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vnpetroleum,
Use the real gas k value as estimted by HYSYS. The ideal gas value is an approximation which was included in the API RP in the days before the real value was available in simulations, or, for those engineers who do not have access to a simulation package. The approximation is acknowledged in Appendix B of API RP 520.
The isentropic exponent k appears in the API capacity equation because k is part of the expression for pressure at sonic velocity ( in the throat of the PSV) as a function of upstream pressure, which is embedded in the API equation.
A good value for k can be obtained from k = ln (p1/p2)/ln (rho 1/rho 2), where subscript 2 refers to isentropic conditions at the throat (sonic velocity). I have checked this expression against HYSYS real k and they match.

Paul

vnpetroleum,

Specific heat at constant volume (C_v) values for most gases is not directly available from thermodyanamic property tables found in books and literature. What is available, is the specific heat at Constant pressure (C_p). k values based on the ideal gas equation are:

k = C_p / (C_p - (R/MW))  

where:

C_p = heat capacity at constant pressure, kJ/kg-K

R = universal gas constant, 8.314 (kJ / kmol-K)

MW = molecular weight, kg / kmol

'k' values based on ideal gas generally give lower values then 'k' values for real gases. When using it in sizing equations as per API 520 for relief valves for gas service a lower 'k' value provides a higher relief rate and relief area compared to a higher 'k' value.

What you get is a more conservative relief valve size with a 'k' value based on ideal gas. The trend today of doing optimum sizing and not conservative over design dictates the usage of the 'k' value based on real gas behaviour at elevated pressures and temperatures.

I still prefer to use the 'k' value based on the ideal gas behaviour and the relief valve sizing excel sheets I have used and also developed consider the 'k' value based on ideal gas behaviour.

However, I have posted a spreadsheet for calculating 'k' values based on 'real gas' behaviour on the forum. This spreadsheet is based on an empirical correlation and has its limitations in comparison to what a simulation software like HYSYS will calculate based on rigorous calculations using EOS like PR or SRK.

The link for the spreadsheet is:

http://www.cheresour...0185#entry30185

Hope this helps.

Regards,
Ankur.

While this may appear to be a simplistic question, it is actually very relevant to the way pressure drops are calculated for gases.  As Art Montemayor has pointed out, in the final analysis you have to use the actual conditions.  But if you ask 90% of the engineers involved in pipelines and reticulation systems for petroleum gases they will tell you exactly the opposite and they will insist that you have to use standard conditions.  My experience of working with these engineers is that they want to have nothing to do with mass flows or actual volumetric flows.

 

The reason for the gas engineers believing that they have to work in flows at standard conditions is that historically they used empirical formulas like Spitzglass, Weymouth and Panhandle - all of which are indeed based on flows in standard conditions.  But if you examine these equations in detail you will quickly see that you have to enter data like the actual temperature, the actual pressure, the actual compressibility factor and all the other data that is necessary for the equation to (internally) convert the flow in standard conditions to the flow at actual conditions.  (The standard conditions are also in there, but are all combined into one constant.)  So, while the gas engineers think they are working in standard conditions, their equations are doing all the conversions for them behind the scenes and they are in fact working in actual conditions.

 

Old habits die hard, and these equations are still used - even though there are better alternatives these days.  It seems that consistency is more important than accuracy.

 

You can prove to yourself that you have to work in actual conditions by examining the Darcy-Weisbach formula, i.e.

ΔP = ƒ(L/D )(ρv²/2)

 

This is applicable to all fluids - gases and liquids - provided that L is small enough that the density (ρ) and velocity (v) can be regarded as constant.  With gases, which are more compressible than liquids, the increments of L have to be taken much smaller than with liquids to meet this requirement.

