSA Data Book [2] we know that gas flowrate (Q) is proportional to a constant times the square root of the static pressure (P1) times the differential pressure (ΔP).

 

 

So the trick is that if you construct a plot of C’ versus time… C’ is constant unless fouling has occurred. Fouling is difficult to find from just looking at data since it’s a chronic problem overtime, and not generally and acute one.

 

Figure 4 shows some SCADA Data using this technique. In order to scale the data, instead of C’ being plotted, the reciprocal of C’ squared was plotted to scale the plot.  This still should be a constant over time, and you can see the inflection point on the red line on this two years of data.

 

Figure 5. C’ Plot over time – Red line indicates a changing C’

 

 

Although this case reflected an asphaltene accumulation over time creating a localized plug. The unfavorable conditions would yield the same results as if it were a uniform scale along the entire pipeline length.

 

In upstream instances for subsurface models it’s difficult to see trends, because of the nature of the reservoir. We start off with high pressures, and flowrates, and these continually change over time.  The compositions oil / gas / water and their relative proportions change as well over time.  With all of these items changing overtime it’s sometimes difficult to detect small changes by simply looking at the data.  As shown in the previous examples if you Benchmark Against PhysicsTM it becomes very clear what’s happening.

 

The last item we will add to the bar graph tool (Figure 6) is a bar for the original design capacity.

 

Figure 6. ACTUAL / MODEL / ORIGINAL DESIGN CONDITIONS

 

 

 

KEY QUESTION – CASE 4

How can we increase the Gas Plant Throughput to full capacity?

 

BACKGROUND

You get a call from one of your Gas Plants in Europe, and it’s having a problem. It used to operate at full capacity, but something has changed and the operators can’t get more than 50% capacity through the plant. The plant processes several Billion Standard Cubic Feet of gas per day, and at current rates will lose over a $1 Billion USD this year if it is not corrected.

 

 

You build a model to compare actual performance with theoretical, and you build the bar charts for major pieces of equipment throughout the gas plant. The plant is a straight refrigeration gas plant that has stabilization of recovered liquids. It’s been there for 20 years operating without a problem. It processes about one third of the gas coming into the country.

 

 

After looking at the model and the graphs it’s clear to you what the problem could be. I flew from Houston to London. Visited the plant. Assessed the situation, and made some setpoint changes with the operators which resolved the issue, and then flew back to Houston the next day.

 

 

Why did I take the time to visit the plant? Below is some sage advice from Norman Lieberman [3].

 

“The tech service engineers employed by my clients, who work with me on these jobs, always ask what special techniques I employ to solve these problems. The procedure I use is the same one I used 10 years ago:

  1. Discuss the problem with the shift operators.
  2. Personally collect field data and carefully observe the operation of the unit.
  3. Develop a theory as to the cause of the malfunction.

 

The error my clients often make is that they develop a theory, usually with process computer simulations, as to the cause of the malfunction. The theory is then reviewed with management and other technical personnel at a large meeting. If no one objects to the theory, it is accepted as the solution to the problem. Typically, no one at the meeting has discussed the problem or the solution with the shift operators, nor has anyone personally observed the process deficiency in the field. Finally, the intended solution is not put to a plant test to see if it is consistent with the problem. This approach to solving refinery process problems by the major oil companies often results in wasting capital resources and engineering man-hours.”

 

 

Figure 7. Gas Plant Process Flow Schematic UNISIM

 

 

The plant schematic is shown in Figure 7, and it’s a straight refrigeration plant with a low temperature separator and liquids stabilizer. The plant must meet a gas dewpoint specification and a maximum liquids vapor pressure – RVP (Reid Vapor Pressure). Figure 8. Presents Benchmarking Against Physics bar charts for the key components in this gas plant. Note that the RVP of 6 psi (42 kPa) is well below specified value of 12 psi (83 kPa) and the plant is operating at 50% of inlet capacity.  The chiller is fully loaded.  The stabilizer overhead compressor is fully loaded. They are meeting all specs.

 

 

Figure 8

 

 

Figure 9 presents the gas plant schematic with bar charts for Benchmarking Against Physics.

 

Figure 9

 

 

So what’s your conclusion? The stabilizer overhead compressor is a bottleneck and current operations are over stabilizing the liquids sending too much vapor overhead (Producing a 6 psi (41 kPa) RVP versus a specification of 12 psi (83 kPa)).  If (they drop) the stabilizer bottoms temperature is reduced they will unload the stabilizer overhead compressor.  If they still are exceeding the stabilizer overhead compressor, then they can reduce the liquids recovery by not chilling the gas as much.  This will reduce the amount of liquid requiring stabilization.  The next limit is the gas dewpoint spec, and they will need to chill the inlet gas sufficiently enough to meet the specification.  Then they can begin increasing the inlet gas rate until they reach design maximum conditions.

 

 

Visiting the site, I found the plant operating as UNISIM predicted and the bar charts show. The operator’s mindset was that they needed to always chill the inlet gas to the maximum to make the company the most money. Then they backed themselves into a corner with the stabilizer operation. I went on rounds with them, and then suggest we try another approach and followed their management of change process to make setpoint changes. I explained the model results with them, and what we should expect to see as each change was slowly made. I have never shutdown a facility while making changes in 38 years because you always go slowly and are physically present to witness the cause and effects of changes. You never want to make a change and then hear the hiss of air leaving shutdown valve actuators and then have the lights go out. Going slowly builds the confidence of the operators in you and what you are doing.  Bottom line is two temperature setpoints were adjusted within limits, and the plant was returned to full capacity in 8 hours.  I then headed back to Houston for more adventures.

 

 

And as promised at the start of this TOTM…

Increase Production Today, Quickly, and for Zero (increase in) CAPEX!”

 

Before you finish reading this article, please think about how to apply the concepts discussed, and TAKE ACTION in your work today. This type of work if fun and exciting…it translates directly to your company’s bottom line. Go forth, be safe, and make it a great day!

 

In the next TOTM I will present other places to look for optimization opportunities within your assets.

 

To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), G5 (Advanced Applications in Gas Processing), PF3 (Concept Selection and Specification of Production Facilities in Field Development Projects), and PF49 (Troubleshooting Oil & Gas Processing Facilities) courses.

