Campbell Tip of the Month

Category: Pipeline

  • Impact of Gas-Oil Ratio (GOR) on Crude Oil Pressure Drop in Gathering Systems

    The use of multiphase flow systems is common practice in the oil and gas industry. Multiphase flow is often encountered in the well tubing, flow lines and gathering systems. For transport of oil and gas (and water) to downstream processing facilities the preference is normally a single pipeline in which both phases are transported simultaneously for economic reasons. Even in gas pipelines where the gas enters the line as a single phase fluid, condensation of liquids can occur due to pressure and temperature changes along the line.

    Modeling and simulation of a multiphase systems, even under steady-state conditions, is complex. There are a few tools designed specifically for modeling and analysis of complex multiphase systems such as PipePhase, PipeSim, OLGA, etc. [1].

    In the June 2008 Tip of the Month (TOTM), we demonstrated how general-purpose process simulation programs can be used to simulate gas dominated two-phase pipelines. In the August 2008 TOTM, we discussed the value of the simple Flanigan correlation and how it can be used to model and analyze the behavior of a wet gas transmission pipeline. The results of the Flanigan correlation were compared with more rigorous calculation methods for multiphase pipelines.

    In this TOTM, we will study the impact of gas-oil ratio (GOR) on pressure drop in crude oil gathering systems. Specifically, pressure drop along a gathering line for nominal pressures of 690, 3450, and 6900 kPag (100, 500, and 1000 psig) and nominal pipe size of 101.6 and 152.4 mm (4 and 6 inches) was calculated using multiphase rigorous method from commercial simulator. The calculated pressure drops are presented in graphical format as a function of the oil stock tank volume flow rate and GOR. Variation of thermo physical properties was considered.

    Case Study

    For the purpose of illustration, we considered a case study for transporting a crude oil of relative density of 0.852 (°API = 34.6) at stock tank condition combined with a gas with relative density of 0.751. The selected GORs were 0 (dead oil), 17.8, 356.5, and 891.3 Sm3 of gas/STm3 of oil (0, 100, 2000, and 5000 scf/STB). The compositions of oil and gas are presented in Table 1. The oil C6+ was characterized as 30 single carbon number (SCN) [2] ranging from SCN6 to SCN35 while the gas C6+ was characterized by 10 SCN ranging from SCN6 to SCN15. For details of the SCN components, see Table 3.2 on page 64 of reference [2]. The mole fraction of SCN components were determined by an exponential decay algorithm [3].

    Table 1. Feed composition at stock condition

    table1

    The following assumptions were made:

    1. Steady state conditions
    2. The line is 1.601 km (1 mile) long with nominal size of 101.6 and 152.4mm (4 and 6 inches), onshore buried line.
    3. Segment lengths and elevation changes are presented in Table 2. This elevation profile is considered to be approximately equivalent to “rolling” terrain.
    4. Pipeline inside surface roughness of 46 microns (0.046 mm, 0.0018 inch)
    5. Line nominal pressure 690, 3450, and 6900 kPag (100, 500, and 1000 psig)
    6. The feed enters the line at 15.6 ˚C and (60 ˚F)
    7. The ground/ambient temperature, is 15.6 ˚C and (60 ˚F)
    8. Water cut is 0 (no water in the feed).
    9. Overall heat transfer coefficients of 2.839 W/m2-˚C (0.5 Btu/hr-ft2-˚F), for onshore buried line (minor effect as inlet temperature = ambient ground temperature).
    10. Simulation software ProMax [4] and using the Soave-Redlich-Kwong (SRK) Equation of State [5] for vapor-liquid equilibrium and Beggs-Brill method for two-phase pressure drop calculation [6].

    Table 2. Line segment length and elevation change

    table2

    Results and Discussions:

    The two phase (oil and gas) flow through the gathering line was simulated by ProMax with SRK EOS for vapor-liquid equilibria and Beggs-Brill for two phase pressure drop calculations. Figures 1A and 1B present the calculated pressure drop per unit length as a function of oil stock tank volume rate and GOR for nominal line diameter of 101.6 mm (4 inches) at nominal line pressure of 690 kPag (100 psig) in SI (international) and FPS (Engineering) system of units, respectively. Figures 1A and 1B indicate that as the GOR increases from 0 to 891 Sm3/STm3 (0 to 5000 scf/STB), the pressure drop increases considerably. Consequently, as the GOR increases, the line capacity decreases.

    Figures 2A, 2B, 3A, and 3B present the results for the same line size but at nominal pressures of 3445 and 6900 kPag (500 and 1000 psig), respectively. Contrary to Figure 1, Figures 2 and 3 indicate that at these higher pressures as the GOR increases, the pressure drop decreases for low GOR value. However, for further increase of GOR the pressure drop increases considerably.

    Similar calculations were repeated for another line with nominal pipe size of 152.4 mm (6 inches) and the simulation results are presented in Figures 4 through 6. Figures 4 through 6 also demonstrate the same impact of GOR on the pressure drop, at higher pressures and low GOR, the pressure drop decreases. However, the impact of low GOR at higher pressures is less compared to the smaller line diameter.

    fig1a

    Figure 1A (SI). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 690 kPag for 101.6 mm pipe diameter

    fig1b

    Figure 1B (FPS). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 100 psig for 4 in pipe diameter

    fig2a

    Figure 2A (SI). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 3445 kPag for 101.6 mm pipe diameter

    fig2b

    Figure 2B (FPS). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 500 psig for 4 in pipe diameter

    fig3a

    Figure 3A (SI). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 6900 kPag for 101.6 mm pipe diameter

    fig3b

    Figure 3B (FPS). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 1000 psig for 4 in pipe diameter

    Conclusions

    The following conclusions can be made based on this case study:

    1. The GOR has a large impact on the capacity of crude oil gathering lines. In general as GOR increases the pressure drop increases which lowers the line capacity.
    2. At high pressures and low GOR, pressure drop is lower than the pressure drop for dead oil (solution gas is zero) because the viscosity of live oil is lower than viscosity of dead oil. This effect is bigger for the smaller line diameter.

    To learn more about similar cases and how to minimize operational problems, we suggest attending our PF 45 (Onshore Gas Gathering Systems: Design and Operation), G4 (Gas Conditioning and Processing), PF81 (CO2 Surface Facilities), PF4 (Oil Production and Processing Facilities), and PL4 (Fundamentals of Onshore and Offshore Pipeline Systems) courses.

    John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

    By: Mahmood Moshfeghian

    Reference:

    1. Ellul, I. R., Saether, G. and Shippen, M. E., “The Modeling of Multiphase Systems under Steady-State and Transient Conditions – A Tutorial,” The Proceeding of Pipeline Simulation Interest Group, Paper PSIG 0403, Palm Spring, California, 2004.
    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. Moshfeghian, M., Maddox, R.N., and A.H. Johannes, “Application of Exponential Decay Distribution of C6+ Cut for Lean Natural Gas Phase Envelope,” J. of Chem. Engr. Japan, Vol 39, No 4, pp.375-382 (2006)
    4. ProMax 3.2, Bryan Research and Engineering, Inc., Bryan, Texas, 2014.
    5. Soave, G., Eng. Sci. Vol. 27, No. 6, p. 1197, 1972.

    Brill, J. P., et al., “Analysis of Two-Phase Tests in Large-Diameter Flow Lines in Prudhoe Bay Field,” SPE Jour, p. 363-78, June 1981.

    fig4a

    Figure 4A (SI). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 690 kPag for 152.4 mm pipe diameter

    fig4b

    Figure 4B (FPS). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 100 psig for 6 in pipe diameter

    fig5a

    Figure 5A (SI). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 3445 kPag for 152.4 mm pipe diameter

    fig5b

    Figure 5B (FPS). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 500 psig for 6 in pipe diameter

    fig6a

    Figure 6A (SI). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 6900 kPag for 152.4 mm pipe diameter

    fig6b

    Figure 6B (FPS). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 1000 psig for 6 in pipe diameter

  • The Importance of Leadership in Process Safety Management

    The first pillar of Risk Based Process Safety Management is “Commitment to Process Safety.”  A formalized mentoring system can ensure workforce involvement, compliance with company and regulatory requirements, increase the competency of personnel and enhance the process safety culture of the entire organization.  Within this element there are several essential features that lead to a more effective process safety culture.

    Providing strong leadership is critical for any organization that strives to manage the risk associated with the activities associated with process safety.  Leadership is a skill that is not necessarily intuitive to managers and mentors.  Leadership is a skill that can be learned.

    In this Tip of the Month (TOTM), we explore process safety leadership.

    This TOTM is part of a paper that was developed by John M. Campbell (JMC) Instructor/Consultants Clyde Young and John Kanengieter for presentation at the Center for Chemical Process Safety (CCPS) 9th Global Conference on Process Safety [1].

    Over the last several years, significant resources have been devoted to examining the issue of process safety culture, and strong leadership has been cited as a key element to enhance a process safety culture.  Study of major accidents within the oil, gas, chemical and allied industries have found that the safety culture of organizations is often proposed as a contributing factor, and development of a culture of process safety as the solution.  Presentations at symposia and conferences point to enhancing culture and providing leadership as necessary to address breakdowns in process safety management systems.

    The first pillar of the Center for Process Safety (CCPS) Guidelines for Risk Based Process Safety (RBPS) is “Commit to Process Safety.”   Supporting this pillar is the element “Process Safety Culture”, which is defined as, “ the combination of group values and behaviors that determine the manner in which process safety is managed.”   One of the four essential features of process safety culture is “strong leadership.”

    Leadership

    What is “leadership”?  It has been described as “organizing or influencing a group to achieve a common goal”.  This would intimate that the leader is a boss or manager, but is a manager necessarily an effective leader?   There is considerable literature about leadership.  This literature includes quotes about leadership, how to find “natural” leaders and how to develop leadership skills.  There are workshops about leadership and even university degrees in leadership.  If there are so many resources dedicated toward understanding and teaching leadership, why is leadership listed as something that needs to be enhanced in symposia, papers and reports that deal with managing process safety in high hazard activities?  It may be because leadership and culture are considered human factors. When associated with process safety, they are known as factors that can lead to loss of the standards of consistently reliable human performance.  These standards are relied on as part of an organization’s defenses against process safety incidents.

    Every person working in the oil, gas, chemical and allied industries should perform their jobs under the guidance of a process safety management system.  CCPS defines a management system as a “formally established and documented set of activities designed to produce specific results in a consistent manner on a sustainable basis.”  Producing specific results in a consistent manner all the time requires that all personnel perform at a high level.  If culture is defined simply as “the way we do things around here”, this is influenced greatly by leadership.  But leadership doesn’t reside in the role of one person.  Leadership needs to be imbedded within the organization with every person.  This is a skill that can be learned by all and dependence on one individual with authority or one person who might be considered a “natural” leader can lead to failure of the system.

    When teams cease to function effectively and breakdowns are discovered in the system to manage process safety, it is highly likely that there is a breakdown in goals, roles and expectations in the team.

    Every person working in or supporting the operation of a high hazard process must be able to recite and explain the goal of every team they work with.  There should never be in any doubt what every team’s goal is.

    Because we may and probably do work on several teams, it is vital that we are clear of our role on each team.  What is my primary function to support achieving the goal? There should never be in any doubt what every person’s role is on that team.

    Does each person on the team have a concisely developed set of expectations for individual and team behavior?  Is there some way for the team to check that the expectations are being met?  What is the procedure for addressing deviation from expectations?

    A PetroSkills client recently asked for a one-day Overview of Risk Based Process Safety Management for Upper Level Management.  Four sessions of this overview have been delivered around the world to the business unit managers and their direct (team members) reports?.  Leadership and working as effective teams are two elements of the session that address the issue of process safety culture in this client’s operations.