 

The density (ρ) and velocity (v) in this equation have to be the actual values at the flowing conditions (this is basic physics for kinetic energy) and shows why I (and Art) have said that in the final analysis you are working with actual conditions.

sure,
I attach the Excel page psvcompare1.xls

this page permits to compare different procedures

Discharging temperature (PSV outlet)
1) calculated with isentropic + adiabatic flash (the default in nozzle.xls)
2) calculated with adiabatic flash
3) calculated with isentropic flash

Calculated Area
1) rigorous numerical solution of isentropic nozzle
2) API formulation for gas and vapors, ideal cp/cv
3) API formulation for gas and vapors, real cp/cv
4) API formulation for gas and vapors, real cp/cv, Zv = 1

all the required properties are calculated by Prode library , for personal/academic use you can download a  free -with limited number of components- copy from www.prode.com

I utilize the unit conversion methods available in Prode Properties to convert from and to different units
UMCR() converts the different values entred by user to the units required by formulation (Pa.a,K,Kg/s)
UMCS() converts the calculated area (cm2) to the units required by the user
StrMw(1) returns the Molecular weight
the formulations are those of API and ISO standards, with different parameters as discuused,
you can easily introduce variants if you wish to compare different methods

the relative errors are calculated as difference with rigorous solution


good luck
Paolo

Carib1 ,
To support your query , attached documents .

Hope this helps

Breizh

I'm going to lock this old thread.  If you'd like, we can put this in the Download section for all.

Dear All,

There has been a very healthy debate on the subject matter and I have spent considerable time and effort researching the subject. Today on the auspicious day of Diwali I am presenting my latest efforts on the subject matter by providing a spreadsheet which is very accurate in not only predicting the water content of sweet natural gas but also for sour natural gas. My Diwali gift to all the esteemed readers and friends of this fantastic forum. Comments are welcome. In the earlier posts on the same topic I have mentioned about this research of mine and wish to thank the authors of the article from which the spreadsheet is inspired. The title of the paper and the link for the article is also mentioned in the spreadsheet.

Comments are welcome.

Regards,
Ankur.

Dear All,
Just find the useful excel sheets to Hysys Users.

Dear all
 
Good day!
 
Inlet design is one of the most commonly neglected aspects of a knockout drum/ separator design, thus  the cause of poor performance.
Improperly selected inlet device may result in :

• Gas jets to the back wall of the vessel.
• Without enough space to diffuse the jet, gas utilizes only part of the mist eliminator. Due to uneven velocity profile liquid carryover occurs.
• The gas jet agitates the accumulated liquid below, generating droplets.
• Turbulence spoils normal gravity settling of larger liquid droplets below the mist eliminator. Additional liquid load increases the likelihood of flooding the mist eliminator.

When a fluid stream strikes a solid surface, it reduces its momentum by exerting a force on that surface, and an equal but opposite force is exerted on itself. If this force is too great, liquid shatter can occur when the surface tension forces of the liquid are overcome. In order to avoid liquid shatter, different types of inlet devices have been developed which can reduce the momentum. The sizing of nozzles is normally judged by the maximum  ρv² value, which depends on the type of inlet device that is used or to be installed.
 
Commonly used values of ρv² throughout the industry for some of these types of inlet devices are as follows:

• No inlet device or simple baffle: ρv² ≤ 1,000 kg/ms²
• Half-open pipe inlet device: ρv² ≤ 1,500 kg/ms²
• V-baffle: ρv² ≤ 2,500 kg/ms²
• Vane-type inlet device: ρv² ≤ 15,000 kg/ms²

Half pipe or better to use perforated pipe
 
 

 
V-baffle  ( rarely used now )
 
 

Vane inlet device
 
Additionally, the best document, I have seen for sizing , selection is Shell DEP 31.22.05.11-Gen " GAS/LIQUID SEPARATORS - TYPE SELECTION AND DESIGN RULES".
 
Few year back I have seen an old version of the document  on an iranian official  site " fumblog.um.ac.ir" . I am not aware whether it is still available or not . However, you may find a similar document on " http://www.namdaran....Separators).pdf "
 
At the end, please do realize that you may suggest a device but the selection would finally be governed by the manufacturer of the device
 
Best regards and God Bless you
 

Books are written that address this question, so it's not easy to summarize in a Post. The daily tasks are highly dependent on the the engineer's role, his skill level, the skills of his colleagues, the industry, the environment (e.g., operating plant, engineering office, etc.), and many other factors. That said, consider these possible descriptions:

A process engineer is responsible for design, procurement, installation, operation, and maintenance of chemical process equipment and systems. He/she prepares flow diagrams, P&IDs, and equipment specifications. He/she sizes equipment, makes material balances, selects technology, chooses materials of construction, prepares equipment data sheets and specifications, and might be considered the overall 'architect' of the chemical plant. From the controls viewpoint, the process engineer determines what must be controlled (e.g., temperature, pressure, flow, level), and often determines the measuring points and final control elements. He/she might develop the initial control loops (on the P&ID).