 

 

Reference:

  1. UNISIM Design R433, Honeywell International Inc., Calgary, Alberta, Canada, 2016
  2. Gas Processors Suppliers Association; “ENGINEERING DATA BOOK” 13th Edition; Tulsa, Oklahoma, USA, 2012.
  3. Norman P Lieberman, N P, “Troubleshooting Process Operations,” 4th Ed., PennWell, Tulsa, Oklahoma, 2009.

2 responses to “Benchmarking Against Physics – Increasing Production Today, Not Tomorrow”

  1. Ken Patel says:

    I enjoyed this – all what we are supposed to be about! I was taught by Crest Engineering – Original Outfit in London – We tackled North Sea Oil by the seat of our pants – Paper and Pen Engineering! No computers except the Mainframe Simulations I ran – 1977!

    I like ‘Out of the Box’ Process Engineering.

    Would like to keep in personal touch.

    Best regards.

  2. Ken Patel says:

    I have two questions – and need a broad brush idea.
    Theoretically if we have Two Subsea Lines.
    FULL CAPACITY – choked
    20″ Slugging Wet Gas Line running 50 miles.
    24″ Slugging Oil Line running the same 50 miles.

    How much more Oil and how much more Gas flow can we get in each line if:

    (1) Gas were to be running dry without slugging
    – take an average terrain with slope, going up, to Shore
    (2) Same goes for Oil.
    (3) If we could get Stabilized NGL at source.

    Thank you in advance.

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Pressure Effect on the Condensate Stabilizer Column Performance – Part 5

This tip is the follow up of January 2017 tip of the month (TOTM) [1] which investigated a non-refluxed condensate stabilizer column having a split design where a portion of the feed is pre-heated by heat exchange with the bottoms product (see also page 352 of Reference [2]). The remainder of the feed is fed to the top tray, similar to a standard “cold feed” stabilizer. That tip simulated the performance of a split-feed and non-refluxed condensate stabilizer column equipped with a stripping sweet gas stream to reduce the H2S content of a sour condensate. It reported that decreasing the top feed-split from 100% to 80 %, decreases the required reboiler duty and the bottom product temperature, but increases the stabilized condensate TVP (True Vapor Pressure) and H2S content.

 

This tip will perform computer simulation to study the effect of pressure on the performance of a non-refluxed stabilizer column with the same split-feed design. This tip will consider stabilizing a sour condensate for RVP (Reid Vapor Pressure) specification of 7 psi (48 kPa). The tip will report the impact of the column operating pressure on the reboiler duty, top and bottom temperatures, stabilized condensate rate, H2S content, and TVP. The tip will present a summary of the computer simulation results and the key diagrams for the same plant.

 

Case Study 

Figure 1 presents a simplified process flow diagram equipped with the preheater, stripping gas and water-draw tray for stabilization and sweetening of raw sour condensates.

 

 

Figure 1. A non-refluxed stabilizer column with feed-split, side water-draw and stripping gas

 

 

The 3-phase separator upstream of the preheater removes essentially all excess/free water. The tip utilizes a feed splitter to heat up a portion of feed in the preheater by the hot stabilized condensate. The stripping sweet gas stream lowers the H2S content of the stabilized condensate and achieves the desired condensate vapor pressure at the specified reboiler temperature. Three cases of column pressures were considered. For the comparison purposes and regardless of the specified 3-phase separator and column pressure, the separated gases from the 3-phase separator and the stabilizer column are compressed and cooled to specified pressure of 255 psia and temperature of 90 °F (1758 kPaa and 32 °C) in the compressor and cooler shown in the top-left side of this figure.

 

Table 1 presents the raw condensate, containing about 21 mole % H2S, and stripping gas compositions, rates and conditions. Table 2 presents the specified column variables. The condensate stabilizer column specifications for the three cases studied are shown in Table 2. For all three cases, the 3-phase separator pressure was 5 psi (35 kPa) higher than the column feed pressure.

 

 

Table 1. Feed and stripping gas compositions, rates and conditions

 

 

Table 2. Condensate stabilizer column specifications

 

 

Simulation Results

 

Based on the data in Tables 1 and 2, and the process flow diagram of Figure 1, the tip performed simulation using the Soave-Redlich-Kwong (SRK) equation of state [3] in ProMax [4] software. Similar to the previous tip the bottom-splitter ratio was adjusted to meet the specified RVP of the stabilized condensate. Table 3 presents the summary of results for the three specified cases. This table indicates that decreasing the specified 3-phase separator and column pressure decreases the bottom and overhead rates, bottom and overhead temperatures, but increases the compressor power and cooler duty. At lower pressures, more gases are separated in the 3-phase separator and feed rate to the column is reduced. Therefore, the bottom product (stabilized condensate) and overhead product rates decrease. In addition, decreasing column pressure increases the relative volatility (the K-value ratio of the key components) making separation of components easier and hence the required reboiler duty decreases.

The results presented in Table 3 are practically independent of the top-feed split which varied from 80 to 100 % with an increment of 2%. The values presented in Table 3 are the calculated arithmetic averages for the top-feed split range of 80 to 100%. The simulation results for those variables which were dependent on the top-feed split are shown graphically in the following figures.

 

 

Table 3. Summary of simulation results for three specified column pressures

 

 

Figure 2 presents the required reboiler duty as a function of the top feed-split and the specified column bottom pressure. This figure indicates that decreasing column bottom pressure decreases reboiler duty and the reboiler duty increases linearly with the top feed-split %. At a specified column pressure, increasing top-feed spilt decreases the hot-feed rate to column and hence increases the reboiler duty. However, the installation of the preheater reduced the required reboiler duty by about 19% for the case of top feed-split of 80% due to the heat recovery from the hot bottom product (stabilized condensate).

Figure 3 presents the variation of stabilized condensate temperature as a function of the top feed-split. Decreasing the top feed-split from 100 to 80%, decreases the bottom product temperature by about 2 °F (1.1 °C) for the three specified bottom pressures. Decreasing the specified column pressure, decreases the bottom temperature because component vapor pressure and temperature are directly related.

Figure 4 presents the variation of stabilized condensate TVP as a function of the top feed-split and column bottom pressure. Decreasing the top feed-split, increases the stabilized condensate TVP by about 2.6 psi (18 kPa) for the three specified bottom pressures.