    A key learning point offered by participants is that a clear understanding of goals, roles and expectations comes from leadership and exhibiting the appropriate leadership role.  Many leave the session with an action item to conduct team work sessions to establish/reaffirm goals, roles and expectations.

    If you would like a copy of the paper presented at the CCPS 9th Global Congress, contact PetroSkills.

    To develop process safety competency attend our PS-4, Process Safety Engineering;  HS-45, Risk Based Process Safety Management; and PS-2, Fundamental of Process Safety courses.  To develop competency in other skills, attend one of our other courses.

    John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

    By Clyde Young

    PetroSkills Instructor/Consultant

    Reference:

    1.     Clyde Young and John Kanengieter, “Process Safety Management Mentoring:  Developing Leaders”, The (CCPS) 9th Global Conference on Process Safety,  the Center for Chemical Process Safety , April, 2013.

     

  • Simple Equations to Approximate Changes to the Properties of Crude Oil with Changing Temperature

    This Tip of the Month describes simple equations to approximate changes to the properties of crude oil with changing temperature.   Changes in crude oil density and specific heat, or heat capacity, can be estimated from graphs and/or more elaborate computer simulation.  The latter generally requires access to a process simulator and characterization data for the crude oil.  A suitable, tuned computer model is likely the most accurate method of estimating the fluid properties, but is not always available.  Direct laboratory measurement is also possible if facilities and oil samples are available and a high degree of accuracy is required.

    Graphs, which were originally generated from empirical data, can be useful, and their accuracy is suitable for most engineering applications. However; use of the data in subsequent calculations requires users to interrupt the calculation, look up a number from a graph, and then proceed with the calculation.  The simple curve-fit equations presented provide the required data suitable for use in spreadsheet or hand calculations.

    Crude Oil Density

    Figure 1A (1) depicts the change in specific gravity with temperature for crude oils of varying API gravity.

    Figure 1A Crude Oil Specific Gravity vs Temperature (1)

    This graph was compared with density-temperature data from Table D-1, API Publication 421, “Monographs on Refinery Environmental Control – Management of Water Discharges” (2).  The colored lines in Figure 1B show the API data superimposed on the original graph. The two show good agreement.

    Figure 1B Data from API Publication 421 superimposed on Figure 1A (data from (2))

    Curve fitting data from Figure 1A resulted in Equation 1 for FPS units.  Equation 2 provides the SI equivalent.

    Figure 1C displays the output from Equation 1 superimposed as colored lines on the original graph (Figure 1A). Although the simple equation does not align perfectly, results are sufficiently accurate for most engineering calculations.  Compared to the data from API Publication 421,  Equation 1 produces a maximum errors of +0.25% and -0.3%.

    Figure 1C Simple Equation 1 superimposed on Figure 1A

     

    Heat Capacity

    A similar approach was used to develop a simple equation for the variation of the Heat Capacity or Specific Heat of crude oil as a function of API gravity and Temperature.  Data, extracted from Figure 2A (1), was regressed to obtain the algorithm presented as Equation 3 for FPS units: Equation 4 for SI units.  Note that the algorithm was developed for a crude oil with a UOP index of 11.8 (indicating intermediate, paraffinic-naphthenic crude).  If the UOP index is known, the correction factor illustrated on the graph could be applied to the output from Equation 3 or 4.

    Figure 2A Heat capacity of crude oil vs Temperature (1)

    The resulting equation is presented as Equation 3 for FPS units, and Equation 4 for SI units.

    Figure 2B Simple Equation 3 superimposed on Figure 2A

     

    Summary

    The simple equations provide approximations for the variation of density and specific heat of crude oils of varying API gravity.  Neither algorithm provides a perfect match with the underlying data.  However; data from varying sources do not always correlate.  Figure 3A (1) depicts an alternate source of density correction for crude oils for varying API gravity and temperature.  Figure 3B shows data from API Publication 421 (colored lines) superimposed on a portion of Figure 3A.  Unfortunately, overlap of the data is limited, but clearly there is a very poor correlation – even the trend as API gravity increases is reversed between the API data and the data presented in Figure 3A.  Note that the trend represented in Figure 1A is supported by data from API 421 Appendix D.  These data (from the API Publication 421) are taken here as being most reliable.  Overall, Figure 1A shows quite good agreement with the API data, so the algorithm (Equations 1 and 2) was developed using data from Figure 1A as the source.

    Figure 3A Crude Oil Density Correction Factor (Hankinson et al, 1979) (1)
    Figure 3B Portion of Figure 3 A with data from API 421 superimposed

    To learn more about similar cases and how to minimize operational problems, we suggest attending our G40 (Process/Facility Fundamentals), G4 (Gas Conditioning and Processing), PF81 (CO2 Surface Facilities), and PF4 (Oil Production and Processing Facilities) courses.

    John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

    By Wes Wright

    Works Cited

    1. Manning, Francis S. and Thompson, Richard E. OILFIELD PROCESSING Volume Two: Crude Oil. Tulsa : PennWell Publishing Company, 1995. ISBN 0-87814-354-8.

    2. Institute, American Petroleum. Management of Water Discharges: Design and Operation of Oil-Water Separtors. Washington : API, 1990. API Publication 421.

     

  • Transportation of Ethane by Pipeline in the Dense Phase

    In the January, February, and March 2012 tips of the month (TOTM) we discussed the transportation of carbon dioxide (CO2) in the dense phase region. We illustrated how thermophysical properties changed in the dense phase and studied their impacts on pressure drop calculations. We showed that the effect of the numerical range of values for the overall heat transfer coefficient on the pipeline temperature is significant.

    In this TOTM, we will study the transportation of ethane by pipeline in the dense phase region. For a case study, a mixture of ethane containing a small fraction of methane and propane was considered. The pressure and temperature profiles along the pipelines are calculated and plotted on the feed phase envelope. In addition, the pump power requirement, pressure and temperature profiles for a single pipeline with a lead pump station are compared with the option of dividing the same line into three equal segments having one lead pump station and two intermediate pump stations.

    Calculation Procedure:

    For a pure compound above critical pressure and critical temperature, the system is often referred to as a “dense fluid” or “super critical fluid” to distinguish it from normal vapor and liquid (see Figure 1 for carbon dioxide in December 2009 TOTM [1]).

    The same step-by-step calculation procedure described in the February 2012 TOTM [2] was used to determine the pressure and temperature profiles in a pipeline.

    In the following section we will illustrate the pressure drop calculations for transporting ethane mixture in the dense phase. For details of pressure drop equations in the gas and liquid phases refer to the January 2012 TOTM [3].

    Case Study:

    For the purpose of illustration, we considered a case study for transporting 31393 kg/h (69209.0 lbm/hr) equivalent to 275000 tonne/y of the cited  ethane  mixture with the composition presented in Table 1. The mixture is available at the pressure and temperature presented in Table 1.

    The following assumptions were made:

    1. Horizontal pipeline, no elevation change
    2. Steady state conditions
    3. The pipeline is 1380 km (858 mile) long with an inside diameter of 208.6 mm (8.212 in), onshore buried line.
    4. Pipeline inside surface roughness of 46 microns (0.046 mm, 0.0018 inch) which is equivalent to inside surface relative roughness (roughness factor), ε/D, of 0.00022.
    5. Delivery Pressure is 5.3 MPa  (769 psia)
    6. The ground/ambient temperature, is 18.3 ˚C, (65 ˚F)
    7. Overall heat transfer coefficient of 2.839 W/m2-˚C (0.5 Btu/hr-ft2-˚F), for onshore buried line.
    8. Pump efficiency is 50%, this is the worst case, the actual pump efficiency is in the range of 50-85%).
    9. Simulation software: ProMax [4] and using Soave-Redlich-Kwong (SRK) [5] Equation of State.

    The block diagrams for two options studied are presented in Figure 1. In option A, only one pump station and a single long segment were considered. In option B, for the same pipeline and feed conditions, one lead and two intermediate pump stations with three equal pipeline segment were considered.  Each segment length is 460 km (286 mile), which is 1/3 of total length and all have the same inside diameter of 208.6 mm (8.212 in).  The delivery pressures for both options are the same.

    Results and Discussions:

    Figure 2 presents the phase envelope for the ethane mixture with the composition presented in Table 1. According to Figure 2, the feed at the pump suction conditions of 0°C (32°F) and 3000 kPa (435 psia) presented in Table 1 is in the liquid phase.  In order to deliver the ethane mixture at pressure of 5.3 MPa (769 psia), for option A this liquid is pumped to a pressure of 13.6 MPa (1972 psia) before entering the pipeline. Due to pumping, the feed temperature rises from 13.6 MPa (1972 psia) before entering the pipeline. A step-by-step pump calculation with increments of 2 MPa (290 psia) for the discharge pressure reveals that the temperature rise is linear with pressure. This significant temperature rise is due to compressibility of ethane mixture.

    Figure 2. Phase envelope for ethane mixture.

    The pumping and pipeline pressure-temperature paths for option A are plotted on the phase envelope and presented in Figure 3. As shown in this figure, the ethane mixture at the pump discharge (pipeline inlet) is in the supercritical region (dense phase).  Figure 3 indicates that as the feed enters the pipeline, its temperature drops rapidly and remains constant and very close to the ambient temperature.

    Figure 3. Phase envelope, pumping path and dense phase pipeline pressure-temperature profile.

    Figure 4 presents the calculated pipeline pressure profiles for options A and B. For option A, the inlet pressure is 13.6 MPa (1972 psia) to assure ethane mixture delivery at 5.3 MPa (769 psia). Similarly, in option B at each pump station the pressure is increased to 8.2 MPa (1189 psia).

    Figure 5 presents the calculated pipeline temperature profiles for options A and B. The constant ambient temperature of 18.3˚C and (65˚F) is also plotted. In option B, the discharge temperature for lead pump station is 11.2˚C, (52.1˚F) which is below the ambient temperature. For both options, the pipeline temperature rapidly approaches the ambient temperature within the first 50 km (31 mile).

    Figures 6 and 7 present the density and velocity profiles along the pipeline, respectively. For a crude oil cross-country pipeline, the velocity is in the range 1.5 to 2.5 m/s (5 to 8 ft/sec). The abrupt change of density, and consequently velocity, along the first 60 km (37.3 mile) is due to the ethane mixture temperature drop, which approaches the ground temperature.

    Figure 4. Pipeline pressure profiles for options A and B.

    Figure 5. Pipeline temperature profiles for options A and B.

    Figure 6. Pipeline fluid density profile for option A

    Figure 7. Pipeline fluid velocity profile for option A

    The total pump power requirements with pump efficiency of 50% for options A and B are 457 and 504 kW (381.6 and 420.8 hp), respectively. This is for screening estimate only. Normally, work is done in collaboration with the pump manufacturers for better efficiencies based on CAPEX and OPEX.

    Conclusions:

    Based on the results obtained for the case study considered in this TOTM, the following conclusions can be made:

    1. During the pumping of ethane mixture, the temperature rises linearly with pressure (Figure 2).
    2. The feed temperature approaches the ground temperature rapidly (Figure 5). This may not be the case for lower overall heat transfer coefficient.
    3.  A single pipeline with only one lead pump station (option A) requires smaller pump power compared to the option of one lead pump and two intermediate pump stations (option B). Due to higher pressure in option A, wall thickness will be higher.
    4. A complete cost analysis of pumping requirement vs pipeline cost should be made to determine the optimum pipeline diameter, wall thickness and power requirement.
    5. The point not considered but worth mentioning is that ethane is very difficult to seal. We would work with pump and seal manufacturers for selecting the correct dry gas seal. This selection could determine the overall system reliability.

