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Tips on Heat Exchanger Process Design
Gives brief of how to decide process design parameters when simulating heat exchangers in design setting. Situations related to cooling water, hot oil / steam, feed / effluent exchangers are covered with real-world examples. Estimating number of shells in series for shell and tube exchangers with temperature cross is included. Tips on setting pressure drop for exchangers are also included.
Analysing Tank / Vessel Transfer
Usually a storage system (bullet for LPG / NGL / propane / Butane or tanks for oil / condensate) and export pumping system have a suction manifold that connects multiple bullets / tanks to multiple export pumps.
Switching from on-spec to off-spec mode or vice-versa can have flows higher than normal. Maintenance of tank / bullet can also have abnormal flow tranfers. Dynamics of such operational modes is studied with AFT Impulse to understand flows.
Pump Shut-off Head Rise
Pump curve is expected to rise continuously to shut-off head. For pump discharge equipment like filters, exchangers we need to specify design pressure. We estimate it before vendor confirms shut-off pressure later.
Initial estimate is based on maximum suction design pressure, normal differential head, maximum density and shut-off head rise (SOHR) of 1.25.
There are instances where shut-off head rise of 1.25 leads to an increase in shut-off pressure beyond piping # rating. I have seen different options tried to limit this.
One case study of produced water degasser transferring water to injection pipeline is attached.
Control Valve Pressure Drop Myths
Control valve pressure drop is one critical activity that is of equal interest to Dynamic simulation guys and detail engineering contractor.
My personal favourite benchmark of engineering understanding of my team is their understanding of control valve pressure drop.
You understand control valve pressure drop and you are good potential dynamic simulation engineer.
Typical minimum control valve pressure drop value suggested in guidelines is 0.7 bar. But this is a trap. Refer slides to understand more.
Petroleum Fluid Characterization
Most of the PVT characterization work is done by Reservoir engineers to estimate reserves / undertand flow characteristic of well etc. Articles by Whitson, Pedersen and many others are great learnings.
Flow assurance engineers, who are process engineers themselves, have tools like Multi-flash / PVTSim to use PVT data. Though I have used these tools I was not doing flow assurance on that assignment and started thinking of how a normal process engineer with access only to general understanding of simulation tool like Hysys / Unisim / Promax shall do it?
This ppt is for working process engineer with PVT report, general simulation tool and responsibility to simulate CPF. All the work is from open literature and references are given.
Understanding Double Block and Bleed
Application: Inlet line of LP Separator
Fluid handled: flashing hydrocarbon liquid with H2S in gas phase ~700 ppmv
Operating pressure = 3.5 barg
Operating temperature = 55 deg C
Line size = 10″
Line rating = 150#
How shall isolation of this line be. Refer image for options.
If vent is required, it would be connected to flare / drain. So do not eliminate options due to that factor. ESDV would be FC, TSO and any other spec that would be needed.
Relief System Design for Tube Rupture
Tube Rupture is one relief valve sizing scenario that has many misunderstandings. It starts from if it is credible at all if 10/13 is considered for exchanger to merely using steady state solution for transient event.
This supporting information from webinar is re-produced for better reach.
Centrifugal Compressor System Design
Deciding pressure ratio stages (process stages) is explained in many sources. What is less mainstream is deciding number of casing stages. Tried to touch this aspect though it is essentially a multidiscipline activity.
Process engineer does not (need not) mention polytropic head in datasheet nor discharge temperature nor work. His/her main role is specifying compressor suction and discharge pressures for which compressor vendor will estimate these quantities.
Process engineers’ confusion is where to start. Tried addressing it with an example (equations are just added to make the point).
Compressor Hot Gas Bypass Valve
dynamic simulation is the still the only sure shot method to decide on HGBV. But hardly 5% know dynamic simulation in an EPC firm.
And yet my Process engineer has to answer Projects team as to why HGBV is added so late in the project. He has to answer to Piping query if small change of volume in suction / discharge after dynamic simulation would need re-simulation.
Is there any way to decide if HGBV would be needed for his/her compressor well in advance without running dynamic simulation?
This post tries to answer that question. Data points are limited. Will try to add more compressor data points as I get them.
Pump NPSH for Dissolved Gases
Centrifugal pumps handling liquids with dissolved gases is a topic that occasionally torments process engineers. Pumps taking suction from blanketed tanks/vessels (hot oil, drain drum, flare KOD) or liquids that have light gases dissolved (amine, glycol, BFW, injection water) are main situations.
Few articles / approaches are available in literature; still the whole calculation is not mainstream as far as I know. This topic is frequently discussed in our Learning Process / Flarenet Whatsapp technical groups.
Below presentation tries to explore NPSH for liquids with dissolved gas and briefly mentions NPSH margin criteria. Calculations from past work are provided for illustration.
NPSHA: How Much is Enough?
We all know NPSHA has to be more than NPSHR by safety margin. NPSHA is estimated by us process engineers based on assumed suction vessel and pump center line (C/L) elevation. NPSHR is provided by pump vendor.
When you calculate NPSHA, you have no idea what pump vendor NPSHR is going to be. If pump vendor NPSHR is more than NPSHA, you will need to elevate suction vessel (frictional resistance is usually not much in pump suction due to low velocity).
How do we know if pump NPSHA we are giving in datasheet is adequate? Can we make a guess for NPSHR and be assured that we need not have to rework?
Attached presentation tries to answer this question. I had data of pump NPSHR and pump C/L. Since we all are talking data these days, thought to use it.
Process Engineering: Situations to Equations
Few years into the industry and most of your colleagues would tell you how college education is of no use in industry. They do not teach us in college anything that we use here.
This is far from true.
I will share few design situations wherein I felt: “Ohh! We could apply that concept from Chemical Engineering Course here”.
Deciding Flare MDMT
Minimum design metal temperature of High Pressure (HP) flare system is a decision that every oil and gas central processing facility (CPF) designer has to make. In absence of any low temperature system (if NGL is not recovered), flare MDMT is typically set by unintended depressurization/blowdown of the facility i.e. so called cold or adiabatic depressurization.
Depressurization is a dynamic event where process system pressure drops with time. Process system temperature also continues to drop as a result of near-isentropic expansion. Flare system also has its own dynamics, where pressure builds-up in the flare system. Temperature of flared gas thus varies with time due to changing P/T profile of process system as well as changing pressure profile in flare system and J-T expansion across BDV/RO.