 

 

Figure 2. Effect of top feed-split on reboiler duty for three column bottom pressures

 

 

Figure 3. Effect of top feed-split on bottom temperature for three column bottom pressures

 

 

Figure 4. Effect of top feed-split on stabilized condensate TVP for three column bottom pressures

 

 

Figure 5 presents the variation of H2S content of the stabilized condensate as a function of top feed-split and column overhead pressure. This figure indicates that decreasing the top feed-split from 100 to 80%, the stabilized condensate H2S content increases by about 44, 50, and 56 ppm for bottom pressure of 195, 225, and 255 psia (1345, 1552, and 1759 kPaa), respectively. This figure also indicates that for all three cases:

  • the H2S content is below the limit of 60 ppm,
  • the higher the top-feed %, the better the H2S stripping,
  • H2S stripping is better and easier at lower pressure.

It should be noted that since the raw feed condensate leaves the feed tank (three phase separator) saturated with water but with no free water, simulation resuts showed no liquid water being trapped in the column. Therefore, no water was removed by the water-draw tray in all cases studied in this tip. Still it is advised to keep the water draw-off capabilities on the top trays.

 

Figure 5. Effect of top feed-split on stabilized condensate H2S content for three column overhead pressures

 

 

 

Conclusions

 

This tip investigated the impact of column and 3-phase separator pressures on the performance of a split-feed and non-refluxed stabilization column by varying the top feed-split from 80 to 100% by an increment of 2 % for three pressure specifications. Based on the simulation results, this tip presents the following observations.

 

Decreasing the 3-phase separator and column pressures,

  1. decreases the required reboiler duty (Figure 2)
  2. decreases the bottom product temperature (Figure 3)
  3. decreases the stabilized condensate TVP (Figure 4)
  4. decreases the stabilized condensate H2S content (Figure 5)
  5. increases the required compressor power and cooler duty (Table 3)
  6. decreases the stabilized condensate rate (Table 3)

 

To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), G5 (Practical Computer Simulation Applications in Gas Processing), and PF4 (Oil Production and Processing Facilities), courses.

 

PetroSkills offers consulting expertise on this subject and many others. For more information about these services, visit our website at http://petroskills.com/consulting, or email us at consulting@PetroSkills.com.

 

By: Dr. Mahmood Moshfeghian

 

Reference:

  1. Moshfeghian, M., January 2017 TOTM, PetroSkills | John M. Campbell, 2017.
  2. Campbell, J.M., Gas Conditioning and Processing, Volume 2: The Equipment Modules, 9th Edition, 2nd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2014.
  3. Soave, G., Chem. Eng. Sci. 27, 1197-1203, 1972.
  4. ProMax 4.0, Bryan Research and Engineering, Inc., Bryan, Texas, 2016.

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Non-Refluxed Split-Feed Condensate Stabilizer Column – Part 4

This tip is the follow up of December 2016 tip of the month (TOTM) [2] which investigated the benefits of having a water-draw and its optimum location in a non-refluxed condensate stabilizer column. That tip simulated the performance of an operating condensate stabilizer column equipped with a side water-draw tray to remove liquid water and a stripping sweet gas stream to reduce the H2S content of a sour condensate. It determined and reported the possible locations of water-draw tray based water partial pressure along the column.

 


 

One option for stabilizer configuration is a split-feed design where a portion of the feed is pre-heated by heat exchange with the bottoms product. The remainder of the feed is fed to the top tray, similar to a standard “cold feed” stabilizer. Figure 1 presents an example of split-feed design. Split-feed provides heat recovery by pre-cooling the stabilized condensate upstream of the product cooler (not shown on this schematic), and reduces the required column reboiler duty (see also page 352 of Reference [1]).

 

Figure 1. A non-refluxed stabilizer column with feed-split, side water-draw and stripping gas

 

In order to lower the reboiler duty, a portion of the cold feed is heated by the hot stabilized condensate in a pre-heater. In addition, a free water knockout the upstream of the preheater is used to remove water and flashed gas from raw condensate. This tip will perform computer simulation to study the benefit of the upstream free water knockout drum and the pre-heater used in the split-feed design.

 

 

Specifically, this tip will determine if a liquid water draw tray is needed and how much reboiler duty is reduced by splitting the feed. This tip will consider stabilizing a sour condensate for Reid Vapor Pressure (RVP) specifications of 7, 7.5, and 8 psi (48, 52, 55 kPa). The tip will study the impact of the top feed-split % on the reboiler duty, bottom temperature, stabilized condensate H2S content and True Vapor Pressure (TVP). The tip will present a summary of the computer simulation results and the key diagrams for the same plant.

 

 

 

Case Study:

 

Table 1 presents the raw condensate, containing about 21 mole % H2S, and stripping gas compositions, rates and conditions. Figure 1 presents a simplified process flow diagram equipped with the preheater, stripping gas and water-draw tray for stabilization of this raw condensate. The 3-phase separator upstream of the preheater removes essentially all excess/free water. The tip utilizes   a feed splitter to heat up a portion of feed in the preheater by the hot stabilized condensate.

 

Table 1. Feed and stripping gas compositions, rates and conditions

 

 

The stripping sweet gas stream lowers the H2S content of the stabilized oil and achieves the desired condensate vapor pressure at the specified reboiler temperature. For each case, the boil-up ratio was adjusted to meet the RVP specification. Table 2 presents the specified column variables.

 

 

Based on the data in Tables 1 and 2, and the process flow diagram of Figure 2, the tip performed simulation using the Soave-Redlich-Kwong (SRK) equation of state [3] in ProMax [4] software.

 

 

Table 2. Condensate stabilizer column specifications

 

 

Simulation Results:

The tip adjusted the boil-up ratio in the reboiler to meet the specified RVP of the stabilized condensate. The resulting boil-up ratio as a function of top feed-split is presented in Figure 2. This figure indicates that to achieve lower RVP at a specified boil-up temperature, higher boil-up ratio is required.