    To learn more about similar cases and how to minimize operational problems, we suggest attending our G40 (Process/Facility Fundamentals), G4 (Gas Conditioning and Processing), PF81 (CO2 Surface Facilities), PF4 (Oil Production and Processing Facilities), and PL4 (Fundamentals of Onshore and Offshore Pipeline Systems) courses.

    John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

    By: Dr. Mahmood Moshfeghian

    Reference:

    1. Bothamley, M.E. and Moshfeghian, M., “Variation of properties in the dese phase region; Part 1 – Pure compounds,” TOTM, http://www.jmcampbell.com/tip-of-the-month/2009/12/variation-of-properties-in-the-dense-phase-region-part-1-pure-compounds/, Dec 2009.
    2. Moshfeghian, M., ”Transportation of CO2 in the Dense Phase,” TOTM, http://www.jmcampbell.com/tip-of-the-month/2012/02/ , Feb 2012
    3. Moshfeghian, M., ”Transportation of CO2 in the Dense Phase,” TOTM, http://www.jmcampbell.com/tip-of-the-month/2012/01/, Jan 2012
    4. ProMax 3.2, Bryan Research and Engineering, Inc, Bryan, Texas, 2014.
    5. Soave, G., Chem. Eng. Sci. 27, 1197-1203, 1972.

  • Debriefing Jobs Provides Several Benefits Associated With Process Safety

    A pillar of Risk Based Process Safety (RBPS) is Learn from Experience.  The work we do and the processes we use to analyze our work provide significant learning opportunities to enhance process safety competency.  This is a derivative of Kolb’s experiential learning cycle [1], but many times we fail to take advantage of the learning opportunities available to us unless there is an incident or a near miss.

    This Tip of the Month (TOTM) will introduce a simple method for debriefing the job tasks we perform to close the loop on this cycle and capture appropriate data to develop competency, work safely and capture near miss/incident data quickly and efficiently.

    Conducting a simplified job hazard analysis will ensure that all hazards are identified, managed, and mitigated prior to performing work.  Performing a simple debrief at the conclusion of the work will ensure that we learn from the experience. By considering every job to be performed a learning opportunity, the experiential learning cycle can be used to identify what was done, how well it was done, and how we might improve in the future.

    This Month’s Tip was recently presented at the Mary K. O’Connor Process Safety Symposium at Texas A&M University [1].

    One of the pillars of the Center for Chemical Process Safety’s (CCPS) Guidelines for Risk Based Process Safety is “Learn from Experience.”  What does this mean?

    The elements of this pillar include:

    • auditing,
    • management review and continuous improvement,
    • measurement and metrics and
    • incident investigation.

    Each of these elements provides findings, lessons and data that are useful for learning and thus changing and enhancing behaviors and attitudes.  The change and enhancement will influence an organization’s culture and ultimately push the organization toward a learning culture.

    These are not the only opportunities available for organizations to learn from experience.  Metrics and audits can allow a general overview of process safety performance.  Incident investigation insures that when reported, incident information is transmitted to all who will benefit from the learning.

    The job hazard analysis process that many organizations use to identify and mitigate hazards provides a tremendous opportunity to capture data and use the experiential learning cycle if the job is debriefed properly after completion.  This paper will provide guidance and explain the benefits that can be derived from debriefing completed jobs.

    At the 2008 symposium, this author presented a paper entitled “Three Simple Things to Improve Process Safety Management.”  One of those simple things was to conduct a formalized Job Hazard Analysis (JHA) for the tasks being performed in the life cycle of a process.  That paper presented a checklist that could be used to guide personnel in the process of conducting a JHA.  (See checklist at end of this paper)

    Many facilities have embraced the concept of conducting JHA.  They may be called something else  (job safety analysis, job safety checklist, job task analysis) but the process is essentially the same.  The job or task is identified and analyzed step by step.  The analysis is to identify hazards that may be involved with each step and then develop strategies to mitigate the hazards.  This sounds simple in theory, but in reality there are many things that can and do go wrong with this process.

    To provide consistency and to make it easier to track that these analyses have been completed, standardized checklists and forms have been created that list the most common hazards that can be found with a job and logically guide the user toward identification and mitigation of hazards.  Experience shows that after these forms and checklists have been used regularly, some personnel have a tendency to try and short cut the process.  This leads to what is known as “pencil whipping” the JHA.  In other words, personnel will complete the checklist or form without actually performing the analysis required.   Familiarity with the forms and checklists may drive personnel to identify common hazards, but do little to mitigate the hazards.  For example, a common checklist item is “slips, trips and fall hazards”.  Personnel will identify that the ground is rutted or that there is ice on the ground, but few will actually smooth the ground or cover the ice with sand to mitigate the hazards identified.

    It is generally agreed among those who supervise personnel performing JHAs that the most important part of the process is not the completion of the forms and checklists but the discussion that happens among a group performing the work.  In order to focus the discussion and insure that all issues are addressed, the JHA checklist at the end of this paper can be used.  The JHA checklist is not intended to replace the checklists and forms that an organization may already have in place.  The JHA checklist can enhance the process by focusing a group’s thoughts on individual checklist items.  By answering each question a work group should be able to identify all issues associated with any job they are conducting.

    As work groups become more familiar with the JHA checklist and the process of discussing and documenting the efforts of the group, a simplified method can be adopted.  By answering six key questions, a group of workers can focus discussion on the issues that are most important.   The six questions and the benefits of using them include:

    What are we doing?  If we can’t answer this question completely and in simple terms, then we should not be doing the job.  A simple explanation will insure that all members of the team are working toward the same goal.

    What is the most dangerous part?  If we can identify the most dangerous part of what we are doing we have identified all hazards, ranked them and determined the most dangerous part.

    What will we do to protect ourselves?  Answering this question ensures that all mitigation measures have been put into place and that all personnel know what is being done.

    How will we know we are changing what we are doing?  To answer this question effectively, we will need to be creative and analytical.  Examination of the work site, knowledge of simultaneous operations, and competency in our job will be required.  Anticipating potential changes will insure that we are not surprised when things do change.

    What will we do about it?  Again, creativity and analytical thinking are critical here.  Many times we hear the phrase, “prior planning prevents poor performance”.  Effectively answering this question insures that performance will be as designed.

    How will we know we are finished?  Review of completed job hazard analysis documents has shown that it may be difficult to determine at what point the job is complete.  If the permit for the job being performed provides a scope of work like, “replace mechanical seal in hot oil pump”, once the seal is replaced, there are numerous tasks that still need to be performed before the job is complete.  Numerous times the JHA does not go beyond analyzing the tasks associated with the scope of work and do not consider additional tasks; like testing, clean up and turnover to operations.

    As previously mentioned, most supervisors believe that the discussion associated with this type of analysis is more important than the completion of the form used to show that the JHA has been performed.

    What about the form though?

    • What happens at the conclusion of the job?
    • Does anyone review the form to determine if all the hazards were found and mitigated?
    • Does anyone follow up with the work group to see if anything happened that made them change the work?
    • How should this review be performed and what are the benefits that will be gained by this?
    • How can we learn from our experience?

    Developing competent personnel is an ongoing process for most organizations.  A great deal of literature exists on the most effective methods of developing competency in adults. Training sessions are delivered using the concept of Kolb’s theory of the experiential learning cycle.  According to Kolb [2], this type of learning can be defined as “the process whereby knowledge is created through the transformation of experience.” [i] In other words, adults learn best when they are actively experiencing something and not just listening to lectures or instructor centered learning.

    Experienced trainers who deliver adult learning sessions use a process of debriefing to allow reflection, reinforce learning and help the learner apply the knowledge to their life.  It is generally acknowledged in the training industry that most real learning takes place in the debrief.  This is the opportunity for learners to reflect and develop knowledge from the activity, in our case the job performed.

    Very simply, debriefing a learning activity should focus on three questions.  What?  So What?  Now What?

    What? is the question that guides the learning toward reflection and what just happened.  This question provides a starting point to discover what everyone involved experienced.

    So What? is the question that leads to drawing conclusions and exploring alternate methods.

    Now What? leads to future planning and continuous improvement initiatives that will be used to strengthen the organization’s culture and work processes.

    If we return to question six of the job hazard analysis process, “How will we know we are done?”, the final answer for this question would be, “When we have completed the debrief of the job performed.”  There are five questions that should be used for debriefing a job.  These five questions, how they relate to the standard debriefing questions and the expected lessons to learn from them include:

    What did we do?  This is the opportunity for reflection and to insure that the job has been completed appropriately.  Each member of the team should come to agreement that what is being described is what was actually done.   This is the What of debriefing.

    Did anything change while doing the job?   Reflection on this question will lead the team to determine if the job was actually performed as it was initially described and analyzed.  This is the question that will also lead to identify incidents for investigation.  If anything unusual occurred during the task, reporting should be more efficient because the incident will be fresh in everyone’s mind.  Capturing these incidents and changes now will help modify future work orders and insure that we learn something from this experience.  This is the So What of the debriefing cycle.

    Did anybody get hurt?  This question should be answered with all personnel examining themselves for strains, pulled muscles, bumps, bruises, cuts, scrapes, twisted joints, twinges in the back and a general self examination for good health.  Any small injury or potential illness should be recorded here.   This will insure that a worker does not leave the job without reporting an injury or illness, and then visit a medical provider later because something cropped up.  Having someone discover they have been injured after leaving the worksite is a problem for managers.  This allows measures to be taken early to manage the injury or illness for reporting purposes.  Here and the next question is where more exploration of the “What” is performed.

    Did anybody come close to getting hurt?  This is the question that will capture near miss incidents quickly.  Near miss reporting programs fail for numerous reasons.  Lack of understanding, lack of motivation, blaming the reporter, and convenience of reporting are reasons that near misses may not be reported.  Reflection and discussion about the completed job will insure that any near miss is reported quickly.  This will lead to creation of a more comprehensive database that can be used to predict trends and identify problems areas in processes.

    What would we do differently?  This is the question that will tie everything together into a plan for the future.  Recommendations and action items should be generated from this final question so that future jobs can be analyzed with more speed and efficiency.  Potential training and competency development issues may be discovered.  Procedures for modification may be identified.  Latent conditions that are not readily apparent may be identified and mitigated before they become active failures.

    The Now What of the debriefing cycle is:

    • Conducting an effective job task analysis and following with an effective debriefing of the job will yield several benefits.
    • Competency gaps of personnel associated with the work will be identified.
    • Training topics and on the job mentoring for personnel with these identified gaps, can be more quickly delivered.
    • Procedural modifications that are necessary to insure that work is performed safely and efficiently will be quickly identified and addressed.
    • Potential process safety incidents will be quickly identified and investigated.
    • Near miss incidents will be reported quickly and the organization’s near miss/incident database will be enhanced.

    The process described in this paper can be expanded to any job and any work group.  Consider an engineering team who is working on the design of a new process to be considered for construction.  Conducting an effective job task analysis in the beginning stages of the project will insure that roles, goals and expectations are addressed and known.  Conducting an effective debrief at the conclusion, or even at selected stages of a project, will enhance the project team’s effectiveness and insure that all team members are always striving to meet the goal of the project.  The attached checklist for engineering projects, at the end of this paper, may be helpful for focusing a team’s efforts.

    Opportunities exist in all phases of operations and in all activities performed to keep processes safe.  It is important that all personnel be aware that learning from experience happens every day and these lessons learned need to be captured and stored for future use.

    To develop process safety competency attend our PS-4, Process Safety Engineering;  HS-45, Risk Based Process Safety Management; and PS-2, Fundamental of Process Safety courses.  To develop competency in other skills, attend one of our other courses.