A process controls engineer is responsible for the determining the controls details. Before distributed control, this function was called 'instrument engineer'. On a micro level, the controls engineer designs the control loops, decides whether advanced control strategies such as cascades are required, and selects the sensors, transmitters, controllers, and final control devices. He/she is responsible for sizing control valves and orifices, determining tuning parameters, and matching control signals. On a macro level, he/she writes Functional Design Specifications and Detailed Design Specifications, and (depending on the person's exact function) might do the actual configuration and programming of the controllers. These days, the controllers are likely to be in a distributed system which may consist of localized PLCs, a centralized control system, or a combination system. The controls engineer helps determine acceptable process ranges, alarm limits, and calibration tolerances.

In summary, adding 'controls' to the job title results in a completely different set of responsibilities. Both jobs are typically filled with a chemical engineer. But the skill sets quickly diverge.

My advice is - do not start looking at specifics until a detailed cycle time analysis of the entire batch process is complete.  Account for every minute of time with a reason, such as pumping raw material A, heating to reaction temperature, reacting according to batch instructions, extra reaction time to obtain quality, waiting on laboratory results, no activity (operator at lunch), pumping out product, rinsing reactor, etc., etc., etc.

Then, decide what can be changed to shorten the cycle time without affecting product quality.  Can two seperate items be done at the same time?  This sometimes gives good cycle time reductions at no cost!  Can transfer rates of the most time consuming steps be reduced with a bigger pump/piping?  Will more heat transfer area speed up some steps?  Will more "delta T" speed up some heat transfer steps?  Will process automation eliminate periods of no activity controlled by operators?  Is "turnaround time" (the time from the end of batch #1 to the beginning of batch n# 2) excessive?  There should be a zero standard for "turnaround time".

I've been in batch processes for 33 years.  The detailed analysis mentioned above is the starting point that drives all downstream investigations, so it must be done and documentation must be painstakingly thorough!  Or, you will get lost somewhere along the way.  Focus on those items that give the biggest bang (largest cycle time reduction) for the buck (lowest cost).

I'm sorry to come to the party so late. I have answered similar questions to this one many times, and being too lazy to write it all out again I thought that it was time for me to consolidate some of my previous answers into a coherent document. And then other deadlines cropped up and suddenly two weeks had flashed by.

Anyway, I have now posted the document on my web page if you still want more information. At least I will be able to answer quickly next time this topic comes up ;-)

Pressure Drop in Pipe Fittings and Valves - A Discussion of the Equivalent Length (Le/D), Resistance Coefficient (K) and Valve Flow Coefficient (Cv) Methods
http://www.katmarsof...essure-drop.htm

take a look at this paper , it should help .

Breizh

Shivshankar,

Norsok P-001 guidelines are for gas piping and not for gas pipelines. These again are way over the top. Engineering companies like Bechtel allow a maximum dry gas velcoity with no solid particles as 100 ft/s (30 m/s) for piping.  BP allows a velocity of 38 m/s for gas or vapour piping including super-heated steam.

I would not design a long-distance natural gas transmission pipe line for a velocity of more than 10 m/s normally unless the client says they can go to a higher velocity and provide this in writing. A recent job for a 600 km sales gas transmission pipeline of 52" inch size had velocities ranging from 9 to 10 m/s through the entire pipeline route.

Regards,
Ankur.

I have many, many spreadsheets from visitors that have tons of great information in them, but aren't necessarily ready to be opened and used as they exist.  Some may need just a little work, some may need a lot of work.  Many of them are well referenced in terms of methodology being used.

However, some people may be working on similar projects and may find these helpful.  My hope would be to release them in the Downloads section (clearly marked as "Development Items") and to see if the community can work together to turn these into polished projects.

I thought I'd start by collecting initial feedback on the idea.  Let me know what you think.  These titles would only be visible to registered members of the community.

Thanks,

munishankar,

What is the unit for oil and grease in your specification? I am assuming it is 8,767 ppm by volume. How come your IGF unit is producing effluent water with so much oil & grease and TSS?

I have serious doubts that you can release waste water with so much oil and grease to disposal wells. The H2S content of 0.1  mg/l also is high.