 

 

Table 3 also presents the summary of results for three specified RVPs of 7, 7.5, and 8 psi (48, 52, 55 kPa). This table indicates that decreasing the specified RVP decreases the stabilized condensate rate. Lower RVP requires higher vaporization of lighter compounds. The simulation results (not shown in this table) indicated that overhead vapor temperature of about 100 °F (37.8 °C) is practically independent of specified RVP and top-feed split. In addition, the results presented in Table 3 are independent of top-feed split which varied from 80 to 100 % with an increment of 2%.

 

 

Figure 2. Effect of top feed-split on the reboiler boil-up ratio for three RVP specifications

 

 

 

Table 3. Summary of simulation results for three specified RVP

 

 

 

Similarly, Figure 3 presents the required reboiler duty as a function of top feed-split and the specified RVP. This figure indicates that decreasing RVP increases reboiler duty and the reboiler duty increases linearly with the top feed-split %. The installation of the preheater reduced the required reboiler duty by about 18% for the case of top feed-split of 80%.

 

 

Figure 4 presents the variation of stabilized condensate temperature as a function of top feed-split. Decreasing the top feed-split from 100 to 80%, the bottom product temperature decreases by about 1.8, 2.1, and 2.4 °F (1, 1.2, and 1.3 °C) for RVP of 7, 7.5, and 8 psi (48, 52, 55 kPa), respectively.

 

 

Figure 5 presents the variation of stabilized condensate true vapor pressure (TVP) as a function of the top feed-split. Decreasing the top feed-split, increases the stabilized condensate TVP by about 2.7, 3.0 and 3.3 psi (19, 21, and 23 kPa) for RVP of 7, 7.5, and 8 psi (48, 52, 55 kPa), respectively.

 

 

Figure 3. Effect of top feed-split on reboiler duty for three RVP Specifications

 

 

Figure 4. Effect of top feed-split on bottom temperature for three RVP specifications

 

 

Figure 5. Effect of top feed-split on stabilized condensate TVP for three RVP specifications

 

 

Figure 6. Effect of top feed-split on stabilized condensate H2S content for three RVP specifications

 

 

Figure 6 presents the variation of H2S content of the stabilized condensate as a function of top feed-split. This figure indicates that decreasing the top feed-split from 100 to 80%, the stabilized condensate H2S content increases by about 56, 80, and 104 ppm for RVP of 7, 7.5, and 8 psi (48, 52, 55 kPa), respectively.

 

 

This figure also indicates that for the specified RVP of 7.5 and 8 psi (52, 55 kPa), the H2S content exceeds the limit of 60 ppm at the top feed-split of 84 and 87 %, respectively. At the higher RVP, the bottom product temperature is cooler and there is not enough heat or stripping gas available to vaporize H2S.

 

 

It should be noted that since the raw feed condensate leaves the feed tank (three phase separator) saturated with water but with no free water, simulation resuts showed no liquid water being trapped in the column. Therefore, no water was removed by the water-draw tray in all cases studied in this tip. This is contrary to the previous tip in which free water was allowed into the column and the traped water was removed by the water-draw tray.

 

 

Conclusions:

 

This tip investigated the impact of the feed-split on the performance of a non-refluxed stabilization column by varying the top feed-split from 80 to 100% by an increment of 2 % for three RVP specifications. Based on the simulation results, this tip presents the following observations:

  1. Lower specified RVP, requires higher vaporization ratio at a given boil-up temperature. (Figure 2).
  2. Decreasing the top feed-split from 100% to 80 %, decreases the required reboiler duty by about 18% (Figure 3).
  3. Decreasing the top feed-split from 100% to 80%, decreases the bottom product temperature (Figure 4).
  4. Decreasing the top feed-split, increases the stabilized condensate TVP (Figure 5).
  5. Decreasing the top feed-split from 100% to 80%, increases the stabilized condensate H2S content (Figure 6).

 

To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), G5 (Practical Computer Simulation Applications in Gas Processing), and PF4 (Oil Production and Processing Facilities), courses.

 

By: Dr. Mahmood Moshfeghian

 

 

Reference:

  1. Campbell, J.M., Gas Conditioning and Processing, Volume 2: The Equipment Modules, 9th Edition, 2nd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2014.
  2. Moshfeghian, M., December 2016 TOTM, PetroSkills | John M. Campbell, 2016.
  1. Soave, G., Chem. Eng. Sci. 27, 1197-1203, 1972.
  1. ProMax 4.0, Bryan Research and Engineering, Inc., Bryan, Texas, 2016.

0 responses to “Non-Refluxed Split-Feed Condensate Stabilizer Column – Part 4”

  1. […] tip is the follow up of January 2017 tip of the month (TOTM) [1] which investigated a non-refluxed condensate stabilizer column having a split design where a […]

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Water-Draw in a Non-Refluxed Condensate Stabilizer Column – Part 3

This tip is the follow up of the previous tips (April and May 2016) [1-2], which investigated the benefits of having a water-draw and its optimum location in a condensate stabilizer column [see chapter 16 of reference 3]. It will simulate the performance of an operating condensate stabilizer column equipped with side water-draw tray to remove liquid water. Recall from the April 2016 TOTM, water can become trapped internally within a condensate stabilizer. The column operating conditions result in water condensing within the column and becoming trapped.  The overhead temperature is too cool and the bottoms temperature is too hot to allow the water to leave the column in either of the product streams.  As a result, liquid water build-up will occur within the column reducing capacity, and depending upon composition, increasing corrosion.  Eventually the water build-up will cause the column to flood, and the major disruption in the tower operations allows the water to be removed.  After this event, the column will operate normally until sufficient time that enough water has accumulated to cause the column to flood once again.  A properly located water-draw off tray will allow for proper column operation and eliminate the operating problems associated with water build up.

 

In this case study, a stripping sweet gas stream is also utilized to achieve low H2S content and stabilized condensate at a specified reboiler temperature. The tip will perform three-phase (vapor, liquid hydrocarbon, and aqueous phases) calculations on the trays with excessive/free water rates. Specifically, it will study possible locations of a water-draw tray based on the profiles for water partial pressure in the vapor phase and water rates in the light and heavy liquid phases. The tip will present a summary of the computer simulation results and the key diagrams for the same plant.