    By Clyde Young

    PetroSkills Instructor/Consultant

    Reference:

    1.    Young, Clyde. ,” Debrief:  The experiential learning cycle, process safety competency, safe work practices, identifying and reporting of near miss/incident data”, Mary K. O’Connor Process Safety Symposium, Texas A&M University, October 29.

    2.    Kolb, David A. Experiential Learning: Experience as the Source of Learning and Development. Prentice-Hall, Inc., Englewood Cliffs, N.J. 1984.

    Job Hazard Analysis Checklist

    1. PROCEDURES

    • ·What are the procedures for the task?
    • ·What is unclear about the procedures?
    • ·What order will we use these procedures?
    • ·What permits are needed for hazard controls?

    2. EQUIPMENT AND TOOLS

    • ·What are the right tools for the job?
    • ·What is the correct way to use them?
    • ·What is the condition of the tool?

    3. POSITIONS OF PEOPLE

    • ·What could we be struck by?
    • ·What could we strike ourselves against?
    • ·What can we get caught in/on/between?
    • ·What are potential trip/fall hazards?
    • ·What are potential hand/finger pinch points?
    • ·What extreme temperatures will we be in/around?
    • ·What are the risks of inhaling, absorbing, swallowing hazardous substances?
    • ·What are the noise levels?
    • ·What electrical current/energized system could we come in contact with?
    • ·What would be a cause for overexerting ourselves?

    4. PERSONAL PROTECTIVE EQUIPMENT

    • ·What is the proper PPE?

    Hard hat, glasses/goggles, ear plugs, gloves, steel toe boots, respiratory system, fire retardant clothing

    5. CHANGING THE COURSE OF WORK

    • ·What would cause us to have to stop or rearrange the job?
    • ·What would cause us to change our tools or equipment?
    • ·What would cause us to have to change our position?
    • ·What would cause us to have to change our PPE?

    YOU HAVE THE RIGHT AND

    THE OBLIGATION TO

    STOP UNSAFE ACTS

    ENGINEERING JOB ANALYSIS

    1. PROCEDURES

    • ·What are the procedures for the task?
    • ·What is unclear about the procedures?
    • ·In what order will we use these procedures?
    • ·What is the proper timeline for these procedures?
    • ·What permits or permissions are needed for job controls?

    2. EQUIPMENT, TOOLS, DOCUMENTS

    • ·What are the right tools for the job? (software, simulators, matrixes, checklists, worksheets…)
    • ·What is the correct way to use them?
    • ·What forms will be needed for the job?
    • ·What documents will we need to produce?
    • ·What else do we need to know?

    3. INTERACTION WITH PEOPLE

    • ·What other departments need to know about this task?
    • ·Who are the personnel that need to know?
    • ·What other departments will supply information for this task?
    • ·Who are the personnel who will supply that information?
    • ·What could prevent other personnel or departments from supplying what we need?
    • ·What could prevent us from supplying what other departments need?

    4.  CHANGING THE COURSE OF WORK

    • ·What would cause us to have to stop or rearrange the job?
    • ·What would cause us to change our tools or equipment?
    • ·What would cause us to have to change our interaction with people?

    YOU HAVE THE RIGHT AND THE OBLIGATION TO

    STOP UNSAFE or UNPRODUCTIVE ACTS

  • Estimating Still Column Top Temperature in TEG Dehydration Unit

    In this Tip of The Month (TOTM), the effect of striping gas rate and TEG circulation ratio on the still column top temperature for regeneration of rich triethylene glycol (TEG) is investigated. Specifically, this study focuses on the variation of still column top temperature with reboiler pressure, TEG circulation ratio and stripping gas rate. By performing a rigorous computer simulation of TEG regeneration at reboiler pressures of 110.3 kPaa (16 psia) and 524.1 kPaa (76 psia), two charts for quick determination of still column top temperature needed for facilities type calculations are developed. In addition, the effect of theoretical number of trays in the stripping gas section is studied.

    Computer Simulation Results:

    In order to study the impact of stripping gas rate and TEG circulation rate on the still column top temperature, the TEG dehydration process was simulated using ProMax [1] software with its Soave-Redlich-Kwong (SRK) [2] equation of state (EOS). The process flow diagram used for these simulations is shown in  Figure 1.

    The water-saturated gas with a water content of 915 mg/std m3 (57 lbm/MMSCF) enters the bottom of the contactor column at 37.8°C (100°F) and 6895 kPaa (1000 psia) at a rate of 2.835×106 std m3/d (100 MMSCFD). The contactor column has three theoretical trays. The lean TEG solution enters at the top of the contactor column and flows down in the column. As shown in Figure 1, the water content of the dried gas is 10 mg/std m3 (0.63 lbm/MMSCF). The rich TEG solution contains 96.1 mass percent TEG entering the still column at 100°C (212°F) and 515 kPaa (74.7 psia). The reboiler temperature was set at 204.4°C (400°F) and boil-up ratio of 0.1 (molar bases). Two theoretical trays in the regenerator (still) column (NR = 2) and two theoretical trays (NS = 2) in the striping gas section were specified. The striping gas enters the bottom of the stripping gas section at 204°C (399°F) and 524 kPaa (76 psia). Methane was used for the stripping gas at a rate of 56.3 std m3/h (1893 scf/hr). The regenerated lean solution contains 99.6 mass percent TEG and the ratio of stripping gas to lean TEG liquid volume rates is 20 std m3 of gas/std m3 of lean TEG solution (2.67 scf/sgal) or a mass ratio of 28.3. The regenerator (still) top temperature is 91.4°C  (196.5°F). If the same stripping gas was sparged directly into the reboiler (NS = 0, no stripping gas section), with everything else remaining the same, the  regenerated solution contains 99.2 mass percent TEG and  the regenerator column top temperature remains practically the same and is 91.1°C  (196°F). For the above case the number of theoretical trays in the still column is increased from 2 to 3 (NR = 3); the lean TEG concentration increased slightly from 99.6 to 99.8 mass percent but the regenerator column top temperature remained the same.

    Using a similar set up as is shown in Figure 1, several simulations were performed for a range of stripping gas rates, for NR=2, NS=0 and for two reboiler pressures of 110.3 and 524 kPaa (16 and 76 psia) and temperature of 204.4°C (400°F). The results of these simulation runs are presented in Figures 2 to 5.

    Figure 1. Sample results using ProMax [1] for TEG dehydration with reboiler P=110.3 kPaa (16 psia) with NR=2 and NS=2

    Figures 2 presents the variation of still column top temperature with circulation ratio (mass basis) and stripping gas rate at top pressure of 101.3 kPaa (14.7 psia) and reboiler pressure of 110.3 kPaa (16 psia) operating at 204.4°C (400°F).

    As was discussed in the August 2013 TOTM, regeneration of TEG at higher reboiler pressure has several advantages such as preventing the emission of harmful contaminants like benzene, toluene, ethylbenzene, xylenes (BTEX), and hydrogen sulfide to the environment [3]. Therefore, similar diagrams as shown in Figure 2 were generated for top pressure of 515.2 kPaa (74.7 psia) and reboiler pressure of 524.1 kPaa (76 psia) at 204.4°C (400°F). Figure 3 presents the variation of still column top temperature for such a high reboiler pressure.

    Fig 2. Variation of still column top temperature with circulation mass ratio and stripping gas rate at top P=101.3 kPaa (14.7 psia) and reboiler P=110.3 kPaa (16 psia) at 204.4°C (400°F)

    Fig 3. Variation of still column top temperature with circulation msss ratio and stripping gas rate at top P=515.2 kPaa (74.7 psia) and reboiler P=524.1 kPaa (76 psia) at 204.4°C (400°F)

    Figures 2 and 3 can be used for a quick determination of the still column top temperature for a given stripping gas rate and TEG circulation ratio either at low or high reboiler pressure. The two reboiler pressures selected in this study are typical operating pressures. For generation of data for Figures 2 and 3, the stripping gas was sparged directly into the reboiler; therefore,  the number of theoretical trays for stripping gas section is zero (NS=0). The corresponding figures in terms of TEG circulation volume ratio are presented in the Appendix (Figures 2A and 3A).

    Generally, either 0, 1, or 2 theoretical trays in the stripping gas section is used. In order to investigate the effect of the number of theoretical trays in the stripping gas section (NS) on the still column top temperature, simulations were performed for the cases of NS=0 and NS=2 for two constant stripping gas rates.

    Figures 4 and 5 present the results of these simulations for low and high reboiler presssures of  110.3 kPaa (16 psia) and 524.1 kPaa (76 psia), respectively. The reboiler temperature for all cases  was set at 204.4°C (400°F).

    Figures 4 and 5 clearly indicate that the still column top temperature is independent of the number of theoretical trays in the stripping gas section. Therefore, Figures 2 and 3 can be used for any number of theoretical trays in the stripping gas section.

    Fig 4. Effect of the number of theoretical trays (NS) on the still column top temperature at various circulation ratio and stripping gas rate at top P=101.3 kPaa (14.7 psia) and reboiler P=110.3 kPaa (16 psia) at 204.4°C (400°F)

    Fig 5. Effect of the number of theoretical trays (NS) on the still column top temperature at various circulation ratio and stripping gas rates at top P=515.2 kPaa (74.7 psia) and reboiler P=524.1 kPaa (76 psia) at 204.4°C (400°F)

    Similar study also showed that the feed gas temperature to the contactor column has no effect on the still column top temperature. The results of this study are shown in Figures 6 and 7 of the Appendix.

    Conclusions:

    In this TOTM, the effect of circulation ratio, stripping gas rate, theoretical number of trays, and the feed gas temperature to the contactor column on the still column top temperature for regeneration of TEG concentration at low and high reboiler pressure operating at 204.4°C (400°F) was studied. Two charts for a quick determination of the still column top temperature at a specified stripping gas rate and circulation ratio to achieve a desired level of lean TEG concentration were prepared and presented in Figures 2 through 3 (see the corresponding figures in the Appendix). These charts are based on the rigorous calculations performed by computer simulations and can be used for facilities type calculations for evaluation and trouble shooting of an operating TEG dehydration unit. In addition, the following observations were made:

    1. The still column top temperature is independent of the number of theoretical trays in the stripping gas section (NS) and feed gas temperature to the contactor column.
    2. As the stripping gas rate increased, the still column top temperature decreased.
    3. As the TEG circulation ratio increased, the still column top temperature decreased.
    4. Pressurized reboiler results in much higher still column top temperature than the atmospheric reboiler.

    To learn more, we suggest attending our G40 (Process/Facility Fundamentals), G4 (Gas Conditioning and Processing), G5 (Gas Conditioning and Processing-Special), and PF81 (CO2 Surface Facilities), PF4 (Oil Production and Processing Facilities), courses.

    John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

    By: Dr. Mahmood Moshfeghian 

    References:

    1. ProMax 3.2, Bryan Research and Engineering, Inc., Bryan, Texas, 2013.
    2. Soave, G., Chem. Eng. Sci. Vol. 27, No. 6, p. 1197, 1972.

    Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2013/08/teg-dehydration-how-does-the-stripping-gas-work-in-lean-teg-regeneration/, Tip of the Month, August 2013.