Normally transport of waste water with oil and grease is done with low shear centrifugal pumps in many installations. These pumps run at low speeds of 750 rpm. The idea of using low shear centrifugal pumps is to prevent shearing and thus formation of water-oil emulsion which is detrimental when further treatment and disposal is required. Low speed postive displacement pump such as progressive cavity pumps which prevent shearing are also acceptable. What finally becomes important in the selection is the CAPEX and OPEX
of the pump. You will have to do this analysis to arrive at the most economic option for your application.

Hope this helps.

Regards,
Ankur.

An additional note.
1. Standard conditions can be a bit different, as pavanayi points out.  Let us accept 60 oF and 1 Atm as standard condition for gases (widely used in USA in the past, versus Nm3 in Europe).
Air out is 200 SCFM * 5 min = 1000 SCF = 1000*.3048^3 std m3= 28.317 std m3.
Ideal gas molar volume is 22.414*(288.71/273.15) = 23.690 std m3/kgmol (22.414 m3 is the molar volume at 0 oC, 1Atm, probably remembered from high school).
Thus air out is 28,317/23.690 = 1.1953 kgmol
2. Initial air in vessel satisfies the relation ni=Pi*V/(RT)
    Final state: nf=Pf*V/(RT)
hence ni-nf = (Pi - Pf)*V/(RT) = 1.1953 kgmol
We can find R, knowing molar volume at 0 oC and 1 Atm: R=22.414*1/1/273.15 = 0.08206 m3*Atm/kgmol/oK.
Thus V=1.1953*0.08206*310.93/((9-4)/1.01325) m3 = 6.18 m3
Note: 9 barg - 4 barg =10.013 bara - 5.013 bara = 5 bar
3. The calculation can be easily done in a spreadsheet, only molar volume at some condition has to be remembered and that 1 ft = 0.3048 m (that is 12 in).

I've noticed that many of you are wanting to share Microsoft Excel files. You can upload them into your messages just like other files. This should simplify the process of sharing the spreadsheets rather than having to email them to one another.

awesome:

Well, for me, pressure loss = NON-RECOVERABLE pressure drop.

For example, frictionnal pressure drop is a pressure loss, as the pressure energy is irreversibly transformed to thermal energy (heat).

Pressure drop due to the difference of static pressure (different elevations) between two points is recoverable (gravitational potential energy), so it is not a loss.

In a nutshell, total pressure drop = recoverable pressure drop (usually a difference of static pressure) + pressure loss (usually caused by friction during fluid flow).
Depending on your system, total pressure drop may be equal to one or the other term only.

thorium90,

 

In the past year and a half I might have interviewed about 20-25 candidates eligible as process engineers and yes it is true that all these candidates had a minimum experience of 4-5 years, because that was the eligibility criteria by my company for selecting a candidate as a process engineer. Higher experience was required for the position of senior process engineer and even higher for a principal / lead process engineer.

Today's cut throat business environment requires that when you invest in human resources, immediate returns are expected which is only possible if you have trained and experienced human resources available to you. Another criteria that discourages many companies from hiring fresh graduates is the fact the loyalty factor has disappeared on part of the younger generation. As soon as freshers get trained and acquire some skills, they start looking for greener pastures. Why would a company invest in training a fresher only to lose him or her as soon as he or she acquires the requisite skills, just because the person has suddenly realized that he or she can take advantage of his or her acquired skills in obtaining a new and more lucrative job.

This was not the case in the past. Young trainees used to join a company and spend a minimum of 8-10 years before moving on. In countries like Japan, an employment was for a lifetime. You joined a company as a fresher in your early twenties and retired from there as an old man. All that is history now.

I don't want to paint a gloomy picture to you but the competition has increased ten-fold since I graduated in the mid-eighties. You want a decent job then you have to qualify on two fronts. The first one would be to be amongst the best in your chosen field and the second one would be to present yourself as the best i.e sell yourself. Thus just being skillful is not enough, you need to be able to present your skills for the world to see. It is an old adage that gloss sells.

Last but not the least, never lose hope and faith. There is a path destined for everybody. We have been created as homo sapiens by the creator and he gave the human race the intelligence and will to overcome any challenges.

Well I hope I haven't overdone the preacher bit as you would find in your local parish.

Regards,

Ankur.