 

 

Case Study:

 

Table 1 presents the raw condensate and stripping gas compositions, rates and conditions. Figure 1 presents a simplified process flow diagram equipped with stripping gas and water-draw tray for stabilization of a raw condensate. The 3-phase separator upstream of the stabilizer column removes essentially all excess/free water. The tip utilizes the front mixer to recombine the light and heavy liquid (excess free water) streams feeding to the column for the simulation purpose only. Table 2 presents the stabilizer column specifications. While location of the water-draw tray is 7 in this figure, the tip also considered tray locations at 5, 8, 9, 17, and 18.

 

The stripping sweet gas stream lowers the H2S content of the stabilized oil and achieves the desired condensate vapor pressure at the specified reboiler temperature. Table 2 presents the specified column variables. Stream 2 is the overhead vapor and stream 3 is the stabilized condensate.

 

Based on the data in Tables 1 and 2, and the process flow diagram of Figure 1, the tip performed simulation using the Soave-Redlich-Kwong (SRK) equation of state [4] in ProMax [5] software.

 

Table 1. Feed and stripping gas compositions, rates and conditions

Table 1. Feed and stripping gas compositions, rates and conditions

 

 

Figure 1. A simplified non-refluxed stabilizer column with side water-draw and stripping gas

Figure 1. A simplified non-refluxed stabilizer column with side water-draw and stripping gas

 

 

Table 2. Condensate stabilizer column specifications

Table 2. Condensate stabilizer column specifications

 

 

 

Simulation Results:

 

The tip varied the boil-up ratio in the reboiler to match the stream 3 plant temperature of 482 °F (250 °C). The resulting boil up-ratio of 76.5 % is presented in Table 2 and Table 3 presents the comparison between simulation results of this work with plant data. Overall, a reasonable agreement is observed.

 

Table 3. Comparison of simulation results with plant data

Table 3. Comparison of simulation results with plant data

 

 

Figure 2 presents the water partial pressure profile for cases of no water-draw and water-draw at either tray of 7, 8 or 9. In this figure, a large spike of water partial pressure is observed on tray 18 for the cases of no water-draw tray. For the case of no water-draw, a bump in water partial pressure is also observed at tray 10 due to the presence of water in the stripping gas that enter the column on tray 10.

 

Similarly Figures 3 and 4 present the water rate profiles in the light (liquid hydrocarbon) and heavy liquid (aqueous) phases, respectively. These figures show the water rate profiles for no water-draw and water-draw tray at either tray 7, 8, or 9. In these figures, a large spike of water rate profile is observed for the case of no water-draw tray.

 

Figure 4 indicates that below any of the water-draw trays the light (hydrocarbon) liquid phase is under-saturated with water and there is no free water (heavy liquid phase). For the case of no water-draw, the light liquid phase remains saturated with water along trays 1 through 18.

 

Figure 5 indicates that the presence of a water-draw tray has no significant impact on the column temperature profile.

 

In all cases the location of a water-draw tray has no impact on Reid Vapor Pressure (RVP) of neither condensate nor the reboiler duty. The calculated RVP for all cases was 8.2 psi (56.6 kPa) and the calculated reboiler duty for all cases was 18.378 MMBtu/hr (5.376 MW). Table 4 presents the impact of water-draw location on the rate of water removed. This table indicates that water rate at either trays of 5 through 9 practically is the same amount.

 

Figure 2. Water partial pressure profile in the stabilizer column for several cases

Figure 2. Water partial pressure profile in the stabilizer column for several cases

 

 

Figure 3. Water rate in light liquid phase in the stabilizer column for several cases

Figure 3. Water rate in light liquid phase in the stabilizer column for several cases

 

 

Figure 4. Water rate in heavy liquid phase in the stabilizer column for several cases

Figure 4. Water rate in heavy liquid phase in the stabilizer column for several cases

 

 

Figure 5. Temperature profiles in the stabilizer column for several cases

Figure 5. Temperature profiles in the stabilizer column for several cases

 

 

Table 5 presents the percent recovery (ratio of a component rate in the condensate to its rate in the feed stream) of selected components in the stabilized condensate. Practically, all ethane and lighter components (N2, C1, CO2, and H2S) leave in the column overhead. Table 5 indicates that the presence of a water-draw tray has some effect on propane and little effect on butane but no effect on other component recoveries.

 

 

Table 4. Impact of water-draw tray location on removal water rate

Table 4. Impact of water-draw tray location on removal water rate

 

 

Table 5. Recovery of selected components in the stabilized condensate

Table 5. Recovery of selected components in the stabilized condensate

 

 

Conclusions:

 

The tip investigated the location of side water-draw and its impact on the performance of the stabilization column. Based on the results obtained, this tip presents the following observations.

  1. The water-draw rate at either trays 5 through 9 is the same.
  2. The water draw tray location had no impact on the RVP of stabilized condensate.
  3. The water draw tray location had no impact on the reboiler duty.
  4. The water draw try improved the propane recovery.

 

To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), G5 (Advanced Applications in Gas Processing), and PF4 (Oil Production and Processing Facilities), courses.

 

PetroSkills offers consulting expertise on this subject and many others. For more information about these services, visit our website at http://petroskills.com/consulting, or email us at consulting@PetroSkills.com.

 

By: Dr. Mahmood Moshfeghian

Reference:

  1. Moshfeghian, M., April 2016 tip of the month, PetroSkills | John M. Campbell, 2016.
  2. Moshfeghian, M., May 2016 tip of the month, PetroSkills | John M. Campbell, 2016.
  3. Campbell, J.M., Gas Conditioning and Processing, Volume 2: The Equipment Modules, 9th Edition, 2nd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2014.
  1. Soave, G., Chem. Eng. Sci. 27, 1197-1203, 1972.

ProMax 4.0, Bryan Research and Engineering, Inc., Bryan, Texas, 2016.

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Charts and Correlation for Estimating Methanol Removal in TEG Gas Dehydration Process

The TEG (triethylene glycol) gas dehydration process removes a considerable amount of methanol from a wet gas stream. if the methanol content of the wet gas is high, the dry gas may still retain high methanol content and can cause operational troubles in the downstream processes.