    Appendix

    Fig 2A. Variation of still column top temperature with circulation volume ratio and stripping gas rate at top P=101.3 kPaa (14.7 psia) and reboiler P=110.3 kPaa (16 psia) at 204.4°C (400°F)

    Fig 3A. Variation of still column top temperature with circulation volume ratio and stripping gas rate at top P=515.2 kPaa (74.7 psia) and reboiler P=524.1 kPaa (76 psia) at 204.4°C (400°F)

    Fig 6. Variation of still column top temperature with circulation mass ratio and feed gas temperature to the contactor column at a specified stripping gas rate at top P=101.3 kPaa (14.7 psia) and reboiler P=110.3 kPaa (16 psia) at 204.4°C (400°F)

    Fig 7. Variation of still column top temperature with circulation mass ratio and feed gas temperature to the contactor column at a specified stripping gas rate at top P=515.2 kPaa (74.7 psia) and reboiler P=524.1 kPaa (76 psia) at 204.4°C (400°F)

  • Onshore Natural Gas Pipeline Transportation Alternatives: Capital Cost Comparisons

    In recent TOTMs (January through April, August, and September 2012 and again in January 2013), we discussed several aspects of the physical behavior and transportation of carbon dioxide (CO2) and natural gas in the dense phase. We illustrated how thermophysical properties change in the dense phase and their impacts on pressure drop calculations. The pressure drop calculation utilizing the liquid phase and vapor phase equations was compared.

    In the August 2012  (TOTM), we studied transportation of rich natural gas in the dense phase region and compared the results with the case of transporting the same gas using a two phase (gas-liquid) option. Our study highlighted the pros and cons of dense phase transportation.

    In September 2012 (TOTM), we analyzed pipeline transportation of a lean natural gas at a wide range of operating pressures from the relatively low pressure typical in many gas transmission pipelines to much higher pressures well into the dense phase region.

    In January 2013 (TOTM), we estimated capital costs (CAPEX) as a tool to compare then selected the operating conditions and associated facilities for a long distance – high volume flow gas transmission pipeline.

    In this month’s Tip of the Month (TOTM), we will revisit the January 2013 (TOTM) and continue to explore alternatives specifically for onshore natural gas transportation in pipelines. This month’s focus is also on the estimation of capital costs as a tool to compare then select the operating pressures and associated facilities for a long distance, high volume flow gas transmission pipeline.

     

    Case Study:

     We will continue to use a similar case study basis as used in the September 2012 TOTM. The gas composition and conditions are presented in Table 1. For simplicity, the calculations and subsequent discussion will be done on the dry basis. The feed gas dew point was reduced to -40 ˚C (-40 ˚F) by passing it through a mechanical refrigeration dew point control plant. The resulting composition and conditions of the lean gas are also presented in Table 1. The lean gas has a Gross Heating Value of 40.33 MJ/Sm3 (1082 BTU/scf). The pipeline parameters are:

    • Length is 1609 km (1000 miles) long
    • Pipeline outside diameter is 1067 mm (42 inches) for cases A through C. Case D outside diameter is: 914 mm (36 in)
    • Steady state conditions are assumed.
    • Pressure at delivery point and suction at each compressor station is 7 MPa  (1015 Psia)
    • This is a horizontal pipeline with no elevation change.
    • Overall Heat Transfer Coefficient: 1.42 W/m2-˚C  (0.25 Btu/hr-ft2-˚F).
    • Ambient temperature is 18.3˚C (65˚F).
    • Compressor polytropic efficiency is 75%.
    • Pressure drop in coolers 35 kPa (5 Psia)
    • Simulation software: ProMax and using Equation of State from Soave-Redlich-Kwong (SRK).

    Table 1

     Four cases of onshore transportation of this natural gas are considered and each is explained briefly below. The number of pipeline segments, segment length, and inlet pressure of each segment for the four cases are presented in Table 2 in the SI (System International) and field (FPS, foot, pound and second) sets of units.

    Hydraulics Simulation Results and Discussions:

    The four cases are simulated using ProMax [3] to determine the pressure and temperature profiles, the compression horsepower,  and the after cooler duties. Table 3 presents a summary of simulation results for the three cases in FPS and SI systems of units.

     

    Case A: High Pressure (Dense Phase)

    This pipeline is a single compressor station configuration. The pipeline inlet pressure is in the dense phase zone. After processing and passing through the first stage scrubber, the lean gas  pressure is raised from 4.24 to 9.363 MPa (615 to 1358 Psia), then cooled to 37.8 ˚C (100 ˚F). The gas is compressed further in the second stage to 20.684 MPa (3000 Psia). The high pressure compressed gas is cooled back to 37.8 ˚C (100 ˚F) and then passed through a separator before entering the long pipeline.

     

    Case B: Intermediate Pressure

    This pipeline has three compressor stations each equally spaced at 536 km (333 miles). The pipeline inlet pressure is near the dense phase zone.  In the first station, the pressure is raised from 4.24 to 12.8 MPa (615 to 1858 Psia) and in the subsequent two stations, the pressure is raised from 7 to 12.8 MPa (1015 to 1858 Psia) in one stage, then cooled to 37.8 ˚C (100 ˚F), and finally passed through a separator before entering each pipeline segment.

     

    Case C: Low Pressure

    This pipeline has five compressor stations equally spaced in 322 km (200-mile)  segments. In the first station, the pressure is raised from 4.24 to 10.9 MPa (615 to 1577 Psia) and in the subsequent four stations, the pressure is raised from 7 to 10.9 MPa (1015 to 1577 Psia) in one stage, then cooled to 37.8 ˚C (100 ˚F), and finally passed through a separator before entering each pipeline segment. The pipeline inlet pressure is well below that for dense phase.

    Table 2

    Case D: High Pressure

    This case is similar to case B except it operates in the dense phase and the outside diameter is 914 mm (36 inches). This pipeline has three compressor stations each equally spaced at 536 km (333 miles). The pipeline inlet pressure is in the dense phase zone.  After processing and passing through the first stage scrubber, the lean gas  pressure is raised from 4.24 to 8.67 MPa (615 to 1257 Psia), then cooled to 37.8 ˚C (100 ˚F). The gas is compressed further in the second stage to 17.72 MPa (2570 Psia). The high pressure compressed gas is cooled back to 37.8 ˚C (100 ˚F) and then passed through a separator before entering the long pipeline. In each subsequent station, the pressure is raised from 7 to 17.7 MPa (1015 to 2565 Psia) in one stage, then cooled to 37.8 ˚C (100 ˚F), and finally passed through a separator before entering each pipeline segment.

    As can be seen in Table 3, Case A with a single compressor station requires the least total compression power and lowest heat duty requirements. The power increase for Case B (with three compressor stations) is about 38%  compared to Case A and 54% and 89% for Cases C (with 5 compressor stations)  and D (with 3 compressor stations), respectively. These increases in power and heat duty requirements are significant.  Similarly, the heat duty increases are about 6, -1, and 59% for case B through D compared to case A, respectively.

     Table 3

     

    Variation of gas  pressures is shown on Figure 1 for Cases A and B. As discussed in the previous TOTM, when the phase diagram and the pressure profiles are cross plotted using the pressure and temperature profiles the pipeline outlet condition remains  to the right of the dew point curve with the gas remaining as single phase.

     

    Mechanical Design (Wall Thickness and Grade)  

    Pipeline wall thickness is an important economic factor. Pipeline materials have historically represented approximately 40% of the Capital Expense (CAPEX) of a pipeline. Construction has historically  accounted for approximately 40% of the CAPEX as well. The estimation of the CAPEX is developed later in this TOTM. Once the wall thickness is determined, then the total weight (tonnage) of the pipeline can be calculated as well as costs for the pipeline steel.

     

    The wall thickness, t, for the three cases is calculated from a variation of the Barlow Equation found in the ASME B31.8 Standard for Gas Transmission Pipelines:

    Formula 1

     

     

    Where,

    • P is maximum allowable operating  pressure, here set to 1.1 times the inlet pressure,
    • OD is outside diameter,
    • E is joint efficiency (assumed to be 1) since the pipeline will be joined with through thickness butt welds and 100% inspected,
    • F is design factor,(ranges from 0.4 to 0.72) and here set  to be 0.72 (for remote area),
    • T is the temperature derating factor and is also 1.0 with the inlet temperatures no more than 37.8 ˚C (100 ˚F).
    • σ is the pipe material yield stress (Grade X70 = 70,000 psi or 448.2 MPa), and
    • CA is the corrosion allowance (assumed to be 0 in or 0 mm, for this dry gas).

    After calculating the wall thickness, the diameter to wall thickness ratio (D/t) is checked against these rules of thumb:

    • Onshore pipelines will have a maximum D/t of 72.

    If the D/t calculated is too high, the wall thickness will be increased to yield the maximum allowed D/t.

    Figure 1

     Using the calculated pipeline inlet pressures from the hydraulics as the starting point, the MAOP, and then the wall thickness can be calculated. The calculated wall thickness is then checked against the maximum D/t criteria. Table 4 summarizes these calculations for the four cases of onshore locations.

    Knowing the wall thickness and diameter allows the weight per lineal length (foot or meter) to be calculated. The total weight of the steel for the 1609 km (1000 mile) long can then be calculated as well. The unit weight is given in kg/m (lbm/ft) and the total weight in metric tonnes (1000 kg) and short tons (2000 pounds). The results of these weight calculations are in Table 5.

     Table 4&5

    Some observations from these calculations are:

    • Decreasing the pipeline diameter from 42 inch to 36 inch does NOT dramatically reduce the total steel tonnage. This is due to the increased pressures needed to flow the same volume of gas in the smaller diameter, hence increasing the wall thickness.
    • Increasing the steel grade (SMYS – Specified Minimum Yield Stress) from X-70 to X-80 would decrease the steel tonnage approximately 14%. As the cost calculations will show, this reduction would lower the cost significantly.
    • The volume of steel combined with the diameter and wall thicknesses will require a major portion of pipe manufacturing capacity. If this were a sanctioned project, pipe steel procurement would need to bid well in advance of the planned construction.
    • Wall thicknesses are NOT raised to next standard API thicknesses. The large quantity of steel needed allows the buyer to dictate a non-standard thickness. The pipe mills will be glad to accommodate such a requirement.

     

    Estimated Capital Costs

     The capital costs (CAPEX) for these estimates are based on two key variables: pipeline wall thickness and the compression power required. Both are dependent of the pipeline pressure profile which is dictated by the number of compressor stations. The estimated cost will be calculated from the following assumptions:

    • Line pipe is priced at US$ 1200 per short ton with a 15% adder for coatings.
    • Pipeline total installed cost is 2.5 times the pipe steel plus coatings cost. This factor  has been surprisingly consistent historically for both onshore and offshore long distance and larger diameter pipelines. Project specific factors such as mountainous terrain for onshore pipelines  can impact this cost multiplier.
    • No additional cost difference is taken into account for this estimate many of the real conditions that are dealt with for the  onshore design  construction. In reality there is a difference that can be significant. These differences are largely dependent on the project location with factors that could include weather and seasonal challenges, terrain for onshore projects, available infrastructure and its impact on logistics, and availability of construction equipment and labor.
    • Compressors and associated equipment (drivers, coolers, and ancillaries) are priced at US$ 1500 per demand horsepower.
    • Onshore compressor stations are priced at US$ 25 million each for site works, buildings and equipment not directly related to gas compression.

     

    With these cost assumptions, an order of magnitude estimate (OME) for the total installed cost (TIC) is developed for the pipeline, then the compressor stations, and finally combined for the total ONSHORE pipeline system in Table 6 – Pipeline Estimate, Table 7 – Compressor Station Estimate, and Table 8 – Total System OME.

     

     Table 6

     

    Table 7&8

    The results are indicative of finding a set of operating pressures, pipe diameter and number of compressor stations that show relatively little change with different combinations of the key parameters (Cases B, C and D). The selection of the “optimum” system configuration will involve more engineering definition, consideration of construction challenges, and evaluation of other parameters such as the operating costs (OPEX), environmental and permitting challenges, and more depth in evaluating the construction plan and costs.

    The total installed costs for this  ONSHORE system declines with decreasing operating pressure (MAOP), although the rate of decline is also decreasing, as more compressor stations are needed. For the onshore systems, the operating cost, particularly fuel costs, may be one of the key deciding parameters for the operating pressure / number of compressor stations decision. It is common for total life cycle costs (OPEX plus CAPEX) to begin rising at some point as the number of compressor stations and total horsepower increases with decreasing operating pressure.

     

    Often, with the operating costs included the “optimum” configuration favors higher operating pressures and fewer compressor stations. The cost adjustments for project location on both CAPEX and OPEX can move to “optimum” configuration either way.

     

     Final Comments:

     We have studied transportation of natural gas in the dense phase region (high pressure) and compared the results with the cases of transporting the same gas using intermediate and low pressures. Our study highlights the following features:

    1. There may be several system configurations (pipe diameter, operating pressures, and number of compressor stations) that show relatively small variation.
    2. As the MAOP increases, the required power and associated cooling duty can significantly decrease.
    3. The decreased costs for compression are offset by increasing pipeline costs. The key is by how much.
    4. Project location can have significant impact on the costs, hence the key decisions are on operating pressures, and the number and power levels at the compressor stations.
    5. With the high power demands of large diameter – high capacity pipelines, the operating costs for fuel can be a key factor in the configuration selection. If the gas at the source is not at high enough pressure, considerable compression power and cooling duty may be required if the decision is to use the dense phase.

     

    In future Tips of the Month, we will consider offshore transportation of natural gas as well as the effect of project location and operating costs on the life cycle costs and the configuration selection.

     

    To learn more, we suggest attending our G40 (Process/Facility Fundamentals), G4 (Gas Conditioning and Processing), G5 (Gas Conditioning and Processing-Special), PF81 (CO2 Surface Facilities), PF4 (Oil Production and Processing Facilities), and PL4 (Fundamentals of Onshore and Offshore Pipeline Systems) courses.

     

    John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

    By: Mahmood Moshfeghian and David Hairston

    References:

    1. Beaubouef, B., “Nord stream completes the world’s longest subsea pipeline,” Offshore, P30, December 2011.
    2. http://www.jmcampbell.com/tip-of-the-month/
    3. ProMax 3.2, Bryan Research and Engineering, Inc., Bryan, Texas, 2012.

  • What is Mentoring?

    What is Mentoring?

    In this Tip of the Month, we explore how process safety competency can be enhanced through mentoring programs.

    This TOTM is the paper that was developed by JMC Instructor/Consultants Clyde Young and Keith Hodges presentation at the Center for Chemical Process Safety (CCPS) 8th Global Conference on Process Safety in April, 2012.  The paper will also be published in the AIChE (American Institute of Chemical Engineering) publication, “Process Safety Progress.”

    Commit to Process Safety is the first pillar mentioned in “Guidelines for Risk Based Process Safety Management”, published by CCPS.  This pillar is supported by five elements.  One of the elements is Process Safety Competency, which is associated with efforts to maintain, improve and broaden knowledge and expertise.

    In Greek mythology, Odysseus, King of Ithaca went to fight in the Trojan Wars. Before he left, he entrusted his son Telemachus to the care of his old and trusted friend MENTOR. It was some ten years before father and son were reunited and during this time the development and care of his son was with Mentor.

    What is often missing from historical accounts is that it is Athene, the Goddess of Wisdom, who appears to Telemachus in the likeness of Mentor and gives advice, encouragement and spiritual insight.

    Since then, the word Mentor has become synonymous with trusted advisor, friend and teacher, a wise person.

    Demographic studies of the oil and gas processing industry indicate that a large number of people are retiring and being replaced by younger, less experienced personnel.  This presents a challenge to the industry.  A wise mountaineer once stated, “Good judgment comes from bad experiences.” With the influx of less experienced personnel, it would be shameful to have their good judgment developed from their bad experiences.  Especially since these bad experiences can be catastrophic.

    Organizations in the industry have spent considerable resources recruiting the best talent available and most have a competency development program that these new workers enter.  The program will generally include a step to have a more experienced person provide feedback on the worker to assess competency in the job. Well-developed and resourced competency development programs will have a Mentor assigned to the worker.

    What does this really mean and how can an organization insure that process safety competency is developed in all personnel, even if process safety activities are not the primary role?

    This TOTM will provide some guidance and best practices for establishing Mentoring programs with an emphasis on developing process safety competency in the younger, less experienced workforce.

    The role of Mentor involves teaching, helping, protecting, challenging, motivating, guiding, coaching, listening, and providing career guidance; it falls short of counseling.  Counseling is the provision of professional psychological help and advice and chosen Mentors would be foolhardy to attempt such a role without extensive training.

    Mentoring is usually a formal or informal relationship between two people, a Mentor (usually and preferably outside the Mentee’s area of supervision) and a Mentee.  The Mentor can also be provided from an external organization. This can be preferable especially if there is any hint of competition between the Mentors and Mentees (e.g. working in the same department as peers).  There are different rules of engagement if the external option is taken and this is outside the scope of this paper.  Peer Mentoring can be a useful option, especially if a peer Mentor has specific skills and qualifications.

    Using a Mentee’s supervisor within a discipline should be avoided as there could be a conflict of interest.  The Mentor may be Mentoring one day and disciplining the next, This is not conducive to building trust, which is an important ingredient in the Mentoring process.

    Mentoring should not be substituted for conventional classroom or computer-moderated training. It enhances traditional training by allowing the Mentee to learn from experienced colleagues within the working environment.

    Choosing a Mentor

    The choice of Mentors is an important aspect of a program and managers should first be satisfied that a Mentor not only has the required technical skills, but also has the ability to convey those to others in an efficient and effective way. Competency associated with Mentoring skills does not necessarily come naturally to everyone with highly competent technical skills.  A key skill to insure effective process safety is communication with all disciplines that could have an impact on the process.

    Mentor Program

    It is foolhardy to think that just putting together a pool of people as Mentors and pairing them with Mentees is going to be an effective way to put a Mentor program together.  It takes planning and needs structure.  There has to be an organizational aim for the program with measurable objectives.  The Mentor should be provided with these and a list of roles and responsibilities, which they should fully comprehend.

    There should be a selection process for Mentors and organizations must recognize that a training program may have to be created for selected Mentors.

    Ideally the Mentee should be able to select the Mentor from a pool of people in the organization; management, the training department or HR should not pair them.  Mentors should have the option to refuse the role should they feel that it would not be appropriate.

    Mentoring and Process Safety

    A Mentoring program is not to be approached in a haphazard fashion if the goal is to develop competent personnel.  A Mentoring program is much like a process safety management system.  The Center for Chemical Process Safety (CCPS) guidelines for Risk Based Process Safety Management (RBPSM) defines a management system as, “A formally established and documented set of activities designed to produce specific results in a consistent manner on a sustainable basis.”  The Mentoring program should be formalized, documented and designed to produce specific results.  The specific results are competent personnel associated with process safety.

    Mentees within a program may have been chosen because they are targeted to fill a key role within the organization.  This role could be a technical position that requires narrow skills in a field or a supervisory position of either engineering personnel or operations personnel.  The competency levels associated with process safety that are required will be highly dependent on the role in the organization.  The Mentor/Mentee relationship should keep this in mind as the process progresses.

    An effective Mentoring program that includes process safety as a key component will yield numerous benefits to the organization.  A Mentor with wide professional and technical expertise should have considerable experience in areas that involve process safety.  A Mentor that truly understands the concepts of risk based process safety will be invaluable to a Mentee with less experience.  Consider the pillars of RBPSM and some of the elements within each pillar.

    Commit to Process Safety

    Elements of this pillar include:

    • Process safety culture
    • Compliance with standards
    • Process safety competency
    • Workforce involvement
    • Stakeholder outreach

    A simple definition of culture is, “How we do things around here.”  Organizations strive to develop a learning culture that seeks hazards and solutions on a continuous basis.  It is imperative that Mentees are provided awareness level training on the organization’s culture and the Mentor will be given training on how to act as the example.  Two significant benefits will come from this.  The Mentors will examine their own actions within the culture and insure that they are setting a good example.  The Mentee will question why and how activities are accomplished and learn his/her role within the organization’s culture, which should accelerate the Mentees contribution through self-awareness.

    It will be difficult for a less experienced worker to learn the things required to insure compliance with all applicable standards.  An effective Mentor should always guide the Mentee toward the correct answer associated with compliance but not necessarily answer the question of compliance.  The guidance and allowing the Mentee to find the answer will insure that the learning associated with compliance will be retained long after the answer is discovered.

    Process safety competency of the Mentee will be enhanced significantly, but only if the Mentor insures that the Mentee is directed to the appropriate resources for this.  The Mentor does not necessarily have to be considered a process safety expert.  The Mentor does have to be aware that some process safety issues require a level of expertise that will be found elsewhere.  And sometimes those resources may be outside the organization.

    For a process safety management system to thrive, staff members at all levels of the organization must take an active role.  The role taken needs to be identified and metrics established to show participation in the role.  A Mentor can provide guidance and suggestions so that the Mentee is consistently working toward the goals of the process safety management system. Appropriately timed reviews of progress associated with established process safety metrics should be scheduled and conducted.

    Stakeholders include outside contractors, shareholders, community members and partners in joint ventures.  A Mentee may be involved with negotiations and planning activities associated with all kinds of stakeholders.  A Mentor’s experience in the industry and the organization can be very useful to insure that all stakeholder interests are addressed.

    Understand Hazards and Risks

    Elements of this pillar include:

    • Process knowledge management
    • Hazard identification and risk analysis

    Development of a Mentee’s competency in this pillar of RBPSM could be the Mentor’s most important role. Insuring that the correct process knowledge is developed and managed appropriately is a critical activity that the Mentee strives for. There is no need for a Mentee to learn from mistakes if a Mentor can provide clear guidance on this pillar.

    It is within these two elements that mistakes can lead to catastrophic events.  Having an incorrectly sized relief valve installed in a process or not anticipating the consequences of failure of controls is not acceptable. The Mentor and Mentee should routinely conduct discussions about these elements.

    Contract services are utilized a great deal for design of new and modified facilities.  A Mentor who has significant experience in this area can provide the Mentee advice and guidance for overseeing these projects.  Oversight by a qualified company representative will insure that all issues associated with a project have been addressed.

    Providing resources during the conduct of Process Hazard Analysis (PHA) studies is a challenge for many organizations. This is especially true considering the demographics of the industry at this time. More experienced personnel have moved on. PHA team members with significant experience are critical to the success of a PHA.  A Mentee who is assigned to a PHA team may or may not work side by side with their Mentor.  If the assigned Mentor is also a member of the PHA team, this may prove advantageous.  As the role of Mentor is to provide guidance and direction to new and developing staff, the PHA is an excellent environment to do just that.  The structure of the PHA provides an opportunity to guide the Mentee in the proper way to identify hazards, develop measures to mitigate those hazards and work as a team member in a formalized setting.

    Manage Risk

    Within this pillar, a Mentee will benefit from the guidance of an experienced Mentor to become proficient at what might be considered the day-to-day activities associated with their job.  Elements are:

    • Procedures
    • Safe work practices
    • Asset integrity
    • Contractors
    • Training and performance
    • Management of change
    • Operational readiness
    • Conduct of operations
    • Emergency management

    Sometimes organizations will assign a younger, less experienced person to a supervisory position in operations to “season” them. Studies have shown that a great number of incidents occur during normal operations.  Having a Mentor with significant operations experience will accelerate the “seasoning” process and insure that the problems associated with day-to-day activities do not lead to a catastrophic incident.