 

Continuing the October 2010 tip of the month (TOTM), in this TOTM we will consider the presence of methanol in the produced oil/water/gas stream and determine the quantitative traces of methanol ending up in the TEG dehydrated gas. To achieve this, we simulated by computer an offshore production facility consisting of oil/water/gas multistage-separation, compression and TEG dehydration processes and determined the methanol concentration in the dried gas. We also studied the effect of wet gas temperature and pressure, the number of theoretical trays in the TEG contactor, the water content spec of dry gas, and lean TEG solution circulation rate on the dried gas methanol content. For this purpose, methanol content in the production stream was assumed to vary from zero to 350 PPM (V).

 

Based on the computer simulation results, the tip develops simple charts and correlations to estimate the methanol removal efficiency in a TEG contactor column under various operating conditions. These charts and correlations are accurate enough for facilities calculations.

 

 

Case Study:

The same case study presented in the October 2010 TOTM is used to study the methanol removal in a TEG contactor column. A simplified process flow diagram (PFD) for the offshore production facility is shown in Figure 1 [1]. The production stream (oil, water, gas, and methanol) was passed through the high pressure separator where free water and gas were separated and the oil was passed through the intermediate and low pressure separators for subsequent gas separation from oil. The separator’s off gas streams were recompressed and cooled to 4830 kPa and 35 °C (700 psia and 95 °F) before entering the TEG contactor for dehydration. The dried gas was compressed further (not shown in the PFD) to 23200 kPa (3365 psia) for reinjection or export purposes. To meet a water content spec of 32 mg/Sm3 (2 lbm/MMscf) or lower, a lean TEG concentration of 99.95 weight percent was used in all of the simulation runs. This tip uses ProMax [2] simulation software with the “PR EOS” property package to perform all of the simulations.

 

To study the impact of methanol (MeOH) concentration and determine the traces in the TEG dehydrated gas, the MeOH content of the production stream feed to the high pressure separator was assumed to vary from 0 to 350 PPM (V). This variation of MeOH content was chosen due to the uncertainty of its concentration in the production stream. The wet compressed gas temperature is an important parameter in the operation of a TEG unit and affects the water content of dried gas and the required lean TEG solution circulation rate and/or the number of required theoretical trays in the contactor. Depending on the design and/or operational problem like scaling on the cooling side of the gas cooler, the wet gas temperature may be higher than 35 °C (95 °F). Therefore, the wet gas temperature was assumed to vary from 35 to 50 °C with 5 °C increments (95 to 122 °F and 9 °F increment). Depending on the requirement, 2 or 3 theroetical trays (xtheoreticalx) were (xwasx) used in the contactor unit. For each case the lean TEG solution rate was varied to meet the desired water content specification for each case.

 

To study the impact of pressure, in addition to pressure of 4830 kPa (700 psia), a wet gas pressure of 7000 kPa (1015 psia) was also used.

 

For each simulation run, the methanol removal efficiency (MRE) was calculated by the following equation.

eq1

 

 

Figure 1. Simple process flow diagram used in this case study [1]

Figure 1. Simple process flow diagram used in this case study [1]

 

Results and Discussions:

Table 1 shows sample calculation results for two theoretical trays and wet gas pressure of 4830 kPa (700 psia) and four temperatures. For each temperature the wet gas methanol content was varied in the range of 0 to 112 PPM on molar basis. In each case, the lean TEG solution circulation rate was adjusted to meet the dry gas water content of 16 mg/Sm3 (1 lbm/MMSCF). Analysis of Table 1 indicates that for each wet gas temperature, the circulation ratio and methanol removal efficiency were independent of the wet gas methanol content. Therefore, for each wet gas temperature, the average circulation ratio and methanol removal efficiency were calculated and presented in Table 2. Similar to the results of tables 1 and 2, seven more tables were generated for 2 and 3 theoretical trays, dry gas water contents of 16 and 32 mg/Sm3 (1 and 2 lbm/MMSCF), pressures of 4830 and 7000 kPa (700 and 1015 psia).

 

Table 1. Sample results for two theoretical trays and wet gas pressure of 4830 kPa (700 psia)

Table 1. Sample results for two theoretical trays and wet gas pressure of 4830 kPa (700 psia)

 

Table 2. Average results for two theoretical trays and wet gas pressure of 4830 kPa (700 psia)

Table 2. Average results for two theoretical trays and wet gas pressure of 4830 kPa (700 psia)

 

The summary for all of the simulation results for methanol removal efficiency as a function of the lean TEG solution circulation ratio for 2 and 3 theoretical number of trays at two pressures, and two dry gas water content specifications are presented in Figures 2 and 3, respectively.

 

Analysis of Figures 2 and 3 indicates that for each pressure the results for dry gas water content of 16 and 32 mg/Sm3 (1 and 2 lbm/MMSCF) specifications follow the same trend and can be represented by a single curve. Therefore, the methanol removal efficiency as a function of the lean TEG solution circulation ratio can be expressed by a single correlation for each pressure and number of theoretical trays independent of the wet gas methanol content and temperature.

 

Figure 2. Average methanol removal efficiency vs circulation ratio for 2 theoretical trays

Figure 2. Average methanol removal efficiency vs circulation ratio for 2 theoretical trays

 

Figure 3. Average methanol removal efficiency vs circulation ratio for 3 theoretical trays

Figure 3. Average methanol removal efficiency vs circulation ratio for 3 theoretical trays

 

A non-linear regression program was used to determine the parameters of the following correlation for the methanol removal efficiency (MRE) as a function of the lean TEG solution circulation ratio.

eq3

 

Where:

MRE         = Methanol Removal Efficiency on the mole basis, %

CR             = Circulation ratio, liter TEG/kg water (gallon TEG/lbm water)

 

Table 3 presents the regressed parameters of “a” and “b” of Equation 1 for two and three theoretical trays and the wet gas pressures of 4830 and 7000 kPa (700 and 1015 psia). The last two rows in Table 3 present the Average Absolute Percent Error (AAPE) and the Maximum Absolute Percent Error (MAPE) for different units of lean TEG solution circulation rate.

 

Table 3. Parameters of Equation 1 for methanol removal efficiency

Table 3. Parameters of Equation 1 for methanol removal efficiency

 

The MRE predictions by Equation 1 were added to Figures 2 and 3 and are presented in Figures 4 and 5. In these two figures the solid lines present the MRE prediction by Equation 1 and symbols represent simulation results. The filled symbols represent the dry gas water content of 16 mg/Sm3 (1 lbm/MMSCF) and the no fill symbols represent the dry gas water content of 32 mg/Sm3 (32 lbm/MMSCF). The analysis of Figures 4 and 5 and the calculated low values of AAPE and MAPE in Table 3 indicate that accuracy of the proposed correlations, compared to the simulation results, is good for estimation of methanol removal efficiency (MRE).