    Working in operations supervision will certainly expose the Mentee to many issues associated with personal interaction. Dealing with people may be one of the most difficult tasks undertaken. Having the ear of a Mentor can be helpful as the Mentee develops his/her skills in this area.

    Learn From Experience

    There is no reason that a young professional cannot learn from the experience of others. To pass along the experience and knowledge that has been gained over the years is the focus of a Mentoring program.   Hopefully, the Mentee will not have to experience a catastrophic incident to learn from experience.

    Elements within this pillar are:

    • Incident investigation
    • Measurement
    • Audits
    • Management review and continuous improvement

    Having a Mentor available to help review near miss reports, incident investigations, audit findings and metrics associated with process safety can provide the Mentee with a “cold eye” review of issues that are the Mentee’s responsibility to address.  Often a wiser, more experienced Mentor will have experienced some of the same things that are being discovered under the Mentee’s watch.  In this case, issues should be able to be addressed quickly and more efficiently.

    Troubleshooting

    All processes within the industries we work have been designed to operate in a specified manner. This manner includes specific temperatures, pressures, flow rates and levels.  Defining these specific parameters establishes “normal” for these processes.  Operating processes in a “normal” manner reduces the likelihood of a catastrophic incident.  Deviation from “normal” is not acceptable and identifying this deviation and taking the steps required to return to normal requires experience and knowledge. This is known as troubleshooting. Process safety management is a system that establishes “normal” and provides directions on maintaining “normal”. Personnel with effective troubleshooting skills will also work efficiently within an organization’s process safety management system.

    A formalized, well established Mentoring program for younger, less experienced personnel entering the business enhances everyone’s troubleshooting skills.  The Mentee has someone (the Mentor) available to query about issues seen and the Mentor is challenged to insure the advice and guidance provided is correct and useful.

    Attaining high-level competency in a job requires training and then performing the job for a period of time.  Accelerating the path to high-level competency is a significant goal of a formalized Mentoring program.

    Conclusion

    At the beginning of this TOTM, it was stated that the word Mentor has become synonymous with trusted advisor, friend and teacher, a wise person. Process safety management has become synonymous for reducing the risk associated with the activities performed in our industries.

    Risk is often viewed differently from individual to individual.  A person’s perception of risk may change with familiarity.  Having a trusted advisor for younger, less experienced personnel, to help identify and provide suggestions for mitigation of hazards, in all their forms, is a strong competency development tool for any organization.  Personnel will be developed quicker and more efficiently. Experienced personnel are one of a company’s most valuable resources.  Acting as a Mentor can be the best use of this resource and will provide a challenge that some people thrive on.

    Any organization that truly strives for a generative safety culture should do whatever it takes to implement a process safety-Mentoring program. The benefits will be seen and reaped for years to come.

    To learn more about managing process safety systems, we suggest attending our PetroSkills HSE course,  HS 45- Risk Based Process Safety Management.

    To enhance process safety engineering skills we suggest any of the JMC foundation courses or our, PS 4 – Process Safety Engineering course.

    John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

    By: Clyde Young and Keith Hodges

     

  • Low Pressure Vs High Pressure Dense Phase Natural Gas Pipeline Transportation

    Capital Cost (CAPEX) Comparisons

    High pressure (or dense phase) is increasingly used for transporting large volumes of carbon dioxide (CO2) and natural gas over long distances. In this month’s – Tip of the Month (TOTM), we continue to explore key aspects of dense phase transportation in pipelines. This month’s focus is on the estimation of capital costs as a tool to compare then select the operating pressures and associated facilities for a long distance – high volume flow gas transmission pipeline.

    In recent TOTMs (January through April 2012 and again in August and September 2012), we discussed several aspects of the physical behavior and transportation of carbon dioxide (CO2) and natural gas in the dense phase. We illustrated how thermophysical properties change in the dense phase and their impacts on pressure drop calculations. The pressure drop calculation utilizing the liquid phase and vapor phase equations were compared.

    In the August 2012  (TOTM), we studied transportation of rich natural gas in the dense phase region and compared the results with the case of transporting the same gas using a two phase (gas-liquid) option. Our study highlighted the pros and cons of dense phase transportation.

    In September 2012 (TOTM), we analyzed pipeline transportation of a lean natural gas at a wide range of operating pressures from the relatively low pressure typical in many gas transmission pipelines to much higher pressures well into the dense phase region.

    Case Study:

    We will continue to use the same case study basis as used in the September 2012 TOTM. The gas composition and conditions are presented in Table 1. For simplicity, the calculations and subsequent discussion will be done on the dry basis. The feed gas dew point was reduced to -40 ˚F (-40 ˚C) by passing it through a mechanical refrigeration dew point control plant. The resulting composition and conditions of the lean gas are also presented in Table 1. The lean gas has a Gross Heating Value of 1082 BTU/scf (40.33 MJ/Sm3), which is in the range typically seen for contract quality natural gas in North America. The pipeline parameters are:

    • Length is 1000 miles (1609 km) long
    • Pipeline outside diameter is 42 inches (1067 mm). Initial inside diameters for the hydraulics analyses are: Case A = 39.0 (991 mm) inches, Case B = 40.0 inches (1016 mm), and Case C = 40.5 inches (1029 mm)
    • Steady state conditions are assumed.
    • Pressure at delivery point and suction at each compressor station is 615 Psia (4.24 MPa)
    • This is a horizontal pipeline with no elevation change.
    • Overall Heat Transfer Coefficient: 0.25 Btu/hr-ft2-˚F (1.42 W/m2-˚C).
    • Simulation software: ProMax and using Equation of State from Soave-Redlich-Kwong (SRK).

    Table 1. Composition and conditions of the feed gas and lean gas

    Table 2. Pipeline specifications for the three casesThree cases for transportation of this natural gas are considered and each is explained briefly below.  The number of pipeline segments, segment length, and inlet pressure of each segment for the three cases are presented in Table 2 in the field (FPS, foot, pound and second) and SI (System International) sets of units.

    Table 2. Pipeline specifications for the three cases


    Hydraulics Simulation Results and Discussions:

    The three cases are simulated using ProMax [3] to determine the pressure and temperature profiles, the compression horsepower, and the after cooler duties. Table 3 presents a summary of simulation results for the three cases in FPS and SI systems of units.

    Case A: High Pressure (Dense Phase)

    This pipeline is a single compressor station configuration. The pipeline inlet pressure is in the dense phase zone. After processing and passing through the first stage scrubber, the lean gas  pressure is raised to 1496 psia (10.32 MPa), then cooled to 100 ˚F (37.8 ˚C). The gas is compressed further in the second stage to 3659 Psia (25.22 MPa). The high pressure compressed gas is cooled back to 100 ˚F (37.8 ˚C) and then passed through a separator before entering the long pipeline.

    Case B: Intermediate Pressure

    This pipeline has three compressor stations each equally spaced at 333 miles. The pipeline inlet pressure is near the dense phase zone.  In each station, the pressure is raised from 615 Psia to 2071 Psia (4.24 to 14.28 MPa) in one stage, then cooled to 100 ˚F (37.8 ˚C), and finally passed through a separator before entering each pipeline segment.

    Case C: Low Pressure

    This pipeline has five compressor stations equally spaced in 200-mile (322 km)  segments. The pipeline inlet pressure is well below that for dense phase. In each station, the pressure is raised from 615 Psia to 1637 Psia (4.24 to 11.28 MPa) in one stage, then cooled to 100 ˚F (37.8 ˚C), and finally passed through a separator before entering each pipeline segment.

    Table 3. Summary of computer simulation results for the three cases.

    As can be seen in this table, Case A with a single compressor station requires the least total compression power and lowest heat duty requirements. The power reduction for Case A is about 51%  compared to Case B (with three compressor stations) and 63% compared to Case C (with 5 compressor stations). These reductions in power and heat duty requirements are significant.  Similarly, the heat duty reduction for Case A is about 39% compared to Case B and 50 % compare to Case C, respectively.

    Variation of gas velocity, pressures, and temperature are shown on Figures 1 through 3 for Cases A and B. As discussed in the previous TOTM, when the phase diagram and the pressure profiles are cross plotted using the pressure and temperature profiles the pipeline outlet condition remains  to the right of the dew point curve with the gas remaining as single phase.

    Figure 1. Variation of gas velocity in the pipeline (Cases A and B)

    Mechanical Design (Wall Thickness and Grade)

    Pipeline wall thickness is an important economic factor. Pipeline materials typically represent approximately 40% of the Capital Expense (CAPEX) of a pipeline. Construction will account for approximately 40% of the CAPEX as well. The estimation of the CAPEX is developed later in this TOTM. Once the wall thickness is determined, then the total weight (tonnage) of the pipeline can be calculated as well as costs for the pipeline steel.

    The wall thickness, t, for the three cases is calculated from a variation of the Barlow Equation found in the ASME B31.8 Standard for Gas Transmission Pipelines:

                                                                                                                         (1)

    Where,

    • P is maximum allowable operating  pressure, here set to 1.05 times the inlet pressure,
    • OD is outside diameter,
    • E is joint efficiency (assumed to be 1) since the pipeline will be joined with through thickness butt welds and 100% inspected,
    • F is design factor,(ranges from 0.4 to 0.72) and here set  to be 0.72 (for remote area),
    • T is the temperature derating factor and is also 1.0 with the inlet temperatures no more than 100 ˚F (37.8 ˚C).
    • σ is the pipe material yield stress (Grade X70 = 70,000 psi or 448.2 MPa), and
    • CA is the corrosion allowance (assumed to be 0 in or 0 mm, for this dry gas).

    After calculating the wall thickness, the diameter to wall thickness ratio (D/t) is checked against these rules of thumb:

    • Onshore pipelines will have a maximum D/t of 72.
    • Offshore pipelines will have a maximum D/t of 42.

    If the D/t calculated is too high, the wall thickness will be increased to yield the maximum allowed D/t.

    Figure 2. Variation of pressure in the pipeline (Cases A and B)

       

    Figure 3. Variation of temperature in the pipeline (Cases A and B)

    Using the calculated pipeline inlet pressures from the hydraulics as the starting point, the MAOP then the wall thickness can be calculated. The calculated wall thickness is then checked against the maximum D/t criteria. Table 4 summarizes these calculations for the three cases for both onshore and offshore locations.

    Knowing the wall thickness and diameter allows the weight per lineal length (foot or meter) to be calculated. The total weight of the steel for the 1000 mile (1609 km) long can then be calculated as well. The unit weight is given in lbm/ft (kg/m) and the total weight in short tons (2000 pounds) and metric tonnes (1000 kg). The results of these weight calculations are in Table 5.

    Some observations from these calculations that can be made are:

    • Increasing the steel grade (SMYS – Specified Minimum Yield Stress) from X-70 to X-80 would decrease the steel tonnage approximately 14%. As the cost calculations will show, this reduction would lower the cost significantly. However, the use of X-80 steels is still not widely accepted in the pipeline industry.
    • The volume of steel combined with the diameter and wall thicknesses will require a major portion of pipe manufacturing capacity. If this were a sanctioned project, pipe steel procurement would need to bid well in advance of the planned construction.
    • Wall thicknesses are NOT raised to next standard API thicknesses. The large quantity of steel needed allows the buyer to dictate a non-standard thickness. The pipe mills will be glad to accommodate such a requirement.

    Table 4: Pressures and Wall Thickness Selections

    Table 5: Pipeline Wall Thickness Selections and Total Steel Weight

    Estimated Capital Costs

    The capital costs (CAPEX) for these estimates are based on two key variables: pipeline wall thickness and the compression power required. Both are dependent of the pipeline pressure profile which is dictated by the number of compressor stations. The estimated cost will be calculated from the following assumptions:

    • Line pipe is priced at US$ 1200 per short ton with a 15% adder for coatings.
    • Pipeline total installed cost is 2.5 times the pipe steel plus coatings cost. This factor is surprising consistent for both onshore and offshore long distance and larger diameter pipelines. Project specific factors such as mountainous terrain for an onshore pipelines, or the requirement to trench an offshore pipeline can impact this cost multiplier.
    • No additional cost difference is taken into account for this estimate between onshore and offshore construction. In reality there is a difference that can be significant. These differences are largely dependent on the project location with factors that could include weather and seasonal challenges, water depth for offshore projects, terrain for onshore projects, available infrastructure and its impact on logistics, and availability of construction equipment and labor.
    • Compressors and associated equipment (drivers, coolers, and ancillaries) are priced at US$ 1500 per demand horsepower.
    • Onshore compressor stations are priced at US$ 25 million each for site works, buildings and equipment not directly related to gas compression.
    • Offshore compressor stations are priced at US$250 million each for the fixed structure, topsides not directly related to gas compression, and a quarters complex. This assumption is sensitive to project location, whether the structure is stand-alone or in a group of structures, water depth, and met-ocean conditions.
    • The offshore pipeline cases originate ONSHORE with the lead compressor station.

    With these cost assumptions, an order of magnitude estimate (OME) for the total installed cost (TIC) is developed for the pipeline, then the compressor stations, and finally combined for the total pipeline system in Table 6 – Pipeline Estimate, Table 7 – Compressor Station Estimate, and Table 8 – Total System OME.

    Table 6: Pipeline Total Installed Cost

    Our estimating assumptions can lead to costs that are the same whether for onshore or offshore pipelines. This is where knowledge of the project becomes vital in adjusting the estimate to account for conditions that can affect the assumptions.

    Table 7: Compressor Stations Total Installed Cost

    The most sensitive variable for the compressor stations calculations is the location of any offshore facilities. Location, water depth and met-ocean conditions can and will impact the estimated cost significantly.

    Table 8: Total System OME

    The total installed costs for an ONSHORE system decline with decreasing operating pressure (MAOP), although the rate of decline is also decreasing as more compressor stations are needed. For the onshore systems, the operating cost, particularly fuel costs, can impact the operating pressure / number of compressor stations decision. It is common for total life cycle costs (OPEX plus CAPEX) to begin rising at some point as the number of compressor stations and total horsepower increases with decreasing operating pressure.

    For an OFFSHORE system, show the lowest total installed cost is with a three compressor station configuration. This “optimum” CAPEX solution will be sensitive to project location as discussed above as well as operating costs. Often, with the operating costs included the “optimum” configuration favors higher operating pressures and fewer compressor stations. The cost adjustments for project location on both CAPEX and OPEX can move to “optimum” configuration either way.

    Final Comments:

    We have studied transportation of natural gas in the dense phase region (high pressure) and compared the results with the cases of transporting the same gas using intermediate and low pressures. Our study highlights the following features:

    1. As the MAOP increases, the required power and associated cooling duty can significantly increase.
    2. The decreased costs for compression are offset by increasing pipeline costs. The key is by how much.
    3. Project location can have significant impact on the costs, hence the key decisions are on operating pressures, onshore versus offshore routing (where possible), and the number and power levels at the compressor stations.
    4. With the high power demands of large diameter – high capacity pipelines, the operating costs for fuel can be a key factor in the configuration selection. If the gas at the source is not at high enough pressure, considerable compression power and cooling duty may be required if the decision is to use the dense phase.

    In future Tip of the Months, we will consider the effect of project location and operating costs on the life cycle costs and the configuration selection.

    To learn more, we suggest attending our G40 (Process/Facility Fundamentals), G4 (Gas Conditioning and Processing), G5 (Gas Conditioning and Processing-Special), PF81 (CO2 Surface Facilities), PF4 (Oil Production and Processing Facilities), and PL 4 (Fundamentals of Onshore and Offshore Pipeline Systems) courses. 

    John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com. 

    By: David Hairston and Mahmood Moshfeghian

    References:

    1. Beaubouef, B., “Nord stream completes the world’s longest subsea pipeline,” Offshore, P30, December 2011.
    2. http://www.jmcampbell.com/tip-of-the-month/
    3. ProMax 3.2, Bryan Research and Engineering, Inc., Bryan, Texas, 2012.

     

  • Should unplanned maintenance jobs be recorded as near misses?

    OSHA mentions “near-misses” as recordable requirements in several passages as: “An unplanned and unforeseeable event that could have resulted, but did not result, in human injury, damage to property or the environment or other form of loss”.  And we know that all industrial maintenance organizations have a history of reactive, run-to-failure-then-run-to fix, maintenance management behaviors.  JMC’s emphasis on process and equipment reliability and operations management helps to bring facilities out of the reactive mode, but reactive maintenance jobs are still all too prevalent.  Some, or many of these reactive jobs “could have resulted, but did not result, in human injury, damage to property or the environment or other form of loss”.

    Safety is everyone’s number one goal.  Most corporate safety programs define near-misses, but few connect the dots between recordable incidents and the degree of reactive, unplanned maintenance work.  The famous safety pyramid is quite familiar, but that’s only the tip of the safety iceberg.  Below the ‘water line’ of recordable first aids (lagging indicators) lie near-misses and, at the base of it all, safe behavior.  These are the leading indicators of our safety performance. This Tip of the Month will tie reactive maintenance and safe behavior together.

    Recent data, compiled by Belgian’s BEMAS, clearly links accidents with injuries to the percent of reactive maintenance work (the opposite of planned and scheduled work).  If, indeed, a company uses the near-miss definition, how can it not require the recording of some, if not all, unplanned maintenance jobs?

    An iceberg is a good metaphor for Safety; most of its mass lies beneath the surface and we see only the tip.  The safety pyramid compares the quantity of accidents in layers with fatality on the top and reportable incidents on the bottom.  But the real basis of safe behavior lies underneath the reporting surface and is comprised of near misses and unsafe behaviors.

    For over 33 years I have been focused on driving down unplanned maintenance jobs through training and consulting on control of work.  We all should know that planned maintenance is simply safer!  But we have been reluctant to tie urgent, reactive jobs to unsafe practices.  In 2012, it’s time to ask “Should unplanned maintenance jobs be recorded as near-misses”?

    In a nearly parallel development path, our emphasis and understanding of safe work environments has also been refined.  With the help of several catastrophic events, like the Texas


    City refinery and, more recently Deepwater Horizons, and many smaller injury-causing accidents, our industry has put safety on the front burner.

    A useful way to look at the safety pyramid (Figure 1) is to draw the dividing plane at what is reported and not reported.  This brings behavior-based safety programs, which we all talk about, into perspective. A key point here is the separation of leading indicators and lagging indicators.  It’s obvious to record Incidents and Accidents after they happen, but less obvious to capture near-misses and instill safe behaviors.

    Recent data (Figure 2), presented by Wim Vancauwenberghe [1] of the Belgian Maintenance Association (BEMAS) at last year’s SMRP (Society for Maintenance and Reliability Professionals) annual conference shows the impact of unplanned maintenance jobs on the rate of accidents with injuries; and subsequent reduction in injuries as the percentage of planned work increases.

    This raises the question in this paper, and it’s time we asked.

    Clearly, not every unplanned maintenance job involves the same level of risk.  We can use a risk-based approach as in Figure 3 to indicate when an unplanned job becomes a near miss.  When we look at the spectrum of behavior from risk averse upwards to reckless, we can begin to establish some range of criteria for defining what to report. Applying the risk spectrum to the nature of unplanned jobs, we would expect risk to increase due to some factors.  Typical risk matrices compare event likelihood to its consequence to determine level of risk.  Should we develop something similar for unplanned jobs?  This tip attempts to describe the conditions that would determine the level of risk in jobs.  Perhaps there are companies who have successfully addressed this issue and, hopefully they will contribute to this discussion.


    OSHA distinguishes Accident, Incident and Near Miss with the definitions in Figure 4.  However, trying to define what ‘could’ have happened in every urgent job opens a Pandora’s Box that probably wouldn’t be very productive.  On the other hand, we could approach the near miss issue by defining ‘failure’ more carefully.  Taking the familiar P-F curve, we might be able to say that earlier definition of failure at the P point and the subsequent maintenance job would inherently be safer than reacting to a failure at the F point.  Figure 5 shows how that might work.  We could say that anytime we have an unexpected complete failure of equipment, it must be reported as a near miss, whereas, if we detect a potential failure and plan and schedule the maintenance action before complete functional failure, it wouldn’t need to be reported because it is not a near miss.

     

    If we are going to require near misses to be reported, another issue is raised:  How is a near miss to be reported?  What do we do with the report?  How much information/data is required on such a report?  If we’re going to require a report, we will have to define what and how much detail is required.

    There are several possible uses for a near miss report.  Whatever decision we take, will impact our staff with more information gathering tasks.  What is it worth?  How can we successfully use the report to lower near misses and effect safer behavior?  Or, drive down reactive maintenance work?

    • Used as a way to ‘speak up’ with the rest of the crew and raise their awareness would not require as much information about what happened,
    • Determine preventability of the near miss with root cause analysis would require a great deal more information.

    The fundamental questions are:

    • How do we raise the awareness of near misses with the target to reduce them?
    • What distinguishes a near miss from an incident?
    • If unplanned maintenance jobs carry higher safety risk, how do we break our reactive maintenance habits?
    • What criteria do we use to define levels of risk?

    In order to determine how the professional SMRP audience would distinguish the reportability of near misses, several situations were presented for the participants to vote using the following choices:

    1. Do it and report as a near miss
    2. Near miss, Speak up!
    3. Risky behavior, don’t tell anyone
    4. No risk, just do it!
    5. Do not proceed without a planned work order

    The sample situations were:

    • Urgent restart of a 100 hp motor after unexpected stoppage
    • Talking on your cell phone while driving
    • Vehicle crossing your path while running a yellow light
    • 5 lb. (2.27 kg) hammer dropped from scaffolding
    • Hurrying to replace hydraulic fitting without lock out, tag out
    • 2 ton lifting sling frayed, but go ahead and use it

    Results of this voting may be published in a subsequent TOTM, or send an email to the author, perry.lovelace@jmcampbell.com.

    In conclusion, we have raised the question and some of the issues around the question “Should unplanned maintenance jobs be recorded as near misses?”  There is not a simple answer and our profession must continue to explore the issues and make efforts to create a safer workplace through planned and scheduled maintenance work.

    To this end, JMC offers training related to reducing unexpected failures:

    • The Operations Management discipline is directly focused on reduction of unplanned events through better control of work,
    • Operator Training broadens facilities operators’ competencies by teaching how facilities work and why certain events happen,
    • Mechanical and Reliability disciplines help identify onset of equipment failures.  Reliable equipment is safer equipment,
    • Many facilities use contractors for maintenance; their safety is also important.  JMC’s Supply Chain and Procurement disciplines concentrate on better contractor relationships in our SC-41 course.

    To learn more about similar cases and how to minimize operational problems, we suggest attending our G40 (Process/Facility Fundamentals), G4 (Gas Conditioning and Processing), G5 (Gas Conditioning and Processing-Special), P81 (CO2 Surface Facilities), PF4 (Oil Production and Processing Facilities), and PL 4 (Fundamentals of Onshore and Offshore Pipeline Systems) courses.

     

    John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email your consulting needs to consulting@jmcampbell.com.

     

    By: Perry Lovelace, Sr. Staff Instructor

    References:

    1. Vancauwenberghe, Wim; The Basics of Safe Maintenance; The Belgian Maintenance Association; 2011.