 

 

Conclusions:

Based on the results obtained for the considered case study, this TOTM presents the following conclusions:

  1. The methanol removal efficiency is independent of the wet gas methanol content (Table 1).
  2. As the wet gas temperature increases, the lean TEG solution circulation ratio increases; therefore, methanol removal efficiency increases (Table 2). The wet gas water content is a strong function of temperature. As temperature increases, the wet gas water content increases; therefore, for a fixed number of trays the required lean TEG solution rate increases.
  3. As the wet gas pressure increases, the absorption of methanol increases; therefore, methanol removal efficiency increases (Figures 4 and 5). The wet gas water content is a function of pressure. As pressure increases, the wet gas water content decreases; therefore, for a fixed number of trays the required lean TEG solution rate decreases but this decrease in rate is offset by higher solubility of methanol at higher pressure.
  4. For the same TEG contactor number of trays and pressure, the methanol removal efficiency as a function of circulation ratio for different dry gas water specifications follows the same trend (Figures 4 and 5).
  5. The tip presents two simple charts (Figures 4-5) and a correlation (Equation 1) along with its parameters (Table 3) for estimating the average methanol removal efficiencies for 2 and 3 theoretical trays and pressures of 4830 and 7000 kPa (700 and 1015 psia), respectively.
  6. Compared to the rigorous computer simulation results, the accuracy of the proposed correlations (Equation 1, and Table 3) to estimate the average methanol removal efficiency provides good agreement to simulation results. This correlation, and Figures 4 or 5 can be used to estimate MeOH removal performance in TEG facilities operating at similar conditions.
  7.  The proposed correlations (Equation 1) and charts (Figures 4-5) are easy to use, and provides a simple approach to estimate MeOH removal efficiency in TEG units without access to a process simulator.

 

Figure 4. Average methanol removal efficiency vs circulation ratio for 2 theoretical trays

Figure 4. Average methanol removal efficiency vs circulation ratio for 2 theoretical trays

 

 

Figure 5. Average methanol removal efficiency vs circulation ratio for 3 theoretical trays

Figure 5. Average methanol removal efficiency vs circulation ratio for 3 theoretical trays

 

To learn more about similar cases and how to minimize operational troubles, we suggest attending our G4 (Gas Conditioning and Processing), G5 (Advanced Applications in Gas Processing), and PF4 (Oil Production and Processing Facilities) courses.

 

PetroSkills | Campbell offers consulting expertise on this subject and many others. For more information about these services, visit our website at http://petroskills.com/consulting, or email us at consulting@PetroSkills.com.

 

References:

  1. Moshfeghian, M., October 2010 tip of the month, PetroSkills | John M. Campbell, 2010.
  2. ProMax 4.0, Bryan Research and Engineering, Inc., Bryan, Texas, 2016.

One response to “Charts and Correlation for Estimating Methanol Removal in TEG Gas Dehydration Process”

  1. Ryan Mulvania says:

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Estimating Methanol Removal in the Gas Sweeting Process

Charts and Correlations for Estimating Methanol Removal in the Gas Sweetening Process

 

The gas-sweetening process by amines like methyldiethanolamine (MDEA) removes a considerable amount of methanol from a sour gas stream. Moreover, if the methanol content of the sour gas is high, the sweet gas may still retain high methanol content and can cause operational troubles in the downstream processes. Provisions of purging reflux (Water Draw) of the regenerator column and its replacement with “Fresh Water” can improve methanol recovery [1, 2].

 

The July 2016 tip of the month (TOTM) considered the presence of methanol in the sour gas stream and determined the quantitative traces of methanol ending up in the sweet gas, flash gas and acid gas streams [2]. It simulated a simplified MDEA gas-sweetening unit by computer and studied the effect of sour gas methanol content, and the rate of replacing condensed reflux with fresh water on the sweet gas methanol content. For the sour gas temperature of 43.3 and 32.2 °C (110 and 90 °F) the tip studied three inlet gas methanol contents of 50, 250, and 500 PPM on mole basis. In each case the tip varied the rate of fresh water replacement from 0 to 100 % by an increment of 20%.

 

The methanol removal efficiency (MRE) on the volume basis is defined by:

eq

 

Table 1 presents the summary of calculated methanol removal efficiency (MRE) based on the simulation results of the July 2016 TOTM [2].

 

Table 1. The effect of purging and sour gas temperature on methanol removal efficiency [2]

Table 1. The effect of purging and sour gas temperature on methanol removal efficiency [2]

 

In continuation of the July 2016 TOTM, this tip will consider the presence of methanol in the sour gas stream and determine the quantitative traces of methanol ending up in the sweet gas, flash gas and acid gas streams. This tip simulates a simplified MDEA gas-sweetening unit by computer simulation [3, 4]. This tip also studies the effect of sour gas methanol content, temperature and the rate of replacing condensed reflux with fresh water on the sweet gas methanol content.

 

For the sour gas temperatures of 43.3, 32.2 and 21.1 °C (110, 90, and 70 °F) the tip studies three inlet sour gas methanol contents of 50, 250, 500 PPM on mole basis. In each case the tip varies the rate of fresh water replacement from 0 to 100 % by an increment of 20%. Similar to the September 2016 TOTM [5] and based on the computer simulation results, the tip develops simple charts and correlations to estimate the methanol removal efficiency under various operating conditions. These charts and correlations are accurate enough for facilities calculations.

 

 

Case Study:

For the purpose of illustration, this tip considers sweetening of a sour gas stream saturated with water using the basic and modified MDEA processes as described in the July 2016 TOTM [2]. In addition to the two sour gas temperatures reported in the July TOTM, this tip also considers a sour gas temperature of 21.1 °C (70 °F). Table 2 presents its composition on the dry basis, gas standard volume rate, pressure, and temperatures. This tip uses ProMax [6] simulation software with the “Amine Sweetening – PR” property package to perform all of the simulations.

 

 

Results and Discussions:

Figures 1 through 3 present the calculated MRE as a function of the reflux rate replacement (RRR) with fresh water for the sour gas temperatures of 43.3, 32.2, and 21.1 °C (110, 90, and 70 °F), respectively. Each figure presents MRE vs replacement rate for the three sour gas methanol contents (50, 250, and 500 PPMV).

 

Table 2. Feed composition on the dry basis, volumetric flow rate and conditions [2]

Table 2. Feed composition on the dry basis, volumetric flow rate and conditions [2]

Figure 1. Methanol removal efficiency vs reflux replacement for sour gas temperature of 43.3 °C (110 °F)

Figure 1. Methanol removal efficiency vs reflux replacement for sour gas temperature of 43.3 °C (110 °F)

 

 

Figure 2. Methanol removal efficiency vs reflux replacement for sour gas temperature of 32.2 °C (90 °F)

Figure 2. Methanol removal efficiency vs reflux replacement for sour gas temperature of 32.2 °C (90 °F)

 

 

Figure 3. Methanol removal efficiency vs reflux replacement for sour gas temperature of 21.1 °C (70 °F)

Figure 3. Methanol removal efficiency vs reflux replacement for sour gas temperature of 21.1 °C (70 °F)

 

 

Since the three curves for different sour gas methanol contents on each figure are close, the effect of the sour gas methanol content on MRE can be neglected. For each sour gas temperature, the calculated arithmetic average of MRE of the three sour gas methanol content are provided in Figure 4. This figure indicates that as the sour gas temperatures decreases the impact of the reflux rate replacement with fresh water diminishes.

 

Figure 4. Average methanol removal efficiency vs reflux replacement

Figure 4. Average methanol removal efficiency vs reflux replacement

 

 

A non-linear regression program was used to determine the parameters of the following correlation for the methanol removal efficiency as a function of the reflux rate replacement % (RRR).

 

eq1

 

Where:

MRE = Methanol Removal Efficiency on the mole basis, %

RRR = Reflux Rate Replacement %

 

Table 3 presents the regressed parameters of A, B and C of Equation 1 for the three considered sour gas temperatures. The last two rows in Table 3 present the Average Absolute Percent Error (AAPE) and the Maximum Absolute Percent Error (MAPE), respectively.

 

A generalized form of this correlation to cover the temperature effect can be expressed as:

 

eq2

 

Where:

MRE   = Methanol removal efficiency on the weight basis

RRR    = Reflux Rate Replacement %

T          = Temperature, ºC (ºF)

 

Table 4 presents the regressed parameters of A1, A2, B1, B2, C1 and C2 of Equation 2 for temperatures in °C and °F. Similarly, the last two rows in Table 4 present the AAPE and the MAPE, respectively.

 

Table 3. Parameters of Equation 1 for methanol removal efficiency

Table 3. Parameters of Equation 1 for methanol removal efficiency

 

 

Table 4. Parameters of Equation 1 for methanol removal efficiency

Table 4. Parameters of Equation 1 for methanol removal efficiency

 

 

The MRE predictions by Equation 2 were added to Figure 4 and is presented as Figure 5. In this figure the solid lines present the MRE prediction by Equation 2 and dashed lines with the filled symbols represents simulation results. The analysis of Figures 5 and the calculated values of AAPE and MAPE in Table 4 indicate that accuracy of the proposed correlations, compared to the simulation results, is very good for estimation of methanol removal efficiency (MRE).

 

fig5

Figure 5. Comparison of model prediction of average methanol removal efficiency vs reflux replacement

 

 

Conclusions:

Based on the results obtained for the considered case study, this TOTM presents the following conclusions:

  1. The impact of the sour gas methanol content on the methanol removal efficiency is small (Figures 1-3), overall only a minor impact, less than 0.5 % point.
  2. As the sour gas temperature decreases, the methanol removal efficiency increases (Figures 1-5), overall only minor impact, less than 3 % points.
  3. Methanol removal efficiency with MDEA sweetening can remove only 89-97% of the methanol in the sour gas feed.  This may still leave more methanol than the gas spec allows.  A separate water wash step may be required.  The fresh water used for the water wash could be recycled as MDEA reflux purge make-up.
  4. The tip presents three simple charts (Figures 1-3) and two correlations (Equations 1 and 2) along with their parameters (Tables 3 and 4) for estimating the average methanol removal efficiencies for the sour gas temperatures of 43.3, 32.2, and 21.1 °C (110, 90, and 70 °F), respectively.
  5. Compared to the rigorous computer simulation, the accuracy of the proposed correlations (Equations 1 and 2) to estimate the average methanol removal efficiency is very good (Tables 3 and 4 and Figure 5) and can be used for facilities calculations.
  6. The proposed correlations (Equations 1 and 2) and charts (Figures 4-5) are easy to use.

To learn more about similar cases and how to minimize operational troubles, we suggest attending our G6 (Gas Treating and Sulfur Recovery), G4 (Gas Conditioning and Processing), G5 (Advanced Applications in Gas Processing), and PF4 (Oil Production and Processing Facilities) courses.

PetroSkills | Campbell offers consulting expertise on this subject and many others. For more information about these services, visit our website at http://petroskills.com/consulting, or email us at consulting@PetroSkills.com.

 

By: Dr. Mahmood Moshfeghian

 

 

 

References:

  1. O’Brien, D., Mejorada, J., Addington, L., “Adjusting Gas Treatment Strategies to Resolve Methanol Issues,” Proceedings of Lawrence Reid Gas Conditioning Conference, Norman, Oklahoma, 2016.
  2. Moshfeghian, M., July 2016 tip of the month, PetroSkills | John M. Campbell, 2016.
  3. Maddox, R.N., and Morgan, D.J., Gas Conditioning and Processing, Volume 4: Gas treating and sulfur Recovery, Campbell Petroleum Series, Norman, Oklahoma, 1998.
  4. Campbell, J.M., Gas Conditioning and Processing, Volume 2: The Equipment Modules, 9th Edition, 1st Printing, Editors Hubbard, R. and Snow –McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2014.
  5. Moshfeghian, M., September 2016 tip of the month, PetroSkills | John M. Campbell, 2016.
  6. ProMax 4.0, Bryan Research and Engineering, Inc., Bryan, Texas, 2016.

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